Hydrocarbon Processing: Petrochemical Processes Handbook 2005

Hydrocarbon Processings Petrochemical Processes 2005 handbooks refl ect the dynamic advancements now available in licensed process technologies catalysts and equipment The petrochemical industry continues to apply energyconserving environmen tally friendly costeffective solutions to produce products that improve the quality of every day life The global petrochemical industry is innovativeputting knowledge into action to create new products that service the needs of current and future markets HPs Petrochemical Processes 2005 handbooks are inclusive catalogs of established and emerging licensed technologies that can be applied to existing and grassroots facili ties Economic stresses drive efforts to conserve energy minimize waste improve product qualities and most important increase yields and create new products A full spectrum of licensed petrochemical technologies is featured These include manu facturing processes for olefi ns aromatics polymers acidssalts aldehydes ketones nitrogen compounds chlorides and cyclocompounds Over 30 licensing companies have submitted process fl ow diagrams and informative process descriptions that include eco nomic data operating conditions number of commercial installations and more To maintain as complete a listing as possible the Petrochemical Processes 2005 hand book is available on CDROM and at our website to certain subscribers Additional copies of the Petrochemical Processes 2005 handbook may be ordered from our website Premier Sponsor Gulf Publishing Company PROGRAM LICENSE AGREEMENT YOU SHOULD READ THE TERMS AND CONDITIONS CAREFULLY BEFORE USING THIS APPLICATION INSTALLING THE PROGRAM INDICATES YOUR ACCEPTANCE OF THESE TERMS AND CONDITIONS CLICK HERE TO READ THE TERMS AND CONDITIONS Acetic acid Acrylonitrile Alkylbenzene Alpha olefins 2 Ammonia 7 Aniline Aromatics Aromatics extraction Aromatics extractive distillation 3 Aromatics recovery Benzene 2 Bisphenol A BTX aromatics 4 Butadiene extraction Butadiene 13 2 Butanediol 14 2 Butene1 Butyraldehyde n and i Cumene 3 Cyclohexane Dimethyl ether DME Dimethyl terephthalate Dimethylformamide EDC 2 Ethanolamines Ethers EthersMTBE Ethyl acetate Ethylbenzene 3 Ethylene 7 Ethylene feed Ethylene glycol 3 Ethylene oxide 3 Ethylene oxideEthylene glycols Formaldehyde 2 Hydrogen Maleic anhydride Methanol 7 Methylamines Mixed xylenes 5 mXylene Octenes Olefins 5 Paraffin normal 2 Paraxylene 6 Paraxylene crystallization Phenol 3 Phthalic anhydride Polyalkylene terephthalates Polycaproamide Polyesters Polyethylene 8 Polypropylene 7 Polystyrene 4 Propylene 7 PVC suspension 2 Pyrolysis gasoline Styrene 3 Styrene acrylonitrile Terephthalic acid Upgrading steam cracker C3 cuts Upgrading steam cracker C4 cuts Urea 2 Ureaformaldehyde VCM by thermal cracking of EDC VCM removal Wet air oxidation Xylene isomerization 3 Processes index Premier Sponsor ABB Lummus Global Aker Kvaerner Axens Axens NA Badger Licensing LLC Basell Polyolefins BASF AG BP BP Chemicals CDTECH Chemical Research Licensing Chisso Corp Chiyoda Corp Davy Process Technology ExxonMobil Chemical ExxonMobil Chemical Technology Licensing LLC GE Plastics GTC Technology Haldor Topsøe AS Hydro Illa International Japan Polypropylene Corp Johnson Matthey Catalysts Johnson Matthey PLC Kellogg Brown Root Inc Linde AG Lonza Group Lurgi AG Lyondell Chemical Co Mitsubishi Chemical Corp Mitsui Chemicals Inc Nippon Petrochemicals Co Ltd Niro Process Technology BV NOVA Chemicals International SA Novolen Technology Holdings CV One Synergy SABIC Scientific Design Company Inc Shell International Chemicals BV Sinopec Research Institute of Petroleum Processing Stamicarbon BV Stone Webster Inc SudChemie Inc Sunoco Technip The Dow Chemical Co Toyo Engineering Corp Uhde GmbH Uhde InventaFischer Union Carbide Corp Univation Technologies UOP LLC Vinnolit Company index Premier Sponsor Processes Axens Axens is a refining petrochemical and natural gas market focused supplier of process tech nology catalysts adsorbents and services backed by nearly 50 years of commercial success Axens is a world leader in several areas such as Petroleum hydrotreating hydroconversion FCC gasoline desulfurization Catalytic Reforming BTX benzene toluene xylenes production purification Selective Hydrogenation of olefin cuts Sulfur recovery catalysts Axens is a fullyowned subsidiary of IFP Alpha olefins Aromatics recovery Benzene BTX aromatics Butene1 Cyclohexane Ethylene feed Mixed xylenes Octenes Paraxylene Paraxylene Propylene Pyrolysis gasoline Upgrading steam cracker C3 cuts Upgrading steam cracker C4 cuts Xylene isomerization PROCESSING PetrochemicalProcesses tet St MOTE ALAA N72 Ue M01 exot0 27 a home processes index company index Acetic acid Application To produce acetic acid using the process ACETICA Metha nol and carbon monoxide CO are reacted with the carbonylation reac tion using a heterogeneous Rh catalyst Methanol feed oT Description Fresh methanol is split into two streams and is contacted steam with reactor offgas in the highpressure absorber 7 and light gases in the lowpressure absorber 8 The methanol exiting the absorbers are recombined and mixed with the recycle liquid from the recycle Process renee rant surge drum 6 This stream is charged to a unique bubblecolumn cooler 2 TO reactor 1 Sp Carbon monoxide is compressed and sparged into the reactor riser BFW The reactor has no mechanical moving parts and is free from leakage CO feed C6 Aj Flue gas maintenance problems The ACETICA Catalyst is an immobilized Rh Fuel complex catalyst on solid support which offers higher activity and op Makeup CHs erates under less water conditions in the system due to heterogeneous system and therefore the system has much less corrosivity Reactor effluent liquid is withdrawn and flashvaporized in the Flash er 2 The vaporized crude acetic acid is sent to the dehydration column 3 to remove water and any light gases Dried acetic acid is routed to Commercial plant One unit is under construction for a Chinese client the finishing column 4 where heavy byproducts are removed in the bottom draw off The finished aceticacid product is treated to remove Reference Acetic Acid Process Catalyzed by lonically Immobilized Rho trace iodide components at the iodide removal unit 5 dium Complex to Solid Resin Support Journal of Chemical Engineering Vapor streams from the dehydration column overhead contacted of Japan Vol 37 4 pp 536545 2004 with methanol in the lowpressure absorber 8 Unconverted CO meth The ChiyodaUOP ACETICA process for the production of acetic ane other light byproducts exiting in the vapor outlets of the highand 2d 8th Annual SaudiJapanese Symposium on Catalysts in Petroleum lowpressure absorbers and heavy byproducts from the finishing column Refining and Petrochemicals KFUPMRI Dhahran Saudi Arabia Nov are sent to the incinerator with scrubber 9 2930 1998 a tethenol mune mptons 0539 Licensor Chiyoda Corp CcOomtmt 0517 Power CO Supply 0 KG kWhmt 129 Water cooling m3mt 137 Steam 100 psig mtmt 17 PROCESSING PetrochemicalProcesses miele IN ce meena JCESSES home processes index company index Acrylonitrile Application A process to produce highpurity acrylonitrile and highpu rity hydrogen cyanide from propylene ammonia and air Recovery of Reaction Quench Recovery Purification byproduct acetonitrile is optional HEN product ar a Spent air Acrylonitrile Description Propylene ammonia and air are fed to a fluidized bed re product actor to produce acrylonitrile ACRN using DuPonts proprietary catalyst HP een system Other useful products from the reaction are hydrogen cyanide steam H2S04 HCN and acetonitrile ACE The reaction is highly exothermic and heat is recovered from the reactor by producing highpressure steam The reactor effluent is quenched and neutralized with a sulfuric solution to remove the excess ammonia The product gas from the quench is absorbed with water to Propylene air recover the ACRN HCN and ACE The aqueous solution of ACRN Ammonia INH S04 ee HCN and ACE is then fractionated and purified into highquality roe products The products recovery and purification is a highly efficient and lowenergy consumption process This ACRN technology minimizes the amount of aqueous effluent a major consideration for all acrylonitrile producers This ACRN technology is based on a highactivity highthroughput catalyst The propylene conversion is 99 with a selectivity of 85 Commercial plants DuPont Chemical Solution Enterprise Beaumont to useful products of ACRN HCN and ACE The DuPont catalyst is a Texas 200000 mtpy mechanically superior catalyst resulting in a low catalyst loss DuPont has developed a Catalyst Bed Management Program CBMP to Licensor Kellogg Brown Root Inc maintain the properties of the catalyst bed inside the reactor at optimal performance throughout the operation The catalyst properties the CBMP and proprietary reactor internals provide an optimal performance of the ACRN reactor resulting in high yields With over 30 years of operating experience DuPont has developed knowhow to increase the onstream factor of the plant This know how includes the effective use of inhibitors to reduce the formation of cyanide and nitrile polymers and effective application of an antifoulant system to increase onstream time for equipment PROCESSING PetrochemicalProcesses home processes index company index Alkylbenzene linear Application The Detal process uses a solid heterogeneous catalyst to produce linear alkylbenzene LAB by alkylating benzene with linear Bhigoe bose olefins made by the Pacol process H me ar recycle LE Benzene Description Linear paraffins are fed to a Pacol reactor 1 to dehydro recycle LAB genate the feed into corresponding linear olefins Reactor effluent is separated into gas and liquid phases in a separator 2 Diolefins in the separator liquid are selectively converted to monoolefins in a DeFine C2 7 reactor 3 Light ends are removed in a stripper 4 and the resulting olefinparaffin mixture is sent to a Detal reactor 5 where the olefins are alkylated with benzene The reactor effluent is sent to a fractionation Heavy section 6 7 for separation and recycle of unreacted benzene to the Linear alkylate Detal reactor and separation and recycle of unreacted paraffins to the paraffin Pacol reactor A rerun column 8 separates the LAB product from the charge Paraffin recycle heavy alkylate bottoms stream Feedstock is typically C1 to C3 normal paraffins of 98 purity LAB product has a typical Bromine Index of less than 10 Yields Based on 100 weight parts of LAB 81 parts of linear paraffins and 34 parts of benzene are charged to a UOP LAB plant Economics Investment US Gulf Coast inside battery limits for the pro duction of 80000 tpy of LAB 1000tpy Commercial plants Twentynine UOP LAB complexes based on the Pa col process have been built Four of these plants use the Detal process Reference Greer D et al Advances in the Manufacture of Linear Alkylbenzene 6th World Surfactants Conference CESIO Berlin Ger many June 2004 Licensor UOP LLC a COC Cee ey a PROCESSING PetrochemicalProcesses home processes index company index Alpha olefins linear Application To produce highpurity alpha olefins C4C9 suitable as copolymers for LLDPE production and as precursors for plasticizer alco aentin hols and polyalphaolefins using the AlphaSelect process and storage Butene1 Description Polymergrade ethylene is oligomerized in the liquidphase reactor 1 with a catalystsolvent system designed for high activity and pny ene Hexene1 selectivity Liquid effluent and spent catalyst are then separated 2 the Octene1 liquid is distilled 3 for recycling unreacted ethylene to the reactor then fractionated 4 into highpurity alphaolefins Spent catalyst is treated to Decene1 remove volatile hydrocarbons and recovered The table below illustrates the superior purities attainable wt with the AlphaSelect process Cya nButene1 99 Solvent nHexene1 98 recycle Catalyst Heavy ends with nOctene1 96 removal spent catalyst nDecene1 92 The process is simple it operates at mild operating temperatures and pressures and only carbon steel equipment is required The catalyst is nontoxic and easily handled Fuel gas 003 Yields Yields are adjustable to meet market requirements and very little Heavy ends 002 high boiling polymer is produced as illustrated Utilities cost USton product 51 Alphaolefin product distribution wt Catalyst chemicals USton product 32 nButene1 3343 nHexene1 3032 Commercial plants The AlphaSelect process is strongly backed by exten nOctene1 1721 P nDecene1 914 sive Axens industrial experience in homogeneous catalysis in particular the Alphabutol process for producing butene1 for which there are 19 Economics Typical case for a 2004 ISBL investment at a Gulf Coast loca units producing 312000 tpy tion producing 65000 tpy of CyC19 alphaolefins is P J a Py 10 2P Licensor Axens Axens NA Investment million US 37 Raw material Ethylene tons per ton of product 115 Byproducts tonton of main products readies here iin ed for EUR es Cea a C2 olefins 01 iste se cal PetrochemicalProcesses miele IN ce meena JCESSES home processes index company index Alpha olefins Application The aSablin process produces aolefins such as butene1 hexane1 octene1 decene1 etc from ethylene in a homogenous Butene1 catalytic reaction The process is based on a highly active bifunctional Ethylene Hexenet catalyst system operating at mild reaction conditions with highest selec tivities to aolefins 2 4 Octene1 Description Ethylene is compressed 6 and introduced to a bubblecol umn type reactor 1 in which a homogenous catalyst system is intro Decene1 duced together with a solvent The gaseous products leaving the reactor eon overhead are cooled in a cooler 2 and cooled in a gasliquid separator solvent for reflux 3 and further cooled 4 and separated in a second gasliquid separator 5 d Unreacted ethylene from the separator 5 is recycled via a com Cot pressor 6 and a heat exchanger 7 together with ethylene makeup to the reactor A liquid stream is withdrawn from the reactor 1 con taining liquid aolefins and catalyst which is removed by the catalyst removal unit 8 The liquid stream from the catalyst removal unit 8 is combined with the liquid stream from the primary separation 5 These combined liquid streams are routed to a separation section in which via a series of columns 9 the aolefins are separated into the Commercial plants One plant of 150000 metric tpy capacity is currently individual components under construction for Jubail United in AlJubail Saudi Arabia By varying the catalyst components ratio the product mixture can oe be adjusted from light products butene1 hexene1 octene1 decene Licensor The technology is jointly licensed by Linde AG and SABIC 1 to heavier products C12 to C9 aolefins Typical yield for light olefins is over 85 wt with high purities that allow typical product applications The light products show excellent properties as comonomers in ethylene polymerization Economics Due to the mild reaction conditions pressure and tempera ture the process is lower in investment than competitive processes Typical utility requirements for a 160000metric tpy plant are 3700 tph cooling water 39 MW fuel gas and 6800 kW electric power i PROCESSING PetrochemicalProcesses miele IN ce a home processes index company index Ammonia Application To produce ammonia from a variety of hydrocarbon feed Desulfurization Reforming Shift stocks ranging from natural gas to heavy naphtha using Topsges low Process steam FA energy ammonia technology Process air DX i 1A Description Natural gas or another hydrocarbon feedstock is compressed eS eS 1 if required desulfurized mixed with steam and then converted into Qe synthesis gas The reforming section comprises a prereformer optional a orm oc but gives particular benefits when the feedstock is higher hydrocarbons I ee a i or naphtha a fired tubular reformer and a secondary reformer where Purge gas st 550 optional O Lo Lo ack process air is added The amount of air is adjusted to obtain an HN i U 3 ratio of 30 as required by the ammonia synthesis reaction The tubular eal Ls aA al i i co steam reformer is Topses proprietary sidewallfired design After the 1 ee reforming section the synthesis gas undergoes high and lowtempera Methanation ture shift conversion carbon dioxide removal and methanation Ammonia Se ee CO Synthesis gas is compressed to the synthesis pressure typically prom ranging from 140 to 220 kgcmg and converted into ammonia in a synthesis loop using radial flow synthesis converters either the two bed S200 the threebed S300 or the S250 concept using an S200 converter followed by a boiler or steam superheater and a onebed S50 converter Ammonia product is condensed and separated by structed within the last decade range from 650 mtpd up to 2050 mtpd refrigeration This process layout is flexible and each ammonia plant will being the worlds largest ammonia plant Design of new plants with be optimized for the local conditions by adjustment of various process even higher capacities are available arameters Topsge supplies all catalysts used in the catalytic process Steps for ammonia production Licensor Haldor Topsze AS Features such as the inclusion of a prereformer installation of a ringtype burner with nozzles for the secondary reformer and upgrading to an S300 ammonia converter are all features that can be applied for existing ammonia plants These features will ease maintenance and improve plant efficiency Commercial plants More than 60 plants use the Topse process con cept Since 1990 50 of the new ammonia production capacity has heen based on the Topsge technology Capacities of the plants con PROCESSING PetrochemicalProcesses miele IN ce a home processes index company index Ammonia KAAPplus Application To produce ammonia from hydrocarbon feedstocks using a Excess air i rL To process highpressure heat exchangebased steam reforming process integrated ate ue oreesor yr condensate A with a lowpressure advanced ammonia synthesis process F Ie 2 stipper THEE P24 5 steam Description The key steps in the KAAPpus process are reforming us aa 3 u a ing the KBR reforming exchanger system KRES cryogenic purification Process steam Process ATR RES of the synthesis gas and lowpressure ammonia synthesis using KAAP Pie synthesis oe To BFW system catalyst fs 7 compressor pe tome Following sulfur removal 1 the feed is mixed with steam heated a L OX ower ammonia and split into two streams One stream flows to the autothermal reformer feo Stripper fea Waste gas product ATR 2 and the other to the tube side of the reforming exchanger 3 XM 7 pre Seer NOR which operates in parallel with the ATR Both convert the hydrocarbon CI CI By 10 ae feed into raw synthesis gas using conventional nickel catalyst In the ATR co abenrbet ot a 0 exchanger feed is partially combusted with excess air to supply the heat needed to vecdltetach UC Recier CD reform the remaining hydrocarbon feed The hot autothermal reformer effluent is fed to the shell side of the KRES reforming exchanger where it combines with the reformed gas exiting the catalystpacked tubes The combined stream flows across the shell side of the reforming exchanger where it supplies heat to the reforming reaction inside the tubes Shellside effluent from the reforming exchanger is cooled in a waste bar with a small catalyst volume Effluent vapors are cooled by ammonia heat boiler where highpressure steam is generated and then flows to refrigeration 13 and unreacted gases are recycled Anhydrous liquid the CO shift converters containing two catalyst types one 4 is a high ammonia is condensed and separated 14 from the effluent temperature catalyst and the other 5 is a lowtemperature catalyst Energy consumption of KBRs KAAPplus process is less than 25 MM Shift reactor effluent is cooled condensed water separated 6 and then Btu LHVshortton Elimination of the primary reformer combined with routed to the gas purification section CO is removed from synthesis lowpressure synthesis provides a capital cost savings of about 10 over gas using a wet CO scrubbing system such as hot potassium carbonate conventional processes or Mee CO moval final purification includes methanation 8 gas Commercial plants Over 200 largescale singletrain ammonia plants drying 9 and cryogenic purification 10 The resulting pure synthesis of KBR design are onstream or neve been contracted worlwide The gas is compressed in a singlecase compressor and mixed with a recycle KAAPplus advanced ammonia technology provides a lowcost lowen stream 11 The gas mixture is fed to the KAAP ammonia converter 12 which uses a rutheniumbased highactivity ammonia synthesis catalyst It provides high conversion at the relatively low pressure of 90 ergy design for ammonia plants minimizes environmental impact re duces maintenance and operating requirements and provides enhanced reliability Two plants use KRES technology and 17 plants use Purifier technology Four 1850mtpd grassroots KAAP plants in Trinidad are in full operation Licensor Kellogg Brown Root Inc Ammonia KAAPplus continued PROCESSING PetrochemicalProcess eee C f CS home processes index company index Ammonia KBR Purifier Application To produce ammonia from hydrocarbon feedstocks and air Air Description The key features of the KBR Purifier Process are mild pri 3 C 5 6 mary reforming secondary reforming with excess air cryogenic purifica Steam il 2 os os tion of syngas and synthesis of ammonia over magnetite catalyst in a horizontal converter Feed WY Aye 1 Desulfurized feed is reacted with steam in the primary reformer 1 with exit temperature at about 700C Primary reformer effluent is re To ug acted with excess air in the secondary reformer 2 with exit at about S OL 900C The air compressor is normally a gasdriven turbine 3 Turbine AW MAW iE 2 S exhaust is fed to the primary reformer and used as preheated combus A Wy i tion air An alternative to the above described conventional reforming is a 15 to use KBRs reforming exchanger system KRES as described in KBRs 11 Or KAAPplus swede Secondary reformer exit gas is cooled by generating highpressure steam 4 The shift reaction is carried out in two catalytic stepshigh temperature 5 and lowtemperature shift 6 Carbon dioxide removal 7 uses licensed processes Following CO removal residual carbon oxides are converted to methane in the methanator 8 Methanator effluent is cooled and water is separated 9 before the raw gas is dried 10 is recycled back to the syngas compressor A small purge is scrubbed Dried synthesis gas flows to the cryogenic purifier 11 where it is with water 15 and recycled to the dryers cooled by feedeffluent heat exchange and fed to a rectifier The syngas is purified in the rectifier column producing a column overhead that is Commercial plants Over 200 singletrain plants of KBR design have essentially a 7525 ratio of hydrogen and nitrogen The column bottoms been contracted worldwide Seventeen of these plants use the KBR Pu is a waste gas that contains unconverted methane from the reforming rifier process section excess nitrogen and argon Both overhead and bottoms are re Licensor Kellogg Brown Root Inc heated in the feedeffluent exchanger The waste gas stream is used to regenerate the dryers and then is burned as fuel in the primary reformer A small lowspeed expander provides the net refrigeration The purified syngas is compressed in the syngas compressor 12 mixed with the loopcycle stream and fed to the converter 13 Convert er effluent is cooled and then chilled by ammonia refrigeration Ammo nia product is separated 14 from unreacted syngas Unreacted syngas PROCESSING PetrochemicalP eee C Oe Sich home processes index company index Ammonia Application The Linde ammonia concept LAC produces ammonia from light hydrocarbons The process is a simplified route to ammonia con Fuel sisting of a modern hydrogen plant standard nitrogen unit and a high fae d efficiency ammonia synthesis loop 9 col olf 3 7 8 4 Description Hydrocarbon feed is preheated and desulfurized 1 Pro AY ott Pel bic f cess steam generated from process condensate in the isothermal shift Feed felt os 6 reactor 5 is added to give a steam ratio of about 27 reformer feed is BFW further preheated 2 Reformer 3 operates with an exit temperature 1 of 850C O Reformed gas is cooled to the shift inlet temperature of 250C by oS nT f Ee generating steam 4 The CO shift reaction is carried out in a single stage in the isothermal shift reactor 5 internally cooled by a spiral b 13 1 wound tube bundle To generate MP steam in the reactor deaerated Air an and reheated process condensate is recycled After further heat recovery final cooling and condensate separation 6 the gas is sent to the pressure swing adsorption PSA unit 7 Loaded adsorbers are regenerated isothermally using a controlled sequence of depressurization and purging steps Nitrogen is produced by the lowtemperature air separation in a to 33C 16 for storage Utility units in the LAC plant are the power cold box 10 Air is filtered compressed and purified before being generation system 17 which provides power for the plant from HP supplied to the cold box Pure nitrogen product is further compressed superheated steam BFW purification unit 18 and the refrigeration and mixed with the hydrogen to give a pure ammonia synthesis gas unit 19 The synthesis gas is compressed to ammoniasynthesis pressure by the Oe a syngas compressor 11 which also recycles unconverted gas through Economics Simplification over conventional processes gives important the ammonia loop Pure syngas eliminates the loop purge and associated Savings such as investment catalystreplacement costs maintenance purge gas treatment system costs etc Total feed requirement process feed plus fuel is approxi The ammonia loop is based on the Ammonia Casale axialradial mately 7 Gcalmetric ton mt ammonia 252 MMBtushort ton de threebed converter with internal heat exchangers 13 giving a high Pending on plant design and location conversion Heat from the ammonia synthesis reaction is used to generate HP steam 14 preheat feed gas 12 and the gas is then cooled and refrigerated to separate ammonia product 15 Unconverted gas h is recycled to the syngas compressor 11 and ammonia product chilled Commercial plants The first complete LAC plant for 1350mtd am monia has been built for GSFC in India Two other LAC plants for 230 and 600mtd ammonia were commissioned in Australia The latest LAC contract is under erection in China and produces hydrogen ammonia and CO2 under import of nitrogen from already existing facilities There are extensive reference lists for Linde hydrogen and nitrogen plants and Ammonia Casale synthesis systems References A Combination of Proven Technologies Nitrogen March April 1994 Licensor Linde AG Ammonia continued PROCESSING PetrochemicalProcesses home processes index company index Ammonia Application To produce ammonia from natural gas LNG LPG or naph ruel CO HP steam from synthesis tha Other hydrocarbonscoal oil residues or methanol purge gas drum are possible feedstocks with an adapted frontend The process uses 5 Ey Ger E 5 Cc Methanation conventional steam reforming synthesis gas generation frontend and rr s tt AN GG a mediumpressure MP ammonia synthesis loop It is optimized with steam LE S onde MM HTshift Co respect to low energy consumption and maximum reliability The larg a i crea CS BFW A est singletrain plant built by Uhde with a conventional synthesis has MPstam oer S LK yy p Combustion air Focess gas removal a nameplate capacity of 2000 metric tons per day mtpd For higher HP steam air Make up eas AW capacities refer to Unde Dual Pressure Process ee qj U Convection bank coils Description The feedstock natural gas as an example is desulfurized Armonia ree een ees mixed with steam and converted into synthesis gas over nickel catalyst oy Neue aol T4f 3 Process air preheater or CF 4 Feed preheater at approximately 40 bar and 800C to 850C in the primary reformer rem a way 5 Combustion air The Uhde steam reformer is a topfired reformer with tubes made of Syngas compressor TH ot liquid Preheater centrifugal high alloy steel and a proprietary cold outlet manifold sys tem which enhances reliability In the secondary reformer process air is admitted to the syngas via a special nozzle system arranged at the circumference of the secondary reformer head that provides a perfect mixture of air and gas Subsequent synthesis loop and allows maximum ammonia conversion rates highpressure HP steam generation and superheating guarantee maximum a 2 Liquid ammonia is separated by condensation from the synthesis process heat usage to achieve an optimized energy efficient process Lo loop and is either subcooled and routed to storage or conveyed at mod CO is converted to CO in the HT and LT shift over standard cata ar erate temperature to subsequent consumers lysts CO is removed in a scrubbing unit which is normally either the Ammonia flash and purge gases are treated in a scrubbing system and BASFaMDEA or the UOPBenfield process Remaining carbonoxides a hydrogen recovery unit not shown and the remains are used as fuel are reconverted to methane in the catalytic methanation to trace ppm levels Commercial plants Seventeen ammonia plants have been commis The ammonia synthesis loop uses two ammonia converters with sioned between 1990 and 2004 with capacities ranging from 600 mtpd three catalyst beds Waste heat is used for steam generation down yp to 2000 mtpd stream the second and third bed Wasteheat steam generators with integrated boiler feedwater preheater are supplied with a special cooled Licensor Unde GmbH tubesheet to minimize skin temperatures and material stresses The con verters themselves have radial catalyst beds with standard small grain click here to email for more information a iron catalyst The radial flow concept minimizes pressure drop in the PROCESSING PetrochemicalProcesses miele IN ce a f home processes index company index Ammonia PURIFIERpus Tear HTS Application To produce ammonia from hydrocarbon feedstocks using a ar wR 4 recrey eam highpressure HP heat exchangebased steam reforming process inte Feed aan condensate grated with cryogenic purification of syngas ZS 7 i ae t Description The key steps in the PURIFIERpus process are reforming orocess steam Lt ae ete SY Us using the KBR reforming exchanger system KRES with excess air cryo Process heater reformer KRES To BFW genic purification of the synthesis gas and synthesis of ammonia over Methanato aH magnetite catalyst in a horizontal converter Following sulfur removal 1 the feed is mixed with steam heated C0 7A a compressor and split into two streams One stream flows to the autothermal refor ae Waste gs Oy u mer ATR 2 and the other to the tube side of the reforming exchanger recovery Ammonia 3 which operates in parallel with the ATR Both convert the hydrocar 1m Unitized chiller oe bon feed into raw synthesis gas using conventional nickel catalyst In an oae sa I Ex the ATR feed is partially combusted with excess air to supply the heat Expander nv 5 mo 16 needed to reform the remaining hydrocarbon feed The hot autother Feethane ie mal reformer effluent is fed to the shell side of the KRES reforming ex m changer where it combines with the reformed gas exiting the catalyst packed tubes The combined stream flows across the shell side of the reforming exchanger where it supplies heat to the reforming reaction inside the tubes column producing a column overhead that is essentially a 7525 ratio Shellside effluent from the reforming exchanger is cooled in a Of hydrogen and nitrogen The column bottoms is a waste gas that con wasteheat boiler where HP steam is generated and then flows to the tains unconverted methane from the reforming section excess nitrogen CO shift converters containing two catalyst types one 4 is a high and argon Both overhead and bottoms are reheated in the feedeff temperature catalyst and the other 5 is a lowtemperature catalyst uent exchanger The waste gas stream is used to regenerate the dryers Shift reactor effluent is cooled condensed water separated 6 and and then is burned as fuel in the primary reformer A small lowspeed then routed to the gas purification section CO is removed from syn expander provides the net refrigeration thesis gas using a wetCO scrubbing system such as hot potassium The purified syngas is compressed in the syngas compressor 12 carbonate or MDEA methyl diethanolamine 7 mixed with the loopcycle stream and fed to the horizontal converter Following CO removal residual carbon oxides are converted tome 13 Converter effluent is cooled and then chilled by ammonia refri thane in the methanator 8 Methanator effluent is cooled and water is geration in a unitized chiller 14 Ammonia product is separated 15 separated 9 before the raw gas is dried 10 Dried synthesis gas flows to the cryogenic purifier 1 1 where it is cooled by feedeffluent heat exchange and fed to a rectifier The syngas is purified in the rectifier from unreacted syngas Unreacted syngas is recycled back to the syngas compressor A small purge is scrubbed with water 16 and recycled to the dryers Commercial plants Over 200 largescale singletrain ammonia plants of KBR design are onstream or have been contracted worldwide The PURIFIERplus ammonia technology provides a lowcost lowenergy design for ammonia plants minimizes environmental impact reduces operating requirements and provides enhanced reliability Two plants use KRES technology and 17 plants use PURIFIER technology Licensor Kellogg Brown Root Inc Ammonia PURIFIERplus continued PROCESSING PetrochemicalProcesses home processes index company index AmmoniaDual pressure process fuel eee Application Production of ammonia from natural gas LNG LPG or mn naphtha The process uses conventional steam reforming synthesis gas NY ae ti R generation in the frontend while the synthesis section comprises a ZN oncethrough section followed by a synthesis loop It is thus optimized HP steam BFW with respect to enable ammonia plants to produce very large capacities a ni MP steam a with proven equipment The first plant based on this process will be Feed NK the SAFCO IV ammonia plant in AlJubail Saudi Arabia which is cur Process air Ammonia aay rr rently under construction This concept provides the basis for even larger Combustion air synthesis loop CO plants 40005000 mtpd Purge i F Cc 7 co Description The feedstock eg natural gas is desulfurized mixed with cs om steam and converted into synthesis gas over nickel catalyst at approxi NH liquid DL ae mately 42 bar and 800850C in the primary reformer The Uhde steam LL BFW reformer is a topfired reformer with tubes made of centrifugal micro MT Makeup gas alloy steel and a proprietary cold outlet manifold which enhances reliability In the secondary reformer process air is admitted to the syngas via ne DL once through a special nozzle system arranged at the circumference of the secondary reformer head that provides a perfect mixture of air and gas HP steam Subsequent highpressure HP steam generation and superheating guarantee maximum process heat recovery to achieve an optimized en tl i ergy efficient process CO conversion is achieved in the HT and LT shift over standard cata lyst while CO is removed either in the BASFaMDEA the UOPBenfield or the UOPAmine Guard process Remaining carbonoxides are recon In the second step the remaining syngas is compressed to the op verted to methane in catalytic methanation to trace ppm levels erating pressure of the ammonia synthesis loop approx 210 bar in The ammonia synthesis loop consists of two stages Makeup gas is op the HP casing of the syngas compressor This HP casing operates at a compressed in a twostage intercooled compressor which is the low much lower than usual temperature The high synthesis loop pressure pressure casing of the syngas compressor Discharge pressure of this a oo is achieved by combination of the chilled second casing of the syngas compressor is about 110 bar An indirectly cooled oncethrough con verter at this location produces one third of the total ammonia Effluent wom ans converters cooree and the major part of the ammonia pro click here to email for more information a uced is separated from the gas compressor and a slightly elevated frontend pressure Liquid ammonia is separated by condensation from the once through section and the synthesis loop and is either subcooled and routed to storage or conveyed at moderate temperature to subse quent consumers Ammonia flash and purge gases are treated in a scrubbing system and a hydrogen recovery unit not shown the remaining gases being used as fuel Economics Typical consumption figures feed fuel range from 67 to 72 Gcal per metric ton of ammonia and will depend on the individual plant concept as well as local conditions Commercial plants The first plant based on this process will be the SAF CO IV ammonia plant with 3300 mtpd currently under construction in AlJubail Saudi Arabia Licensor Uhde GmbH AmmoniaDual pressure process continued PROCESSING PetrochemicalProc eee OTe ATLANTIC Ud Mele cists iets home processes index company index Aniline Application A process for the production of highquality aniline from Vent benzene and nitric acid Benzene Reaction anne Description Aniline is produced by the nitration of benzene with nitric Nitric acid acid to mononitrobenzene MNB which is subsequently hydrogenated LH to aniline In the DuPontKBR process benzene is nitrated with mixed Sulfuric Ji conden acid nitric and sulfuric at high efficiency to produce mononitrobenzene ace EE MNB in the unique dehydrating nitration DHN system The DHN sys tem uses an inert gas to remove the water of nitration from the reaction Tars mixture thus eliminating the energyintensive and highcost sulfuric Wash NB acid concentration system C er emt oa As the inert gas passes through the system it becomes humidified water ne removing the water of reaction from the reaction mixture Most of the Hydrogen P energy required for the gas humidification comes from the heat of Dehydrating striming hydrogenation nitration The wet gas is condensed and the inert gas is recycled to the nitrator The condensed organic phase is recycled to the nitrator while the aqueous phase is sent to effluent treatment The reaction mixture is phase separated and the sulfuric acid is returned to the nitrator The crude MNB is washed to remove residual acid and the impurities formed during the nitration reaction The product is then distilled 4 plant located in Beaumont Texas In addition DuPonts aniline technol and residual benzene is recovered and recycled Purified MNB is fed gy is used in three commercial units and one new license was awarded together with hydrogen into a liquid phase plugflow hydrogenation in 2004 with a total aniline capacity of 300000 tpy reactor that contains a DuPont proprietary catalyst The supported noble metal catalyst has a high selectivity and the MNB conversion per pass is Licensor Kellogg Brown Root Inc 100 The reaction conditions are optimized to achieve essentially quantitative yields and the reactor effluent is MNBfree The reactor product is sent to a dehydration column to remove the water of reaction followed by a purification column to produce highquality aniline product Commercial plants DuPont produces aniline using this technology for the merchant market with a total production capacity of 160000 py at PROCESSING PetrochemicalProcesses PROCESSING JLGOIOU home processes index company index Aromatics Application The technology produces benzene and xylenes from tolu H Toluene ene and Cg streams This technology features a proprietary zeolite cata lyst and can accommodate varying ratios of feedstock while maintain Reactor ing high activity and selectivity Light aromatics Catalyst Description The technology encompasses three main processing ar Heavy aromatics Cot Nomreacted eas splitter reactor and stabilizer sections The heavyaromatics stream sol hydrogen to Cots feed is fed to the splitter The overhead Cg aromatic product is ue recycle sone the feed to the transalkylation reactor section The splitter bottoms Separator is exchanged with other streams for heat recovery before leaving the eave syste mM Reactor Stabilizer The aromatic product is mixed with toluene and hydrogen vapor ever ized and fed to the reactor The reactor gaseous product is primarily unreacted hydrogen which is recycled to the reactor The liquid prod Heavy aromatics uct stream is subsequently stabilized to remove further light aromatic components The resulting aromatics are routed to product fraction ation to produce the final benzene and xylenes products The reactor is charged with zeolite catalyst which exhibits both long life and good flexibility to feed stream variations including substantial C19 aromatics Depending on feed compositions and light components Significant decrease in energy consumption due to efficient heat present the xylene yield can vary from 25 to 32 and Cy conversion Integration scheme oO oO from 53 to 67 Commercial plants Two commercial plants are using GTTransAlk tech Process advantages include nology and catalyst othe evel cost fixedbed reactor design drop in replacement for Licensor GTC Technology using catalyst manufactured by SudChemie e Very high selectivity benzene purity is 999 without extraction Inc e Physically stable catalyst with long cycle life e Flexible to handle up to 90 C t components in feed with high conversion Catalyst is resistant to impurities common to this service Operating parameters are moderate a Decreased hydrogen consumption due to low cracking rates Aromatics extraction Application The Sulfolane process recovers highpurity C6 C9 aromat ics from hydrocarbon mixtures such as reformed petroleum naphtha reformate pyrolysis gasoline pygas or coke oven light oil COLO by extractive distillation with or without liquidliquid extraction Description Fresh feed enters the extractor 1 and fl ows upward coun tercurrent to a stream of lean solvent As the feed fl ows through the extractor aromatics are selectively dissolved in the solvent A raffi nate stream very low in aromatics content is withdrawn from the top of the extractor The rich solvent loaded with aromatics exits the bottom of the extractor and enters the stripper 2 The lighter nonaromatics tak en overhead are recycled to the extractor to displace higher molecular weight nonaromatics from the solvent The bottoms stream from the stripper substantially free of nonaro matic impurities is sent to the recovery column 3 where the aromatic product is separated from the solvent Because of the large difference in boiling point between the solvent and the heaviest aromatic compo nent this separation is accomplished easily with minimal energy input Lean solvent from the bottom of the recovery column is returned to the extractor The extract is recovered overhead and sent on to dis tillation columns downstream for recovery of the individual benzene toluene and xylene products The raffi nate stream exits the top of the extractor and is directed to the raffi nate wash column 4 In the wash column the raffi nate is contacted with water to remove dissolved sol vent The solventrich water is vaporized in the water stripper 5 and then used as stripping steam in the recovery column The raffi nate product exits the top of the raffi nate wash column The raffi nate prod uct is commonly used for gasoline blending or ethylene production The solvent used in the Sulfolane process was developed by Shell Oil Co in the early 1960s and is still the most effi cient solvent available for recovery of aromatics Economics The purity and recovery performance of an aromatics extrac tion unit is largely a function of energy consumption In general higher solvent circulation rates result in better performance but at the expense of higher energy consumption The Sulfolane process demonstrates the lowest solventtofeed ratio and the lowest energy consumption of any commercial aromatics extraction technology A typical Sulfolane unit consumes 275 300 kcal of energy per kilogram of extract produced even when operating at 9999 wt benzene purity and 9995 wt recovery Estimated inside battery limits ISBL costs based on unit processing 158000 mtpy of BT reformate feedstock with 68 LV aromatics US Gulf Coast site in 2003 Investment US million 85 Utilities per mt of feed Electricity kWh 62 Steam mt 048 Watercooling m3 135 Commercial plants In 1962 Shell commercialized the Sulfolane process in its refi neries in England and Italy The success of the Sulfolane pro cess led to an agreement in 1965 whereby UOP became the exclusive licensor of the Sulfolane process Many of the process improvements incorporated in modern Sulfolane units are based on design features and operating techniques developed by UOP UOP has licensed a total of 134 Sulfolane units throughout the world Licensor UOP LLC PROCESSING PetrochemicalProcesses miele IN ce a Mele erste home processes index company index Aromatics extractive distillation Application The Distapex process uses extractive distillation for recov ering individual aromatics from a heart cut containing the desired aro anaractive eee oar istillation column Raffinate matic compound column Description The feedstock ie the heart cut with the aromatic compo nent to be recovered is routed to the middle of the extractive distillation purearomatic column 1 The solvent NMP is supplied at the top of the column In component the presence of the solvent the aromatic component and the nonaro Aromatics matics are separated in the column The aromatic component passes together with the solvent to the bottom and is routed to the stripper 3 It is separated from the solvent under vacuum The overhead aromatic component leaves the plant as solvent aromatic pure product and the solvent is circulated to the extractive distillation column 1 Balen High heat utilization is obtained by intensive heat exchange of the circulating solvent Necessary additional heat is supplied by medium pressure steam at 1214 bar The nonaromatics still containing small quantities of solvent are obtained at the top of the extractive distillation column 1 This solvent is recovered in the raffinate column 2 and returned to the solvent re Installations The Distapex process is applied in more than 25 reference cycle plants Benzene recovery from pyrolysis gasoline is usually above 995 at feed concentration above 80 Reference G Krekel G Birke A Glasmacher et al Developments in Aromatics Separation Erdé Erdgas Kohle May 2000 Economics A typical investment for a Distapex plant to recover 200000 tpy benzene is approximately 85 million Licensor Lurgi AG Typical energy consumption figures of the Distapex plant calculated per ton of benzene produced are Steam 1214 bar ton 06 Electric power kWh 4 Water cooling m3 24 Solvent loss kg 001 iste se cal PetrochemicalProcesses miele IN ce meena JCESSES home processes index company index Aromatics extractive distillation Application The Sulfolane process recovers highpurity aromatics from hydrocarbon mixtures by extractive distillation ED with liquidliquid ex traction or with extractive distillation ED Typically if just benzene or colon Recovery toluene is the desired product then ED without liquidliquid extraction is the more suitable option D D Description Extractive distillation is used to separate closeboiling com Raffinate ponents using a solvent that alters the volatility between the compo to storage Sorat toe nents An ED Sulfolane unit consists of two primary columns they are the ED column and the solvent recovery column Aromatic feed is pre heated with lean solvent and enters a central stage of the ED column Fresh feed 1 The lean solvent is introduced near the top of the ED column Non Steam aromatics are separated from the top of this column and sent to storage generar The ED column bottoms contain solvent and highly purified aromatics that are sent to the solvent recovery column 2 In this column aromat ics are separated from solvent under vacuum with steam stripping The overhead aromatics product is sent to the BT fractionation section Lean solvent is separated from the bottom of the column and recirculated back to the ED column oo Commercial plants In 1962 Shell commercialized the Sulfolane process Economics The solvent used in the Sulfolane process exhibits higher jn its refineries in England and Italy The success of the Sulfolane pro selectivity and capacity for aromatics than any other commercial sol Gace Jed to an agreement in 1965 whereby UOP became the exclusive vent Using the Sulfalane process minimizes concern about trace nitro jicansor of the Sulfolane process Many of the process improvements gen contamination that occurs with nitrogenbased solvents Estimated incorporated in modern Sulfolane units are based on design features inside battery limits ISBL costs based on a unit processing 158000 sn4q operating techniques developed by UOP UOP has licensed a total mtpy of BT reformate feedstock with 68 LV aromatics US Gulf Coast 134 sylfolane units throughout the world site in 2003 Investment US million 68 Licensor UOP LLC Utilities per mt of feed Electricity kWh 27 Steam mt 035 Water cooling m 25 hietesciit cal PetrochemicalProcesses z home processes index company index Aromatics extractive distillation Application Recovery of highpurity aromatics from reformate pyrolysis extractive Nonaromatics gasoline or cokeoven light oil using extractive distillation distillation column Description In Uhdes proprietary extractive distillation ED Morphylane SZ L process a singlecompound solvent NFormylmorpholine NFM alters the TI C a vapor pressure of the components being separated The vapor pressure of Aromatics A the aromatics is lowered more than that of the less soluble nonaromatics fraction Nonaromatics vapors leave the top of the ED column with some solv ent which is recovered in a small column that can either be mounted on a Aromatics the main column or installed separately S L ee Bottom product of the ED column is fed to the stripper to separate 1 pure aromatics from the solvent After intensive heat exchange the lean QD QO solvent is recycled to the ED column NFM perfectly satisfies the neces sary solvent properties needed for this process including high selectivity BE NenE ee entbaroiiatics thermal stability and a suitable boiling point Economics Pygas feedstock Production vield Benzene Benzenetoluene Commercial plants More than 55 Morphylane plants total capacity Benzene y 9995 9995 of more than 6 MMtpy vvality 9998 References Emmrich G F Ennenbach and U Ranke Krupp Uhde Quality 30 wt pom NA 80 wt ppm NA Processes for Aromatics Recovery European Petrochemical Technology Toluene e 600 wt ppm NA Conference June 2122 1999 London Consumption Emmrich G U Ranke and H Gehrke Working with an extractive dis Steam 475 kgt ED feed 680 kgt ED feed tillation process Petroleum Technology Quarterly Summer 2001 p 125 k with lowaromati ntent 20 wt Reformate feedstoc ow aro atics content 20 wt Licensor Uhde GmbH Quality Benzene 10 wt ppm NA Consumption Steam 320 kgt ED feed ORCC Lm Clu E LCL a Maximum content of nonaromatics Including benzenetoluene splitter PROCESSING PetrochemicalProcesses PROCESSING LAVA ETeTe Le home processes index company index Aromatics recovery Application Recovery via extraction of high purity CgCg aromatics Extractor Water wash Stripper Recovery from pyrolysis gasoline reformate coke oven light oil and kerosene frac tower Raffinate tions lanes Extract recycle Description Hydrocarbon feed is pumped to the liquidliquid extraction column 1 where the aromatics are dissolved selectively in the sulfolane waterbased solvent and separated from the insoluble nonaromatics Feed paraffins olefins and naphthenes The nonaromatic raffinate phase exits at the top of the column and is sent to the wash tower 2 The wash tower recovers dissolved and entrained sulfolane by water extrac Aromatics vo To water tion and the raffinate is sent to storage Water containing sulfolane is stripper sent to the water stripper Rich solvent Water The solvent phase leaving the extractor contains aromatics and small amounts of non aromatics The latter are removed in the stripper 3 Lean solvent and recycled to the extraction column The aromaticenriched solvent is pumped from the stripper to the recovery tower 4 where the aromat ics are vacuum distilled from the solvent and sent to downstream clay treatment and distillation Meanwhile the solvent is returned to the extractor and the process repeats itself Commercial plants Over 20 licensed units are in operation Yields Overall aromatics recoveries are 99 while solvent losses are extremely small less than 0006 bbb of feed Licensor Axens Axens NA Economics For 2005 US Gulf Coast location CC pyrolysis CgCo Feed gasoline reformate Feed bpsd 8000 15000 Aromatics wt 6488 6072 Utilities per bbl feed Cooling 10 Btu 014016 01012 Steam MP Ib 180210 188225 Power kWh 0608 11 ISBL Investment 10 US 1518 1720 eta cal PetrochemicalProcesses home processes index company index Benzene Application To produce highpurity benzene and heavier aromatics from toluene and heavier aromatics using the Detol process Description Feed and hydrogen are heated and passed over the catalyst rLts 1 Benzene and unconverted toluene andor xylene and heavier aro rete eee Fuel gas matics are condensed 2 and stabilized 3 To meet acid wash color specifications stabilizer bottoms are passed through a fixedbed clay treater then distilled 4 to produce the desired Benzene specification benzene The cryogenic purification of recycle hydrogen to C7 Aromatic cAI Xylenes reduce the makeup hydrogen requirement is optional 6 2 Unconverted toluene andor xylenes and heavier aromatics are recycled oo Recycle toluene and C aromatics Yields Aromatic yield is 990 mol of fresh toluene or heavier aromatic charge Typical yields for production of benzene and xylenes are Type production Benzene Xylene feed wt Nonaromatics 32 23 Benzene 113 Toluene 473 07 aromas 495 ana Commercial plants Twelve plants with capacities ranging from about 12 g aromatics a aye Products wt of feed million to 100 million galy have been licensed Benzene 757 369 C aromatics ot 377 Licensor ABB Lummus Global 545C minimum freeze point 1000 ppm nonaromatics maximum Economics Basis of ISBL US Gulf Coast Estimated investment bpsd 6700 Typical utility requirements per bb feed Electricity kWh 58 Fuel MMBtu 031 Water cooling gal 450 Steam Ib 144 No credi ee Cu mC Cr et j No credit taken for vent gas streams PROCESSING PetrochemicalProcesses miele IN ce mere ate aaiter LCL eksyS ets home processes index company index Benzene Application Produce benzene via the hydrodealkylation of C7C aromatics Hydrogen makeup Light ends Description Fresh C7Cg to C feed is mixed with recycle hydrogen 0 makeup hydrogen and C7 aromatics from the recycle tower The mix 3 Benzene ture is heated by exchange 1 with reactor effluent and by a furnace 2 that also generates highpressure steam for better heat recovery 7 Tr Tight temperature control is maintained in the reactor 3 to arrive La My 12 7 13 at high yields using a multipoint hydrogen quench 4 In this way con version is controlled at the optimum level which depends on reactor 6 f throughput operating conditions and feed composition 1 Purge By recycling the diphenyl 5 its total production is minimized to the advantage of increased benzene production The reactor effluent is eeu CY G Go recycle cooled by exchange with feed followed by cooling water or air 6 and sent to the flash drum 7 where hydrogenrich gas separates from the condensed liquid The gas phase is compressed 8 and returned to the reactor as quench recycle Hy Part of the stream is washed countercurrently with a feed sidestream in the vent H absorber 9 for benzene recovery The absorber overhead flows to the hydrogen purification unit 10 where hydrogen purity is Economics Basis US Gulf Coast 2005 increased to 90t so it can be recycled to the reactor The stabilizer 11 Toluene feed metric tpy 120700 removes light ends mostly methane and ethane from the flash drum Benzene product metric tpy 100000 liquid The bottoms are sent to the benzene column 12 where high Ween ooling 650 purity benzene is produced overhead The bottoms stream containing Flow inhr 208 unreacted toluene and heavier aromatics is pumped to the recycle col Temperature differential C 111 umn 13 Toluene Cg aromatics and diphenyl are distilled overhead and Fuel heat release million kcalhr 83 recycled to the reactor A small purge stream prevents the heavy compo 420 barg steam production kghr 3859 nents from building up in the process ISBL investment 10 USD 4045 Yields Benzene yields are close to the theoretical owing to several tech Commercial plants Thirtyfive plants have been licensed worldwide for niques used such as proprietary reactor design heavy aromatic diphe nyl recycle and multipoint hydrogen quench i processing a variety of feedstocks including toluene mixed aromatics reformate and pyrolysis gasoline Licensor Axens Axens NA Benzene continued tistesstttcial PetrochemicalProcesses miele IN ce a f home processes index company index Bisphenol A Recycle Phenol Application The Badger BPA technology is used to produce highpu acetone 3 rity bishenol A BPA product suitable for polycarbonate and epoxy resin applications The product is produced over ionexchange resin from phenol and acetone in a process featuring proprietary purifica a a os o tion technology Description Acetone and excess phenol are reacted by condensation in acet Solvent an ion exchange resincatalyzed reactor system 1 to produce pp BPA oe Water Adduct Molten BPA BPA prills water and various byproducts The crude distillation column 2 removes water and unreacted acetone from the reactor effluent Acetone and lights are adsorbed into phenol in the lights adsorber 3 to produce a recycle acetone stream The bottoms of the crude column is sent to the Purge Mother liquor crystallization feed preconcentrator 4 which distills phenol and con centrates BPA to a level suitable for crystallization Wastewater BPA is separated from byproducts in a proprietary solvent crystal pada lization and recovery system 5 to produce the adduct of pp BPA and phenol Mother liquor from the purification system is distilled in the solvent recovery column 6 to recover dissolved solvent The solvent free mother liquor stream is recycled to the reaction system A purge Product quality Typical values for BPA quality are from the mother liquor is sent to the purge recovery system 7 along Freezing point C 157 with the recovered process water to recover phenol The recovered BPA wiw wt 9995 purified adduct is processed in a BPA finishing system 8 to remove Methanol color APHA 5 phenol from product and the resulting molten BPA is solidified in the prill tower 9 to produce product prills suitable for the merchant BPA ommercial plants The first plant among the largest in the world began market operation in 1992 at the Deer Park Houston plant now owned and oper ated by Resolution Performance Products LLC Since that time two other Process features The unique crystallization system produces a stable worldscale plants were licensed to the AsiaPacific market crystal that is efficiently separated from its mother liquor These plants are extremely reliable and have been engineered to meet the operating Licensor Badger Licensing LLC standards of the most demanding refining and chemical companies The catalyst system uses a unique upflow design that is beneficial to catalyst re ane performance High capacity operation has been fully demon i iste se cal PetrochemicalProcesses miele IN ce meena MO etsyslete eae home processes index company index BTX aromatics Application To produce high yields of benzene toluene xylenes and Renaormand Regenerator hydrogen from naphthas via the CCR Aromizing process coupled with a heaters 7 1Y RegenC continuous catalyst regeneration technology Benzene and tolu yr pi A ene cuts are fed directly to an aromatics extraction unit The xylenes H i Booster fraction obtained by fractionation and subsequent treatment by the an C NX a compressor Arofining process for diolefins and olefins removal is ideal for para aN Hydrogen xylene and orthoxylene production rich gas Y Y al Description This process features moving bed reactors and a continu Y Y YY ous catalyst regeneration system coupled with a hard smoothflowing U AN OT 9 DA catalyst Feed enters the reactor 1 passes radially through the moving pW L system catalyst bed exits at the reactor bottom and proceeds in the same man ner through the 23 remaining reactors 2 The robust latest genera eee CE GQ Morar to tion AR 501 505 catalyst moves downward through each reactor Recycle stabization Leaving the reactor the catalyst is gaslifted to the next reactors feed ee hopper where it is distributed for entry After the last reactor an inert gas lift system isolates and transports the catalyst to the recentlyin troduced RegenC regeneration section 3 Coke is removed catalyst is returned to its original state and sent to the first reactor the cycle begins again A recovery system 4 separates hydrogen for use in downstream Economics The ISBL investment for a typical 25000bpsd CCR Aromiz units and the Aromizate is sent to a stabilization section The unit is fully ing unit with a RegenC regenerator 2004 Gulf Coast location automated and operating controls are integrated into a DCS requiring Investment including initial catalyst inventory only a minimum of supervisory and maintenance effort US million 53 j 0 Typical utility requirements Mields Feed Products Fuel 106 kealh 76 TBP cut C 80150 Hydrogen 41 Steam HP th net export 17 Paraffins 57 Cot 87 Electricity kWhh 5900 Naphthenes 37 Benzene 85 Catalyst operating cost ton feed 05 Aromatics 6 Toluene 263 Exclusive of noble metals Xylenes 261 Total aromatics 743 Commercial plants Sixtyfour CCR reforming units have been licensed including seven plants in operation and four under design Licensor Axens Axens NA BTX aromatics continued iste se cal PetrochemicalProcesses miele IN ce meena JCESSES home processes index company index BTX aromatics Application An aromatics process based on extractive distillation GT BTX efficiently recovers benzene toluene and xylenes from refinery or Lean a petrochemical aromatics streams such as catalytic reformate or pyrolysis solvent Raffinate gasoline Hydrocarbon feed Extractive Description Hydrocarbon feed is preheated with hot circulating sol distillation ae vent and fed at a midpoint into the extractive distillation column a downstream EDC Lean solvent is fed at an upper point to selectively extract solvent fractionation the aromatics into the column bottoms in a vaporliquid distillation G recovery Water operation Nonaromatic hydrocarbons exit the column top and pass column through a condenser A portion of the overhead stream is returned a Steam to the column top as reflux to wash out any entrained solvent The solvent balance of the overhead stream Is the raffinate product requiring no tf further treatment Rich solvent from the bottom of the EDC is routed to the solvent recovery column SRC where the aromatics are stripped overhead Stripping steam from a closedloop water circuit facilitates hydrocarbon stripping The SRC operates under vacuum to reduce the boiling point at the column base Lean solvent from the bottom of the SRC is passed through heat Economics exchange before returning to the EDC A small portion of the lean New unit Expansion of conventional circulating solvent is processed in a solventregeneration step to remove BTX recovery unit heavy decomposition products which are purged daily Capital cost MM 3900 Ht reformate 4000 incremental The process advantages over conventional liquidliquid extraction Simple pretax payout yr 22 12 processes include lower capital and operating costs and simplicity of ROI 44 85 operation Advantages over other extractive processes include superior a solvent system fewer equipment pieces small equipment and expanded Commercial plants Fourteen commercial licenses are in place eecstoc range Des ent alows se QTSHoOS aoMatKS Reference Benzene reducon in motor gesoineobgaton or op configurations portunity Hydrocarbon Processing Process Optimization Confer ence April 1997 Improve BTX processing economics Hydrocarbon Processing March 1998 Licensor GTC Technology BTX aromatics continued PROCESSING PetrochemicalProcesses home processes index company index BTX aromatics Application To produce reformate which is concentrated in benzene toluene and xylenes BTX from naphtha and condensate feedstocks via a highseverity reforming operation with a hydrogen byproduct The CCR Platforming Process is licensed by UOP Net H rich gas Description The process consists of a reactor section continuous cata Waphtha feed rieleee lyst regeneration section CCR and product recovery section Stacked catalyat rom treating radial flow reactors 1 facilitate catalyst transfer to and from the CCR fe 6 C Spent catalyst regeneration section 2 A charge heater and interheaters 3 catalyst fl x are used to achieve optimum conversion and selectivity for the endo i thermic reaction y Light ends Reactor effluent is separated into liquid and vapor products 4 Fo Liquid product is sent to a stabilizer 5 to remove light ends Vapor O from the separator is compressed and sent to a gasrecovery section C aromatics 6 to separate 90pure hydrogen byproduct A fuel gas byproduct of LPG can also be produced UOPs latest R270 series catalyst maximizes aromatics yields Yields Typical yields from lean Middle East naphtha Ha wt 43 units in design and construction Total operating capacity represents Benzene wt 17 over 39 million bpd Toluene wt 299 pa Xylenes wt 304 Ast wt 131 Licensor UOP LLC Economics Capital investment per mtpy of feed US 5075 Utilities per metric ton feedrate Electricity kWh 12 Steam HP mt 016 Water cooling m 20 Fuel MMkcal 013 Commercial plants There are 173 units In operation and 30 additional click ia ia ee LL for atlas information a tistesstttcial PetrochemicalProcesses S home processes index company index BTX aromatics Application To produce petrochemicalgrade benzene toluene and xy lenes BTX via the aromatization of propane and butanes using the BP Stripper offgas UOP Cyclar process 4 Description The process consists of a reactor section continuous cata C Aromatic lyst regeneration CCR section and productrecovery section Stacked product radialflow reactors 1 facilitate catalyst transfer to and from the CCR Net fuel gas catalyst regeneration section 2 A charge heater and interheaters 3 Hydrogen achieve optimum conversion and selectivity for the endothermic reac Lom Ol tion Reactor effluent is separated into liquid and vapor products 4 reactor The liquid product is sent to a stripper column 5 to remove light satu Comp rates from the C aromatic product Vapor from the separator is com pressed and sent to a gas recovery unit 6 The compressed vapor is then separated into a 95 pure hydrogen coproduct a fuelgas stream Besniied peor containing light byproducts and a recycled stream of unconverted LPG Yields Total aromatics yields as a wt of fresh feed range from 61 for propane to 66 for mixed butanes feed Hydrogen yield is approxi mately 7wt fresh feed Typical product distribution is 27 benzene 43 toluene 22 Cg aromatics and 8 Cg aromatics process 1000 bpd of C3 or Cy feedstock at either high or lowpressure Economics US Gulf Coast inside battery limits basis assuming gas tur VE 2 wide range of operating conditions A second unit capable of bine driver is used for product compressor processing C3 and C feedstock was commissioned in 2000 and oper ates at design capacities Investment US per metric ton mt of feed 175208 Typical utility requirements unit per mt of feed Reference Doolan PC and P R Pujado Make aromatics from LPG Electricity kWh 0013 Hydrocarbon Processing September 1989 pp 7276 Steam MP mt credit 07 Gosling C D et al Process LPG to BTX products Hydrocarbon Steam LP mt 013 Processing December 1991 Water cooling mt 19 Fuel MMkcal 2 Licensor UOP LLC Boiler feedwater mt 055 Commercial plants In 1990 the first Cyclar unit was commissioned at the BP refinery at Grangemouth Scotland This unit was designed to a Ce Oe a hieteseii cal PetrochemicalProcesses home processes index company index Butadiene extraction Application To produce a polymergrade butadiene product from Butanebutylene product mixedC streams by extractive distillation using acetonitrile ACN as 2 caivent naked Butane P the solvent ak ed 6 Wash water to toa washer SANE Description This butadiene extraction process was originally developed recove Ce Solvent bleed y by Shell Chemicals It is offered under license agreement by Kellogg C 1 system Brown Root who has updated and optimized the process to reduce feed Extractive Recovered solvent capital and operating costs distillation be Light ends The process scheme consists of contacting mixedC feed with lean ee solvent in the extractive distillation column 1 The raffinate butenes and section 13 Butadiene butanes which are not absorbed flow overhead to the wash column 5 teay 2 for solvent recovery The butadienerich solvent flows to the stripper 31 stripping Heavyends ends system 3 where the butadiene is separated from the solvent Raw section ee butadiene is purified to meet specifications in the purification section Lean solvent Tean solvent 4 Heavy ends Cy acetylenes are also separated in the stripper system 3 as a side product and further processed in the heavyends stripping section 5 The solvent recovery step 6 maintains solvent quality and recovers solvent from various product streams Use of acetonitrile is advantageous to other solvent systems for a Component Value Units number of reasons ACNs lower boiling point results in lower operating 13 Butadiene 995 7 wt minimum otal acetylenes 20 ppm wt maximum temperatures resulting in low fouling rates and long runlengths Only Methyl acetylene 10 opm wt maximum lowpressure steam is required for reboilers The low molecular weight Vinyl acetylenes 10 ppm wt maximum and low molar volume of ACN combined with its high selectivity to Propadiene 10 ppm wt maximum butadiene result in low solvent circulation rates and smaller equipment u2 Butadiene 10 ppm wt maximum sizes The low viscosity of ACN increases tower efficiencies and Cs hydrocarbons 200 ppm wt maximum reduces column size and cost ACN is also very stable noncorrosive Gommercial plants Over 35 butadiene units have been constructed us and biodegradable The basic process is noncorrosive and requires only ing the Shell ACN technology Unit capacities range from 20 Mtpy to carbon steel materials of construction over 225 Mtpy Yields This process can exceed 98 recovery of the butadiene con Licensor Kellogg Brown Root Inc tained in the feed as product This product will meet all butadiene de rivative requirements with typical specifications shown below i rg eg cg mere a RECS Rag aaa FPA i PROCESSING PetrochemicalPro eee C Oe Sich home processes index company index 13 Butadiene Extraction from mixed C Application To produce highpurity butadiene from a mixed C stream a 1 3Butadiene typically a byproduct stream from an ethylene plant using liquid feeds product liquids cracker The BASFLummus process uses nmethylpyrrolidone ata ae Lean NMP as the solvent NMP Lean solvent Description The mixed C feed stream is fed into the first extractive NMP distillation column 1 which produces an overhead butenes stream raf solvent CG Reavies finate1 that is essentially free of butadiene and acetylenes The bottoms stream from this column is stripped free of butenes in Mixed the top half of the rectifier 2 A side stream containing butadiene and C feed a small amount of acetylenic compounds vinyl and ethylacetylene is po C acetylenes withdrawn from the rectifier and fed into the second extractive distillation aera column 3 The Cy acetylenes which have higher solubilities in NMP than 13 butadiene are removed by the solvent in the bottoms and returned Lean NMP solvent to heat recovery to the rectifier A crude butadiene BD stream from the overhead of the second extractive distillation column is fed into the BD purification train Both extractive distillation columns have a number of trays above the solvent addition point to allow for the removal of solvent traces from the overheads fed to a water scrubber to remove a small amount of NMP from the The bottoms of the rectifier containing BD C4 acetylenes and exiting gases The scrubbed gases containing the Cy acetylenes are Cs hydrocarbons in NMP is preheated and fed into the degasser the purged to disposal solvent stripping column 4 In this column solvent vapors are used as In the propyne column 5 the propyne C3 acetylene is removed the stripping medium to remove all light hydrocarbons from NMP as overhead and sent to disposal The bottoms are fed to the second The hotstripped solvent from the bottom of the degasser passes distillation column the 13butadiene column 6 which produces pure through the heat economizers a train of heat exchangers and is fed to BD as overhead and a small stream containing 12butadiene and Cs the extractive distillation columns hydrocarbons as bottoms The hydrocarbon column dot pons Feaving eet IMP ddl ater and Ped to the Yield Typically more than 98 of the 13butadiene contained in the bottom of the rectifier via a recycle gas compressor feed is recovered as product Hydrocarbons having higher solubilities in the solvent than 13 butadiene accumulate in the middle zone of the degasser and are drawn off as a side stream This side stream after dilution with raffinate1 is Economics Unit based on a 100000 metric tpy ISBL US Gulf Coast Investment US million 30 Utilities per ton BD Steam ton 2 Water cooling m3 150 Electricity kWh 150 Commercial plants Currently 27 plants are in operation using the BASF butadiene extraction process Five additional projects are in the design or construction phase Licensor BASFAGABB Lummus Global 13 Butadiene Extraction from mixed C4 continued tistesstttcial PetrochemicalProcesses home processes index company index Butadiene 13 Application The KLP process selectively hydrogenates acetylenes in crude butadiene streams from steam crackers to their corresponding diene or olefin to recover 13butadiene The KLP process can be used in Hy Raffinate1 1 3Butadiene new installations to eliminate the costly secondstage extractive distilla tion step or as a retrofit to increase product quality or throughput el wg oo feed Description In the KLP process the Cy stream is mixed with an essen 7 tially stoichiometric amount of hydrogen and fed to two fixedbed reac tors in series containing KLP60 catalyst The reaction pressure is high enough to maintain the reaction mixture in the liquid phase The KLP reactor effluent then flows to a distillation column to remove hydrogen Heavies Solvent 0 eb and and a small amount of heavies formed in the process The KLP effluent fp stream is processed in a singlestage extractive distillation unit to sepa KLP Onestage BD extraction rate and recover highpurity 13butadiene Yields The combination of the KLP process with butadiene extraction can provide over 100 recovery of the butadiene contained in the feed as product The recovery is enhanced by the conversion of vinylacetylene to 13butadiene Total acetylene levels in the product of less than 10 wtppm are achievable The process also offers improved safety in op erations by eliminating concentrated acetylene byproduct streams Economics The capital investment and operating costs for the combi nation of the KLP process with butadiene extraction are similar or less than twostage extraction processes Commercial plants Eight KLP units are in operation These units repre sent nearly one million metric toy of operating capacity Licensor UOP LLC a COC Cee ey j tistesstttcial PetrochemicalProcesses miele IN ce meena eerste lets ae home processes index company index Butanediol 14 Application To produce 14butanediol BDO or mixture of BDO with tetrahydrofuran THF andor gammabutyrolactone GBL from normal Tetrahydrofuran butane using a fluidbed oxidation and fixedbed hydrogenation reactor Hydrogen product combination Tail gas to Description BP Chemicals has combined its 40 years of experience in incinerator tea slet fluidbed oxidation technology with Lurgi AGs 30 years of hydrogena tion expertise to jointly develop a direct dualreactor process called GEMINOX Air and nbutane are introduced into a fluidbed catalytic reactor 1 ere The fluidbed reactor provides a uniform temperature profile for optimum catalyst performance Reaction gases are cooled and filtered to remove Heavies small entrained catalyst particles and then routed to the recovery section Air to fuel Reactor effluent is contacted with water in a scrubber 2 where essentially 100 of the reactormade maleic anhydride is recovered as maleic acid The process has the capability of coproducing maleic anhydride MAH with the addition of the appropriate purification equipment Scrubber overhead gases are sent to an incinerator for safe disposal The resulting maleic acid from the scrubber is then sent directly to the fixedbed catalytic hydrogenation reactor 3 Reactor yields exceed cost savings and lower operating costs The unique product flexibility 94 BDO By adjustments to the hydrogenation reactor and recovery afforded by this process also allows the user to quickly meet changing purification sections mixtures of BDO with THF andor GBL can be customer and market needs directly produced at comparable overall yields and economics oon The hydrogenation reactor effluent is then sent through a series of commercial Prantsi ere ust worerscale 60000tpy ec BDO distillation steps 4 5 and 6 to produce final market quality products plant in Lima Ohio has been successtully operating since July 2000 Two unique process features are Licensor BP Chemicals and Lurgi AG No continuous liquid waste stream to treatthe water separated in the product purification section is recycled back to the aqueous MAH scrubber 2 No pretreatment of the two catalysts is necessary Economics The GEMINOX technology uses fewer processing steps as a found in competing BDO technologies leading to significant capital tistesstttcial PetrochemicalProce eee C fOCESSES home processes index company index Butanediol 14 Application To produce 14 butanediol BDO from butane via maleic Makeup H anhydride and hydrogen using ester hydrogenation MF MeOH recycle Description Maleic anhydride is first esterified with methanol in a reac Makeup 0H 14 tion column 1 to form the intermediate dimethyl maleate The metha MeOH nol and water overhead stream is separated in the methanol column 2 and water discharged MAL 8 The ester is then fed directly to the lowpressure vaporphase H recycle p H50 roduct THF hydrogenation system where it vaporized into an excess of hydrogen in 2 the vaporizer 3 and fed to a fixedbed reactor 4 containing a copper Product catalyst The reaction product is cooled 5 and condensed 6 with the Heavies 6 G G BDO hydrogen being recycled by the centrifugal circulator 7 The condensed product flows to the lights column 8 where it is distilled to produce a small coproduct tetrahydrofuran THF stream recycle The heavies column 9 removes methanol which is recycled to the methanol column 2 The product column 10 produces highquality butanediol BDO Unreacted ester and gamma butyralactone GBL are recycled to the vaporizer 3 to maximize process efficiency The process can be adapted to produce higher quantities of co product THF and to extract the GBL as a coproduct if required Economics per ton of BDO equivalent Maleic anhydride 1125 Hydrogen 0116 Methanol 0050 Electric power kWh 164 Steam t 36 Water cooling m 326 Commercial plants Since 1989 six plants have been licensed with a total capacity of 300000 tpy Licensor Davy Process Technology UK PROCESSING PetrochemicalProcesses miele IN ce a Mele erste home processes index company index Butene1 Application To produce highpurity butene1 that is suitable for copo lymers in LLDPE production via the Alphabutol ethylene dimerization oe preparation process developed by IFPAxens in cooperation with SABIC Butene1 Description Polymergrade ethylene is oligomerized in the liquidphase reactor 1 with a catalyst system that has high activity and selectivity Liquid effluent and spent catalyst are then separated 2 the liquid is dis Ethylene tilled 3 for recycling of unreacted ethylene to the reactor and fraction feed ated 4 into highpurity butene1 Spent catalyst is treated to remove volatile hydrocarbons and recovered The Alphabutol process features are simple processing high an turndown ease of operation low operating pressure and temperature liquidphase operation and carbon steel equipment The technology has 2 Heavy ends with advantages over other production or supply sources uniformly high Removal spent catalyst quality product low impurities reliable feedstock source low capital costs high turndown and ease of production Yields LLDPE copolymer grade butene1 is produced with a purity ex ceeding 995 wt Typical product specification Is 0 ies butenes butanes ie wn Commercial plants There are 19 licensed units producing 312000 Ethylene 005 wt tpy Sixteen units are in operation C6 olefins 100 ppmw Ethers as DME 2ppmw Licensor Axens Axens NA Sulfur chlorine 1ppmw Dienes acetylenes 5ppmw each CO CO Oz HO0 MeOH 5ppmw each Economics Case for a 2004 ISBL investment at a Gulf Coast location for producing 20000tpy of butene1 is Investment million US 8 Raw material Ethylene tons per ton of butene1 11 Byproducts Cs tons per ton of butene1 008 Typical operating cost US per ton of butene1 38 Butyraldehyde n and Application To produce normal and isobutyraldehyde from propylene and synthesis gas CO H using the LP Oxo SELECTOR Technology reactor Product isomer utilizing a lowpressure rhodiumcatalyzed oxo process removal section separation Description The process reacts propylene with a 11 syngas at low pres Meal iso Butyraldehyce sure 20 kgcm2g in the presence of a rhodium catalyst complexed with a ligand 1 Depending on the desired selectivity the oxonation preeriene reaction produces normal and isobutyraldehyde with typical ni ratios of either 101 or 221 Several different ligand systems are commercially Syngas available which can produce selectivity ratios of up to 301 and as low as 21 The butyraldehyde product is removed from the catalyst solution 2 and purified by distillation 3 Nbutyraldehyde is separated from the iso A nButyraldehyde The SELECTOR Technology is characterized by its simple flow sheet and lowoperating pressure This results in low capital and maintenance expenses and product cost and high plant availability Mild reaction conditions minimize byproduct formation Low byproduct formation also contributes to higher process efficiencies and product qualities Technology for hydrogenation to normal or isobutanols or aldoliza tion and hydrogenation to 2ethylhexanol exists and has been widely spent catalysts can be reactivated onsite The technology is also prac licensed One version of the SELECTOR Technology has been licensed to ticed by Union Carbide Corp at its Texas City and Taft plants produce a mixture of alcohols predominantly 2 propylheptanol from an vl nbutene feedstock and another version to produce higher alcohols up Licensees Twentythree worldwide since 1978 to C15 from Fischer Tropsch produced olefins Licensor Davy Process Technology Ltd UK and Union Carbide Corp a Economics Typical performance data per ton of mixed butyraldehyde subsidiary of The Dow Chemical Co U5 Feedstocks Propylene kg contained in chemical grade 600 Synthesis gas CO H2 Nm 639 Commercial plants The LP Oxo SELECTOR Technology has been licensed for 23 plants worldwide and is now used to produce more than 60 of the worlds butyraldehyde capacity Plants range in size from 30000 b Med rk Cri kGi on iin a to 350000 tpy The rhodiumbased catalyst has a long service life and iste se cal PetrochemicalProcesses miele IN ce a f home processes index company index Cumene Application To produce cumene from benzene and any grade of propyleneincluding lowerquality refinery propylenepropane mix Propylene turesusing the Badger process and a new generation of zeolite cata Benzene recycle Cumene lysts from ExxonMobil ppp Description The process includes a fixedbed alkylation reactor a fixed aa bed transalkylation reactor and a distillation section Liquid propylene and benzene are premixed and fed to the alkylation reactor 1 where propylene is completely reacted Separately recycled polyisopropylben zene PIPB is premixed with benzene and fed to the transalkylation reac tor 2 where PIPB reacts to form additional cumene The transalkylation nea and alkylation effluents are fed to the distillation section The distillation Coa section consists of as many as four columns in series The depropanizer Transalkylation Benzene Cumene PIPB 3 recovers propane overhead as LPG The benzene column 4 recov reactor column column column ers excess benzene for recycle to the reactors The cumene column 5 recovers cumene product overhead The PIPB column 6 recovers PIPB overhead for recycle to the transalkylation reactor Utility requirements per ton of cumene product Process features The process allows a substantial increase in capacity Heat MMkcal import 032 for existing SPA AICl3 or other zeolite cumene plants while improv Steam ton export 060 ing product purity feedstock consumption and utility consumption The utilities can be optimized for specific site conditionseconomics and The new catalyst is environmentally inert does not produce byproduct integrated with an associated phenol plant oligomers or coke and can operate at the lowest benzene to propylene a ratios of any available technology with proven commercial cycle lengths Commercial plants The first commercial application of this process came of over seven years Expected catalyst life is well over five years onstream In 996 At present there are 12 plants operating with a com bined capacity exceeding 52 million mtpy In addition four grassroots Yield and product purity This process is essentially stoichiometric and Plants and an AICI3 revamp are in the design phase Fifty percent of the product purity above 9997 weight has been regularly achieved in worldwide and 75 of Zeolite cumene production are from plants using commercial operation the Badger process Economics Estimated ISBL investment for a 300000mtpy unit on the Licensor Badger Licensing LLC US Gulf Coast 2004 construction basis is US15 million PROCESSING PetrochemicalProcesses at MOTH PLUCIS home processesindex company index Cumene Application Advanced technology to produce highpurity cumene from propylene and benzene using patented catalytic distillation CD EeHZERE ence technology The CDCumene process uses a specially formulated zeolite p alkylation catalyst packaged in a proprietary CD structure and another specially formulated zeolite transalkylation catalyst in loose form Description The CD column 1 combines reaction and fractionation in a singleunit operation Alkylation takes place isothermally and at Propylene low temprature CD also promotes the continuous removal of reaction products from reaction zones These factors limit byproduct impurities and enhance product purity and yield Low operating temperatures and pressures also decrease capital investment improve operational safety 4 and minimize fugitive emissions In the mixedphase CD reaction system propylene concentration pire in the liquid phase is kept extremely low 01 wt due to the higher volatility of propylene to benzene This minimizes propylene oligomerization the primary cause of catalyst deactivation and results in catalyst run lengths of 3 to 6 years The vaporliquid equilibrium effect provides propylene dilution unachievable in fixedbed systems even with expensive reactor pumparound andor benzene recycle arrangements Economics Based on a 300000mtpy cumene plant located in the US Overhead vapor from the CD column 1 is condensed and returned Gulf Coast the ISBL investment is about US15 million as reflux after removing propane and lights P The CD column bottom Typical operating requirements per metric ton of cumene section strips benzene from cumene and heavies The distillation train Propylene 0353 separates cumene product and recovers polyisopropylbenzenes PIPB Benzene 0650 and some heavy aromatics H from the net bottoms PIPB reacts with Yield 997 Utilities benzene in the transalkylator 2 for maximum cumene yield Operating Electricity KWh 8 conditions are mild and noncorrosive standard carbon steel can be used Heat import 10 kcal 05 for all equipment Steam export mt 10 Water cooling m 12 Yields 100000 metric tons mt of cumene are produced from 65000 mt of benzene and 35300 mt of propylene giving a product yield of over 997 Cumene product is at least 9995 pure and has a Bro a mine Index of less than 2 without clay treatment Commercial plants Formosa Chemicals Fibre Corporation Taiwan 540000 mtpy Licensor CDTECH a partnership between ABB Lummus Global and Chemical Research Licensing Cumene continued PROCESSING PetrochemicalProcesses PROCESSING EATATROIO LS home processesindex company index Cumene Application To produce highquality cumene isopropylbenzene by alkylating benzene with propylene typically refinery or chemical Benzene Recycle benzene nen grade using liquidphase QMax process based on zeolitic catalyst technology DIPB Description Benzene is alkylated to cumene over a zeolite cata lyst in a fixedbed liquidphase reactor Fresh benzene is combined Propyfene with recycle benzene and fed to the alkylation reactor 1 The ben zene feed flows in series through the beds while fresh propylene teed is distributed equally between the beds This reaction is highly exothermic and heat is removed by recycling a portion of reactor effluent to the reactor inlet and injecting cooled reactor effluent between the beds Heavies In the fractionation section propane that accompanies the propylene feedstock is recovered as LPG product from the overhead of the depropanizer column 2 unreacted benzene is recovered from the overhead of the benzene column 4 and cumene product is taken as overhead from the cumene column 5 Diisopropylbenzene DIPB is recovered in the overhead of the DIPB column 6 and recycled to the Egonomies Basis ISBL US Gulf Coast transalkylation reactor 3 where it is transalkylated with benzene over a Investment UStpy 4090 second zeolite catalyst to produce additional cumene A small quantity es Raw materials utilities per metric ton of cumene of heavy byproduct is recovered from the bottom of the DIPB column Propylene tons 035 6 and is typically blended to fuel oil The cumene product has a high Benzene tons 066 purity 99969997 wt and cumene yields of 997 wt and higher Electricity kW 12 are achieved Steam tons import 07 The zeolite catalyst is noncorrosive and operates at mild conditions Water cooling m 3 thus carbonsteel construction is possible Catalyst cycle lengths are two The QMax design is typically tailored to provide optimal utility years and longer The catalyst is fully regenerable for an ultimate catalyst advantage for the plant site such as minimizing heat input for stand lite of six years and longer Existing plants that use SPA or AICI3 catalyst alone operation or recovering heat as steam for usage in a nearby can be revamped to gain the advantages of QMax cumene technology phenol plant while increasing plant capacity Commercial plants Seven QMax units are in operation with a total cumene capacity of 23 million tpy and two additional units are either in design or under construction Licensor UOP LLC Cumene continued iste se cal PetrochemicalProcesses miele IN ce meena MO etsyslete eae home processes index company index Cyclohexane Application Produce highpurity cyclohexane by liquidphase catalytic HP purge gas hydrogenation of benzene sg Catalyst q Description The main reactor 1 converts essentially all the feed isother ir mally in the liquid phase at a thermodynamicallyfavorable low temper aan separator ature using a continuouslyinjected soluble catalyst The catalysts high Steam LP purae aas activity allows use of low hydrogen partial pressure which results in few Benzene ne er side reactions eg isomerization or hydrocracking The heat of reac Hydrogen 2 cw Cyclohexane tion vaporizes cyclohexane product and using pumparound circulation t BFW through an exchanger also generates steam 2 With the heat of reaction Finishing HP separator being immediately removed by vaporization accurate temperature con reactor or stripper trol is assured A vaporphase fixedbed finishing reactor 3 completes the catalytic hydrogenation of any residual benzene This step reduces resid ee cea oo ual benzene in the cyclohexane product to very low levels Depending on Optional the purity of the hydrogen makeup gas the stabilization section includes either an LP separator 4 or a small stabilizer to remove the light ends A prime advantage of the liquidphase process is its substantially lower cost compared to vapor phase processes investment is particularly low because a single inexpensive main reactor chamber is used compared to multiplebed or tubular reactors used in vapor phase processes Quench Commercial plants Thirtythree cyclohexane units have been licensed gas and unreacted benzene recycles are not necessary and better heat recovery generates both the cyclohexane vapor for the finishing step and Licensor Axens Axens NA a greater amount of steam These advantages result in lower investment and operating costs Operational flexibility and reliability are excellent changes in feedstock quality and flows are easily handled Should the catalyst be deactivated by feed quality upsets fresh catalyst can be injected without shutting down Yield 1075 kg of cyclohexane is produced from 1 kg of benzene Economics Basis 200000tpy cyclohexane complex ISBL 2005 Gulf Coast location with PSA hydrogen is US8 million Catalyst cost is US 12metric ton of product hietesciit cal PetrochemicalProcesses eee eer Dimethyl ether DME Application To produce dimethyl ether DME from methanol using Toyo Engineering Corps TECs DME synthesis technology based on metha nol dehydration process Feedstock can be crude methanol as well as refined methanol 2 Description If feed is crude methanol water is separated out in the o 1 methanol column 1 The treated feed methanol is sent to a DME Reac i tor 2 after vaporization in 3 The synthesis pressure is 1020 MPaG The inlet temperature is 220250C and the outlet is 300350C OG Ye Methanol onepass conversion to DME is 7085 in the reactor The a f reactor effluents DME with byproduct water and unconverted metha OME nolare fed to a DME column 4 after heat recovery and cooling Meter OH Inthe DME column 4 DME is separated from the top and condensed crude The DME is cooled in a chilling unit 5 and stored in a DME tank 6 as a methanol product Water and methanol are discharged from the bottom and fed to a methanol column 1 for methanol recovery The purified methanol from this column is recycled to the reactor after mixing with feedstock methanol Economics The methanol consumption for DME production is approxi mately 14 tonmethanol per tonDME Commercial plants A 10000tpy unit was commissioned in August 2003 in China and is the first fuel DME facility A second 110000tpy facility is scheduled to start up in the third quarter of 2005 in China and will be the largest DME plant Reference Mii T Commercial DME plant for fuel use First Interna tional DME Conference Paris France Oct 1214 2004 Licensor Toyo Engineering Corp TEC a Ce Oe j PROCESSING PetrochemicalProc eee C Oe Sich home processes index company index Dimethyl terephthalate Application To increase capacity and reduce energy usage of existing Atm ray or grassroots dimethyl terephthalate DMT production facilities using variations of GTDMT proprietary technology ff Description The common production method of DMT from paraxylene Methanol ww and methanol is through successive oxidations in four major steps oxi dation esterification distillation and crystallization A mixture of pxy is aac DMT lene and methyl ptoluate MPT is oxidized with air using a heavymetal MPT pete pure catalyst All organics are recovered from the offgas and recycled to the Air ni E DMT system The acid mixture from the oxidation is esterified with methanol Heavy design oxidation and produces a mixture of esters The crude ester mixture is distilled to boilers recovery ond remove all heavy boilers and residue produced lighter esters are recy isomer removal cled to the oxidation section Raw DMT is then sent to the crystallization Residue section to remove DMT isomers and aromatic aldehydes The technology improvements enhance the traditional processing in each section The adaptations include changes in process configurations and operating conditions alterating the separation schemes revising the recovery arrangement increasing the value of the byproducts and reducing the overall plant recycles GTC Technology offers complete implementation of the technology on operating costs and overall plant reviews for selective improvements to reduce operating 7 Operating reviews to reduce operating downtime and extend and overall production costs Some separate improvements available online factors are 8 Advanced control models for improved operability 1 Oxidation optimization reduces byproduct formation thus lowering a pxylene consumption Economics Based on process modifications an existing DMT plant can 2 Recoveries of byproducts for sale such as methyl benzoate MeBz Increase production with an investment of 200 to 600tpy of addi and acetic and formic acid tional capacity A new plant will have an investment reduction of about 3 Improved esterifier reactor design enables higher throughputs 29 equipment cost Raw material consumption per ton of product and improves methanol usage with the complete modification is 605 tons of paraxylene and 360 tons 4 Enhanced isomer removal minimizes DMT losses of methanol 5 Improved crystallization schemes for reduced energy lowers methanol handling and losses improves purity and operating flexibility h 6 Integration of steam usage in the plant for considerable savings Commercial plants GTDMT technology is used by seven DMT produc ers Licensor GTC Technology Dimethyl terephthalate continued PROCESSING PetrochemicalProcesses were Ne AMMO TC Ae AT Ore Ut aa f E home processes index company index Dimethylformamide Application To produce dimethylformamide DMF from dimethylamine DMA and carbon monoxide CO Synthesis DMA recovery ee Vaporization DME Description Anhydrous DMA and CO are continuously fed to a spe cialized reactor 1 operating at moderate conditions and containing a catalyst dissolved in solvent The reactor products are sent to a separa tion system where crude product is vaporized 2 to separate the spent Catalyst catalyst Excess DMA and catalyst solvent are stripped 3 from the crude product and recycled back to the reaction system Vacuum distillation DMA 4 followed by further purification 5 produces a highquality solvent a and fibergrade DMF product A saleable byproduct stream is also pro duced Spent Byproduct Yields Greater than 95 on raw materials CO yield is a function of its catalyst quality Economics Typical performance data per ton of product Dimethylamine t 063 Carbon monoxide t 041 Steam t 13 Water cooling m 100 Electricity kWh 10 Commercial plants Thirteen plants in eight countries use this process with a production capacity exceeding 100000 mtpy Licensor Davy Process Technology UK a COC Cee ey a PROCESSING PetrochemicalProcesses PROCESSING home processes index company index EDC via oxygenlean oxychlorination Application The modern Vinnolit oxychlorination process produces a ethylene dichloride EDC by an exothermic reaction from feedstocks Quench Vent to incineration including ethylene anhydrous hydrogen chloride HCI and oxygen An Hy hydrous HC can be used from the VCM process as well as from other Boilerfeedwater steam processes such as isocyanates MDI TDI chlorinated methanes chlori aa nated ethanes epichlorohydrin etc Decanter Oxygen can be supplied from an air separation plant as well as from HCI the costeffective pressure swing adsorption PSA process The Vinnolit Oxygen oxychlorination process is also able to handle ethylene andor anhydrous Ethylene acre HCl containing vent streams from direct chlorination acetaldehyde heads monochloroacetic acid and other processes Compressor column To wastewater treatment Description The exothermic reaction is catalyzed by a copper chloride Hydrogenation OCreactor Cat filtration EDC catalyst in a singlestep fluidizedbed reactor at temperatures of 220C reactor condensation Heat of reaction is recovered by producing 10 bar g steam or heating other heattransfer fluids The small amount of catalyst fines that pass through the highly efficient cyclone system are removed by a newly developed hotgas catalyst filter or alternatively by wastewater treatment that meets even the strictest regulations for copper dioxins and furanes The environmentally Safety The oxygen is mixed with anhydrous HCI outside the reactor friendly process uses recycle gas which is fed back to the reactor after and is fed independently of the ethylene into the fluidized bed The condensing EDC and water oxygen concentration in the recycle stream is approximately 05 vol After removal of carbon dioxide CO3 and chloralchloroethanol which is well outside the explosion range the crude EDC is purified in the EDC distillation unit it can be used as Environment friendly A highly efficient hotgas filtration system furnace feed or sales EDC separates the small quantities of catalyst fines Besides the EDC removal via steam stripping no additional wastewater treatment is required Process features and economics are oo The charter for European Council for Vinyl Manufacturers ECVM Low manufacturing costs The unlimited catalyst lifetime is combined easily met EDC 5gt of EDC purification capacity copper 1gt with the low losses via the highly efficient cyclone system less than 15g og oxychlorination capacity dioxinlike components 1yg TEQt of catalyst per metric ton mton of EDC produced High rawmaterial yields oxychlorination capacity 985 ethylene 99 anhydrous HCI and 94 oxygen high crude EDC purity 995 and the possibility of using lowcost oxygen from PSA h units ensure a highly competitive process with low production costs Reliability A stable temperature control combined with an excellent heat transfer and a uniform temperature profile no hot spots in the fluidized bed easily achieves an onstream time 99 per year A specially designed rawmaterial sparger system allows operation spans of two years without maintenance Larger heattransfer area allows a higher steam temperature and pressure in the cooling coils which improves the safety margin to the critical surface temperature where hydrochloric acid dewpoint corrosion may occur Flexibility A turndown ratio as low as 20 capacity utilization can be achieved as well as quick load changes Commercial plants The process is used in 20 reactors at 15 sites with annual single reactor capacities up to 320000 mtons of EDC alone as HClconsuming plant or as part of the balanced VCM process In some cases it has replaced other oxychlorination technologies from different licensors by replacing existing reactors or existing units Two new oxy chlorination trains were successfully commissioned in September 2004 one oxychlorination unit is under design Licensor Vinnolit Contractor Uhde GmbH EDC via oxygenlean oxychlorination continued PROCESSING PetrochemicalProcesses home processes index company index EDC via hightemperature chlorination Application Vinnolits new hightemperature direct chlorination DC re boll actor provides an energy efficient technology for the production of fur Column reboiler eco nace feed and sales ethylene dichloride EDC without distillation from ant gas to chlorine and ethylene oxychlorination Description The liquid phase reaction of ethylene and chlorine releases approximately 220 kJmol of produced EDC EDC column In asimple carbon steel ushaped loop reactor chlorine and ethylene are separately dissolved in EDC before the reaction takes place In combination with the special Vinnolit catalyst this method significantly Chlorine minimizes byproduct formation Ethylene Furmace feed 0 Downstream of the reaction zone the lower static pressure permits the reactor content to boil and applies the thermosyphon effect for circulation EDC vapor leaves the horizontal vessel and either enters the reboiler of a column eg reboiler of highboilheads andor vacuum column or a heat exchanger which condenses the EDC vapor The reaction heat is transferred to the column indirectly A fraction of the condensed EDC is fed back to the reactor and the rest is directly sent to the EDC cracker without further distillation Because of the high yields the Vinnolit DC reactor can be operated Low capital costs A simple design with a minimized number of in the standalone mode However if the reactor is part of a complete equipment results in low unit investment costs VCM plant offgas can be sent to the oxychlorination reactor to recover Energy savings Vinnolits DC process significantly reduces the the remaining small quantities of ethylene If salesEDC specification is steam consumption ina balanced EDCVCM plant The saving of steam the target only a small stripper column is required to eliminate traces is approximately 600 kg per metric ton mton of EDC produced The of HCl reaction heat can preferably be used in the EDC distillation Process features and economics are Simple process The HTCboiling reactor is simple due to elimination a of washing equipment wastewater treatment and EDC distillation Low manufacturing costs High raw material yields 999 for N 5 f ew Catalyst The Vinnolit DC catalyst guarantees a furnace feed ethylene and 998 for chlorine and a product quality which requires oy wt arte EDC quality of 999 without any distillation Catalyst makeup is not no further treatment ensure a highly competitive process with low production costs The HTC high temperature chlorination boiling reactor is simple because no EDC washing wastewater treatment and h click here to email for more Information a EDC distillation facilities are necessary required Operability and maintainability A corrosion inhibiting catalyst system and simple equipment without major moving parts keep the maintenance costs low Less plot area The plot area requirement for the DC boiling reactor unit is very small and can be accommodated to customers needs Commercial plants The DCprocess andor DCcatalyst are used for the annual production of more than 65 million mtons of EDC One unit with an annual capacity of 320000 mtons of EDC has been successfully commissioned in September 2004 Another unit with the latest plant design is currently under construction Licensor Vinnolit Contractor Uhde GmbH EDC via hightemperature chlorination continued PROCESSING PetrochemicalProcesses E home processes index company index Ethanolamines Application To produce monoMEA diDEA and triethanolamines TEA from ethylene oxide and ammonia DEA Synthesis Dehydration Description Ammonia solution recycled amines and ethylene oxide are NH Product fed continuously to a reaction system 1 that operates under mild con eee purification ditions and simultaneously produces MEA DEA and TEA Product ratios MEA LL can be varied to maximize MEA DEA or TEA production The correct selection of the NH3EO ratio and recycling of amines produces the de Reyer TEA sired product mix The reactor products are sent to a separation system where ammonia 2 and water are separated and recycled to the reac Ammonia tion system Vacuum distillation 4567 is used to produce pure MEA 7 DEA and TEA A saleable heavies tar byproduct is also produced Techni Ethylene oxide cal grade TEA 85 wt can also be produced if required Tar byproduct Yields Greater than 98 on raw materials Economics Typical performance data per ton amines MEADEATEA product ratio of 34 Ethylene oxide t 082 Ammonia t 019 Steam t 5 Water cooling m 300 Electricity kWh 30 Commercial plants One 20000mtpy original capacity facility Licensor Davy Process Technology UK a Ce Oe j iste se cal PetrochemicalProcesses miele IN ce OTC ATTA AT TOre ted R01C erste ers eae home processes index company index EthersETBE Application The Uhde Edeleanu ETBE process combines ethanol and isobutene to produce the highoctane oxygenate ethyl tertiary butyl ni Sener we Fthanoliwater ether ETBE reactor was separation BB raffinate Feeds C cuts from steam cracker and FCC units with isobutene con tents ranging from 12 to 30 Products ETBE and other tertiary alkyl ethers are primarily used in gas Ethanolwater oline blending as an octane enhancer to improve hydrocarbon com Se bustion efficiency Moreover blending of ETBE to the gasoline pool will lower vapor pressure Rvp C4 feedstock Description The Uhde Edeleanu technology features a twostage re Ethanol actor system of which the first reactor is operated in the recycle mode ETBE product With this method a slight expansion of the catalyst bed is achieved that ensures very uniform concentration profiles in the reactor and most important avoids hot spot formation Undesired side reactions such as the formation of diethyl ether DEE are minimized The reactor inlet temperature ranges from 50C at startofrun to about 65C at endofrun conditions One important feature of the two Uttility requirements C feed containing 21 isobutene per metric stage system is that the catalyst can be replaced in each reactor sepa ton of ETBE Steam LP kg 110 rately without shutting down the ETBE unit Steam MP kg 1000 The catalyst used in this process is a cationexchange resin and is available Electricity kWh 35 from several manufacturers Isobutene conversions of 94 are typical for Water cooling m 24 FCC feedstocks Higher conversions are attainable when processing steam Commercial plants The Uhde Edeleanu proprietary ETBE process has cracker Ca cuts that contain isobutene concentrations of about 25 7 been successfully applied in two refineries converting existing MTBE ETBE Is recovered as the bottoms product of the distillation unit The units Another MTBE plant is in the conversion stage ethanolrich Cy distillate is sent to the ethanol recovery section Water is used to extract excess ethanol and recycle it back to process At the top Licensor Unde GmbH of the ethanolwater separation column an ethanolwater azeotrope is recycled to the reactor section The isobutenedepleted C stream may be sent to a raffinate stripper or to a molsievebased unit to remove oxygenates such as DEE ETBE ethanol and tertbutanol PROCESSING PetrochemicalProcesses tet St MOTE ALAA N72 Ue M01 exot0 27 a home processesindex company index EthersMTBE Application The Uhde Edeleanu MTBE process combines methanol and isobutene to produce the highoctane oxygenatemethy tertiary MIBE reactor Debutanizer wae onan butyl ether MTBE BB raffinate Feeds Ccuts from steam cracker and FCC units with isobutene con tents range from 12 to 30 Products MTBE and other tertiary alkyl ethers are primarily used in gas oline blending as an octane enhancer to improve hydrocarbon combus tion efficiency Cy feedstock Description The technology features a twostage reactor system of which the first reactor is operated in the recycle mode With this meth Methanol od a slight expansion of the catalyst bed is achieved which ensures very MTBE product uniform concentration profiles within the reactor and most important avoids hot spot formation Undesired side reactions such as the forma tion of dimethyl ether DME are minimized The reactor inlet temperature ranges from 45C at startofrun to about 60C at endofrun conditions One important factor of the two stage system is that the catalyst may be replaced in each reactor sepa through a debutanizer column with structured packings containing ad rately without shutting down the MTBE unit ditional catalyst This reactive distillation technique is particularly suited The catalyst used in this process is a cationexchange resin and is when the raffinatestream from the MTBE unit will be used to produce available from several catalyst manufacturers Isobutene conversions of 2 highpurity butene1 product 97 are typical for FCC feedstocks Higher conversions are attainable For a C4 cut containing 22 isobutene the isobutene conversion when processing steamcracker C cuts that contain isobutene concen May exceed 98 at a selectivity for MTBE of 995 trations of 25 ae Utility requirements C feed containing 21 isobutene per metric ton MTBE is recovered as the bottoms product of the distillation unit 4 MTBE The methanolrich C distillate is sent to the methanolrecovery section Steam LP kg 900 Water is used to extract excess methanol and recycle it back to process Steam MP kg 100 The isobutenedepleted C stream may be sent to a raffinate stripper Electricity kWh 35 or to a molsievebased unit to remove other oxygenates such as DME Water cooling m 15 MTBE methanol and tertbutanol Very high isobutene conversion in excess of 99 can be achieved Commercial plants The Uhde Edeleanu proprietary MTBE process has been successfully applied in five refineries The accumulated licensed capacity exceeds 1 MMtpy Licensor Uhde GmbH EthersMTBE continued PROCESSING PetrochemicalProcesses aides ee ett JULI home processes index company index Ethyl acetate Application To produce ethyl acetate from ethanol without acetic acid or other cofeeds Dehydrogenation Refining Hydrogen Description Ethanol is heated and passed through a catalytic dehydro genation reactor 1 where part of the ethanol is dehydrogenated to form ethyl acetate and hydrogen The product is cooled in an integrated heatexchanger system hydrogen is separated from the crude prod uct The hydrogen is mainly exported Crude product is passed through Loy Ethanol feed a second catalytic reactor 2 to allow polishing and remove minor byproducts such as carbonyls ee The polished product is passed to a distillation train 3 where a hydrogenation Ethyl acetate novel distillation arrangement allows the ethanolethyl acetate water product azeotrope to be broken Products from this distillation scheme are Rea eeenensl unreacted ethanol which is recycled and ethyl acetate product The process is characterized by lowoperating temperatures and pressures which allow all equipment to be constructed from either carbon steel or lowgrade stainless steels It allows ethyl acetate to be made without requiring acetic acid as a feed material The process is appropriate for both synthetic ethanol and fermentation ethanol as the feed The synthetic ethanol can be impure ethanol without significantly Licensor Davy Process Technology UK affecting the conversion or selectivity The product ethyl acetate is greater than 9995 Economics Typical performance data per ton of ethyl acetate pro duced Feedstock 112 tons of ethanol Product 45 kg of hydrogen Commercial plants The technology has been developed during the mid to late 1990s The first commercial plant is a 50000tpy plant in South Africa using synthetic ethanol Licensees One since 1998 PROCESSING PetrochemicalProcesses miele IN ce a home processes index company index Ethylbenzene Application Advanced technology to produce highpurity ethylbenzene EB alkylating benzene with ethylene using patented catalytic distilla oral aa ane cokan calunn tion CD technology The CDTECH EB process uses a specially formu pp lated zeolite alkylation catalyst packaged in a proprietary CD structure Benzene Ethyfbenzene The process is able to handle a wide range in ethylene feed composi tionfrom 10 to 100 ethylene 2 Description The CD alkylator stripper 1 operates as a distillation col Ethylene XK umn Alkylation and distillation occur in the alkylator in the presence of a zeolite catalyst packaged in patented structured packing Unreacted ethylene and benzene vapor from the alkylator top are condensed and fed to the finishing reactor 2 where the remaining ethylene reacts over Flux oil zeolite catlayst pellets The alkylator stripper bottoms is fractionated 4 5 into EB product polyethylbenzenes and flux oil The polyethylben Polyethylbenzenes zenes are transalkylated with benzene over zeolite catalyst pellets in the transalkylator 3 to produce additional EB The ethylene can be polymer grade or with only minor differences in the process scheme dilute eth ylene containing as little as 10 mol ethylene as in FCC offgas Reactors are designed for 3 to 6 years of uninterrupted runlength The process does not produce any hazardous effluent Low operating temperatures Benzene kg 738 allow using carbon steel for all equipment Water cooling m2 Yields and product quality Both the alkylation and transalkylation reactions ree oil me export be are highly selectiveproducing few byproducts The EB product has high purity 999 wt minimum and is suitable for styreneunit feed Xylene Commercial plants Three commercial plants are in operation in Argen make is less than 10 ppm The process has an overall yield of 997 tina and Canada with capacities from 140000 to 816000 mtpy They process ethylene feedstocks with purities ranging from 75 ethylene to Economics The EB process features consistent product yields high polymergrade ethylene An 850000mtpy unit using dilute ethylene is procuct ey lowenergy consumption low investment cost and easy currrently under construction reliable operation Investment 500000 tpy ISBL Gulf Coast US 17 million Raw materials and utilities based on one metric ton of eB Sa click here to email for more information a Licensor CDTECH a partnership between ABB Lummus Global and Chemical Research Licensing Ethylbenzene continued PROCESSING PetrochemicalProc eee C Oe Sich home processes index company index Ethylbenzene Application To produce ethylbenzene EB from benzene and a poly mergrade ethylene or an ethyleneethane feedstock using the Bad Senzene ger EBMax process and proprietary ExxonMobil alkylation and trans Light alkylation catalysts The technology can be applied in the design of column EB product grassroots units upgrading of existing vaporphase technology plants Alkylation or conversion of aluminum chloride technology EB plants to zeolite eal technology Gi Description Ethylene reacts with benzene in either a totally liquidfilled Ethylene or mixedphase alkylation reactor 1 containing multiple fixedbeds of Residue ExxonMobils proprietary catalyst forming EB and very small quantities Recycle PEB of polyethylbenzenes PEB In the transalkylation reactor 2 PEB is con verted to EB by reaction with benzene over ExxonMobils transalkylation Transalkylation Benzene rr pee catalyst PEB and benzene recovered from the crude EB enter the trans reactor column column column alkylation reactor Effluents from the alkylation and transalkylation reactors are fed to the benzene column 3 where unreacted benzene is recovered from crude EB The fresh benzene feedstock and a small vent stream from the benzene column are fed to the lights column 4 to reject light im Product quality The EB product contains less than 100 ppm of Cg plus purities The lights column bottoms is returned to the benzene column Cg impurities Product purities of 9995 to 9999 are expected The bottoms from the benzene column is fed to the EB column 5 to recover EB product The bottoms from the EB column is fed to the PEB Economics column 6 where recyclable alkylbenzenes are recovered as a distillate Raw materials and steam tons per ton of EB product Ethylene 0265 and diphenyl compounds are rejected in a bottoms stream that can be Benzene 0739 used as fuel Steam highpressure used 098 Steam medium and lowpressured generated 139 Catalysts Cycle lengths in excess of four years are expected for the Utilit be optimized Fic sit dit alkylation and transalkylation catalysts Process equipment is fabri INTES CAP DE OPUMIZER TOF SPECIIC SIE CONGHIONS cated entirely from carbon steel Capital investment is reduced as a Commercial plants Since the commercialization of the Badger EB tech consequence of the high activity and extraordinary selectivity of the nology in 1980 45 licenses have been granted The total licensed capac alkylation catalyst and the ability of both the alkylation and transalkyl ation catalysts to operate with very low quantities of excess benzene i ity for the Badger EB technology exceeds 17 million mtpy The capacity for the EBMax technology exceeds 106 million mtpy Licensor Badger Licensing LLC Ethylbenzene continued PROCESSING PetrochemicalP eee C OCESSES home processes index company index Ethylbenzene Application Stateoftheart technology to produce highpurity ethylben Ethylbenzene zene EB by liquidphase alkylation of benzene with ethylene The Lum musUOP EBOne process uses specially formulated proprietary zeolite catalyst from UOP The process can handle a wide range of ethylene feed compositions ranging from chemical 70 to polymer grade 100 ix A rn Description Benzene and ethylene are combined over a proprietary zeo Ethylene EX ye lite catalyst in a fixedbed liquidphase reactor Fresh benzene is combined DX S with recycle benzene and fed to the alkylation reactor 1 The combined DX CO benzene feed flows in series through the beds while fresh ethylene feed BX Flux oil is distributed equally between the beds The reaction is highly exothermic a and heat is removed between the reaction stages by generating steam Polyethylbenzene Unreacted benzene is recovered from the overhead of the benzene col Benzene umn 3 and EB product is taken as overhead from the EB column 4 Recycle benzene A small amount of polyethylbenzene PEB is recovered in the over head of the PEB column 5 and recycled back to the transalkylation reactor 2 where it is combined with benzene over a second proprietary zeolite catalyst to produce additional EB product A small amount of flux oil is recovered from the bottom of the PEB column 5 and is usually Investment ISBL Gulf Coast USmtpy 3045 burned as fuel Raw material and utilities per metric ton of EB The catalysts are noncorrosive and operate at mild conditions al ptnylene mons oe lowing for all carbonsteel construction The reactors can be designed Utilities US 4 for 26 year catalyst cycle length and the catalyst is fully regenerable Additional utility savings can be realized via heat integration with The process does not produce any hazardous effluent downstream LummusUOP Classic SM or SMART SM styrene unit Yields and product quality Both the alkylation and transalkylation reactions Commercial plants Nineteen EBOne units are in operation throughout are highly selective producing few byproducts The EB product has a high the world with a total EB capacity of 57 million mtpy Unit capacities purity 999 wt minimum and is suitable for styreneunit feed Xylene range from 65000 to 725000 mtpy Ethylene feedstock purity ranges make is less than 10 ppm The process has an overall yield of 997 from 80 to 100 Nine additional units are either in design or under Economics The EBOne process features consistently high product yields construction the largest unit is 770000 mtpy over the entire catalyst life cycle highproduct purity lowenergy con sumption low investment cost and simple reliable operation Licensor ABB Lummus Global and UOP LLC Ethylbenzene continued PROCESSING PetrochemicalP eee C ICAIF TOCESSES home processes index company index Ethylene Application To produce polymergrade ethylene 9995 vol Major SRT cracking Acid gas byproducts are propylene chemical or polymergrade a butadienerich Feed furnace hidrooen Cg to Cg aromaticsrich pyrolysis gasoline and highpurity ee a ee o stm 3 stm en removal Description Hydrocarbon feedstock is preheated and cracked in the oh tg v4 CC presence of steam in tubular SRT short residence time pyrolysis furnaces 1 This approach features extremely high olefin yields long runlength Pyrolysis fuel ol and mechanical integrity The products exit the furnace at 1500F to ed Ethylene H2 Propylene Mixed Cs 1600F and are rapidly quenched in the transfer line exchangers 2 Methane H that generate super highpressure SHP steam The latest generation Chilling E E f furnace design is the SRT VI train and Furnace effluent after quench flows to the gasoline fractionator Ethane p Pyrolysis 3 where the heavy oil fraction is removed from the gasoline and lighter 5 no gasoline fraction liquids cracking only Further cooling of furnace effluents is accomplished by a direct water quench in the quench tower 4 Raw gas from the quench tower is compressed in a multistage centrifugal compressor 5 to greater than 500 psig The compressed gas is then dried 6 and chilled Hydrogen is recovered in the chilling train 7 which feeds the demethanizer 8 The demethanizer operates at about A revised flow scheme eliminates 25 of the equipment from this 100 psia providing increased energy efficiency The bottoms from the conventional flowsheet It uses CDHydro hydrogenation for the selective demethanizer go to the deethanizer 9 hydrogenation of C through Cy acetylenes and dienes in a single tower Acetylene in the deethanizer overhead is hydrogenated 10 or reduces the crackedgas discharge pressure to 250 psig uses a single recovered The ethyleneethane stream is fractionated 11 and polymer refrigeration system to replace the three separate systems and applies grade ethylene is recovered Ethane leaving the bottom of the ethylene metathesis to produce up to 13 of the propylene product catalytically fractionator is recycled and cracked to extinction rather than by thermal cracking thereby lowering energy consumption The deethanizer bottoms and condensate stripper bottoms from by 15 he char mpression m ar ropanized 12 Methy ve ropa diene are hy wrogens 4 nine demopanicer ising Coho Energy consumption Energy consumptions are 3300 kcalkg of ethylene catalytic distillation hydrogenation technology The depropanizer bottoms produced for ethane cracking and 5000 kcalkg of ethylene for naphtha is separated into mixed Cy and light gasoline streams 14 Polymergrade propylene is recovered in a propylene fractionator 13 feedstocks Energy consumption can be as low as 4000 kcalkg of ethyl ene for naphtha feedstocks with gas turbine integration As noted above the new flow scheme reduces energy consumption by 14 Commercial plants Approximately 40 of the worlds ethylene plants use Lummus ethylene technology Many existing units have been sig nificantly expanded above 150 of nameplate using Lummus MCET maximum capacity expansion technology approach Licensor ABB Lummus Global Ethylene continued PROCESSING PetrochemicalProcesses miele IN ce a home processes index company index Ethylene Application High performance steamcracking and recovery to produce polymergrade ethylene and propylene butadienerich mixed Cys aro fen HP superheated maticrich pyrolysis gasoline hydrogen and fuel streams Cracking feed 6 9 6 steam stocks range from ethane through vacuum gas oils Hydrocarbon SO D feedstock Description Kellogg Brown Roots proprietary Selective Cracking ocesssteam ri q x W Optimum REcovery SCORE olefins technology represents the integra audlgestve al Ethylene tion of the technologies of the former MW Kellogg and Brown Root Tailgas companies combined with olefins technology developed by ExxonMobil Propylene 2 vi py a Hydrogen Chemical Co through a longterm worldwide licensing agreement fill ExxonMobil brings innovative technology as well as the benefits of ex ae 18 g g 13 12 tensive operating experience to further improve operability reliability and reduce production costs Pyrolysis gasoline The SCORE pyrolysis furnace portfolio features the straight tube SC1 Propane recycle Ethane recycle design which has a low reaction time in the range of 01 seconds and low operating pressures The design and operating conditions produce higher olefin yields The portfolio includes a range of designs to satisfy any requirements The pyrolysis furnace 1 effluent is processed for heat and product recovery in an efficient reliable lowcost recovery section The recovery Cracked gases are cooled and fractionated to remove fuel oil and section design can be optimized for specific applications andor selected water 25 then compressed 6 processed for acidgas removal 8 based on operating company preferences Flowschemes based on de and dried 9 The C3 and lighter material is separated as an overhead ethanizerfirst depropanizerfirst and demethanizerfirst configurations Product in the depropanizer 10 and acetylene is hydrogenated in the are available The depropanizerfirst flowscheme primarily applicable to acetylene converter 11 The acetylene converter effluent is processed in liquid crackers is shown above The similar but simpler deethanizerfirst the demethanizer system 1214 to separate the fuel gas and hydrogen scheme is appropriate for ethane through ethanepropane gas crackers products The demethanizer bottoms is sent to the deethanizer 15 from These two schemes use frontend acetylene converter systems which which the overhead flows to the Csplitter 16 which produces the minimize greenoil production and allow using lowpressure recovery Polymergrade ethylene product and the ethane stream which is typically towers KBR also has extensive experience with the demethanizer recycled to the furnaces as a feedstock The deethanizer bottoms flows first flowscheme which can be offered to clients preferring that to the C3splitter 18 where the polymergrade propylene is recovered technology as the overhead product The C3splitter bottoms product propane is typically recycled to the furnaces as a feedstock The depropanizer bottoms product C4s and heavier flow to the debutanizer 19 for recovery of the mixedC4 product and aromaticrich pyrolysis gasoline Yields Ethylene yields to 84 for ethane 38 for naphtha and 32 for gas oils may be achieved depending upon feedstock characteristics Commercial plants KBR has been involved in over 140 ethylene projects worldwide with singletrain ethylene capacities up to 13 million tpy in cluding 21 new grassroots ethylene plants since 1990 Licensor Kellogg Brown Root Inc Ethylene continued PROCESSING PetrochemicalP eee C Oe Sich home processes index company index Ethylene Application To produce polymergrade ethylene and propylene by ther F a as eed Dilution Cracked 6 C mal cracking of hydrocarbon fractionsfrom ethane through naphtha steam gas 2 H up to hydrocracker residue Byproducts are a butadienerich C stream ialc2 comp i ression a C6Cg gasoline stream rich in aromatics and fuel oil O 2 Description Fresh feedstock and recycle streams are preheated and I 4 cracked in the presence of dilution steam in highly selective PyroCrack removal furnaces 1 PyroCrack furnaces are optimized with respect to residence Fueloil 4 time temperature and pressure profiles for the actual feedstock and the H required feedstock flexibility thus achieving the highest olefin yields Mixed Cs Borer CHa Ethylene Furnace effluent is cooled in transfer line exchangers 2 generating HP steam and by direct quenching with oil for liquid feedstocks The cracked gas stream is cooled and purified in the primary Pyrolysis ae ie Propane E Fthane fractionator 3 and quench water tower 5 Waste heat is recovered by gasoline Ct recycle recycle a circulating oil cycle generating dilution steam 4 and by a water cycle 5 to provide heat to reboilers and process heaters The cracked gas from the quench tower is compressed 6 in a 4 or 5stage compressor and dried in gas and liquid adsorbers 8 COz and H2S are removed ina causticwash system located before the final compressor stage hydrocarbon condensates from the hot section forms an aromatic The compressed cracked gas is further cooled 9 and fed to the sich gasoline product recovery section frontend deethanizer 10 isothermal frontend C hydrogenation 11 cold train 12 demethanizer 13 and the heat Economics Ethylene yields vary between 25 35 45 and 83 for pumped lowpressure ethylene fractionatior 14 which is integrated gas oils naphtha LPG and ethane respectively The related specific energy with the ethylene refrigeration cycle This wellproven Linde process is consumption range is 600054004600 and 3800 kcalkg ethylene highly optimized resulting in high flexibility easy operation low energy Typical installation costs for a worldscale ISBL gas naphtha cracker on a consumption low investment costs and long intervals between major Gulf Coast basis are 500 750 USton installed ethylene capacity turnarounds typically five years The C3 from the deethanizer bottoms 10 is depropanized 15 Commercial plants Over 15 million tons of ethylene are produced in hydrogenated 16 to remove methyl acetylene and propadiene 16 ore than 40 plants worldwide Many plants have been expanded in and fractionated to recover polymer grade propylene C components Capacity up to 50 and more are separated from heavier components in the debutanizer 18 to recover a Cy product and a Cs stream The Cs together with the Recent awards for worldscale ethylene plants include Borouge in Abu Dhabi Optimal in Malaysia Amir Kabir and Marun in Iran and TVK II in Hungary The Marun plant is one of the worlds largest crackers with a capacity of 11 million mtpy ethylene and 200000 mtpy propylene Licensor Linde AG Ethylene continued iste se cal PetrochemicalP eee C Oe Sich home processes index company index Ethylene Application To produce polymergrade ethylene and propylene by ther mally cracking paraffinic feedstocks ethane through hydrocracked resi Feed nl stock steam due Two main process technologies are used pal 1 USC ultra selective crackingPyrolysis and quench systems 4 2 ARSHRS advanced recovery system with heatintegrated rectifier ere simplification Cold fractionation Plants are characterized by high operational reliability rapid startups l Ethylene and ability to meet environmental requirements Propylene Description Feeds are sent to USC cracking furnaces 1 Contaminants Iq removal may be installed upstream A portion of the cracking heat may be supplied by gas turbine exhaust Pyrolysis occurs within the temperature time requirements specific to the feedstock and product requirements HMethane Rapid quenching preserves higholefin yield and the waste heat gener Ethane recycle Propane recycle ates highpressure steam Lowertemperature waste heat is recovered and pyrolysis fuel oil and gasoline distillate fractionated 2 Cracked gas C and lighter is then compressed 3 scrubbed with caustic to remove acid gases and dried prior to fractionation C and lighter components are separated from the Cy and heavier components in the lowfouling C3s are combined and hydrogenated to remove methyl acetylene frontend dual pressure depropanizer 4 Overhead vapor is hydroge and propadiene 10 Polymer or chemicalgrade propylene is then nated to remove acetylene 5 and is routed to the ARSHRS 6 produced overhead from the C3 superfractionator 11 ARS minimizes refrigeration energy by using distributed distillation C and heavier coproducts are further separated in a sequence and simultaneous heat and mass transfer in the dephlegmator exclusive of distillation steps Ethane and propane are typically recycle cracked arrangement with Air Products or HRS system Two C streams of varying Refrigeration is supplied by cascade ethylenepropylene systems composition are produced Hydrogen and methane are separated Specific advantages of ARS technology are 1 reduced chilling overhead train refrigeration requirements due to chillingprefractionation in the The heavier C stream is deethanized 7 and C overhead passes to dephlegmator or HRS system 2 reduced methane content in feed to the MP ethyleneethane fractionator 9 integrated with C refrigeration demethanizer 3 partial deethanizer bypassing 4 dual feed ethylene system The lighter C stream is routed directly to the ethyleneethane fractionator lower reflux ratio and 5 reduced refrigeration demand fractionator 9 Polymergrade ethylene product is sent overhead from approx 75 the ethyleneethane fractionator Acetylene recovery may optionally be installed upstream of the ethyleneethane fractionator 8 Economics Ethylene yields range from 57 ethane high conversion to 28 heavy hydrogenated gas oils Corresponding specific energy consumptions range from 3000 kcalkg to 6000 kcalkg Commercial plants Over 120 ethylene units have been built by Stone Webster Expansion techniques based on ARSHRS technology have increased original capacities by as much as 100 Licensor Stone Webster Inc a Shaw Group Co Ethylene continued iste se cal PetrochemicalProcesses home processes index company index Ethylene Application Thermal cracking of a wide range of feedstocks into light pretees olefins and aromatics using proprietary cracking coils Cracked gas nn oa Feedstocks Ethane through to heavy feeds up to 600C EP Products Cracked gas rich in ethylene propylene butadiene and BTX Feed Description Thermal cracking occurs in presence of steam at high tem en peratures in cracking coils located centrally in the firebox Coil outlet Bee temperatures vary up to 880C depending on feed quality and cracking Process steam severity The proprietary cracking coils are the GK5 GK6 and SMK coils They feature high selectivity to ethylene and propylene together with low coking rates long run lengths GKSMK Cracked gases from the furnace pass through a transferline Se exchanger TLE system where heat is recovered to generate high pressure steam The primary TLEs are linear or special S and T type exchangers The selected exchanger type ensures low to very low fouling rates and thus extends run lengths Heat from the flue gases is recovered in the convection section to preheat feed and process steam and to superheat generated HP Steam The technology may be applied to retrofit furnaces Furnace performance is optimized using proprietary SPYRO programs NO abatement technology is incorporated Performance data Ethane conversion 6575 Naphtha cracking severity as PE 040070 Overall thermal efficiency 9295 Coil residence time sec GK5GK6 coils 015025 SMK coil 035040 Oncethrough ethylene yields depend on feed characteristics and severity and range from 58 for ethane to 36 for liquid feeds Commercial plants Over 450 installations since the mid1960s Licensor Technip Ce eur Lice a PROCESSING PetrochemicalProcesses home processes index company index Ethylene Application To produce polymergrade ethylene and propylene a bu tadienerich C cut an aromatic CgCg richraw pyrolysis gasoline and I dryi highpurity hydrogen by using the TPAR process for gas separation and gas removal drying Cbroadcut product purification from raw cracked gas Compressors gs Deethani Acetylene hyd Acetylene Description Effluents from cracking furnaces are cooled and processed Feed core stripper acetylene recovery for tar and heavygasoline removal A multistage compressor driven by a steam turbine compresses the Hydrogenrich Demethanizer cooled gas LP and HP condensates are stripped in two separate strippers imlees stripper co Ret where medium gasoline is produced and part of the C3 cut is recovered iL respectively A caustic scrubber removes acid gases P 9 EthyleneEthane Compressed gas at 450 psig is dried and then chilled A multi Fractionation a on stream heat exchanger chills the tail gas to 265F Liquid condensates Propylene J Pygas are separated at various temperatures such as 30F 65F 100F and gasoline C cut rear Ethylene 140F and are reheated against incoming cracked gas The partially vaporized streams are sent to a deethanizer stripper operating at about 320 psig The bottoms C3 stream is sent to propylene and heavys recovery The overhead is reheated and enters an adiabatic acetylene Economics The advantages of this process are low equipment costs hydrogenation reactor which transforms the acetylene selectively to viz the deethanizer system and ethyleneethane separation and reli ethylene and ethane As an alternate a solventrecovery process canbe ability of the acetylene hydrogenation due to low excess hydrogen at applied without reheating the gas the reactor inlet The refrigeration compressor benefits from low specific Reactor effluent is chilled and lightends are separated from the Power and suction volume while the crackedgas compressor processes Chydrocarbons The demethanizer overhead is processed for ethylene aboveambienttemperature gas recovery while the bottoms IS sent to ethyleneethane separation An Commercial plants Technip is commercializing the TPAR process on a open heatpump splitter is applied thus sending ethylene product to the a casebycase basis gas pipeline from the discharge of the ethylenerefrigerant compressor Dilute ethylene for chemical applications such as styrene production Licensor Technip can be withdrawn downstream of the hydrogenation reactor The ethylene content Is typically 60 vol Catalyst suppliers have tested the hydrogenation step and commercially available frontend catalysts are suitable for this application Se OR eT Cm Cm CE a PROCESSING PetrochemicalProcesses miele IN ce a JCESSE home processes index company index Ethylene Application The MaxEne process increases the ethylene yield from naphtha crackers by raising the concentration of normal paraffins n Adsorbent Desorbent paraffins in the naphthacracker feed The MaxEne process is the new Rotary Extract est application of UOPs Sorbex technology The process uses adsorptive valve column separation to separate C5C naphtha into a rich nparaffins stream 5 and a stream depleted of nparaffins Normal paraffins to cracker Description The separation takes place in an adsorption chamber 2 lbesorbens that is divided into a number of beds Each bed contains proprietary Feed shapeselective adsorbent Also each bed in the chamber is connected panne Raffinate to a rotary valve 1 The rotary valve is used along with the shapese sinus column lective adsorbent to simulate a countercurrent moving bed adsorptive Nonnormal hydrocarbons separation Four streams are distributed by the rotary valve to and from cull hth to reformer for gasoline the adsorbent chamber The streams are as follows ee or aromatics production e Feed The naphtha feed contains a mixture of hydrocarbons e Extract This stream contains nparaffin and a liquid desorbent Naphtha rich in nparaffin is recovered by fractionation 3 and is sent to the naphtha cracker Raffinate This stream contains nonnormal paraffin and a liquid desorbent Naphtha depleted in nparaffin is recovered by fraction ation 4 and Is sent to a refinery or an aromatics complex Desorbent This stream contains a liquid desorbent that is recycled from the fractionation section to the chamber The rotary valve is used to periodically switch the position of the liquid feed and withdrawal points in the adsorbent chamber The process operates in a continuous mode at low temperatures in a liquid phase Economics Capital costs and economics depend on feed composition as well as the desired increase in ethylene and propylene production in the steam cracker Licensor UOP LLC PROCESSING PetrochemicalProcesses miele IN ce a home processes index company index Ethylene feed pretreatment mercury arsenic and lead removal Organometallic Arsenic and Application Upgrade natural gas condensate and other contaminated a ae caer br streams to highervalue ethylene plant feedstocks Mercury arsenic and lead contamination in potential ethylene plant feedstocks precludes their use despite attractive yield patterns The contaminants poison cat alysts cause corrosion in equipment and have undesirable environmen we tal implications For example mercury compounds poison hydrotreating a CMG CMG Mercury catalysts and if present in the steamcracker feed are distributed in ad 273 trap the CCs cuts A condensate containing mercury may have negative addedvalue as a gas field product Distilled Steam gs feedstock Description Three RAM processes are available to remove arsenic cw RAM 1 arsenic mercury and lead RAM Il and arsenic mercury and sulfur from liquid hydrocarbons RAM Ill Described above is the RAM Il process Feed is heated by exchange with reactor efflu ent and steam 1 It is then hydrolyzed in the first catalytic reactor 2 in which organometallic mercury compounds are converted to elemental mercury and organic arsenic compounds are converted to arsenicmetal complexes and trapped in the bed Lead if any is 500 ppb 10 ppb arsenic and 120 ppb lead excluding basic engineering also trapped on the bed The second reactor 3 contains a specific detailed engineering offsites contractor fees mercurytrapping mass There is no release of the contaminants to Clear oxygenfree Aerated condensate the environment and spent catalyst and trapping material can be condensate with particulate matter disposed of in an environmentally acceptable manner Investment USbpd 130 180 Utilities USbpd 008 023 Contaminant pica eeveedstoc Product Catalyst cost USbpd 003 003 Arconie tab 00 i Commercial plants Fifteen RAM units have been licensed worldwide ee ee en eens far References Dicillon B L Savary J Cosyns Q Debuisschert and P Trav of less than one ppb can be achieved ers Mercury and Arsenic Removal from Ethylene Plant Feedstocks Sec Economics The ISBL 2004 investment at a Gulf Coast location for two condensates each containing 50ppb average mercury content max Sa click here to email for more information a ond European Petrochemicals Technology Conference Prague 2000 Licensor Axens Axens NA Ethylene feed pretreatmentmercury arsenic and lead removal continued PROCESSING PetrochemicalProcesses miele IN ce a Mele erste home processes index company index Ethylene glycol mono MEG Application To produce monoethylene glycol MEG from ethylene ox ide EO Purge CO recycle Description EO in an aqueous solution is reacted with CO in the pres ence of a homogeneous catalyst to form ethylene carbonate 1 The Water Mc ethylene carbonate subsequently is reacted with water to form MEG 2 and CO 3 The net consumption of CO in the process is nil since all CO converted to ethylene carbonate is released again in the ethylene carbonate hydrolysis reaction Unconverted CO from the ethylene car bonate reaction is recovered 2 and recycled together with CO re es leased in the ethylene carbonate hydrolysis reaction The product from the hydrolysis reaction is distilled to remove CO residual water 4 In subsequent distillation columns highpurity MEG is Catalyst recycle recovered 5 and small amounts of coproduced diethylene glycol are eTUESCSES Residue removed 6 The homogeneous catalyst used in the process concentrates in the bottom of column 5 and is recycled back to the reaction section The process has a MEG yield of 99 Compared to the thermal glycol process steam consumption and wastewater production are relatively low the latter because no contaminated process steam is generated MEG quality and performance of the MEG product in derivatives polyesters manufacturing have been demonstrated to be at least as good as and fully compatible with MEG produced via the thermal process Commercial plants The first commercial plant is currently under con struction in Taiwan Two other process licenses have been awarded The combination of this process with the Shell EO process is licensed under the name Shell OMEGA process Licensor Shell International Chemicals BV Contact ctamsterdamshellcom PROCESSING PetrochemicalProcesses PROCESSING Meiaern nT Lefe Ue AOL ess SsieiS home processesindex company index Ethylene glycol Application To produce ethylene glycols MEG DEG TEG from ethyl ene oxide EO using Dows Meteor process Recycled 1 Steam water MEG Description In the Meteor Process an EOwater mixture is preheated and fed directly to an adiabatic reactor 1 which can operate with or EOiwater without a catalyst An excess of water is provided to achieve high selec tivities to monoethylene glycol MEG Diethylene DEG and triethylene TEG glycols are produced as coproducts In a catalyzed mode higher selectivities to MEG can be obtained thereby reducing DEG production to onehalf that produced in the uncatalyzed mode The reactor is spe Steam cially designed to fully react all of the EO and to minimize backmixing ic Steam which promotes enhanced selectivity to MEG Excess water from the reactor effluent is efficiently removed in a DEGTEG multieffect evaporation system 2 The lasteffect evaporator overhead produces lowpressure steam which is a good lowlevel energy source for other chemical units or other parts of the EOMEG process The concentrated waterglycols stream from the evaporation system is fed to the water column 3 where the remaining water and light ends are stripped from the crude glycols The waterfree crude glycol stream is fed to the MEG refining column 3 where polyestergrade MEG suitable for Commercial plants Since 1954 18 UCCdesigned glycol plants have polyester fiber and PET production is recovered DEG and TEG exiting been started up or are under construction the base of the MEG refining column can be recovered as highpurity products by subsequent fractionation Licensor Union Carbide Corp a subsidiary of The Dow Chemical Co Economics The conversion of EO to glycols is essentially complete The reaction not only generates the desired MEG but also produces DEG and TEG that can be recovered as coproducts The production of more DEG and TEG may be desirable if the manufacturer has a specific use for these products or if market conditions provide a good price for DEG and TEG relative to MEG A catalyzed process will produce less heavy glycols The ability to operate in catalyzed or uncatalyzed mode provides flex ibility to the manufacturer to meet changing market demands i eee eer Ethylene glycols Application To produce ethylene glycols MEG DEG and TEG from eth ylene oxide EO Water ee Description Purified EO or a waterEO mixture is combined with re cycle water and heated to reaction conditions In the tubular reactor Steam Water MEG DEG EG 1 essentially all EO is thermally converted into monoethylene glycol MEG with diethylene glycol DEG and triethylene glycol TEG as co products in minor amounts Excess water required to achieve a high selectivity to MEG is evaporated in a multistage evaporator 2 3 4 The last evaporator produces lowpressure steam that is used as a heat ing medium at various locations in the plant The resulting crude glycols mixture is subsequently purified and fractionated in a series of vacuum columns 5 6 7 8 The selectivity to MEG can be influenced by adjusting the glycol reactor feed composition Most MEG plants are integrated with EO plants In such an integrated EOMEG facility the steam system can be optimized to fully exploit the benefits of highselectivity catalyst applied in the EO plant However standalone MEG plants have been designed and built The quality of glycols manufactured by this process ranks amongst the highest in the world It consistently meets the most stringent specifica tions of polyester fiber and PET producers Commercial plants Since 1958 more than 60 Shelldesigned MEG plants have been commissioned or are under construction Licensor Shell International Chemicals BV The combination of this process with the Shell EO process is licensed under the name Shell MASTER process a COC Cee ey a eee eer Ethylene oxide Application To produce ethylene oxide EO from ethylene using oxygen compres as the oxidizing agent Description Ethylene and oxygen in a diluent gas made up of a mixture of mainly methane or nitrogen along with carbon dioxide and argon are fed to a tubular catalytic reactor 1 The temperature of reaction is Ethylene controlled by adjusting the pressure of the steam which is generated in 0 compressor the shell side of the reactor and removes the heat of reaction The EO cam y A produced is removed from the reaction gas by scrubbing with water 2 Purified after heat exchange with the circulating reactor feed gas C Byproduct CO is removed from the scrubbed reaction gas 3 4 a 3 steam before it is recompressed and returned to the reaction system where ethylene and oxygen concentrations are restored before returning to mn the EO reactor 550 product The EO is steam stripped 5 from the scrubbing solution and re covered as a more concentrated water solution 6 for feed to an EO purification system 7 8 where purified product is made along with a high aldehyde EO product Product quality The EO product meets the low aldehyde specification of 10 ppm maximum which is required for EO derivatives production Product yield The ethylene yield to purified EO is 12 kg per kg ethylene feed In addition a significant amount of technicalgrade glycol may be recovered by processing waste streams Commercial plants Nearly 50 purified EO projects have been completed or are being designed This represents a total design capacity of about 4 million metric tons of purified EO with the largest plants exceeding 200000 mtpy Licensor Scientific Design Company Inc a COC Cee ey j tistesstttcial PetrochemicalProcesses miele IN ce a home processes index company index Ethylene oxide Application To produce ethylene oxide EO from ethylene and oxygen in a direct oxidation process Ethylene Description In the direct oxidation process ethylene and oxygen are Oxygen mixed with recycle gas and passed through a multitubular catalytic re a actor 1 to selectively produce EO A special silvercontaining highse Steam lectivity catalyst is used that has been improved significantly over the years Methane is used as ballast gas Heat generated by the reaction x D is recovered by boiling water at elevated pressure on the reactors shell side the resulting highpressure steam is used for heating purposes at ee various locations within the process aqueous EO EO contained in the reactor productgas is absorbed in water 2 and further concentrated in a stripper 3 Small amounts of coabsorbed r ethylene and methane are recovered from the crude EO 4 and recycled Steam back to the EO reactor The crude EO can be further concentrated into highpurity EO 5 or routed to the glycols plant as EOwater feed EO reactor productgas after EO recovery is mixed with fresh feed and returned to the EO reactor Part of the recycle gas is passed through an activated carbonate solution 6 7 to recover COz a byproduct of the Commercial plants Since 1958 more than 60 Shelldesigned plants EO reaction that has various commercial applications have been commissioned or are under construction Approximately 40 Most EO plants are integrated with fibergrade monoethylene of the global capacity of EO equivalents is produced in Shelldesigned glycol MEG production facilities In such an integrated EOMEG facility plants the steam system can be optimized to fully exploit the benefits of high selectivity catalyst Licensor Shell International Chemicals BV When only highpurity EO is required as a product a small amount The Shell EO process is licensed under the name Shell MASTER pro of technicalgrade MEG inevitably is coproduced cess when combined with the Shell ethylene glycols process and under the name Shell OMEGA process when combined with the Shell process Yields Modern plants are typically designed for and operate ata molar for selective MEG production via ethylene carbonate intermediate EO catalyst selectivity approaching 90 with fresh catalyst and 8687 as an average over 3 years catalyst life resulting in an average EO pro duction of about 14 tons per ton of ethylene However the technol ogy is flexible and the plant can be designed tailormade to customer i requirements or different operating time between catalyst changes PROCESSING PetrochemicalProcesses miele IN ce a home processes index company index Ethylene oxide Application To produce ethylene oxide EO from the direct oxidation of Carbon ethylene using the Dow Meteor process Ethylene dioxide Description The Meteor Process a technology first commercialized in water Oxygen fi 1994 is a simpler safer process for the production of EO having lower Steam capital investment requirements and lower operating costs In the Meteor Steam JG Process ethylene and oxygen are mixed with methaneballast recycle gas and passed through a singletrain multitubular catalytic reactor 1 to selec 2 Ethylene tively produce EO Use of a single reactor is one example of how the Meteor oxide process is a simpler safer technology with lower facility investment costs The special highproductivity Meteor EO catalyst provides very a high efficiencies while operating at high loadings Heat generated by the reaction is removed and recovered by the direct boiling of water to aa generate steam on the shell side of the reactor Heat is recovered from the reactor outlet gas before it enters the EO absorber 2 where EO is scrubbed from the gas by water The EOcontaining water from the EO absorber is concentrated by stripping 3 The cycle gas exiting the absorber is fed to the CO removal section 45 where CO which is coproduced in the EO reactor is removed via activated hot potassium carbonate treatment The CO lean cycle gas is recycled by compression Commercial plants Union Carbide was the first to commercialize the back to the EO reactor direct oxidation process for EO in the 1930s Since 1954 18 Union Car Most EO plants are integrated with glycol production facilities bidedesigned plants have been started up or are under construction When producing glycols the EO stream 3 is suitable for feeding directly Three million tons of EO equivalents per year approximately 20 of to a Meteor glycol process When EO is the desired final product the total world capacity are produced in Union Carbidedesigned plants EO stream 3 can be fed to a single purification column to produce highpurity EO This process is extremely flexible and can provide the Licensor Union Carbide Corp a subsidiary of The Dow Chemical Co full range of product mix between glycols and purified EO Economics The process requires a lower capital investment and has lower fixed costs due to process simplicity and the need for fewer equip ment items Lower operating costs are also achieved through the high Pion losdings EO catalyst which has very high efficiencies at very i PROCESSING PetrochemicalProcesses miele IN ce at home sprocesses index company index Ethylene oxideEthylene glycols Application To produce ethylene glycols EGs and ethylene oxide EO OS en from ethylene using oxygen as the oxidizing agent purified Modern EOEG plants are highly integrated units where EO pro LC EO duced in the EO reaction system can be recovered as glycols MEG DEG l and TEG with a coproduct of purified EO if desired Process integra q tion allows for a significant utilities savings as well as the recovery of all Ethylene Recycle I bleed streams as highgrade product which would otherwise have been 0 recovered as a lesser grade product The integrated plant recovers all Steam I MEG as fibergrade product and EO product as lowaldehyde product Lod The total recovery of the EO from the reaction system is 997 with only a small loss as heavy glycol residue ih Description Ethylene and oxygen in a diluent gas made up of a mixture Sau of mainly methane or nitrogen along with carbon dioxide CO and ar Jc gon are fed to a tubular catalytic reactor 1 The temperature of reaction steam is controlled by adjusting the pressure of the steam which is generated in the shell side of the reactor and removes the heat of reaction The EO Ethylene oxide section produced is removed from the reaction gas by scrubbing with water 2 Glycolsection i ststidSCSdYSC Ce TCU after heat exchange with the circulating reactor feed gas Byproduct CO is removed from the scrubbed reaction gas 3 4 before it is recompressed and returned to the reaction system where bel bef Wese 2 ethylene and oxygen concentrations are restored before returning to the EO reactor E steam The EO is steam stripped 5 from the scrubbing solution and recov ered as a more concentrated water solution 6 that is suitable for use as a feed to a glycol plant 8 or to an EO purification system 7 The stripped he Recycle water water solution is cooled and returned to the scrubber The glycol plant feed along with any high aldehyde EO bleeds from the EO purification section are sent to the glycol reactor 9 and then to a multieffect evaporation train 10 11 12 for removal of the bulk of the water from the glycols The glycols are then dried 13 and sent to the glycol distillation train 14 15 16 where the MEG DEG and TEG h products are recovered and purified Product quality The SD process has set the industry standard for fiber grade MEG quality When EO is produced as a coproduct it meets the low aldehyde specification requirement of 10ppm aldehyde maximum which is required for EO derivative units Yield The ethylene yield to glycols is 181 kg of total glycols per kg of ethylene The ethylene yield for that portion of the production going to purified EO is 131 kg of EO product kg of ethylene Commercial plants Over 90 EOEG plants using SD technology have been built The worlds largest MEG plant with a capacity of 700000 mtpy of MEG is presently in design and follows the startup of a 600000 mtpy plant in October 2004 Licensor Scientific Design Company Inc Ethylene oxideEthylene glycols continued tistesstttcial PetrochemicalProcesses miele IN ce a home processes index company index Formaldehyde Application To produce aqueous formaldehyde AF or urea formal dehyde precondensate UFC from methanol using the Hador Topsge Tail gas Formaldehyde SR process comprising two reactors in series Recycle Description Air and recycle gas are compressed by the blower 1 and fe 10 then mixed with liquid methanol that is injected through spray nozzles CO The mixture is preheated to about 200C by heat exchange with hot cir Air culating oil in the heat exchanger 2 after which the gas is successively S 26 9 passed to the two series reactors 3 and 4 BFW bi eeeire Additional methanol is injected into the gas between the two reactors The reactors contain many tubes filled with FK2 catalyst Methanol eS at where methanol and oxygen react to make formaldehyde Reaction heat is removed by a bath of boiling heattransfer oil Hot oil vapor is condensed in the wasteheat boiler 5 thus generating steam at up to 40 bar pressure Before entering the absorber 7 the reacted gas is cooled in the after cooler 6 and reheats the circulating oil from the processgas heater 2 In the absorber the formaldehyde is absorbed in water or urea solution Heat is removed by one or two cooling circuits 8 9 From the lower circuit 8 product in the form of either AF or UFC is withdrawn Scrubbed gas from the absorber is split in two streamsrecycle gas and Longer catalyst life 3036 months in reactor 18 months in tail gas The tail gas is vented after any organic impurities are catalytically eactor ll 7 incinerated in the reactor 10 Thus the tailgas purity conforms to the Lower electricity consumption and higher steam production environmental standards for any country e Higher conversion of methanol therefore less methanol in With regard for the catalyst the percentage of methanol that can Product be added to a formaldehyde reactor is limited to about 9vol Using The Haldor Topsee Formaldehyde SR process is wellsuited to two reactors in series higher production yields are achievable with the xPand existing formaldehyde plantsup to 100 capacity increase same gas flow than what would be possible in a plant with only one ay be achieved reactor or a plant with two reactors In parallel Utility requirements Per 1000 kg of 37wt formaldehyde Advantages of series reactors vs single or parallel reactors are e Lower capital cost due to reduced size of equipment and piping Product 55 wt AF 85 wt UFC Methanol kg 420 425 425 430 70 urea solution kg 220 Process water kg 250 72 Water cooling m3 42 38 Electricity kWh 49 52 Commercial plants Three commercial SR units built all are operating successfully Three additional units are under construction Licensor Haldor Topsøe AS Formaldehyde continued PROCESSING PetrochemicalProcesses home processes index company index Formaldehyde Application Formaldehyde as a liquid solution of 3752 wt is primar ae Offgas ily used in the production of polyoxymethylene POM and hexamethy Steam Absorption water 4 lenetetramine as well as synthetic resins in the wood industry ae solution Steam 7 ep Description Formaldehyde solutions are produced by oxidation with NX A Steam 7 1 Cc XA methanol in the air In the UIF process the reaction occurs on the sur WV face of a silvercrystal catalyst at temperatures of 620C680C where hl rows X 0 wer the methanol is dehydrated and partly oxidized Or e Oo Cc NN CH30H CH0 H Ah 84 kJmol ti x O i NZ wer CH30H 2 O7 CH20 H20 Ah 159 kJmol EHO OQ LN formaldehyde or Methanol ureaformaldehyde The methanolwater mixture adjusted for density balance and stored NZ precondensate in the preparation tank is continuously fed by pump to the methanol LUN evaporator 1 The required process air is sucked in by a blower via a filter and air scrubber into the methanol evaporator From here the methanolwaterair mixture enters the reactor 2 where the conversion of methanol to formaldehyde occurs Because the reaction s exothermic the required temperature is selfmaintained To produce ureaformaldehyde precondensate an aqueous urea once the ignition has been executed solution in place of absorption water is fed into the absorption tower The reaction gases emerging from catalysis contain formaldehyde water nitrogen hydrogen and carbon dioxide as well as nonconverted methanol Economics Due to the wastegas recycling system the methanol content They are cooled to 150C in a wasteheat boiler directly connected to the in the formaldehyde solution can be reduced to less than 1 wt and reactor The amount of heat released in the boiler is sufficient for heating formic acid less than 90 ppm the methanol evaporator The reaction gases enter a 4stage absorption Typical consumption figures per 1000 kg of formaldehyde solution tower 3 where absorption of formaldehyde occurs in counterflow via 37 wt are aqueous formaldehyde solution and cold demineralized water The final Methanol kg 445 formaldehyde solution is removed from the first absorption stage Water kg 390 Waste gas from the absorption tower with a heating value of ieee kWh 38 3 ater cooling m 40 approximately 2000 kJm is burned in a post connected thermal combustion unit The released heat can be used to produce highpressure Licensor Uhde InventaFischer steam or thermal oil heating By recycling a part of waste gas to the reactor formaldehyde i concentrations up to 52 wt in the final solution can be reached ri el PROCESSING PetrochemicalProcesses miele IN ce a f home processes index company index Hydrogen Recycle H Process steam Application Production of hydrogen H from hydrocarbon HC feed rtineryoffgeses Coneaion stocks by steam reforming Pumping Prereformer Makeup fuel PSA purge gas optional Feedstocks Ranging from natural gas to heavy naphtha as well as po Export steam tential refinery offgases Many recent refinery hydrogen plants have multiple feedstock flexibility either in terms of backup or alternative PEN or mixed feed Automatic feedstock changeover has also successfully CSecton Reformer WoL WPS steam Vent steam been applied by Technip in several modern plants with multiple feed pons en we cng stock flexibility oi BFW a convertion Purge gas 0 Description The generic flowsheet consists of feed pretreatment pre O ae reforming optional steamHC reforming shift conversion and hydro Flue gas sy To steam gen purification by pressure swing adsorption PSA However it is often aa tailored to satisfy specific requirements DMw Recycle Hy Feed pretreatment normally involves removal of sulfur chlorine and ee Aearacen other catalyst poisons after preheating to 350400C steam system products The treated feed gas mixed with process steam is reformed in a fired reformer with adiadatic prereformer upstream if used after necessary superheating The net reforming reactions are strongly endothermic Heat is supplied by combusting PSA purge gas supplemented by make criteria and steam export requirements Recent advances include inte up fuel in multiple burners in a topfired furnace gration of hydrogen recovery and generation and recuperative post Reforming severity is optimized for each specific case Waste heat reforming also for capacity retrofits from reformed gas is recovered through steam generation before the watergas shift conversion Most of the carbon monoxide CO is further Commercial plants Technip has been involved in over 240 hydrogen converted to hydrogen Process condensate resulting from heat recovery plants worldwide and cooling is separated and generally reused in the steam system after Licensor Techni necessary treatment The entire steam generation is usually on natural Pp circulation which adds to higher reliability The gas flows to the PSA unit that provides highpurity hydrogen product up to 1 ppm CO at near inlet pressures Typical specific energy consumption based on feed fuel export steam ranges between 3 GcalKNm and 35 GcalKNm 330370 Btu scf LHV depending upon the feedstock plant capacity optimization PROCESSING PetrochemicalProcesses miele IN ce a Mele erste home processes index company index Maleic anhydride Application To produce maleic anhydride from nbutane using a fluid Tail cas to fuel use bed reactor system and an organic solvent for continuous anhydrous uP steam or Sore product recovery steam generation Light ends Description Nbutane and air are fed to a fluidbed catalytic reactor 1 LJ a to produce maleic anhydride The fluidbed reactor eliminates hot spots and permits operation at close to the stoichiometric reaction mixture sw J Pure mle This results in a greatly reduced air rate relative to fixedbed processes 1 and translates into savings in investment and compressor power and vs J large increases in steam generation The fluidbed system permits online catalyst additionremoval to adjust catalyst activity and reduces down Butane iL Js Crude maleic time for catalyst change out anmyeride to The recovery area uses a patented organic solvent to remove the ia maleic anhydride from the reactor effluent gas Aconventional absorption Ar Ce Heavy byproducts 2stripping 3 scheme operates on a continuous basis Crude maleic anhydride is distilled to separate light 4 and heavy 5 impurities A slipstream of recycle solvent is treated to eliminate any heavy byproducts that may be formed The continuous nonaqueous product recovery system results in superior product quality and large savings in steam consumption It also reduces investment product degradation loss and byproduct formation and wastewater Economics The ALMA process produces highquality product with at tractive economics The fluidbed process is especially suited for large singletrain plants Commercial plants Nine commercial plants have been licensed with a to tal capacity of 200000 mtpy The largest commercial installation is Lonzas 55000mtpy plant in Ravenna Italy Second generation process optimiza tions and catalyst have elevated the plant performances since 1998 Licensor ABB Lummus GlobalLonza Group iste se cal PetrochemicalPro eee C Oe Sich home processes index company index Methanolsteammethane reforming Application To produce methanol from natural or associated gas feed CO stocks using advanced tubular reforming followed by boiling water reac Steam SOtstCSstStSSY tor synthesis This technology is an option for capacities up to approxi mately 3000 mtpd methanol for cases where carbon dioxide CO is eee available Topsge also offers technology for largerscale methanol facili reactor ties up to 10000 mtpd per production train and technology to modify Sulfur Sulfur y Steam pu ammonia capacity into methanol production te aor H d ir or Steam reformer H Makeup 7 Description The gas feedstock is compressed if required desulfurized reformer compressor 1 and process steam is added Process steam used is a combination of Natural gas 3 D steam from the process condensate stripper and superheated medium Condensate Pf pressure steam from the header The mixture of natural gas and steam Product methanol SS raw is preheated prereformed 2 and sent to the tubular reformer 3 The HO ao methanol Raw prereformer uses waste heat from the fluegas section of the tubular water a ab Sorace reformer for the reforming reaction thus reducing the total load on the tubular reformer Due to high outlet temperature exit gas from the tubular reformer has a low concentration of methane which is an in ert in the synthesis The synthesis gas obtainable with this technology typically contains surplus hydrogen which will be used as fuel in the reformer furnace If CO is available the synthesis gas composition can reactor feed cools effluent from the synthesis reactor Further cooling is be adjusted hereby minimizing the hydrogen surplus Carbon dioxide obtained by air or water cooling Raw methanol is separated and sent can preferably be added downstream of the prereformer directly to the distillation section 5 featuring a very efficient three The flue gas generated in the tubular reformer is used for preheat column layout Recycle gas is sent to the recirculator compressor 8 of reformer and prereformer feed natural gas preheat steam superheat after a small purge to remove inert compound buildup and preheat of combustion air The synthesis gas generated in the tubular Topse supplies a complete range of catalysts for methanol reformer is cooled by highpressure steam generation 4 preheat of production The total energy consumption for this process scheme boiler feed water and reboiling in the distillation section 5 is about 72 Gcalton methanol without CO addition With CO After final cooling by air or cooling water the synthesis gas is addition the total energy consumption can be reduced to 70 Gcalton compressed 6 and sent to the synthesis loop 7 The synthesis loop methanol is comprised of a straighttubed boiling water reactor which is more efficient than adiabatic reactors Reaction heat is removed from the reactor by generating MP steam This steam is used for stripping of h process condensate and thereafter as process steam Preheating the Economics Tubular reforming technology is attractive at capacities 2500 3000 mtpd methanol where the economy of scale of alterna tive technologies such as twostep or autothermal reforming cannot be fully utilized Commercial plants The most recent plant is a 3030mtpd methanol facility with CO2 import The plant was commissioned in 2004 Licensor Haldor Topsøe AS Methanol steammethane reforming continued Meee PROCESSING PetrochemicalProcesses home processes index company index Methanolautothermal reforming AT R Oxygensteam Application To produce methanol from natural or associated gas feed Natural gas Saturator Steam stocks using autothermal reforming ATR followed by boiling water re a Methanol actor synthesis This technology is well suited for very largescale plants C as well as for the production of methanol to olefins or fuelgrade meth Steam Cl anol Topsge also offers technology for smaller methanol facilities and 7 LL technology to modify ammonia capacity into methanol production Hydro Sulfur Pre ee genator removal reformer 7 Description The gas feedstock is compressed if required desulfurized 5 1 and sent to a saturator 2 where the natural gas is saturated with eas Condensate 6 process condensate and excess water from the distillation section Re oe Lh Off gas cycling of process condensate and excess water minimizes the water re quirement Lowgrade mediumpressure steam is used in the saturator ee Raw methanol thus saving highpressure steam The mixture of natural gas and steam is preheated prereformed 3 and sent to the autothermal reformer 4 Autothermal reforming features a standalone oxygenfired reformer and thus the costintensive primary tubular reformer may be omitted completely The autothermal reformer can operate at any pressure The Operating pressure Is normally selected between 30 and 40 kgcrng reactor feed cools effluent from the synthesis reactor Further cooling synthesis gas generated in the autothermal reformer is cooled is by air or water cooling Raw methanol is separated and sent directly by highpressure steam generation 5 preheat of boiler feed water to the distillation section featuring a very efficient threecolumn layout reboiling in the distillation section and preheat of demineralized water Recycle gas is sent to the recirculator compressor 9 after a purge The synthesis gas obtainable with this technology is typically deficient to remove inert compound buildup The purge is sent to a hydrogen mn hydrogen Therefore the synthesis gas composition must be adjusted recovery unit where hydrogen is separated and recycled to the synthesis by recycling recovered hydrogen 6 from the synthesis loop After final gas compressor cooling by air or cooling water the recycle hydrogen is added to the Topsge supplies a complete range of catalysts for methanol produc synthesis gas which is compressed in a singlestep compressor 7 and tion The total energy consumption for this process scheme is about 71 aaa cunitiess loon ccmmprised of a straighttubed boiling water Gcalton methanol Total energy consumption for production of fuel grade reactor which is more efficient than adiabatic reactors Reaction heat methanol is approximately 68 Gcalton methanol is removed from the reactor by generation of mediumpressure steam ars iif i ti i This steam is used for heating in the saturator 2 Preheating the Economics For largescale plants the total investment including an oxygen plant is approximately 10 lower than for a conventional plant based on tubular steam reforming Licensor Haldor Topsøe AS Methanol autothermal reforming ATR continued iste se cal PetrochemicalP eee C Oe Sich home processes index company index Methanoltwostep reforming Application To produce methanol from natural or associated gas feed Steam Saturator Oxygen stocks using twostep reforming followed by lowpressure synthesis This technology is well suited for worldscale plants Topsge also offers Pre Geer Steam Methanol technology for smaller as well as very large methanol facilities up to reformer reformer rear 10000 tpd and technology to modify ammonia capacity into methanol Hydro Sulfur 4 Steam TOL production genator removal re eo Description The gas feedstock is compressed if required desulfurized wow i i seat 1 and sent to a saturator 2 where process steam is generated All Natural gas 3 5 process condensate is reused in the saturator resulting in a lower water Condensate BS requirement The mixture of natural gas and steam is preheated and Product methanol Light ends to fuel sent to the primary reformer 3 Exit gas from the primary reformer goes Raw Raw directly to an oxygenblown ndary reformer 4 The oxygen amount SB Brethanol methanol y to an oxygenblown secondary reforme yg 4 FH i and the balance between primary and secondary reformer are adjusted water 7 7 16 Storage so that an almost stoichiometric synthesis gas with a low inert content is obtained The primary reformer is relatively small and the reforming section operates at about 35 kgcm2g The flue gas heat content preheats reformer feed Likewise the heat content of the process gas is used to produce superheated highpressure steam 5 boiler feedwater preheating preheating process condensate about 70 Gcalton including energy for oxygen production going to the saturator and reboiling in the distillation section 6 E After final cooling by air or cooling water the synthesis gas is conomics Total investments including an oxygen plant are approxi compressed in a onestage compressor 7 and sent to the synthesis loop mately 1 0 lower for large plants than for a conventional plant based 8 comprised of three adiabatic reactors with heat exchangers between on straight steam reforming the reactors Reaction heat from the loop is used to heat saturator water Commercial plants The most recent largescale plant is a 3030tpd fa Steam provides additional heat for the saturator system Effluent from ility in Iran This plant was commissioned in 2004 the last reactor is cooled by preheating feed to the first reactor by air or water cooling Raw methanol is separated and sent directly to the Licensor Haldor Topsge AS distillation 6 featuring a very efficient threecolumn layout Recycle gas is sent to the recirculator compressor 9 after a small purge to remove inert compound buildup Topsge supplies a complete range of catalysts that can be used in the methanol plant Total energy consumption for this process scheme is iste se cal PetrochemicalP eee C Oe Sich home processes index company index Methanol Application To produce methanol in a singletrain plant from natural Fired Oxygen gas or oilassociated gas with capacities up to 10000 mtpd It is also heater Suu well suited to increase capacities of existing steamreformingbased 5 methanol plants 8 Y 4 Description Natural gas is preheated and desulfurized After desulfur As 6 x A mroces ization the gas is saturated with a mixture of preheated process water 1 condensate from the distillation section and process condensate in the saturator Fuel refoener toner uP steam to The gas is further preheated and mixed with steam as required for LL S oxygen plant the prereforming process In the prereformer the gas is converted to Natural I H CO and CHy Final preheating of the gas is achieved in the fired gas 4 LLP steam heater In the autothermal reformer the gas is reformed with steam wae Gas rv II and O The product gas contains H CO CO and a small amount of aor reactor ae Pressure He unconverted CH and inerts together with undercomposed steam The methanol BFW reformed gas leaving the autothermal reformer represents a consider able amount of heat which is recovered as HP steam for preheating energy and energy for providing heat for the reboilers in the distilla tion section The reformed gas is mixed with hydrogen from the pressure swing adsorption PSA unit to adjust the synthesis gas composition Synthesis the optimum reaction route The reactor outlet gas is cooled to about gas is pressurized to 510 MPa by a singlecasing synthesis gas 40C to separate methanol and water from the gases by preheating compressor and is mixed with recycle gas from the synthesis loop This BFW and recycle gas Condensed raw methanol is separated from the gas mixture is preheated in the trim heater in the gascooled methanol unreacted gas and routed to the distillation unit The major portion reactor In the Lurgi watercooled methanol reactor the catalyst is fixed Of the gas is recycled back to the synthesis reactors to achieve a high in vertical tubes surrounded by boiling water The reaction occurs under Overall conversion The excellent performance of the Lurgi combined almost isothermal condition which ensures a high conversion and converter LCC methanol synthesis reduces the recycle ratio to about eliminates the danger of catalyst damage from excessive temperature 2 A small portion of the recycle gas is withdrawn as purge gas to Exact reaction temperature control is done by pressure control of the lessen inerts accumulation in the loop steam drum generating HP steam In the energysaving threecolumn distillation section lowboiling The preconverted gas is routed to the shell side of the gas and highboiling byproducts are removed Pure methanol is routed to cooled methanol reactor which is filled with catalyst The final conversion to methanol is achieved at reduced temperatures along the tank farm and the process water is preheated in the fired heater and used as makeup water for the saturator Economics Energy consumption for a standalone plant including utili ties and oxygen plant is about 30 GJmetric ton of methanol Total in stalled cost for a 5000mtpd plant including utilities and oxygen plant is about US350 million depending on location Commercial plants Thirtyfive methanol plants have been built using Lurgis LowPressure methanol technology One MegaMethanol plant is in operation two are under construction and three MegaMethanol con tracts have been awarded with capacities up to 6750 mtpd of metha nol Licensor Lurgi AG Methanol continued tistesstttcial PetrochemicalProcesses PROCESSING home processes index company index Methanol Application The One Synergy process is improved lowpressure metha HP steam nol process to produce methanol The new method produces metha nol from natural or associated gas using twostage steam reforming an 1 Methanol product followed by compression synthesis and distillation Capacities ranging mi I a F from 5000 to 7000 mtpd are practical in a single stream Carbon di Natural 7 ll ie Hic8 oxide CO can be used as a supplementary feedstock to adjust the gas Hh fei 4 TT stoichiometric ratio of the synthesis gas CO BFW Description Gas feedstock is compressed if required desulfurized Water from 2 Pi Steam C C 1 and sent to the optional saturator 2 where some process steam is distillation AY C 8 7 generated The saturator is used where maximum water recovery is im 9 a portant Further process steam is added and the mixture is preheated O oy reformer and sent to the prereformer 3 using the CatalyticRichGas process 0 cnide Steam raised in the methanol converter is added along with avail CO optional methanol able CO and the partially reformed mixture is preheated and sent to the reformer 4 Highgrade heat in the reformed gas is recovered as highpressure steam 5 boiler feedwater preheat and for reboil heat in the distillation system 6 The highpressure steam is used to drive the main compressors in the plant After final cooling the synthesis gas is compressed 7 and sent to compounds These impurities are removed in a twocolumn distillation the synthesis loop The loop can operate at pressures between 70 to 100 system 6 The first column removes the light ends such as ethers esters bar The converter design does impact the loop pressure with radialflow acetone and dissolved noncondensable gases The second column designs enabling low loop pressure even at the largest plant size Low removes water higher alcohols and similar organic heavy ends oo oe unthosis loop comarees e crculator 3 oo rss Economics Recent trends have been to build methanol plants in re operates around 200C to 270C depending on the converter type gions offering lowcost gas such as Chile Trinidad and the Arabian Reaction heat from the loop is recovered as steam and is used directly Gulf In these regions total economics favor low investment rather than as process steam for the reformer lowenergy consumption Recent plants have an energy efficiency of A 7278 Gcalton A guideline figure to construct a 5000mtpd plant is purge is taken from the synthesis loop to remove inerts nitrogen US370400 million methane as well as surplus hydrogen associated with nonstoichiometric operation The purge is used as fuel for the reformer Crude methanol trom the separator contains water as well as traces of ethanol and other Commercial plants Thirteen plants with capacities ranging from 2000 to 3000 mtpd as well as 50 smaller plants have been built using the Synetix LPM methanol technology Two 5000mtpd plants are under construction Licensor One Synergy a consortium of Davy Process Technology John son Matthey Catalysts and Aker Kvaerner Methanol continued PROCESSING PetrochemicalProcesses PROCESSING home processes index company index Methanol Application To produce FederalGrade AA refined methanol from natu ce ral gasbased synthesis gas and naphtha using Toyo Engineering Corps TECs Synthesis Gas Generation technologies and proprietary MRFZ re 0 aT lhl actor incorporated in the Johnson Mattheys JMs process In a natural A Me gasbased plant the synthesis gas is produced by reforming natural gas XxX with steam andor oxygen using highactivity steam reforming ISOP x Tr ro ot catalyst X a 10 Crude methanol Description Syngas preparation section The feedstock is first preheated Nae A and sulfur compounds are removed in a desulfurizer 1 Steam is add NX ed and the feedstocksteam mixture is preheated again A part of the Steam av ae Methanol feed is reformed adiabatically in prereformer 2 The half of feedstock steam mixture is distributed into catalyst tubes of the steam reformer Fusel oil 3 and the rest is sent to TECs proprietary heat exchanger reformer TAFX 4 installed in parallel with 3 as the primary reforming The Process water heat required for TAFX is supplied by the effluent stream of secondary reformer 5 Depending on plant capacity the TAFX 4 andor the sec ondary reformer 5 can be eliminated Methanol purification section The crude methanol is fed to a twocolumn Methanol synthesis section The synthesis loop is comprised of a circula gictillation system which consists of a small topping column 11 and a tor combined with compressor 6 MRF2 reactor 7 feedeffluent refining column 12 to obtain highpurity Federal Grade AA methanol heat exchanger 8 methanol condenser 9 and separator 10 Cur rently MRFZ reactor is the only reactor in the world capable of produc Economics In typical natural gas applications approximately 30 GJ ing 50006000 td methanol in a singlereactor vessel The opera tonmethanol including utilities is required tion pressure is 510 MPa The syngas enters the MRFZ reactor 7 at 220240C and leaves at 260280C normally Installations Toyo has accumulated experience with the licensing of 20 JM proprietary methanol synthesis catalyst is packed in the shell side methanol plant projects of the reactor Reaction heat is recovered and used to efficiently gener Reference US Patent 6100303 ate steam in the tube side Reactor effluent gas is cooled to condense the crude methanol The crude methanol is separated in a separator Licensor Toyo Engineering Corp TECJohnson Matthey PLC 10 The unreacted gas is circulated for further conversion A purge is taken from the recycling gas used as fuels in the reformer 3 i PROCESSING PetrochemicalProcesses miele IN ce a f home processes index company index M et h a n O Feedstock Feed Application Production of highpurity methanol from hydrocarbon fi feedstocks such as natural gas process offgases and LPG up to heavy b ae naphtha The process uses conventional steamreforming synthesis gas csturator ane CLS S generation and a lowpressure methanol synthesis loop technology It i r 4 is optimized with respect to low energy consumption and maximum te J a reliability The largest singletrain plant built by Uhde has a nameplate ss a rey capacity of 1250 mtpd Description The methanol plant consists of the process steps feed puri dstilaton a ee fication steam reforming syngas compression methanol synthesis and 2 oF crude methanol distillation The feed is desulfurized and mixed with pro 2 cess steam before entering the steam reformer This steam reformer is a y Condenser e topfired box type furnace with a cold outlet header system developed Jc a Methanol by Uhde The reforming reaction occurs over a nickel catalyst Outlet Product Separator reformed gas is a mixture of Hy CO CO and residual methane It is cooled from approximately 880C to ambient temperature Most of the ree heat from the synthesis gas is recovered by steam generation BFW pre heating heating of crude methanol distillation and demineralized water preheating regulating steam pressure To avoid inert buildup in the loop a purge is Also heat from the flue gas is recovered by feedfeedsteam withdrawn from the recycle gas and is used as fuel for the reformer preheating steam generation and superheating as well as combustion Crude methanol that is condensed downstream of the methanol air preheating After final cooling the synthesis gas is compressed to reactor is separated from unreacted gas in the separator and routed the synthesis pressure which ranges from 30100 bara depending on via an expansion drum to the crude methanol distillation Water and plant capacity before entering the synthesis loop small amount of byproducts formed in the synthesis and contained in The synthesis loop consists of a recycle compressor feedeffluent the crude methanol are removed by an energysaving threecolumn exchanger methanol reactor final cooler and crude methanol separator distillation system Uhdes methanol reactor is an isothermal tubular reactor with a copper catalyst contained in vertical tubes and boiling water on the shell side Economics Typical consumption figures feed fuel range from 7 to 8 The heat of methanol reaction is removed by partial evaporation of Gcal per metric ton of methanol and will depend on the individual plant the boiler feedwater thus generating 114 metric tons of MP steam concept per metric ton of methanol Advantages of this reactor type are low byproduct formation due to almost isothermal reaction conditions high level heat of reaction recovery and easy temperature control by Commercial plants Eleven plants have been built and revamped world wide using Uhdes methanol technology Licensor Uhde GmbH is a licensee of Johnson Matthey Catalysts Low Pressure Methanol LPM Process Methanol continued PROCESSING PetrochemicalProcesses home processes index company index Methylamines Application To produce mono MMA di DMA and trimethylamines Synthesi NH Product Methanol TMA from methanol and ammonia yntnests recovery purification recovery Description Anhydrous liquid ammonia recycled amines and metha TMA nol are continuously vaporized 1 superheated 3 and fed to a cat alystpacked converter 2 The converter utilizing a highactivity low Recycle ie MMA byproduct amination catalyst simultaneously produces MMA DMA x Dehydration and TMA Product ratios can be varied to maximize MMA DMA or Ammonia TMA production The correct selection of the NC ratio and recycling of eT amines produces the desired product mix Most of the exothermic reac TaRoeaon Methanol tion heat is recovered in feed preheating 3 OMA The reactor products are sent to a separation system where the am i monia 4 is separated and recycled to the reaction system Water from Waste the dehydration column 6 is used in extractive distillation 5 to break water the TMA azeotropes and produce pure anhydrous TMA The product column 7 separates the waterfree amines into pure anhydrous MMA and DMA Methanol recovery 8 improves efficiency and extends catalyst life by allowing greater methanol slip exit from the converter Addition of a methanolrecovery column to existing plants can help to increase pro Commercial plants Twentysix companies in 18 countries use this pro duction rates cess with a production capacity exceeding 300000 mtpy Anhydrous MMA DMA and TMA can be used directly in down stream processes such as MDEA DMF DMAC choline chloride andor Licensor Davy Process Technology UK diluted to any commercial specification Yields Greater than 98 on raw materials Economics Typical performance data per ton of product amines having MMADMATMA product ratio of 13 V3 V3 Methanol t 138 Ammonia t 040 Steam t 88 Water ling m3 500 Electricity KWh 20 a COC Cee ey a PROCESSING PetrochemicalProcesses miele IN ce a JCESSE home processes index company index Mixed xylenes Application To convert Cot heavy aromatics alone or in conjunction with toluene or benzene cofeed primarily to mixed xylenes using Make yatogen Offgas to ExxonMobil Chemicals TransPlus process fuel system Description Fresh feed ranging from 100 C aromatics to mixtures of Cg aromatics with either toluene or benzene are converted primarily to xylenes in the TransPlus process Coboiling C aromatics components up to 435F NBP can be included in the Cg feed In this process liquid feed along with hydrogenrich recycle gas are sent to the reactor 2 aed after being heated to reaction temperature through feedeffluent heat 3 Cot product exchangers 3 and the charge heater 1 Primary reactions occurring are the dealkylation of alkylaromatics Fresh toluene cS Toluene and C recycle transalkylation and disproportionation producing benzenetoluene ae and Cg aromatics containing over 95 xylenes The thermodynamic Fresh Cy aromatics equilibrium of the resulting product aromatics is mainly dependent on the ratio of methyl groups to aromatic rings in the reactor feed Hydrogenrich gas from the highpressure separator 5 is recycled back to the reactor with makeup hydrogen 6 Unconverted toluene and Co aromatics are recycled to extinction The ability of TransPlus to process feeds rich in Co aromatics enhances the product slate toward xylenes Owing to its unique catalyst long cycle lengths are possible Economics Favorable operating conditions relative to other alternative technologies will result in lower capital and operating costs for grassroots units and higher throughput potential in retrofit applications Commercial plants The first commercial unit was started up in Taiwan in 1997 Performance of this unit has been excellent Licensor ExxonMobil Chemical Technology Licensing LLC retrofit ap plications Axens Axens NA grassroots applications iste se cal PetrochemicalProcesses 7 home processes index company index Mixed xylenes Application To selectively convert toluene to mixed xylene and highpu rity benzene using ExxonMobil Chemicals Toluene DisProportionation Hydrogen makeup Hydrogen recycle To fuel system 3rd Generation MTDP3 process 3 Description Dry toluene feed and up to 25 wt Cg aromatics along with hydrogenrich recycle gas are pumped through feed effluent heat exchangers and the charge heater into the MTDP3 reactor 1 Toluene disproportionation occurs in the vapor phase to produce the mixed xy Su lene and benzene product Hydrogenrich gas from the highpressure Toluene Stabilizer separator 2 is recycled back to the reactor together with makeup hy feed OC drogen Unconverted toluene is recycled to extinction Reactor yields wt Feed Product Product Cs and lighter 13 Reactor Separator eae Benzene 198 Toluene 1000 520 Ethylbenzene 06 pXylene 63 mXylene 128 oXylene 54 Ct aromatics 18 Water cooling 10C rise cmhr 03 1000 1000 25 Toluene conversion wt 48 Catalyst fill IbIb feed converted 153 10 Maintenance per year as of investment 20 Operating conditions MTDP3 operates at high space velocity and low Hhydrocarbon mole ratio These conditions could potentially result in Commercial plants Four MTDP3 licensees since 1995 increased throughput without reactor andor compressor replacement Reference Oil Gas Journal Oct 12 1992 pp 6067 in retrofit applications The thirdgeneration catalyst offers long operat ing cycles and is regenerable Licensor ExxonMobil Chemical Technology Licensing LLC retrofit ap ou lications A A NA ots applications Economics Estimated onsite battery limit investment for 1997 open shop plications Axens Axens NA grassro PP construction at US Gulf Coast location is 1860 per bpsd capacity Typical utility requirements per bbl feed converted Fuel 10 kealhr 878 click here to email for more informatio PROCESSING PetrochemicalProcesses aides ee ett JULI home processes index company index Mixed xylenes Application To convert Cot heavy aromatics alone or in conjunction with toluene or benzene cofeed primarily to mixed xylenes using Marcu Offgas to ExxonMobil Chemicals TransPlus process 0S fuel system Description Fresh feed ranging from 100 Co aromatics to mixtures of Cg aromatics with either toluene or benzene are converted primarily to xylenes in the TransPlus process Coboiling C aromatics compo nents up to 435F NBP can be included in the Cot feed In this process liquid feed along with hydrogenrich recycle gas are sent to the reactor aed 2 after being heated to reaction temperature through feedeffluent 3 C product heat exchangers 3 and the charge heater 1 imi Primary reactions occurring are the dealkylation of alkylaromatics Fresh toluene cS Toluene and C recycle transalkylation and disproportionation producing benzenetoluene a and Cg aromatics containing over 95 xylenes The thermodynamic Fresh Co aromatics equilibrium of the resulting product aromatics is mainly dependent on the ratio of methyl groups to aromatic rings in the reactor feed Hydrogenrich gas from the highpressure separator 5 is recycled back to the reactor with makeup hydrogen 6 Unconverted toluene and Cg aromatics are recycled to extinction The ability of TransPlus to process feeds rich in Co aromatics enhances the product slate toward xylenes Owing to its unique catalyst long cycle lengths are possible Economics Favorable operating conditions relative to other alternative technologies will result in lower capital and operating costs for grassroots units and higher throughput potential in retrofit applications Commercial plants The first commercial unit was started up in Taiwan in 1997 There are five TransPlus references Licensor ExxonMobil Chemical retrofit applications Axens Axens NA grassroots applications PROCESSING PetrochemicalProcesses miele IN ce a Wet home processes index company index Mixed xylenes Application In a modern UOP aromatics complex the TAC9 process is integrated into the flow scheme to selectively convert CoC 19 aromat ics into xylenes rather than sending them to the gasoline pool or selling Purge gas To fuel gas them as a solvent Description The TAC9 process consists of a fixedbed reactor and prod J uct separation section The feed is combined with hydrogenrich recycle liquid to gas preheated in a combined feed exchanger 1 and heated in a fired debutanizgr heater 2 The hot feed vapor goes to a reactor 3 The reactor effluent 1 J is cooled in a combined feed exchanger and sent to a product separa Makeup tor 4 Hydrogenrich gas is taken off the top of the separator mixed ene hydrogen soduct to with makeup hydrogen gas and recycled back to the reactor Liquid ied Recycle gas e fractionation from the bottom of the separator is sent to a stripper column 5 The stripper overhead gas is exported to the fuel gas system The overhead liquid may be sent to a debutanizer column or a stabilizer The stabilized product is sent to the product fractionation section of the UOP aromat ics complex Economics The current generation of TAC9 catalyst has demonstrated the ability to operate for several years without regeneration ISBL costs based on a unit processing 306400 mtpy of feed consisting of 100 wt CoC19 US Gulf Coast site in 2003 Investment US million 116 Utilities per mt of feed Electricity kWh 31 Steam mt 007 Water cooling m 16 Fuel MMkcal 013 Commercial plants Three commercial units have been brought on stream with feed rates ranging from 210000 mtpy to 850000 mtpy Licensor UOP LLC PROCESSING PetrochemicalProcesses aides ae Aisi home processes index company index Mixed xylenes Application The Tatoray process produces mixed xylenes and petro chemical grade benzene by disproportionation of toluene and transalk lyation of toluene and Co aromatics Purge gas To fuel gas Description The Tatoray process consists of a fixedbed reactor and product separation section The fresh feed is combined with hydrogen C rich recycle gas preheated in a combined feed exchanger 1 and heated facie in a fired heater 2 The hot feed vapor goes to the reactor 3 The debutanizer reactor effluent is cooled in a combined feed exchanger and sent to a product separator 4 Makeup J Hydrogenrich gas is taken off the top of the separator mixed with Toulene and bydiogen makeup hydrogen gas and recycled back to the reactor Liquid from the feed Recycle gas product to Br bottom of the separator is sent to a stripper column 5 The stripper 2 overhead gas is exported to the fuel gas system The overhead liquid may be sent to a debutanizer column The products from the bottom of the stripper are recycled back to the BT fractionation section of the aromatics complex The Tatoray process unit is capable of processing feedstocks ranging from 100 wt toluene to 100 wt Ag The optimal concentration of A in the feed is typically 4060 wt The Tatoray process provides an Commercial plants UOP has licensed a total of 44 Tatoray units 40 of ideal sic to produce additional mixed xylenes from toluene and heavy these units are in operation and 4 are in various stages of construction aromatics Economics The process is designed to function at a much higher level Licensor UOP LLC of conversion per pass This high conversion minimizes the size of the BT columns and the size of Tatoray process unit as well as the utility consumption of all of these units Estimated ISBL costs based on a unit processing feed capacity of 355000 mtpy US Gulf Coast site in 2003 Investment US million 113 Utilities per mt of feed Electricity kWh 175 Steam mt 011 Water cooling M 25 Fuel MMkcal 004 PROCESSING PetrochemicalProcesses miele IN ce meena MO etsyslete eae home processes index company index mXylene Application The MX Sorbex process recovers metaxylene mxylene from mixed xylenes UOPs innovative Sorbex technology uses adsorp tive separation for highly efficient and selective recovery at high purity Atsorbent chamber of molecular species that cannot be separated by conventional frac Desorbent Rotary valve tionation 1 5 Extract column Description The process simulates a moving bed of adsorbent with con Extra Mxylene tinuous countercurrent flow of liquid feed over a solid bed of adsor bent Feed and products enter and leave the adsorbent bed continuous Feed ly at nearly constant compositions A rotary valve is used to periodically paffinate column switch the positions of the feedentry and productwithdrawal points as the composition profile moves down the adsorbent bed Raffinate to storage The fresh feed is pumped to the adsorbent chamber 2 via the ro Mixed xylenes feed tary valve 1 Mxylene is separated from the feed in the adsorbent chamber and leaves via the rotary valve to the extract column 3 The dilute extract is then fractionated to produce 995 wt mxylene as a bottoms product The desorbent is taken from the overhead and recircu lated back to the adsorbent chamber All the other components present in the feed are rejected in the adsorbent chamber and removed via the Investment US million 300 rotary valve to the raffinate column 4 The dilute raffinate is then frac Utilities per mt of mxylene produced tionated to recover desorbent as the overhead product and recirculated Electricity kWh 87 back to the adsorbent chamber Steam mt 40 Water cooling m 38 Economics The MX Sorbex process has been developed to meet in creased demand for purified isophthalic acid PIA The growth in de Commercial plants Five MX Sorbex units are currently in operation and mand for PIA is linked to the copolymer requirement for PET bottle resin another unit is in design These units represent an aggregate production applications a market that continues to rapidly expand The processhas Of 335000 mtpy of mxylene become the new industry standard due to its superior environmental Licensor UOP LLC safety and lower cost materials of construction Estimated ISBL costs based on unit production of 50000 mtpy of mxylene US Gulf Coast site in 2003 ba PROCESSING PetrochemicalProcesses miele IN ce meena eerste lets ae home processesindex company index Octenes Application The DimersolX process transforms butenes to octenes which are ultimately used in the manufacture of plasticizers via iso Reaction Catalyst Separation nonanol isonony alcohol and diisononyl phthalate units section removal section Description Butenes enter the DimersolX process which comprises Octenes three sections In the reactor section dimerization takes place in multiple liquidphase reactors 1 using homogeneous catalysis and an efficient 2 recycle mixing system The catalyst is generated in situ by the reaction of components injected in the recycle loop The catalyst in the reactor effluent is deactivated in the neutralization section and separated 2 Catalyst The stabilization section 3 separates unreacted olefin monomer and Caustic Purge saturates from product dimers while the second column 4 separates eutenes oe C the octenes A third column can be added to separate dodecenes Process Yields Nearly 80 conversion of nbutenes can be attained and se lectivities toward octenes are about 85 The typical Cg product is a mixture having a minimum of 985 octene isomers with the following distribution nOctenes 7 Methylheptenes 28 Reference Convers A D Commereuc and B Torck Homogeneous Dimethylhexenes 35 oy Catalysis IFP Conference DimersolX octenes exhibit a low degree of branching resulting in higher downstream oxonation reaction yields and rates and better Licensor Axens Axens NA plasticizer quality Economics Basis ISBL 2004 for a Gulf Coast location using 50000 tpy of a raffinate2 C cut containing 75 nbutenes Investment US million 6 Typical operating cost US 60 per metric ton of octenes Commercial plants Thirtyfive Dimersol units treating various olefinic C3 and C cuts have been licensed Typical octenes production capacities range from 20000 tpy up to 90000 tpy hietesciit cal PetrochemicalProces eee C f Sets home processes index company index Olefinsprogressive separation for vata olefins recovery and raw crackedgas cK furnaces PS purification ss sete Application To produce polymergrade ethylene and propylene a buta Final dienerich C4 cut an aromatic CeCg rich raw pyrolysis gasoline and a Compression netioes highpurity hydrogen by steam pyrolysis of hydrocarbons ranging from Fuel oil removal Ethylene ethane to vacuum gas oils Mb Fuel gas 2 Feedstocks For either gaseous ethanepropane or liquid C4naphtha le Co spliter gasoil feeds this technology is based on Technips proprietary Pyroly 1st MP sis Furnaces and progressive separation This method allows producing aeipoer Ond MP Propylene olefins at low energy consumption with particularly low environmental cree Ree ethane impact Hydrocarbon feedstocks are preheated also to recover heat and mine Pe C3 splitter then cracked by combining with steam in tubular Pyrolysis Furnace 1 at an outlet temperature ranging from 1500F to 1600F The furnace Deethanizer C4 product technology can be either an SMK type for gas cracking or GK type Cond stripping rec Propane for liquid cracking The GK type design can be oriented to a high olefins yield with very flexible propyleneethylene ratios GK6 TYPE or aan om to a high BTX production GK3 type This specific approach allows long eee run length excellent mechanical integrity and attractive economics The hydrocarbon mixture at the furnace outlet is quenched rapidly in the transfer line exchangers 2 TLE or SLE generating highpressure steam In liquid crackers cracked gas flows to a primary fractionator 3 Compressed gas at 450 psig is dried and chilled A double demetha after direct quench with oil where fuel oil is separated from gasoline nizing stripping system 89 operating at medium pressure and reboiled and lighter components and then to a quench water tower 4 for wa by cracked gas minimizes the refrigeration required heat integration ter recovery to be used as dilution steam and heavy gasoline produc as well as the investment cost for separating methane top and C cut tion endpoint control bottoms A dual column conceptabsorber 10 conceptis applied A multistage compressor driven by a steam turbine compresses the between the secondary demethanizer overheads and the chilled cracked cooled gas LP and HP condensate are stripped in two separate strippers 56 where medium gasoline is produced and part of the C3 cut is re covered respectively A caustic scrubber 7 removes acid gases that minimizes the ethylene losses with a low energy requirement High purity hydrogen is produced in a cold box 11 The bottoms from the two demethanizers of different quality are sent to the deethanizer 12 The Technip progressive separation allows the deethanizer reflux ratio to be reduced The deethanizer overhead is selectively hydrogenated for acetylene conversion prior to the ethylene splitter 13 where ethylene is separated from ethane The residual eth ane is recycled for further cracking The HP stripper and deethanizer bottoms of different quality are fed to a twocolumn dual pressure depropanizing system 1415 for C3 cut separation from the C4 cut and heavies thus giving a low fouling tendency at minimum energy consumption The methylacetylene and propadiene in the C3 cut are hydroge nated to propylene in a liquidphase reactor Polymergrade propylene is separated from propane in a C3 splitter 16 The residual propane is either recycled for further cracking or exported C4s and light gasoline are separated in a debutanizer 17 Gas expansion heat recovery and external cascade using ethylene and propylene systems supply refrigeration The main features of Tech nips patented technology are Optimization of olefins yields and selection of feedstocks Reduced external refrigeration in the separation sections Autostable process heat integration acts as feed forward sys tem Simple process control large usage of stripperabsorbers towers single specification instead of distillation tower antagonistic top bottom specifications Economics Ultimate range of ethylene yields vary from 83 ethane to around 25 vacuum gas oils 35 for the intermediate fullrange naphtha These correspond to the respective total olefins yields ethylene propylene from 84 ethane to 38 vacuum gas oils and 49 for an intermediate fullrange naphtha The specific energy consump tion range is 3100 kcal kg ethylene ethane to 5500 kcalkg ethyl ene gas oil and 4700 kcal kg ethylene for an intermediate fullrange naphtha Commercial plants Technip has been awarded four ethylene plants ranging from 500 kty up to 1400 kty using either ethane or liquid feed stocks While over 300 cracking furnaces have been built and 15 units operate worldwide numerous expansions over the nominal capacity based on progressive separation techniques are under way with up to an 80 increase in capacity For ethane cracking frontend hydrogena tion scheme is also available Licensor Technip Olefinsprogressive separation for olefins recovery and raw crackedgas purification continued PROCESSING PetrochemicalProcesses miele IN ce a Wet home processes index company index Olefinsbutenes extractive distillation Application Separation of pure C olefins from olefinicparaffinic C4 mix Cx parattins tures via extractive distillation using a selective solvent BUTENEX is the Uhde technology to separate light olefins from various C feedstocks which include ethylene cracker and FCC sources a LL a MI 7 Description In the extractive distillation ED process a singlecom VL C olefins pound solvent NFormylmorpholine NFM or NFM in a mixture with C AK Extractive NM further morpholine derivatives alters the vapor pressure of the com fraction distillation Stripper ponents being separated The vapor pressure of the olefins is lowered column Sa more than that of the less soluble paraffins Paraffinic vapors leave the top of the ED column and solvent with olefins leaves the bottom of the ED column C C The bottom product of the ED column is fed to the stripper to eee separate pure olefins mixtures from the solvent After intensive heat exchange the lean solvent is recycled to the ED column The solvent which can be either NFM or a mixture including NFM perfectly satisfies the solvent properties needed for this process including high selectivity thermal stability and a suitable boiling point Economics Consumption per metric ton of FCC C fraction feedstock Steam tt 0508 Water cooling AT 10C mt 150 Electric power kWht 250 Product purity nButene content 99 wt min Solvent content 1 wtppm max Commercial plants Two commercial plants for the recovery of nbu tenes have been installed since 1998 Licensor Unde GmbH PROCESSING PetrochemicalProcesses j home processes index company index Olefins by dehydrogenation Application The Uhde STeam Active Reforming STAR process produces HP steam a propylene as feedstock for polypropylene propylene oxide cumene Air acrylonitrile or other propylene derivatives and b butylenes as feed Fuel gas tt stock for methyl tertiary butyl ether MTBE alkylate isooctane polybu ae Raw gas tylenes or other butylene derivatives pat rn compression Fuel gas rerormer Feed Liquefied petroleum gas LPG from gas fields gas condensate Opair a fields and refineries Oxy i P reactor separation Product Propylene polymer or chemicalgrade isobutylene nbutylenes highpurity hydrogen H may also be produced as a byproduct Hydrocarbon feed Olefin Boiler feed water product Description The fresh paraffin feedstock is combined with paraffin re SReeeernneern cycle and internally generated steam After preheating the feed is sent eae nae Hydrocarbon to the reaction section This section consists of an externally fired tubular recycle fixedbed reactor Uhde reformer connected in series with an adiabat ic fixedbed oxyreactor secondary reformer type In the reformer the endothermic dehydrogenation reaction takes place over a proprietary noble metal catalyst separated from unconverted paraffins in the fractionation section In the adiabatic oxyreactor part of the hydrogen from the interme Apart from lightends which are internally used as fuel gas the diate product leaving the reformer is selectively converted with added a i olefin is the only product Highpurity Hy may optionally be recoverd oxygen or air thereby forming steam This is followed by further dehy from lightends in the gas separation section drogenation over the same noblemetal catalyst Exothermic selective H conversion in the oxyreactor increases olefin product spacetime yield Economics Typical specific consumption figures for polymergrade and supplies heat for further endothermic dehydrogenation The reac propylene production are shown per metric ton of propylene product tion takes place at temperatures between 500C600C and at4 bar6 jncluding production of oxygen and all steam required bar Propane kgmetric ton 1200 The Uhde reformer is topfired and has a proprietary cold out Fuel gasGJmetric ton 64 let manifold system to enhance reliability Heat recovery utilizes process Circul cooling water m3metric ton 220 heat for highpressure steam generation feed preheat and for heat re Electrical energy kWhmetric ton 180 quired in the fractionation section After cooling and condensate separation the product is subse quently compressed lightends are separated and the olefin product is aie ee mm sir eel a Commercial plants Two commercial plants using the STAR process for dehydrogenation of isobutane to isobutylene have been commissioned in the US and Argentina More than 60 Uhde reformers and 25 Uhde secondary reformers have been constructed worldwide References HeinritzAdrian M N Thiagarajan S Wenzel and H Gehrke STARUhdes dehydrogenation technology an alternative route to C3 and C4olefins ERTC Petrochemical 2003 Paris France March 2003 Thiagarajan N U Ranke and F Ennenbach Propanebutane de hydrogenation by steam active reforming Achema 2000 Frankfurt Germany May 2000 Licensor Uhde GmbH Olefins by dehydrogenation continued tistesstttcial PetrochemicalProcesses miele IN ce meena JCESSES home processes index company index Olefins Application To produce ethylene propylene and butenes from natural gas or equivalent via methanol using the UOPHydro MTO methanol Reactorregeneration Productrecovery to olefins process section CH I Description This process consists of a reactor section a continuous cat Product alyst regeneration section and product recovery section One or more eee fluidizedbed reactors 1 are used with continuous catalyst transfer to fae Sere and from the continuous catalyst regenerator 2 The robust regener able MTO100 catalyst is based on a nonzeolitic molecular sieve Raw Water 98 Purity nondewatered methanol is fed to the lowpressure reactor 1 which propylene offers very high 99 conversion of the methanol with very high se 2 lectivity to ethylene and propylene The recovery section design depends on product use but will contain a product water recovery and recycle MeOH system 3 a CO removal system 4 a dryer 5 a deethanizer 6 an Air Ca product acetylene saturation unit 7 a demethanizer 8 and a depropanizer 9 The process can produce polymergrade ethylene and propylene by adding simple fractionation to the recovery section Yields The process gives very high total olefins yields A typical product yield structure is shown based on 5204 mtd raw methanol feedrate to an MTO plant Economics The MTO process competes favorably with conventional Metric tpd liquid crackers due to lower capital investment It is also an ideal ve Ethylene 887 hicle to debottleneck existing ethylene plants and unlike conventional Propylene 882 steam crackers the MTO process is a continuous reactor system with Total light olefins 1762 no fired heaters Butenes 272 Commercial plants Hydro operated a demonstration unit that was in Cs 100 stalled in Norway in 1995 The first commercial MTO unit is planned for Fuel gas 88 Startup in 2008 in Nigeria Other water coke CO 2980 Licensor UOP LLCHydro The process is flexible Ethylene to propylene product weight ratio can be modified between the range of 075 to 13 by altering reactor re rane severity The total yield of olefins varies slightly throughout Sa click here to email for more information a iste se cal PetrochemicalProcesses miele IN ce a Mele erste home processes index company index Olefins catalytic Application To selectively convert vacuum gas oils and the resulting Product vapors blends of each into CCz olefins aromaticrich highoctane gasoline and distillate using deep catalytic cracking DCC methods Reactor Description DCC is a fluidized process to selectively crack a wide va Flue gas Vapor and catalyst riety of feedstocks into light olefins Propylene yields over 24 wt are distributor achievable with paraffinic feeds A traditional reactorregenerator unit Stripper design uses a catalyst with physical properties similar to traditional FCC Regenerator catalyst The DCC unit may be operated in two operational modes max imum propylene Type or maximum isoolefins Type Il Each opera Reactor riser tional mode utilizes unique catalyst as well as reaction conditions DCC Combustion air p maximum propylene uses both riser and bed cracking at severe reactor conditions while Type Il utilizes only riser cracking like a modern FCC Regenmcanerndpive Riser steam unit at milder conditions feed nozzlesFIT The overall flow scheme of DCC is very similar to a conventional FCC However innovations in catalyst development process variable selection and severity enables the DCC to produce significantly more olefins than FCC in a maximum olefins mode of operation Products wt FF DCC Typel DCC Type ll FCC Ethylene 61 23 09 Reference Dharra et al Increase light olefins production Hydrocar Propylene 205 143 68 bon Processing April 2004 Butylene 143 146 110 in which IC 54 61 33 Licensor Stone Webster Inc A Shaw Group CoResearch Institute of Amylene 98 85 Petroleum Processing Sinopec in which IC 65 43 This technology is suitable for revamps as well as grassroot applications Commercial plants Currently eight units are in operation seven in Chi na and one in Thailand Another plant for Saudi Aramco presently in design will be the largest DCC unit in the world PROCESSING PetrochemicalProcesses eRe tie AMMO TC Ae AT Ore Ut aa f E home processes index company index Normal paraffins CC3 Application The Molex process recovers normal Ci 9C3 paraffins from kerosine using UOPs innovative Sorbex adsorptive separation techn ology Makeup hydrogen Light ends Description Straightrun kerosine is fed to a stripper 1 and a rerun Light kerosine Reeve gs column 2 to remove light and heavy materials The remaining heartcut Normal paraffin kerosine is heated in a charge heater 3 and then treated in a Union Streightrun P fining reactor 4 to remove impurities The reactor effluent is sent to a kerosine product separator 5 to separate gas for recycle and then the liquid is natfinate sent to a product stripper 6 to remove light ends The bottoms stream from the product stripper is sent to a Molex unit 7 to recover normal paraffins kerosize Feedstock is typically straightrun kerosine with 1850 normal paraffin content Product purity is typically greater than 99 wt Economics Investment US Gulf Coast battery limits for the production of 100000 tpy of normal paraffins 700 tpy Commercial plants Twentyeight Molex units have been built Reference McPhee A Upgrading Kerosene to Valuable Petrochemi cals 24th Annual DeWitt World Petrochemical Review Houston Texas US March 1999 Licensor UOP LLC a Ce Oe j PROCESSING PetrochemicalProcesses home processesindex company index Paraffin normal Application Efficient lowcost recovery and purification processes for the production of LABgrade andor highpurity nparaffin products Ammonia LAB grade nparaffins product from kerosine Desorbent il ar Molecular Description The ExxonMobil Chemical EMC process offers commer 1 am sieve beds Adsorption ee VY YY Desorption cially proven technologies for efficient recovery and purification of high 1 A purity nparaffin from kerosine feedstock Kerosine feedstocks are in letfuel 2 Wy a VY troduced to the recovery section where the nparaffins are efficiently to ci Zi Zi Desorption Y Y recovered from the kerosine stream in a vaporphase fixedbed molecu Adsorption ar 1 to det Tue lar sieve adsorption process In the process the nparaffins are selec VY 1 ee 7 tively adsorbed on a molecular sieve and subsequently desorbed with a is sieve beds highly effective desorbent SS High pany The non nparaffin hydrocarbons are rejected and returned to the Recovery Purification rere refinery The process provides a unique environment allowing the solid section section adsorbent to be very tolerant of sulfur compounds which are typically present in kerosine feedstock The adsorbent is therefore able to last long cycle lengths with a total life up to 20 years as commercially demonstrated by ExxonMobil In most cases due to the high sulfur tolerance the kerosine feedstock will not require hydrotreating pretreatment which significantly reduces capital investment and operating cost The recovery te bn iri section produces LABgrade nparaffin product Product qualty Typical properties of highpurity nparaffin product Highpurity specialtygrade nparaffin products are produced in Ay tet m 100 the ExxonMobil Purification process The LABgrade product from the Bromine Midex mg100g 20 recovery process is further processed in a purification section where Sulfur wt ppm 1 residual aromatics and other impurities are further reduced Purification is accomplished in a liquidphase fixedbed adsorption system The Yield Typically over 99 of the nparaffin contained in the kerosine impurities are selectively adsorbed ona molecular sieve andsubsequently stream is recovered removed with a hydrocarbon desorbent The highpurity nparaffins product is the highest quality available in the market ExxonMobil Fommercia Plants Pxoniopl shea has Ay years oF schicer in the commercially produces and markets nparaffin product with aromatics production of nparattins ana Is the second largest producer in the content below 100 wtppm The ExxonMobil nparaffin technologies offer the industrys lowest capital and operating cost solutions and click here to email for more information a highest purity products for nparaffin producers world ExxonMobils nparaffin plant at Baytown Texas produces high purity product in a single train at a nameplate capacity of 250000 tpy Licensor Kellogg Brown Root Inc Paraffin normal continued hietesciit cal PetrochemicalProcesses PROCESSING AVL UEE TELE home processes index company index Paraxylene Application Suite of advanced aromatics technologies combined in the Raffinate most effective manner to meet customers investment and production 4 objectives for paraxylene and benzene and are licensed under the name Benzene Paramax extraction Paraxylene Cc Description Aromatics are produced from naphtha in the Aromizing Toluene section 1 and separated by conventional distillation The xylene frac MH Cyt a Tol tion is sent to the Eluxyl unit 2 which produces 999 paraxylene via simulated countercurrent adsorption The PXdepleted raffinate is 2 C 5 isomerized back to equilibrium in the isomerization section 3 with ei ae Reforming 7 ther EB dealkylationtype XyMax processes or EB isomerizationtype 1 G S CotCio Oparis catalysts Highpurity benzene and toluene are separated from Cyt nonaromatic compounds with extractive distillation Morphylane 3 Heavy processes 4 Toluene and Cg to C aromatics are converted to more Cot Crot aromatics valued benzene and mixed xylenes in the TransPlus process 5 leading to incremental paraxylene production Eluxyl technology has the industrially proven ability to meet ultimate single train PX purity and capacities as high as 750000 mtpy Proprietary hybrid Eluxyl configurations integrate an intermediate purity adsorption section with a singlestage crystallization ideal for retrofits Axens is the Investment million US 430 licensor of all the technologies involved in the Paramax suite Annual utilities aaa and chemical a1 Mobil and Uhde technologies licensed by Axens for grassroots applications operating cost million USyr Production Typical paraxylene single train complex from naphtha to rommersial Pa of pet olene tnd three unite that sre Moperation Si paraxylene featuring Aromizing Eluxyl XyMax and TransPlus units isomerization units use the Oparis catalyst and 19 ExxonMobil EB deal Feed60175 Arab light naphtha Thousand PY kylating units have been put into operation Three TransPlus units are Paraxylene 600 currently in operation Net producer of hydrogen 08 Reference Dupraz C et al Maximizing paraxylene production with ParamaxX Hotier G and Methivier A Paraxylene Production with Economics The ISBL 2004 Gulf Coast location erected cost including 5 first load of catalysts and chemicals with 30 allowance for offsites the Eluxyl Process AIChE 2002 Spring Meeting New Orleans March 2002 Licensor Axens Axens NA Paraxylene continued PROCESSING PetrochemicalP eee OTe ATLANTIC Ud Mele cists iets home processes index company index Paraxylene Application To selectively convert toluene to highpurity 90 para k xylenerich PX xylenes and benzene using ExxonMobil Chemicals tech le Hydrogen recycle To fuel system nologies PxMax and ASTDP Description Dry toluene feed and hydrogenrich recycle gas are pumped through feedeffluent exchangers and charge heater and into the reac tor 1 Selective toluene disproportionation STDP occurs in the vapor i phase to produce the paraxylenerich xylene and benzene coproduct cw Byproduct yields are small Reactor effluent is cooled by heat exchange Toluene Stabilizer and liquid products are separated from the recycle gas Hydrogenrich ie gas from the separator 2 is recycled back to the reactor together with makeup hydrogen Liquid product is stripped of remaining light gas in the stabilizer 3 and sent to product fractionation Unconverted toluene Product is recycled to extinction Reactor Separator fractionation The PxMax technology uses catalyst which is exsitu selectivated by pretreatment during catalyst manufacture The ASTDP technology uses catalyst which is insitu coke selectivated Both technologies provide significantly higher selectivity and longer operating cycles than other STDP technologies Operating costs associated with downstream recovery are also reduced by the high paraxylene purity from PxMax and ASTDP Licensor ExxonMobil Chemical retrofit applications Axens Axens NA wg grassroots applications Operating conditions PxMax operates at lower startofcycle temperatures and lower hydrogen to hydrocarbon recycle ratios than other STDP technologies resulting in longer cycles and lower utilities By eliminating the insitu selectivation step the PxMax version of this technology results in simplified operation and lower capital costs Both catalysts offer long operating cycles and are regenerable Commercial plants There are seven MSTDP units predecessor technology to PxMax and ASTDP and four units using PxMax technology The first two PxMax units started up in 1996 and 1997 at Chalmette Refinings Louisiana Refinery and Mobil Chemicals Beaumont plant respectively PROCESSING PetrochemicalProcesses miele IN ce meena JCESSES home processes index company index Paraxylene Application To selectively convert toluene to highpurity 90 para bivd k xylenerich PX xylenes and benzene using ExxonMobil Chemicals tech ee Hydrogen recycle To fuel system nologies PxMax and ASTDP Description Dry toluene feed and hydrogenrich recycle gas are pumped through feedeffluent exchangers and charge heater and into the reac tor 1 Selective toluene disproportionation STDP occurs in the vapor i phase to produce the paraxylenerich xylene and benzene coproduct cw Byproduct yields are small Reactor effluent is cooled by heat exchange Toluene Stabilizer and liquid products are separated from the recycle gas Hydrogenrich feed C gas from the separator 2 is recycled back to the reactor together with makeup hydrogen Liquid product is stripped of remaining light gas in the stabilizer 3 and sent to product fractionation Unconverted toluene Product is recycled to extinction Reactor Separator eee The PxMax technology uses catalyst which is exsitu selectivated by pretreatment during catalyst manufacture The ASTDP technology uses catalyst which is insitu coke selectivated Both technologies provide significantly higher selectivity and longer operating cycles than other STDP technologies Operating costs associated with downstream recovery are also reduced by the high paraxylene purity from PxMax and ASTDP Licensor ExxonMobil Chemical Technology Licensing LLC retrofit ap wa plications Axens Axens NA grassroots applications Operating conditions PxMax operates at lower startofcycle tempera tures and lower hydrogen to hydrocarbon recycle ratios than other STDP technologies resulting in longer cycles and lower utilities By eliminating the insitu selectivation step the PxMax version of this technology re sults in simplified operation and lower capital costs Both catalysts offer long operating cycles and are regenerable Commercial plants There are seven MSTDP units predecessor tech nology to PxMax and ASTDP and four units using PxMax technology The first two PxMax units started up in 1996 and 1997 at Chalmette Refinings Louisiana Refinery and Mobil Chemicals Beaumont plant respectively PROCESSING PetrochemicalProcesses miele IN ce a home processes index company index Paraxylene ese Benzene Application A UOP aromatics complex is a combination of process units here which are used to convert petroleum naphtha and pyrolysis gasoline Hydrogen into the basic petrochemical intermediates benzene toluene paraxy lene andor orthoxylene J Description The configuration of an aromatics complex depends upon the available feedstock the desired product slate and the balance be Aromatics C tween performance and capital investment A fully integrated modern i Paragiene complex contains a number of UOP process technologies Naphtha Light ends The naphtha feed is first sent to a UOP naphtha hydrotreating unit 1 to remove sulfur and nitrogen compounds and then sent to a CCR Platforming unit 2 to reform paraffins and naphthenes to aromatics The reformate produced in the CCR Platforming unit is sent to a debutanizer column which strips off the light ends The debutanizer bottoms are sent to a reformate splitter 3 The C7 fraction from the overhead of the reformate splitter is sent to a Sulfolane unit 4 The Cg fraction from the bottom of the reformate splitter is sent to a xylene fractionation section The Sulfolane unit extracts the aromatics and then individual highpurity benzene and toluene products are recovered in a BT fractionation section 5 6 Toluene is usually blended with Co aromatics Ag from the at very high recovery The raffinate from the Parex unit is almost overhead of the heavy aromatics column 7 and charged to a Tatoray entirely depleted of paraxylene and is sent to an Isomar unit 12 In unit 8 for production of additional xylenes and benzene Toluene and the Isomar unit additional paraxylene is produced by reestablishing heavy aromatics can also be charged to a THDA unit 9 for production an equilibrium distribution of xylene isomers The effluent from the of additional benzene lsomar unit is sent to a deheptanizer column 13 The bottoms from The Cg fraction from the bottom of the reformate splitter is the deheptanizer are recycled back to the xylene splitter column charged to a xylene splitter column 10 The bottom of the xylene splitter column is sent to the oxylene column 14 to separate high Economics A summary of the investment cost and the utility consump purity oxylene product and the bottoms are sent to the heavy aromatics tion for a typical paraxylene aromatics complex to process 1336 million column 7 mtpy of naphtha feed is indicated below The estimated ISBL erected The xylene splitter overhead is sent directly to a Parex unit 11 where 999 wt pure paraxylene is recovered by adsorptive separation cost for the unit assumes construction on a US Gulf coast site in 2003 Investment US million 274 Products mtpy Benzene 226000 paraxylene 700000 Pure hydrogen 47000 Utilities per mt of feed Electricity kWh 643 Steam mt 02 Water cooling m3 359 Fuel Gcal 25 Commercial plants UOP is the worlds leading licensor of process tech nology for aromatics production UOP has licensed more than 600 sepa rate process units for aromatics production including over 200 CCR Platforming units 134 Sulfolane units 80 Parex units 61 Isomar units 44 Tatoray units and 38 THDA units UOP has designed 80 integrated aromatics complexes which produce both benzene and paraxylene These complexes range in paraxylene production capacity from 21000 to 12 million mtpy Licensor UOP LLC Paraxylene continued PROCESSING PetrochemicalProcesses miele IN ce OTC ATTA AT TOre ted R01C erste ers eae home processes index company index Paraxylene Application To produce a desired xylene isomer or isomers from a mix ture of Cg aromatics using the UOP Isomar and Parex processes pe makeup or Description Fresh feed containing an equilibrium mixture of Cg aromatic isomers is fed to a xylene splitter 1 Bottoms from the splitter are then separated 2 into an overhead product of oxylene and a byproduct of 5 Cg aromatics Overhead from the splitter is sent to a UOP Parex process unit 3 to recover ultrahighpurity pxylene If desired highpurity m eeu Bee xylene may also be recovered using the MX Sorbex process Remaining igen components are recycled to the UOP Isomar process unit reactor 4 where they are catalytically converted back toward an equilibrium mix ture of Cg aromatic isomers Hydrogenrich recycle gas is separated 5 from the reactor effluent before fractionation 6 to remove lightcracked Co aromatics byproducts overhead The remaining Cg aromatics are then combined with the fresh feed and sent to the xylene splitter 1 The feedstock consists of a mixture of Cg aromatics typically derived from catalytically reformed naphtha hydrotreated pyrolysis gasoline or an LPG aromatization unit The feed may contain up to 40 ethylbenzene which is converted either to xylenes or benzene by the Isomar reactor at a highconversion rate per pass Feedstocks may be pure solvent extracts Composition Fresh feed wt units Product wt units or fractional heartcuts containing up to 25 nonaromatics Hydrogen ey eeazene 4 may be supplied from a catalytic reforming unit or any suitable source Pn exylene 410 Chemical hydrogen consumption is minimal oXylene 195 196 oXylene product purity of up to 99 is possible depending on the composition of the feed and fractionation efficiency The Parex unit Economics Estimated inside battery limits ISBL erected and utility costs is capable of producing 999 pure pxylene with per pass recovery are given for a Parexlsomar complex which includes the xylene splitter greater than 97 column and the oxylene column US Gulf Coast fourth quarter 2002 Investment US per mt of feed 94108 Operating conditions Moderate temperature and pressure requirements Utilities US per mt of pxylene product 30 permit using carbon and lowalloy steel and conventional process equip ment Yields Typical mass balance for the Parexlsomar complex Commercial plants Since 1971 UOP has licensed 80 Parex units and 61 Isomar units Licensor UOP LLC Paraxylene continued PROCESSING PetrochemicalProcesses home processes index company index Paraxylene Application The PXPlus XP Process converts toluene to paraxylene and benzene The paraxylene is purified to 999 wt via singlestage crys fhannence tallization and a wash column The benzene purity is 545grade by frac enzene tionation Hydrogen Paraxylene Description The PXPlus XP Process is composed of three processing steps Toluene 1 Selective toluene disproportionation via the PXPlus Process 2 Fractionation for recovery of recycle toluene and benzene prod uct Recycle toluene Heavies Mother liquor 3 The BadgerNiro paraxylene crystallization process where single stage crystallization and crystal wash columns are used In the PXPlus technology fresh toluene is combined with recycle gas heated and fed to a fixedbed reactor The paraselective catalyst produces xylene product with 90 paraxylene in the xylenes Reactor effluent flows to a separator where the recycle gas is recovered and the Yields liquid product i sent toa stripper Toluene conversion per pass 30 In the fractionation section stripper bottoms are fed to a benzene Paraxylene yield wt 40 column where the benzene product is recovered and the unconverted Benzene yield wt 45 toluene is fractionated for recycle The toluene column bottoms are sent Light ends wt 6 to a rerun column where the paraxylene concentrated fraction is taken Paraxylene recovery 935 overhead Paraxylene purity wt 999 In the BadgerNiro crystallization unit the xylenes are fed toa Economics Capital investment per mty of paraxylene product singlestage crystallization section that uses continuous suspension EEC US 200 crystallization In this section the paraxylene is purified with a single a Utilities per mt of paraxylene product refrigerant compressor system and the mother liquor rejected The Electricity kWh 87 purified paraxylene is fed to a Niro wash column section where Steam HP mt 07 ultrahighpurity paraxylene is produced by countercurrent crystal Steam LP mt 007 washing Water cooling m 15 Components of this flexible technology are especially suited for Fuel MMkcal 2 capacity expansion of existing paraxylene production facilities a COC Cee ey a Commercial plants Two PXPlus units are in operation another unit is in design and construction Two BadgerNiro licensed and process pack ages were produced for three BadgerNiro crystallization projects Licensor UOP LLC Stone Webster Inc and Niro Process Technology BV Paraxylene continued PROCESSING PetrochemicalProcesses PROCESSING home processes index company index Paraxylene crystallization Application CrystPX is suspension crystallization technology to improve I Highpurity paraxylene f Paraxylene crystallization production of paraxylene increasing capacities increasing purity levels production section i recovery section achievable simplifying operation scheme and significantly lowering Scrapedsurface Filtrate recycle capital investment The technology optimizes current equipment and crystallizer i Scrapedtsurface Hee ondary design techniques to deliver efficient and reliable production utilizing ep O H centrifuge ft lean flexible attainable equipment and feed streams centrifuge or 1 il Description Suspension crystallization of paraxylene PX in the xylene i isomer mixture is used to produce paraxylene crystals The technology ee i Crystals to feed drum uses an optimized arrangement of equipment to obtain the required td recovery and product purity Washing the paraxylene crystal with the ea final product in a high efficiency pushercentrifuge system produces the Feed paraxylene product When paraxylene content in the feed is enriched above equilibrium Paraxylene wash for example streams originating from selective toluene conversion pro ns cesses the proprietary crystallization process technology is even more economical to produce highpurity paraxylene product at high recover ies The process technology takes advantage of recent advances in crys tallization techniques and improvements in equipment to create this ec onomically attractive method for paraxylene recovery and purification Crystallization equipment is simple easy to procure and opera Design uses only crystallizers and centrifuges in the primary opera tionally trouble free tion This simplicity of equipment promotes low maintenance costs easy Compact design requires small plot size and lowest capital invest incremental expansions and controlled flexibility Highpurity paraxylene ment is produced in the front section of the process at warm temperatures System is flexible to meet market requirements for paraxylene pu taking advantage of the high concentration of paraxylene already in rity the feed At the back end of the process high paraxylene recovery is System is easily amenable to future requirement for incremental obtained through a series of crystallizers operated successively at colder Capacity increases temperatures This scheme minimizes the need for recycling excessive Feed concentration of paraxylene is used efficiently amounts of filtrate thus reducing overall energy requirements Technology is flexible to process a range of feed concentrations 7595 wt paraxylene in a 1stage refrigeration system Process advantages include ose coy purity and recovery 998 wt purity at up to i Design variations are used to recover paraxylene efficiently from feedstocks 22 PX in a multistage system competitive with adsorp tionbased systems Economics Technoeconomic comparison of CrystPX to conventional technologies basis 90 PX feed purity 400000 tpy of 998 wt PX CrystPX Other crystallization technologies Investment cost MM 260 400 Paraxylene recovery 95 95 Electricity consumption kWhton PX 50 80 Operation mode Continuous Batch Licensor GTC Technology in alliance with Lyondell Chemical Co Paraxylene crystallization continued iste se cal PetrochemicalProcesses miele IN ce a f home processes index company index Phenol Application Improved technology to produce highest quality phenol wresh vena Recycle cumene and acetone from cumene Refined alpha methyl styrene AMS produc cumene ca a eee tion is optional High yield is achieved at low operating and capital costs product without tar cracking é i Recycle Recycled 4 Description Fresh and recycle cumene is oxidized 1 with air to form eS acetone pu 12 cumene hydroperoxide CHP using new oxidizer treatment technology a Light waste toe lyarocarbon to reduce organic acid formation and improve selectivity Overhead va Air 3 pors are cooled and condensed to recover cumene Spent air is treated Phencl to absorb and recover residual hydrocarbons phenol product Oxidate is concentrated in a multistage cumene stripping system c 2 Concentrated CHP flows directly to the cleavage unit where it is a decomposed under precisely controlled conditions using new two Catalyst stage Advanced Cleavage Technology 3a and 3b Cleavage conditions re hydrcarbos are optimized to permit CHP decomposition without producing heavy byproducts Cleavage effluent is neutralized 4 before the mixture is fractionated Neutralized cleavage effluent is first split into separate acetone cumeneAMSwater and phenolheavier fractions 5 Overheads from the splitter are then fractionated to remove aldehydes 6 and cumene Commercial plants GE Plastics Mt Vernon Indiana 300000 metric AMSwater 7 to produce highpurity acetone 9975 wt Splitter tonsyr mtpy revamped in 1992 Formosa Chemicals Fibre Corpora bottoms is fractionated under vacuum to produce a crude phenol distillate tion Taiwan 400000 mtpy revamped in 2001 to double the original 8 and a heavy waste hydrocarbon stream Hydrocarbon impurities plant capacity Lummus has more than 50 years of phenolplant design are removed from the crude phenol by hydroextractive distillation 9 experience followed by catalytic phenol treatment 10 and vacuum distillation 11 to produce ultrahighpurity phenol 9999 wt Licensor ABB Lummus GlobalGE Plastics Illa International Phenol is recovered from the acetone finishing column bottoms 12 by extraction with caustic AMS in the raffinate is then concentrated 13 hydrogenated 14 and recovered as cumene for recycle to oxidation Refined AMS production is optional Yields 100000 tons of phenol and 61500 tons of acetone are produced a i in ti i from 131600 tons of cumene giving a product yield of over 99 hietesciit cal PetrochemicalP eee C fOCESSES home processes index company index Phenol Application A highyield process to produce highpurity phenol and ac etone from cumene with optional byproduct recovery of alpha methyl Acetone styrene AMS and acetophenone AP Phenol Description Cumene is oxidized 1 with air at high efficiency 95 Catalyst to produce cumene hydroperoxide CHP which is concentrated 2 Air and cleaved 3 under highyield conditions 99 to phenol and ac 3 etone in the presence of an acid catalyst The catalyst is removed and the cleavage mixture is fractionated to produce highpurity products 48 suitable for all applications AMS is hydrogenated to cumene omens Hydro and recycled to oxidation or optionally recovered as a pure byproduct Waste oils to fuel Phenol and acetone are purified A small aqueous effluent is pretreat recovery cy ed to allow efficient biotreatment of plant wastewater With AMS hy Wastewatst drogenation 131 tons of cumene will produce 1 ton of phenol and AMS optional 0615 tons of acetone This highyield process produces very high quality phenol and acetone products with very little heavy and light end byproducts With over 40 years of continuous technological devel opment the Kellogg Brown Root KBR phenol process features low cumene and energy consumptions coupled with unsurpassed safety and environmental systems Commercial plants Thirty plants worldwide have been built or are now under construction with a total phenol capacity of over 28 MMtpy KBR has licensed 7 grassroots plants in 10 years with a total capacity of 10 MMtpy Three new licenses were awarded in 2004 with two startups scheduled for 2005 More than 50 of the worlds phenol is produced via the KBR process Reference Hydrocarbon Engineering DecemberJanuary 1999 Licensor Kellogg Brown Root Inc iste se cal PetrochemicalProcesses ae APL ASARMI ATCA UL ACMA RNS LEE home processes index company index Phenol Application The SunocoUOP phenol process produces highquality phenol and acetone by liquidphase peroxidation of cumene Description Key process steps Spent alr Oxidation and concentration 1 Cumene is oxidized to cumene hydroperoxide CHP A small amount of dimethylphenylcarbinol DMPC Cumene a ae ener Acetone is also formed but lowpressure and lowtemperature oxidation results Air neutralization purification Phenol in very high selectivity of CHP CHP is then concentrated and unreacted 2 3 Residue cumene is recycled back to the oxidation section Decomposition and neutralization 2 CHP is decomposed to phenol H AMS and acetone accompanied by dehydration of DMPC to alphamethylstyrene a part AMS catalyzed by mineral acid This unique design achieves a very high 4 selectivity to phenol acetone and AMS without using recycle acetone The high total yields from oxidation and decomposition combine to achieve 131 wt cumenewt phenol without tar cracking Decomposed catalyst is neutralized Phenol and acetone purification 3 Phenol and acetone are separated and purified A small amount of byproduct is rejected as heavy residue h Le molified izati AMS hydrogenation or AMS refining 4 AMS is hydrogenated back recycle tot e decomposition cleavage section simplified neutra IZation i re and no tar cracking make the SunocoUOP Phenol process easier to to cumene and recycled to oxidation or AMS is refined for sale oe operate Cumene peroxidation is the preferred route to phenol accounting for more than 90 of world production The SunocoUOP Phenol process Commercial plants The SunocoUOP Phenol process is currently used in features low feedstock consumption 131 wt cumeneAwt phenol 11 plants worldwide having total phenol capacity of more than 1 mil without tar cracking avoiding the expense and impurities associated ign mtpy Four additional process units with a total design capacity of with tar cracking High phenol and acetone product qualities are achieved 690000 mtpy are in design and construction through a combination of minimizing impurity formation and efficient purification techniques Optimized design results in low investment Licensor Sunoco and UOP LLC cost along with low utility and chemicals consumption for low variable cost of production Design options for byproduct alphamethylstyrene AMS allow producers to select the best alternative for their market hydrogenate AMS back to cumene or refine AMS for sale No acetone iste se cal PetrochemicalProcesses miele IN ce a f home processes index company index Phthalic anhydride Application To produce phthalic anhydride PA from oxylene naph BiinereatedlHP steam thalene or mixtures of both feedstocks using a fixedbed vapor phase HP steam process originally known as the von Heyden Process Bol eniecaiater tsein Description Air is heated and loaded with evaporated 1 oxylene and Catalytic to ndneegee or naphthalene The hydrocarbonair mixture enters a multitubular re oxidation Vacuum actor 2 containing catalyst An agitated salt melt removes the heat of 1 unit reaction and maintains constant temperature conditions Reaction heat HP Hot Ee PA generates highpressure steam Steam oil a product Modern plants operate with oxylene feedstock loadings of 90100 tP i 7 9 gNm3 air The loadings of 100 gNm air in an adiabactic postreactor a 6 BFW is recommended which is installed in the enlarged gas cooler casing Steam Mend Hot 3 Reactor effluent gas is precooled in a gas cooler 3 before part hot oil i 10 of the PA vapor is condensed to a liquid in the precondensor 4 and oXylene Steam Liquid CTU4E Lights HB residue is continuously discharged to the crude PA tank 5 The remainder of Naphthalene FMT condenser P column ieee about 65 g PAm in the reaction gas is condensed as solid sublimate in switch condensors 6 on specially designed finned tubes The switch condensors are periodically cooled and heated in a discontinuous operation of an automated switching cycle using heat transfer oil circuits Yield 110112 kg PA from 100 kg of pure oxylene 9799 kg PA from During the heating phase solid PA is melted from the condensor tubes 100 kg of pure naphthalene and discharged as a liquid to crude PA tank Effluent gas is vented to the a atmosphere after water scrubbing andor incineration Economics Excellent energy utilization and minimized offgas volume The crude PA is thermally pretreated 7 and then fed to the vacuum ae due to high hydrocarbonair ratio Plants can be designed to operate distillation system Low boiling LB impurities are removed in the lights independently of external power supply and export electric energy or column 8 as LB residues The highboiling HB residue from the pure T1P steam PA column 9 is sent to the residue bollout vessel for PA recovery Pure Commercial plants More than 110 plants with typical production ca PA obtained as a distillate can be stored either in the molten state or pacities of 2000075000 tpy with a maximum capacity of 140000 flaked and bagged toy have been designed and built by Lurgi Catalyst Special highperformance catalysts oxidize oxylene as well as Licensor BASF AG and Lurgi AG naphthalene and mixtures of both feedstocks in any proportions All eee ene aa PROCESSING PetrochemicalP eee C Oe Sich home processes index company index Polyalkylene terephthalatesPET PBT PTT PEN CHO he pcs omni i Application New process to produce polyesters from the polyalkylene 1 terephthalate family from terephthalic acid PTA or dimethyl terephthal aye Diol ate DMT and diols using the UIF proprietary tworeactor 2R process tower reactor eS consisting of tower reactor ESPREE and DISCAGE finisher or alterna es tively a solidstate finishing Catalyst A Description A slurry composed of a dicarboxylic acid and a diol is pre SI E pared at a low mole ratio The slurry is fed to the tower reactors bottom wy Fo f M Meltphase where the main esterification occurs under pressure or under vacuum at commer finisher temperatures ranging between 170C to 270C This reaction may be orAldil 1v055130 catalyzed or autocatalyzed or DMTdiol Monomer is transferred via a pressurized pipe to the reactor top Polyesters where reaction side products are flashed out Higher conversion rates 1 Ea LV 075 130 9799 are achieved by a cascade of four to six reaction cups at finishing decreasing pressures and increasing temperatures Stirring and intermix are done by reaction vapors while passing through the cups A precondensate with iVs of 028 to 035 is obtained after surfaceactive film evaporationdone as a twin assembly under vacuum and higher temperature Energy cost can be reduced by more than 20 Additionally the end The prepolymer may be finished in the melt phase with UIFs Products quality is improved due to eliminating intermediate product DISCAGE reactor or in a solidstating unit to obtain the required end lines it offers narrow residence time distribution as well as intensive product features surface renewal and fast reaction di A process column separates side reaction low boilers trom the Commercial plants Four commercial units with a total operating capac iol which is then recycled back to the reaction Spray condensers and ty of 1000 mtod and lot unit of 1 mtod vacuum units recover unreacted feedstock and recycle the diol thus YON MEP ane One PHOT AIT OT AMPS improving the economics of this process References Compact continuous process for high viscosity PBT Poly Economics This new process reduces conversion cost by more than ester 2000 Fifth World Congress Zurich 25 as compared to conventionalhistorical processes by its compact design low energy input shortterm reaction and agitatorless design A product yield of more than 995 is attainable 2R singlestream PET process A new highly economic polyester tech nology International Fibre Journal vol 192004 issue 4 pp 6467 Licensor Uhde InventaFischer Polyalkylene terephthalatesPET PBT PTT PEN continued PROCESSING PetrochemicalProcesses miele IN ce OTe aT AAT TOre ted ROL erste lero home processes index company index Polycaproamide Application Uhde InventaFischers VKtube process polymerizes scap rolactam LC monomer to produce polycaproamide nylon6 chips preparation Polymerization Extraction Drying Description Liquid LC is continuously polymerized in a VKtube 1 in Caprolactam OM V the presence of water stabilizer and modifying additives at elevated T U temperatures The polymerization process has proven to be very reliable Refeeding 1 fn 3 Aa e easy to operate and economical Prepolymerization is available to reduce VKtube 2 4 Y reactor volume for large capacity units The polycaproamide chips are 6 formed from the melt using strand cutters and are conveyed to the ex Chips LH production ps 5 X traction column 2 O The chipscontaining about 9 of monomer and cyclic Ds N Le oligomersare treated with hot water in the extraction column The C Final PAG chips extractables are removed to a very large extent to achieve a good ettes polymer quality and high performance when processed further oy Wet chips are sent to the centrifuge 3 and dried by hot dry nitrogen in a twozone dryer 4 5 The nitrogen gas is regenerated in separate cycles In the bottom zone of the dryer the chips are cooled via a heat exchanger The drying unit can be extended to a solidstate postcondensation ie drying and solidstate postcondensation occurs in one process stage Licensor Uhde InventaFischer Thus high viscosity chips for industrial yarns films and extrusion molded parts can be produced Low utility and energy consumption are achieved by using closed circuits of water and nitrogen as well as by recovering heat The recovery process for the recycling of the extractables reduces raw material cost Extract water is concentrated and directly refed 6 to the polymerization unit Alternatively the concentrated extract is fed to a separate specially designed continuous repolymerization unit Batch and continuous process units are available to meet all potential requirements regarding polymer grades as well as regarding flexibility in output rates and capacities Special attention is devoted during plant design to attain minimal operating expenses for raw material utilities and personnel PROCESSING PetrochemicalProcesses aides ae LeU Reet home processes index company index Polyesters polyethylene terephthalate Dx Reaction vapors Application To produce polyesters for resin and textile applications from v v terephthalic acid PTA or dimethyl terephthalate DMT and diols eth CatalystTi0 RZ ro ylene glycol EG or others using the UlFproprietary fourreactor4R 4k 2 H0 to wee process including DISCAGEfinisher PTA IPA evcaalst finisher diol slurry cl LOTT additives M ay HM Description A slurry composed of PTA and EG or molten DMT and EG Ne ZF PAC is fed to the first esterificationesterinterchange reactor 1 in which ba M main reaction occurs at elevated pressure and temperatures 200C DMT WU 270C Reaction vaporswater or methanolare sent to a lowhigh diol boiler separation column High boilers are reused as feedstock M en The oligomer is sent to a second cascaded stirred reactor 2 Diol recycling LJ SO chips operating at a lower pressure and a higher temperature The reaction Ecterification Prepolycondensation Polycondensation conversion continues to more than 97 Catalyst and additives may be added Reaction vapors are sent to the process column 5 The oligomer is then prepolymerized by a third cascaded reactor 3 under sub atmospheric pressure and increased temperature to obtain a degree of polycondensation 20 Final polycondensation up to intrinsic viscosities of i V 09 is done in the DISCAGEfinisher 4 Pelletizing or direct have been built worldwide Presently 700 mtpd lines are in operation as melt conversion usage is optional singletrain lines including a single finisher EG is recovered by condensing process vapors at vacuum conditions Vacuum generation may be done either by water vapor as a motive Licensor Uhde InventaFischer stream or by the diol EG The average product yield exceeds 99 Economics Typical utility requirements per metric ton of PET are Electricity kWh 550 Fuel oil kg 610 Nitrogen Nm 08 Air Nm 90 Commerical plants Thirteen lines with processing capacities ranging from 100 to 700 mtpd are operating more than 50 polyester CP plants 3 PROCESSING PetrochemicalProcesses miele IN ce a f home processes index company index Polyethylene HDPE Application To produce highdensity polyethylene HDPE using the t raletrereheearent 1 1 rom scrubber stirredtank heavydiluent Hostaen process rom retigerant To scrubber Description The Hostaen process is a slurry polymerization method with ciel Reactors Fosctor Heater two reactors parallel or in series Switching from a single reaction to a re 1 2 3 2 action in cascade enables producing top quality unimodal and bimodal Z la Cyclone Catalyst polyethylene PE from narrow to broad molecular weight distribution storage 41 MWD with the same catalyst vessel Coarse Polymerization occurs in a dispersing medium such as nhexane Ethylene Comonomer Hydrogen 6 a ger using a very highactivity Ziegler catalyst No deactivation and catalyst Prsretneces tte Cooler 7 bed removal is necessary because a very low level of catalyst residue remains Purified hexane Collecting vessel dryer Screen in the polymer For unimodalgrade production the catalyst the dispers ratedhewre f cennenc ing medium monomer and hydrogen are fed to the reactor 1 2 where polymerization occurs In the case of bimodal grade production the recovery To tank farm catalyst is only fed to the first reactor 1 the second step polymerization occurs under different reaction conditions with respect to the first reac tor Also ethylene butene and further dispersing medium are fed to the second reactor 2 Reactor conditions are controlled continuously thus a very highquality PE is manufactured Finally the HDPE slurry from the second reactor is sent to the post applications such as blowmolding large containers small bottles ex reactor 3 to reduce dissolved monomer and no monomer recycling is trusion molding film pipes tapes and monofilaments functional pack needed In the decanter 4 the polymer is separated from the dispers aging and injection molding crates waste bins transport containers ing medium The polymer containing the remaining hexane is dried in a fluidized bed dryer 3 and then nelletized in the extrusion section The Economics Consumption per metric ton of PE based on given product separated and collected dispersing medium of the fluid separation step mix 6 with the dissolved cocatalyst and comonomer is recycled to the po Ethylene and comonomer t 1015 lymerization reactors A small part of the dispersing medium is distilled Steam ke kWh oo to maintain the composition of the diluent Water cooling water AT 10C mt 175 Products The cascade technology enables the manufacturing of tai Commercial plants There are 33 Hostalen plants in operation or under lormade products with a definite MWD from narrow to broad MWD The melt flow index may vary from 02 bimodal product to over 50 sr gsm opi yr mumbercacsng anton een i unimodal product Homopolymers and copolymers are used in various construction with a total licensed capacity of nearly 55 million tpy Indi vidual capacity can range up to 400000 tpy for a singleline installation Licensor Basell Polyolefins Polyethylene HDPE continued iste se cal PetrochemicalProcesses miele IN ce meena JCESSES home processes index company index Polyethylene LDPE Application The highpressure Lupotech TS or TM tubular reactor pro cess is used to produce lowdensity polyethylene LDPE homopolymers and EVA copolymers Singletrain capacity of up to 400000 tpy can be provided Description Ethylene initiator and if applicable comonomers are fed to BS the process and compressed to pressures up to 3100 bar before entering 74 the tubular reactor In the TS mode the complete feed enters the reactor Ethylene y at the inlet after the preheater in the TM mode part of the gas is cooled and quenches the reactor contents at various points of injection Comonomer The polymer properties MI p MWD are controlled by the initiator y pressure temperature profile and comonomer content After the reac tor excess ethylene is recovered and recycled to the reactor feed stream The polymer melt is mixed with additives in an extruder to yield the final product A range of products can be obtained using the Lupotech T process ranging from standard LDPE grades to EVA copolymers or Nbutylac rylate modified copolymer The products can be applied in shrink film extrusion injection molding extrusion blow molding pipe extrusion pipe coating tapes and monofilaments Commercial plants Many Lupotech T plants have been installed after There is no limit to the number of reactor grades that can be pro the first plant in 1955 with a total licensed capacity of 44 million tons duced The product mix can be adjusted to match market demand and Basell operates LDPE plants in Europe with a total capacity of close to 1 economical product ranges Advantages for the tubular reactor design million tpy The newest stateoftheart Lupotech TS unit at Basells site with low residence time are easy and quick transitions startup andshut in Aubette France was commissioned in 2000 with a capacity of 320 down thousand tons it is the largest singleline LDPE plant Reactor grades from MI 015 to 50 and from density 0917 to 0934 gcm with comonomer content up to 30 can be prepared Licensor Basell Polyolefins Economics Consumption per metric ton of PE Ethylene t 1010 Electricity kWh 7001000 Steam t 12 export credit Nitrogen Nm 4 PROCESSING PetrochemicalProcesses miele IN ce a home sprocesses index company index Polyethylene Application New generation Spherilene gasphase technology with sim plified process flow scheme to produce linearlowdensity polyethylene 5 4 LLDPE medium density polyethylene MDPE and highdensity polyeth da oo ylene HDPE of narrow unimodal molecular weight distribution as well Y Y as bimodal molecular weight distribution using only a single Ziegler Y Y Natta titaniumbased catalyst family with full online swing capability 1 if steam Roma without shutdowns r W 9 Description Catalyst components are mixed and fed directly to a pre cw al 3 contact vessel 1 where the catalyst is activated under controlled condi Catalyst 2 PE tions The activated catalyst system flows continuously into the first gas eee L pellets phase reactor GPR 3 A cooler on the circulation gas loop 2 removes Nitrogen Monomers Comonomers see the reaction heat Hydrogen Gas phase Finishing Product containing still active catalyst is continuously discharged Catalyst activation two Pe Teen setup ene tet LD from the first GPR via a proprietary device to a second GPR 5 with simi lar configuration Resultant discharged gas is recovered and no gas from the first GPR enters the second GPR due to a proprietary lockhopper system 4 The second GPR is independently supplied with necessary monomer comonomer and hydrogen to maintain reaction conditions ruly in ndent from the first GPR This gives Soherilene process the abilty WO oreduce tral bimodal HDPE grades and the added freedom to Products Product density range is very wide from approximately 0915 obtain inverse comonomer distribution in the final product by selec gcc LLDPE to 960 gcc HDPE including full access to the MDPE tively feeding comonomer only where necessary Pressure and tempera 9 0930 to 0940 gcc Melt index MI capability ranges from 001 ture in the GPRs are also independently controlled while no additional to 100 g10 min Because of the dual GPR setup Spherilene technol teed of catalytic components to the second GPR Is required ogy enables production of premium bimodal grades MI density In ges The polymer in spherical form with particle size ranging from ap phase with inverse comonomer distribution hitherto available only proximately 05 mm to 3 mm is then discharged in a receiver recovering V4 More Investmentintensive slurry technologies Commercially proven the resultant gas 6 and to a proprietary unit for monomer stripping grades include bimodal HDPE for Pressure PIpe markets with PE100 cer and neutralization of any remaining catalyst activity 7 Residual hydro tification and bimodal HDPE grades for highstrength film markets Tra carbons in the polymer are stripped out and recycled back to reaction ditional HDPE grades for injection molding and extrusion applications a The polymer is dried by a closedloop nitrogen system 8 and with no full range of LLDPE products for cast and blown film extrusion coating and injection molding applications as well as MDPE products for roto molding geomembranes textile and raffia are available Economics Consumption per metric ton of LLDPE Ethylene and comonomer t 1005 Electricity kWh 410 Steam kg 200 Water cooling T 10C mt 150 Commercial plants Licensed from 1992 nine plants using Spherilene process and technology have been licensed with a total capacity of 18 million tpy Singleline capacities in operation range from 100000 to 300000 tpy with current process design available for plants up to 400000 tpy in singleline capacity Licensor Basell Polyolefins Polyethylene continued tistesstttcial PetrochemicalProcesses miele IN ce a f home processes index company index Polyethylene Application The Innovene G gas phase process produces linearlow density polyethylene LLDPE and highdensity polyethylene HDPE us ing either ZieglerNatta chromium or metallocene catalysts YY Logs a Cyclone Description ZieglerNatta and metallocene catalysts are directly injected Beas into the reactor from storage whereas chromium catalysts are injected Reactor following activation of the catalyst via BP proprietary technology The BP catalyst portfolio enables the production of a fullrange of PE products Compressor comaiviene with the same swing reactor using these three main catalyst families GH Hydrogen Accurate control of all the product properties such as density and om melt index is achieved by continuous and automatic adjustment of the Mee A process gas composition and operating conditions The reactor 1 is LJ separator designed to ensure good mixing and a uniform temperature Operating conditions within the bed are mild the pressure is about 20 bar g and Pump the temperature between 75C and 110C Polymer particles grow in the fluidized bed reactor where the fluidization gas is a mixture of ethylene comonomer hydrogen and nitrogen Fine particles leaving the reactor with the exit gas are collected by cyclones 2 which are unique to the Innovene gasphase technology and recycled to the reactor This feature Economics The lowpressure technology and ease of operation ensures ensures that fine particles do not circulate in the reaction loop where that the Innovene process is inherently safe bestinclass environmen they could foul the compressor exchanger and reactor grid The cyclones tally and economically attractive with regard to both investment capex also prevent product contamination during transitions Unreacted gas is and opex cooled 3 and separated from any liquid 4 compressed 5 and returned to the reactor maintaining the growing polymer particles at the desired Products A wide range of LLDPE and HDPE products can be produced temperature Catalysts are incorporated into the final product without within the same reactor LLDPE is used in film injection molding and any catalyst removal step extrusion applications and can be made using either butene or hexene The reactor and almost all other equipment is made from carbon as the comonomer Narrow molecular weight HDPE provides superior in steel Polymer powder is withdrawn from the reactor via a proprietary jection molding and rotational molding grades whereas broad molecular lateral discharge system and separated from associated process gas in Weight HDPE is used for blow molding pipe film and other extrusion a simple degassing stage using hot recirculating nitrogen The powder applications is then pneumatically conveyed to the finishing section where additives are incorporated before pelletization and storage Commercial plants Thirtyfive reactor lines are operating in design or under construction worldwide representing around 6 MMtpy produc tion with capacities ranging from 50000 tpy to 350000 tpy Designs up to 450000 tpy are also available Licensor BP Polyethylene continued PROCESSING PetrochemicalP eee C ICAIF TOCESSES home processes index company index Polyethylene Application To produce lowdensity polyethylene LDPE homopolymers and EVA copolymers using the highpressure free radical process Large scale tubular reactors with a capacity in the range of 130400 Mtpy as well Meee Compressor as stirred autoclave reactors with capacity around 100 Mtpy can be used comonomers ee Description A variety of LDPE homopolymers and copolymers can be O 6 produced on these large reactors for various applications including films Compressors Us molding and extrusion coating The melt index polymer density and molecular weight distribution are controlled with temperature profile Separators pressure initiator and comonomer concentration Autoclave reactors can give narrow or broad molecular weight distribution depending on the selected reactor conditions whereas tubular reactors are typically used to produce narrow molecular weight distribution polymers Silo Gaseous ethylene is supplied to the battery limits and boosted to 300 bar by the primary compressor This makeup gas together with the recycle gas stream is compressed to reactor pressure in the secondary compressor The tubular reactors operate at pressures up to 3000 bar whereas autoclaves normally operate below 2000 bar The polymer is separated in a high and lowpressure separator nonreacted gas is recycled from both separators Molten polymer from the lowpressure Economics separator is fed into the extruder polymer pellets are then transferred Raw materials and utilities per metric ton of pelletized polymer to storage silos Ethylene tonton 1008 The main advantages for the highpressure process compared to cect wn oSe other PE processes are short residence time and the ability to switch from Nitrogen Nm3t 5 homopolymers to copolymers incorporating polar comonomers in the same reactor The highpressure process produces longchain branched Commercial plants Affiliates of ExxonMobil Chemical Technology Licens products from ethylene without expensive comonomers that are required ing LLC operate 22 highpressure reactors on a worldwide basis with a by other processes to reduce product density Also the highpressure capacity of approximately 14 MMtpy Homopolymers and a variety of process allows fast and efficient transition for a broad range of polymers copolymers are produced Since 1996 ExxonMobil Chemical Technol ogy Licensing LLC has sold licenses with a total installed capacity either Products Polymer density in the range 0912 up to 0935 for homo Vinylacetate content up 1030 W196 i ae in operation or under construction of approximately 1 million tpy Licensor ExxonMobil Chemical Technology Licensing LLC Polyethylene continued PROCESSING PetrochemicalProcesses miele IN ce meena MO etsyslete eae home processes index company index Polyethylene Applications To produce high density polyethylene HDPE and medium Polymerization Separation Pelletizing Silo storage density polyethylene MDPE under lowpressure slurry process CX and drying Stabilizer and packing process Ethylene 2 Catalyst A Description The CX process uses two polymerization reactors in se 1 ries The products have bimodal molecularweight distribution MWD 3 a where MWD and composition distribution is freely and easily controlled VY by adjusting the operating conditions of two reactors without changing oo Sa the catalyst 4 This process produces a wide melt index range by applying inno vative catalyst chemistry combined with a sophisticated polymerization process An allround catalyst and simple polymerization operation pro vide easy product changeover that reduces transition time and yields 4 negligible offspec product from the transition Mitsui has also devel oped new catalyst that contributes better morphology of the polymer powder and ethylene consumption Ethylene hydrogen comonomer and a superhigh activity cata lyst are fed into the reactors 1 Polymerization reaction occurs under a slurry state The automatic polymer property control system plays Economics Typical consumption per metric ton of natural HDPE pellets very effective role in productquality control Slurry from the reactors is pumped to the separation system 2 The wetcake is dried into powder Ethylene and comonomer kg 1004 Electricity kWh 345 in the dryer system 3 As much as 90 of the solvent is separated from Steam kg 340 the slurry and is directly recycled to the reactors without any treatment Water cooling t 190 The dry powder is pelletized in the pelletizing system 4 along with re quired stabilizers Commercial plants Fortyone reaction lines of CX process are in opera tion or construction worldwide with a total production capacity of over Products Broad range of homopolymer and copolymer can be pro 45 million tpy duced including PE100 pipe grade Licensor Mitsui Chemicals Inc Melt index 001 to 50 Molecularweight distribution Freely controlled from Comonomer distribution narrow to very wide Density 093 to 097 tistesstttcial PetrochemicalProcesses miele IN ce meena JCESSES home processes index company index Polyethylene Application The SCLAIRTECH technology PE process can produce linearlowdensity mediumdensity and highdensity polyethylene PE with narrow to broad molecular weight distribution using either Ziegler Comonomers es Natta ZN or proprietary singlesite catalyst SSC Description Ethylene and comonomer are dissolved in solvent then fed into a reactor Butene1 octene1 or both together can be used as co eo monomer The reactor system operates in a solution phase and due to inherent low residence time less than 2 minutes it offers a tremen Purification Reactor Reactor dous flexibility for grade transitions and significant versatility for meet Ethylene feed system ing product needs of a diverse market cee To finishing High conversions maximize production and eliminate any potential Woy eee for runaway reactions A hydrocarbon solvent is used to keep the con Pellitizer tents of the reactor in solution and also aids in heat removal The solvent is flashed and recovered along with the energy captured from the heat of reaction and circulated back to the reactor Molten polymer is sent to a simple extruder and pelletizer assembly Products SCLAIRTECH process can produce PE products with density range of 09050965 kgm melt index MI from 02 to in excess of nomers such as octene1 allows producers to participate in premium 150 and narrow to broad molecular weight distribution MWD This markets resulting in higher business returns allows producers to participate in the majority of the polyethylene mar ket segments including among low medium and highdensity films Commercial plants The first SCLAIRTECH plant was built in 1960 Cur rotational injection and blow molding applications rently more than 12 plants worldwide are either operating in design Products made with this technology offer exceptional quality as OF under construction with this technology representing about 3 million measured by low gel superior opticals and lottolot consistency along tpy total capacity with high performance characteristics for demanding applications Licensor NOVA Chemicals International SA Economics This technology offers advantaged economics for producers isc aiRTECH is a trademark of NOVA Chemicals desirous of participating in a broad range of market segments andor niche applications due to its ability to transition quickly and cover a large product envelope on a single line An ability to incorporate como iste se cal PetrochemicalProcesses miele IN ce meena JCESSES home processes index company index Polyethylene Application To produce linear lowdensity polyethylene LLDPE to high density polyethylene HDPE using the lowpressure gasphase UNIPOL PE process Description A wide range of polyethylenes is made in a gasphase flu idizedbed reactor using proprietary solid and slurry catalysts The prod 3 a uct is in a dry freeflowing granular form substantially free of fines as it leaves the reactor and is converted to pellet form for sale Melt index st and molecular weight distribution are controlled by selecting the proper Ethylene and aaa ra catalyst type and adjusting operating conditions Polymer density is con trolled by adjusting comonomer content of the product High produc tivity of conventional and metallocene catalysts eliminates the need for Polyethylene to catalyst removal resin loading The simple and direct nature of this process results in low investment and operating costs low levels of environmental pollution minimal potential fire and explosion hazards and easy operation and maintenance Gaseous ethylene comonomer and catalyst are fed to a reactor 1 containing a fluidized bed of growing polymer particles and operating near 25 kgcm2 and approximately 100C A conventional singlestage ow or broad Melt index may be varied from less than 01 to greater centrifugal compressor 2 circulates reaction gas which fluidizes the than 200 Grades suitable for film blowmolding pipe rotomolding reaction bed provides raw material for the polymerization reaction and and extrusion applications are produced removes the heat of reaction from the bed Circulating gas is cooled in one oo a conventional heat exchanger 3 Commercial plants Ninetysix reaction lines are in operation under con The granular product flows intermittently into product discharge struction or in the design phase worldwide with singleline capacities tanks 4 where unreacted gas is separated from the product and ranging from 40000 tpy to more than 450000 tpy returned to the reactor Hydrocarbons remaining with the product are Licgensor Univation Technologies removed by purging with nitrogen The granular product is subsequently pelletized in a lowenergy system 5 with the appropriate additives for each application Products Polymer density is easily controlled from 0915 to 0970 gcm i Depending on catalyst type molecular weight distribution is either nar tistesstttcial PetrochemicalProcesses Polypropylene Application Spheripo process technology produces propylenebased polymers including homopolymer PP and many families of random and heterophasic impact and specialty impact copolymers er 4 Description In the Spheripo process homopolymer and random co polymer polymerization takes place in liquid propylene within a tubular Catalyst sf loop reactor 1 Heterophasic impact copolymerization can be achieved Steam by adding a gasphase reactor 3 in series Removal of catalyst residue and amorphous polymer is not required Propylene Unreacted monomer is flashed in a twostage pressure system 2 4 and recycled back to the reactors This improves yield and minimizes A energy consumption Dissolved monomer is removed from the polymer i steam Ft P by a steam sparge 5 The process can use lowerassay chemicalgrade yrene to storage propylene 94 or the typical polymerizationgrade 995 Yields Polymer yields of 4000060000 kgkg of supported catalyst are obtained The polymer has a controlled particle size distribution and an isotactic index of 9099 Economics The Spheripo process offers a broad range of products with excellent quality ve lowcopital and operating monte P Commercial plants Soheripo technology is used for about 50 of the total global PP capacity There are 94 Spheripol process plants operating Consumption per metric ton of PP aye Propylene and comonomer t 10021005 worldwide with total capacity of about 17 million tpy Singleline design Catalyst kg 00160025 capacity is available in a range from 40000 to 550000 tpy Electricity kWh 80 Steam kg 280 Licensor Basell Polyolefins Water cooling mt 90 In case of copolymer production an additional 20 kWh is required Products The process can produce a broad range of propylenebased polymers including homopolymer PP various families of random copo lymers and terpolymers heterophasic impact and speciality impact co polymers up to 25 bonded ethylene as well as highstiffness high clarity copolymers a Ce Oe a PROCESSING PetrochemicalProcesses miele IN ce a home processes index company index Polypropylene Application To produce polypropylenebased polymers including ho pees Additional Finishing steaming mopolymer polypropylene random heterophasic impact and specialty copolymerization drying and additivation dual composition copolymers using Spherizone process technology Ps PS Description The Spherizone process is Basells new proprietary gasloop y ov OH reactor technology based on a MultiZone Circulating Reactor MZCR EB 0 concept Inside the reactor 1 the growing polymeric granule is continu Additives ously recirculating between two interrelated zones where two distinct Reagent and different fluodynamic regimes are realized 6 steam 3 In the first zone 1a the polymer is kept in a fast fluidization regime Catalyst and E when leaving this zone the gas is separated and the polymer crosses the cocatalyst second zone 1b in a packed bed mode and is then reintroduced in the Reagent o stogen first zone A complete and massive solid recirculation is obtained be cw Pellets tween the two zones Ethylene The fluodynamic peculiar regime of the second zone where the polymer enters as dense phase in plug flow altering the gas compo sition with respect to the chain terminator hydrogen and to the co monomer This is accomplished by injecting monomers from the exter nal system 2 in one or more points of the second zone 1b and so two or more different polymers MFR andor comonomer type and content to the reaction While the polymer is dried by a closedloop nitrogen can grow on the same granule system 7 and now free from volatile substances the polymer is sent to While the granules recycle through the multiple zones different additives incorporation step 8 polymers are generated in an alternate and cyclic way via continuous E ics R terial and utilit tric t f polymerization This allows the most intimate mixing of different poly conomics aw Material ang uum reduirements per metic ton mers giving a substantial homogeneity of the final product product Unreacted monomer is flashed at intermediate pressure 3 and re Propylene plus comonomer for copolymers kg 10021005 cycled back to the loop reactor while polymer can be fed to a fluidized aan 9 vb Wea gasphase reactor 4 operated in series optional where additional co Steam ke 120 polymer can be added to the product from the gas loop Water cooling m2 85 From the intermediate separatorsecond reactor the polymer is dis In case of high impact copolymer production an additional 20 kWh is required charged to a receiver 5 the unreacted gas is recovered while the poly mer is sent to a proprietary unit for monomer steam stripping and cata lyst deactivation 6 The removed residual hydrocarbons are recycled Products The process can produce a broad range of propylenebased polymers including mono and bimodal mediumwidevery wide mo lecular weight distribution homopolymer PP high stiffness homopoly mers random copolymers and terpolymers highclarity random copo lymers as well as two compositions homopolymerrandom copolymer twinrandom or randomheterophasic copolymer Conventional hetero phasic impact copolymers with improved stiffnessimpact balance can be produced with the second additional gas phase reactor with ethyl enepropylene rubber content up to 40 Commercial plants A retrofitted 160000 tpy plant is in operation at the Basell site in Brindisi since 2002 and 3 licenses for a total capacity of 1 million ton have been granted during 2004 The largest unit license is a 450000tpy singleline plant Technology owner Basell Polyolefins Polypropylene continued tistesstttcial PetrochemicalProcesses miele IN ce a f home processes index company index Polypropylene Application To produce polypropylene PP homopolymer random co eae 2 Condenser pronvienetiecrcle polymer and impact copolymer using the BP Innovene gasphase pro to reactor cess with proprietary 4th generation supported catalyst Cece e re Powder Catalyst W CoD EB Description Catalyst in mineraloilslurry is metered into the reactor to Propylene separation Propylene oe recovery gether with cocatalyst and modifier The proprietary supported catalyst developed by BP has control morphology superhigh activity and very or flare pe Power high sterospecifity The resulting PP product is characterized by narrow 2 Condenser deactivation particle size distribution good powder flowability minimum catalyst powder residues noncorrosiveness excellent color and low odor transfer Propylene HEE The horizontal stirredbed reactor 1 is unique in the industry in that aes ee it approaches plugflow type of performance which contributes to two 4 Ethylene major advantages First it minimizes catalyst bypassing which enables the L Pode process to produce very highperformance impact copolymer Second it L 7 Q makes product transitions very quick and sharp which minimizes offspec transition materials The reactor is not a fluidized bed and powder mixing is accomplished by very mild agitation provided by a proprietarydesigned horizontal agitator Monomer leaving the reactor is partially condensed 2 and recycled The condensed liquid together with fresh makeup Products A wide range of polypropylene products homopolymer ran monomer is sprayed onto the stirred reactor powder bed to provide dom copolymer and impact copolymer can be produced to serve many evaporative cooling remove the heat of polymerization and control the applications including injection molding blow molding thermoform bed temperature Uncondensed gas is returned to the reactor ing film extrusion sheet and fiber Impact copolymer produced using For impact copolymer production a second reactor 4 in series is this process exhibits a superior balance of stiffness and impact resistance required A reliable and effective gaslock system 3 transfers powder over a broad temperature range from the first homopolymer reactor to the second copolymer reactor and prevents cross contamination of reactants between reactors Commercial plants Fourteen plants are either in operation or in de This is critically important when producing the highest quality impact signconstruction worldwide with capacities ranging from 65000 to copolymer In most respects the operation of the second reactor system 350000 mtpy is similar to that of the first except that ethylene in addition to propylene is fed to the second reactor Powder from the reactor is transferred and Licensor BP depressurized in a gaspowder separation system 5 and into a purge column 6 for catalyst deactivation The deactivated powder is then pelletized 7 with additives into the final products PROCESSING PetrochemicalProcesses miele IN ce a home processes index company index Polypropylene Application A process to produce homopolymer polypropylene and a mropvl lel ethylenepropylene random and impact copolymers using Chisso Gas Y eeEr Phase Technology utilizing horizontal plugflow reactor Propylene Gatalet Powdergas Description The process features a horizontal agitated reactor and a Separation highperformance catalyst specifically developed by the licensor The catalyst has a controlled morphology very high activity and very high oe selectivity The process provides low energy consumption superior Gas mee ethylenepropylene impact copolymer properties minimum transition oe products high polymer throughput and a high operating factor Each ae 2 MS Moist nitrogen process step has been simplified consequently the technology offers a ennites a low initial capital investment and reduced manufacturing costs while providing product uniformity excellent quality control and wide range cr 6 Pelletized of polymer design especially for comonomer products ai product Particles of polypropylene are continuously formed at low pressure in the reactor 1 in the presence of catalyst Evaporated monomer is partially condensed and recycled The liquid monomer with fresh propylene is sprayed onto the stirred powder bed to provide evapora tive cooling The powder is passed through a gaslock system 2 to a second reactor 3 This acts in a similar manner to the first except that Chisso offers processing designs for singleproduction with capacities ethylene as well as propylene is fed to the system for impact copoly reaching 400000 tpy mer production The horizontal reactor makes the powder residence Licensor Japan Polypropylene Cor time distribution approach that of plugflow The stirred bed is well Th i to i yP thi tech i ven Chisso to Ja suited to handling some high ethylene copolymers that may not flow TIQITS EO TICENSE EUS TECNNOVOGY WETE GIVEN TTOM KNISSO 10 wa or fluidize well pan Polypropylene Corp which is a PP joint venture between Chisso The powder is released periodically to a gaspowder separation sys and Mitsubishi Chemical Corp tem 4 It is depressurized to a purge column 5 where moist nitro gen deactivates the catalyst and removes any remaining monomer The monomer is concentrated and recovered The powder is converted into a variety of pelletized resins 6 tailored for specific market applications Commercial plants Ten polypropylene plants are in operation or under construction with capacities ranging from 65000 tpy to 360000 tpy PROCESSING PetrochemicalProcesses at ACIASALILL PLUCIS home processes index company index Polypropylene Applications To produce polypropylene PP including homopolymer random copolymer and impact copolymer ee ene cer Description The process with a combination of the most advanced Stabilizer highyield and highstereospecificity catalyst is a nonsolvent nondeash ing process It eliminates atactic polymers and catalyst residue removal Propvlene i a The process can produce various grades of PP with outstanding product ey S quality Polymer yields of 20000 to 100000 kgkg of supported catalyst rie Bl are obtained and the total isotactic index of polymer can reach 98 to Y al 99 With new catalysts based on diether technology 5th generation O catalyst RKCatalyst and RHCatalyst wider meltindex ranged polymers can be produced compare with those produced with 4th generation catalyst due to the high hydrogen response of RKRHCatalyst shipping The reactor polymer has narrow and controlled particle size distribution that stabilizes plant operation and also permits easy shipment as powder Due to the proprietary design of the gasphase reactor no fouling is observed during the operation and consequently reactor cleaning after producing impact copolymer is not required In addition combination of the flexibility of the gasphase reactor and Product The process can produce a broad range of polypropylene poly highperformance catalysts allow processing impact copolymer with a mers including homopolymer random copolymer and impact copoly highethylene content mer which become highquality grades that can cover various applica In the process homopolymer and random copolymer polymerization tions occurs in the looptype reactor or vesseltype reactor 1 For impact copolymer production copolymerization is performed in a gasphase Economics Typical consumption per metric ton of natural propylene reactor 2 after homopolymerization The polymer is discharged from homopolymer pellets a gasphase reactor and transferred to the separator 3 Unreacted gas Propylene and ethylene for copolymerkg 1005 accompanying the polymer is removed by the separator and recycled to Electricity kWh 320 the reactor system The polymer powder is then transferred to the dryer Steam kg 310 a Water cooling t 100 system 4 where remaining propylene is removed and recovered The dry powder is pelletized by the pelletizing system 5 along with required stabilizers Commercial plants Twentyfive reactor lines are in operation engineer ing design or construction worldwide with a total production capacity of over 25 MMtpy Licensor Mitsui Chemicals Inc Polypropylene continued PROCESSING PetrochemicalProcesses miele IN ce a f home processes index company index Polypropylene Application To produce homopolymer random copolymer and impact copolymer polypropylene using the Dow gasphase UNIPOL PP process Description A wide range of polypropylene is made in a gasphase flu 6 idizedbed reactor using proprietary catalysts Melt index isotactic level 3 and molecular weight distribution are controlled by utilizing the proper q catalyst adjusting operating conditions and adding molecularweight control agents Random copolymers are produced by adding ethylene or butene to the reactor Ethylene addition to a second reactor in series is used to produce the rubber phase of impact copolymers The UNIPOL PP process simple yet capable design results in low Polypropylene to investment and operating costs low environmental impact minimal po Propylene la resin loading tential fire and explosion hazards and easy operation and maintenance oJoOorv To produce homopolymers and random copolymers gaseous propylene comonomer and catalyst are fed to a reactor 1 containing a fluidized bed of growing polymer particles and operating near 35 kgcm and ap proximately 70C A conventional singlestage centrifugal compressor 2 circulates the reaction gas which fluidizes the reaction bed provides raw materials for the polymerization reaction and removes the heat of the reaction from the bed Circulating gas is cooled in a conventional nitrogen Granular products are pelletized in systems available from mul heat exchanger 3 Granular product flows intermittently into product tiple vendors 9 Dow has ongoing development programs with these discharge tanks 4 unreacted gas is separated from the product and suppliers to optimize their systems for UNIPOL PP resins guaranteeing returned to the reactor low energy input and high product quality Controlled rheology high To make impact copolymers the polypropylene resin formed in meltflow grades are produced in the pelleting system through the ad the first reactor 1 is transferred into the second reactor 5 Gaseous dition of selected peroxides propylene and ethylene with no additional catalyst are fed into the sec ond reactor to produce the polymeric rubber phase within the existing Products Homopolymers can be produced with melt flows from less polypropylene particles The second reactor operates in the same man than 01 to 3000 dgmin and isotactic content in excess of 99 Ran ner as the initial reactor but at approximately half the pressure with a 0m copolymers can be produced with up to 12 wt ethylene or up to centrifugal compressor 6 circulating gas through a heat exchanger 7 21 wt butene over a wide melt flow range 01 to 100 dgmin A and back to the fluidbed reactor Polypropylene product is removed by product discharge tanks 8 and unreacted gas is returned to the reactor h Hydrocarbons remaining in the product are removed by purging with full range of impact copolymers can be polymerized with excellent stiff ness for even the most demanding applications Products from narrow to broad molecularweight distribution can be manufactured in grades proven advantage for film injection molding blow molding extrusion and textile applications Commercial plants Nearly 40 reaction lines are in operation with ca pacities ranging from 80000 to 260000 tpy and plants in design up to 500000 tpy Total worldwide production of polypropylene with this technology is nearly 6 million tpy Licensor The Dow Chemical Co Univation Technologies is the licensor of the UNIPOL PE process Polypropylene continued tistesstttcial PetrochemicalProcesses PROCESSING JLGOIOU home processes index company index Polystyrene expandable Application To produce expandable polystyrene EPS via the suspension Initiator and Suspending agents process using BP ChemicalsABB Lummus Global technology chemical additives To atmosphere Description The BPLummus styrene polymerization technology for the iv Centrifuge io manufacture of regular and flameretardant grades of EPS is a onestep 3 Dryer batch suspension reaction followed by continuous dewatering drying Styrene L 4 Screening and size classification Water Slender Additives Styrene monomer water initiators suspending agents nucleating Reactor agents and other minor ingredients are added to the reactor 1 The contents are then subjected to a timetemperature profile under agitation Slurry Effluent The suspending agent and agitation disperse the monomer to form beads eo To At the appropriate time a premeasured quantity of pentane is introduced Nin y Airy At eae into the reactor Polymerization is then continued to essentially 100 ie a conversion After cooling the EPS beads and water are discharged to a C eee holding tank 2 From this point the process becomes continuous The beadwater Slurry is centrifuged 3 where most of the mother liquor is removed The beads are conveyed to a pneumatic dryer 4 where the remaining moisture is removed The dry beads are then screened 5 yielding as many as four product Commercial plants Three commercial production units are in operation cuts External lubricants are added in a proprietary blending operation one in France one in Germany and one in China for a total capacity of 6 and the finished product is conveyed to shipping containers 200000 metric tons Economics The BPLummus process is one of the most modern technolo Licensor ABB Lummus GlobalBP Chemicals gies for EPS production Computer control is used to produce product uni formity while minimizing plant energy requirements BP provides ongoing process research for product improvement and new product potential Raw materials and utilities based on one metric ton of EPS Styrene and pentane kg 10001015 Process chemicals kg 2549 Demineralized water kg 1000 Electricity kWh 150 Steam mt 042 Water cooling m 120 iste se cal PetrochemicalProcesses miele IN ce a f home processes index company index Polystyrene high impact Application To produce a wide range of general purpose and high impact polystyrenes PS via the bulk continuous process using the BP Styrene Preheater ChemicalsABB Lummus Global technology q Styrene recycle rinder iti Description The production of general purpose PS GPPS and high catia impact PS HIPS is essentially the same with the exception of the initial Styrene purge rubberdissolution step for HIPS The production of HIPS begins with the granulating and dissolving 8 of rubber and other additives in styrene monomer 1 and then eee transferring the rubber solution to a storage tank 2 For general Been eee Prepoly 6 2 purpose product controlled amounts of ingredients are fed directly to storage reactor a Pevol the feed preheater 3 From this point on the production steps for GPPS and HIPS are the Storage same The feed mixture is preheated 3 and continuously fed to the 9 prepolymerizer 4 where the rubber morphology is established Following prepolymerization the polymer mixture is pumped to the polymerization reactor 5 of proprietary design At the exit of the reactor the polymerization is essentially complete The mixture is then preheated 6 in preparation for devolatization Raw materials and utilities based on one metric ton of polystyrene The devolatilizer 7 is held under a very high vacuum to remove GPPS HIPS Styrene and mineral oil kg 1011 937 unreacted monomer and solvent from the polymer melt The monomer Rubber kg 73 is distilled in the styrene recovery unit 8 and recycled back to the Additives 1 2 prepolymizer The polymer melt is then pumped through a die head Electricity kWh 97 110 9 to form strands a waterbath 10 to cool the strands a pelletizer Fuel 10 kcal 127 127 11 to form pellets and is screened to remove large pellets and fines eee cooing m The resultant product is airconveyed to bulk storage and packaging eam XS facilities Commercial plants Plants in France Germany and Sweden are in op Economics The BPLummus process offers one of the most modern eration with a total capacity of approximately 450000 mtpy of GPPS oor and HIPS Another 300000 mtpy GPPS and HIPS unit will start up in technologies for GPPS and HIPS production A broad product line is China in 2005 available with a consistently high product quality BP provides ongoing process research for product improvement and new product potential i Licensor ABB Lummus GlobalBP Chemicals Polystyrene high impact continued PROCESSING PetrochemicalProcesses E home processes index company index Polystyrene general purpose GPPS Application To produce a wide range of general purpose polystyrene GPPS with excellent high clarity and suitable properties to process PS Styrene foam via direct injection extrusion by the continuous bulk polymeriza Solvent tion process using Toyo Engineering Corp TECMitsui Chemicals Inc Additives en technology React Description Styrene monomer a small amount of solvent and additives Devolatilizers are fed to the specially designed reactor 1 where the polymerization is Condensers carried out The polymerization temperature of the reactor is carefully Nactn controlled at a constant level to keep the desired conversion rate The Recovered monomer 4 heat of polymerization is easily removed by a specially designed heat transfer system Storage Pelletizer Die head At the exit of the reactor the polymerization is essentially complete 6 5 The mixture is then preheated 2 and transferred to the devolatilizers 3 where volatile components are separated from the polymer solution by evaporation under vacuum The residuals are condensed 4 and recycled back to the process The molten polymer is pumped through a die 5 and cut into pellets by a pelletizer 6 Economics Basis 50000 mtpy GPPS US Gulf Coast Investment million US 14 Raw materials consumption per one metric ton of GPPS kg 1009 Utilities consumption per one metric ton of GPPS US 105 Installations Six plants in Japan Korea China India and Russia are in operation with a total capacity of 200000 metric tpy Licensor Toyo Engineering CorpTEC Mitsui Chemicals Inc a COC Cee ey a iste se cal PetrochemicalProcesses miele IN ce a Mele erste home processes index company index Polystyrene highimpact HIPS Application To produce a wide range of highimpact polystyrene HIPS with wellbalanced mechanical properties and processability via Solvent the continuous bulk polymerization process using Toyo Engineering Additives mrepolymer Corp TECMitsui Chemicals Inc technology The process has a swing ee production feature and is also capable of producing general purpose 5 polystyrene GPPS Preheaters Description Styrene monomer ground rubber chips and small amount of additives are fed to the rubber dissolver 1 The rubber chips com 2 Devolatilizers pletely dissolved in styrene This rubber solution is sent to a rubbersolu Weche aaouie Condensers tionfeed tank 2 The rubber solution from the tank is sent to the pre Vacuum polymerizer 3 where it is prepolymerized and the rubber morphology ee ee 7 is established The prepolymerized solution is then polymerized in specially de Storage Pelletizer Die head signed reactors 4 arranged in series The polymerization temperature of the reactors is carefully controlled at a constant level to maintain the desired conversion rate The heat of the polymerization Is easily removed by a specially designed heattransfer system The polymerization product a highly viscous solution is preheated 5 and transferred to the devolatilizers 6 Volatile components are separat ed from the polymer solution by evaporation under vacuum The residu als are condensed 7 and recycled to the process The molten polymer is pumped through a die 8 and cut into pellets by a pelletizer 9 Economics Basis 50000metric toy HIPS unit US Gulf Coast Investment million US 21 Raw materials consumption per one metric ton of HIPS kg 1009 Utilities consumption per one metric ton of HIPS US 8 Installations Six plants in Japan Korea China and India are in operation with a total capacity of 190000 metric tpy Licensor Toyo Engineering Corp TECMitsui Chemicals Inc PROCESSING PetrochemicalProcesses miele IN ce a f home processes index company index Propylene and isobutylene Application Technology for dehydrogenation of propane or isobutane Propane to make highpurity propylene or isobutylene The CATOFIN process a uses specially formulated proprietary catalyst from SUdChemie Description The CATOFIN reaction system consists of parallel fixedbed o On reheat Exhaust air reactors and a regeneration air system The reactors are cycled through a sequence consisting of reaction regeneration and evacuationpurge Fuel gas Steam steps Multiple reactors are used so that the reactor feedproduct system Propylene and regeneration air system operate in a continuous manner Fresh propane feed is combined with recycle feed from the bottom of the product splitter 6 vaporized raised to reaction temperature in 13 4 a charge heater 1 and fed to the reactors 2 Reaction takes place at vacuum conditions to maximize feed conversion and olefin selectivity C and After cooling the reactor effluent gas is compressed 3 and sent Recycle propane meaner to the recovery section 4 where inert gases hydrogen and light hydrocarbons are separated from the compressed reactor effluent and C and heavier are rejected The ethane propane and propylene components are then sent to the product purification section deethanizer 5 and product splitter 6 where propylene product is separated from unreacted propane The propane is recycled to the reactors Economics Where a large amount of low value LPG is available the After a suitable period of onstream operation feed to an individual CATOFIN process is the most economical way to convert it to high value reactor is discontinued and the reactor is reheatedregenerated Reheat product The large singletrain capacity possible with CATOFIN units the regeneration air heated in the regeneration air heater 7 is passedthrough largest to date is for 455000 mtpy propylene minimizes the investment the reactors The regeneration air serves to restore the temperature profile costmt of product of the bed to its initial onstream condition in addition to burning coke Investment ISBL Gulf Coast USmtpy 400500 off the catalyst When reheatregeneration is completed the reactor is reevacuated for the next onstream period Raw material and utilities per mt of propylene Propane mt 117118 Yields and product quality Propylene produced by the CATOFIN process rd 5 530 is typically used for the production of polypropylene where purity de mands are the most stringent 9995 The consumption of propane Of 100 is 117 metric ton mt per mt of propylene product i Commercial plants Currently 11 CATOFIN dehydrogenation plants are onstream producing over 2600000 mtpy of isobutylene and 700000 mtpy of propylene Licensor ABB Lummus Global Propylene and isobutylene continued PROCESSING PetrochemicalProcesses at POCHIEMICAITTOCESSES home processes index company index Propylene Application To produce propylene from ethylene and butenes using Lummus olefin conversion technology OCT Other OCT process con Guard bed Metathesis reactor Ethylene column Propylene column figurations involve interconversion of light olefins and production of Ethylene feed Recycle ethylene Lights purge C5Cs monoolefins Propylene Description Ethylene feedstream plus recycle ethylene and butenes nD LA feedstream plus recycle butenes are introduced into the fixedbed OTT metathesis reactor The catalyst promotes reaction of ethylene and 2 butene to form propylene and simultaneously isomerizes 1butene to 2butene Effluent from the metathesis reactor is fractionated to yield highpurity polymerizationgrade propylene as well as ethylene and butenes for recycle and small byproduct streams Due to the unique C plus nature of the catalyst system the mixed C feed stream can contain a ast significant amount of isobutylene without impacting performance of C feed the OCT process A variation of OCTAutomet Technologycan be used to generate ethylene propylene and the comonomerhexene 1by metathesis of nbutenes Yields OCT process selectivity to propylene is typically greater than 98 Overall conversion of nbutenes is 8592 Ethylene and butenes feed Cooling duty Btu 1033 streams can come from steam crackers or many refinery sources and in Nitrogen scf 21 varying concentrations Alternatively butenes can come from ethylene Catalyst cost est per yr US 325000 dimerization which is also licensed by Lummus Maintenance per yr as of investment 15 In the Automet Technology butenes yield about 10 ethylene 38 propylene and 47 hexene1 The balance is C and heavier material Commercial plants Lyondell Petrochemical Co Channelview Texas uses both the OCT technology and ethylene dimerization technology Economics Based ona 300000mtpy propylene plant US Gulf Coast two other plants have used related technology Two plants have recently mid2000 assuming 86 nbutenes in teedstream Started up a 690 MM lbyr unit for BASF Fina Petrochemical in Port Investment total direct field cost US205 million Arthur Texas and a 320 MM lbyr unit for Mitsui Petrochemical in Osa Utilities required per pound of product ka Japan Six other plants are under design or construction bringing Fuel gas fired Btu 340 Electricity kWh 36 Steam 50 psig saturated Btu 704 the worldwide propylene capacity via OCT to over 2 million mtpy The Automet Technology is in operation on a semicommercial scale at the Tianjin Petrochemical Co in Tianjin China Licensor ABB Lummus Global Propylene continued PROCESSING PetrochemicalProcesses miele IN ce meena eerste lets ae home processesindex company index Propylene Application To produce polymergrade propylene plus either an isobutylenerich stream or MTBE by upgrading lowvalue pyrolysis Cy Methanol cuts or butenerich streams via selective hydrogen and Meta4 process Hydrogen for MTBE only es This process is particularly profitable when butadiene markets are isobutene weak and propylene demand is strong Raw Cs from rich cut steam cracker Butadiene or MTBE Description Crude C streams are converted into propylene and an ee ysctoninen seat isobutylenerich stream in three IFP process steps 1 butadiene and Cy acetylenes selective hydrogenation and butenes hydroisomerization 2 2Butenesrich isobutylene removal via distillation or MTBE production and 3 metath esis Meta4 Ethylene Propylene The hydroisomerization step features complete C acetylenes and Unreacted Cs butadiene conversion to butenes maximum 2butenes production 3 and C flexibility to process different feeds polymerfree product and no residual hydrogen The second step separates isobutylene either by conventional distillation or by reacting the isobutylene with methanol to produce MTBE The CCR Meta4 process features are a hard highly active and robust catalyst low catalyst inventory low operating temperature and pressure outstanding yields liquidphase operation and continuous Of movingbed continuous catalyst regeneration technology is industri operation and catalyst regeneration ally proven in Axens CCR Octanizing and Aromizing reformers Yields Process selectivity to propylene is typically greater than 98 Reference Chodorge J A J Cosyns D Commereuc Q Debuisschert Overall conversion of 2butenes can reach 90 and P Travers Maximizing propylene and the Meta4 process Oil Gas 2000 Economics ISBL 2004 investment for a Gulf Coast location of a Meta A process producing 180000 tpy propylene is US19 million Typical Licensor Axens Axens NA operating cost is 18 per metric ton of propylene Commercial plants Over 100 C hydrogenation units have been built using Axens technology The CCR Meta4 technology has been devel oped jointly with the Chinese Petroleum Corp and demonstrated on real feedstock at Kaohsiung Taiwan industrial complex The same type PROCESSING PetrochemicalP aoe i a OLENA re AO cis toot home processes index company index Propylene Application To produce propylene and ethylene from lowvalue light hydrocarbon streams from ethylene plants and refineries with feeds in the carbon number range of Cy to Cg such as steam cracker CCs Feed olefins catcracker naphthas or coker gasolines ccna Cicareyde IF Light gas upertiex Description The SUPERFLEX process is a proprietary technology pat converter c eer ented by ARCO Chemical Technology Inc now Lyondell Chemical oot Fuel oil v Propylene Co and is exclusively offered worldwide for license by Kellogg Brown Flue gas system e C Root It uses a fluidized catalytic reactor system with a proprietary Catalyst Y ff y catalyst to convert lowvalue feedstocks to predominately propylene Yi and ethylene products The catalyst is very robust thus no feed pre Oil wash treatment is required for typical contaminants such as sulfur water Regn air oxygenates or nitrogen Attractive feedstocks include Cy and C olefin rich streams from ethylene plants FCC naphthas or Cys thermally cracked naphthas from visbreakers or cokers BTX or MTBE raffinates Cs olefinrich streams removed from motor gasolines and Fischer Tropsch light liquids The fluidized reactor system is similar to a refinery FCC unit and consists of a fluidized reactorregenerator vessel air compression catalyst handling fluegas handling and feed andeffluentheatrecovery Yields The technology produces up to 70 wt propylene plus ethylene Using this reactor system with continuous catalyst regeneration allows with a propylene yield about twice that of ethylene from typical C4 and higher operating temperatures than with competing fixedbed reactors C raffinate streams Some typical yields are so that a substantial portion of the paraffins as well as olefins are Pyrolysis Pyrolysis converted This allows for flexibility in the amounts of paraffins in the Feedstock FCCLCN Coker LN Cas Css feeds to SUPERFLEX and the ability to recycle unconverted feed to eumate yield wt a uel gas 136 116 72 120 extinction Ethylene 200 198 225 221 The cooled reactor effluent can be processed for the ultimate Propylene 401 387 482 438 production of polymergrade olefins Several design options are avail Propane 66 70 53 65 able including fully dedicated recovery facilities recovery in a nearby Ce gasoline 197 229 168 156 existing ethylene plant recovery section to minimize capital investment Ultimate yield with Cas and Css recycled or processing in a partial recovery unit to recover recycle streams and concentrate olefinrich streams for further processing in nearby plants Commercial plants The first SUPERFLEX licensee with a propylene pro duction of 250000 mtpy is Sasol Technology Engineering is underway and completion of the unit in South Africa is scheduled for 2005 Licensor Kellogg Brown Root Inc Propylene continued PROCESSING PetrochemicalProcesses miele IN ce a home processes index company index Propylene methanol to propylene MTP Logs Methanol wel Application To produce propylene from natural gas via methanol This route delivers dedicated propylene from nonpetroleum sources ié in DME Product dependently from steam crackers and FCCs Preeacior conditioning Propylene Description Methanol feed from a MegaMethanol plant is sent to an a LPG adiabatic DME prereactor where methanol is converted to DME and a water The highactivity highselectivity catalyst nearly achieves thermo dynamic equilibrium The methanolwaterDME stream is routed to the first MTP reactor stage where the steam is added MethanolDME are Water recycle eae converted by more than 99 with propylene as the predominant hy ar drocarbon product Additional reaction proceeds in the second and third y MTP reactor stages Process conditions in the three MTP reactor stages wargs are chosen to guarantee similar reaction conditions and maximum total Reactor section Purification section propylene yield The product mixture is then cooled and product gas organic liquid and water are separated The product gas is compressed and traces of water CO and DME are removed by standard techniques The cleaned gas is then further processed yielding chemical or polymergrade propylene as specified terest and d ati derat dit of 160 mt f Several olefincontaining streams are recycled to the MTP reactor as IMNETEST ANG GEPFECiaHON assuMing a Moderate creat me ror ne the byproduct gasoline additional propylene sources To avoid accumulation of inert materials within the loop a small purge removes lightends further purge streams Technology status From January 2002 until March 2004 a demonstra of Cy and CsC Highgrade gasoline is produced as a byproduct tion unit was operating at the Statoil methanol plant at Tjeldbergodden Water is recycled to steam generation excess water from the Norway This unit has confirmed the lab results The catalyst is commer methanol conversion is purged This process water can be used for cially available Lurgi offers the process on commercial terms irrigation after appropriate and inexpensive treatment References Koempel H W Liebner and M Wagner MTPAn Economics Current studies and projects are based on a combined economical route to dedicated propylene Second ICISLOR World MegaMethanolMTP plant with a capacity of 5000 mtpd of methanol Olefin Conference Amsterdam Feb 1112 2003 1667 million mtpy yielding approximately 519000 mtpy of propylene Koempel H W Liebner and M Rothaemel Progress report on and 143000 mtpy of gasoline Based on a natural gas cost of 05 MMBtu net production cost h for propylene will be 166 mt Including owners cost capitalized MTP with focus on DME AIChE Spring National meeting New Orleans April 25 29 2004 Licensor Lurgi AG Propylene methanol to propylene MTP continued tistesstttcial MHA NANO Ae eee OTe ATLANTIC Ud Mele cists iets home processesindex company index Propylene Applications To primarily produce propylene from Cy to Cg olefins sup plied by steam crackers refineries andor methanoltoolefins MTO Lightolefin plants via olefin cracking covery Olefinic C C feed Description The ATOFINAUOP Olefin Cracking Process was jointly de veloped by Total Petrochemicals formerly ATOFINA and UOP to convert C byproduct lowvalue Cy to Cg olefins to propylene and ethylene The process fea tures fixedbed reactors operating at temperatures between 500C and Depropanizer 600C and pressures between 1 and 5 bars gauge eee This process uses a proprietary zeolitic catalyst and provides high yields of propylene Usage of this catalyst minimizes reactor size and operating costs by allowing operation at highspace velocities and high conversions C byproducts and selectivities without requiring an inert diluent stream A swingreactor system is used for catalyst regeneration Separation facilities depend on how the unit is integrated into the processing system The process is designed to utilize olefinic feedstocks from steam crackers refinery FCC and coker units and MTO units with typical C to Cg olefin and paraffin compositions The catalyst exhibits little sensitivity to common impurities such as dienes oxygenates sulfur compounds and nitrogen compounds Commercial plants Total Petrochemicals operate a demonstration unit that was installed in an affiliated refinery in Belgium in 1998 Engineer Economics Capital and operating costs depend on how the process is ing is in progress for the first commercial unit integrated with steam cracking refinery or other facilities Yields Product yields are dependent on feedstock composition The pro Licensor UOP LLC cess provides propyleneethylene production at ratios of nearly 41 Case studies of olefin cracking integration with naphtha crackers have shown 30 higher propylene production compared to conventional naphtha cracker processing Reference Vermeiren W J Andersen R James D Wei Meeting the changing needs of the light olefins market Hydrocarbon Engineering October 2003 tistesstttcial PetrochemicalProcesses home processes index company index Propylene Application To produce polymergrade propylene from propane using the Oleflex process in a propylene production complex pe I 7 fC Description The complex consists of a reactor section continuous cata a expander lyst regeneration CCR section product separation section and fraction HG ation section Four radialflow reactors 1 are used to achieve optimum mropvI conversion and selectivity for the endothermic reaction Catalyst activity Cp hoo sHp robyene is maintained by continuously regenerating catalyst 2 Reactor effluent is compressed 3 dried 4 and sent to a cryogenic separation system 3 7 5 A net hydrogen stream is recovered at approximately 90 mol hy eee drogen purity The olefin product is sent to a selective hydrogenation H Recycle process 6 where dienes and acetylenes are removed The propylene Propane Net H stream stream goes to a deethanizer 7 where lightends are removed prior to ae the propanepropylene splitter 8 Unconverted feedstock is recycled ea back to the depropanizer 9 where it combines with fresh feed before being sent back to the reactor section Yields Propylene yield from propane is approximately 85 wt of fresh feed Hydrogen yield is about 36 wt of fresh feed Economics US Gulf Coast inside battery limits are based on an Ole Commercial plants Eleven Oleflex units are in operation to produce flex complex unit for production of 350000 mtpy of polymergrade propylene and isobutylene Six of these units produce propylene These propylene The utility summary is net utilities assuming all light ends are Units represent 125 million mtpy of propylene production Three ad used as fuel ditional Oleflex units for propylene production are in design or under j oo construction Inside battery limits investment million 145 Total project investment million 210 Licensor UOP LLC Typical net utility requirements per ton of propylene product Electricity kWh 200 Water cooling m 50 Net fuel gas MMkcal export credit 12 Catalyst and chemical cost metric ton product 14 me Ce Cm ely j PROCESSING PetrochemicalProcesses miele IN ce a JCESSE home processes index company index PVC suspension Application A process to produce polyvinyl chloride PVC from vinyl chloride monomer VCM using suspension polymerization Many types 6 vv older of PVC grades are produced including commodity high Kvalue low 9 Kvalue matted type and copolymer PVC The PVC possesses excellent Recovery VCM product qualities such as easy processability and good heat stability Fresh VCM Additives Description PVC is produced by batch polymerization of VCM dispersed Water in water Standard reactor sizes are 60 80 100 or 130 m The stirred reactor 1 is charged with water additives and VCM oO O Centrifuge During polymerization reaction the temperature is controlled at a de Effluent fined temperature depending on the grade by cooling water or chilled Reactor water At the end of the reaction the contents are discharged into a Blowdown Slurry PVC product blowdown tank 2 where most of the unreacted VCM is flashed off The tank Stripping tank reactor is rinsed and sprayed with an antifouling agent and is ready for Dryer the following batch The PVC slurry containing VCM is continuously fed to the stripping col umn 3 The column has a proprietary design and effectively recovers VCM from the PVC slurry without any deterioration of PVC quality After strip ping the slurry is dewatered 4 and dried effectively by the proprietary dryer 5 It is then passed to storage silos for tanker loading or bagging Licensor Chisso Corp Recovered VCM is held in a gas holder 6 then compressed cooled and condensed to be reused for the following polymerization batch Economics Raw materials and utilities per ton of PVC VCM t 1003 Electricity kWh 160 Steam t 07 Additives for pipe grade US 12 Commercial plants The process has been successfully licensed 15 times worldwide Total capacity of the Chisso process in the world is more than 15 million tpy In addition Chisso VCM removal technology has been licensed to many PVC producers worldwide tistesstttcial PetrochemicalProcesses miele IN ce a home processes index company index PVC suspension Application Production of suspension polyvinyl chloride PVC resins Liquid RVCM from vinyl chloride monomer VCM using the Vinnolit process RVCM recovery Description The Vinnolit PVC process uses a new highperformance re Fresh actor 1 which is available in sizes up to 150 m A closed and clean VCM a 6 reactor technology is applied thus opening of the reactors is not neces sary except for occasional inspections Equally important highpressure Cooling Dried oo water HEE PVC to water cleaning is not necessary All process operations of this unit are 5 Natural gas storage controlled by a distributed process control system DCS vere or steam The batchwise polymerization occurs in the following operation se Disper aa 2 air quence agent ee OT e Prepare the reactor which includes applying a highly effective an Catalyst centrifuge Air heater tifouling agent SPVC process e Charge reaction solutions including dispersing agents additives polymerization and degassing SPVC process drying chemicals VCM and water into the reactor e Exothermic conversion from VCM to PVC e Discharge of the PVC slurry into the blowdown tank e Flush the reactor internals The PVC slurry and unreacted VCM from the polymerization reactors are fed to the blowdown tankthe intermediate buffer between the dis Raw materials and utilities per metric ton of PVC continuous polymerization and the continuous degassing and drying unit VCM t 1001 In the blowdown tank 2 unreacted VCM is flashed out of the PVC Steam t 08 slurry From the blowdown tank the slurry is fed through heat recu Electricity kWh 170 perator 3 to the sievetray type Vinnolit degassing column 4 VCM is Additive costs for pipe grade US 14 stripped out with steam The VCM concentration of the slurry leaving Productivity tmy up to 600 the degassing column is less than 1 ppm The unreacted VCM Is lique Vi i fied in the VEN recovery unit and charged back to polymerization After Commercial plants Vinnolit is producing up to 620000 PVC metric tpy Total capacity of the Vinnolit process in the world is about one million dewatering the SUSPENSION In the centrifuge 5 the wet PVC cake S fed metric tpy Vinnolit cyclone dryer has been licensed to many PVC pro in the Vinnolit cyclone drying system 6 The solid particles and air are ducers worldwide separated in the cyclone separator 7 Economics Chilled water for polymerization is not required High pro i ductivity is achieved by using an innercooler reactor Licensor Vinnolit Contractor Uhde GmbH PVC suspension continued erase ca PetrochemicalProcesses home processes index S company index Upgrading pyrolysis gasoline Application Increase the value of steam cracker pyrolysis gasoline py gas using conversion distillation and selective hydrogenation process C5 dimenwation ao es Pygas the CCg fraction issuing from steam crackers is a potential and recovery source of products such as dicyclopentadiene DCPD isoprene cyclo pune de pentane benzene toluene and xylenes ee Gs from isoprene extraction li Gs t RG Description To produce DCPD and isoprene pygas is depentanized and a FG Sar Crs HS the Cs fraction is processed thermally to dimerize cyclopentadiene to cracking DCPD which separates easily 1 from the Css via distillation Isoprene 7 3 can be recovered by extractive distillation and distillation The remaining Css and the CgCog cut are fed to the first stage 2 catalytic hydrogena hydro hydro tion unit where olefins and diolefins are eliminated Boe The Ces are recycled to the steam cracker or an isomerization unit Hy Hy en Sulfur and nitrogen compounds are removed in the second stage 3 Cot hydrogenation units The BTX cut is ideal for processing in an aromatics optional complex Yields For the new generation catalysts recovery and product quality parameters are as follows 5 to Cy aromatics aia 300 References Debuisschert Q P Travers and V Coupard Optimizing enzene recovery 7 Diene value 0 Pyrolysis Gasoline Upgrading Hydrocarbon Engineering June 2002 Sulfur p om mg100g 0 Commercial plants Over 90 1st stage and 60 2nd stage pygas hydroge Thiophene ppm 02 nation units have been licensed C cut Bromine Index mg100g 20 Ce cut acid wash color 1 Licensor Axens Axens NA Economics Based on a 1 million metric toy naphtha steam cracker pro ducing a 620000 tpy pygas stream ISBL Gulf Coast location in 2004 Investment US metric ton of feed 40 Utilities catalysts US metric ton 10 a COC Cee ey a PROCESSING PetrochemicalProcesses miele IN ce a home processes index company index Styrene Application To produce polymergrade styrene monomer SM by dehy Recycle benzene drogenating ethylbenzene EB to form styrene using the LummusUOP Styrene Classic styrene process for new plants and the LummusUOP SMART moron process for revamps involving plant capacity expansion Inhibitor Description In the Classic SM process EB is catalytically dehydroge Toluene nated to styrene in the presence of steam The vapor phase reaction is carried out at high temperature under vacuum The EB fresh and Fuel gas Tar recycle is combined with superheated steam and the mixture is de hydrogenated in a multistage reactor system 1 A heater reheats the Etniyibenizelic oan Hydrocarbons process gas between stages Reactor effluents are cooled to recover se waste heat and condense the hydrocarbons and steam Uncondensed offgascontaining mostly hydrogenis compressed and is used as Superheater AirO C2 fuel Condensed hydrocarbons from an oilwater separator 2 are sent SMART only Condensate to the distillation section Process condensate is stripped to remove dissolved aromatics A fractionation train 34 separates highpurity styrene product un converted EB which is recycled and the relatively minor byproduct tar which is used as fuel Toluene is produced 56 as a minor byproduct and benzene 6 is normally recycled to the upstream EB process nates the costly interstage reheater and reduces superheated steam Typical SM product purity ranges from 9985 to 9995 Thepro requirements For existing SM producers revamping to SMART SM cess provides highproduct yield due to a unique combination of catalyst May be the most costeffective route to increased capacity and operating conditions used in the reactors and the use of a highly affective solymerization inhibitor in the fractionation columns Economics Classic 500000 mtpy ISBL US Gulf Coast The SMART SM process is the same as Classic SM except that oxi Investment US million 78 dative reheat technology is used between the dehydrogenation stages Ethylbenzene tonton SM 1055 of the multistage reactor system 1 Specially designed reactors are Utilities USmton SM 29 used to achieve the oxidation and dehydragenation reactions In oxi Commercial plants Currently 36 operating plants incorporate the dative reheat oxygen introduced to oxidize part of the hydrogen LummusUOP Classic Styrene technology Seven operating facilities produced over a proprietary catalyst to reheat the process gas and to remove the equilibrium constraint for the dehydrogenation reaction The process achieves up to about 80 EB conversion per pass elimi are using the SMART process technology Many future units using the SMART process are expected to be retrofits of conventional units since the technology is ideally suited for revamps Licensor ABB Lummus Global and UOP LLC Styrene continued PROCESSING PetrochemicalProcesses miele IN ce a home processes index company index Styrene Application Process to manufacture styrene monomer SM by dehydro genating ethylbenzene EB to styrene Feedstock EB is produced by al Styrene product Recycle F Benzenetoluene byproduct kylating benzene with ethylene using the MobilBadger EBMax process soyone Description EB is dehydrogenated to styrene over potassium promoted ironoxide catalyst in the presence of steam The endothermic reaction is Offgas we to fuel done under vacuum conditions and high temperature At 10 weight ratio Heavies to fuel of steam to EB feed and a moderate EB conversion reaction selectivity to et fi styrene is over 97 Byproducts benzene and toluene are recovered via 6 distillation with the benzene fraction being recycled to the EB unit e Vaporized fresh and recycle EB are mixed with superheated steam 1 and fed to a multistage adiabatic reactor system 2 Between dehydrogenation Steam 3 4 Sein stages heat is added to drive the EB conversion to economic levels typically between 60 and 75 Heat can be added either indirectly using cooling Clean conventional means such as a steam heat exchanger or directly using a water condensate proprietary Direct Heating Technology developed by Shell Oil Reactor effluent is cooled in a series of exchangers 3 to recover waste heat and to condense 4 the hydrocarbons and steam Uncondensed offgasprimarily hydrogenis compressed 5 and then directed to an absorber system 6 for recovery of trace aromatics Following aromatics Economics recovery the hydrogenrich offgas is consumed as fuel by process Ethylbenzene consumption per ton of SM 1052 heaters Condensed hydrocarbons and crude styrene are sent to the Net energy input kcal per ton of SM 125 distillation section while process condensate is stripped 7 to remove Water cooling m per ton of SM 150 dissolved aromatics and gases The clean process condensate is returned Note Raw material and utility requirements presented are representative each plant is optimized based on specific raw material and utility costs as boiler feedwater to offsite steam boilers The distillation train first separates the benzenetoluene byproduct from Commercial plants The technology has been selected for use in over main crude styrene stream 8 Unconverted EB is separated from styrene 9 40 units having design capacities single train ranging from 320 to 850 and recycled to the reaction section Various heat recovery schemes are used Mmtpy The aggregate capacity of these units exceeds 8 MMmtpy to conserve energy from the EBSM column system In the final purification step 10 trace Cg components and heavies are separated from the finished Licensor Badger Licensing LLC SM To minimize polymerization in distillation equipment a dinitrophenolic type inhibitor is cofed with the crude feed from the reaction section Typical SM purity ranges between 9990 and 9995 ee Cu mC Cr et iste se cal PetrochemicalProcesses miele IN ce meena JCESSES home processes index company index Styrene Application To directly recover styrene from raw pyrolysis gasoline de rived from steam cracking of naphtha gas oils and NGLs using the GT To hydrogenation Styrene process oS ay Styrene C Description Raw pyrolysis gasoline is prefractionated into a heartcut Cg concentrate Styrene stream The resulting styrene concentrate is fed to an extractivedistillation Extractive product column and mixed with a selective solvent which extracts styrene to the C5Cot distillation Pyrolysis tower bottoms The rich solvent mixture is routed to a solventrecovery gasoline column which recycles lean solvent to the extractivedistillation column Solvent and recovers the styrene overhead A final purification step produces a ae 999 styrene product containing less than 50 ppm phenyl acetylene Rich solvent The extractivedistillation column overhead can be further processed on to recover a highquality mixed xylene stream A typical worldscale Lean solvent cracker could produce approximately 25000 tpy styrene and 75000 Prefractionator Styrene recovery Purification toy mixed xylenes from pyrolysis gasoline The styrene is a highpurity product suitable for polymerization at a very attractive cost compared with conventional styrene production routes If desired the mixed xylenes can also be extracted from the pygas upgrading their value as chemical feedstock The process is economically attractive for typical pygas and supplemental feeds which Styrene product sales value t 700 contain 15000 tpy or more styrene processing cost t 100 ross margin MMyr 875 Traditional pygas processing schemes destroy styrene in the firststage Pretax ROI 43 hydrogenation unit Hydrotreated pygas is then fractionated to extract benzene and toluene With the GFStyrene process this fractionation Commercial plants One license has been placed is made upstream of the hydrotreaters which avoids some hydrogen i consumption and catalyst fouling by styrene polymers In many cases Reference Generate more revenues from pygas processing Hydro most of the existing fractionation equipment can be reused in the bon Processing June 1997 styrenerecovery mode of operation Licensor GTC Technology Economics Styrene recovery considering styrene upgrade only basis 25000tpy styrene capacity Mirenevace oon okt SE PROCESSING PetrochemicalProcesses E home processes index company index Styrene acrylonitrile SAN copolymer Application To produce a wide range of styrene acrylonitrile SAN co polymer with excellent chemical resistance heat resistance and suitable Styrene property for compounding with ABS via the continuous bulk polymer eceninile ization process using Toyo Engineering Corp TECMitsui Chemicals Solvent 2 Inc technology Asertves a Description Styrene monomer acrylonitrile a small amount of solvent Reactor oe and additives are fed to the specially designed reactor 1 where the polymerization of the fed mixture is carried out The polymerization Vy Condenser temperature of the reactor is carefully controlled at a constant level to 7 maintain the desired conversion rate The heat of the polymerization is eee ened easily removed by a specially designed heattransfer system At the exit a of the reactor the polymerization is essentially complete Sprage The mixture is preheated 2 and transferred to the devolatilizer 3 Volatile components are separated from the polymer solution by evapo ration under vacuum The residuals are condensed 4 and recycled to the process The molten polymer is pumped through a die 5 and cut into pellets by a pelletizer 6 Economics Basis 50000 metric toy SAN US Gulf Coast Investment million US 16 Raw materials consumption per one metric ton of SAN kg 1009 Utilities consumption per one metric ton of SAN US 18 Installations Seventeen plants in Japan Korea Taiwan China and Thai land are in operation with a total capacity of 508000 metric tpy Licensor Toyo Engineering Corp TEC Mitsui Chemicals Inc a COC Cee ey a PROCESSING PetrochemicalP eee C OCESSES home processes index company index Terephthalic acid E PTA Application E PTA Eastman polymergrade terephthalic acid is an ex cellent raw material for engineering plastics and packaging materials Vent gas to off gas cleaning bottles other food containers including hot fill as well as films The potions Neeeotem onion process is proven to be suitable for the production of all kinds of polyes ter fibers and containers without limitation at international firstgrade E PTA quality Acetic 1 6 EPTA process Description The general flow diagram to produce E PTA using East t man Chemicals proprietary process comprises three different main sec Air tionscrude terephthalic acid CTA polymergrade terephthalic acid E al CTA residue PTA and catalyst recovery Solvent catalyst recycle 4 Fiiate to incineration Crude terephthalic acid 123 CTA is produced by the catalytic treatment oxidation of pxylene with air in the liquid phase using acetic acid as a a solvent 1 The feed mixpxylene solvent and catalysttogether with compressed air is continuously fed to the reactor which is a bubblecolumn oxidizer It operates at moderate temperature and offers an extremely high yield The oxidizer product is known as crude terephthalic acid CTA due to the high level of impurities contained Many impurities are fairly soluble in the solvent In the CTA separation step 2 impurities can be effectively removed from the product by is separated from the solvent and dried for further processing in the exchanging the reaction liquor with lean solvent from the solvent Polyesterproduction facilities recovery system The reactor overhead vapor mainly reaction water Catalyst recovery 4 After exchanging the liquor in the CTA acetic acid and nitrogen is sent to the solventrecovery system 3 where Separation the suspended solids are separated and removed as CTA water is separated from the solvent by distillation After recovering its residue which can be burned in a fluidizedbed incinerator or if energy the offgas is sent to a regenerative thermal oxidation unit for desirable used as land fill The soluble impurities are removed from the further cleaning filtrate within the filtrate treatment section and the dissolved catalyst is Polymergrade terephthalic acid 56 The crude acid is purified ecycled to obtain E PIA in a postoxidation step at elevated temperature Economics The advanced Eastman E PTA technology uses fewer pro conditions The post oxidizers serve as reactors to increase conversion cessing steps In combination with the outstanding mildoxidation tech of the partially oxidized compounds to terephthalic acid The level of 4 carboxy benzaldehyde 4CBA ptoluic acid pTAthe main impurities in terephthalic acidis significantly lowered In a final step 6 E PTA Sa click here to email for more information a nology this technology leads to considerable capital cost savings and lower production cost than in other technologies Commercial plants Commercial plants are operating in the US Europe and Asia Pacific The latest plant with a capacity of 660000 tpy for Zhejiang Hualian Sunshine PetroChemical Co Ltd in Shaoxing China is under construc tion and will be started up in April 2005 increasing the worldwide ca pacity to 21 million tpy Licensor Lurgi AG Terephthalic acid E PTA continued iste se cal PetrochemicalProcesses miele IN ce meena MO etsyslete eae home processes index company index Upgrading steam cracker C cuts Application To purify propylenepropane cuts from pyrolysis processes via selective catalytic hydrogenation of methylacetylene and propadiene im purities MAPD Steam cracker C3 effluents typically contain over 90 Hydrogen propylene with propane and MAPD making up the balance Although distillation can be used to remove MAPD it is often not practical or Main Finishing economical for achieving a propylene product meeting the partspermillion levels required by chemical and polymergrade propylene specifications Furthermore distillation results in propylene losses Selective hydrogenation is the route most commonly employed as it not only achieves the tight MAPD specifications but it produces more propylene 5 Description The C3 cut is joined by recycled C3s and makeup hydrogen a prior to entering the main reactor 1 There the MAPD is catalytically oe Hydrogenated Css hydrogenated forming propylene and traces of propane A single reactor suffices for polymergrade propylene MAPD content 10 ppm when a C3 splitter is used A finishing reactor 2 can be used to reduce MAPD content to five or even one ppm A second reactor is advantageous when making chemicalgrade propylene With a typical specification of 95 propylene 5 propane and 5 ppm MAPD a costly C3 splitter system is avoided Economics Based on a 1million tpy capacity steam cracker ISBL Gulf Coast location in 2004 Yields The highly selective active and stable catalyst LD 273 provides Investment USmetric ton of 49 propylene the tyerca yess shown pelow compare its predecessor LD 265 Utilities catalysts USmetric ton of propylene 024 which Is used In most of the units worlawide Feed Product A Performance Commercial plants Over 100 C3 hydrogenation units have been licensed with LD273 wt Ethane 010 011 Licensor Axens Axens NA Propane 328 421 Propylene 9403 9555 1 Propadiene 123 1 ppm Methylacetylene 133 1 ppm Cyclopropane 003 003 Ce 0 012 Propylene yield 1016 11 PROCESSING PetrochemicalProcesses miele IN ce meena JCESSES home processes index company index Upgrading steam cracker C cuts Application Increase the value of steam cracker C cuts via lowtemper ature selective hydrogenation and hydroisomerization catalysis Several Hydrogen Highpurity Cs olefins options exist removal of ethyl and vinyl acetylenes to facilitate butadi ene extraction processing downstream conversion of 1 3 butadiene to Main Finishing maximize 1butene or 2butene production production of highpurity reactor reactor isobutylene from crude C cuts total C cut hydrogenation and total C hydrogenation of combined C3C and CC cuts for recycle to cracking furnaces or LPG production fat Description Each application uses a specific process catalyst and op erating conditions The basic process for maximizing 1butene consists of sending a combined butadienerich Cy cut recycled Cys makeup iw hydrogen to the main reactor 1 where acetylenes and 13 butadiene in the case of hydroisomerization to a specified product distribution CW are hydrogenated A finishing reactor 2 is used if required Reactions take place in the liquid phase at relatively low temperatures to provide significant advantages in the area of heat removal approach to equi librium catalyst life and reaction homogeneity Information here is for the Cy selective hydrogenation process employed to maximize 1butene Distillation is used to separate the products The process is different in Cis 2butene 388 927 the case of high purity isobutylene production where a reactor and dis 1 3 Butadiene 4858 13 ppm tillation column operate on the Cy stream simultaneously 1 2 Butadiene 015 0 Vinylacetylene 061 0 Yields In the example below a highly selective active and stable cat Ethylacetylene 015 005 alyst LD 271 provides the typical yields shown below 50 of the Economics Based on a 160000tpy crude C feed ISBL Gulf Coast 1 3 butadiene converts to 1butene location in 2004 Feed Product with LD271 wt Investment US 31 million C3s 003 003 Utilities catalysts Water cooling m3h 500 Isobutane 062 063 Electrical power kWhh 250 nButane 342 571 1Butene 1293 3722 Isobutene 2451 2444 Trans 2butene 511 2265 Commercial plants A total of over 50 C4 hydrogenation units have been licensed for this process application Licensor Axens Axens NA Upgrading steam cracker C4 cuts continued iste se cal PetrochemicalProcesses miele IN ce OTC ae AAT TC re ted RO1C erste ers ene home processes index company index Urea Application To produce urea from ammonia NH3 and carbon di f oxide CO using the Stamicarbon CO stripping Urea 2000plus a a Technology vs Stm A H ap rl ve Description Ammonia and CO react at synthesis pressure of 140 a bar to urea and carbamate Fig 1 The conversion of ammonia as CO dy SZ bh o well as CO in the synthesis section is 80 resulting in an extreme 7 low recycle flow of carbamate Because of the highammonia ef ficiency NO pure ammonia is recycled in this process The synthesis temperature of 185C is low and consequently corrosion in the plant is negligible Because of the elevation difference within the synthesis section in ternal synthesis recycle is based on gravity flow Result Electrical energy requirement is very low Synthesisgas condensation in the pool reac tor generates steam which is used in downstream sections within the plant Process steam consumption is low Processing inerts are vented to the atmosphere after washing thus ammonia emissions from the plant are virtually zero Because of the high conversions in the synthesis the recycle section of the plant is very small An evaporation stage with vacuum condensa tion system produces urea melt with the required concentration either for the Stamicarbon fluidizedbed granulation or for prilling Process wa ter produced in the plant is treated in a desorbtionhydrolyzer section This section produces an effluent which is suitable for use as boiler feedwater Stamicarbon licenses several proprietary technologies Fluidizedbedgranulation e Urea 2000Plus Technology for capacities up to 5000 metric tpd e Stamicarbon fluidized bed urea granulation Fig 2 e UAN technology e Several revamp technologies e Proprietary material Safurex Economics Depending on heat exchange options included within the design the raw material and utility consumptions per metric ton of urea melt are Ammonia kg 566 Carbon dioxide kg 733 Steam 110 bar 510C kg 6901 Electric power kWh 14 Water cooling m3 50 1 Includes steam for CO2 compressor drive and steam for desorbtionhydrolyzes section Commercial plants More than 200 plants based on Stamicarbons CO2 stripping technology are in operation The largest singleline unit with Urea 2000plus technology produces more than 3250 metric tpd Highlights in 2005 include Three urea plants with Stamicarbons new Granulation technology are under construction One Urea 2000plus Technology plant with a complete synthesis in Safurex is under construction More than six major capacity increase revamps are under con struction Licensor Stamicarbon BV Urea continued tistesstttcial PetrochemicalProcesses miele IN ce meena eerste lets ae home processesindex company index Urea Application To produce urea from ammonia and carbon dioxide COz ee eee i using the CO stripping process Stripper condenser decomposer decomposer Evaporator Description Ammonia and carbon dioxide react at 155 bar to synthesize A A 5 urea and carbamate The reactor conversion rate is very high under the awe Ms Ai 1 g L NC ratio of 37 with a temperature of 182185C Unconverted mate ll i co i to priting rials in synthesis solution are efficiently separated by CO stripping The 7 Cal T bh T ds toe milder operating condition and using twophase stainless steel prevent ss MA inelt corrosion problems Gas from the stripper is condensed in vertical sub up mn fond ara merged carbamate condenser Using an HP Ejector for internal synthesis a absorber eee recycle major synthesis equipment is located on the ground level Ured The urea solution from synthesis section is sent to MP decomposer Nis slurry Pa at 17 bar and LP decomposer at 25 bar for further purification No pure oe ammonia recycle is required due to the high separation efficiency in the pump Carbamate pump stripper The vacuum evaporator unit produces urea melt at the required concentration either for prilling or granulation The vent scrubber and process condensate treatment unit treat all emission streams thus the plant is pollution free Process condensate is hydrolyzed and reused as boiler feedwater Commercial plants More than 100 plants including urea granulation Toyo Engineering Corp TEC has a spoutfluid bed granulation plants have been designed and constructed based on TEC technology technology to produce granular ureatypically 24 mm size Due to proprietary granulator electric power consumption is the lowest among Licensor Toyo Engineering Corp TEC granulation processes Economics Raw materials and utilities consumptions per metric ton of orilled urea are Ammonia kg 566 Carbon dioxide kg 733 Steam 110 bar 510C 690 Electric power kWh 20 Water cooling m 75 1 Includes steam for COz compressor turbine and steam for process condensate treatment PROCESSING PetrochemicalProcesses miele IN ce a Mele erste home processes index company index Ureaformaldehyde Application Ureaformaldehyde resins are used as adhesives in the woodworking industry and are typically used in the production of ply ial ae i wood and particle board They are available as concentrated solutions or we LD in powder form as a result of the spraydrying process Description The reaction mechanisms of the major components are i wr Formaldehyde and urea are by polyaddition wo LP HNCONHCH0 HNCONHCH0H hea wer S Ah 24 kJmol Prey NJ Dt The hydroxymethyl compounds undergo further slow reaction by o polycondensation LI wet OD CT Formaldehyde Pl NM One H2NCONHHNCONHCHOH S OS PN op Urresin Lg HNCONHCHNHCONHH0 which is also responsible for the viscosity increase during the storage The formation of methylene bridges can be accelerated by raising storage temperatures The technology is based on batchwise production of the aqueous solution short intermediate storage and continuously Licensor Uhde InventaFischer operating spray drying in a connected stage After cooling the resin in the reactor the resin is pumped to the buffer tank of the connected spray dryer plant Usually the complete batch processing takes 45 h The ureaformaldehyde resin solution can be dried in a spray dryer based on cocurrent flow principle This process costeffectively produces highquality glues at large quantities The product is a lowformaldehyde resin adhesive suitable for veneering plywood and particle board production by the hot pressing process The quality of the bonding complies with the requirements of DIN 68705 Part 2 respectively to DIN 68763V20 For particle board a perforate value according to DIN EN 120 of less 10 mg HCHO100 g dry board will be maintained PROCESSING PetrochemicalP eee C Oe Sich home processes index company index VCM by thermal cracking of EDC Application Vinnolits mediumpressure EDCcracker provides an energy efficient cracking technology operating at moderate cracking pressure Furnace feedEDC Gite Dente et iad epmitenieau oh with the benefit of low byproduct formation and long operation cycles Steam between cleaning intervals Steam mae Description In the cracking furnace feed EDC ethylene dichloride from the EDC purification section or from the EDC storage facility is cracked to vinyl chloride and hydrogen chloride HCI at approximately 490C and at 15 MPa g Prior to cracking the feed EDC is preheated in the quench ie overhead exchanger and in the radiation coils of the EDCcracker The To HCl column hot reaction gases downstream of the EDC cracking furnace are cooled in the EDCevaporator by vaporizing the feed EDC Additional cooling of the reaction gas occurs in the quench tower Fractions of the quench cle overhead stream are condensed in the steam generator in the feed EDC cee a Moa repair preheater of the quench column prior to entering the HClcolumn The quenchbottom product is filtered and fed through a highefficiency flash system to remove coke Process features and economics Processing benefits of the VINNOLIT EDC cracking process consist of EDC cracking furnace and external EDC the vinyl chloride monomer VCM distillation unit evaporation and include Low maintenance cost The natural EDC circulation in EDC vaporizer Energy savings More than 50 savings of electrical energy minimizes maintenance costs no pumps no sealing problems and no compared to lowpressure cracking furnace technology are available plugging because of 125 bar g condensation pressure in the HCI column further reduced fuel consumption by using the heat of the cracking gas to heat Commercial plants The process is used in 19 plants with an annual and evaporate EDC nets savings of 500 kg 20 bar g steamton VCM production of around 38 million metric tons mtons of VCM A single Furthermore steam is generated via flue gas from the furnace EDC is stream plant with an annual capacity of 400000 mtons of VCM was preheated on quench top prior to entering the furnace commissioned in a record time of two months in September 2004 One Operation The continuous operation time is approximately two VCM plant with an annual capacity of 300000 mtons of VCM is under years without decoking The high conversion rate is 55 due to the Construction vapor EDCfeed No iron enters the radiant section As the coke carryover with the product stream is avoided the Sa click here to email for more information a Vinnolit desuperheated quench system allows a long operation time of Licensor Vinnolit Contractor Uhde GmbH VCM by thermal cracking of EDC continued iste se cal PetrochemicalProcesses home processes index company index VCM removal Application Adding a stripping column to existing polyvinyl chloride vem PVC plants to remove vinyl chloride monomer VCM from PVC slurry PVC slurry Vem gas holder The recovered VCM can be reused in the PVC process without any de from gas terioration of PVC polymer quality a oo recovery Description PVC slurry discharged from reactors contains significant Vacuum amounts of VCM 30000 ppm even after initial flashing This process pump effectively removes the remaining VCM so that the monomer is recov Steam ered and reused Recycling of raw materials drastically reduces VCM a slur emissions from the following dryer There is no significant change in Blowdean 1 tank PVC quality after stripping Residual VCM level in the PVC product can rant Slurry Stripping Slurry To dryer be lowered below 1 ppm and in some cases below 01 ppm feed column discharge The PVC slurry containing VCM is continuously fed to the stripping p Pump pump column 1 The slurry passes countercurrently to steam which is fed BL BL into the base of the column The proprietary internals of the column are specially designed to ensure intimate contact between the steam and the PVC slurry and to ensure that no PVC particles remain inside the col umn All process operations including grade change are automatically done in a completely closed system While steam stripping is widely used this proprietary technology Licensor Chisso Corp which involves sophisticated design and knowhow of the column of fers attractive benefits to existing PVC plant sites The process design is compact with a small area requirement and low investment cost The size of the column is 25 th to 30 th Economics Steam 130 kgt of PVC Commercial plants Chisso has licensed the technology to many PVC producers worldwide More than 100 columns of the Chisso process are under operation or construction and total capacity exceeds 5 million tpy of PVC ee Cu mC Cr et a iste se cal PetrochemicalProcesses miele IN ce a mele etstsie home processes index company index Wet air oxidation WAO spent caustic Application To oxidize sodium sulfide Na2S component in the caustic Ofigas scrubber effluent of olefin plants with air using wet air oxidation WAO process developed by Nippon Petrochemicals Co Ltd NPCC the li Wash tower steam condensate cense being available from Toyo Engineering Corp TEC Spent caustic Description Conventional wet oxidation processes adopt a plugflow type of reactor systemwhich usually has problems such as ee e Plugflow reactors require higher reaction temperature for the oxi treatment unit dation reaction and need a feed preheater Clogging problems in the outlets of the reactor and preheater often occur e High processing temperatures cause corrosion problems High grade construction materials such as nickel or nickel alloy are needed Air for the reactor De weHa Soe NPCC process conversely uses a complete mixing type of reactor 1 and has several advantages such as e Mild and uniform reactor conditions can be maintained by com plete mixing with very fine bubbles generated by a special nozzle ap plication No preheater is required e A lowergrade construction material such as stainless steel is used Commercial plants Many olefins plants worldwide use this WAO process for the reactor Fourteen processing units have been designed by TEC since 1989 e Less clogging problems and easier operation are due to the simple flow scheme Licensor Nippon Petrochemicals Co Ltd Economics Typical performance data Base Spent caustic flowrate tph 25 NaS Inlet wt 2 Outlet wt ppm less than 10 Utilities Electric power kWhh 175 Steam HP kgh 750 Water coolingmh 55 Washwater mh 2 PROCESSING PetrochemicalProcesses miele IN ce a home processes index company index Xylene isomerization Application To selectively isomerize a paraxylene depletedCg aromat ics mixture to greater than equilibrium paraxylene concentration using Oitges cw ExxonMobil Chemicals XyMax and Advanced MHAI processes Simulta a neously ethylbenzene EB and nonaromatics in the feed are converted Rey he to benzene and light paraffins respectively Conversion of EB is typically s 6080 2 Lt aromatics Description The paradepleted liquid Cg aromatics raffinate stream from en Fractionator the paraxylene separation unit along with hydrogenrich recycle gas are pumped through feedeffluent exchangers and the charge heater Liquid feed C Recycle Steam 1 and into the reactor 2 Vapor then flows down through the fixed aromatics paral comp Isomerate C dualbed catalyst system Dealkylation of EB and cracking of nonaro ortho depleted J Se iia matics preferentially occurs in the top bed The bottom bed promotes isomerization of xylenes while minimizing loss of xylenes from side reac au ee tions The reactor effluent is cooled by heat exchange and the resulting liquid and vapor phases are separated in the product separator 3 The liquid is then sent to a fractionator 4 for recovery of benzene and tolu ene from the isomerate Two enhanced isomerization catalyst technologies have been significant savings in associated paraxylene recovery facilities Both tech developed by ExxonMobil Chemical The first technology referred to nologies offer long operating cycles as Advanced Mobil High Activity Isomerization AMHAI provides higher selectivity and lower operating costs compared to isomerization Commercial plants The AMHAI Process Was first commercialized In processes used in the past The AMHAI technology offers increased 1999 Five AMHA units are currently in operation The first commer operating flexibility in terms of a greater range of EB conversion and a cial unit using XyMax technology was brought onstream in 2000 Since lower temperature requirement The second technology referred to as then two additional applications of the XyMax technology have been XyMax further increases yield performance and debottleneck potential licensed Including other ExxonMobil xylene isomerization technologies This technology can operate at even higher EB conversion with higher there are a total of 22 units in operation selectivity and significantly lower xylene loss Licensor ExxonMobil Chemical Technology Licensing LLC retrofit ap Operating conditions XyMax and AMHAI units operate with a high plications Axens Axens NA grassroots applications space velocity and a low hydrogentohydrocarbon ratio which results in increased debottleneck potential and decreased utilities costs By con verting a high portion of EB in the feed these technologies can provide PROCESSING PetrochemicalProcesses home processes index company index Xylene isomerization Application To selectively isomerize a paraxylene depletedCg aromat of ics mixture to greater than equilibrium paraxylene concentration using aa cw ExxonMobil Chemicals XyMax and Advanced MHAI processes Simul A Gas taneously ethylbenzene EB and nonaromatics in the feed are con verted to benzene and light paraffins respectively Conversion of EB is Reactor typically 6080 Lt aromatics Description The paradepleted liquid Cg aromatics raffinate stream from cw Fractionator the paraxylene separation unit along with hydrogenrich recycle gas are pumped through feedeffluent exchangers and the charge heater 1 and retrace et Recycle team into the reactor 2 Vapor then flows down through the fixed dualbed aromatics para comp Pee catalyst system Dealkylation of EB and cracking of nonaromatics prefer ortho depleted ortho rich entially occurs in the top bed The bottom bed promotes isomerization of hyd k xylenes while minimizing loss of xylenes from side reactions The reactor eel effluent is cooled by heat exchange and the resulting liquid and vapor phas es are separated in the product separator 3 The liquid is then sent to a fractionator 4 for recovery of benzene and toluene from the isomerate Two enhanced isomerization catalyst technologies have been developed by ExxonMobil Chemical The first technology referred to Both technologies offer long operating cycles as Advanced Mobil High Activity Isomerization AMHAI provides higher selectivity and lower operating costs compared to isomerization Commercial plants The AMHAI Process Was first commercialized in 1999 processes used in the past The AMHAI technology offers increased Seven AMHAI units are currently in operation The first commercial unit operating flexibility in terms of a greater range of EB conversion and a using XyMax technology was brought onstream in 2000 Since then five lower temperature requirement The second technology referred to as additional total of six applications of the XyMax technology have been XyMax further increases yield performance and debottleneck potential eee netuaing omer moronMobialene isomerization technologies This technology can operate at even higher EB conversion with higher ere are a total uns IN operation selectivity and significantly lower xylene loss Licensor ExxonMobil Chemical retrofit applications Axens Axens NA Operating conditions XyMax and AMHAI units operate with a high grassroots applications space velocity and a low hydrogentohydrocarbon ratio which results in increased debottleneck potential and decreased utilities costs By converting a high portion of EB in the feed these technologies can ee eee provide significant savings in associated paraxylene recovery facilities id PROCESSING PetrochemicalProcesses PROCESSING JLGOIOU home processes index company index Xylene isomerization Application The Isomar process isomerizes Cg aromatics to mixed xy lenes to maximize the recovery of paraxylene in a UOP aromatics com plex Depending on the type of catalyst used ethylbenzene EB is also ae ene converted into xylenes or benzene re Description The Isomar process reestablishes an equilibrium distribution Overhead of xylene isomers essentially creating additional paraxylene from the re Fa tO maining ortho and metaxylenes The feed typically contains less than 1 wt of paraxylene and is first combined with hydrogenrich recycle gas 1 J and makeup gas The combined feed is then preheated by an exchanger Maieten 1 with reactor effluent heated in a fired heater 2 and raised to the eg Recycle gas Product to Xylene reactor operating temperature The hot feed vapor is then sent to the fractionation reactor 3 where it is passed radially through a fixedbed catalyst The reactor effluent is cooled by exchanger with the combined feed and then sent to the product separator 4 Hydrogenrich gas is taken off the top of the product separator and recycled back to the reac tor Liquid from the bottom of the products separator is charged to the deheptanizer column 5 The Cz overhead from the deheptanizer is Investment US million 29 cooled and separated into gas and liquid products The gas is exported Utilities per mt of feed to the fuel gas system and the liquid is sent to a debutanizer column Electricity kWh 32 or a Stripper The Cg fraction from the bottom of the deheptanizer is Steam mt 0065 recycled back to a xylene column Water cooling m 36 There are two broad categories of xylene isomerization catalysts Fuel Geal 0096 EB isomerization catalysts which convert ethylbenzene into additional ommercial plants UOP has licensed more isomerization units than any xylenes and EB dealkylation catalysts which convert ethylbenzene to other licensor in the world The first lsomar unit went onstream in 1968 valuable benzene coproduct The selection of the isomerization catalyst since that time UOP has licensed a total of 61 Isomar units depends on the configuration of the UOP aromatics complex the composition of the feedstocks and the desired product slate Licensor UOP LLC Economics A summary of the investment cost and the utility consump tion for a typical lsomar unit processing capacity of 184 million mtpy is shown below The estimated inside battery limits ISBL erected cost for a the unit assumes construction on a US Gulf Coast site in 2003 Gulf Publishing Company provides this program and licenses its use throughout the world You assume responsibility for the selection of the program to achieve your intended results and for the installation use and results obtained from the program LICENSE You may 1 Use the program on a single machine 2 Copy the program into any machine readable or printed form for backup or modification purposes in support of your use of the program on 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