Haber-Bosch Process and Its Evolution in Ammonia Production

The HaberBosch Heritage The Ammonia Production Technology Haber Bosch Mittasch Max Appl 50th Anniversary of the IFA Technical Conference September 25 26th 1997 Sevilla Spain Introduction Based on the fundamental research work of Fritz Haber Carl Bosch and his engineering team devel oped the ammonia synthesis to technical operability using the promoted ironbased catalyst found by Alwin Mittasch and coworkers Since then there has been no fundamental change in the synthesis reaction itself Even today every plant has the same basic con figuration as this first plant A hydrogennitrogen mix ture reacts on the iron catalyst todays formula dif fers little from the original at elevated temperature in the range 400 500 C originally up to 600 C operating pressures above 100 bar and the uncon verted part of the synthesis gas is recirculated after removal of the ammonia formed and supplemented with fresh synthesis gas to compensate for the amount of nitrogen and hydrogen converted to ammonia 3H2 N2 2NH3 1 H0298 924 kJmole F0298 328 kJmole Of course progress made in mechanical and chemi cal engineering and increased theoretical knowledge have led to improvements in efficiency converter design and energy recovery in the synthesis section but really dramatic changes happened over the years in the technology of synthesis gas generation As the synthesis is the very heart of every ammonia production and is also from an historical point of view the most interesting section it is probably appropri ate to start our review with this section The synthesis The ammonia equilibrium and the recycle concept The reaction proceeds with a reduction in volume and is also exothermic so the equilibrium concentrations of ammonia are higher at high pressure and low tem perature but at the turn of the last century a quan titative knowledge of chemical equilibrium was not available and this might explain why early experi ments aimed at the ammonia synthesis were unsuc cessful A famous victim of the lack of thermodynamic data was Wilhelm Ostwald He offered in 1900 BASF a process in which nitrogen and hydrogen were passed over heated iron wire at atmospheric pressure claim ing several percent of ammonia a concentration which was far beyond equilibrium BASF found the reason for his erroneous data and irony of history he withdraw his patent application not knowing how important that application could have been later when indeed iron became the basis of the commer cial ammonia synthesis catalyst First systematic measurements were made by Haber in 190405 but they yielded too high figures as a con sequence of problems with exact analysis of the low concentrations values attained at atmospheric pres sure and 1000 C using iron for catalysis As this fig ures did not comply with the Heat Theorem W Nernst made own measurements at 75 bar which were actually the first experiments at elevated pressure From the results he concluded that a technical pro cess which he probably anticipated as a oncethrough process should not be feasible as the much higher pressures needed in this case seemed to be beyond the technical possibilities of the time Haber continued with his investigations now also including pressure experiments From the more reliable equilibrium data now avail able it was obvious that at normal pressure the reac tion temperature should be kept well below 300 C in order to obtain even a small percentage of ammo nia For this temperature level no catalyst was 2 Figure 1 Equilibrium conversion and space time yield for NH3 and SO3 production available By increasing the pressure for example to 75 bar the equilibrium conditions were improved but with catalysts active at 600 C only low ammonia con centrations were attained So Haber concluded that much higher pressures had to be employed and that perhaps more importantly a recycle process had to be used an actually new process concept at that time and thus he overcame his collegues preoccupation which resulted from the unfavorable equilibrium con centrations and the concept of a oncethrough process The amount of ammonia formed in a single pass of the synthesis gas over the catalyst is indeed much too small to be of interest for an economic production Haber therefore recycled the unconverted synthesis gas after separating the ammonia formed by conden sation under synthesis pressure and supplementing it with fresh synthesis gas to make up for the portion which was converted to ammonia the gas was recir culated by means of a circulation compressor to the catalyst containing reactor Habers recycle idea changed the static conception of process engineering in favor of a more dynamic approach For the first time reaction kinetics were considered as well as the thermodynamics of the system In addition to the chemical equilibrium Haber recognized that for the technical realization reaction rate was a determining factor Instead of simple yield in a oncethrough process he concentrated on space time yield Figure 1 illustrates this consideration of equilibrium concentration in combination with space time yield by a comparison of the ammonia synthe sis with the SO2 oxidation process Also anticipated by him was the preheat of the syn thesis gas to reaction by heat exchange with the hot effluent gas from the reactor In 1908 Haber approached the BASF to find support for his work and to discuss the possibilities for the realisation of a technical process Early in 1909 he discovered in finely distributed osmium a catalyst which yielded 8 Vol of ammonia at 175 bar and 600 C A success ful demonstration in April 1909 of a small labscale ammonia plant convinced the representatives of BASF and the companys board decided to pursue the technical development of this process with all avail able resources In BASF then Carl Bosch entrusted with extraordi nary authority became project leader and succeeded together with a team of dedicated and very able coworkers to develop in an unprecendented effort a commercial process in less than five years The pro duction facilities for 30 td were erected on a new site near the village Oppau now a part of the city of Lud wigshafen the first production was in September 1913 and full capacity was reached in 1914 The ammonia catalyst In BASF Alwin Mittasch was responsible for the cat alyst search Osmium used by Haber showed excel lent catalytic activity but was difficult to handle the main disadvantage however was that the worlds stock of this rare material was only a few kilograms Mittasch started a systematic screening program covering nearly all elements of the periodic table Until 1910 more that 2500 different formulas were tested in 6500 runs For these experiments special small test reactors containing easily removable car tridges holding about 2 g of catalyst were developed In November 1909 a sample of magnetite from a place in Sweden showed exceptionally good yields which was surprising because other magnetite types were total failures Mittasch concluded that certain impur ities in this Gallivara magnetite were important for 3 Figure 2 Ammonia equilibrium and catalyst volume its good performance So he investigated the influence of various individual additives which in todays ter minology are called promoters By 1911 the catalyst problem had been solved Iron with a few percent alu mina and a pinch of potassium yielded a catalyst with acceptable reproducibility and performance and tol erable lifetime But the research program was continued until 1922 to be certain about the optimum composition The only additional result was that the further addition of calcium gave a certain improvement All magnetite based catalysts on the market today have a similar composition to that of the original BASF catalyst Also the catalyst preparation remained practically the same Melting natural magnetite from Sweden with addition of the various promoters cooling the melt breaking the solidified melt into small particles fol lowed by screening to obtain a fraction with suitable particle size From these early days until today an enormous amount of academic research was dedicated to elu cidate the mechanism of the synthesis to study the microstructure of the catalyst and to explain the effect of the promoters Besides the scientific interest there was of course some hope to find an improved cata lyst which could operate at far lower temperatures and thus at lower pressures saving compression energy which is in a modern plant still 300 kWht NH3 In principle one can operate with the classic magnetite catalyst at 35 45 bar in the temperature range of 350 to 450 C but needing a trainload of cat alyst about 450 m3 1300 t for a plant of 1350 td NH3 to achieve very low ammonia concentrations which would require removal by waterscrubbing instead of condensation by refrigeration M W Kel logg proposed such a process in the early 1980s but didnt succeed with commercialization For a real low pressure catalyst operating at front end pressure to need no compression an operation temperature well below 300 C would be required To illustrate this sit uation figure 2 shows ammonia equilibrium and cat alyst volume With the modern spectroscopic tools of Surface Sci ence rather detailed information on the reaction mechanism at the catalyst surface was obtained Kinetics of nitrogen and hydrogen adsorption and desorption were investigated and adsorbed interme diate species could be identified The results allowed to explain for the most part the mechanism of ammo nia synthesis in the pressure range of industrial inter est This success has many fathers outstanding con tributions were made by Brill Ertl Somorjai Bou dard Nielsen Scholtze Schlögl and many others The rate determining step is the dissociative adsorption of the nitrogen at the catalyst surface and the most active sites are the crystal faces 111 and 211 which is probably caused that these are the only surfaces which expose C7 sites which means iron atoms with seven nearest neighbors The primary function of the Al2O3 is to prevent sin tering by acting as a spacer between the small iron platelets and it may in part also contribute to stabi lize the Fe111 facets The promoting effect of the potassium is probably based on two factors One is the lowering of the activation energy of the dissociative adsorption of nitrogen by an electronic charge trans fer effect from potassium to iron which increases the nitrogen bond strength to the iron and weakens the nitrogennitrogen bond The other factor consists in reducing the adsorption energy of ammonia thus eas ing the desorption of the formed ammonia which avoids blocking the surface and hindering the nitro gen adsorption Commonly the term ammonia catalyst is used for the oxidic form consisting of magnetite and the promot ers Actually this is only the catalyst precursor which is transformed into the active catalyst consisting of α iron and the promoters by reduction with synthesis gas In the 1980s prereduced ammonia catalysts found acceptance in the market as they avoid the relatively long insitu reduction which causes additional down time and considerable feedstock consumption with out production These catalysts are reduced at the vendors facilities and subsequently passivated at tem peratures around 100 C using nitrogen with a small amount of air A notable improvement of the magnetite system was the introduction of cobalt as an additional component by ICI in 1984 The cobalt enhanced formula was first used in an ammonia plant in Canada using ICI Catalcos AMV process with a synthesis pressure of 90 bar With similar kinetic characteristics the vol umetric activity is about two times higher than that of the standard iron catalyst In October 1990 Kellogg commercialized the Kellogg Advanced Ammonia Process using a catalyst com posed of ruthenium on a graphite support which is 4 claimed to be 10 20 times as active as the traditional iron catalyst According to original patents asigned to BP the new catalyst is prepared by subliming ruthe niumcarbonyl Ru3CO12 onto the carboncontain ing support which is impregnated with rubidium nitrate The catalyst has a considerably higher surface than the conventional catalyst and according to the patent example it should contain 5 Ru and 10 Rb by weight This catalyst works best at a lower than stoichiometric HN ratio of the feed gas and it is also less susceptible to selfinhibition by NH3 and has an excellent low pressure activity The potential of ruthenium to displace iron in new plants will depend on whether the benefits of its use are sufficient to compensate the higher costs In com mon with the iron catalyst it will also be poisoned by oxygen compounds Even with some further poten tial improvements it seems unlikely to reach an activ ity level which is sufficiently high at low temperature to allow an operation of the ammonia synthesis loop at the pressure level of the syngas generation The ammonia converter and the synthesis loop configuration With the catalyst at hand the next step was to con struct somewhat larger test reactors for catalyst charges of about 1 kg Surprisingly these reactors rup tured after only 80 hours Further studies showed that the internal surface had totally lost its tensile strength This phenomenon had apparently propagated from the inner surface outward until the residual unaffected material was so thin that rupture occurred With the aid of microscopic investigations by thin sec tion technique Bosch found the explanation Decar bonization of the carbon steel had occurred but sur prisingly the result was not soft iron but rather a hard and embrittled material Hydrogen diffusing into the steel caused decarbonization by methane formation This methane entrapped under high pressure within the structure of the material led to crack formation on the grain boundaries which finally resulted in embrittlement Systematic laboratory investigations and material tests demonstrated that all carbon steels will be attacked by hydrogen at high temperatures and that the destruction is just a matter of time Boschs unconventional solution to the embrittlement problem was to use a carbon steel pressure shell with a soft iron liner To prevent the hydrogen which had penetrated this liner from attacking the pressure shell measures had to be taken to release it safely to nor mal pressure This was achieved by providing small channels on the outer side of the liner which was in tight contact with the inner wall of the pressure shell and by drilling small holes later known as Bosch Holes through the pressure shell through which hydrogen could escape to the atmosphere These holes had no effect on the strength of the shell and the resulting losses of hydrogen were negligible Fig ure 3 gives a sketch of such a pilot plant converter Bosch did not content himself with his linerhole con cept but looked further for alternative solutions for the embrittlement problem He intitiated in the late 1920s research in the steel industry to develop steels resistant to hydrogen under pressure Special alloy components as for example molybdenum chromium tungsten and others form stable carbides and enhance the resistance of steel against this sort of attack con siderably This problem and the related physical hydrogen attack is not restricted to the synthesis but has to be considered carefully also in the synthesis gas production section because of the temperatures and 5 Figure 3 First pilot plant converter with soft iron lining and external heating hydrogen partial pressures involved there Extensive research and careful evaluation of operation experi ences have made it possible to prevent largely hydro gen attack in modern ammonia plants by proper selec tion of hydrogentolerant alloys with the right con tent of metals which form stable carbides Of fundamental significance in this respect was the work of Nelson who produced curves for the stability of various steels as a function of operation temperature and hydrogen partial pressure In the small reactors heat losses predominated and continuous direct external heating by gas was neces sary and this led to deterioration of the pressure shells after short operation times even without hydrogen attack With increasing converter dimensions in the commercial plant heating was only necessary for start up Bosch developed an internal heating by the socalled inversed flame introducing at the top of the reactor a small amount of air igniting with an electrically heated wire Later this was replaced by an electric resistance heater Subsequently introduced flushing with nitrogen as shown in figure 4 and later with cold synthesis gas kept the pressure vessel walls cool and rendered the linerhole concept redundant Subsequent reactor designs in the technical plant included internal heat exchangers and later the cat alyst was placed in separate tubes which were cooled by the feed gas Another improvement was the intro duction of an externally insulated catalyst basket Because of the low concentrations aqueous ammonia was separated from the loop by water scrubbing Con verters with catalyst tubes had a better temperature control and this led together with an increased pres sure to higher ammonia concentrations which now allowed from 1926 onwards the direct production of liquid ammonia In 1942 the first quench converter was installed and this design gradually has replaced then the converters with the catalyst tubes Soon after the first world war development started also in other countries partly on basis of BASFs pio neering work Luigi Casale built 1920 the first plant in Italy and based on developments by M G Claude the first French plant started to produce in 1922 Both the Casale and the Claude process operated under extreme high pressure In contrast to this Uhde con structed a plant based on coke oven gas operating under extreme low pressure Mont Cenis process Futher developments were by G Fauser who worked together with Montecatini During the 1920s several plants were built in the USA some based on Euro pean some on American Technology The successful US company was Nitrogen Engineering Corporation NEC the predecessor of Chemico Mechanical design was now already rather advanced but for the process design of converter and loop so far empirical data in form of charts were used as no suitable mathematical expressions for the reaction kinetics were at hand When better experimental data for the reaction kinetics and other process variables became available in the 1940s and 1950s layout of converters received a better quantitative chemical engineering basis Figure 5 shows reaction rate of ammonia formation and equilibrium When the tem perature is increased under otherwise constant con ditions the reaction rate increases to a maximum to decrease with further temperature increase and becomes zero when reaching equilibrium tempera ture Joining these points will result in a line giving for each NH3 concentration the temperature for the maximum rate This curve runs about parallel to the 6 Figure 4 Converter with pressure shell cooling by nitrogen equilibrium line and at a about 30 50 C lower tem perature To maintain the maximum ammonia forma tion rate the reaction temperature must decrease as the ammonia concentration increases For optimal catalyst usage the reactor temperature profile after a initial adiabatic heating zone in the first part of the catalyst should follow this ideal line For a long time converters were always compared to this ideal for optimum use of highpressure vessel volume Today the objective is rather to maximize heat recovery at the highest possible level and to minimize investment costs for the total synthesis loop In any case it is necessary to remove the heat of reac tion as the conversion proceeds to keep the temper ature at an optimal level For the removal of the reaction two principal configurations are possible Tubular converters have cooling tubes within the cat alyst bed through which the cooling medium usually cooler feed gas flows cocurrently or countercur rently to the gas flow in the catalyst bed Alternatively the catalyst can be placed within tubes with the cool ing medium flowing on the outside The tube cooled converters dominated until the early fifties but are largely outdated today Well known examples were the TVA converter countercurrent and the NECChemico design cocurrent with best approx imation to the maximum rate curve An interesting revival of this principle is the ICI tube cooled con verter used in the LCA process and also for metha nol production In the multibed converters the catalyst volume is divided into several beds in which the reaction pro ceeds adiabatically Between the individual catalyst layers heat is removed either by injection of colder synthesis gas quench converters or by indirect cool ing with synthesis gas or via boiler feed water heat ing or steam raising indirectly cooled multibed con verter 7 Figure 5 Reaction rate of ammonia formation Figure 6 Quench converter In the quench converters only a fraction of the recy cle gas enters the first catalyst layer at about 400 C The catalyst volume of the bed is chosen so that the gas will leave it at around 500 C Before entering the next catalyst bed the gas temperature is quenched by injection of cooler 125 200 C recycle gas The same thing is done at subsequent beds In this way the reaction profile describes a zigzag path around the maximum reaction rate line A schematic drawing of a quench converter together with its tempera turelocation and temperatureammonia concentra tion profile is presented in figure 6 The catalyst beds may be separated by grids designed as mixing devices for main gas flow and quenchgas cold shot or be just defined by the location of cold gas injection tubes as for example in the ICI lozenge converter A disadvantage is that not all of the recycle gas will pass over the whole catalyst volume with the conse quence that a considerable amount of the ammonia formation occurs at higher ammonia concentration and therefore at reduced reaction rate This means that a larger catalyst volume will be needed compared to an indirect cooled multibed converter On the other hand no extra space is required for interbed heat exchangers so that the total volume will remain about the same as for the indirect cooled variant As the quench concept was well suited for large capac ity converters it had a triumphant success in the early generation of large single stream ammonia plants con structed in the 1960s and 1970s Mechanical simplic ity and very good temperature control contributed to the widespread acceptance Multibed converters with indirect cooling In convert ers of this category the cooling between the individ ual beds is effected by indirect heat exchange with a cooling medium which may be cooler synthesis gas andor boiler feed water warming and steam raising The heat exchanger may be installed together with the catalyst beds inside one single pressure shell but an attractive alternative too preferentially for large capacities is to accommodate the individual catalyst beds in separate vessels and have separate heat exchangers This approach is especially chosen when using the reaction heat for raising high pressure steam The indirect cooling principle is applied today in almost all large new ammonia plants and also in revamps an increasing number of quench converters are modified to the indirect cooling mode Axial flow through the catalyst in the converters as exclusively used until the early 1970s face a general problem With increasing capacity the depth of the catalyst beds will increase as for technical and eco nomical reasons it is not possible to enlarge the pres sure vessel diameter above a certain size In order to compensate for the increasing pressure drop axial flow converters with usual space velocities of 10 15000 h1 have to use relatively large catalyst par ticles and a particle size of 6 10 mm has become stan dard But this grain size has compared to finer cata lyst a considerably lower activity which decreases approximately in a linerar inverse relation Two fac tors are responsible for the lower activity of the larger particles Firstly the larger grain size retards on account of the longer pores the diffusion from the interior to the bulk gas stream and this will inhibit the dissociative nitrogen adsorption and by this the reac tion rate Secondly the reduction of an individual cat alyst particle starts from the outside and proceeds to the interior The water formed by removing the oxy gen from the iron oxide in the interior of the grains will pass over already reduced catalysts on its way to the outer surface of the particle This induces some recrystallization leading to the lower activity The effect is considerable going from a partide size of 1 mm to one of 8 mm the inner surface will decrease from 1116 to 3 8 m2g 8 Figure 7 Topsoe Series 200 indirect cooled converter radial flow Haldor Topsøes company solved the dilemma with the pressure drop and small catalyst particles with a radial flow pattern using a grain size of 15 3 mm Fig ure 7 MW Kellogg chose another approach with its hori zontal crossflow converter Fig ure 8 The catalyst beds are arranged side by side in a car tridge which can be removed for catalyst loading and unloading through a fullbore closure of the horizontal pressure shell Today each new worldsize ammonia plant employs the indirect cooling concept raising high pressure steam up to 125 bar Generally after the first bed an inletoutlet heatexchanger is placed and after the sec ond or further beds the reaction heat is used to raise high pressure steam Brown and Root formerly C F Braun or Uhde Figure 9 accommodate the cat alyst in several vessels Figure 9 is a simplified flow sheet of Uhdes synthesis loop Actually the concept of separate vessels for the catalyst beds with heat exchange after the first and waste heat boiler after the 9 Figure 8 Indirect Cooled Horizontal Converter of M W Kellogg Figure 9 Uhdes synthesis loop with two pressure vessels and three catalyst beds second nowadays they use also a third one followed by a boiler too was already introduced by C F Braun at time when most plants still used quench convert ers The Ammonia Casale ACAR Converter has a mixed flow pattern In each catalyst layer the gas flows through the top zone predominantly axially but tra verses the lower part in radial direction This simpli fies the design by avoiding special sealing of the top end of the bed to prevent bypassing Today computerized mathematical models are used for converter and loop layout In principle these models use two differential equations which describe the steady state behavior of the reaction in the converter The first gives a concentrationlocation relationship within the catalyst bed for the reactants and the ammonia It reflects the reaction kinetic expression The second models the temperatureposi tion relationship for the synthesis gas catalyst and vessel internals The form of this equation is specific to the type of the converter The kinetics of the intrinsic reaction that means the reaction on the catalyst surface without any mass transport restrictions are derived from measurements on very fine catalyst particles The first useful expres sions for engineering purposes to describe the reac tion rate was the TemkinPyshew equation proposed in 1940 It was widely applied but today there are improved versions and other equations available Additional terms are included to model the influence of oxygencontaining impurities on the reaction rate Although oxygencontaining compounds may be regarded as a temporary poison severe exposure for an extended period of time leads to permanent dam age For practical application these equations have to be modified to make allowance for transport phenom ena heat and mass transfer and this is done by socalled pore effectiveness factors 10 Figure 10 Simplified flow sheet of a cokebased ammonia plant Synthesis Gas Preparation The classical route based on coke The pilot plant experiments at BASF for the ammo nia synthesis were based on hydrogen from the chlo rinealkali electrolysis When the capacity of this gas source was exhausted water gas served as an indepen dent hydrogen feedstock using the cryogenic Linde Fränkl process for the separation In this process car bon monoxide is condensed out of the water gas at 200 C and 25 bar Nitrogen was provided by an air separation unit and nitrogen was also used in an indi rect liquid nitrogen circulation system in the cryogenic hydrogen separation The residual content of 15 CO in the hydrogen was removed by conversion to sodium formate in a gas scrubber operated with a l0 sodium hydroxide solution and at 230 C and 200 bar The initial operation of the commercial plant commis sioned in September 1913 was based on hydrogen and nitrogen produced by this cryogenic separation but after a few months on line it became apparent that the Linde refrigeration process was not reliable and economic enough for the production on large scale A new catalytic process the shift conversion was introduced In this reaction found by W Wild in BASF already in 1912 the gas is passed together with a surplus of steam over an iron oxide chromium oxide catalyst at about 350 to 450 C The carbon monox ide reacts with water to form hydrogen and carbon dioxide The use of the shift reaction permitted a great simplification of the synthesis gas preparation Instead of using the refrigeration processes producer gas a mixture of 60 nitrogen and 40 carbon monoxide was generated by reacting air with red hot coke and mixed with the parallel generated water gas supplied by the alternating air blowing and steaming process and this mixture was converted in the shift reaction to yield a gas consisting of hydrogen nitrogen car bon dioxide and a small amount of residual carbon monoxide The carbon dioxide could then be removed satisfactorily by water scrubbing at 25 bar The removal of the residual carbon monoxide by scrub bing with hot caustic soda solution with formation of sodium formate used in the initial cryogenic route was corrosive and troublesome It could now be replaced by copper liquor scrubbing Water gas production from lignite started in 1926 in Leuna using a process developed by Winkler This process in which coal is gasified continuously with oxygen and steam in a fluidized bed was a spinoff of the research work on the removal of sulfur from ammonia synthesis gas Figure 10 is a simplified flow sheet of a coke based HaberBosch plant as it was operated in the 1930s and 1940s at BASF and elsewhere In the 1950s BASF developed and introduced continuously operated water gas generators using oxygen or oxygen enriched air from which the slag could withdrawn in liquid form A new age with hydrocarbons The plants continued to be based on coal for synthe sis gas generation until the 1950s With growing avail ability of cheap hydrocarbon feedstocks and novel cost saving gasification processes a new age dawned in the ammonia industry The development started in the USA where steam reforming was introduced a pro cess originally developed in the 1930s by BASF and greatly improved by ICI which extended it also to naphtha Before natural gas became available in large quantities in Europe too partial oxidation of heavy oil fractions was used in several plants with process technology developed by Texaco 1940 and Shell 1950 After several oil crisis coal gasification research and development was resumed with the result that for this route a few technically proven pro cesses are available today The chemical reaction of water oxygen air or any combination of these reactants with fossil feedstocks is generally described as gasification In a simplified way it can be viewed as the reduction of water by means of carbon and carbon monoxide It yields a gas mixture made up of carbon monoxide and hydrogen in various proportions along with carbon dioxide and where air is introduced some nitrogen CHx H2O CO H2 x2H2 H 0 2 CHx 12 O2 CO x2H2 H 0 3 Reaction 2 is endothermic and needs an external source of energy supply whereas reaction 3 is exo thermic and can be carried out adiabatically For the initial carbon dioxide content in the raw gas from the gasification the shift reaction equilibrium is respon sible which at the high temperature is rather on the CO side CO H2O CO2 H2 H0298 412 kJmol 4 11 This shift reaction in which actually CO reduces water to yield additional hydrogen is favored by low temperature and is therefore purposely made to pro ceed on a catalyst in a separate step at a temperature lower than the preceding gas generation step With coke the reaction 2 corresponds to the non catalytic classic water gas process With light hydro carbons reaction 2 is called steam reforming and is made to proceed over a nickel catalyst The reaction 3 is commonly called partial oxidation and in prin ciple applicable for any fossil feedstock from coal to natural gas As can be seen from the stoichiometric equation the hydrogen contributed by the feedstock itself increases with its hydrogen content which ranges from a minimum of CH01 in coke to a maximum of CH4 in methane Syngas preparation via steam reforming The steam reforming process is restricted to light hydrocarbons ranging from natural gas methane to light naphtha For higher hydrocarbons such as fuel oil or vacuum residue this technology is not applicable on account of impurities as sulfur and heavy metals which would poison the sensitive nickel catalyst In addition cracking reactions are more likely to occur on the catalyst depositing carbon which might block the catalysts pores and also restrict the gas flow As the nickel catalysts are highly sensitive to sulfur com pounds these catalysts poisons have to be removed prior to the reforming reaction For this purpose any organic sulfur compounds contained in the hydrocar bon feedstock are first hydrogenated on a cobalt molybdenum catalyst to hydrocarbon and hydrogen sulfide which is then absorbed with zinc oxide to form zinc sulfide RSH H2 H2S RH 5 H2S ZnO ZnS H2O 6 For ammonia production the steam reforming is per formed in two steps First the hydrocarbon steam mixture is passed through highalloyed nickelchro mium tubes filled with a catalyst containing finely dis persed nickel on a carrier The heat needed for the endothermic reaction is supplied by gas burners in a furnace box The reaction in this primary reformer is controlled to achieve only a partial conversion of around 65 leaving about 14 methane dry basis content in the effluent gas at a temperature of 750 to 800 C The gas is then introduced into the socalled secondary reformer a refractory lined vessel also with a nickel catalyst where it is mixed with a con trolled amount of air introduced through a burner This raises the temperature sufficiently to complete the reforming of the residual methane adiabatically It also introduces the right amount of nitrogen to achieve the correct stoichiometric ratio in the final synthesis gas The overall reaction in the secondary reformer may be described as some sort of a partial oxidation but the stoichiometric equation 7 does not give a clue to the actual reactions taking place 2CH4 O2 4N2 2CO 4H2 4N2 H0298 714 kJmol 7 The gas leaves the secondary reformer at 950 to 1000 C and a methane content of 03 to 15 It is cooled down to 350 400 C using the removed heat for high pressure steam generation In the first steam reforming based plants the shift conversion used only the classical chromiumiron catalyst achieving around 2 residual CO For CO2 removal in this early plants the traditional water scrubbing was applied and the final purification was still performed by copper liquor In the early 1960s copperzincalu mina catalysts became available for a second conver sion step at temperatures of about 200 C whereby the residual CO concentration could be lowered to 02 03 This allowed to eliminate the copper liq uor scrubbing removing the residual concentrations of CO and CO2 by methanation In this highly exo thermic reaction which is performed at about 300 C on a nickel catalyst hydrogen reacts with carbon monoxide to methane and water it is the reverse of the steam reforming reaction of methane equation 8 and 9 CO 3H2 CH4 H2O H0298 2063 kJmol 8 CO2 4H2 CH4 2H2O H0298 1651 kJmol 9 With aqueous monoethanolamine MEA a new sol vent for CO2 removal was introduced in 1943 This process has been used extensively in many ammonia plants until hot potash and other solvents with lower heat requirement were developed The plants with capacities up to 300 td used reciprocating compres sors for compression As natural gas is usually delivered under elevated pres sure and because the reforming reaction entails an 12 increase in total volume significant savings of com pression energy are possible if the process is performed under higher pressure But there is also a disadvan tage in raising the pressure level of reforming as the equilibrium is shifted to lower conversions which can be compensated by higher temperatures As all the heat in the primary reformer has to be transferred through the tube wall the wall temperatures will rise and approach the material limits Originally HK 40 tubes with a content of 20 nickel and 25 chro mium were commonly used With new grades as HP modified with higher nickel content and stabilized with niobium and the recently introduced Micro Alloys which addition ally contain titanium and zirkonium higher wall temperatures and thus higher pressures up to 44 bar in the pri mary reformer have become possible The steam surplus applied in the reformer could thus also be reduced from a steam to carbon ratio of 4 and higher to about 3 or slightly below and this was assisted by improved catalysts with enhanced activity and better heat trans fer characteristic For naphtha reforming a higher steam surplus is necessary Fancy catalyst shapes as wagon wheels sixshooters shamrock or fourhole have replaced the old Raschig rings The stability of the standard catalyst supports as calcium aluminate magesium aluminate and αalumina has been improved and it has become a widely accepted pratice to install in the first third of the catalyst tube where the bulk of the reforming reac tion takes place a potassium promoted catalyst which was developed by ICI originally for naphtha steam reforming in order to prevent carbon deposition by cracking reactions From the various primary reformer designs the top fired concept with a single radiation box dominates in the larger plants the sidefired design in which only 2 tube rows can be accommo dated in the radiation box allows only a linear exten sion and additional fire boxes connected to a common flue gas duct The secondary reformers have been optimized regarding hydrodynamics and burner design using computational fluid dynamics Figure 11 shows an example of a topfired reforming furnace together with the secondary reformer The reduction of the steamtocarbon ratio was a bigger problem for the HT shift than in the reformer step as the gas mixture became a higher oxidative potential and tended to overreduce the ironoxide from magnetite to FeO and in extreme cases partially to metallic iron Under these conditions the Boudu ard reaction will become significant and carbon accu mulation in the catalyst particles leads to breaking In addition the FischerTropsch reaction leads to the formation of methane and higher hydrocarbons Cop per promotions of the iron catalyst suppresses these side reaction The nasty problem of methanol and amine formation in the LT shift is largely solved by 13 Figure 11 Topfired primary reformer and secondary reformer Uhde design Figure 12 CO2 Loading characteristics of various solvents improved formulations of the copperzincalumina and a new development is the intermediate temper ature shift catalyst operated quasi isothermal in a tubular reactor for example in the ICI LCA ammo nia process or the Linde ammonia process LAC Large progress in the CO2 removal systems was made in the last decade The original MEA systems had a heat consumption for solvent regeneration over 200 kJkmol a corrosion inhibitor system called amine guard III brought it down to about 120 kJkmol but this is still nearly 5 times as high as the most advanced system the BASF aMDEA Process which uses an aqueous solution of monomethyldiethanolamine together with a special promotor which enhances the mass transfer Other low energy systems are the Benfield LoHeat Process which is a hot potash system or the Selexol Process which uses a mixture of gly col dimethylethers a pure physical solvent In phys ical solvents a prominent example was water in the old plants the solubility of the CO2 is according to Henrys law direct proportional to the CO2 partial pressure and regeneration can be achieved by flash ing without application of heat In contrast to this the MEA is a chemical solvent the solubility is only slightly dependent on the CO2 partial pressure and approaches a saturation value MEA forms a stable salt with the carbon dioxide and a high amount of heat is required in the stripper to decompose it BASFs aMDEA Process is about in between the characteristic can be adjusted in a flex ible way by the concentration of the activator so that the major part of the dissolved carbon dioxide can be released by simple flashing and only a smaller propor tion has to be stripped out by heat Figure 12 shows CO2 loading characteristics of various solvents The tubular steam reformer has become a very reliable apparatus and the former problems with tube and trans fer line failures and catalyst difficulties are largely his tory But the tubular furnace and its associated convec tion bank is a rather expensive item and contributes sub stantially to the investment cost of the total ammonia plant So in some modern concepts the size was reduced by shifting some of the load to the secondary reformer necessitating an overstoichiometric amount of process air The surplus of nitrogen introduced in this way can be removed downstream by the use of a cryogenic unit CF Braun was the first contractor which introduced this concept in the socalled Purifier Process Some con tractors have gone so far to bypass some of the natu ral gas around the tubular reformer and feeding it directly to the secondary reformer which likewise needs surplus of process air or oxygen enriched air But there are additional reasons for breaking away fur ther from the fired furnace concept The temperature level of the flue gas from a traditional reformer is usu ally higher than 1000 C and the process gas at the out let of the secondary reformer is also around 1000 C It is thus from a thermodynamic point of view waste ful to use this high temperature level simply to raise and superheat high pressure steam The boiling point of HP steam is only 325 C and the first heat exchanger in the flue gas duct preheats process air in the conservative plants to only 500 C 600 700 C in more modern installations Recycling highlevel heat from the sec ondary reformer and making use of it for the primary reforming reaction is thermodynamically the better option Concepts which use this heat in an exchanger reformer have been successfully developed and com mercially demonstrated The first to come out with this concept in a real production plant was ICI with its GHR Gas Heated Reformer The hot process gas from the secondary reformer is the sole heat source A surplus of process air of around 50 is needed in the secon 14 Figure 13 ICI GasHeated Reformer dary reformer to achieve a closed heat balance Figure 13 is a simplfied drawing of the ICI GasHeated Reformer Quite recently ICI has come out with a modified design the AGHR with the A standing for advanced The bayonet tubes are replaced by normal tubes attached to a bottom tube sheet using a special packing which allows some expansion Thus the del icate double tubesheet is now eliminated In the Kellogg Exchanger Reformer System abbre viated KRES the gas flow pattern is different The tubes are open at the lower end and the reformed gas mixes with the hotter effluent of the secondary reformer The mixed gas stream flows upward on the shell side to heat the reformer tubes Thus primary reforming and secondary reforming reaction proceed in parallel in contrast to the ICI concept where the two reactions proceed in series The Kellogg process uses enriched air The complete elimination of the fired tubular furnace leads to a drastic reduction of NOx emission because there is only flue gas from much smaller fired heaters required for feed and process air preheat An even more progressive exchanger reformer presently operating in a demo plant is Uhdes CAR Combined autothermal reformer which not only replaces the catalytic sec ondary reforming step by a non catalytic partial oxy dation step but also combines this with the exchanger reformer in one single vessel Syngas from heavy oil fractions via partial oxidation In partial oxidation heavy oil fractions react accord ing to equation 2 with an amount of oxygen insuf ficient for total combustion The reaction is noncat alytic and proceeds in an empty vessel lined with alu mina refractory The reactants oil and oxygen along with a minor amount of steam are introduced through a nozzle at the top of the generator vessel The noz zle consists of concentric pipes so that the reactants are fed separately and react only after mixing at the burner tip in the space below The temperature in the generator is between 1200 and 1400 C Owing to the insufficient mixing with oxygen about 2 of the total hydrocarbon feed is transformed into soot which is removed by water scrubbing The separation of the soot from the water and its further treatment differs in the Shell and the Texaco Process the two commer cially available partial oxidation concepts The gas ification pressure can be as high as 80 bar After gas cooling by further waste heat recovery the hydrogen sulfide formed during gasification is removed along with carbon dioxide by scrubbing with chilled methanol below 30 C in the Rectisol pro 15 Figure 14 Ammonia syngas by partial oxidation of heavy hydrocarbons Texaco cess Then as in the steam reforming route the gas undergoes the CO shift reaction Because of the higher carbon monoxide content much more reaction heat is produced which makes it necessary to distrib ute the catalyst on several beds with intermediate cooling The carbon dioxide formed in the shift con version is removed in a second stage of the Rectisol unit both have a common methanol regeneration system The H2Srich carbon dioxide fraction from the first stage of the regenerator is fed to a Claus plant where elemental sulfur is produced In the final pur ification the gas is washed with liquid nitrogen which absorbs the residual carbon monoxide methane and a portion of the argon which was introduced into the process in the oxygen feed The conditions in this stage are set so that the stoichiometric nitrogen requirement is allowed to evaporate into the gas stream from the liquid nitrogen wash The process needs of course an air separation plant to produce oxygen usually around 985 pure and to supply the liquid nitrogen Figure 14 is a simplified flowsheet of synthesis gas preparation by partial oxidation of heavy fuel oil using the Texaco Syngas Generation Process The Shell process uses of a waste heat boiler for raw gas cooling whereas Texaco prefers for ammonia plants a water quench for this purpose which has the advantage that this intro duces the steam for the subsequent shift conver sion which different from Shell is performed without prior removal of the sulfur compounds using a sulfur tolerant shift catalyst Besides some optimiza tions there are no funda mental new develop ments in the individual process steps Some pro posed changes in the pro cess sequence for exam ple methanation instead of liquid nitrogen wash or the use of air instead of pure oxygen are not realized so far Though other CO2 removal systems as Selexol or Purisol NMethylpyrrol idon and alternative sulfur recovery processes are suitable too Rectisol and Claus Process remain the preferred options Synthesis gas by coal gasification There is no chance for a widespread use of coal as feed stock for ammonia in the near future but a few remarks should be made regarding the present status of coal gas ification technology Proven gasification processes are the Texaco Process the KoppersTotzek Process and the Lurgi Coal Gasification The Shell gasification not yet in use for ammonia production but successfully applied for other productions is an option too Texacos concept is very similar to its partial oxydation process for heavy fuel oil feeding a 70 coalwater paste into the generator KoppersTotzek is an entrained flow con cept too but feeding coal dust In the Lurgi process the coarse grounded coal is gasified in a moving bed at comparably low temperature using higher quantities of steam as the others Shells process differs consid erably from its oil gasification process in flow pattern and feeds coal dust Texaco Lurgi and Shell operate under pressure whereas the KoppersTotzek gasifier is under atmospheric pressure but a pressure version called PRENFLOW is presently tested in a demo plant Continuous slag removal either in solid or mol 16 Figure 15 Ammonia plant temperature profile ten form is indeed the fundamental technical problem with coalbased systems and the technical solutions dif fer considerably Gas cooling is achieved by quench and or waste heat boiler entrained coal dust is removed by water scrubbing The following process steps for shift conversion CO2 removal and final purification are largely the same as in partial oxdiation of heavy fuel oil Energy integration and ammonia plant concept The integrated steam reforming ammonia plant In the old days an ammonia plant was more or less just a combination with respect to mass flow and energy management was handled within the separate process sections which were often sited separately as they usually consisted of several parallel units A revolu tionary breakthrough came in the mid of the 1960s with the steam reforming ammonia plants The new impulses came more from the engineering and con tractor companies than from the ammonia plant industry itself Engineering contractors have been working since the thirties in the oil refining sector The growing oil demand stimulated the development of machinery vessel and pipe fabrication instrumenta tion and energy utilization leading to singletrain units of considerable size By applying the experience gained in this field it was possible to create within a few years in the mid 1960s the modern largescale ammonia concept To use a singletrain for large capacities no parallel lines and to be as far as possible energetically selfsufficient no energy import through a high degree of energy inte gration with process steps with surplus supplying those with deficit was the design philosophy for the new steam reforming ammonia plants pioneered by M W Kellogg and some others It certainly had also a revolutionary effect on the economics of ammonia production making possible an immense growth in world capacity in the subsequent years The basic 17 Table 1 Main energy sources and sinks in the steam reforming ammonia Process Process section Originating Contribution Reforming Primary reforming duty Demand Flue gas Surplus Process gas Surplus Shift conversion Heat of reaction Surplus CO2 removal Heat for solvent regeneration Demand Methanation Heat of reaction Surplus Synthesis Heat of reaction Surplus Machinery Drivers Demand Unavoidable loss Stack and general Demand Balance Auxiliary boiler or import Deficit Export Surplus reaction sequence has not changed since then Figure 15 shows the process sections and the relevant gas temperature levels in a steam reforming ammonia plant Highlevel surplus energy is available from the flue gas and the process gas streams of various sections while there is a need for heat in other places such as the process steam for the reforming reaction and in the solvent regenerator of the carbon dioxide removal unit Table 1 Because a considerable amount of mechanical energy is needed to drive compressors pumps and fans it seemed most appropriate to use steam turbine drives since plenty of steam could be generated from waste heat As the temperature level was high enough to raise HP steam of 100 bar it was possible to use the process steam first to generate mechanical energy in a turbine to drive the synthe sis gas compressor before extracting it at the pressure level of the primary reforming section The earlier plants were in deficit and they needed an auxiliary boiler which was integrated in the flue gas duct This situation was partially caused by inadequate waste heat recovery and low efficiency in some of the energy consumers Typically the furnace flue gas was discharged up the stack at unnecessarily high temper atures because there was no combustion air preheat and too much heat was rejected from the synthesis loop while the efficiency of the mechanical drivers was low and the heat demand in the carbon dioxide removal unit regenerator was high A very important feature of this new concept was the use of a centrifugal compressor for synthesis gas com pression and loop recycle One advantage of the cen trifugal compressors is that they can handle very large volumes which allows also for the compression duties a single line approach The lower energetic efficiency compared to the reciprocating compressors of which in the past several had to be used in parallel is more than compensated by the lower investment and the easy energy integration In the first and also the sec ond generation of plants built to this concept max imum use was made of direct steam turbine drives not only for the major machines such as synthesis gas air and refrigeration compressors but even for relatively small pumps and fans The outcome was a rather com plex steam system and one may be tempted to describe an ammonia plant as a sophisticated power station making ammonia as a byproduct The plants produce more steam than ammonia even today the most modern plants still produce about three times as much In recent years electrical drives have swung back into favor for the smaller machines In most modern plants total energy demand feedfuelpower has been drastically reduced On the demand side important savings have been achieved in the carbon dioxide removal section by switching from old heatthirsty processes like MEA 18 Figure 16 Simplified flow sheet of a modern steam reforming ammonia plant CF Braun Purifier Process scrubbing to lowenergy processes like the newer ver sions of the Benfield process or aMDEA Fuel is saved by air preheat and feed by hydrogen recovery from the purge gas of the synloop by cryogenic mem brane or pressure swing adsorption technology In the synthesis loop the mechanical energy needed for feed compression refrigeration and recycle has been reduced and throughout the process catalyst volumes and geometry have been optimized for maximum activity and minimum pressure drop On the supply side available energy has been increased by greater heat recovery and the combined effect of that and the savings on the demand side have pushed the energy balance into surplus Because there is no longer an auxiliary boiler there is nothing in the plant that can be turned down to bring the energy sit uation into perfect balance therefore the overall sav ings have not in fact translated into an actual reduc tion in gross energy input to the plant in the form of natural gas they can only be realized by exporting steam or power and it is only the net energy consump tion that has been reduced But under favorable cir cumstances this situation can be used in a very advan tageous way If there is a substantial outlet on the site for export steam it can be very economic depend ing on the price of natural gas and the value assigned to steam to increase the steam export deliberately by using additional fuel because the net energy con sumption of the plant is simultaneously reduced It is only possible to reduce the gross energy demand that is to reduce the natural gas input to the plant by reducing fuel consumption because the feedstock requirement is stoichiometric So the only way is to cut the firing in the reforming furnace by shifting reform ing duty to the secondary reformer as we had already discussed earlier or to choose a more radical aproach by the use of an exchanger reformer instead of the fired furnace ICIs GasHeated Reformer GHR system the KRES of M W Kellogg and the Tandem Reformer now marketed by Brown Root or the even more advanced Combined Autothermal Reformer CAR of Uhde But none of these designs necessarily achieves any significant improvement over the net energy consumption of the most advanced con ventional concepts under the best conditions For the cases in which export of steam andor power is welcome there is the very elegant possibility of inte grating a gas turbine into the process to drive the air compressor The hot exhaust of 500 550 C contains well enough oxygen to serve as preheated combustion air for firing the primary reformer The gas turbine does not even have to be particularly efficient because any heat left in the exhaust gas down to the flue gas temperature level of 150 C is used in the fur nace Thus an overall efficiency of about 90 can be achieved 19 Boiler makers provide today largely reliable designs for highduty waste heat boilers after secondary reformer and in the synthesis loop in which up to 15 t steamt NH3 are produced corresponding roughly to a recovery of 90 of the reaction enthalpy of the synthesis Centrifugal compressors have become much more reliable though their efficiency has not increased spectacularly in recent years Some improvements were made in turndown capability in improving the surge characteristic New developments are dry seals instead of oil seals and another poten tial improvement already successfully introduced in nonammonia applications is the magnetic bearing Although the introduction of the singletrain inte grated large plant concept in the 1960s revolutionized the energyeconomics of ammonia production it is surprising that since then the total consumption has been reduced by about 30 from roughly 40 to 28 GJt An example of a modern plant shows Figure 16 From this enormous reduction in energy consumption the question may come up what is the theoretical min imum energy consumption for ammonia production via steam reforming of natural gas Based on pure methane we may formulate the following stoichio metric equation CH4 03035 O2 1131 N2 1 393 H2O CO2 2262 NH3 10 H0298 86 kJmol F0298 101 kJmol So from a mere thermodynamic point of view in an ideal engine or fuel cell heat and power should be obtained from this reaction But because there is a high degree of irreversibility in the real process a con siderable amount of energy is necessary to produce the ammonia from methane air and water The stoi chiometric quantity of methane derived from the fore going equation is 583 Nm3 per mt NH3 which corre sponds to 209 GJ LHV per tonne of ammonia which with some reason could be taken as minimum value Of course if one assumes full recovery of the reaction heat then the minimum would be the heating value of ammonia which is 186 GJ LHV per mt NH3 Energy and exergy anal ysis First and Second Law of Thermodynamics respectively identify the process steps in which the biggest losses occur The biggest energy loss is in the turbines and com pressors whereas the exergy loss is greatest in the reforming section almost 70 Based on exergy the thermody namic efficiency for the ammonia production based on steam reform ing of natural gas is almost 70 It has become rather common to measure modern ammonia concepts above all by their energy consump tion Yet these comparisons need some caution in interpretation without a precise knowledge of design bases physical state of the produced ammonia and state of the utilities used eg cooling water temper ature nitrogen content in natural gas or conversion factors used for evaluating imported or exported steam and power misleading conclusions may be drawn In many cases too the degree of accuracy of such figures is overestimated The best energy consumption values for ammonia plants using steam reforming of natural gas are around 28 GJtNH3 Industrial figures reported for plants with highduty primary reforming and stoichometric pro cess air and for those with reduced primary reform ing and excess air show practical no difference 20 Figure 17 Flow diagram of ICIs LCA Ammonia Process Core unit for 450 mtpd New steam reforming ammonia process configurations The ICI Leading Concept Ammonia Process LCA a radical breakaway from the philosophy of the highly integrated large plant which has been so suc cessful for more than 25 years had its industrial debut in 1988 at ICIs own location in Severside Eng land The process consists of a core unit with all the essential process steps Figure 17 and a separate util ity unit which comprises utility boiler and electric gen erator CO2 recovery cooling water system demi water and boiler feed water conditioning and ammo nia refrigeration Feed gas is purified in a hydrodesulfurization oper ating at lower than usual temperatures and passes a saturator to supply a part of the process steam the balance is injected as steam Heated in an inletout let exchanger to 425 C the mixed feed enters the ICI Gas Heated Reformer GHR at 41 bar passing to the secondary reformer at 715 C The shell side entrance temperature of the GHR secondary reformer exit is 970 C falling to 540 C at the exit of the GHR Methane levels exit GHR and secondary reformer are 25 and 067 respectively dry basis Overall steam to carbon ratio is 25 to 27 The gas cooled down to 265 C in the inletoutlet exchanger enters a single stage shift conversion using a special copper zincalumina based catalyst operating in quasiisother mal fashion in a reactor with cooling tubes circulat ing hot water whereby the absorbed heat is used for the feed gas saturation as described above CO2 removal and further purification is effected by a PSA System followed by methanation and drying The syn thesis operates at 82 bar in a proprietary tubular con verter loaded with a cobalt enhanced formula of the classical iron catalyst Purge gas is recycled to the PSA unit and pure CO2 is recovered from the PSA waste gas by an aMDEA wash Very little steam is gener ated in the synloop and from waste gases and some natural gas in an utility boiler in the utility section 60 bar and all drivers are electric The original inten tion was to design a small capacity ammonia plant which can compete with modern large capacity plants in consumption and specific investment and to achieve with lower energy integration a higher flex ibility for start up and reduced load operation need ing a minimal staffing The basic plant features GHR isothermal shift and synthesis can principally be applied for larger capacities too The flow sheet energy consumption is 293 GJt NH3 In the context of the LCA process some discussion on the economics of scale came up Within the same sort of process configuration specific investment will be reduced by increasing capacity at least to a point where limitations for equipment size and transport might play a role and specific investment would then increase again after having reached a minimum In any case for the traditional modern steam reforming ammonia plant a capacity of 2000 td is not beyond the optimum On the other hand it cannot be excluded that concepts as the LCA with no elaborate steam system and a modular and prefabrication construction may come close to the specific investment of world size plants but with regard to the other fixed costs eg staffing some question marks remain Kellogg has combined the ruthenium catalyst based synthesis loop KAAP with its exchanger reformer system KRES to an optimized integrated ammo nia plant concept Ammonia 2000 intended for the use in worldscale singletrain plants in the 1850 td range Desulfurized gas is mixed with steam and then split into two streams in approximate proportion 21 These streams are separately heated in a fired heater The smaller of the two enters the exchanger reformer at 550 550 C while the remainder is passed directly to the autothermal reformer at 600 640 C The exchanger reformer and the autothermal reformer use conventional nickelbased primary and secondary reforming catalysts respectively To satisfy the stoi chiometry and the heat balance the autothermal reformer is fed with enriched air 30 O2 The required heat for the endothermic reaction in the tubes of the exchanger reformer comes from the gases on the shell side comprising a mixture of the efflu ent from the autothermal reformer and the gas emerg ing from the tubes The shell side gas leaves the ves sel with 40 bar The synthesis proceeds at about 90 bar in a 4bed radialflow converter hot wall design with interbed exchangers The first bed is charged with conventional ironbased catalyst for bulk conversion and the other beds with Kelloggs high activity ruthe niumbased catalyst allowing to attain an exit ammo nia concentration in excess of 20 The other pro cess steps are more along the traditional lines The overall energy claimed for this process can be as 1ow as 272 GJt NH3 Another recently launched process is the Linde Ammonia Concept LAC which consists essentially of a hydrogen plant with only a PSA unit to purify the synthesis gas a standard cryogenic nitrogen unit and 21 an ammonia synthesis loop The concept is similar to KTIs PARC process for small capacities The first project with a capacity of 1350 td is presently exe cuted in India The single isothermal shift conversion uses Lindes spiralwound reactor which has been suc cessfully used for methanol plants and hydrogenation in ten plants around the world In the loop a Casale three bed converter with two interbed exchangers is used As in ICIs LCA process pure carbon dioxide can be recovered by scrubbing the off gas from the PSA unit for which Linde also uses the BASF aMDEA process The process consumes about 285 GJt NH3 or with inclusion of pure CO2 recov ery 293 GJt NH3 The status of ammonia plants based on heavy fuel oil and coal For lack of economic incentive not much optimiza tion and development work has been dedicated in the last few years to the field of partial oxidation of higher hydrocarbon fractions The gasification of these plants usually does not consist of a single line Compared to a steam reformer furnace there are more production interuptions because of periodic burner changes and cleaning operations in the gasification units For this reason most installations have a standby unit In addi tion to that the maximum capacity of single gas gene rator corresponds only to 10001100 td of ammonia Therefore world size ammonia plants have 3 4 par tial oxidation generators Generally the degree of energy integration is lower than in the steam reform ing process because in the absence of a large fired fur nace there is no large amount of hot flue gas and con sequently less waste heat is available So in this pro cess route a separate auxiliary boiler is usually necessary to provide steam for mechanical energy and power generation Nevertheless in modern concepts some efforts have been made to bring the energy con sumption down Whereas older plant concepts had values of around 38 GJt NH3 for a concept with the traditional use of 985 oxygen quite recently a fig ure of 335 GJt NH3 was claimed in a commercial bid To reduce investment cost and energy consumption it has been recommended to use air or enriched air 22 Table 2 Comparison of ammonia production 1940 1990 Multiline plant Modern Single Train BASF 1940 1991 Feedstock Coke Natural gas Capacity td 800 1800 Plot size m3 35000 18000 Steel t 30000 13000 Personal 1800 100 Investment 1990 Mio DM 1000 300 Energy consumption GJt NH3 88 28 Table 3 Ammonia production cost from various feedstock in 1996 in NW Europe 1800 td new plant Feedstock Natural gas Vacuum residue Coal Process Steam Partial Partial Reforming Oxidation Oxidation Feedstock price DMGJ 43 30 27 Total energy consumption GJt NH3 285 38 485 Feedstock energy costs DMt NH3 123 114 131 Other cash costs DMt NH 50 65 100 Total cash costs DMt NH3 173 179 231 Capitalrelated costs DMt NH3 100 143 260 Total cost DMt NH3 273 322 491 Investment Mio DM 350 500 900 For capitalrelated costs a debtequity ratio of 60 40 is assumed With 6 depreciation 8 interest on debts and 16 ROI on equity total capitalrelated charges are 172 on investment instead of pure oxygen Topsøe proposed the use of enriched air 43 and methanation instead of liq uid nitrogen wash For CO2H2S removal Selexol is applied The shift reaction proceeds over a sulfur resistant catalyst in a threebed configuration bring ing the residual carbon monoxide content down to 055 For the loop a Series 200 converter is chosen The partial oxidation step can be designed according to either the Texaco or the Shell process An overall consumption of 348 GJmt NH3 is stated Foster Wheeler suggests the use of highly preheated air in a Texaco generator operating at 70 bar The gas purification train comprises soot scrubbing followed by shift conversion acid gas removal and methana tion The gas is dried by molecular sieves and finally fed to a cryogenic unit to remove the surplus nitro gen and residual methane argon and carbon monox ide traces The rejected nitrogen is expanded in a tur bine which helps to drive the air compressor A spe cial design consideration was the following Conventional air separation uses fractional distilla tion of oxygen and nitrogen at a difference in boiling points of only 13 C In the cryogenic unit of the Fos terWheeler process a lesser quantity of nitrogen because the stoichoimetrically needed proportion remains in the gas is separated from hydrogen at a much higher boiling point differential 57 C This should save capital investment and energy consump tion against the traditional approach A figure of 356 376 GJmt NH3 is given for heavy oil feedstock For coal based plants the economic incentive for extensive RD is even lower than with fuel oil The major part of coal fed ammonia plants most of them of rather small size are located in China and use still the water gas route A few ammonia plants based on more modern coal gasification processes as the Tex aco Process the KoppersTotzek Process and the Lurgi Coal Gasification are of larger size and oper ate in South Africa India and Japan Also in coal based ammonia plants the gas generation consists of several lines Depending on the gasification process the maximum capacity of a single gasifier corresponds to an ammonia production between 500 td Koppers Totzek Texaco and 800 td Lurgi gasifier Regard ing the degree of energy integration the situation is at best as in the partial oxidation of fuel oils but in any case much lower than in a steam reforming plant Lurgis moving bed gasification produces a gas with a rather high content of methane which after separ ation in the cryogenic step is processed in a small steam reforming unit Shift conversion Rectisol unit liquid nitrogen wash are the other essential steps in the synthesis gas preparation The gasification needs 32 34 GJt NH3 power and steam generation con sumes 18 22 GJt NH3 resulting in a total energy consumption of 5056 GJt NH3 For the KoppersTot zek route a figure of 515 GJt is reported Ube Indus tries commissioned a 1000 td ammonia plant in 1984 using Texacos coal gasification process An energy consumption of 455 GJt NH3 is stated which is lower than the normally quoted figure of 485 GJt NH3 for this technology Economics of ammonia production The enormous technical and economical progress made from the old plants using coke and water gas technology to the modern steam reforming ammonia plant with natural gas feedstock may be seen from the table 2 Table 3 gives an estimate for ammonia pro duction in 1996 cost in northwest Europe for differ ent feedstocks using todays best and proven techno logical standards for each process From table 3 it is obvious that at present there is no chance for the other feedstocks to compete against steam reforming of natural gas Only under very spe cial circumstances in cooperation with a refinery for example partial oxidation of heavy oil fractions might be economically justified It should be noted however that the average energy consumption of the steam reforming plants presently in operation is noticeable higher than the example of the modern lowenergy concept used in table 3 The combined cost of feedstock and energy for a steam reforming plant both are natural gas is the principal determinant of the overall production costs The price of gas and by extension the price of ammonia is to a greater or lesser extent linked to the price of crude oil The present interfuel relation ship between gas and oil pricing might be distorted by the need for cleaner and less polluting fuels result ing from increasing environmental awareness In this respect natural gas is so advantageous in relation to other fossil fuels that demand could well be pushed up in the coming years 23 The received opinion is that although gas supply can be increased upward pressure on prices is necessary to make that happen The forecast is for higher pro duction costs in Western Europe and the USA In the low gas cost areas such as the Arab GuIf Trinidad and Indonesia competing usages for natural gas are not expected to grow to any great extent and in such loca tions feedstock costs for the ammonia producers are expected to rise only moderately The biggest share of the proven world reserves of natural gas have for mer USSR followed by the Middle East In the very long term coal has prospects which might be drawn from consumption and world reserves of fos sil feedstocks At present consumption rates coal will cover the demand for 235 years natural gas for 66 years and oil for 43 years But at least for the medium term natural gas can continue as preferred feedstock Future perspectives for the ammonia production technology Albeit world population and thus the demand of fer tilizers is increasing 87 of the ammonia produc tion is consumed in this sector building of new ammonia plants did not keep up adequately with this Of course the main increase of demand is in the developing countries but in most cases there is not sufficient capital available for the investments needed In the industrial countries with sufficient food sup ply the fertilizer consumption is at best stagnant for ecological concerns presently exercising the collective minds in the Western World The sometimes from environmentalists propagated ecological agriculture is no alternative as manure and biomass are not suf ficient in effect and quantity to supply the necessary nitrogen and in addition they have the same problem with nitrate runoff Direct biological fixation is presently restricted to the legumes by their symbiotic relationship with the Rhi zobium bacteria which settle in the root nodules of the plants Intensive genetic engineering research has provided so far a lot of insights in the mechanism of this biological fixation but a real breakthrough for practical agricultural application has not yet hap pened The enzyme nitrogenase practically performs an ammonia synthesis in the bacteria and for the syn thesis of the nitrogenase the socalled NIF gene is responsible One option for example would be to broaden the host spectrum of the Rhizobium bacte ria by genetic manipulation Other possibilities are to transfer the NIF gene to other bacteria which have a broader host spectrum but have no own nitrogen fix ation ability or to insert the NIF gene directly into plants One important point to consider especially with the option of constructing a nitrogen fixing plant is the energy balance of the plant Because of the low efficiency a considerable amount of the photosynthe sis product would be consumed to supply the energy needed This would consequently lead to a reduction in yield which is estimated by some researchers to be as high as 18 The possibility of converting atmospheric nitrogen into ammonia in homogeneous solution using metal organic complexes was first raised around 1966 The prospects for this route are not judged to be very promising in terms of energy consumption and also with respect to the cost of these very sophisticated cat 24 alyst systems Photochemical methods of producing ammonia at ambient temperature and atmospheric pressure in the presence of a catalyst have been reported but the yields are far too low to be econom ically attractive So for the foreseeable future we have to rely on the conventional ammonia synthesis reaction combining hydrogen and nitrogen using a catalyst at elevated temperature and pressure in a recycle process as con ceived in laboratory by Fritz Haber and made oper able on an industrialscale by C Bosch and on the use of the known routes to produce the hydrogen and nitrogen needed which are in fact what consumes all the energy In conclusion it is possible to sum up the prospects by the following broad predictions Natural gas will remain the preferred feedstock for at least the next 15 years Coal gasification will not play a major role in ammonia production in that period The present ammonia technology will not change fundamentally at least in the next 15 years Even if there are radical unforeseeable developments they will take time to develop to commercial intro duction With the available concepts the margins of additional improvements have become rather small after years of intensive research and devel opment Thus only minor improvements of individ ual steps catalysts and equipment might be expected A further significant reduction in the energy con sumption of the natural gasbased steam reform ing ammonia process is unlikely figures between 27 and 28 GJt NH3 are already close to theoreti cal minimum In the medium term the bulk of ammonia produc tion will still be produced in worldscale plants of 1000 2000 td NH3 Small capacity plants will be limited to locations where special logistical finan cial or feedstock conditions favor them New developments in ammonia technology will mainly reduce investment costs and increase oper ational reliability Smaller integrated process units eg exchanger reformer CAR contribute to this reduction and give additional savings by simplify ing piping and instrumentation Improved reliabil ity may result from advances in catalyst and equip ment quality and from improved instrumentation and computer control It is very likely that genetic engineering will suc ceed in modifying some classical crops for biolog ical nitrogen fixation and that application in large scale will occur predominantly in areas with still strongly growing population to secure the increas ing food demand This development may be pushed by the fact that compared to the classical fertilizer route less capital and less energy would be needed This may happen within the next 20 years but time estimates are always risky A famous example Man will not fly for 50 years Wilbur Wright 1901 But even with the introduction of this new approach traditional ammonia synthesis will con tinue to operate in parallel because it might be nec essary to supplement the biological nitrogen fixa tion with classical fertilizers In addition the exist ing ammonia plants represent a considerable capital investment and a great number of them may reli ably operate for at least another 20 30 years from a mere technical point of view References AVSlack GRussel James Ammonia Part IIII Marcel Decker New York 1073 1974 1977 JR Jennings Ed Catalytic Ammonia Synthesis Plenum Press New York London 1991 ANielsen Ed Ammonia Catalysis and Manufacture Sprin gerVerlag Berlin Heidelberg New York 1995 SATopham The History of the Catalytic Synthesis of Ammo nia in Catalysis ed RAnderson and MBoudart Volume 7 p 1 50 SpringerVerlag Berlin Heidelberg New York 1980 MAppl Ammonia Methanol Hydrogen Carbon Monoxide Modern Production Technologies Ed Alexander More British Sulphur Publishing 1977 ISBN 1873 38726 MAppl Ammonia Synthesis and the Development of Cataly tic and High Pressure Processes DrH L Roy Memorial Lec ture IIChE Conference Hyderabad Dec 1986 Indian Chemi cal Engineer XXIX 1 229 1987 MAppl The HaberBosch Process and the Development of Chemical Engineering in A Century of Chemical Engineering ed WFurter Plenum Publishing Corporation 2051 1982 MAppl A Brief History of Ammonia Production from the Early Days to the Present Nitrogen 100 4758 MarApr 1976 LConnock Ammonia and Methanol from Coal Nitrogen 226 4756 MarApr 1997 25