haber ammonia technology

8/14/2019 HABER Ammonia Technology

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increase in total volume, significant savings of com-pression energy are possible if the process is performedunder 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 canbe compensated by higher temperatures. As all theheat in the primary reformer has to be transferredthrough the tube wall, the wall temperatures will riseand approach the material limits. Originally HK 40tubes with a content of 20% nickel and 25% chro-mium were commonly used. With new grades as HPmodified with higher nickel content andstabilized with niobium and the recentlyintroduced Micro Alloys which addition-ally contain titanium and zirkoniumhigher wall temperatures and thushigher pressures up to 44 bar in the pri-mary reformer have become possible.The steam surplus applied in thereformer could thus also be reduced froma steam to carbon ratio of 4 and higher toabout 3 or slightly below, and this wasassisted by improved catalysts withenhanced activity and better heat trans-fer characteristic. For naphtha reforming

a higher steam surplus is necessary.

Fancy catalyst shapes as wagon wheels, six-shooters,shamrock or four-hole have replaced the old Raschigrings. The stability of the standard catalyst supportsas calcium aluminate, magesium aluminate and -alumina has been improved and it has become awidely 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 whichwas developed by ICI originally for naphtha steamreforming in order to prevent carbon deposition bycracking reactions. From the various primary reformerdesigns the top fired concept with a single radiationbox dominates in the larger plants, the side-fireddesign 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 commonflue gas duct. The secondary reformers have beenoptimized regarding hydrodynamics and burnerdesign using computational fluid dynamics. Figure 11shows an example of a top-fired reforming furnacetogether with the secondary reformer.

The reduction of the steam-to-carbon ratio was abigger problem for the HT shift than in the reformerstep, as the gas mixture became a higher oxidativepotential and tended to over-reduce the iron-oxidefrom magnetite to FeO and in extreme cases partiallyto 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 Fischer-Tropsch reaction leads to theformation of methane and higher hydrocarbons. Cop-per promotions of the iron catalyst suppresses theseside reaction. The nasty problem of methanol andamine formation in the LT shift is largely solved by

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Figure 11: Top-fired primary reformer and secondary reform er ( U hde design)

F igur e 12: C O 2 L oadin g characteri stics of vari ous solvents


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improved formulations of the copper/zinc/alumina,and a new development is the intermediate temper-ature shift catalyst, operated quasi isothermal in atubular reactor, for example in the ICI LCA ammo-nia process or the Linde ammonia process (LAC).

Large progress in the CO 2 removal systems was made

in the last decade. The original MEA systems had aheat consumption for solvent regeneration over 200kJ/kmol, a corrosion inhibitor system called amineguard III brought it down to about 120 kJ/kmol, butthis is still nearly 5 times as high as the most advancedsystem, the BASF aMDEA Process, which uses anaqueous solution of monomethyl-diethanolaminetogether with a special promotor which enhances themass transfer. Other low energy systems are theBenfield LoHeat Process, which is a hot potash systemor 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 theold plants, the solubility of the CO 2 is according toHenrys law direct proportional to the CO 2 partialpressure and regeneration can be achieved by flash-ing, without application of heat.

In contrast to this the MEA is a chemical solvent, thesolubility is only slightly dependent on the CO 2partial pressure and approaches a saturation value.

MEA forms a stable salt with the carbon dioxide anda high amount of heat is required in the stripper todecompose it. BASFs aMDEA Process is about inbetween, the characteristic can be adjusted in a flex-ible way by the concentration of the activator, so thatthe major part of the dissolved carbon dioxide can bereleased by simple flashing and only a smaller propor-tion has to be stripped out by heat. Figure 12 showsCO 2 loading characteristics of various solvents.

The tubular steam reformer has become a very reliableapparatus 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 ammoniaplant. So in some modern concepts the size was reducedby shifting some of the load to the secondary reformernecessitating an overstoichiometric amount of processair. The surplus of nitrogen introduced in this way, canbe removed downstream by the use of a cryogenic unit.

C.F. Braun was the first contractor which introducedthis concept in the so-called Purifier Process. Some con-tractors have gone so far to by-pass some of the natu-

ral gas around the tubular reformer and feeding it

directly to the secondary reformer which likewise needssurplus of process air or oxygen enriched air.

But there are additional reasons for breaking away fur-ther from the fired furnace concept. The temperaturelevel 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 andsuperheat high pressure steam. The boiling point of HPsteam is only 325 C and the first heat exchanger in theflue gas duct preheats process air in the conservativeplants to only 500 C (600 700 C in more moderninstallations). Recycling high-level heat from the sec-ondary reformer and making use of it for the primaryreforming reaction is thermodynamically the betteroption. Concepts which use this heat in an exchangerreformer have been successfully developed and com-mercially demonstrated. The first to come out with thisconcept in a real production plant was ICI with its GHR

(Gas Heated Reformer). The hot process gas from thesecondary reformer is the sole heat source. A surplusof process air of around 50% is needed in the secon-

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Figure 13: ICI Gas-Heated Reformer


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dary reformer to achieve a closed heat balance. Figure13 is a simplfied drawing of the ICI Gas-HeatedReformer.

Quite recently ICI has come out with a modifieddesign, the AGHR, with the A standing foradvanced. The bayonet tubes are replaced by normal

tubes attached to a bottom tube sheet using a specialpacking 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. Thetubes are open at the lower end and the reformed gasmixes with the hotter effluent of the secondaryreformer. The mixed gas stream flows up-ward on theshell side to heat the reformer tubes. Thus primaryreforming and secondary reforming reaction proceedin parallel in contrast to the ICI concept where thetwo reactions proceed in series. The Kellogg processuses enriched air. The complete elimination of thefired tubular furnace leads to a drastic reduction of NO x emission, because there is only flue gas frommuch smaller fired heaters required for feed andprocess air preheat. An even more progressiveexchanger reformer presently operating in a demo-plant is Uhdes CAR (Combined autothermalreformer) which not only replaces the catalytic sec-

ondary reforming step by a non catalytic partial oxy-

dation step but also combines this with the exchangerreformer in one single vessel.

Syngas from heavy oil fractions viapartial 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 non-cat-alytic and proceeds in an empty vessel lined with alu-mina refractory. The reactants, oil and oxygen, alongwith a minor amount of steam, are introduced througha nozzle at the top of the generator vessel. The noz-zle consists of concentric pipes so that the reactantsare fed separately and react only after mixing at theburner tip in the space below. The temperature in thegenerator is between 1200 and 1400 C. Owing to theinsufficient mixing with oxygen, about 2% of the totalhydrocarbon feed is transformed into soot, which isremoved by water scrubbing. The separation of thesoot from the water and its further treatment differsin 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, thehydrogen sulfide formed during gasification is

removed along with carbon dioxide by scrubbing withchilled methanol below 30 C in the Rectisol pro-

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Figure 14: A mmon ia syngas by partial oxidation of heavy hydrocarbons (Texaco)


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cess. Then, as in the steam reforming route, the gasundergoes the CO shift reaction. Because of thehigher carbon monoxide content much more reactionheat is produced, which makes it necessary to distrib-ute the catalyst on several beds with intermediatecooling. The carbon dioxide formed in the shift con-version is removed in a second stage of the Rectisol

unit; both have a common methanol regenerationsystem. The H 2S-rich carbon dioxide fraction from thefirst 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, whichabsorbs the residual carbon monoxide, methane anda portion of the argon (which was introduced into theprocess in the oxygen feed). The conditions in thisstage are set so that the stoichiometric nitrogenrequirement is allowed to evaporate into the gasstream from the liquid nitrogen wash. The processneeds, of course, an air separation plant to produceoxygen, usually around 98.5% pure, and to supply theliquid nitrogen. Figure 14 is a simplified flowsheet of synthesis gas preparation by partial oxidation of heavyfuel oil using the Texaco Syngas Generation Process.The Shell process uses of a waste heat boiler for rawgas cooling whereas Texaco prefers for ammoniaplants a water quench for this purpose which has theadvantage that this intro-duces the steam for the

subsequent shift conver-sion which differentfrom Shell is performedwithout prior removal of the sulfur compoundsusing a sulfur tolerantshift catalyst.

Besides some optimiza-tions there are no funda-mental new develop-ments in the individualprocess steps. Some pro-posed changes in the pro-cess sequence, for exam-ple methanation insteadof liquid nitrogen wash,or the use of air insteadof pure oxygen are notrealized so far. Thoughother CO 2 removal

systems as Selexol orPurisol (N-Methylpyrrol-idon ) and alternative

sulfur recovery processes are suitable too, Rectisoland Claus Process remain the preferred options.

Synthesis gas by coal gasificationThere is no chance for a wide-spread use of coal as feed-stock for ammonia in the near future, but a few remarksshould be made regarding the present status of coal gas-ification technology. Proven gasification processes arethe Texaco Process, the Koppers-Totzek Process, andthe Lurgi Coal Gasification. The Shell gasification, notyet in use for ammonia production , but successfullyapplied for other productions is an option , too. Texacosconcept is very similar to its partial oxydation processfor heavy fuel oil feeding a 70% coal-water paste intothe generator. Koppers-Totzek 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 quantitiesof steam as the others. Shells process differs consid-erably from its oil gasification process in flow patternand feeds coal dust. Texaco, Lurgi, and Shell operateunder pressure, whereas the Koppers-Totzek gasifieris under atmospheric pressure, but a pressure version,called PRENFLOW is presently tested in a demo-plant. Continuous slag removal either in solid or mol-

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Figure 15: A mmonia plant temperature profi le


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ten form is, indeed, the fundamental technical problemwith coal-based systems and the technical solutions dif-fer considerably. Gas cooling is achieved by quench andor waste heat boiler, entrained coal dust is removed bywater scrubbing. The following process steps for shiftconversion, CO 2 removal and final purification arelargely the same as in partial oxdiation of heavy fuel

oil.

Energy integration andammonia plant concept

The integrated steam reformingammonia plantIn the old days an ammonia plant was more or less justa combination with respect to mass flow and energymanagement was handled within the separate processsections, which were often sited separately, as theyusually consisted of several parallel units. A revolu-tionary break-through came in the mid of the 1960swith the steam reforming ammonia plants. The newimpulses came more from the engineering and con-tractor companies than from the ammonia plantindustry itself. Engineering contractors have beenworking 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 single-train unitsof considerable size.

By applying the experience gained in this field it waspossible to create within a few years in the mid 1960sthe modern large-scale ammonia concept. To use asingle-train for large capacities (no parallel lines) andto be as far as possible energetically self-sufficient (noenergy import) through a high degree of energy inte-gration (with process steps with surplus supplyingthose with deficit) was the design philosophy for thenew steam reforming ammonia plants pioneered byM. W. Kellogg and some others. It certainly had alsoa revolutionary effect on the economics of ammoniaproduction, making possible an immense growth inworld capacity in the subsequent years. The basic

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Table 1: Main energy sources and sinks in the steam reforming ammonia Process

Process section Originating Contribution

Reforming Primary reforming duty DemandFlue gas SurplusProcess gas Surplus

Shift conversion Heat of reaction SurplusCO 2 removal Heat for solvent regeneration DemandMethanation Heat of reaction SurplusSynthesis Heat of reaction SurplusMachinery Drivers DemandUnavoidable loss Stack and general DemandBalance (Auxiliary boiler or import) Deficit

(Export) Surplus


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reaction sequence has not changed since then. Figure15 shows the process sections and the relevant gastemperature levels in a steam reforming ammoniaplant.

High-level surplus energy is available from the fluegas and the process gas streams of various sections,

while there is a need for heat in other places such asthe process steam for the reforming reaction and inthe solvent regenerator of the carbon dioxide removalunit (Table 1). Because a considerable amount of mechanical energy is needed to drive compressors,pumps and fans, it seemed most appropriate to usesteam turbine drives, since plenty of steam could begenerated from waste heat. As the temperature levelwas high enough to raise HP steam of 100 bar, it waspossible to use the process steam first to generatemechanical energy in a turbine to drive the synthe-sis gas compressor before extracting it at the pressurelevel of the primary reforming section.

The earlier plants were in deficit, and they needed anauxiliary boiler, which was integrated in the flue gasduct. This situation was partially caused by inadequatewaste heat recovery and low efficiency in some of theenergy consumers. Typically, the furnace flue gas wasdischarged up the stack at unnecessarily high temper-atures because there was no combustion air pre-heat

and too much heat was rejected from the synthesisloop, while the efficiency of the mechanical drivers

was low and the heat demand in the carbon dioxideremoval unit regenerator was high.

A very important feature of this new concept was theuse 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 dutiesa single line approach. The lower energetic efficiencycompared to the reciprocating compressors of whichin the past several had to be used in parallel is morethan compensated by the lower investment and theeasy 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 notonly for the major machines such as synthesis gas, airand refrigeration compressors but even for relativelysmall pumps and fans. The outcome was a rather com-plex steam system and one may be tempted todescribe an ammonia plant as a sophisticated powerstation making ammonia as a by-product. The plantsproduce more steam than ammonia, even today, themost modern plants still produce about three timesas much. In recent years electrical drives have swungback into favor for the smaller machines.

In most modern plants total energy demand(feed/fuel/power) has been drastically reduced. On

the demand side important savings have beenachieved in the carbon dioxide removal section byswitching from old, heat-thirsty processes like MEA

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Figur e 16: Sim pli fi ed fl ow sheet of a mo dern steamreform ing ammon ia plant (C .F. Braun Pur if ier Process)


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scrubbing to low-energy processes like the newer ver-sions of the Benfield process or aMDEA. Fuel issaved by air preheat and feed by hydrogen recoveryfrom the purge gas of the synloop by cryogenic, mem-brane or pressure swing adsorption technology. In thesynthesis loop the mechanical energy needed for feedcompression, refrigeration and recycle has been

reduced, and throughout the process catalyst volumesand geometry have been optimized for maximumactivity and minimum pressure drop.

On the supply side, available energy has beenincreased by greater heat recovery, and the combinedeffect of that and the savings on the demand side havepushed the energy balance into surplus. Because thereis no longer an auxiliary boiler, there is nothing in theplant 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 exportingsteam 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 sitefor export steam, it can be very economic (depend-ing on the price of natural gas and the value assignedto 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 feedstockrequirement is stoichiometric. So the only way is to cutthe firing in the reforming furnace by shifting reform-ing duty to the secondary reformer, as we had alreadydiscussed earlier or to choose a more radical aproach

by the use of an exchanger reformer instead of thefired furnace: ICIs Gas-Heated Reformer (GHR)system, the KRES of M. W. Kellogg and the TandemReformer (now marketed by Brown & Root), or theeven more advanced Combined AutothermalReformer (CAR) of Uhde. But none of these designsnecessarily achieves any significant improvement overthe net energy consumption of the most advanced con-ventional concepts under the best conditions.

For the cases in which export of steam and/or poweris welcome there is the very elegant possibility of inte-grating a gas turbine into the process to drive the aircompressor. The hot exhaust of 500 550 C containswell enough oxygen to serve as preheated combustionair for firing the primary reformer. The gas turbinedoes not even have to be particularly efficient,because any heat left in the exhaust gas down to theflue gas temperature level of 150 C is used in the fur-nace. Thus an overall efficiency of about 90% can beachieved.

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Boiler makers provide today largely reliable designsfor high-duty waste heat boilers after secondaryreformer and in the synthesis loop, in which up to 1.5t steam/t NH 3 are produced, corresponding roughlyto a recovery of 90% of the reaction enthalpy of thesynthesis. Centrifugal compressors have become muchmore reliable, though their efficiency has not

increased spectacularly in recent years. Someimprovements were made in turn-down capability inimproving the surge characteristic. New developmentsare dry seals instead of oil seals and another poten-tial improvement, already successfully introduced innon-ammonia applications, is the magnetic bearing.

Although the introduction of the single-train inte-grated large plant concept in the 1960s revolutionizedthe energy-economics of ammonia production, it issurprising that since then the total consumption hasbeen reduced by about 30%, from roughly 40 to 28GJ/t. An example of a modern plant shows Figure 16.

From this enormous reduction in energy consumptionthe question may come up, what is the theoretical min-imum energy consumption for ammonia productionvia steam reforming of natural gas. Based on puremethane, we may formulate the following stoichio-metric equation:

CH 4 + 0.3035 O 2 + 1.131 N 2 + 1. 393 H 2O CO 2 + 2.262 NH 3

(10)

H 0298 = 86 kJ/mol; F0298 = 101 kJ/mol

So from a mere thermodynamic point of view, in anideal engine or fuel cell heat and power should beobtained from this reaction. But because there is ahigh degree of irreversibility in the real process a con-siderable amount of energy is necessary to producethe ammonia from methane, air and water. The stoi-chiometric quantity of methane derived from the fore-

going equation is 583 Nm3

per mt NH 3, which corre-sponds to 20.9 GJ (LHV) per tonne of ammonia, whichwith some reason could be taken as minimum value.Of course, if one assumes full recovery of the reactionheat, then the minimum would be the heating value of ammonia, which is 18.6 GJ (LHV) per mt NH 3.

Energy and exergy anal-ysis (First and SecondLaw of Thermodynamicsrespectively) identify theprocess steps in whichthe biggest losses occur.The biggest energy loss isin the turbines and com-pressors, whereas theexergy loss is greatest inthe reforming section,almost 70%. Based onexergy the thermody-namic efficiency for the

ammonia productionbased on steam reform-ing of natural gas isalmost 70%.

It has become rather common to measure modernammonia concepts above all by their energy consump-tion. Yet these comparisons need some caution ininterpretation; without a precise knowledge of designbases, physical state of the produced ammonia andstate of the utilities used, e.g. cooling water temper-ature, nitrogen content in natural gas, or conversionfactors used for evaluating imported or exportedsteam and power, misleading conclusions may bedrawn. In many cases, too, the degree of accuracy of such figures is overestimated.

The best energy consumption values for ammoniaplants using steam reforming of natural gas are around28 GJ/tNH 3. Industrial figures reported for plants withhigh-duty primary reforming and stoichometric pro-

cess air and for those with reduced primary reform-ing and excess air show practical no difference.

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Figure 17: Flow d iagram of I CI s L CA A mmon ia Process (Cor e unit) f or 450 mtpd


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New steam reforming ammoniaprocess configurationsThe ICI Leading Concept Ammonia Process LCA a radical break-away from the philosophy of thehighly 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 theessential process steps (Figure 17) and a separate util-ity unit which comprises utility boiler and electric gen-erator, CO 2 recovery, cooling water system, demiwater and boiler feed water conditioning and ammo-nia refrigeration.Feed gas is purified in a hydro-desulfurization oper-ating at lower than usual temperatures and passes asaturator to supply a part of the process steam, thebalance is injected as steam. Heated in an inlet/out-let exchanger to 425 C the mixed feed enters the ICIGas Heated Reformer (GHR) at 41 bar, passing to thesecondary reformer at 715 C. The shell side entrancetemperature 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 are25% and 0.67% respectively (dry basis). Overallsteam to carbon ratio is 2.5 to 2.7. The gas, cooleddown to 265 C in the inlet/outlet exchanger, entersa single stage shift conversion, using a special copper-

zinc-alumina based catalyst operating in quasi-isother-mal fashion in a reactor with cooling tubes, circulat-ing hot water, whereby the absorbed heat is used forthe feed gas saturation as described above. CO 2removal and further purification is effected by a PSASystem, 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 theclassical iron catalyst. Purge gas is recycled to the PSAunit and pure CO 2 is recovered from the PSA wastegas by an aMDEA wash. Very little steam is gener-ated in the synloop, and from waste gases and somenatural gas in an utility boiler in the utility section (60bar) and all drivers are electric. The original inten-tion was to design a small capacity ammonia plantwhich can compete with modern large capacity plantsin consumption and specific investment, and toachieve 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 sheetenergy consumption is 29.3 GJ/t NH 3.

In the context of the LCA process some discussion onthe economics of scale came up. Within the same sortof process configuration specific investment will bereduced by increasing capacity, at least to a pointwhere limitations for equipment size and transportmight play a role and specific investment would thenincrease again after having reached a minimum. In

any case for the traditional modern steam reformingammonia plant, a capacity of 2000 t/d is not beyondthe optimum. On the other hand it cannot be excludedthat concepts as the LCA with no elaborate steamsystem and a modular and prefabrication constructionmay come close to the specific investment of worldsize plants, but with regard to the other fixed costs,e.g. staffing, some question marks remain.

Kellogg has combined the ruthenium catalyst basedsynthesis loop (KAAP) with its exchanger reformersystem ( KRES) to an optimized integrated ammo-nia plant concept (Ammonia 2000) intended for theuse in world-scale single-train plants in the 1850 t/drange. Desulfurized gas is mixed with steam and thensplit into two streams in approximate proportion 2 : 1.These streams are separately heated in a fired heater.The smaller of the two enters the exchanger reformerat 550 550 C, while the remainder is passed directlyto the autothermal reformer at 600 640 C. Theexchanger reformer and the autothermal reformer use

conventional nickel-based primary and secondaryreforming catalysts respectively. To satisfy the stoi-chiometry and the heat balance, the autothermalreformer is fed with enriched air (30% O 2). Therequired heat for the endothermic reaction in thetubes of the exchanger reformer comes from the gaseson 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 barin a 4-bed radial-flow converter (hot wall design) withinter-bed exchangers. The first bed is charged withconventional iron-based catalyst for bulk conversionand the other beds with Kelloggs high activity ruthe-nium-based catalyst, allowing to attain an exit ammo-nia concentration in excess of 20%. The other pro-cess steps are more along the traditional lines. Theoverall energy claimed for this process can be as 1owas 27.2 GJ/t NH 3.

Another recently launched process is the Linde

Ammonia Concept (LAC) which consists essentiallyof a hydrogen plant with only a PSA unit to purify thesynthesis gas, a standard cryogenic nitrogen unit and

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an ammonia synthesis loop. The concept is similar toKTIs PARC process for small capacities. The firstproject with a capacity of 1350 t/d is presently exe-cuted in India. The single isothermal shift conversionuses Lindes spiral-wound reactor, which has been suc-cessfully used for methanol plants and hydrogenationin ten plants around the world. In the loop a Casale

three bed converter with two interbed exchangers isused. As in ICIs LCA process, pure carbon dioxidecan be recovered by scrubbing the off gas from thePSA unit, for which Linde also uses the BASFaMDEA process. The process consumes about28.5 GJ/t NH 3, or, with inclusion of pure CO 2 recov-ery 29.3 GJ/t NH 3.

The status of ammonia plants basedon heavy fuel oil and coal

For lack of economic incentive, not much optimiza-tion and development work has been dedicated in thelast few years to the field of partial oxidation of higherhydrocarbon fractions. The gasification of these plantsusually does not consist of a single line. Compared to

a steam reformer furnace there are more productioninteruptions because of periodic burner changes andcleaning operations in the gasification units. For thisreason most installations have a standby unit. In addi-tion to that the maximum capacity of single gas gene-rator corresponds only to 1000-1100 t/d 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 usuallynecessary to provide steam for mechanical energy andpower generation. Nevertheless, in modern conceptssome efforts have been made to bring the energy con-sumption down. Whereas older plant concepts hadvalues of around 38 GJ/t NH 3, for a concept with thetraditional use of 98.5 %+ oxygen quite recently a fig-ure of 33.5 GJ/t NH 3 was claimed in a commercial bid.

To reduce investment cost and energy consumptionit has been recommended to use air or enriched air

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Table 2: Comparison of ammonia production 1940 1990

Multi-line plant Modern Single TrainBASF 1940 1991

Feedstock Coke Natural gasCapacity t/d 800 1800Plot size m 3 35 000 18 000Steel t 30 000 13 000Personal 1800 100Investment (1990) Mio DM 1000 300Energy consumption GJ/t NH 3 88 28

Table 3: Ammonia production cost from various feedstock in 1996 in NW Europe (1800 t/d, new plant)

Feedstock Natural gas Vacuum residue CoalProcess Steam Partial Partial

Reforming Oxidation Oxidation

Feedstock price DM/GJ 4.3 3.0 2.7Total energy consumption GJ/t NH 3 28.5 38 48.5Feedstock & energy costs DM/t NH 3 123 114 131Other cash costs DM/t NH 50 65 100Total cash costs DM/t NH 3 173 179 231Capital-related costs DM/t NH 3 100 143 260Total cost DM/t NH 3 273 322 491

Investment Mio DM 350 500 900For capi tal-related costs a debt/equi ty rati o of 60 : 40 is assumed. Wi th 6 % depreciation, 8 % interest on d ebts and 16% RO I on equity, total capital-related charges are 17.2% on i nvestment.


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instead of pure oxygen. Topse proposed the use of enriched air (43%) and methanation instead of liq-uid nitrogen wash. For CO 2/H 2S removal Selexol isapplied. The shift reaction proceeds over a sulfur-resistant catalyst in a three-bed configuration, bring-ing the residual carbon monoxide content down to0.55%. For the loop a Series 200 converter is chosen.

The partial oxidation step can be designed accordingto either the Texaco or the Shell process. An overallconsumption of 34.8 GJ/mt NH 3 is stated.

Foster Wheeler suggests the use of highly preheatedair in a Texaco generator operating at 70 bar. The gaspurification train comprises soot scrubbing followedby shift conversion, acid gas removal and methana-tion. The gas is dried by molecular sieves and finallyfed 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 boilingpoints of only 13 C. In the cryogenic unit of the Fos-ter-Wheeler process a lesser quantity of nitrogen(because the stoichoimetrically needed proportionremains in the gas) is separated from hydrogen at amuch higher boiling point differential (57 C). This

should save capital investment and energy consump-tion against the traditional approach. A figure of 35.6 37.6 GJ/mt NH 3 is given for heavy oil feedstock.

For coal based plants the economic incentive forextensive R&D is even lower than with fuel oil. Themajor part of coal fed ammonia plants -most of themof rather small size are located in China and use stillthe water gas route. A few ammonia plants based onmore modern coal gasification processes as the Tex-aco Process, the Koppers-Totzek Process, and theLurgi Coal Gasification are of larger size and oper-ate in South Africa, India and Japan. Also in coalbased ammonia plants the gas generation consists of several lines. Depending on the gasification processthe maximum capacity of a single gasifier correspondsto an ammonia production between 500 t/d (Koppers-Totzek, Texaco) and 800 t/d (Lurgi gasifier). Regard-ing the degree of energy integration the situation isat best as in the partial oxidation of fuel oils, but inany case much lower than in a steam reforming plant.

Lurgis moving bed gasification produces a gas witha 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 inthe synthesis gas preparation. The gasification needs32 34 GJ/t NH 3, power and steam generation con-sumes 18 22 GJ/t NH 3, resulting in a total energyconsumption of 50-56 GJ/t NH 3. For the Koppers-Tot-zek route a figure of 51.5 GJ/t is reported. Ube Indus-

tries commissioned a 1000 t/d ammonia plant in 1984using Texacos coal gasification process. An energyconsumption of 45.5 GJ/t NH 3 is stated, which is lowerthan the normally quoted figure of 48.5 GJ/t NH 3 forthis technology.

Economics of ammoniaproduction

The enormous technical and economical progressmade from the old plants using coke and water gastechnology to the modern steam reforming ammoniaplant with natural gas feedstock may be seen from thetable 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 againststeam reforming of natural gas. Only under very spe-cial circumstances in cooperation with a refinery, forexample partial oxidation of heavy oil fractionsmight be economically justified. It should be noted,however, that the average energy consumption of thesteam reforming plants presently in operation isnoticeable higher than the example of the modernlow-energy concept used in table 3.

The combined cost of feedstock and energy for asteam reforming plant both are natural gas is theprincipal 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 tothe price of crude oil. The present interfuel relation-ship between gas and oil pricing might be distortedby the need for cleaner and less polluting fuels result-ing from increasing environmental awareness. In thisrespect, natural gas is so advantageous in relation toother fossil fuels that demand could well be pushed

up in the coming years.

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The received opinion is that, although gas supply canbe increased, upward pressure on prices is necessaryto make that happen. The forecast is for higher pro-duction costs in Western Europe and the USA. In thelow gas cost areas such as the Arab GuIf, Trinidad andIndonesia, competing usages for natural gas are notexpected to grow to any great extent, and in such loca-

tions feedstock costs for the ammonia producers areexpected to rise only moderately. The biggest shareof 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 mightbe drawn from consumption and world reserves of fos-sil feedstocks. At present consumption rates coal willcover the demand for 235 years, natural gas for 66years and oil for 43 years. But at least for the mediumterm natural gas can continue as preferred feedstock.

Future perspectives for theammonia 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 thedeveloping countries, but in most cases there is notsufficient capital available for the investments needed.In the industrial countries with sufficient food sup-ply the fertilizer consumption is at best stagnant forecological concerns presently exercising the collectiveminds in the Western World. The sometimes fromenvironmentalists propagated ecological agricultureis no alternative as manure and biomass are not suf-ficient in effect and quantity to supply the necessarynitrogen and in addition they have the same problemwith nitrate run-off.

Direct biological fixation is presently restricted to thelegumes by their symbiotic relationship with the Rhi-zobium bacteria, which settle in the root nodules of the plants. Intensive genetic engineering research hasprovided so far a lot of insights in the mechanism of this biological fixation but a real breakthrough forpractical agricultural application has not yet hap-

pened. The enzyme nitrogenase practically performsan ammonia synthesis in the bacteria, and for the syn-thesis of the nitrogenase the so-called NIF gene isresponsible. One option, for example, would be tobroaden the host spectrum of the Rhizobium bacte-ria by genetic manipulation. Other possibilities are totransfer the NIF gene to other bacteria which havea broader host spectrum but have no own nitrogen fix-ation ability or to insert the NIF gene directly intoplants. One important point to consider especiallywith the option of constructing a nitrogen fixing plant,is the energy balance of the plant. Because of the lowefficiency, a considerable amount of the photosynthe-sis product would be consumed to supply the energyneeded. This would consequently lead to a reductionin yield, which is estimated by some researchers to beas high as 18%.

The possibility of converting atmospheric nitrogeninto ammonia in homogeneous solution using metal-organic complexes was first raised around 1966. The

prospects for this route are not judged to be verypromising in terms of energy consumption and alsowith respect to the cost of these very sophisticated cat-

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alyst systems. Photochemical methods of producingammonia at ambient temperature and atmosphericpressure in the presence of a catalyst have beenreported, 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, combininghydrogen and nitrogen using a catalyst at elevatedtemperature and pressure in a recycle process, as con-ceived in laboratory by Fritz Haber and made oper-able on an industrial-scale by C. Bosch, and on the useof the known routes to produce the hydrogen andnitrogen needed (which are in fact, what consumes allthe energy).

In conclusion it is possible to sum up the prospects bythe following broad predictions:

Natural gas will remain the preferred feedstock forat least the next 15 years. Coal gasification will notplay a major role in ammonia production in thatperiod.

The present ammonia technology will not changefundamentally, at least in the next 15 years. Evenif there are radical, unforeseeable developments,they will take time to develop to commercial intro-

duction. With the available concepts, the marginsof additional improvements have become rathersmall after years of intensive research and devel-opment. Thus only minor improvements of individ-ual steps, catalysts and equipment might beexpected.

A further significant reduction in the energy con-sumption of the natural gas-based steam reform-ing ammonia process is unlikely; figures between27 and 28 GJ/t NH 3 are already close to theoreti-cal minimum.

In the medium term the bulk of ammonia produc-tion will still be produced in world-scale plants of 1,000 2,000 t/d NH 3. Small capacity plants will belimited to locations where special logistical, finan-cial or feedstock conditions favor them.

New developments in ammonia technology willmainly reduce investment costs and increase oper-

ational reliability. Smaller integrated process units(e. g. exchanger reformer, CAR) contribute to thisreduction 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 instrumentationand 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 largescale will occur predominantly in areas with stillstrongly growing population to secure the increas-ing food demand. This development may be pushedby the fact that compared to the classical fertilizerroute less capital and less energy would be needed.This may happen within the next 20 years, but timeestimates are always risky. (A famous example:Man will not fly for 50 years, Wilbur Wright1901). But even with the introduction of this newapproach, 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 capitalinvestment and a great number of them may reli-ably operate for at least another 20 30 years froma mere technical point of view.

References:

A. V. Slack, G.Russel James, Ammonia, Part I-III, MarcelDecker, New York, 1073, 1974, 1977

J.R. Jennings (Ed.), Catalytic Ammonia Synthesis, PlenumPress, New York London, 1991

A.Nielsen (Ed.), Ammonia Catalysis and Manufacture, Sprin-ger-Verlag, Berlin Heidelberg New York, 1995

S. A. Topham, The History of the Catalytic Synthesis of Ammo-nia in Catalysis (ed. R.Anderson and M.Boudart), Volume 7,p. 1 50, Springer-Verlag, Berlin Heidelberg New York, 1980

M. Appl, Ammonia, Methanol, Hydrogen, Carbon Monoxide Modern Production Technologies, Ed. Alexander More, BritishSulphur Publishing, 1977, ISBN 1873 387 26

M. Appl, Ammonia Synthesis and the Development of Cataly-tic and High Pressure Processes, Dr. H. L. Roy Memorial Lec-ture, IIChE Conference Hyderabad (Dec 1986); Indian Chemi-cal Engineer XXIX (1), 2-29 (1987)

M. Appl, The Haber-Bosch Process and the Development of Chemical Engineering in A Century of Chemical Engineering(ed. W.Furter), Plenum Publishing Corporation, 20-51 (1982)

M.Appl, A Brief History of Ammonia Production from the Early

Days to the Present, Nitrogen (100), 47-58, Mar-Apr 1976)

L.Connock, Ammonia and Methanol from Coal, Nitrogen (226)47-56 (Mar-Apr 1997)

haber ammonia technology

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