EP0000993B1 - Procédé pour la production d'ammoniac - Google Patents

Procédé pour la production d'ammoniac Download PDF

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EP0000993B1
EP0000993B1 EP78300276A EP78300276A EP0000993B1 EP 0000993 B1 EP0000993 B1 EP 0000993B1 EP 78300276 A EP78300276 A EP 78300276A EP 78300276 A EP78300276 A EP 78300276A EP 0000993 B1 EP0000993 B1 EP 0000993B1
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gas
ammonia
hydrogen
synthesis
stage
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EP0000993A1 (fr
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Alwyn Pinto
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Imperial Chemical Industries Ltd
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    • C—CHEMISTRY; METALLURGY
    • C01—INORGANIC CHEMISTRY
    • C01B—NON-METALLIC ELEMENTS; COMPOUNDS THEREOF; METALLOIDS OR COMPOUNDS THEREOF NOT COVERED BY SUBCLASS C01C
    • C01B3/00—Hydrogen; Gaseous mixtures containing hydrogen; Separation of hydrogen from mixtures containing it; Purification of hydrogen
    • C01B3/02—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen
    • C01B3/025—Preparation or purification of gas mixtures for ammonia synthesis
    • C—CHEMISTRY; METALLURGY
    • C01—INORGANIC CHEMISTRY
    • C01B—NON-METALLIC ELEMENTS; COMPOUNDS THEREOF; METALLOIDS OR COMPOUNDS THEREOF NOT COVERED BY SUBCLASS C01C
    • C01B3/00—Hydrogen; Gaseous mixtures containing hydrogen; Separation of hydrogen from mixtures containing it; Purification of hydrogen
    • C01B3/02—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen
    • C01B3/32—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen by reaction of gaseous or liquid organic compounds with gasifying agents, e.g. water, carbon dioxide, air
    • C01B3/34—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen by reaction of gaseous or liquid organic compounds with gasifying agents, e.g. water, carbon dioxide, air by reaction of hydrocarbons with gasifying agents
    • C01B3/38—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen by reaction of gaseous or liquid organic compounds with gasifying agents, e.g. water, carbon dioxide, air by reaction of hydrocarbons with gasifying agents using catalysts
    • C01B3/382—Multi-step processes
    • C—CHEMISTRY; METALLURGY
    • C01—INORGANIC CHEMISTRY
    • C01C—AMMONIA; CYANOGEN; COMPOUNDS THEREOF
    • C01C1/00—Ammonia; Compounds thereof
    • C01C1/02—Preparation, purification or separation of ammonia
    • C01C1/04—Preparation of ammonia by synthesis in the gas phase
    • C01C1/0405—Preparation of ammonia by synthesis in the gas phase from N2 and H2 in presence of a catalyst
    • C—CHEMISTRY; METALLURGY
    • C01—INORGANIC CHEMISTRY
    • C01C—AMMONIA; CYANOGEN; COMPOUNDS THEREOF
    • C01C1/00—Ammonia; Compounds thereof
    • C01C1/02—Preparation, purification or separation of ammonia
    • C01C1/04—Preparation of ammonia by synthesis in the gas phase
    • C01C1/0405—Preparation of ammonia by synthesis in the gas phase from N2 and H2 in presence of a catalyst
    • C01C1/0411—Preparation of ammonia by synthesis in the gas phase from N2 and H2 in presence of a catalyst characterised by the catalyst
    • C—CHEMISTRY; METALLURGY
    • C01—INORGANIC CHEMISTRY
    • C01C—AMMONIA; CYANOGEN; COMPOUNDS THEREOF
    • C01C1/00—Ammonia; Compounds thereof
    • C01C1/02—Preparation, purification or separation of ammonia
    • C01C1/04—Preparation of ammonia by synthesis in the gas phase
    • C01C1/0405—Preparation of ammonia by synthesis in the gas phase from N2 and H2 in presence of a catalyst
    • C01C1/0458—Separation of NH3
    • C01C1/0464—Separation of NH3 by absorption in liquids, e.g. water
    • C—CHEMISTRY; METALLURGY
    • C01—INORGANIC CHEMISTRY
    • C01B—NON-METALLIC ELEMENTS; COMPOUNDS THEREOF; METALLOIDS OR COMPOUNDS THEREOF NOT COVERED BY SUBCLASS C01C
    • C01B2203/00—Integrated processes for the production of hydrogen or synthesis gas
    • C01B2203/14—Details of the flowsheet
    • C01B2203/142—At least two reforming, decomposition or partial oxidation steps in series
    • C01B2203/143—Three or more reforming, decomposition or partial oxidation steps in series
    • C—CHEMISTRY; METALLURGY
    • C01—INORGANIC CHEMISTRY
    • C01B—NON-METALLIC ELEMENTS; COMPOUNDS THEREOF; METALLOIDS OR COMPOUNDS THEREOF NOT COVERED BY SUBCLASS C01C
    • C01B2203/00—Integrated processes for the production of hydrogen or synthesis gas
    • C01B2203/14—Details of the flowsheet
    • C01B2203/146—At least two purification steps in series
    • C01B2203/147—Three or more purification steps in series
    • C—CHEMISTRY; METALLURGY
    • C01—INORGANIC CHEMISTRY
    • C01B—NON-METALLIC ELEMENTS; COMPOUNDS THEREOF; METALLOIDS OR COMPOUNDS THEREOF NOT COVERED BY SUBCLASS C01C
    • C01B2203/00—Integrated processes for the production of hydrogen or synthesis gas
    • C01B2203/80—Aspect of integrated processes for the production of hydrogen or synthesis gas not covered by groups C01B2203/02 - C01B2203/1695
    • C01B2203/82—Several process steps of C01B2203/02 - C01B2203/08 integrated into a single apparatus
    • Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
    • Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
    • Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
    • Y02P20/00—Technologies relating to chemical industry
    • Y02P20/50—Improvements relating to the production of bulk chemicals
    • Y02P20/52—Improvements relating to the production of bulk chemicals using catalysts, e.g. selective catalysts

Definitions

  • This invention relates to an ammonia production process in which the energy consumption can be smaller than in previously used processes.
  • the commonest ammonia production process comprises the following reactions in sequence:
  • Reaction 1 proceeds incompletely in economic conditions of steam ratio, temperature and pressure but the small percentage of unreacted methane can be largely converted in reaction 2, which is carried out by feeding preheated air into the hot gas leaving reaction 1 and bringing the reaction mixture towards equilibrium over a catalyst.
  • the conditions in reaction 1 are chosen so that the quantity of air required in reaction 2 to provide oxygen also introduces the correct proportion of nitrogen required in reaction 5. It has, however, been proposed (UK specifications GB-A 1156002, GB-A 1156003) to have substantially more methane in the gas leaving reaction 1, to use more air in reaction 2 and to remove the resulting excess of nitrogen cryogenically between reactions 4 and 5.
  • ammonia is recovered from reacted synthesis gas by absorption in water.
  • Such a method of ammonia recovery is described in above-cited UK specification GB-A 1186939 and a more detailed description is given in UK specification GB-A 1115952. Neither of these references discloses the application of that method in the process sequence according to the present invention.
  • step will be used in referring to the major operations listed above and the term “stage” in referring to detailed operations within the steps.
  • the hydrocarbon feedstock is suitably methane or other steam reformable hydrocarbon such as a normally gaseous or liquid hydrocarbon boiling at up to about 220°C.
  • the primary catalytic reforming can be in one stage, over a catalyst externally heated or, when the feedstock is of a higher molecular weight than methane and especially when it is normally liquid, can be in two stages, in the first of which the feedstock is catalytically converted with steam to a gas of high methane content at an outlet temperature under 650°C and in the second of which that gas is reacted in the externally heated process.
  • Various types of supported nickel catalyst are available for these hydrocarbon-steam reactions.
  • the feedstock preferably should be substantially sulphur-free (under 0.5 ppm S) and may have been subjected to a preliminary catalytic hydrodesulphurisation and H 2 S-removal treatment.
  • External heating can be effected by having the catalyst in tubes surrounded by a furnace or in adiabatic beds preceded by heating zones.
  • the hydrocarbon steam mixture is preheated, preferably to a temperature in the range 600-700°C which is higher than commonly used heretofore.
  • the hydrocarbon is methane or ethane it can be used directly at such a high preheat temperature.
  • the preliminary conversion to gas of high methane content is useful in permitting the preferred high pre-heat temperature.
  • the pressure at the outlet of step (a) is preferably in the range 30-120, especially 40-80 bars abs. and the temperature is in the range 750-850°C.
  • the steam ratio is preferably in the range 2.5-3.5 molecules of total steam per atom of carbon in the hydrocarbon if all the steam and hydrocarbon are fed to this stage.
  • the relatively high pressure gives rise to efficient heat transfer into the tubes and makes possible a saving of compression power further downstream.
  • the relatively low steam ratio decreases the quantity of heat that has to be fed in and removed later.
  • the temperature range is one for which thermally efficient steam reforming furnaces are available; at these temperatures the methane content of the product gas is typically up to 30, for example in the range 10 to 20% by volume on a dry basis and this is preferred, although it is considerably higher than is normally thought. suitable in the primary reformer gas of an ammonia synthesis plant.
  • the feed to step (b) includes the gas from step (a) (which may, if desired be further heated before entering step (b)) and an oxygen-containing gas, which is conveniently air and preferably is preheated to a temperature in the range 400-700°C. Further steam can be fed to step (b) or further hydrocarbon feedstock if it is desired to minimise the total steam ratio without having too low a steam ratio in step (a), but usually all the hydrocarbon and steam are fed to step (a).
  • the outlet pressure is conveniently about the same as at the outlet of step (a), apart from the pressure drop through the secondary reforming catalyst.
  • the outlet temperature is preferably in the range 950-1050°C and the outlet methane content in the range 0.2 to 10% v/v on a dry basis.
  • Such a temperature is rather higher than has been proposed for processes using an over-stoichiometric air feed rate and has the effect of reacting a greater proportion of the hydrocarbon feedstock in the internally-heated secondary reformer and thus less in the thermally less efficient externally heated primary reformer, and also of permitting a lower steam ratio than the preferred high pressure would otherwise demand.
  • the excess of air results in a gas containing preferably 2.2 to 2.7 molecules of hydrogen equivalent (i.e. total of H 2 and CO) per molecule of nitrogen.
  • the catalyst in step (b) can be for example a supported nickel catalyst or chromium oxide catalyst or a combination thereof. Usually the reaction mixture is brought substantially to equilibrium.
  • step (b) By introducing the excess air into step (b) it becomes possible to operate step (a) to produce a gas of higher methane content, that is, at a lower temperature and/or lower steam ratio and/or higher pressure.
  • step (c) the gas from step (b) is cooled with recovery of useful heat to the inlet temperature of the "shift" catalyst over which the reaction of carbon monoxide with steam occurs.
  • this temperature is in the range 300-400°C, especially 320­350°C, appropriate to iron-chrome shift catalysts.
  • the reaction over the iron-chrome catalyst is exothermic (outlet temperature 400-450°C) and the outlet gas is again cooled with recovery of useful heat and subjected to further shift reaction, preferably over a copper-containing catalyst, for which the inlet temperature is suitably 200 ⁇ 240° C and the outlet temperature 240-270°C.
  • Such a low-temperature shift stage produces a gas containing usually 0.2 to 0.6% v/v of carbon monoxide on a dry basis.
  • the steam to dry gas molar ratio in the gas entering the low temperature shift catalyst is kept down to a level that avoids damage to the catalyst, preferably in the range 0.1 to 0.3 when the pressure is over 40 bar abs.
  • the steam to carbon ratio in steps (a) and (b) should not be too high, but it can be readily attained using ratios in the range 2.5 to 3.5 (methane feedstock) or 2.4 to 3.2 (feedstock of empirical formula CH 2 ) or in intermediate ranges for hydrocarbons of intermediate composition.
  • the higher nitrogen content due to excess air helps to keep down the steam to dry gas ratio.
  • the steam to dry gas ratio can be kept down by recycling synthesis gas from a downstream point after removal of steam and carbon dioxide and possibly after compression.
  • shift conversion systems such as iron chrome/C0 2 removal/iron chrome or systems based on catalysts containing other Group VI and Group VIII metals and oxides or on alkali metal compounds of weak acids could be used if desired.
  • Step (d) removal of carbon oxides, is usually carried out in a first stage in which carbon dioxide is substantially removed, and a second in which residual carbon monoxide and carbon dioxide are removed down to a very low level such that the ammonia synthesis catalyst is not significantly poisoned by them. If desired, any carbon monoxide remaining after shift step (c) can be selectively oxidised to carbon dioxide.
  • the first step of step (d) can be carried out using any liquid absorbent.
  • Well-established chemical systems such as Benfield's potassium carbonate or diethanolamine-potassium carbon, "Vetrocoke”, “Catacarb” or amine systems such as monoethanolamine can be used. These have, however, the disadvantage of consuming a substantial quantity of steam in the regeneration of the liquid absorbent, a requirement that is especially inconvenient when the steam to dry gas ratio of the shifted gas is at the preferred low levels.
  • step (a) which is substantially maintained apart from inevitable pressure drops in subsequent steps, makes possible the use of "physical" absorbents, the preferred examples of which can be regenerated merely by lowering pressure.
  • Suitable absorbents used in industrially developed processes are tetramethylene sulfone ("Sulfinol”) propylene carbonate (Fluor), N-methyt-2-pyrroiidone (“Purisol”), methanol (“Rectisol”) and the dimethyl ether of polyethyleneglycol (“Selexol”).
  • part or all the carbon dioxide can be removed by absorption in anhydrous or aqueous ammonia.
  • Such a procedure is especially useful if the ammonia to be produced by the process of the invention is to be used for urea synthesis or for making ammonium sulphate by the calcium sulphate process.
  • the bulk of the carbon dioxide can be removed in a physical absorbent and the remainder in a chemical solvent as mentioned above or in ammonia.
  • the latter procedure can be designed to suit any desired relative outputs of ammonia, carbon dioxide and urea.
  • the second stage of step (d) can be carried out by contacting the gas with a carbon oxides absorbent such as copper liquor but is most conveniently effected by catalytic methanation, for example over a supported nickel catalyst at an outlet temperature in the range 250-400°C. This decreases the carbon oxides content to a few parts per million by volume but produces water, which can be removed by cooling, separation and passage over a water-absorbent such as alumina or a molecular sieve.
  • a carbon oxides absorbent such as copper liquor but is most conveniently effected by catalytic methanation, for example over a supported nickel catalyst at an outlet temperature in the range 250-400°C.
  • This decreases the carbon oxides content to a few parts per million by volume but produces water, which can be removed by cooling, separation and passage over a water-absorbent such as alumina or a molecular sieve.
  • the dried gas from step (d) contains nitrogen, hydrogen in less than the stoichiometric ratio from ammonia synthesis, a small quantity (usually under 1% v/v) of methane and fractional percentages of noble gases introduced with the secondary reformer air, and is thus ready for use as an ammonia synthesis gas. It may be compressed to any convenient synthesis pressure, for example in the range 120-400 bar abs. Owing, however, to its unusually high pressure, it is preferably passed to the synthesis with less than 50% compression and preferably, no more than the increase in pressure (for example up to 20%) encountered in a circulator in a recycle-type synthesis.
  • the pressure increase can be rather larger without seriously impairing the energy economy of the process, preferably such as can be provided by a single machine acting as compressor and circulator and typically up to 100 bar and more conveniently in the range 20-80 bar.
  • the compressorcirculator is a single-barrel machine.
  • the compressor and circulator functions can be separated within the machine so as, for example to provide for mixing of recycle gas with fresh synthesis gas outside the barrel. As an alternative such mixing can be effected within an undivided barrel.
  • the larger pressure increase is advantageous in the process described below employing cryogenic separation of a hydrogen-rich stream from reacted synthesis gas because it permits the use of gas expansion to provide part of the cooling.
  • a side stream of synthesis gas is taken downstream of the compressor section of the machine or of the circulator section and passed through a hydrogen separation unit, whereafter the hydrogen rich fraction is returned to the process at the inlet of the compressor or to an intermediate level therein, depending on the. extent to which its pressure has been decreased in the separation unit.
  • the "fresh" synthesis gas from step (d) can be fed through a succession of catalytic stages and ammonia removal stages but, as in most ammonia synthesis processes is preferably mixed with synthesis gas recycled from an ammonia removal stage.
  • the attainable pass conversion over the synthesis catalyst is relatively low, giving an ammonia outlet concentration in the range 8 to 12% v/v.
  • the ratio of recycled gas to fresh gas is suitably in the range 4 to 6.
  • the catalyst used in the ammonia synthesis can be of the usual composition, namely iron with promoting quantities of non-reducible oxides such as those of potassium, calcium, aluminium and others such as of beryllium, cerium or silicon.
  • the iron catalyst preferably contains also cobalt, suitably to the extent of 1-20% w/w calculated as C 03 0 4 on the total oxidic composition from which the catalyst is made by reduction and in which the iron oxide is assumed to be all Fe 3 0 4 .
  • the catalyst can be in the form of particles in the sieve range 18 to 4 ASTM (1--4.7 mm) especially 10 to 5 (2-4 mm), if it is desired to maximise their available contact surface or larger, for example up to 20 mm; the arrangement of the catalyst in the synthesis reactor preferably therefore may afford short gas flow paths, such as radial or secantial flow in a cylindrical reactor.
  • the outlet temperature of the synthesis catalyst is preferably in the range up to 500°C, especially 300-450 0 C for example 350-430°C. This is lower than has been usual, in order to obtain a more favourable synthesis equilibrium.
  • the catalyst volume is suitably in the range 100-200 m 3 per 1000 metric tons per day output; this is higher than has been usual but can be tolerated because at the low operating temperature and pressure the reactor can be of simple construction for example of the hot-wall type.
  • the recovery of ammonia from reacted synthesis gas can be carried out by ordinary air-cooling or water-cooling if the pressure is high enough, but at preferred pressures in the range 40-80 bar abs. is best carried out by absorption in water. Absorption in an acid or on a solid such as zinc chloride can be used if convenient. Absorption in water is conveniently carried out in two or more stages, in the first of which the gas contacts a relatively strong ammonia solution (for example 15 to 30% w/w) and in the last pure water or a weak ammonia solution (for example up to 10% w/w).
  • a relatively strong ammonia solution for example 15 to 30% w/w
  • a weak ammonia solution for example up to 10% w/w.
  • the water temperatures and flow rates are preferably such that the gas leaving the absorption contains less than 2.0% v/v of ammonia: this has the effect of increasing the pass conversion over the catalyst and thus compensating for the lower pressure.
  • the gas is dried in order to prevent deactivation of the catalyst by water.
  • the aqueous ammonia product can be used as such or distilled to recover anhydrous ammonia from it.
  • the fresh synthesis gas contains nitrogen in excess of the stoichiometric proportion, noble gases and also methane to an extent dependent on the incompleteness of the secondary reforming reaction and of the shift reaction, the continued removal of ammonia from it, especially in a recycle process, results in a substantial concentration of non-reacting gases. It is preferred to treat the gas mixture to remove such gases before their concentration reaches 10% v/v especially 5.0% v/v. This treatment could be applied to the whole of the reacted gas after removal of ammonia or even to the mixture of reacted gas and fresh synthesis gas, but it is preferred to apply it only to a side stream, because then any failure of the treatment plant does not cause a shutdown of the whole production process.
  • the side stream can conveniently be taken from the gas downstream of the ammonia separation and treated for hydrogen separation, whereafter the hydrogen is returned to the circulating synthesis gas. It could be taken before ammonia separation but the treatment would then involve also ammonia recovery.
  • the hydrogen separation treatment involves a pressure-drop and may involve also a pressure let-down through an expansion engine in order to decrease the gas temperature for cryogen separation; consequently the hydrogen stream has to be compressed to return it.
  • the side stream is taken from the effluent of the circulator, where the gas pressure in the system is highest, and the separated hydrogen stream is returned to the inlet of the circulator, where the gas pressure is lowest. Part or all of the separated hydrogen stream can be recycled to the low temperature shift inlet.
  • the hydrogen separation treatment can be by any suitable means, for example by cryogenic fractionation, molecular sieve adsorption of gases other than hydrogen or palladium membrane diffusion.
  • the hydrogen stream returned to the synthesis can be substantially (over 90% v/v) pure but in any event should contain at least 3 molecules of hydrogen per nitrogen molecule.
  • the non-reactive gases discarded from the hydrogen separation treatment should of course be substantially free of hydrogen, since any discarded hydrogen represents wasted energy.
  • the separation treatment can be designed and operated to separate a methane-rich stream and that stream can be used as process feed or furnace fuel for step (a) or fed to step (b).
  • a typical side stream flow rate is in the range 15 to 30% of total gas flow.
  • a preferred cryogenic purification of the hydrogen stream in step (f) comprises the stages:
  • stage f the hydrogen-depleted phase will evaporate in stage f) and possibly in part in stage g). It is then discarded possibly by discharge to atmosphere but possibly to a use for example as an auxiliary coolant or as a working fluid in a heat engine or as a fuel, depending on its composition and on local requirements.
  • the gas entering stage a) is suitably at a temperature in the range 0-50 0 C, typical of gas leaving an ammonia separation system or leaving a circulator.
  • the steam raised at 100-200 bar is let down in one or more pass-out turbines exhausting at a pressure suitably of 45-90 bar and the exhaust is partly used as feed to step (a), partly let down in one or more further turbines exhausting at low pressures or to condenser.
  • the turbines drive the process air compressor, the circulator and (preferably by way of electrical generators) the smaller machines such as liquid pumps, furnace combustion air pump and flue gas fan.
  • the waste heat recoverable from furnace flue gas is less per metric ton of ammonia per day than in processes in common use but at the same time, in the absence of a compression stage, less energy is to be found from waste heat. Consequently the process represents a new balance of energetic provision and use.
  • the plant combination is also novel.
  • a heated mixture of natural gas and steam is fed at 10 into the catalyst filled primary reforming tubes heated in furnace 14.
  • the resulting hot gas consisting of carbon oxides, hydrogen, methane and excess steam is fed to secondary reformer 16 and heated air at 18 is fed into that gas by means of a suitable nozzle.
  • a flame is formed in the space at the top of reformer 16 but the temperature falls as the methane-steam reaction comes to equilibrium over the catalyst.
  • the resulting gas has a very low methane content as a result of the relatively high air feed rate into 16 and also contains nitrogen in excess of what will react with the hydrogen that will be present after the subsequent stages.
  • the gas leaving 16 is cooled at 20, by passage through a high pressure boiler, a boiler feed water heater and possibly other heat exchanges such as with natural gas and process air, to the inlet temperature of high temperature shift catalyst 22.
  • the shifted gas is cooled at 24, which like 20 also includes a high pressure boiler and other heat exchanges, to the inlet temperature of low temperature shift catalyst 26.
  • Shifted gas from 26 is cooled at 28, which represents low grade heat recoveries such as boiler feed water heating and possibly carbon dioxide absorbent regeneration and includes cooling to below the dewpoint of steam such that water is separated from the gas in separator 30.
  • the resulting relatively dry gas is contacted in packed tower 32 with a carbon dioxide absorbent solution fed in at 34.
  • Carbon dioxide-loaded solution is taken off at 36 and passed to a regeneration tower (not shown) and returned to point 34.
  • Gas substantially free of carbon dioxide leaves the top of tower 32, is heated in feed/effluent heat exchanger 38 and enters methanator 40 in which residual carbon monoxide and carbon dioxide are converted to methane.
  • Methanated gas is cooled by heat exchange with feed gas at 38 and united at mixing point 44 with a steam of recycled synthesis gas to be described.
  • the mixed gas is cooled further at 42 and passed to separator 46, in which the bulk of its water content is removed, optional mixing point 47 (see path B below), where a recycle hydrogen stream is fed in, and drier 48 in which it is dehydrated to the low level of water required in the synthesis.
  • Dehydrated gas is passed into circulator 50 in which its pressure is increased sufficiently to maintain circulation, united at 52 with a recycle hydrogen stream to be described, heated to catalyst inlet temperature in feed/effluent heat exchanger 54 and passed into ammonia synthesis reactor 56.
  • reactor 56 is shown with a single catalyst bed and gas feed point, with an external feed/effluent heat exchanger, without internal cooling and without external heat recoveries. In practice a more complicated reactor would be used).
  • In 56 incomplete conversion to ammonia occurs.
  • the reacted gas is cooled by stages to a temperature at which ammonia can be separated from it.
  • the cooling stages include feed/effluent heat exchange at 54 and possibly other heat exchanges such as steam raising and boiler feed water heating, ending with final cooling with discard of heat represented generally by 58.
  • Ammonia is recovered as a liquid in separator 60, which can be of the simple type with baffle plates if the gas pressure is high enough and cooling at 58 is to below the dewpoint of ammonia; alternatively and more conveniently separator 60 is a packed absorber down which water is fed countercurrently to the ammonia-containing gas.
  • separator 60 is a packed absorber down which water is fed countercurrently to the ammonia-containing gas.
  • anhydrous ammonia is run off to store or aqueous ammonia is run off to a distillation system producing anhydrous ammonia overhead and water, possibly also a relatively strong ammonia solution, as bottoms, for recycle to the absorber.
  • Reacted gas leaving the top of separator 60 is then fed to point 44 and united with methanated gas.
  • the purge gas stream taken at 64
  • Figure 1 shows two methods for treating reacted synthesis gas for recovery of hydrogen.
  • a purge stream (25% v/v of total flow), is diverted at 64 and fed to hydrogen recovery unit 66.
  • Unit 66 typically includes a drier and a cryogenic fractionation plant in which the gas is cooled to minus 188°C to condense out nitrogen, methane and noble gases.
  • the uncondensed fraction consisting of substantially pure hydrogen is fed via blower 70 to point 52 where it is united with synthesis gas about to be heated and fed to synthesis reactor 56.
  • the condensed fraction is reevaporated in order to cool purge gas entering unit 66 and leaves at 68. If it contains sufficient methane it can be used as fuel in furnace 14. Alternatively it can be fractionated during evaporation in order to produce a methane stream to be used as furnace fuel or as process feed at 10.
  • path B a side stream (25% of total flow) of mixed gas is taken at 51 at the outlet of circulator 50 and passed into hydrogen recovery unit 66.
  • the hydrogen fraction from 66 is fed to point 47 upstream of drier 48.
  • Path B has the advantage that the pressure difference across circulator 50 is utilised to drive the gas through unit 66, so that no blower 70 is needed.
  • the path B process is especially suitable when circulator 50 includes a compression stage: then the hydrogen stream is fed to the inlet of that stage instead of at 47.
  • line 110 carries a mixture of make-up fresh ammonia synthesis gas and unreacted gas from an ammonia separator or absorber.
  • the gas after treatment in the process of the invention will pass out at 118 to a final drier and then via the usual heat exchangers to the ammonia synthesis catalyst.
  • At point 112 there enters a side stream of hydrogen-enriched gas made by the process of the invention.
  • the mixed stream is compressed at 114 to at least the extent required to maintain circulation.
  • a side stream is taken off and may be. washed with absorber feed water to remove residual ammonia from it. It is passed through drier 120 charged with solid drying agent such as a molecular sieve.
  • This stream is chilled at 122 by indirect heat exchange with circulated refrigerant, passed through first cold box 124 and then let down in pressure in engine 126 with performance of work W.
  • chiller 122 can be omitted if the pressure increase at 114 is relatively large.
  • drier 120 can be omitted if cold box 124 is operated with switching gas paths.
  • untreated gas following path A until sufficient condensation of water and ammonia or of ammonia only if the gas has previously been dried, as taken place on the walls of pipe A to increase its pressure drop measurably.
  • paths A and C are interchanged so that untreated gas passes through a clean pipe and the dry gas carries away the condensed water and ammonia.
  • paths A and B are interchanged, so that the condensed ammonia is carried back into stream 110.
  • Gas cooled by expansion in engine 126 is then passed through cold box 128 to below the dewpoint of at least one of its components and thence into separator 130.
  • Uncondensed gas enriched in hydrogen passes overhead at 132 and constitutes one of the cold streams entering cold box 128.
  • Bottoms liquid passing out at 134 is let down in pressure through valve 135 and passes into cold box 128, where it is allowed to evaporate, constituting the other cold stream.
  • the resulting two gas streams having been warmed in cold box 128 constitute the cool streams heat exchanged with untreated gas in cold box 124 before cooling by expansion in engine 126.
  • the hydrogen-enriched gas warmed further in cold box 124 is passed back to the synthesis gas stream at 112, at the low pressure side of compressor/circulator 114.
  • the hydrogen-depleted stream is passed out at 138.
  • Table 1 shows the gas pressure, temperature, flow rate and composition at the significant positions in a process according to figure 1 for the manufacture of 1000 metric tons per day of ammonia (as NH 3 ).
  • the H 2 :N 2 ratio of the methanated gas is 2.34, as a result of the rate of feed or air to the secondary reformer, which rate is about 28% greater than in a process producing a gas with an H 2 :N 2 ratio of 3.0.
  • the primary reformer to produce a gas containing 17.2% v/v of methane on a dry basis, that is, to operate at the unusually high pressure of 45.2 bar abs.
  • the temperature of the natural gas/steam mixture entering the primary reformer is 640°C, which is higher than usual and is attained at the cost of more expensive alloy used in the construction of the required preheater, but usefully increases the conversion of methane in the primary reformer.
  • the unusually low H 2 :N 2 ratio has the effect that hydrogen is more completely reacted (other things being equal) in the synthesis reactor.
  • the recycle rate and purge rate are both higher than usual, but the high purge rate does not result in loss of hydrogen because hydrogen is recovered from it and returned to the synthesis.
  • the catalyst volume is 200 m 3 using a promoted iron catalyst or 120 m 3 using a promoted iron-cobalt catalyst, as explained further below.
  • the side stream taken at 116 is 20% v/v of the effluent from 114 and has the % v/v composition 71.2 H Z , 23.7 N 2 , 3.9 CH 4 , 1.2 Ar.
  • the hydrogen-enriched stream returned at 112 then has the % v/v composition 85.6 H 2 , 13.6 N 2 , 0.4 CH 4 , 0.4 Ar.
  • the pressure at the inlet of circulator 114 is 46 bar and the outlet 51 bar in this typical process.
  • ammonia synthesis data in the table are based on activity measurements on a catalyst made by reducing the following oxide composition, weight percent, as made by co-fusion and solidification.
  • This catalyst was compared with the corresponding cobalt-free catalyst by passing over them in parallel reactors a stoichiometric ammonia synthesis gas and adjusting the gas space velocity to produce an outlet gas containing the same volume percentage of ammonia.
  • the activity of the cobalt-containing catalyst relative to that of the cobalt-free catalyst taken as 100 at various ammonia synthesis temperatures was:
  • the cobalt-containing catalyst maintains its activity approximately as well as the cobalt-free catalyst.
  • the volume of the cobalt-containing catalyst required for a given ammonia output is typically under 75% that of a corresponding cobalt-free catalyst and as low as 50% at the lower temperature at which the cobalt-containing catalyst is active.

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Claims (10)

1. Procédé de production de l'ammoniac caractérisé en ce que:
(a) on effectue un réformage catalytique primaire, à une pression supérieure à la pression atmosphérique, d'une charge formée par un hydrocarbure avec de la vapeur d'eau, dans des conditions de rapport vapeur-carbone, de pression et de température fournissant un gaz contenant des oxydes de carbone, de l'hydrogène et au moins 10% v/v de méthane;
(b) on effectue un réformage catalytique secondaire du gaz provenant du stade (a) par introduction d'air et en amenant le mélange vers un équilibre, de façon à produire un gaz contenant del'azote, des oxydes de carbone, de l'hydrogène et une quantité réduite de méthane, la quantité d'air étant suffisante pour introduire une molécule d'azote pour 2,0 à 2,9 molécules d'hydrogène;
(c) on convertit par voie catalytique le monoxyde de carbone en dioxyde de carbone et en hydrogène avec de la vapeur d'eau;
(d) on élimine les oxydes de carbone pour fournir un gaz de synthèse de l'ammoniac contenant de l'azote et de l'hydrogène;
(e) on fait réagir le gaz de synthèse pour produire de l'ammoniac et on récupère l'ammoniac à partir du gaz ayant réagi;
(f) on traite le gaz de synthèse après réaction pour effectuer la synthèse de l'ammoniac afin de séparer un courant enrichi en hydrogène et on renvoie le courant enrichi à la synthèse; et
(g) on rejette les gaz non réactifs présents dans le gaz de synthèse.
2. Procédé suivant la revendication 1, caractérisé en ce qu'au cours du stade (a), le rapport vapeur-carbone est compris dans une gamme allant de 2,5 à 3,5, la pression de sortie est comprise dans une gamme allant de 30 à 120 bars en pression absolue, et la température de sortie est comprise dans une gamme allant de 750 à 850°C.
3. Procédé suivant la revendication 1 ou 2, caractérisé en ce qu'au cours du stade (b), la température de sortie est comprise dans une gamme allant de 950 à 1050°C, la teneur en méthane à la sortie est comprise dans une gamme allant de 0,2 à 10% v/v sur une base sèche, et le rapport d'équivalent d'hydrogène à l'azote dans le gaz de sortie est compris dans une gamme allant de 2,2 à 2,7.
4. Procédé suivant l'une quelconque des revendications précédentes, caractérisé en ce que la pression au cours du stade (a) et au cours des stades ultérieurs est d'au moins 40 bars en pression absolue et le stade (c) comprend une opération de conversion à basse température, avec un rapport molaire entre la vapeur et le gaz sec compris dans une gamme allant de 0,1 à 0,3.
5. Procédé suivant l'une quelconque des revendications précédentes, caractérisé, en ce que la pression au cours du stade (a) et des stades ultérieurs est d'au moins 40 bars en pression absolue, et en ce que le stade (d) comprend l'élimination du dioxyde de carbone en deux opérations, au cours de la première desquelles on utilise comme agent d'absorption une amine pouvant être régénérée sans chauffage, tandis qu'au cours de la seconde on utilise un agent d'absorption exigeant de la chaleur pour la régénération.
6. Procédé suivant l'une quelconque des revendications précédentes, caractérisé en ce que pour le stade (e) la pression du gaz de synthèse est augmentée d'une valeur ne dépassant pas 50%.
7. Procédé suivant l'une quelconque des revendications précédentes, caractérisé en ce qu'au cours du stade (e) le catalyseur de synthèse de l'ammoniac renferme du fer, des composés d'activation et du cobalt, la teneur en cobalt étant de 1 à 20% p/p calculée en C0304 sur la base de la composition oxydée totale à partir de laquelle le catalyseur est préparé, l'oxyde de fer étant supposé être en totalité du Fe304, la température de sortie du catalyseur de synthèse étant comprise dans une gamme allant de 350 à 430°C.
8. Procédé suivant l'une quelconque des revendications précédentes, caractérisé en ce qu'au cours du stade (e) de l'ammoniac est récupéré à partir du gaz de synthèse ayant réagi par absorption dans l'eau.
9. Procédé suivant l'une quelconque des revendications précédentes, caractérisé en ce que le stade (e) est mis en oeuvre par pompage du gaz de synthèse dans une boucle comprenant le catalyseur de synthèse, un appareil de séparation de l'ammoniac et un appareil de circulation et en ce qu'au cours du stade (f) un courant partiel du gaz de synthèse est, après réaction pour effectuer la synthèse de l'ammoniac et après récupération de l'ammoniac à partir de ce courant, prélevé au gaz quittant l'appareil de circulation et traité pour séparer un gaz enrichi en hydrogène, le gaz enrichi en hydrogène étant renvoyé à l'entrée de l'appareil de circulation.
10. Procédé suivant l'une quelconque des revendications précédentes, caractérisé en ce qu'au cours du stade (f) un courant de gaz partiel est traité:
(a) en le soumettant au cours d'une première opération de refroidissement indirect à un échange thermique avec un ou plusieurs courants frais qui sont décrits;
(b) en refroidissant le produit de l'opération (a) par détente dans un moteur;
(c) en soumettant l'effluent du moteur, au cours d'une seconde opération de refroidissement indirect, à un échange thermique avec un ou plusieurs courants froids qui sont décrits, afin de réduire ainsi sa température jusqu'au-dessous du point de rosée de l'azote;
(d) en séparant une phase liquid contenant de l'azote, du méthane et éventuellement des gaz nobles;
(e) en transférant la phase gazeuse enrichie en hydrogène provenant de l'opération (d) dans une relation d'échange thermique à l'opération (c) pour former l'un des courants froids;
(f) en faisant passer la phase liquide appauvrie en hydrogène provenant de l'opération (d) dans une relation d'échange thermique dans l'opération (c) pour former l'un de ces courants froids;
(g) en faisant passer les courants ainsi réchauffés au cours des opérations (e) et (f) dans une relation d'échange thermique lors de l'opération (a) pour former les courants froids;
(h) en renvoyant la phase aqueuse enrichie en hydrogène à la boucle de synthèse.
EP78300276A 1977-08-22 1978-08-10 Procédé pour la production d'ammoniac Expired EP0000993B1 (fr)

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EP (1) EP0000993B1 (fr)
JP (1) JPS6058165B2 (fr)
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DE (1) DE2862117D1 (fr)
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Cited By (2)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
EP0106076B1 (fr) * 1982-09-13 1990-02-07 The M. W. Kellogg Company Préparation de gaz de synthèse d'ammoniac
DE102007017403A1 (de) * 2007-04-13 2009-06-10 Lurgi Gmbh Verfahren und Vorrichtung zum Herstellen von NH3-Synthesegas

Families Citing this family (41)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US4409196A (en) * 1979-04-24 1983-10-11 Foster Wheeler Energy Corporation Synthesis gas for ammonia production
DK2481A (da) * 1980-01-08 1981-07-09 Montedison Spa Fremgangsmaade til syntese af ammoniak ud fra carbonhydrider
EP0049967A1 (fr) * 1980-10-14 1982-04-21 Imperial Chemical Industries Plc Procédé de production de l'ammoniac
EP0080270B1 (fr) * 1981-11-19 1985-09-04 Imperial Chemical Industries Plc Procédé de synthèse et réacteur
DE3363367D1 (en) * 1982-04-14 1986-06-12 Ici Plc Ammonia production process
US4485070A (en) * 1983-03-18 1984-11-27 Phillips Petroleum Company Control apparatus for an ammonia manufacturing process
US4783516A (en) * 1983-03-31 1988-11-08 Union Carbide Corporation Polysilane precursors containing olefinic groups for silicon carbide
US4524056A (en) * 1983-07-05 1985-06-18 Foster Wheeler Energy Corporation Process for the production of ammonia
US4592903A (en) * 1983-11-10 1986-06-03 Exxon Research & Engineering Co. Low severity hydrocarbon steam reforming process
WO1987002347A1 (fr) * 1985-10-21 1987-04-23 Union Carbide Corporation Recuperation amelioree d'hydrogene dans les ecoulements de gaz de rejet
GB8410517D0 (en) * 1984-04-25 1984-05-31 Ici Plc Ammonia synthesis
US4689317A (en) * 1984-08-03 1987-08-25 Imperial Chemical Industries Plc Catalyst precursor for ammonia synthesis and process for its production
EP0174079B1 (fr) * 1984-08-03 1990-08-22 Imperial Chemical Industries Plc Catalyseurs de synthèse de l'ammoniac à base d'oxides de fer, d' aluminium et de cobalt
US4568530A (en) * 1984-10-16 1986-02-04 The M. W. Kellogg Company Ammonia synthesis
EP0200315A3 (fr) * 1985-03-25 1988-05-25 Imperial Chemical Industries Plc Catalyseurs
DE3719780A1 (de) * 1987-06-13 1988-12-22 Uhde Gmbh Verfahren zur herstellung von ammoniak aus erdgas
GB8724474D0 (en) * 1987-10-19 1987-11-25 Ici Plc Ammonia synthesis gas
US4846851A (en) * 1987-10-27 1989-07-11 Air Products And Chemicals, Inc. Purification of ammonia syngas
DE3743595A1 (de) * 1987-12-22 1989-07-13 Wacker Chemie Gmbh Verfahren zur verringerung des halogengehaltes von halogenhaltigen polycarbosilanen und polysilanen
DE4010602A1 (de) * 1989-04-05 1990-10-11 Piesteritz Agrochemie Verfahren zur stofflichen nutzung von restgasen aus einer edelgasanlage
DE4010603A1 (de) * 1989-04-05 1990-10-11 Piesteritz Agrochemie Verfahren zur stofflichen nutzung von produktentspannungsgas
DD295741A7 (de) * 1989-04-05 1991-11-14 Stickstoffwerke Ag Wittenberg,De Verfahren zur steuerung des methan- und edelgasgehaltes eines ammoniak-wasserstoffrueckgewinnung-edelgas-komplexes
US4981676A (en) * 1989-11-13 1991-01-01 Minet Ronald G Catalytic ceramic membrane steam/hydrocarbon reformer
DK173917B1 (da) * 1998-07-02 2002-02-18 Topsoe Haldor As Fremgangsmåde til fremstilling af ammoniak
US6002059A (en) * 1998-10-28 1999-12-14 Mobil Oil Corporation Process for upgrading natural gas
US6685754B2 (en) 2001-03-06 2004-02-03 Alchemix Corporation Method for the production of hydrogen-containing gaseous mixtures
US6620398B2 (en) 2001-03-06 2003-09-16 Alchemix Corporation Method for the production of ammonia
US6663681B2 (en) 2001-03-06 2003-12-16 Alchemix Corporation Method for the production of hydrogen and applications thereof
EP2022754A1 (fr) * 2007-08-08 2009-02-11 Ammonia Casale S.A. Procédé de production de gaz synthétique d'ammoniac
US7867460B2 (en) * 2007-11-26 2011-01-11 Kellogg Brown & Root Llc Efficiency of ammonia processes
US8617270B2 (en) * 2008-12-03 2013-12-31 Kellogg Brown & Root Llc Systems and methods for improving ammonia synthesis efficiency
EP2316792A1 (fr) * 2009-10-27 2011-05-04 Ammonia Casale S.A. Procédé de production d'ammoniac
US8987402B2 (en) 2010-02-26 2015-03-24 General Electric Company Stoichiometric silicon carbide fibers from thermo-chemically cured polysilazanes
US8889093B2 (en) * 2010-09-16 2014-11-18 Kellogg Brown & Root Llc High pressure cyrogenic process and system for producing ammonia products
US8992842B2 (en) * 2012-04-17 2015-03-31 Roger Gordon Systems and methods of making ammonia using hydrogen and nitrogen gases
JP6082915B2 (ja) * 2012-10-15 2017-02-22 日本パイオニクス株式会社 アンモニア及び水素の回収方法
EP3299336A1 (fr) * 2016-09-23 2018-03-28 Casale SA Procede de production d'acide nitrique
CN106865568B (zh) * 2017-04-14 2023-08-29 山西兰花科技创业股份有限公司 一种利用分子筛再生气对双甲催化剂升温系统及其工艺
EP3630682A4 (fr) * 2017-05-26 2021-08-11 Starfire Energy Élimination de nh3 gazeux d'un flux de produit de réacteur de nh3
EP4237375A1 (fr) * 2020-10-30 2023-09-06 Casale Sa Contrôle d'une boucle de synthèse d'ammoniac à charge partielle
ES2995034R1 (es) * 2023-07-28 2025-05-13 Sener Ing & Sist Sistema de producción de amoníaco

Family Cites Families (13)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US2795559A (en) * 1954-04-01 1957-06-11 Texas Co Production of hydrogen-nitrogen mixtures
US3081268A (en) * 1961-02-07 1963-03-12 Jr Walton H Marshall Ammonia synthesis gas process
NL136782C (fr) * 1964-08-07
GB1057020A (en) * 1964-12-17 1967-02-01 Petrocarbon Dev Ltd Recovery of hydrogen from a gas mixture
GB1156003A (en) * 1965-10-22 1969-06-25 Braun & Co C F Ammonia Synthesis Gas Purification Process and Apparatus.
US3442613A (en) * 1965-10-22 1969-05-06 Braun & Co C F Hydrocarbon reforming for production of a synthesis gas from which ammonia can be prepared
US3441393A (en) * 1966-01-19 1969-04-29 Pullman Inc Process for the production of hydrogen-rich gas
GB1186939A (en) * 1967-09-08 1970-04-08 Whessoe Ltd Improvements relating to the Production of Ammonia
DE1815532A1 (de) * 1968-12-18 1970-06-25 Linde Ag Verfahren zum Erzeugen von Kaelte
FR2129148A5 (fr) * 1971-03-17 1972-10-27 Air Liquide
US3947551A (en) * 1974-05-29 1976-03-30 The Benfield Corporation Ammonia synthesis
GB1460681A (en) * 1975-02-06 1977-01-06 Petrocarbon Dev Ltd Treatment of ammonia synthesis purge gas
US4172885A (en) 1978-03-20 1979-10-30 Monsanto Company Process for the recovery of hydrogen from ammonia purge gases

Cited By (3)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
EP0106076B1 (fr) * 1982-09-13 1990-02-07 The M. W. Kellogg Company Préparation de gaz de synthèse d'ammoniac
DE102007017403A1 (de) * 2007-04-13 2009-06-10 Lurgi Gmbh Verfahren und Vorrichtung zum Herstellen von NH3-Synthesegas
DE102007017403B4 (de) * 2007-04-13 2012-05-31 Lurgi Gmbh Verfahren und Vorrichtung zum Herstellen von NH3-Synthesegas

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IT1098269B (it) 1985-09-07
EP0000993A1 (fr) 1979-03-07
JPS5460298A (en) 1979-05-15
CA1124032A (fr) 1982-05-25
JPS6058165B2 (ja) 1985-12-18
US4298588A (en) 1981-11-03
IT7826942A0 (it) 1978-08-22
NO158616B (no) 1988-07-04
NO782769L (no) 1979-02-23
NO158616C (no) 1988-10-12
DE2862117D1 (en) 1983-01-13

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