NZ203799A - Catalytic processes for the production of ammonia - Google Patents

Description

No.: Date: 0379 Priority Date(s): VK~ H .TSvt.

Complete Specification Filed: Class: |P<U '< > ■ 1 ■ > Publication Date: .. P.O. Journal, No: 1 3 DEC 1985 NEW ZEALAND PATENTS ACT, 1953 COMPLETE SPECIFICATION AMMONIA PRODUCTION PROCESS {7 We, IMPERIAL CHEMICAL INDUSTRIES PLC a British Company of Imperial Chemical House, Millbank, London SW1P 3JF, England hereby declare the invention for which T/ we pray that a patent may be granted to pW/us, and the method by which it is to be performed, to be particularly described in and by the following statement: - _ 1 _ (followed by page la) I 20379? la. 33 12.2% This invention relates to an ammonia production process, in particular to such a process capable of operation at a relatively low rate of energy consumption per unit quantity of product. 5 A typical amnion i a production process comprises (a) primary catalytically reforming at 3uperatmospheric pressure a hydrocarbon feedstock with, steam to give a gas containing carbon oxides, hydrogen and methane; (b) secondary catalytically reforming the gas from step (a) 10 by introducing air and bringing the mixture towards equilibrium, whereby to produce a gas containing nitrogen, carbon oxides, hydrogen and a. decreased quantity of methane; (c) converting carbon monoxide catalytically with steam to carbon dioxide and hydrogen; (d) . removing carbon oxides to give a nitrogen-hydrogen ammonia synthesis gas; (e) reacting the synthesis gas to produce ammonia and recovering ammonia from the reacted gas; and (f) discarding non-reactive gases present in the synthesis 20 gas.

U.S. Patent Specification No 4,298,588 describes such a process in which energy input is decreased by: (i) operating step (a) in conditions of steam-to—carbon ratio, pressure and temperature to produce a gas 25 containing at least 10% V/v methane and using in step (b) a quantity of air in excess of what would £03739 2 B 53356 introduce 1 molecule of nitrogen per 3 molecules of hydrogen; and (ii) treating synthesis gas after reaction to synthesise ammonia to separate a stream enriched in hydrogen 5 and returning the enriched stream to the synthesis.

In New Zealand Patent Specification No 19 8358 ai process comprising steps (a) to (f) and (i) to (ii) above is described having the features of X. controlling the rate of flow of the stream enriched in hydrogen so that the hydrogen to nitrogen molar ratio of the gas entering the synthesis catalyst is In the range 1.0 to 2.5; and T. operating step (a) in at least one adiabatic catalyst bed and providing the endothermic heat of reaction 15 by preheating, -whereby the temperature of the react ing gas falls as it proceeds through the catalyst bed.

According to the first aspect of the invention an ammonia production process comprises steps (a) to (f) above and is characterised by using in step (b) air enriched with oxygen up to an 20 oxygen content not over 50% T/v The oxygen enrichment is moderate and, except in a process involving carbon oxides removal as methanol, does not introduce less than 1 molecule of nitrogen per 3 molecules of hydrogen. Preferably it introduces an excess of nitrogen, such that after steps (c) and (d) there are 2.0 to 25 2.9, especially 2.2 to 2.8, molecules of hydrogen per molecule of nitrogen.

It has been generally considered that, as a process for producing ammonia synthesis gas, steam hydrocarbon reforming is preferable to hydrocarbon partial oxidation because it does not 30 require an air separation plant and because it converts feedstock to a gas containing a higher proportion of hydrogen to carbon oxides. However, the thermal, efficiency of the hydrocarbon steam reaction heated by internal combustion can be higher than when heated externally and operation at higher pressure is made easier 35 by reacting more hydrocarbon in the secondary reformer (that is, 3 2 037 by internal combustion) and less in the primary reformer. We have now realised that these advantages can be obtained by using in a steam reforming system enriched air produced in an air separation plant much simpler than that used in making substan-5 tially pure oxygen. Various other advantages will be described.

The oxygen content of the enriched air is preferably over 25% V/v for example in the range 30 - 45 > especially 30 -40% V/v. Such a mixture is conveniently the by-product stream from a process designed primarily to produce pure nitrogen. One 10 such process comprises compressing air, cooling it, fractionating it in a single stage to give nitrogen overhead and an oxygen enriched mixture as bottoms and expanding the nitrogen and the mixture to cool the inlet compressed air. The working pressure of such a process is typically in the range 5 - 10 bar abs. The 15 nitrogen can be recovered as liquid or gas and can be used in the ammonia production process in ways described below. The oxygen-containing mixture can be delivered at the working pressure or less, but usually will require compression since secondary reforming is carried out preferably at a pressure higher than that used 20 for the air separation. Suitable processes are described in Kirk-Othmer's Encyclopedia of Chemical Technology, 3rd Edition, vol. 15, pages 935 - 936 and in "High purity nitrogen plants", leaflet FD 38/6, J8a issued by Petrocarbon Developments Ltd.

Another such process ia a selective adsorption 25 system. This can use a zeolite, by which nitrogen is preferentially adsorbed from air, or a carbon adsorbent, by which oxygen is adsorbed more rapidly than nitrogen, a nitrogen stream is passed out and an oxygen-enriched mixture (typically 35% 02> 69% Hg) is obtained by desorption. A process of this type is described by 30 Knoblauch in Chemical Engineering, 6 November 1978, 87 - 89.

Other processes such as membrane diffusion can be used, if desired.

Whoa the oxygen enriched air is produced from such a preferred process the nitrogen product is preferably stored, for 35 injection into the ammonia production plant in start-up, stand-by 4 -B-32256- or shut down phases, in particular for purposes such as a catalyst heating medium during start-up, a flushing or circulating gas during periods of interrupted production when the plant is to be kept hot and ready for rapidly restarting pro-5 duction, a blanketing gas during catalyst shut down or waiting periods and a dilution gas during catalyst activation by reduction,, The combination of such a nitrogen and enriched-air production step with a process involving steps (a) to (f) is thus a technological unity. If nitrogen storage facilities are filled, 10 nitrogen can be vented to atmosphere, possibly with cold recovery in the ammonia recovery section of the production process and/or with changed operation of air separation so as to deliver oxygen-enriched air at increased pressure. Addition of some nitrogen further downstream, for example to the gas after carbon oxides 15 removal or to circulating synthesis gas or purge gas is not excluded, but not usually necessary.

As a result of using oxygen-enriched air at step (b) the following preferred modes of operation become practicable: I. If step (a) employs preheating followed by adiabatic . 20 endothermic reaction, fewer such stages, preferably one only, can be used; II. If the hydrocarbon feedstock contains more than 2 carbon atoms, step (a) can be of the preheat/ adiabatic type known as WCRGW; (Modes I and II are especially preferred because they make the conventional reforming furnace unnecessary).

HI. If step (a) uses externally heated catalyst beds, the outlet methane content can be higher and a higher pressure, lower steam ratio or lower temperature 30 used; 17. The nitrogen excess can be less and thus the size of the side stream from which hydrogen is recovered after synthesis can be less; V. Since the air separation step includes or readily can include removal of water and corrosive impurities, 203799 any compressor used to introduce enriched air in step (t>) need not include a cooler and water separator: thus the isentropic heat of compression is made use of in preheating the enriched 5 air; VI. The synthesis gas generation pressure can be high enough to permit removal of excess nitrogen, or some of-it, before the fresh synthesis gas enters the ammonia synthesis loop; VII. Part of the synthesis gas can be used for methanol synthesis without resulting in an ammonia synthesis gas excessively rich in nitrogen.

In addition, if step (b) includes selective oxidation of carbon monoxide,, this can use the enriched air and then, involves 15 a smaller introduction of nitrogen than when air is used as the oxidant.

The'invention is especially useful when step (a) is carried out by preheating the reactants and then allowing them to react in one or more adiabatic catalyst beds. This arises 20 from I or II above, but also from the possibility of starting -op the whole plant rapidly or shutting it down hot under nitrogen according to the demand for ammonia. This has not been possible using the conventional primary reforming furnace because such a furnace, a brick-lined box, can be heated or cooled only very 25 slowly, and thus it has been preferred to continue operation and put ammonia into store, rather than to cool the furnace. In principle rapid start-up or shut down would be possible using the process of New Zealand Patent Specification No 198358 or the Fluor process (US 3743488 , 3795485)» but has been thought uneconomic in the 30 absence of an on-site nitrogen supply. If desired, the air separation plant can be operated without the other steps in ammonia production, for example during start-up or during a waiting period after stored nitrogen has been used up, to provide the required nitrogen. In such events the oxygen-enriched air is stored or 35 disposed of. 203799 6 D 32206 The pressure at which steps (a) to (d) are operated is preferably at least 10 bar abs. and especially at least 30 bar abs. to make most advantageous use of the oxygen-enriched air.

The upper limit is likely to be 120, conveniently 80 bar abs.

Such pressures apply to the outlet of step (a), and the pressures at subsequent steps are lower as the result of resistance to gas flow in reactors and pipes. In or after step (d), or possibly before a methanation stage forming the last stage of step (d), the gas is compressed if its pressure is not high enough for ammonia 10 synthesis. The extent of such compression is preferably not more than by 20 - 80 bar, and can be by as little as 25% or less such as occurs in a synthesis gas circulating pump. When compression is by such a preferred limited extent, it is preferred to remove excess nitrogen from synthesis gas after ajnmnnia synthesis. 15 The power for the compressors in the air separation plant, for secondary reformer air feed and for synthesis gas, and for various punrps and other machinery, is conveniently derived from engines driven by steam produced in heat recovery from hot gases in the process. If desired such power drives can be partly 20 or wholly electrical.

Step (a) is carried out at an outlet temperature under 800°C, preferably much lower, for example in the range 450 - 650°C. As a result, in an externally heated tubular catalyst system the life of the tubes can be very long, even at the higher pressures 25 specified above. More conveniently step (a) is carried out by preheating followed by adiabatic reaction, in mode I or H mentioned above. A catalyst having adequate low temperature activity should be used, such as a co-precipitated nickel/alumina "CRG" catalyst or a catalyst comprising an active metal on an oxidic secondary sup-30 port on a metal or alloy or highly calcined ceramic primary support as described in New Zealand Patent Specification No.194074 or European published application 45126. Because the secondary reformer, fed with oxygen-enriched air, can react substantial quantities of methane without introducing too great an excess of nitrogen, the outlet 35 methane content of step (a) can be in the range 40 - 80% V/v_on a 2037 7 -3- S3S56- dry basis: in effect step (a) is acting thus as a "chemical preheat er" for the reactants of step (b). The steam to carbon molar ratio in step (a) is typically in the range 2.5 to 8.0, the higher ratios being used at higher pressures in the specified range.

The preheater for the primary reforming step or steps is preferably a pressurised furnace. The pressure of the gases brought into heat exchange with the reactants is suitably at least 5 "bar abs and preferably within 30 bar abs of the pressure of the reactants. By this means the life of the tubes through 10 which the reactants flow in the furnace can be very usefully lengthened and/or the tubes can be made of thinner or cheaper metal. If the heating fluid is combustion gas a useful energy recovery as expansion engine power and waste heat is possible. The heating fluid can be at a pressure dictated by its source, 15 for example it may be helium heated in a nuclear reactor. If desired it can be combustion gas from a solid fuel, especially if combusted in a fluidised bed. Conveniently it is secondary reformer outlet gas.

Whether pressurised or not, the heating fluid after 20 leaving the heat exchange surfaces is passed in heat exchange with one or more fluids such as saturated steam, process air, boiling water, boiler feed water, hydrocarbon feed or combustion air, in decreasing order of grade of heat recovery.

Step (b) is normally carried out in an adiabatic reactor 25 over a refractory-supported Ni or Co catalyst. The primary reformer gas fed to it can be further preheated or can contain added steam or hydrocarbon. The oxygen enriched air is fed at a temperature preferably in the range 400 - 800°C preferably at least partly as the result of using a compressor with limited, 30 if any, cooling. The outlet temperature of step (b) is preferably in the range 800 - 950°C. The outlet methane content is preferably in the range 0.2 to 1.006 on a dry basis but can be greater, for example up to 3«0^, if it is desired not to introduce too much nitrogen and if provision is made to utilise the fuel 35 value of the methane in the non-reactive gas discarded in step (f), 202,19$ 8 -B as in the process of the second aspect of the invention described below or in New Zealand Patent Specification No 198538. ' The gas leaving step (b) can if desired be used as the source of heat for step (a). Very usefully it can be cooled with 5 recovery of heat in ways similar to those applied to preheater heating fluid, except that heat exchange with air is usually avoided for reasons of safety, and that cooling is not below carbon monoxide shift conversion inlet temperature,, In conventional processing this temperature is in the range 300 - 400°C, 10 especially 320 - 350°C, for "high temperature shift", usually over an iron-chrome catalyst. The outlet temperature is typically in the range 400 - 450°C, whereafter the gas is cooled with heat recovery as above to 200 - 240°C and passed to low temperature shift over a copper-containing catalyst at an outlet temperature 15 in the range 240 - 270°C. The final CO content is up to 0.5% V/v on a dry basis, and can be followed by methanation.

As an alternative to such processing the secondary reformer gas can be cooled to 250 - 325°C, with appropriately greater recovery of heat, and passed to shift at an outlet temper-20 ature up to 400°C, especially up to 350°C. This results in a higgler final CO content (up to 2.0% T/v on a dry basis) than low temperature shift, but is preferable at higher pressures (over 30 bar abs) because there is less risk of condensation of steam on the catalyst. The catalyst can be supported copper, suitably 25 with zinc oxide and one or more refractory oxides such as alumina.

If the alternative shift is used it is preferred to remove carbon monoxide finally by cooling and water removal, then selective oxidation. The selective oxidation catalyst is suitably supported platinum (0.01 to 2.0% W/w) containing possibly 30 one or more of manganese, iron, cobalt or nickel as a promoter. A description of a suitable selective oxidation process is given in UE 1555826 and in the articles by Colby et al (23rd Symposium on safety in ammonia plants and related facilities, Am. Inst.

Chem. Engrs. Conv., Miami, November 1978) and Bonacci et al. (im. 35 Inst. Chem. Engrs. Symposium, Denver, August 1977)* 203799 9 B 32256 After low temperature shift or selective oxidation the gas is cooled, water (if still present) is removed from it and the gas is contacted with a regenerable liquid absorbent to remove carbon dioxide. Many processes for doing this are well 5 established and reference is made to our U.S. Patent Specification No 4,298,588 for a survey of them..

The gas now contains a fractional percentage of carbon dioxide and, if produced by low temperature shift, of carbon monoxide. These gases are rendered harmless preferably by methan-10 ation, typically using a supported nickel catalyst at an inlet temperature of 250 - 350°C. The gas is then cooled and dried and then compressed to synthesis pressure. If desired, it can be compressed before methanation.

The conditions of ammonia synthesis can be generally as 15 described in our U.S. Patent Specification No 4,298,588. As a result of the relatively low synthesis pressure, preferably in the range 40 - 120 bar abs, the synthesis catalyst outlet temperature is preferably also low, for example in the range 300 - 450°C, to obtain a more favourable equilibrium. The catalyst volume is 20 typically 100 - 200 m^ per 1000 metric tons per day ammonia output, and is chosen preferably to give an ammonia content of 10 -15% VA in reacted synthesis gas. Recovery of ammonia is preferably by condensation using moderate refrigeration, to for example between +2 and -10°Co Separation of the hydrogen-enriched anr) 25 the methane/nitrogen stream from reacted synthesis gas can be by a cryogenic, adsorptive or diffusion method.

According to the second aspect of the invention an ammonia production process comprises (a) primary catalytically reforming a hydrocarbon feedstock 30 with steam at superatmospheric pressure and in conditions of steam-to-carbon ratio, pressure and temper-•ure to produce a gas containing carbon oxides, hydrogen and at least 10% v/v of methane on a dry basis; (b) secondary catalytically reforming the gas from step (a) 35 by introducing air and bringing the mixture towards 203719 ~B- -j22T)G equilibrium, whereby to produce a gas containing nitrogen, carbon oxides, hydrogen and a decreased quantity of methane, the quantity of air used being in excess of what would introduce 1 molecule of 5 nitrogen per 3 molecules of hydrogen; (c) converting carbon monoxide catalytically with steam to carbon dioxide and hydrogen; (d) removing carbon oxides to give fresh nitrogen-hydrogen ammonia synthesis gas; (e) ~ reacting the synthesis gas to produce ammonia and recovering ammonia from the reacted gas; and (f) treating synthesis gas after reaction to synthesise ammonia to separate a stream enriched in hydrogen, returning the enriched stream to the synthesis and 15 purging the residual stream after separation of the stream enriched in hydrogen; and is characterised by X. controlling the rate of flow of the stream enriched in hydrogen so that the hydrogen to nitrogen molar ratio of the gas 20 entering the synthesis catalyst is in the range 1.0 to 2.5; and T. operating step (a) in an externally heated catalyst and operating step (b) to give an outlet methane content such that in step (f) the purged residual stream contains methane amounting to 5 - 15% by carbon atoms of the hydrocarbon fed to step (a). 25 A process of this type involving synthesis of ammonia using hydrogen deficient , synthesis gas and production of synthesis gas by adiabatic hydrocarbon steam reforming is described in our New Zealand Patent Specification No 198538 mentioned above. Vfe have: now realised that the preferred form of that process in which a 30 small proportion of hydrocarbon is not converted to synthesis gas can be on the basis of an externally heated reforming step operated in advantageous conditions.

The H2 :U2 molar ratio is preferably in the range 1.5 "to 2.3 in the gas entering the synthesis catalyst in step (e). Ttfhat-35 ever its ratio within the defined broad or preferred range, it is -■ E "-shims' v 11 2 037 B 322^6- 9 maintained preferably within 20% of the ratio in the fresh synthesis gas produced in step (d). By this means the rate of flow of the hydrogen recovery stream and thus the power consumption are limited.

The required molar ratio in fresh synthesis gas can be attained without excessive catalyst outlet temperatures provided the steam to carbon ratio in the primary and secondary reforming steps is high enough. To make synthesis gas at a pressure over 30 bar abs, especially in the range 40 - SO bar abs, a 10 steam ratio of at least 3 especially in the range 4 - 8 is preferably used. la the gas leaving step (a) the methane content is preferably in the range 25 - 35% V/v 031 a &ry basis. The methane content of the gas leaving step (b) is preferably in the range 1.5 to 3% on a dry basis. Such methane contents are substantially 15 higher than have previously been considered suitable for ammonia production. They are specified, however, because it is now realised (l) that the plant for removing excess nitrogen from synthesis gas can also remove methane; (2) the methane finally purged is not wasted but is used as fuel in the reformer furnace. As a result 20 the primary reformer outlet temperature need not be over 750°C and can be under 700°C and the secondary outlet temperature need not be over 900°. If desired, step (b) can be fed with oxygen-enriched air, as in the first aspect of the invention. In step (a) the at earn/hydrocarbon reaction can take place at an outlet temperature 25 as low as 550 - 650°C. Consequently a catalyst having adequate low temperature activity should be chosen. £ very suitable catalyst comprises nickel an a refractory secondary support on a metal or alloy primary support, or an a highly calcined ceramic support, as referenced above„ The steps of converting carbon monoxide catalytically with steam and removing carbon oxides can be conventional as described above in relation to the first aspect of the invention. Especially since it is preferred to operate the primary and secondary reforming steps at relatively low temperatures, resulting in a 35 rather higher methane content than was previously considered 203799 12 -B- 32256- suitable for ammonia synthesis gas, it is preferred to remove carbon monoxide finally by selective oxidation.. This leaves carbon dioxide in the gas, and this can be removed largely by contact with the liquid absorbent, as disclosed in our U.S. Patent 5 Specification No 4,298,588. Residual carbon dioxide can then be removed by methanation or adsorption or treatment with non-regenerable alkali.

The conventional sequence of high temperature and low temperature shift can be used, or a single shift stage as des-10 cribed above. In the single stage the inlet steam-to-gas volume ratio can be at least 0.8 which, with suitable temperature control, enables the outlet CO content to be low enough (up to 0.5% V/v on a dry basis) for final removal by methanation. Alternatively the CO content can be up to 2.0% Y/v whereafter it is removed by 15 selective oxidation, as described above.

The conditions of ammonia synthesis can be generally , as in the process of the first aspect of the invention.

In the accompanying drawings two flowsheets are set outs Figure 1, illustrating the first aspect of the 20 invention; and Figure 2, illustrating the second aspect of the invention.

In the process of figure 1, a hydrocarbon feedstock, natural gas, is desulphurised by known means (not shown) and fed 25 at 10 to the lower portion of packed tower 12, in which it rises through a falling stream of hot water fed in at 14 from a source to be described* The resulting water-saturated gas is mixed, if necessary, with steam at 16. (in an alternative process, shown by the dotted line, towers 12 and 38 are not used and all the 30 steam is added as such at 16). The mixture is preheated to 640°C in furnace 18 fired at 20 with natural gas which, for this purpose, need not be thoroughly, if at all, desulphurised. The heated gas is then passed over a supported nickel catalyst in insulated reactor 22. The endothermic methane/steam reaction 35 CH4 + ^0 > 00 + 3% 203799 13 S 322^6- takes place and the temperature falls, reaching 523°C at the catalyst outlet. The resulting gas is then reheated to 700°C in furnace 24! and passed into secondary reformer 30* Here it encounters a steam of hot oxygen enriched air (32% V/v C^j 600°C) 5 derived from air separation plant 27 in which a feed of compressed air 26 is resolved into a substantially pure nitrogen stream sent to storage 28 and enriched air stream which is compressed at 31 and further heated at 32. The use of nitrogen stream 29 will be referred to below. In secondary reformer 30 the temperature 10 rises initially as hydrogen burns with a flame, but over the catalyst further methane/steam reaction takes place and the temperature falls to 924°C at the catalyst bed outlet. The temperature and rate of feed of enriched air are chosen so that the gas leaving 30 contains nitrogen in excess of what can react 15 later with hydrogen to produce ammonia. Furnaces 18 and 24/ can include flue gas heat recoveries such as expansion turbines, combustion air preheaters and boiler feed water heaters but for the sake of clarity these are not shown.

Gas leaving secondary reformer 30 is cooled at 34> which 20 represents heat recovery by high pressure steam generation and one or more of boiler feed water heating and natural gas preheating. The cooled gas, now at about 370°C, is passed into high temperature shift reactor 35 and. there it reacts exothermal ly over an iron chrome catalyst. It is then cooled in heat exchange 36, which 25 usually includes a high pressure boiler and feed water heater, then passed into low temperature shift reactor 37» in which it contacts a copper containing catalyst and its carbon monoxide is almost completely reacted. The shifted gas is cooled with low grade heat recovery at 38 and contacted with water in packed 30 tower 39» in which it becomes cooled and depleted of part of its content of steam. The resulting heated water is passed at 14 in to tower 12 already mentioned. The cool water fed to tower 39 at 40 is derived in part from tower 12 in which heated water from the bottom of tower 39 is cooled by evaporation and partly 35 from supplementary water fed in at 42 from external supplies or ®37 14 B--32256 from point 50 or 58 to be described.

Water-depleted gas leaving tower 39 overhead is passed to cooling, water removal and COg removal units, which are conventional and are indicated generally by item 48• (in the alter-5 native process following the dotted line all the water removal is effected at 48 and stream 50 is larger). At 50 the water contains dissolved carbon dioxide but with simple purification can be fed to point 42. At 52 the carbon dioxide can be expanded in an engine to recover energy. After unit 48 the gas contains residual 10 CO and COg, and these are made harmless by preheating the gas and reacting it over a supported nickel catalyst in methanation reactor 54- The gas is then cooled, largely freed of water in catchpot 56 and thoroughly dried in regenerable adsorption unit 60. Water taken at 58 from catchpot 56 can be used at point 42. 15 The dried gas is compressed at 62, mixed at 64 with recycle gas to be described, heated to synthesis inlet temper^ ature and fed to reactor 66 (this reactor is shown with a single catalyst bed but in practice would include a plurality of beds and conventional means for feed gas preheating and temperature 20 control. It is, however, preferred in any event to have feed gas preheater 67 upstream of part of the catalyst, so that hot gas from the downstream-most bed can pass to external, heat recovery 68 without cooling). After heat recovery 68 the gas is cooled by conventional means (not shown) including moderate re-25 frigeration, to below the dewpoint to ammonia and passed to catchpot 70 from which liquid product ammonia is run off at 72. TJnreacted gas passes out overhead; at this stage is contains less hydrogen per nitrogen molecule than the gas fed to reactor 68, because ammonia formation removes 3 hydrogen molecules per 30 nitrogen molecule, but at 74 it receives a feed of hydrogen-rich gas to be described below. The mixed gas is fed to circulator 76, which increases its pressure by 10 - 20% and is then divided at 78 into a synthesis recycle stream (which is fed to point 64) and a hydrogen recovery stream. This stream is fed to separation 35 section 80. Here it is washed with water to remove ammonia and 2 037 d y&tjg dried. Part of the dried gas is taken off at 81 to regenerate absorber 60, the remainder of the gas is resolved cryogenically or by adsorption or selective diffusion into the hydrogen-rich stream fed to point 74 and a waste stream 86, which may have fuel 5 value. The aqueous ammonia is distilled tinder pressure and the resulting anhydrous ammonia is fed out at 84 to the main product offtake 72.

Table 1 sets out the process conditions, gas compositions and hourly flow rates in a process for making 775 metric tons per 10 day of ammonia from a natural gas of average composition CH^ containing 2.4% V/v of nitrogen and 0.1% V/v of C02* This process follows the dotted paths an the flowsheet.

Daring operation of the process, nitrogen, as liquid or compressed gas, accumulates in reservoir 28. Should the plant 15 have to be shut down, this nitrogen (after evaporation if it is in liquid form) is piped by lines (indicated by 29 generally) to the inlets of catalytic reactors. Such nitrogen can be cold or, if a short shut down period is expected, can be preheated to catalyst operating temperature, nitrogen flow is maintained 20 until process gases have been displaced. If the plant is to be restarted from cold, nitrogen from 29 is preheated and passed through, whereafter burners 20 are lit and the various reactors are brought up to operating temperature. Such nitrogen flow can be on a once through or recycle basis depending on the 25 capacity of reservoir 28. o o TABLE 1 Position Temp °C Pressure bar abs Gas composition % by mole Flow Rate kg mol h ^ CO C02 =2 CH. 4 Ar N2 °2 NH, 3 HgO 22 inlet 640 .6 - - 0.55 24.76* - 0.31 - 74.3 4152.3 22 outlet 523 .0) 64.62 inlet 700 .0J 0.13 2.93 11.53 .50 — ■ 0.30 — — 4378.5 600 40.0 - 0.03 - - 0.79 67.18 32.0 - - 1577.5 outlet 924 32.9*) inlet 370 32.9J 7.84 6.59 33.47 0.43 0.17 14.93 — — 37.07 7184.5 outlet 32.3 37 inlet 210 32.3 1.93 11.98 39.36 0.43 0.17 14o93 — 31.18 7184.5 37 outlet 229 31.9 0.17 13.76 41.13 0.43 0.17 14.93 - - 29.40 7184.5 54 inlet 300 31.9 0.30 0.10 72.29 0.75 0.31 26.24 - - - 4087.7 60 outlet 30o9j 62 outlet 85 J — — 71.92 1.18 0.31 26.59 — — — 4002.2 66 inlet 249 87.7 - - 65.55 6.68 2.07 21.89 - 3.81 - 21609 66 outlet 395") 70 inlet -J 85.6 57.42 7.33 2.27 19.19 13.80 19713 M cr> bl VJI M l\ i V. I C\ sD o o TABLE 1 (continued) Position Temp °C Pressure bar abs Gas composition )& by mols Flow Rate kg mol h ^ CO <U o 0 % CH. 4 Ar W2 °2 NH, 3 ¥> 70 overhead -10 85.6 - ■ - 63.39 8.08 2.50 21.13 - 4.85 • - 17850.0 70 bottoms -10 85.6 - - 0.23 0.14 0.02 0.09 - 99.52 - 1863.2 76 inlet 29 83.5) 80 inlet 36 880 5|. — — 63.34 6„6q 2.07 21.89 — 3.81 — 22712 82 - - 91.43 0,33 0.43 7»82 - - - - - - - - 42.1 86 - - 14.80 22.12 6.08 57.02 - — — 200.7 * At this position but not thereafter the percentage under "CH^" in fact relates to the natural gas hydrocarbon of average composition CH^ 2 ©37 18 %-}22:0 In the process of figure 2 natural gas is desulphurised by known means (not shown) and fed at 110 to the lower portion of packed tower 112, in which it rises through a falling stream of hot water fed in at 114 from a source to be described. The result-5 ing water-saturated gas is mixed, if necessary, with steam at 116. (in an alternative process, shown by the dotted lines, tower 112 and corresponding desaturator tower 138 are not used and all the steam is added as such at 116). The mixture is preheated in the convective section of furnace 118 fired at 120 with synthesis 10 residual gas 86 and natural gas 111, which for this purpose need not be thoroughly, if at all, desulphurised. The heated gas is then passed over a supported nickel catalyst in heated tubes 122. The endothermic methane/steam reaction CH^ + HgO > 00 + 502 takes place, the temperature reaching 629°C at the catalyst outlet. The resulting gas is passed into secondary reformer 130. Here it encounters a stream of hot air (700°C) fed in at 132. The temperature rises initially as hydrogen bums with a flame, but over the catalyst further methane/steam reaction takes place and the 20 temperature falls to 857°C at the catalyst bed outlet. The temperature and irate of feed of air are chosen so that the gas leaving 130 contains nitrogen in excess of what can react later with hydrogen to produce ammonia. It also contains methane to an extent that would normally be regarded as excessive in ammonia 25 synthesis gas: this is preferred because the feedstock economy due to more complete methane reaction would entail extra energy consumption in compressing air and in removing nitrogen later or, alternatively or additionally, would require higher fuel consumption in furnace 118. Furnace 118 includes also flue gas heat 30 recoveries such as combustion air preheaters and boiler feed water heaters but for the sake of clarity these are not shown.

Gas leaving secondary reformer 130 is cooled at 134, which represents heat recovery by high pressure steam generation and one or more of boiler feed water heating and natural gas pre-35 heating. The cooled gas, now at about 300°C, ia passed into 2 03 7 19 B-5225& shift reactor 136 and there it reacts exothermally over a copper-containing catalyst and becomes heated to 335°C. The gas is cooled with heat recovery in boiler 137- It is contacted with water in packed tower 138 and there cooled and depleted of part 5 of its content of steam. The resulting heated water is passed at 114 into tower 112 already mentioned. The cool water fed to tower 138 at 140 is derived in part from tower 112 in which heated water from the bottom of tower 138 is cooled by evaporation and partly from supplementary water fed in at 142 from external 10 supplies or from point 150 or 158 to be described.

Water-depleted gas leaving tower 138 overhead is reacted with air fed at 146 over a noble metal catalyst in selective oxidation unit 144. (in the alternative process following the dotted line item 137 includes also further cooling and water removal). 15 The CO-free gas leaving 144 is passed to cooling, water-removal and COg-removal units, which are conventional and are indicated generally by item 148. At 150 the water contains dissolved carbon dioxide but with simple purification can be fed to point 142. Carbon dioxide passed out at 152 can be expanded in an 20 engine to recover energy. After unit 148 the gas contains residual C02» and this is made harmless by preheating the gas and reacting it over a supported nickel catalyst in methanation reactor 154. The gas is then cooled, largely freed of water in catchpot 156 and thoroughly dried by adsorption in unit 160. 25 Water taken at 158 from catchpot 156 can be used at point 142.

The ammonia synthesis section flowsheet, items 62 - 84, is the same as in figure 1. Stream 86, however, is substantially larger than in the process of figure 1 and is fed to the burners of furnace 118 as a significant part of its fuel supply. 30 Table 2 sets out the process conditions, gas compositions and hourly flow rates in a process for making 1000 metric tons per day of ammonia from a natural gas of average composition CH, QO v v contai.-ni.Tig 2.4% /v of nitrogen and 0.1% /v of C02. This process follows the dotted paths on the flowsheet. The purged residual stream 86 used to fuel furnace 118 contains methane flowing at 037 9 9 -B-32S56 146.4 kg mol h , which is 10% by carbon atoms of the hydrocarbon fed to primary reforming at point 110.

PA/EHG/foP #-^reil 1902 TABLE 2 Position Temp °C Pressure bar abs Gas composition % by mols Flow Rate kg mol h ^ CO C02 *2 CH. 4 Ax N2 °2 NH, 3 ^0 22 inlet 540 51.27 - 0.02 - 16.59* - . 0.38 - - a 3,01 8476 22 outlet 629 50.46 0.52 4.27 17.63 .39 - 0.35 - - 66.85 9275 inlet 700\ 50.46 0.52 4% 27 17.63 .39 0.35 - 66.85 9275 700j 55.22 - 0.03 - - 0.92 78.06 .99 - 2401 outlet 857^ 36 inlet 300} 49.85 3.95 .92 26.89 1.13 0.17 14.89 — — 47.06 12809 36 outlet 335.5 49.25 0.35 9.51 .49 1.13 0.17 14.89 - - 43.46 12809 44 outlet 49.25 - 9.80 .29 1.12 0.18 .44 - - 43ol7 12894 54 inlet 300 49.25 - 0.05 64.20 2.38 0.38 32.76 - - 0.17 6075 60 outlet 48.23") 62 outlet 136 100.72J — — 64.3 2.44 0o 38 32.87 — - — 6054 66 inlet 295 100.00 - - 61.79 3.84 1.57 28.1 - 4.7 - 30157 66 outlet 438 93.02*) 70 inlet - 1,0 90o50) 54.00 4.18 1.71 26.15 13.96 27706 ro ta VM 1 D iW O vn 0 01 VI sD \D o • TABLE 2 (continued) Position Temp °C Pressure bar abs Gas composition % by mols Flow Rate kg mol h~^ CO co2 *2 CH, 4 Ar N2 °2 NH, 3 HgO 70 overhead =rl.08 89.98 - - 58.7 6 4.54 . 1.86 28.4 6 - 6.37 - 25447 70 bottoms -1.08 89.98 •*. - 0.27 0„09 0o02 0.15 - 99.47 - 2258 76 in!et 24.77 89.88^ f27566 80 inlet 38.0 100.72J ' — — 61.16 4.19 1.87 26.89 r .88 — ^3463 82 89.88 - - 90.0 - 2.0 8.0 - - - 2118 84 - - - - - - - - 194 86 .4 4.76 - - 18.85 12.72 1.98 66.44 - — - 1151 * At thia position but not thereafter the percentage under "CH^" in fact relates to the natural gas hydrocarbon of average composition CH, OD. p»oo

Claims (3)

1. 203799 WHAT f/WE CLAIM IS: 23 -B 322^6 1. An ammonia production process which comprises (a) primary catalytically reforming at superatmospheric pressure a hydrocarbon feedstock with steam to give a gas containing carbon oxides, hydrogen and methane; (b) secondary catalytically reforming the gas from step (a) by introducing air and bringing the mixture towards equilibrium, whereby to produce a gas containing nitrogen, carbon oxides, hydrogen and a decreased quantity of methane; (c) converting carbon monoxide catalytically with steam to carbon dioxide and hydrogen; (d) removing carbon oxides to give a nitrogen-hydrogen ammonia synthesis gas; (e) reacting the synthesis gas to produce ammonia and recovering ammonia from the reacted gas; and (f) discarding nan— reactive gases present in the synthesis gas: and is characterised by using in step (b) air enriched with oxygen up to an oxygen content not over 50% T/v.
2. 2. A process according to claim 1 in which enriched air is fed to step (b) at a rate such that it introduces nitrogen corresponding to a hydrogen/nitrogen molar ratio after steps (c) and (d) in the range 2.2 to 2.8.
3. 3. A process according to claim 2 in which excess nitrogen is removed from the synthesis gas after ammonia synthesis'. 4* A process according to any one of the preceding claims in which the enriched air is derived from a preliminary step of air separation into a pure nitrogen stream and a by-product oxygen enriched air stream.;5. A process according to any one of the preceding claims in which step (a) is carried out by preheating the reactants and allowing them to react in one or more adiabatic catalyst beds.;6. A method of operating a process according to claim 5 which comprises, in phases of start-up, stand-by or shut-down,;injecting into the process plant gaseous nitrogen produced by the preliminary air separation step. L."';203799;24 3- 32256;7* An ammonia production process which comprises (a) primary catalytically reforming a hydrocarbon feed stock with steam at superatmospheric pressure and in conditions of ateam-to-carbon ratio, pressure and temperature to produce a gas containing carbon oxides, hydrogen and at least 1096 V/V °£ methane on a dry basis; (b) secondary catalytically reforming the gas from step (a) by introducing air and bringing the mixture towards equilibrium, whereby to produce a gas containing nitrogen, carbon oxides, hydrogen and a decreased quantity of methane, the quantity of air used being in excess of what would introduce 1 molecule of nitrogen per 3 molecules of hydrogen; (c) converting carbon monoxide catalytically with steam to carbon dioxide and hydrogen; (d) removing carbon oxides to give fresh nitrogen-hydrogen ammonia synthesis gas; (e) reacting the synthesis gas to produce ammonia and re covering from the reacted gas; and (f) treating synthesis gas after reaction to synthesise ammonia to separate a stream enriched in hydrogen, returning the enriched stream to the synthesis and purging the residual stream after separation of the stream enriched in hydrogen: and is characterised by X. controlling the rate of flow of the stream enriched in hydrogen so that the hydrogen to nitrogen molar ratio of the gas entering the synthesis catalyst is in the range 1.0 to 2.5; and Y. operating step (a) in an externally heated catalyst and operating step (b) to give an outlet methane content such that in step (f) the purged residual stream contains methane amounting to 5 - 15% by carbon atoms of the hydrocarbon fed to step (a). 8. A process according to claim 7 in which in the gas leaving step (a) the methane content is 25-35% and in the gas leaving -A r- „ „ v • - // ■ . ' - : '' ' 2 037 99 25 5^52254- step (b) the methane content is 1.5 to both v/v an a dry basis. 9. A process according to claim 8 in which step (a) is operated at an outlet temperature in the range 550 - 650°C. 10. A process according to claim 8 or claim 9 in which step (b) is operated at an outlet temperature not over 900°C. IJM. OMED THIS Uj DAY OP A. J. PARK & SOS ,->ER - AGEtlrft FOB THE APFLtCANfK