CA1129626A

CA1129626A - Processes for separating at least one gas from a gaseous feed mixture by using at least two permeator stages in series

Info

Publication number: CA1129626A
Application number: CA323,747A
Authority: CA
Inventors: Tommy E. Graham; Donald L.G. Maclean
Original assignee: Monsanto Co
Current assignee: Monsanto Co
Priority date: 1978-03-20
Filing date: 1979-03-19
Publication date: 1982-08-17
Also published as: DE2953814A1; AU517235B2; FR2420366A1; FR2420366B1; NL7902101A; RO84536B; GB2016945A; PL214236A1; NL182796B; BR7901688A; AU4521879A; RO78306A; PL123383B1; KR840000966B1; DE2910661A1; JPS54130484A; GB2016945B; IT1113311B; JPS6022966B2; IT7921114A0

Abstract

ABSTRACT OF THE DISCLOSURE

Processes are disclosed for separating at least one gas from a gaseous feed mixture containing at least one other gas comprising passing the gaseous feed mixture to at least two permeator stages in series. Each permeator stage contains a separation membrane which has a feed side and a permeate exit side and exhibits selective permeator of the at least one gas as compared to the permeation of the at least other gas of the gaseous mixture. A total pressure differential is maintained across the thickness of the separation membrane in each permeator stage to provide a driving force for the permeation of the at least one gas across the separation membrane. The ratio of total pressure on the feed side to total pressure on the permeate exit side of the separation membrane for at least one permeator stage is least than the ratio of total pressure on the feed side to total pressure on the permeator exit side of the separation membrane for at least one subsequent, i,e., downstream, permeator stage. The at least one gas of the gaseous feed mixture permeates through the membrane, and a permeating gas containing the at least one gas is obtained on the permeator exit side of each of the permeator stages. Between permeator stages, the non-permeating gas from one permeator stage is passed to the feed side of the next permeator stage. In one practical aspect, the process for separating at least one gas is employed to treat an ammonia synthesis purge stream from an ammonia synthesis loop. The separation membrane in the permeators is selective to the permeation of hydrogen as compared to the permeation of inert contaminants such as methane and argon.
A hydrogen-containing permeate is obtained, and the permeate from at least one permeator stage is recycled to the ammonia synthesis zone in the ammonia synthesis loop.

Description

9~2~

This invention relates to processes for separating at least one gas from a gaseous feed mixture containing at least one other gas by selective penmeation t~rough a separation membrane The need to separate at least one gas ~rom a gaseous mi~ture is often encountered in modern so~iety. For instance, the removal of contaminants in waste gas streams may b~ xequired from an environmental standpoint, and if the contaminants are useful, the removal and reco~ery of the contaminants may be economicall~ desirable~ Moreover, the recovery of one or more gases from a gaseous mîxture may be a necessary procedure in chemical processing operations Accordingly, many procedures have been d~veloped to ef~ect gas separations such as selective condensation, adsorption-desorption, absorption-desorption, and the like. One o~
the more recent proposals to eect gas separations is by selective p~rmeation t~rough semi~permeable membranes, i~e~, i separation m~mbranes.
According to current theories, gas separations effect~d by separa~ion membranes may be by several mechanisms One ~` group of such mechanism~ include Knudsen flow, or dif~usion, and the like which invol~e the passage of gases through pores (i.e., continuous ~low channels for gas flow in communication with both the feed and eæit suraces of the membrane~ in the separation membrane. In another postulated mechanism for gas separa~ions, the passage of ~ gas through the membrane may be by interac~ion wi~h the material of the membrane In order to effect the penmeation of a gas through a separation mem~rane, a driving force must be provided Generally, this driving force îs provîded by ' .

L~29~26

-2 07-0403 maintaining a total pressure differential across the thickness of the separation membrane~ Hence, the permeate exit side of the separation membrane is often at a substantially lower pressure than the feed side of the separation membrane. The use of substantial total pressure differentials is especially prevalent in connection with g~s separation operations in which the permeation is by interaction with the material of the separation membrane in order to provide economically attractive fluxes of the permeating gas per unit of available membrane surace area.
If, for instance, the permeating gas is to be discharged to the environment or utilized at low pressure9 e.g., as a burner feed, the use of a substantial total pressure differential across the separation membrane may be wholly acceptable. However, it is oten desired to employ the permeating gas in a chemical process operating at superatmospheric pressure. For example, the gaseous feed mixture to a separation membrane ma~ be an off-stream, e g , a purge stream, from a superatmospheric synthesis process using a cyclic reaction loop such as an ammonia or methanol synthesis process~ At least one of the unreacted reactants in the of~-s~ream may be recovered by permeation through a separation membrane and returned to the super~
abmospheric synthesis process to enhance conversion yields of the process. Thus, compression costs are incurred in returning the permeating gas to the synthesis process.
These compression costs may off-set any savings which may have been realized due ~o the recovery and returning to ~~ the synthesis process of the permeating gas.
Various methods for using a plurality of permeator stages to effect the separation of a gas from a gaseous feed mixture have baen suggested. For instance, United States Patents Nos. 2,617,493 (November 11, 1952) and

3,713,271 (January 30, 1973) disclose cascade-type permeator stages in which the permeating gas ~rom one permea~or stage is passed to the feed side of a su~sequent permeator stage. United States Patent No~ 3,339 7 341 (Septem~er 5, 1967~ discloses in connection with Figure 8 ~129626 two permeator stages in seri~s in which the non-permeating gas from the firs~ permeator stage is passed to the feed side of th~ subsequent permeator stage; however~ the ratio of total pressure at the feed side to total pressure at the permeate exit side of the subsequent permeator stage is disclosed to be lower than that ratio in the first permeator stage. In West German published patent application DT 26 52 432 (May 26, 1977) two permeator stages are disclosed in which ~he non-permeating gas from the irst permeator stage is passed to the feed side of the subsequent permeator stage; however, the total pressure at the feed side of each permeator stage is disclosed to be the same and the total pressure at the permeate exit side of each permeator stage is disclosed to be the same, United States Patent No. 3,836,457 (September 17, 1974) discloses a staged reverse osmosis system for purifying or concentrating aqueous solution~ in which the concentrated aqueous solution is passed ~o t:he feed side of a subsequent reverse osmosis stage and the feed side of the subsequent stage is operated at a higher total pressure than a preceding stage; howe~er, no disclosure is provided pertaining to the separation of gases.
Gardner, et al.~ in "Hollow Fiber Permeator for Separating Gases", emical En~ineering Progress, October, 1977, pages 76 to 78, suggest that one application for separation membranes is in treating an ammonia synthesis purge stream to recover hydrogen. Gardner, et al., do not disclose the use of permeator stages in series.
By this inven~ion processes are provided for separating at least one gas from a gaseous feed mixture containing at leas~ one other gas by selective permeation through a separation membrane in which processes desirable ~mounts of permeating gas can be obtained while requiring a reduced amount of compression to provide the permeating gas at advantageous elevated pressures. In accordance with the processes of this invention a gaseous feed mîxture is passed to at least two permeator stages in series. Each of the permeator stages contain a separation membrane 1 1 2 '~

-4- 07-0403A

having a feed side and a permeate exit side and exhibiting selectivity to the permeability of the at least one gas as compared to the permeabilîty of the at least one other gas. A total pressure differential is maintained across t~ separation membrane to provide the driving force to ef~ect the desired permeation of the at least one gas.
Between permeator stages, the non-permeating gas from the feed side of the separation membrane of on~ permeator staga is passed to the feed side of the separation membrane of the next permeator stage~ The ratio of total pressure on the feed side to total pressure on the permeate exit side of the separation membrane for at least one penmeator stage Chereinafter low total pressure ratio permeator stage~ is less than the ratio of to~al pressure on the lS feed side to total pressure on the permeate exit side of the separation m~mbrane for at least one subsequent, i.e., downstream, permeator stage (hereinafter high total pressure ratio permeator stage2.
In a highly advantageous use, permeator stages in accordance with the invention are utilized to treat a purge stream from an ammonia synthesis loop, Each permeation stage contains a separation membrane which exhi~its selective permeation of hydrogen as compared to the permeation o inert contaminants in the purge stream. Hydrogen which permeates through the separation membrane of at least on~
perme~tor stage can ~e rec~cled to the ammonia synthesis reaction zone, The recover~ from ~he purge stream and recycling of hydrogen to the ammonia synthesis reaction can result in an en~anced conversion of ~ydrogen values to ammonia~ This enhanced co~version of hydrogen values can ~e achieved ~ven ~en process equipment design limîta~ion do not permit an increase in the amount of ammonia produced; however, often increased ammonia production can ~e obtained, T~e at least two permeator stages of this invention provid~ significant advantages in t~at at least one low total pressure ratio permeator ~tage separates the at least one g~s ~rom the gaseous ~e~d mixtur~ whil~ ena~ling the permeating gas from that ~tage to ~e at a desira~le ~ 1~9~20 -5- 07~0403A

total pressure which may require little, if any, recompression for use in a chemical process, Thus, for instance, in an ammonia synthesis process, an enhanced CGnversion of hydrogen values provided by ~his invention can involve little, if any, additional energy consumption over similar ammonia processes which do not employ the process of this invention, and in some instances the energy consumption per unit ammonia produced is decreased.
The non~permeating gas from the at least one low total pressure ratio permeator stage is passed to at least one high total pressure ratio permeator stage in which additional amounts of the at least one gas are separated.
Although the permeating gas from this permeator stage may be at a lower total pressure than that of the permeating gas from the at least one low total pressure ratio permeator stage, the weight amount of permeating gas which requires additional recompression is only a portion of the permeating gases from all of the permeator stages. Thus, less recompression is required than if all of the permeating gases were at the lower total pressure. By the use of the process of this invention, it is possibl~ to enhance recovery of the at least one gas of the gaseous mixture withou~ unduly increasing permeatîng gas recompression costs. Moreover, the total available separating membrane area for a given recovery of the at least one gas is reduced using the processes of this învention in comparison to t~e total available membrane area required for ths given reco~ery of ~he at least one gas if only low total pressure ratio permeators in parallel flow relationship were employed~
~ ccording to current theor~, th~ rate at which a moiet~ permeates through a separatîon membrane is dependent in part on the driv7ng force for that moi~ty~
~ith respect ~o membrane separations in which the moiety is gaseous and passes from a feed gas mixture to a permeating gas on the exit side of the membrane, th~
driving force is the differential in fugacity for that moiety. Generally, fugacities for ideal gases are ~ 129626 -6- 07-0403A

approximated by par~ial pressures and thus, conventionally, in gas separations, the driving force is referred to in terms of partial pressure differen~ials. The partial pressure o~ a moi~y in a gas mixture can be defined as the concentration of the moiety in the gas mixture on a molecular basis times the total pressure of the gas mixture.
Often, the concentration of the moiety on a molecular basis is approximated by the volume concentration of the moie~y.
In view of the effect of the concPntration of the moiety in the gas and the to~al pressure of the gas on the partial pressure, ~hese parameters can be varied jointly or separately to provide suitable partial pressure differentials across the membrane ~o provide desirable fluxes of the moiety. For instance, with the moiety concentrations on the feed side and on the permeate exit side and the total pressure differential across ~he membrane remaining constantJ
but varying the total pressures on the feed and permeate exit sides, a greater partial pressure differential of the moiety is provided at lower total pressures on the feed side and permeate exit side of the membrane, Thu3, in accordance with this invention, the at least one low total pressure ratio permeator stage can be operated such that a suitable partial pressure differential for the at least one gas is maintained across the separation m~mbrane to provide, for ins~ance, a permeating gas contaîning up to about 70 percent of the at least one gas in the gaseous feed mîxture whereîn the permeating gas is at a desirable total pressure for being used in a chemical process without requiring undue rec~mpression. In certain instances, it may be desirable to compress th~ gaseous feed mîxture such that the permeating gas from this permeator stage îs at a total pressure suitable for direct reintroduction into the chemîcal process. In such instances, the gaseous feed stream may of~en be compressed to at least about 20 atmospheres above, say, about 25 to 100 atmospheres above, the original pressure of the gaseous feed stream.
It is clear that th~ non-permeating gas from the low total pressure ratio permeator stage will contain substantial 112g62'o amounts of the at least one gas, for instance, at least about 20 percent of the at least one gas in th~ gaseous feed mixture. While additional amounts of the at least one gas can often be recovered in the low total pressure ratio permeator stage, e.g., by increasing the available separation membrane area, it is preferred that this permeator stage not be operated to maximize its recovery of the at least one gas. Rather, this permeator stage is preferably operated predominantly on a flux-limiting basis.
In a flux-limiting basis operation, the separation is conducted u~der conditions such that when the flux of the at least one gas through the membrane significantly decreases, the separation operation is terminated, e.g., by p~ssing thP non-permeating gas rom the permeator~
Flux-limiting basis operations are in contrast to unwanted permeate~limiting basis operations, In unwanted permeate-limiting basis operations, the separation is continued to provide a suitable recovery of a high proportion of the moiety from the feed mixture without undue permeation of the undesired moieties in the fl_ed mix~ure. Generally1 in any commlercially practical mlembrane separation operation, both flux-limiting basis and unwanted permate-limiting basis considerations will be involved, Often, in a predominantly flux-limiting mode of operation, it is desired that the percent of the difference in partial pressures of the at least one gas (A~ between the gaseous eed mixture CppA feed~ and the non-permeating gas CppA non~permeating) divided by the difference ~etween the partial pressure of the at least one gas in the gaseous feed mixture and the minimum partial pressure of the at least one gas on the permeate Pxit side of the membrane CppA permeate min.~ is up to about 90, say, about 20 or 30 to 90, often a~out 30 to 85. On the other hand, in a predominantly unwanted permeate-limiting basis mode of operation, this relatîonship will often be at least about 85 or ~0 percent.
As stated above, the low total pressure ratio permeator stage is preferabl~ opera~ed on a predominantly 1129~6 flux-limiting basis in order to provide a permeating gas at a desirable total pressure, For a given total pressure differential across the separation membrane and a given separation membrane, a high purge stream flow rate per unit S of available membrane surface area can be employed and a greater amount of the at least one gas permeates the membrane per unit area per unit time than if the permeator stage were operated on an u~wanted permeate-limiting basis.
Generally, sufficient membrane area is provided in the low total pressure ratio permeator stages to permeate at least about 20, preferably about 30 to 70, percent of the at least one gas in the gaseous feed mixture, Since the low total pressure ratio permeator stages are preferably flux limited, particularly desirable separation membranes exhibit high permeabilities for the permeation of the at least one gas, but need not exhibit as high a selectiv.i~y to the pe:nmeability of ~he at least one gas as compared to ~he permeability of the at least one other gas in the gaseous mixture as the selectivity required of a membrane in a predominantly unwanted permeate-limited mode of operation or if the separation were conducted in a single permeator stage to provide the same overall recovery of the at least one gas, The non-penmeating gas from the at least one low total pressure ratio permeator s~age is passed to the feed side of at least one high total pressure ratio permeator stage to recover additional amounts of the at least one gas.
The amount of the at least one gas in ~he permeating gas from this permeator s~age is frequentl~ at least about 10, say, at least about 15 percent of the amount of the at least one gas in the gaseous feed mixture. The amount of the at least one gas in the total permeating gas fro~ all of the permeator stages is preferably at least about 50, e,g., at least about 60, say, a~out 6Q to 95, percent of the at least one gas in the gaseous feed mi~ture.
The at least one high total pressure ratio permeator stage can be operated on a predomînantly flux-limited basis or a predominantl~ unwanted permeate~limited basis.

~.~29626 ~9- 07-0403A

The gas fed to the high total pressure ratio permeator stage can be at any suitable total pressure, For instance, the non-pe~meating gas from the low total pressure ra~io permeator stage can be compressed or decompressed, or can remain at substantially the same pressure depending upon the desired total pressure differential across the separation mem~rane, the total pressure of the permeating gas, and the like. Often, due to strengths obtainable in some suitable separation membranes, the total pressure of the gas fed to the high total pressure ratio permeator stage is decompressed to enable achieving a desirable total pressure differential across the membrane.
The ratio of the total pressure on the feed side to the total pressure on the permeate exit side of the at least one low total pressure ratio permeator stage is less than that ratio for the at least one high total pressure ratio permeator sta~e. Often, the total pressure ra~io of at Least one low total pressure ratio permeator stage is at least about 10 or 15, say, about 15 to 9~, preferably, about 20 to 95, percent less th,an the total pressure ratio of at least one high total pressure ratio permeator stage.
Generally the total pressure drop across at least one high total pressure ratio permeator stage is wQthin about 10 to 500, say, about 15 to 250, percent of the total pressure drop across at least one low total pressure ratio permeator stage. In one aspect of this invention, the total pressure on the parmeate exit side of ~he higher total pressure ratio permeator s~age is at a lower total pressure than the total pressure on the permeate exit side of the lower total pressure ratio permeator stage, An~ suitable number of permeator stages may be employed so long as at least one low total pressure ratio permeator stage and at least one high total pressure ratio permeator stage are provided. Each permeator stage may be comprised of one or more separate permeators wherein plural permeators are arranged in substantially parallel flow relationships.
Preferably, the first permeator stage is a low total press~re ratio permeator stage, Often, the last permeator 11~96~

stage is a high total pressure ratio permeator stage.
Most frequently, two permeator stages are u~ilized, however, in some instances three or more permeator stages mar be desirable. Generally, little ben~fit is achieved in the use of permeator stages above about ive. Preferably, if any permeator stage is operated on a predominantly unwanted permeate-limiting basis, that permeator stage is the last permeator stage.
The effective membrane surface area ci.e., the membrane area a~ailable to effect separation) for each permeator stage should be sufficient to allow a desired amount o~
the at least one gas to permeate. The amount of effective membrane surface area to be employed is influenced ~y, for instance, the permeation rate o~ the at least one gas through the membrane under the separation conditions, i.e., temperature, absolute pressure, total pressure differential across the membrane, and partial pressure dif~eren~ials of the at least one gas across t~e membrane, Advantageous total pressure differentials across separation membranes are at least abouk 10, say, at least about 2~, atmosp~eres 2 and ma~ be up ~o 100 or 200 atmospheres or more. However, the pressure differential shou].d not be so great as to ~nduly stress the membranes suc~ t~at it ruptures or is prone to easily rupturing.
2S A permeator containing the separation membrane may be of any suitable design for gas separations, e.g~, plate and frame, or ha~ing spiral wound film membranes> tu~ular membranes, hollow fi~er membranes, or t~ like. Preferably, the permeator compxises hollow fi~er membranes due to the high membrane surface area per unit volume w~ich can b~
obtained, ~en the membranes are in tubular or hollow fiber form, a plurality of the membranes can be substantially parallelly arranged in bundle form and the gaseous feed mixture can be contacted with either the ou~side (shell side) or the inside (bore side) of the membranes.
Preferably, the gaseous feed mixture is contacted with the shell side of the membranes since passage of the gaseous feed mixture through the bore side of the membranes may 9~26 ~ 07-0403A

involve substantially greater pressure losses. With shell side feed, the shell side effluent from the permeator can often be at less than about 1 or 5, often within less than about 0.5, atmospheres below the pressure of the gaseous feed mixture fed to the permeator and thus be at an advantageous pressure for subsequent processing or energy recovery, e.g., by the use of turbines. Since the concentration of ~he at lea~t one gas on the feed side o~
the membrane is continually diminishing as the at least one gas permeates to the permeate exit side of the membrane which has increasing concentration of the at least one gas, the partial pressure differential of the at least one gas across the membrane is continually changing~ Therefore, flow patterns in the permeator can be utilized to provide desirable recoveries o the at least one gas from the gaseous feed mixture. For instance, the flows of the gaseous feed mixture and the permeating gas can be concurrent or countercurrent~
With bundles of hollow fiber and tubular mem~ranes, the shell side feed can be radîal, i.e , the feed stream transversely flows past the membrane either to the insid~
or, usually the outside of the bundle, or the flow can be axial, i.e., the feed stream disperses within the bundle and generally flows in the direction in which the hollo~
fibers or tubular m~mbranes are oriented Any suitable material ma~ be emplo~ed for t~e separation m~m~rane as is well-known in the art Typical membrane materials include organic polymers or organic polymer mixed with inorganics, e,g~, fillers, reinforcements, and the like. Metallic and metal-containing membranes may also be used.
The ammonia synthesis process is descri~ed in more detail in order that this aspect invention can be full~
appreciated~ Ammonia is synthesized b~ the catalytic reac~ion of hydrogen and nitrogen~ The hydrogen feedstock for the ammonia synthesis is generally obtained from prîmary reforming of hydrocarbon, etg,, natural gas The e~fluent from the primary reforming thus contains impuriti~s such as methane, carbon oxides, i~e , car~on dioxide and car~on 1 .1.296~5 -12- 07-0403A

monoxide, wa~er and the like. Current practîce provides for the removal of impurities from the reformer effluent which may be harmful to the ammonia synthesis catalyst such as the carbon oxides, sulur compounds and the like; however, impurities such as methane are generally not completely removed from the reformer effluent since they are not directly harmful to the ammonia synthesis reaction and are expensive to remove. The nitrogen feedstock is usually obtained rom air with the removal of oxyg~n, e.g,, b~
combustion wlth fuel to produce water or carbon dioxide and water, followed by removal of the water and carbon di~xide, if present, or by liquifaction. The resultant nîtrogen stream contains minor amounts of ~mpurities such as argon which are present in small amounts in air, Since they are not directly harmful to the ammonia synthesis reaction, these impurities are not generalLy removed from the nitrogen feedstock due to economic considerations, Thus, even though the predominant components of the synthesis eed gas are hydrogen and nitrogen, at least on~ of methane and argon are present as contaminants in the synthesis feed gas, Me~hane is often present in amounts of up to about 5, e,g,, about 0.1 to 3, volume percent, and argon is often present in amounts of up to about 0,5, e,g., about 0,1 to 0,5, mos~
often about 0,3, volume per~ent based on the synthesis feed gas. Othe~ contaminan~s which may be present include water and helium.
The ratio of hydrogen to nitrogen which is present in the synthesis feed gas is preferabl~ such that the mole ratio of hydrogen to nitrogen of the reaction gas introdueed into the ammonia synthesis reaction zone is su~stantially constant to prevent a build-up of either ~ydrogen or nitrogen in the ammonia synthesis loop. However, the mole ratio of hydrogen to nitrogen in the reaction gas may ~e greater or less than the stoichiometric ratio such that the excess of hydrogen or nitrogen over that required for the reaction to ammonia on a stoichiometric basis shifts the equilibrium in favor of ammonia production. In such situations, the mole ratio of hydrogen to nitrogen may ~e from about 2 or 2~5:1 to 9~26 about 3.5 or 4:1. Higher or lower mole ratios could be employed; however, since a purge stream must be removed from the synthesis loop to prevent undue build-up of contaminants, considerable increases in the loss of valuable nitrogen or hydrogen would be incurred. The processes of this invention do minimize the increases in loss of hydrogen through the purge stream when the reaction gas has a greater than 3:1 mole ratio of hydrogen to nitrogen because of the recovery and recycling of hydrogen from the purge stream.
Generally, the mole ratio of hydrogen to nitrogen in the reaction gas is about 2.8:1 to 3.5:1, say, about 2.9:1 to 3.3:1. Frequently, the mole ratio of hydrogen to nitrogen in the reaction gas introduced into the ammonia synthesis reaction zone is substantially that mole ratio required for the reaction of hydrogen and nitrogen on a stoichiometric basis, e.g., about 2~95:1 to 3.05:1. Generally, nitrogen does not permeate the membrane to a significan-t extent, and the permeating gas contains little, if any, nitrogen. How-ever, any nitrogen which is recovered and recycled in the permeating gas represents a savings with respect to the nitro-gen Eeedstock demands. The mole ratio of hydrogen to nitro-gen in the synthesis feed yas is thus usua:Lly slightly less than the mole ratio of hydrogen to nitrogen in the reaction gas such that desirable hydrogen to nitrogen ratios are provided when combined with the permeating gas which is re-covered from the purge stream. In typical ammonia plants in accordance with this invention, the mole ratio of hydrogen ~~ to nitrogen in the synthesis feed gas may be about 2.7:1 to 3.2:1, say, about 2.3:1 to 3.0:1.
The reaction between hydrogen and nitrogen to produce ammonia is exothermic and is an equilibrium reaction. The ammonia synthesis may be conducted using any suitable pro-cedure such as the Haber-Bosch, modified Haber-Bosch, Fauser and Mont Cenis systems. See, the Encyclopedia of Chemical Technology, Second Edition, Volume 2, pages 258, et seq., for 1~2962~`
- 13a -various of the processes for synthesizing ammonia from hydrogen and nitrogen. In general, these processes employ super-atmospheric ammonia synthesis pressures of at least about ~I

6 ~ ~

100 atmospheres absolute and promoted iron synthesis catalysts. The ammonia synthesis reaction zone is generally cooled to maintain reaction temperatures of about 150 or 200 to 600C. The use of high synthesis pressures shifts the equilibrium in favor of the formation o ammonia.
Although some ammonia synthesis pressures which have been employed are as high as 500 or more atmospheres absolute, most present day ammonia plants utilize synthesis pressures of about 100 to 300 or 350 atmospheres absolute, especially about 125 to 275 atmospheres absolute, Typically, the ammonia synthesis feed gas is compressed in at least two stages in order to facilitate achieving synthesîs pressures.
Generally, the pressure of the feed gas prior to at least one compression stage is within at least about 100, say, within about 10 or 20, atmospheres below the synthesis pressure. The lowest pressure in the ammonia synthesis loop is preferably within about 5 or lO atmospheres below the synthesis pressure. A recycle compressor is generally emplo~ed to circulate the gases in the synthesis loop and to main~ain the desired synthesis pressure in the ammonia synthesis reaction zone.
The conversion to ammonia based on hydrogen entering the ammonia syn~hesis reaction zon~ is often a~out 5 to 30, e~g., about 8 to 20 percent~ In many commercial plants, the ammonia concentration of the reaction effluent exiting the ammonia synthesis reaction zone is about lO to 25, e.g., about 10 to 15 or 20, volume percent~ Thus the reaction effluent from the ammonia synthesis reaction zone contains substantial amounts of hydrogen and nitrogen. Accordingly, ammonia is condensed from the reaction effluent, and the reaction effluent containing the valuable hydrogen is recycled in an ammonia synthesis loop to the ammonia synthesis reaction zone to provide an attracti~e convQrsion of hydrogen in the feed to ammonia. Frequently, the reactor feed gas fed to the ammonia synthesis reaction zone contains about 0.5 to 5, say, about 1 to 4, volume percent ammonia and less than about 25 volume percent inert contaminants, say, about 4 to 15 volume percent inert contaminants :1~12962~

T~us, th~ reactor feed gas may comprise about 2 to lS
volume percent methane, about 2 to lO volume percent argon, and helium, if present in the reformer feed, e.g., in an amount of about 0.1 to 5 volume percent.
Ammonia in the reaction effluent from the ammonia synt~esis reaction zone is removed from the synthesis loop.
A preferred method for removing the ammonia is by chilling the ammonia-containing reaction effluent to coalesce ammonia which can be removed as a liquid product After removal of the ammonia the gas în the synthesis loop still may contain ammonia, e.g,, up to about S volume percent ammonia. The coalescing of ammonia from the gas in the ammonia synthesis loop is preferably conducted subsequent to the recycle compression. Two or more ammonia coalescers may be employed in the synthesis loop to enhance ammonia recovery.
The compressed synthesis feed gas may be introduced into the ammonia synthesis loop at any suitable location, e.g,, before or after the recycle compressor, and before or after the ammonia removal, In many instances, however, it is preerred to introduce the compressed synthesis feed gas into the ammonia synthesis loop prior to coalescîng ammonia since the coalescing can remove water vapor and thus ensure that the reaction gas fed to the ammonia synthesis reaction zone h~s a low oxygen-containing compound content to prevent catalyst poisoning.
Difficulties occur in that the inert contaminants such as methane, argon, etc., in t~e hydrogen and nitrogen feedstocks do not take part in the ammonia synthesis reaction and must be removed from the ammonia synthesis loop in an amount suficient ~o prevent an undue build-up of these inert contaminants in the ammonia synthesis loop.
Conveniently, the removal of these inert contaminants is effected by removing a purge stream Erom the ammonia synthesis loop. The purge stream will contain the same concentration of hydrogen and nitrogen as the recycling reaction effluent. Hence~ recovery o the valuable hydrogen from the purge stream for return to the ammonia synthesis catalyst zone may be highly desirable, Frequently, the ~.29~
-16~ 07~0403A

reactor feed gas contains less than about 25, say, about 4 to 15, volume percent inert contaminants. The purge stream often comprises up to about 3, say, about 0.5 to 2.5, volume percent of the gases in the synthesis loop at the point from which the purge is taken. The purge stream may, o course, be a greater portion of the gases in the synthesis loop;
however, such large purge amounts result in increases in the weight amounts of nitrogen and, possibly, hydrogen exhausted from the ammonia synthesis system, The volume of the purge stream is usually sufficient to maintain the concentrations of methane and argon substantially constant.
It is generally preferred to remove the purge stream from the gases in the ammonia synthesis loop upstream of the introduction of the compressed synthesis ~eed gas to prevent purging the fresh hydrogen and nitrogen feed. The purge stream may be removed from the synthesis loop upstream of the ammonia removal, or the purge stream may be removed ~rom the synthesis loop downstream of the ammonia removal from the synthesis loop. Usually the gases in the ammonia synthesis loop downstream from the ammonia removal contain reduced, but still significant, amounts of ammonia.
In the case in which the purge stream is removed rom the synthesis loop upstream of the ammonia removal, the ammonia concentration in the purge stream is often at least about 5 volume percent, say, up to about 30, e.g., about 8 to 25, or e~en 10 to 15 or 20, volume percent. Convenlently, the purge stream is chilled to coalesce ammonia, and the separated liquid ammonia can provide additional ammonia product The purge stre~m still contains sîgnifîcant amounts of ammonia, e.g., often at least about 0 5 or 1 volume psrcent ammonia. This procedure is par~icularly desirable w~en modifying existing ammonia synthesis plants to produce ammonia in accordance with this in~ention since existing ammonia synthesis plants generally emplo~ an ammonia coalescer to remove ammonia from the purge stream. The amount of ammonia in the purge stream ma~ be further reduced b~ scrubbing with water or by diffusion of the ammonia through the separation membrane, ~hus, the ammonia .L 1 ~ 3 ~ ~ S

concentration of the non-permeating gas from the last permeator stage may be sufficîently low that it is suitable for use as, e.g., fuel,or can be vented to the environment, especially after recovering energy provided by the higher pressure of the purge stream. On the other hand, the purge stre~m may ~e passed to the permeator stages without removal of ammonia, or ammonia can be removed from the non-permeating gas between permeator stages, e.g., by chilling and coalescing and/or b~ water scrub~ing.
In the case in which the purge stream is removed from the synthesis loop downstream of the ammonia r~moval, the ammonia concentration in the purge stream is often at least about 0 5 up to about 5 volume percent. Tn view of the low ammonia concentration in the purge stream, removal of ammonia from the purge stream prior to contacting the separation membrane sometimes is not done. Additional ammonia is recovered from the purge stream by permeation through the separation membrane, and the non-permeating gas from the last permeator stage may be suitable for use as, e,g., fuel or can be vented to the atmosphere, especially after recovering energy provided by the high pressure of the purge stseam, The purge stre~m may, if necessary, be subjected to heat exchange to pro~ide suitable temperatures for efIecting hydrogen separation by the use of separation m~mbranes Often, the purge stream to ~e contacted with t~e separation membran~ o~ a permeator stage is at least a~out 10C, say?
about 15 to 50C, preferably, about 25 to 40C, Higher temperatures may be employed dPpending upon the physical stability and thè selectivity of separation of t~e mem~rane at the higher temperatures.
The purge stream is contacted with a separation membrane which exhibits selectivit~ to the permPat;on o~
hydrogen as compared to the permeation of each of m~t~ane and argon. In viPw of the generally substantîally lower volume concen~rations of methane and argon in the purge stream as compared to the volume concentration o~ hydrogen in the purge stream, suitable separation mem~ranes need -11~96~

not exhibit high selectivit~ of separation of hydrogen from each of methane and argon in order to provide an enhanced ammonia synthesis process~ Generally, the selectivity of separation of a membrane is described in terms of the ratio of the permeability of the fast permeating gas Chydrogen) to the permeability of the slow permeating gas (methan~ or argon~ whereîn the permeability of the gas through the mem~rane can be defined as t~e volume of gas, standard temperature and pressure, which passes through a membrane per square centimeter of surface area, per second, for a partial pressure drop of 1 centimeter of mercur~ across the membrane, This ratio is referred to as a separation factor for the membrane, For sake of uniformit~, the permeabîlities and separation factors mentioned hRrein are determined a~
a~out 25C and a pressure drop o a~out 3.4 atmospheres across the membrane with the feed side of the membrane ~eing about 3.4 atmospheres absolute unless otherwise indicated.
Often, the separation factor of the m~mbrane for the separation o~ ~lydrogen from methane is at least about l Separation ~actors ~or hydrogen over methane of 100 or greater ma~ be provided by certain membranes; however, lit~le advantage ma~ be ohtained us~ng such highl~ selecti~e mem~ranes, Often the membrane may be sPlected on its ability to quirkly permeate hydrogen rather than on its selectivity of separation. Consequently, membranes exhibiting a separation factor for hydrogen over methane of about 10 to 80 are adequate. Clearly, the higher the permeability of hydrogen through a membrane, the less available membrane surface area which is required to pass a desired amount of hydrogen through the membrane, Particularly desirable membranes exhibit hydrogen permeabilities of at least abou~ 1 x 10-6, preferably at least about 20 x 10-6, cubic centimeters of hydrogen per square centimeters of m~mbrane surface area per second at a partial pressure drop of 1 centimeter of mercury across the membrane.

1129~6 The volume ratio of the permeating to non-permeating gases from each of the permeator stages as well as the composition of each of the permeating and non-permeating gases which may be employed in accordance with the method o~ this invention can be ~aried over a wide range. By way of illustration, Table I provides typical approximate concentrations of the significant components in the gases passed to a low total pressure ratio permeation stage and to a high total pressure ratio permeator stage and those in the permeating and non-permeating gases from each stage.
The permeating gas from each of the permeator stages contains valuable hydrogen and can be recycled such that the hydrogen can be utilized in the ammonia synthesis. In accordanc.e with the processes of this invention the total pressure under which the permeating gas exits each permeator stage is taken advantage of, for instance, by returning the permeating gas to the synthesis feed gas at a point where the permeating gas is at substantially the same pressure as the s~nthesis gas. Thus recompression costs are minimized.
The stream into which a permeating gas is introduced can be selected partially on the basis of the operating pressure differentials across t:he separation membrane which can be employed. Since the compression of the synthesis feed gas is condu~ted in severaLl stages, or steps, some limitation exists as to the pressure differentials across the membrane which are available in a given ammonia synthesis system, especially in ammonia synthesis systems which are retrofitted with separation membrane hydrogen reco~ery sy~tems to enable conducting the ammonia synthesis processPs of this ~nvention. The processes of ~his invention, however, are sufficiently flexible, since ~he separation of h~drogen is conducted in a plurality of permeator stages, that a permeating gas stream having a desirable total pressure can be provided. In general, the pressure differential for a given ammonia synthesis system in a~cordance with this invention is selected to pro~ide the largest operating total pressure differential across the separation membrane (within the range of suîtable operating 9 6 ,~ ~ -20- 07-0403A

~_ _ _ ~1 --D I O O ~1 o u~ o Lf~ o S l I Lf~ U) ~1 ~) C~ ~ ~C) ~ ~ C~l ~ ~ I O ~ C" O O o ~ U~ o t ~1 ~
~ 1:4 l ~V
Ei _ I .
D ~1 l I O O C~l O O o o U) l ~1 I ~ U'l ~ ~ LO `;t 1~ C~
i~ a) i I I I I i a) l l l l l O ~:: I ~ o o C~ o t~l O O C`~ I` O
Z ID l c~ l r~
U~ U~ __ _ r ~ h ID ~ I O ~ c~ l Ifl ~ ~1 J~ ~1 I~D
U~ ~1 l ~ C5~ ~ -I ~ ~r a~ I I~ O ~v I v I ~v o Q~ I
~1 l D O O c.~ c~ 0 1:4 u~ o C~ C) O
O C~ O I P~ o~ d `D td t~ t~
S~l l h ~I h ~: ~ P-~ I ,1 E~ E~
D ~ ~ i - - ~
D ID r-l ¦ P: t~ O ~ ID ~3 0 C`l o El td I a~ ~r~l 0 ~1 ~1 ~1 a) ~1 ~v I ~ ~ ~ O
E~ l4 ~ tq ~ D~D C~ ~q ~ ~ ~D a) 1--1 1--l ~D tQ r~l U ID I ~d td td t~
I:'l O 1-1 O-rl ~ 4 O-rl P _ ~d J~
~D O O
~ E~a~ ~ ~ o o c~l o h ~ c~ ~1 u~ ,~ u~
ID O1~ ~ ~ ~D o rl O O ~ C~l O
t~ ~ ~ U~ C ~ 1 ~iv :~:
O
C~ ~ o U~ o o It~~1 1~ C~
~ O ~ ~ ~ O O O ~ U~ O
L.~ ~

a) Iv ~D ~ ~D Q~ ~
O bo~ ~oO ~ ~ ~ bo~
h S~ C7 ~ rl 5~ h O ~ r~
3 ra _ :~ z ¢ ~ ~D~ z 1 ~9~ 21- 07-0403A

pressure differentials for a given separation membrane) which provides a permeating gas at a suitable pressure for introduction into a synthesis feed gas stream. Usually, the permeating gas is at a slightly higher pressure than the gas stream into which it is introduced, e.g., abou~ 0.1 to 5 atmosph~res higher. A reduction in the total pressure of the permeating gas or, preferably, the gas passing the permeator stage should only be employed when no suitable gas stream is available which is at a pressure which permits a suitable total pressure differential across the membrane to be achieved, e.g., the available synthesis feed gas streams are at pressures too high to permit a desired hydrogen flux through t~e separation membrane or at pressures so low that the separation membrane can not physically withstand the pressure differential. Alternatively, a desirable total pressure differential may be provided across the separation membrane and the permeating gas compressed to suitable pressures for introduction into>
e.g., the synthesis feed gas stream or synthesis loop, or the puxge stream may be compressed prior to contacting the separation membrane to provide a desired pressure diferential across the separation membrane and a permeating gas at a suitable pressure for reintroduction into the ammonia synthesis system.
The non-pe~meating gas from the last permeator stage may be utilized in any suitable manner, e.g., used as fuel.
Since the non-permeating gas is at high pressure, significant energy can be recovered from this gas by, for instance, the use of a ~urbine, and the like.
The invention will be further described with reference to the drawings in which:

FIGURE 1 is a simplified schematic flow diagram of an ammonia synthesis plant having two permeator stages in accordance with ~his invention wherein the permeating gas rom the first permeator stage is combined with the synthesis eed gas between compression stages for recycling to the ammonia synthesis reaction zone and the permeating 11~9626 -22- 07-Q403A

gas from the second permeator stage is com~ined with the synthesis feed gas prior to its compression to the super~
atmospheric synthesis pressure.

FIGURE 2 is a simplified schematic flow diagram of an ammonia synthesis plant having two permeator stages and is similar to the plant depicted in Figure 1 except that the permeating gas from the second permeator stage is compressed and combined with the permeating gas to ~e recycled to the ammonia synthesis reaction zone~

FIGURE 3 is a simplified schematic flow diagram of an ammonia synthesis plant having tWQ perme~tor stages in accordance with this invention wherein ammonia is removed from the ammonia synthesis loop prior to the removal o~ the purge stream, The permeating gas from the first permeator stage contains ammonia and is combined with dry synthesîs feed gas between compressor stages, The non-permeating gas from the first permeator stage is passed through a scru~ber to remove ammonia and then to the second permeator stage rom which permeator s~age the permeating gas is combined with the synthesis feed gas pri.or to its compression to the superatmospheric synthesis pre~,sure, FIGU~E 4 is a simpli~ied schematic flow diagram of an ammonia synthesis plant having two permeator stages and is similar to the plan~ depicted in Figur~ 3 except that ~he purge stream is compressed prior to ~eing passed to the first permeator stage and the permeating gas from th~ ~irst permeator stage is directly introduced into the ammonîa synthesis loop, FIGURE 5 is a schematic cross-~ection of a holLow fiber mem~rane-containing permeator which may ~e employed in an ammonia synthesis plant in accordance with this invention.

11'~9626 In Figures 1 and 2 and 3 and 4, like reference numerals indicate like features.
With reference to Figure 1, a synthesis feed gas comprising hydrogen and nitrogen in approximately a 3:1 mole ratio is introduced via line 10 into the ammonia synthesis system, The synthesis feed gas is compressed in several stages to superatmospherîc ammonia synthesis pressures. As depicted, compressor 12 partiall~ elevates the pressure of the synthesis feed gas toward ~e super~
atmospheric synthesis pressure. The partially compressed feed gas is chilled in coalescer 14 to condense and remove (:~ia line 15~ water vapor from the synt~esis feed gas. The thusly dried synthesis feed gas is passed to compressor 16 where it is compressed to above the superatmospheric pressure in the ammonia synthesis loop, It is ~o ~e realized that one or both of compressors 12 and 16 may ~e comprîsed of two or more compressor stages, The effluent from compressor 16 is transported via conduit 18 to the ammonia synthesis loop. The conventional synthesis loop is depicted in t:hat the s~n~hesis feed ga~
passes through rec~cle compressor 20 and ammonia coalescer 22 from which product a~onia is withdrawn via line 23 before entering ammonia synthesis reaction zone 24 for con~ersion to ammonia. Positioning coalescer 22 prior to the ammonia synthesis reaction zone 24 insures that any water vapor which may be present in the reactor feed gas is reduced to provide less than about 10 ppmv total oxygen-containing compounds in the reactor feed gas, The reaction ef~luent from the ammonia synthesis reactîon zone is cooled in heat exchanger 26 to a t~mperature in t~e range of a~out 0 to 100C, The heat ~ransfer medium in heat exc~anger 26 ma~
~e the reactor feed gas from ammonia coalescer 22 which in turn is heated to a suitable temperature for introduction into the ammonia synthesis reaction zone. T~e efluent from heat exchanger 26 is recycled via line 28 Cammonia synthesis loop~ to the recycle compressor 20, A purge stream is withdrawn rom line 28 via line 3Q, The volume of the purge s~ream is sufficient to main~ain an 11~9~26 acceptable level of inert contaminants in the ammonia synthesis loop and ammonia synthesis reaction zone. As depicted, the purge stream is removed prior to the condensation of the ammonia product from the gases in the ammonia synthesis loop. Thus, the purge stream contains substantial quantities of ammonia. As depicted in Figure 1 ammonia is removed from the purga stream by passing the purge stream throug~ chiller-coalescer 32 which removes ~mmonia by condensation followed by water scrubber 33 which absorbs ammonia and generally provides a gas containing less than about 0.1 volume percent ammonia~ The purge stream may, if necessary, be subjected to heat exchange to provide a temperature of, say, about 25 ~o 40C. The purge stream is passed via line 34 to first permeator 36.
A schematic cross-section of an axially, shell-side fed permeator such as may be employed in the system of Figure 1 i9 provided in Figure 5. With reference to Figure 5, within casing 100 is positioned a plurality of hollow fiber membranes which are arranged in bundle generally designated by the numeral 102, One end of the bundle is emBedded in header 104 such tha~ the bores of the hollow fibers communicate through the header. The header is positioned in casing 100 such that essentially the only fluid communication t~rough the header is through the bores of the hollow fi~ers. The opposite ends of the hollow fibers are sealed in end seal 106, Th~ purge stream enters the casing through feed port 108, disperses within bundle 102 and passes to shell exît port 110 positioned at the opposite end of the casing~ ~ydrogen permeates to the bores of the hollow fibers, and passes via t~e bores through header 104, The permeating gas exits casing 100 through permeate exit port 112. While Figure 5 depicts a hollow fiber membrane~containing permeator in which onl~ one end of the holl~w fibers is open, it is apparent that both ends 35 of the hollow fi~ers can be open.
~ ith reference to Figure 1, a first permeating gas~
i,e,, a hydrogen-rich stream, exits permeator 36 via line 38~ Th2 pressure drop across the membrane is such that the ll~9S~

permeating gas is at a pressure substantially the same as the pressure of the synthesis feed gas exiting compressor 12, and the first permeating gas is combined with the synthesis feed gas exiting compressor 12 in order to be recycled to ammonia synthesis reaction zone 24, The first permeating gas is introduced into the synthesis feed gas upstream of coalescer 14 such that water vapor which is introduced into the purge stream in scrubber 33 and permeated through the separation mem~rane, can be removed The non-permeating gas is withdrawn from the feed side of first permeator 36 and is passed via line 40 to permeator 42. The non-permeating gas contains hydrogen as well as nitrogen, methane and argon. A second permeating gas exits second permeator 42 via line 44. The second permeating gas is at a pressure substantially the same as the pressure of the synthesis feed gas entering compressor 12, and the second permeating gas is combined with the synthesis eed ~as entering compressor 12 to be recycled to ammonia synthesis reaction zone 24. The non-permeating gas from t~e second permeator exits via line 46 and can be treated in an additional permeator (not depicted), exhausted to the environment, or used, for instance, as a fuel.
The ammonia synthesis syst~m of Figure 2 îs substantially the same as the system depicted in Figure 1 except t~at the second permeating gas from second permeator 42 is compressed in compressor 48 to a total pressure slightly above the pressure of the first permeatîng gas in line 38. The compressed second penmeating gas is passed via line 50 to line 38 whereat it is combined wit~ the first permeating gas ~eîng recycled to ammonia synthes~s reaction zone 24~ This method may find applicatîon when retrofitting separation membrane ~ydrogen recovery systems in exîsting ammonia plants in order to utilize a process of this invention~ For instance, îf încreased ammonia production is desired in an ammonia plant, ~ut compressor 12 is at its maximum capacity, thîs ~ottleneck can be obvîated ~y utilizing a compressor to increase the pressure of the second permeatin~ gas such that it can be introduced into the synt~esis feed gas 1~.2962~
-26- 07-0403~

without increasing the load through compressor 12, Also, the difference between the pressure on the feed side of second permeator 42 and the feed side of compressor 12, in some ammonia plants, may bP too great to be withstood by a membrane which may be employed. Accordingly, a lesser pressure drop can be utilized across the membrane in the second permeator while maintaining the feed side o~ the second permeator at elevated pressures~ In view of the use of high pressures on the feed side of the second permeator, little compression is required to elevate the pressure of the second permeating gas for introduction into the ammonia synthesis system.
In the ammonia synthesis system depicted in Figure 3 the purge stream is withdrawn from the ammon~a synthesis loop downstream of the ammonia removal and upstream of the introduction of the fresh synthesis ~eed gas into the synthesis loop, A synthesis feed gas comprising hydrogen and nitrogen is introduced via line 200 into the ammonia synthesis sys~em. The synthesis feed gas contains moisture and is thereore Eed in~.o adsorber 202 in which essentially all the water contained in ~he synthesis feed gas is removed such that t~e total oxygen-containing compound content of the synthesis feed gas is less than about lO ppmv. The thusly dried gas is transported through ~he line 204 to compressor 206 in which the synthesis feed gas is partially co~pressed to su~stantially the superatmospheric synthesis pressure, The partiall~ compressed synthesis feed gas is compressed to a~ove the superatmosp~eric pressure in the ammonia synthesis loop in compressor 208, Each of compressors 206 and 208 can be a multistage compressor, The effluent from compressor 208 is passed via line 210 into the ammonia synthesis loop where the synthesis feed gas is com~ined with the gas circulating in the synthesis loop to provide the reac~or feed gas. The reactor feed gas is compressed in recycle compressor 2127 heated in heat exchanger 214, and introduced into ammonla synthesis reactîon æone 216, reaction effluent from the ammonia synthesis reaction zone -11~96~

is used as the ex~hange medium in heat exchanger 214 and is cooled. The reaction effluent then passed to ammonia coalescer 218 from which product ammonia is withdrawn via line 220. Since the fresh synt~esis f~ed gas has not been combined with the reaction effluent, a lesser weight of gas needs to be refrigerated to condense the ammonia. The overhead from ammonia coalescer 218 is recirculated in conduit 222 (ammonia synthesis loop) to ammonia synthesis reaction zone 216. Since the ammonia is removed rom the synthesis loop prior to recompression in recycle compressor 212, less energy is expended in recirculating the gas in the synt~esis loop.
A purge stream is withdrawn from the synthesis loop via line 224, Since the purge stream is at a low temperature because o~ the cooling to condense the ammonia product, the purge stream is heated in heat exchanger 226 to suitable temperatures for efecting the separation of hydrogen and ammonia, e.g,, about 25 to 40C. The warmed purge stre~m is passed to irst permeator 228.
First permeator 228 may be of any suitable design including the design of the separator depicted in Figure 5.
The permeating gas exits permeator 228 via line 230 at substantiall~ the pressu~e of the effluent from compressor 206. The permeating gas passes to and is com~ined with the synthesis feed gas exitin~ compressor 206 in order to be recycled to ammonia synthesis reaction zone 216. Since the synthesis feed ga~ has been dried, the presence of ~mmonia due to permeation through the separation membrane can be tolerated in the synthesis feed gas to ~e compressed.
The non-permeating gas from first permeator 228 is passed via line 232 to water scrubber 234 ~o remove ammonia.
Sincs a significant amount of the hydrogen has been separated fr~m the purge stream in permeator 228, the water scrubber can be of less volume than would ~e necessary if the scru~ber were positioned upstream of the first permea~or, The non-permeating gas having ammonia r~moved is passed through line 236 to second permeator 238. A second permeating gas is obtained ro~ second permeator 238 at a 112962~

pressure substantially the same as the pressure of the synthesis feed gas in line 200. The second permeating gas is passed through line 240 to line 200 whereat it is combined with the synthesis feed gas in order to be recycled to ammonia synthesis reaction 20ne 216. Since the second permeating gas may contain water vapor from water scrubber 234, it is added to the synthesis feed gas prior to the synthesis feed gas being dried in adsorber 202~ The non-permeating gas from second permeator 238 exits via line 242~
The ammonia synthesis system of Figure 4 is substantially the same as the system depicted in Figure 3 except that the purge stream in line 224 is compressed in compressor 22~ to sufficiently elevated pressures that the first permeating gas is at a pressure suitable for being directly introduced back into the ammonia synthesis loop via line 231. Also, the second permeating gas from second permeator 238 can be at a higher total pressure than the second permeating gas in the ammonia system depicted in Figure 3 even though the total pressure dif~erentiaLs are essentially the same.
Accordingly, in the system of Figure 4, the second permeating gas is combined with the synthesis feed gas downstream of compressor 206. Instead of adsor~er 2Q2 as in the system of Figur~ 3, coalescer 207 is provided between compressors 206 and 203 in order to r~move water vapor from t~e combined synthesis feed gas and second permeating gas~
The following example is provided in illustration of a process in accordance with th~^s invention~ ~11 parts and percentages are b~ volume unless otherwise noted.
Ammonia is synthesi~ed from nitrogen and hydrogen employing an ~mmonia synthesis plant similar to that depicted in Figure 1. The hydrogen feedstock is obtained by primary reforming of natural gas and the synthesis feed gas is obtained by introducing air and the primary reformer effluent into a secondary reformer. The effluent from the secondary reformer is treated in a shift converter, a carbon dioxide absorber and a methanator to provide approximately 52,000 kilograms per hour of a synthesis feed gas containing about 25 7 mole percent nitrogen, 73.1 mole percent hydrogen, ~.12962~

0.6 mole percent methane, 0.4 mole percent argon, and O.2 mole percent wa~er. The synthesis feed gas is obtained a~ abou~ 28 atmospheres absolute and 50C. The synthesis feed gas is compressed to about 70 atmospheres absolute, cooled to about 8C to condense water. The syn~hesis dried feed gas îs ~urther compressed to about 133 a~mospheres a~solute and îs introduced into and combined with the gas in the ammonia synthesis loop. In the ammonia synthesis loop the combined gases are compressed an additional 6 or 7 atmospheres and are treated in an ammonia coalescer which removes a~out 44,500 kilograms o ammonia per hour. The gases are heated to about 135 to 140C. Approximately 310,000 kilograms per hour of gas comprising about 66.5 mole percent hydrogen, 22 mole percent nitrogen, 6.8 mole percent methane, 3.5 mole percent argon, and 1,2 mole percent ammonia are introduced into a Kellogg-type ammonia synthesis converter utilizing a promoted iron $mmonia synthesis catalyst. A
reaction effluent gas at a temperature at about 280C is obtained from the sgnthesis con~erter and contains about 11.4 percent ammonia. The effluent is cooled to a~out 43C. A
purge stream of about 2~1 percent of the gases in the synthesis loop is removed, and the remaining gases are returned to the s~nthesis loop compressor.
T~e purge stream is chil-led to a~out ~23C and about 1000 kilograms per hour of liquid ammonia are condensed and removed ~rom t~e purge stream~ T~e purge stream contains about 1,2 volume percent ammonia. The purge stream is then scrubbed with water at about 25C at a water rate of about 2000 kilograms per hour. The purge stream contains less than about 1~0 ppm~ ammonîa.
The purge stream is heated to about 3~G and then passed to the first permeator which consists of 25 hollow fi~er membrane-containing permeators in parallel. The permeators are similar to tha~ depicted in Figure 5 and each permeator contains about 93 square meters of ef~ective sur~ace area. The m~m~ranes are comprised of anisotropic polysulfone su~stantîally prepared in accordance with the method disclosed in Example 64 of West German published 1~.2~62~

patent application DT 27 50 874 except that the spinning solution contains about 30 weight percent solids; the spinning jet dimensions are about 458 microns outside diameter, 127 microns inside diameter, and 76 microns diameter injection bore; the inj~ction fluid is a mixture of 60 volume percent dimethylacetamide in water. The last godet bath is at a temperature of about 50C; and the fibers are washed for 24 hours with no subsequent storage in water. Appropriate polymer solution and injection fluid rates are employed such that the dimensions of the hollow fibers are about 450 microns outside diameter and about 120 microns inside diameter. The permeator exhibits a separation factor of hydrogen over methane of about 30 and a permeability of about 50 x 10-6 cubic centimeters of hy~rogen per square centimeters of surface area per second per centimeter of mercury pressure drop. A pressure drop of about 65 atmospheres is maintained across the mem~rane, and approximately 1100 kilograms per hour o a first permeating gas is obtained from the bore side of the first permeator stage. The first permeating gas comprises 90.3 volume percent hydrogen, 6,2 volume percent nitrogen, 2,4 volume percent methane, and 1.2 volume percent argon, The first permeating gas is introduced into the feed gas exiting the first compressor prior to the condensation of water from the combined synthesis feed gas and first permeating gas stre~.
The non-permeating gas from the first permeator stage is at a pressure of a~out 136 atmospheres absolute and contains about 43.8 volume percent hydrogen, 35,4 volume percent nitrogen, 13.7 volume percent me~hane, and 7.1 volume percent argon~ This non-permeating gas enters the second permeator stage comprising 7 permeators Cas described above2 in parallel. About 496 kilograms per hour of a second permeating gas is obtained from t~e bore side of the second permeator stage which is at a pressure of about 30 atmospheres absolute a~d comprises 88 volume percent hydrogen, 7.4 volume percen~ nitrogen, 1.5 volume percent argon, and 2.4 volume percent methane. The non~permeating 1129~26 -31~ 07~Q403A

gas from the second permeator stage is at a pressure of about 136 atmospheres absolute and contains about 24.2 volume percent hydrogen, 47.7 volume percent nitrogen, 18.5 volume percent methane, and 9.6 volume percent argon~
About 86.3 percent of the hydrogen in the purge stre~m is recycled to the 2mmonia synthesis reaction zone.

Claims

THE EMBODIMENTS OF THE INVENTION IN WHICH AN EXCLUSIVE
PROPERTY OR PRIVILEGE IS CLAIMED ARE DEFINED AS FOLLOWS

1. A process for separating at least one gas from a gaseous feed mixture containing at least one other gas comprising passing the gaseous feed mixture to at least two permeator stages in series, each permeator stage comprising a separation membrane exhibiting selectivity to the permeability of said at least one gas as compared to the permeability of said at least one other gas in the gaseous feed mixture and having a feed side and a permeate exit side in which the permeate exit side is at a lower total pressure than the total pressure on the feed side, wherein between permeator stages, the non-permeating gas from the feed side of one permeator stage is passed to the feed side of the next permeator stage; and wherein at least one permeator stage has a lower ratio of total pressure on the feed side to total pressure on the permeate exit side than the ratio of total pressure on the feed side to total pressure on the permeate exit side of at least one subsequent permeator stage, and the total pressure on said permeate exit side of said at least one subsequent permeator stage is lower than the total pressure on said permeate exit side of said at least one permeator stage.

2. The process of claim 1 wherein the total pressure differential across the at least one permeator stage having a low ratio of total pressure on the feed side to total pressure on the permeator exit side is at least about 20 atmospheres.

3. The process of claim 1 wherein the total pressure differential across the at least one permeator stage having a higher ratio of total pressure on the feed side to total pressure on the permeate exit side is at least about 20 atmospheres.

4. The process of claim 1 wherein the at least one permeator stage having a lower ratio of total pressure on the feed side to total pressure on the permeate exit side is operated predominantly on a flux-limiting basis.

5. The process of claim 4 wherein the percent of difference in partial pressures of said at least one gas between the gaseous feed mixture passed to said lower ratio of total pressure permeator stage and the non-permeating gas from said permeator stage divided by the difference between the partial pressure of said at least one gas in the gaseous feed mixture passed to said permeator stage and the minimum partial pressure of said at least one gas on the permeate exit side of said permeator stage is about 20 to 90.

6. The process of claim 1 or 5 wherein at least about 20 percent of the at least one gas in the gaseous feed mixture permeates to the permeate exit side of the at least one permeator stage having a lower ratio of total pressure on the feed side to total pressure on the permeate exit side.

7. The process of claim 1 or 5 wherein the last permeator stage is said permeator stage having a higher ratio of total pressure on the feed side to total pressure on the permeate exit side.

8. The process of claim 1 or 5 wherein the amount of said at least one gas in the permeating gas from the at least one permeator stage having a higher ratio of total pressure on the feed side to total pressure on the permeate exit side is at least about 10 percent of the at least one gas in the gaseous feed mixture.

9. The process of claim 1 or 5 wherein the ratio of the total pressure on the feed side to total pressure on the permeate exit side across the at least one permeator stage having a lower ratio of total pressure on the feed side to total pressure on the permeate exit side is at least about 10 percent less than that ratio across the at least one permeator stage having a higher ratio of total pressure on the feed side to total pressure on the permeate exit side.

10. A process for separating at least one gas from a gaseous feed mixture containing at least one other gas comprising passing the gaseous feed mixture to at least two permeator stages in series, each permeator stage comprising a separation membrane exhibiting selectivity to the permeability of said at least one gas as compared to the permeability of said at least one other gas in the gaseous feed mixture and having a feed side and a permeate exit side in which the permeate exit side is at a lower total pressure than the total pressure on the feed side, wherein between permeator stages, the non-permeating gas from the feed side of one permeator stage is passed to the feed side of the next permeator stage; wherein at least one permeator stage has a lower ratio of total pressure on the feed side to total pressure on the permeate exit side than the ratio of total pressure on the feed side to total pressure on the permeate exit side of at least one subsequent permeator stage, and wherein said at least one permeator stage having a lower ratio of total pressure on the feed side to total pressure on the permeate exit side is operated predominantly on a flux-limiting basis.

11. The process of claim 10 wherein the percent of difference in partial pressures of said at least one gas between the gaseous feed mixture passed to said lower ratio of total pressure permeator stage and the non-permeating gas from said permeator stage divided by the difference between the partial pressure of said at least one gas in the gaseous feed mixture passed to said permeator stage and the minimum partial pressure of said at least one gas on the permeate exit side of said permeator stage is about 20 to 90.

12. The process of claim 11 wherein the total pressure differential across the at least one permeator stage having a low ratio of total pressure on the feed side to total pressure on the permeate exit side is at least about 20 atmospheres.

13. The process of claim 11 or 12 wherein the total pressure differential across the at least one permeator stage having a higher ratio of total pressure on the feed side to total pressure on the permeate exit side is at least about 20 atmospheres.

14. The process of claim 11 wherein at least about 20 percent of the at least one gas in the gaseous feed mixture permeates to the permeate exit side of the at least one permeator stage having a lower ratio of total pressure on the feed side to total pressure on the permeate exit side.

15. The process of claim 11 or 14 wherein the amount of said at least one gas in the permeating gas from the at least one permeator stage having a higher ratio of total pressure on the feed side to total pressure on the permeate exit side is at least about 10 percent of the at least one gas in the gaseous feed mixture.

16. The process of claim 11 or 14 wherein the ratio of the total pressure on the feed side to total pressure on the permeate exit side across the at least one permeator stage having a lower ratio of total pressure on the feed side to total pressure on the permeate exit side is at least about 10 percent less than that ratio across the at least one permeator stage having a higher ratio of total pressure on the feed side to total pressure on the permeate exit side.

17. The process of claim 1 or 11 wherein the total pressure on the feed side of the permeator stage having the higher ratio of total pressure on the feed side to total pressure on the permeate exit side, is up to substantially the same total pressure as the total pressure on the feed side of the permeate stage having the lower ratio of total pressure on the feed side to total pressure on the permeate exit side.

18. In a process for synthesizing ammonia from hydrogen and nitrogen comprising introducing a synthesis feed gas comprising nitrogen, hydrogen and at least one of methane and argon as an inert contaminant at substantially a superatmospheric synthesis pressure into an ammonia synthesis loop, reacting hydrogen and nitrogen in the ammonia synthesis loop at a superatmospheric synthesis pressure to produce ammonia in an ammonia synthesis reaction zone, removing ammonia from the ammonia synthesis loop, and removing a purge stream from the ammonia synthesis loop in an amount sufficient to maintain the concentration of inert contaminants less than about 25 percent wherein the purge stream is passed to a permeator comprising a separation membrane having a feed side and a permeate exit side and exhibiting selectivity to the permeability of hydrogen as compared to the permeability of each of methane and argon, a permeating gas is recovered at the permeate exit side of the separation membrane and is recycled to the ammonia synthesis reaction zone, the improvement wherein the permeator comprises at least two permeator stages in series, each permeator stage having a feed side and a permeate exit side in which the permeate exit side is at a lower total pressure than the total pressure on the feed side, wherein between permeator stages, the non-permeating gas from the feed side of one permeator stage is passed to the feed side of the next permeator stage; wherein at least one permeator stage has a lower ratio of total pressure on the feed side to total pressure on the permeate exit side than the ratio of total pressure on the feed side to total pressure on the permeate exit side of at least one subsequent permeator stage;
and wherein the permeating gas from at least one permeator stage is recycled to the ammonia synthesis reaction zone.

19. The process of claim 18 wherein the permeating gas from at least one permeator stage having a lower ratio of total pressure on the feed side to total pressure on the permeate exit side and the permeating gas from at least one permeator stage having a higher ratio of total pressure on the feed side to total pressure on the permeate exit side are recycled to the ammonia synthesis reaction zone.

20. The process of claim 19 wherein the permeating gas from at least one permeator stage is combined with the synthesis feed gas at a point where the permeating gas from said permeator stage is at a slightly higher pressure than the synthesis feed gas.

21. The process of claim 20 wherein the synthesis feed gas is compressed in at least two stages to said super-atmospheric synthesis pressure and the permeating gas from said at least one permeator stage having a lower ratio of total pressure on the feed side to total pressure on the permeate exit side is introduced into said synthesis feed gas between two of the compression stages.

22. The process of claim 21 wherein the permeating gas from said at least one permeator stage having a higher ratio of total pressure on the feed side to total pressure on the permeate exit side is introduced into said synthesis feed gas at a point where the permeating gas is at a slightly higher pressure than the synthesis feed gas, which point is prior to the introduction of the permeating gas from the at least one permeator stage having a lower ratio of total pressure on the feed side to total pressure on the permeate exit side, and at least one compressor stage is between the point where the permeating gas from said at least one permeator stage having a higher ratio is introduced into the synthesis feed gas and the point where the permeating gas from the at least one permeator stage having a lower pressure ratio is introduced.

23. The process of claim 21 wherein the purge stream passed to said at least one permeator stage having a lower ratio of total pressure on the feed side to total pressure on the permeate exit side is at a total pressure of at least substantially said superatmospheric synthesis pressure.

24. The process of claim 18 or 23 wherein the total pressure differential across the at least one permeator stage having a low ratio of total pressure on the feed side to total pressure on the permeate exit side is at least about 20 atmospheres.

25. The process of claim 18 or 23 wherein the total pressure differential across the at least one permeator stage having a higher ratio of total pressure on the feed side to total pressure on the permeate exit side is at least about 20 atmospheres.

26. The process of claim 18 or 23 wherein the permeating gas from the at least one permeator stage having a higher ratio of total pressure on the feed side to total pressure on the permeate exit side is at a lower total pressure than the permeating gas from the at least one permeator stage having a lower ratio of total pressure on the feed side to total pressure on the permeate exit side.

27. The process of claim 18 wherein the at least one permeator stage having a lower ratio of total pressure on the feed side to total pressure on the permeate exit side is operated predominantly on a flux-limiting basis.

28. The process of claim 27 wherein the percent of difference in hydrogen partial pressures between the purge gas passed to said permeator stage and the non-permeating gas from said permeator stage divided by the difference between the hydrogen partial pressure of the purge gas passed to said permeator stage and the minimum hydrogen partial pressure on the permeate exit side of said permeator stage is about 20 to 90.

29. The process of claim 18 or 27 wherein at least about 20 percent of the hydrogen in the purge stream permeates to the permeate exit side of the at least one permeator stage having a lower ratio of total pressure on the feed side to total pressure on the permeate exit side.

30. The process of claims 18, 22 or 27 wherein the last permeator stage is said permeator stage having a higher ratio of total pressure on the feed side to total pressure on the permeate exit side.

31. The process of claims 18, 22 or 27 wherein the amount of hydrogen in the permeating gas from the at least one permeator stage having a higher ratio of total pressure on the feed side to total pressure on the permeate exit side is at least about 10 percent of the hydrogen in the purge stream.

32. The process of claims 18, 19 or 22 wherein ammonia is removed from the purge stream prior to at least one permeator stage to provide a purge stream containing less than about 0.5 volume percent ammonia.

33. The process of claim 18, 19 or 22 wherein ammonia is removed prior to the first permeator stage.

34. The process of claim 18, 19 or 22 wherein ammonia is removed between two permeator stages.

35. The process of claim 18, 19 or 22 wherein ammonia is removed from the purge stream prior to the first permeator stage to provide a purge stream containing less than about 0.5 volume percent ammonia.

36. The process of claim 18, 19 or 22 wherein ammonia is removed from the purge stream between two permeator stages to provide a purge stream containing less than about 0.5 volume percent ammonia.