Chapter 15 Membrane Permeation Processes INTRODUCTION, 1238 History and Status, 1239 PROCESS TECHNOLOGY,1242 Transport Mechanisms, 1242 Design and Operating Considerations, 1245 Membrane and Module Configurations, 1246 Flow Arrangements, 1250 Simulation andDesign Calculations, 1252 Fields of Application, 1258 APPLICATION CASE STUDIES, 1259 Hydrogen, 1259 Carbon Dioxide, Hydrogen Sulfide, and W ater Removal, 1270 Helium Removal from Natural Gas, 1281 Air Separation, 1282 Solvent Vapors, 1288 REFERENCES, 1291 INTRODUCTION Membranetechnology,asappliedtogases,involvestheseparationofindividualcompo- nentsonthebasisof thedifferenceintheir ratesof permeationthroughathinmembranebar- rier.Therateofpermeationforeachcomponentisdeterminedbythecharacteristicsofthe component, thecharacteristicsofthemembrane,andthepartialpressuredifferentialofthe gaseouscomponent acrossthemembrane.Sinceseparationisbasedonadifferenceinthe ratesof permeationratherthanonanabsolutebarriertoonecomponent,therecoveredcom- ponentthatflowsthroughthemembrane(thepermeate)isnever100%pure.Also,sincea 1238 MembranePermeationProcesses1239 finitepartialpressuredifferentialisrequiredasthedrivingforce,someportionoftheperme- atingcomponent remainsintheresiduegas,and100%recoveryisnotpossible.Asthese generalizationswouldsuggest,theprocessisparticularlysuitableforbulkremoval opera- tionsratherthanfortheremoval oftraceimpuritiesfromgasstreams.Itshouldbenoted, however, thatrelativelyhighproductpuritiesandhighrecoveriesarepossiblewithmem- branesystems(atincreasedcost)bytheuseofmultiplestagesandrecyclesystemsorwhen usedincombinationwithothertechnologies. Theresiduegasproductnormallyleavestheunitatapressureclosetothatofthefeed, whi l ethepermeat eproduct , whi chmust passt hroughthemembr ane, l eavesatamuch reducedpressure.Theprincipal(and/orhighestpurity)productmaybeeitherthepermeate (e.g.,theproductionofhydrogenfromdilutegasstreams)ortheresiduegas(e.g.,thepurifi- cationofnaturalgasbytheremoval ofexcesscarbondioxidefromhighpressurefeed),and theprocessmaybeconsideredeitherseparationorpurification. Gaspurificationandseparationbymembranepermeationhasmanyadvantages,including 9 Lowcapitalinvestment 9 Easeof operation.Processcanbeoperatedunattended 9 Goodweightandspaceefficiency 9 Easeofscaleup.However, thereislittleeconomyofscale(seedisadvantagesbelow) 9 Mi ni mal associatedhardware 9 Nomovi ngparts 9 Easeof installation 9 Flexibility 9 Mi ni mal utilityrequirements 9 Lowenvi ronment al impact 9 Reliability 9 Easeofincorporationofnewmembranedevelopments. Userscaninstallthenextgenera- tionofmembranesintoexistingequi pment atthescheduledmembranereplacementtime (Schell,1983) Theprincipaldisadvantagesare 9 Acleanfeedisrequired.Particulates,andinmostcasesentrainedliquids,mustberemoved. Filtrationtoremoveparticlesdowntoonemicroninsizeispreferred. 9 Becauseof theirmodul arnature,thereislittle economyofscaleassociatedwithlarger mem- braneinstallations. 9 Becausemembranesusepressureasthedrivingforceof theprocess,theremaybeaconsid- erableenergyrequi rement forgascompression. History andStatus Sincethe19thcenturyithasbeenknownthatcertainpol ymermembranescanseparate gasesbypermeation.Asearlyas1831,Mitchellreportedthatdifferentgasespermeatemem- branesatdifferentrates.Graham, in1866,di scussedthemechani smofpermeat i onand demonst rat edexperi ment al l ythatmixturesofgasescanbeseparatedusingrubbermem- branes.In1950,Wel l erandSteinerreportedonpermeationprocessesofindustrialimpor- tance,theseparationofoxygenfromairandtherecoveryofhel i umfromnaturalgas.How- ever,theselectivityandproductionratesofthemembranesavailableatthetimewerepoor 1240Gas Purification andthelargemembraneareasrequiredmademembranepermeationeconomicallyunattrac- tive.In1960,LoebandSourirajandeveloped atechniquetocastcelluloseacetateintoafilm thathadanactivethinsurfacelayerandahighlyporoussupportinglayer.Theseasymmetric membranesprovidedincreasedpermeationrateswhileretainingtheirselectivityforspecific gases.Thedevelopmentof thesemembranesimprovedtheeconomicsof membraneapplica- tionsand led toincreasedinterestin membrane technology. Duringthe1970s,considerableresearchanddevelopmentalworkwasdevotedtomem- branes.Manypotentialapplicationswereidentified,butcommercializationwasslow.In 1977,Monsantodemonstrated itsfirstfullscalemembraneseparatorat TexasCity,Texas,in ahydrogen/carbonmonoxideratioadjustmentapplication(BurmasterandCarter,1983).In 1979,MonsantocommercializeditshollowfibermembranemoduleasthePrismseparator. From1979to1982Prismseparatorswereevaluated inseveral refineryhydrogen purification applications(Bollingeretal.,1982).Thesuccessof thesepilottestsestablishedthecommer- cialviabilityof gasseparationwithmembranes.ThefirstlargescalecommercialCO2mem- braneseparationprojectwastheinstallationoftwomembraneseparationfacilitiesatthe Sacroctertiaryoil recovery projectin West Texasin1983.Upto80 MMscfdof gashasbeen processedin thesefacilities(Parro,1984). Sincetheearly1980s,membranetechnologyhasadvancedrapidlyandcontinuesto advance.Inadditiontocelluloseacetateandpolysulfone,thepolymersusedinmakinggas separationmembranesincludepolyimides,polyamides,polyaramid,polydimethylsiloxane, siliconpolycarbonate,neoprene,siliconerubber,andothers.Todaymembranescanbe designedtowithstanda2,000psipressuredifferential.Membranesusedinhydrogenorcar- bondioxideapplicationsoperateattemperaturesupto200~whilethoseusedinsolvent applicationscanoperateat temperaturesuptoabout 400~(Baker,1985). Improvementsinmanufacturingmethodshaveresultedinimprovedmembraneperfor- manceandeconomics.Afluxincreaseof5%andaseparationfactorincreaseof20%have resultedfromimprovedmanufacturingmethods(Hamaker,1991),whileduringthemid- 1980s,airseparationmembranesbecamefrom twotofour timesmoreefficient.Notonlyare thepolymersrapidlychanging,butcapitalcostsarealsocomingdown.Advancesinmem- branesreducedtheinstalledcostof amembraneplantbyabout40%duringthe1980s(Spill- man,1989).Becauseoftherapidchangesintechnologyandcosts,economicstudiesand costdatabecomeoutdatedsoonafterpublication,andthereadershouldusecautioninusing costdatapresentedlaterinthischapterastechnicaladvancesarecontinuingtoimproveper- formanceand reducecosts. Thedevelopmentalandcommercialsuccessesof theearly1980s,andtheperceivedlarge market,attractedmanycompaniesintothefield.Asthetechnologymaturedandthemarket becameextremelycompetitive,somecompaniesdroppedoutandotherschangedownership. CompaniesofferingcommercialscalemembranesystemsarelistedinTable15-1.Thetable alsoidentifiesthe principalareasof applicationfor each company' sproducts. In thefieldof airseparation,theimprovingeconomicsof membrane-based processeshave encouragedlargeindustrialgassuppliersto joinforceswithmembranesuppliers.Thistrend ispointed out byPrasadet al.(1994)who provide the followingchronology: 1985:UnionCarbideIndustrialGases,Inc.andAlbanyInternationalMembraneVenture forma joint venture. 1986:InnovativeMembraneSystems(formerlyAlbanyInternationalMembraneVenture) becomesawhollyownedsubsidiaryof UnionCarbideIndustrialGases,Inc.(nowPraxair, Inc.). MembranePermeat i onProcesses1241 Table 15-1 Commercial-Scale Membrane Suppliers Company Application A/G Technology (AVIR) Air Products(Permea) AsahiGlass(HISEP) Cynara (Dow) Dow(Generon) DuPont Grace Membrane Systems Hoescht Celanese (Separex) International Permeation Membrane Technology and Research NipponKokanK.K. Osaka Gas Oxygen Enrichment Co. Perma Pure Techmashexport (USSR) Teijin Ltd. Toyobo Ube Industries Union Carbide(Linde) UOP/UnionCarbide Note: Air CO2H202N2 XXX XXXX XX XX X XXX X *Includes solvent vapor recovery,dehumidification,and~or heliumrecovery membranes. Source:Spillman(1989) and Prasad (1994) Other* X X 1988: British Oxygen Co.(BOC)and Dow/Generon Membrane Systems form a joint venture. 1989: L' Air Liquideand DuPont form a joint venture. 1990: Medal AL becomes a whollyownedsubsidiary of L' Air Liquide,Inc. 1991: Permea, Inc.becomes a wholly ownedsubsidiary of Air Productsand Chemicals, Inc. Thecommercial applicationof membrane-based hydrogen processes alsoexpanded rapid- lybecauseofthehighpermeabilityofselectedmembranesforhydrogen,thehighvalueof purehydrogen,andthewiderangeofhydrogencontaininggasstreamsinrefinery,petro- chemical,andindustrialchemicaloperations.AccordingtoKorosandFleming(1993),the majorsuppliersofmembranesystemsforhydrogenapplicationsarePermea,Medal,UOP, Ube,andSeparex.In1996, Medalreportedthatmorethan60oftheirhydrogenrecovery systems were in operation or under construction(Medal,1994). 1242Gas Purification PROCESS TECHNOLOGY Transport Mechanisms Itisgenerallyagreedthatasolution-diffusionmechanismgovernsthetransportofgases throughallcommerciallyimportantnonporousmembranes.Themechanisminvolvesthefol- lowing:(a)adsorptionof thegasatonesurfaceof themembrane,(b)solutionof thegasinto themembrane, (c)diffusionofthegasthroughthemembrane,(d)releaseofthegasfrom solutionattheoppositesurface,and(e)desorptionofthegasfromthesurface.Sincethese stepsarenotnecessarilyindependent,thetermpermeationisusedtodescribetheoverall transportof gasesthroughamembrane. Initssimplestform,thesolution-diffusionmodelconsidersonlystepsb,c,andd.This modelisbasedontwoassumptions:(1)theconcentrationofacomponentinamembraneat itssurfaceisdirectlyproportionaltothepartialpressureofthecomponentinthegasphase adjacenttothesurface,and(2)therateatwhichacomponentpassesthroughamembraneis proportionaltotheconcentrationgradient(concentration/distance)inthemembrane.These twoassumptionsrepresentHenry' slawandFi ck' sfirstlawofdiffusion,respectively,and canbestatedasfollows: c i -kip i(Henry'slaw)(15-1) Ji -"-Di(dci/dx)(Fick' slaw)(15-2) Where:c i - - local concentrationof iinthemembrane J i -steadystatefluxof i k i =solubility coefficient Pi=partialpressureof iinthegas D i =local diffusivity x=distancethroughactivemembrane Combi ni ngandintegratingequations15-1and15-2overthefullmembranethickness acrossthemembrane, yields Ji =PiApi//(15-3) Similarly,forahollowtubularmembrane,suchasahollowfiber,thesteadystaterateof gaspermeationis(Stem,1986) Ji =Pi(2rtLApi/ln(Ro/Ri))(15-4) Where:Pi=kiDi,thepermeabilitycoefficient Api =Pi(feed)-P i ( p e r me a t e ) (thepartialpressuredifferenceacrossthemembrane) 1 =membranethickness(x =1) Ro =effectiveouter radiusof thetube R i =effectiveinner radiusof thetube L=lengthof tube Me mbr ane Pe r me at i onPr oces s es 1243 Anyconsistentsetofunitsmaybeusedinequations15-1-15-4.Permeabilitycoefficient dataareoftengiveninBarrers.OneBarrer =10-1~3 {STP })(cm)/(cm2)(sec)(mmHg).Since commerci al membranesnormallyconsistofaverythinactivelayeronathickerporoussub- strate,theeffectivethicknessmaynotbeaccuratelyknown,anditismoreconvenienttouse thepermeationrate,Pi/l,fortheoverallmembrane, asacorrelatingfactor.Typicalengineer- ingunitsforthepermeationrateare(scf)/(ftZ)(hr)(100psi).One(scf)/(ftZ)(hr)(100p s i ) - 1.55 10 -5(cm3){ STP})/(cm2)(sec)(cmHg).Theunitsof flux,J,are(scf)/(ftZ)(hr). TheHenry' sl aw/ Fi ck' slawmodelpreviouslydescribedisasimplificationoftheactual permeat i onmechani sm, andmorecompl exmodelshavebeenproposed.Thedualsystem model, forexample, isamorepreciserepresentationformanycases.Itassumesthatgas moleculeswhichdissolveinthedenseregionsofthemembranesurfacefollowHenry' slaw; whilemoleculesthatadsorbonthewallsofmicroscopiccavitiesinthemembranesurface followLangmui r ' sadsorptionisotherms.Equationsbasedonthedualsystemmodelhave beendevel opedandpresent edbyLeeetal.(1988).Addi t i onal di scussi onsoft ransportmechani smsareprovidedbyLaceyandLoeb(1972)andStern(1986). Permeat i onratesforspecificcomponent sandmembranesarenotconstants.Theyvary withtemperature,pressure,andthepresenceofothercomponent sinthegas.Detailedperme- ationratedataforcommerci al membranesarenormal l yconsideredproprietary; however,somecomparativedatahavebeenpublished.Tabl e15-2liststypicalpermeationratedatafor anumberofmembranesandgases.ThecelluloseacetatedataarefromMazurandChan (1982),andtheotherdataarebasedonapaperbyToml i nsonandFinn(1990). ThepermeationdatagiveninTabl e15-2arerelativevaluesbasedonapermeationrateof 1.0foroxygenincelluloseacetateandpolysulfone.However, datapresentedbySchelland Hoernschemeyer(1982)forthepermeationratesofvariousgasesincelluloseacetateindi- catethatactualvalues,expressedas(scf)/(ft2)(hr)(100psi),areclosetotherelativevalues listedinthetable.Forexample, theyshowthepermeationrateforoxygen,measuredatroom temperatureand100psig,toactuallybeabout1.0(scf)/(ftZ)(hr)(100psi),andtheactualper- meationratesforothergasesincelluloseacetatetobesimilartothoselisted. Table 15-2 Permeation Rates of Gases Through Membranes RelativePermeationRate (1) MembraneHzN202CH4COzH20HeHzSCOC2H6 Polysulfone Cellulose Acetate Polyamide DowProduct Permea Product PDMS (2) 130.210.226 120.1810.26 90.050.50.05 13683293 220.42.30.49 649281604 10015100.30.1 Notes: 1.Permeationratesarerelative,based on1 f oroxygenin polysulf oneand celluloseacetate. 2.PDMS= poly(dimethylsiloxane);siliconerubber. Sources:Tomlinsonand Finn(1990) and Mazur andChan(1982) 1244Gas Purification Leeetal.(1995)evaluatedavailabledataonthepermeabilityof gasesincelluloseacetate toobtainvaluesforuseindesigningsystemsforremovingCO2fromnaturalgas.Theycon- cludedthat the followingwould be reasonablevaluesfor computersimulationof the process: Permeabilityratefor CO 2 :910-5(cm3){ STP})/(cm2)(sec)(cmHg) (or approximately5.8(scf)/(ftZ)(hr)(100psi) Selectivity for CO2 / CH 4 =20 Selectivity for Nz / CH 4 =1 Sel ect i vi t yf or ( C2+) / CH 4 --0. 4 AdditionalpermeabilitydataforcelluloseacetatemembranesareprovidedbyLietal. (1990),Donohue et al.(1989),and Ettouney et al.(1995). Gaseswithhighpermeabilityratesareoftencalled"fast"gases.ThedatainTable15-2 indicatethatlowmolecularweightandhighlypolargasestendtobefastgases;whilethe slowergasesarenonpolarand/orhighermolecularweight.Detaileddataonorganicsolvent permeationratesare given byBakeret al.(1986). Theseparationfactorof amembrane,c~ij, isdefined as: Oqj :Pi/Pj( 15- 5)andisanindicationofamembrane' sabilitytoseparatespeciesiand j.Aseparationfactor greaterthan1indicatesthat,atequalpartialpressures,componentipermeatesthroughthe membranefasterthan j.Veryhigh(orverylow)separationfactorsresultineasyseparations. Noseparationispossibleif oqj =1. Separationfactorsareinfluencedbymembranematerials,feedcomposition,temperature, andpressure.SeparationfactorsforvarioussystemsarepresentedinTable15-3.Notethat theseparationfactorfor oxygen andnitrogen isconsiderablylower thanthosefor other com- ponentpairsgenerallyconsideredviableforcommercialmembraneseparation.Oxygenand nitrogenareverysimilarinmoleculesizeandsolubility,whichmakestheirseparationvia Table 15-3 Separation Factors for Various Component Systems in Commercial Membranes ComponentSystem Rangesof Separation Factors Typical Separation Factors inCellulose Acetate CO2 / CH 410-5025 O2/N 23-12 H2/CH445-20045 H2/CO35-80 Hz] N 245-20045 Hz S/ CH 440-6050 He/CH460-10060 CO2/ C2H644-5250 Sources:Stookeyet al.(1986)and Spillmanet al.(1988) MembranePermeationProcesses1245 membranesoneofthemoredifficultprocesses.However,duetotheabundanceof"free" feedstocka high recoveryisnot required. Theseparationfactorsforsolvent/airprocesseshaveaverywiderangedependingonthe membranematerial.Forexample,theseparationfactorfortoluene/nitrogencanbeaslowas 40fornitrile rubberorashighas10,000for neoprene(Bakeretal.,1986). Design and Operating Considerations Theseparationefficiencyofamembraneforagivengasmixturewilldependonthegas composition,thepressuredifferencebetweenthefeedandthepermeate,andtheseparation factorforthetwocomponentsatthespecificoperatingconditions.Thehighertheseparation factor,thegreater theselectivity of themembraneandthehigher the productpurity. Thegascompositionandpressuredifferentialbecomeveryimportantwhenthemoreper- meablegasinthefeedhasalowconcentration.Sincethepartialpressureofthefastcompo- nentonthepermeatesidecannotexceeditspartialpressureonthefeedside,highfeedgas pressuresandlowpermeatepressuresarerequiredtoobtainefficientseparationsevenwith highseparationfactors.Thedifferentialpressureacrossthemembranerelatesdirectlytothe membranearearequired.Compressioncostsontheotherhandareafunctionofpressure ratio.Therefore,operationathighpressurewithasubstantialpressuredifferentialacrossthe membranebutwithareasonablylowpressureratio,iseconomicallyadvantageouswhere recompressionof thepermeateisrequired. Thegeneraleffectsofvaryingkeyoperatingfactors,whileholdingotherconditionscon- stant,foratypicalsingle-stagemembrane permeationsystem,canbesummarizedasfollows: 1) Increasingthe overalldifferentialpressureacrossthemembraneleadstoanincreasein per- meateflowrateandadecreaseintheconcentrationofthefastgasinthepermeatestream. Thedifferentialpressureatanypointinamodule canbeaffectedbypressuredropin either thefeedor permeateflowchannels.If thepermeateflowrateissufficientlyhighit cangen- eratea backpressurethat reducesthedifferentialpressureandrateof permeation. 2)Increasingthefeedgasflowratedecreasesthepercentrecoveryofthefastgasasperme- ateanddecreasesthepurityof theresidue.However,increasingthefeedrateincreasesthe purityof the permeateandincreasesthepercent recoveryof theslowgasin the residue. 3)Decreasingthefeedgasflowratebelowacriticalvaluedecreasestheseparationefficien- cyduetoaboundarylayereffect.Theconcentrationofthefastgasisdepletedinthefeed gasadjacenttothemembranesurface,reducingitspartialpressureandthereforeitsrateof permeation.Sincetherateofpermeationoftheslowgasisnotaffected(orisincreased) thisreducespermeatepurity.Thecriticalflowrateisdeterminedbythedegreeofmixing atthemembranesurface,andthisisafunctionof gasvelocity;gaspropertiessuchasvis- cosity,density,anddiffusivity;andmodule design. 4)Increasingthetemperatureraisesmostpermeabilitiesbyabout10to15%per10~and haslittle effectonseparationfactors(SchellandHoernschemeyer,1982). 5)Increasingthemembraneareaincreasesthepurityoftheresidue;whiledecreasingthe membraneareaincreasesthepurityof thepermeate. Inhydrocarbonmembranesystems,permeationoffastgases(e.g.,H 2 orCO2)increases theconcentrationofheavycomponents(e.g.,C3+hydrocarbons)intheremaininggas.The increasedconcentrationofreadilycondensiblecomponents,aswellasthedecreaseintem- peraturethatfrequentlyaccompaniespermeation, maycausethecondensationofliquid 1246Gas Purification hydrocarbonswithinthemembraneunit,interferingwithitsoperation.Typicalremediesfor thisproblem arepreheatingthegasandremovingheavyhydrocarbonsfromthegasaheadof themembraneunit.Evenwheremembranescanphysicallytoleratethecondensate,perfor- mancewillnormallysuffer.Thecondensedliquid(usuallymadeupprimarilyofslowper- meatingcomponents)cancoverthemembranesurface,forminganadditionalbarriertoper- meation.Alternatively, theliquidmaywetthemembranecreatingagasleakagepath,which canresult in permeate contamination. Membranelifeisanimportantfactoraffectingprocesseconomics.Membranestypically requirereplacementeverythreetosevenyears.Itisthereforenecessarytoconsiderfactors thatmayshortenorpossiblyextendtheirlife.Funketal.(1986)investigatedtheeffectof impuritiesinthegasonacelluloseacetatemembraneinacidgasservice.Theyfoundthat variouscomponentsinthegasaffectthepermeability,thetensilestrength,andtheelastic modulusof themembrane.Someof theadverse effectsof individual componentscanbemit- igated byproper design(e.g.,temperature control),butother impuritiesmayrequireremoval priortomembraneprocessing.Therefore,thefeedgasshouldbeevaluatedforthepresence ofparticulates,entrainedliquids,oilmist(compressors),condensiblecomponents,alltrace compounds,andwater.Pretreatment systemsthatmay beneededinclude: l)Ahighefficiencyseparator to remove particlesandoil mist 2)Aliquid knock-out drumto remove liquid hydrocarbons andwater 3)Preheator reheatstepsto raisethegastemperaturesufficientlyaboveitswaterandhydro- carbondewpoint to prevent condensationin the module 4)Componentremovalstepstoeliminatecompounds,suchassolvents,BTX,orammonia, thatmaycausemembranedeterioration Thedesignof pretreatment equipment mustalsoaccountfor bothnormaloperating andupset conditions. Sincemanyof theprocessvariablesarehighlyinterdependent,theoverall designof mem- branesystems requiresconsiderationof all requirementsandconditions.Thisisillustrated in Fi gur e15-1,whichshowstheeffectsof percent hydrogeninthe permeate,percent hydrogen recovery,andpermeatepressureontherelativecostsofrecoveringhydrogeninatypical application.Thesethreefactors,aswellasothers,mustbeevaluated tooptimizeoverall eco- nomicsforeachspecificcase.Mostmembranesystemsuppliershavecomputerprogramsto optimizethesystemdesign(PoffenbargerandGastinne,1989;Spillman,1989;MacLeanet al.,1983). Membrane andModule Configurations Membranes Thekeyrequirementsforamembranetobeusedinaneconomicalgaspurificationorsep- arationprocessare: 1) Highpermeability for the component to beremoved 2)Highselectivity for the component to beremoved in relation to other components 3)Highmembranestability inthepresenceof allgascomponentswhichwillcomeintocon- tactwiththemembrane 4)Uni formi t ywfreedom from pinholesor other defects MembranePermeationProcesses1 2 4 71 . 2 5 - 0 1. 15 > 0 o ff o ~1. 05 m @ r r%Hz R e c o v e r y9 5 4 0~ ' ~ . ,,. . ,~ 07, , ,-%%%% " ' - C' - . _ -,_- , . " - , _ ' . . / I-" - ' , , , ~ >. . - " ' . , . - " ' , , , " , 1 " i1 4 0 p s i a0 . 9 5 IIIII1IJ!I 9 0 9 5 1 O0%H y d r o g e n i nP e r me a t eFigure15-1. Ef f ects ofhydrogen permeatepu rity , permeatepressu re, andpercent hydrogen recoveryonrelativecost ofhydrogen recoveryf rom a7 5%H2,815psia refinerygasstream.(/-leyd, 19 8 6 )5)Loweffective thicknessof theactive portion of themembraneto ensureahigh permeation rate 6)Physicalstrength towithstandthe required operating conditions Items1 through3arerelatedprimarilytothepolymersusedinfabricatingthemembrane; whileitems4through6relatetothefabricationmethod.Akeybreakthroughinthedevelop- mentof fabricationmethodswasprovided byLoebandSourirajan(1960).Theydeveloped a techniqueforcastingasymmetriccelluloseacetatemembraneswithauniform,verythin (0.1-1.0micron)skinonastrongporoussubstrate(100-200micronsthick).Sincethisorigi- nalwork,whichwasactuallyaimedatthedevelopmentofreverseosmosismembranes,the basicapproachhasbeenappliedtoavarietyof polymericmaterialsandtosheetandhollow fiber configurationsforusein bothgasandliquid phaseoperations. Devel opment workhasalsoinvestigatedalternativeasymmet ri cmembranesystems, including(a)anultrathinnonporousfilmlaminatedtoamuchthickermicroporousbacking (whichmaybeadifferentmaterial)and(b)averythinnonporousfilmappliedasacoating toathickermicroporoussubstrate(Stern,1986).Acomplexmembranestructurereportedly usedintheMonsantoPrismseparator isa"skinned"asymmetrichollowfiberof polysulfone coated withathinfilmof silicone rubber(about1 micronthick).The polysulfone skin(about 0.1micronthick)istheactiveseparator,whilethesiliconerubberservestosealanydefects inthebasemembranewithoutaffectingtheintrinsicpermeabilityofthemembrane(Koros andChern,1987). KorosandFleming(1993)presentacomprehensive reviewofmembrane-basedgassepa- rationwithemphasisonmembranematerials,formationtechniques,andmoduledesigns. Themostpopularmoduleconfigurationsarethehollowfiberandspiral-wounddesignsdue to their high packingdensity. 1248Gas Purification Spiral-Wound Configuration Inthespiral-woundconfigurationanenvelopeisformedwithtwomembranesheetssepa- ratedbyaporoussupportmaterial.Typically themoduleconsistsofseveralsuchenvelopes. Thematerialbetweenthemembranes(permeatechannelspacer)supportsthemagainstthe operatingpressureanddefinesthepermeateflowchannel.Theenvelopeissealedonthree sides.Thefourthsideissealedtoaperforatedpermeatecollection tube,andtheenvelopeis wrappedaroundthe collection tubewithanet-like spacer sheet that hastwofunctions: 1) It keepsadjacentmembranesapart to formafeed channel. 2)Itpromotesturbulenceofthefeedgasmixtureasitpassesthroughthemodule,thus reducingconcentration polarization. Duringoperation,thefeedgasmixtureentersonefaceof themodule,travelsaxially along thefeedchannelspacerandmembranesurface,andexitstheother faceasaresidueor reten- tate.Themorepermeablegasespassthroughthemembranesandtravelinaspiralpath inwardwithintheenvelopethroughthepermeatechannelspaceruntiltheyreachtheperfo- ratedcollection tubeandfinally exit aspermeate(Fi gure15-2).Thefeedchannelspacer isa keyfeatureofthespiral-woundmoduleand,asshownbyDaCostaetal.(1991,1994),its designsignificantlyaffectsmoduleperformance.Typicallythemoduleshaveabout1,000 squarefeetofsurfacepercubicfootofvolumeandare4-12inchesindiameterby36-42 incheslong.Uptosix modules maybehousedin asingle pressurevessel shell. Hollow Fiber Configuration Thehollowfiberconfigurationconsistsofthousandsofhollowfiberspackagedinbun- dlesmountedinapressurevesselresemblingashellandtubeheatexchanger.Forhigh pressureapplicationsthefiberdiameterisusuallyontheorderof100~tmIDand150-200 ~tmOD.Thebundlesarecappedononeendandhaveanopentubesheetontheotherend Figure 15-2. Diagram of a spiral-wound membrane permeation el ement. (Courtesyof Membrane Technologyand Research, Inc.) MembranePermeationProcesses1249 (Figure 15-3).Thebundlestypically have3,000squarefeetofmembranesurfacepercubic footof modulevolumeandthemoduledimensionsrangefrom4-12in.indiameter by4-20 ftlong.Thefeedgasisintroducedontheshellsidebecausehollowfibersaremuchstronger undercompressionthanexpansion.Thefasterpermeatinggasesmigrateintothefiberbore andexit via theopenendof the bundle.For lowpressureapplications thefibershaveadiam- etergreaterthan400lamandthefeedgasenterstheboresidewhilethepermeateexitsvia theshell side.Thisconfiguration reducespressuredroponthe feedside. Notallmembranematerialscanbemadeintoathinselectivelayeronaporoussubstrate inahollowfiberform.Consequently,spiral-woundmembranes,whichcanbemadefroma wider rangeof materials,usually havehigher permeation rates.However,thisisoffset bythe muchhigher packingdensityof hollowfiber modules, resultinginsimilar overall productivi- typerunitmodulevolumeforthetwoconfigurations.Thissituationcouldchangeifdevel- opmentsin polymer science lead tomoreeffective thinfilmsina hollow fiber form. Numerousmechanicaldesignsofmoduleshavebeendeveloped forboththespiral-wound andhollow-fiberconcepts.Thevariousdesignsareaimedatoptimizingsuchfeaturesas: membraneareaperunitvolume,gasflowdistributionandpressuredrop,sealsandfasten- ings,andassemblytechnology.Witheitherspiral-woundorhollow-fibersystems,large Figure 15-3. Diagram of ahollow-fiber membrane permeation element. (Courtesyof Permea) 1250Gas Purification commercialinstallationsnormallyrequirealargenumberof individualmodules.Thisisevi- dentinFi gure15-4,whichshowsaUOPAdvanced MembraneSystem for removingcarbon dioxide from naturalgas. Athird moduledesign--plateand framehutilizesastackof diskshapedmembranessepa- ratedbysheetsof porousfilterpaperandmembranesupportplates.Theflowpatternisvery muchlikethatofaplateandframefilter.Thisdesignisnotcompetitiveforlargecommer- cialapplicationsbecauseoftherelativelysmallmembraneareaperunitvolumeattainable. However,plateandframedesignsareusedfor producingoxygen enrichedairinsmallmed- icalapplications. Flow Arrangements Toobtainthespecifiedflowratesandpurities,theoptimumarrangementofthemodules needstobedetermined.Seriesflow,Fi gure15-5,providesforhighrecoveriesforagiven feedrate.Thepermeatepurityvariesfrommoduletomodule,whichmakesitpossibleto producemultiplepermeateproducts.Thesubstantialvelocitychangesthatoccurwhenlarge portionsof thefeedarerecoveredaspermeatecanbeaccommodatedbyusingsuccessively smallerelements.Theseriesflowarrangementhasaslightlyhigherpressuredropbetween the feedand the residue thana parallelflowarrangement. Parallelflow,Fi gure15-6,allowsforhigherfeedratesforthesamerecovery.Turndown isaccomplishedbytakingelementsout of service.In designingparallelflowsystems,partic- ularcaremustbepaidtothegasdistributionsystems.Non-uniformgasflowtotheelements Figure 15-4. Large commercial installation of UOP Advanced Membrane System for removing carbon dioxide f rom natural g as. (Courtesyof UOt~MembranePermeationProcesses1251 Residue Gas (Non-Permeate Gas) Feed Gas PermeateGas F ig u re15-5. Series flowconfiguration.Feed Gas F ig u re15-6. Parallel flowconfiguration.Residue Gas (Non-Permeate Gas) Permeate Gas canresultin declinesin product purityand recovery.High recovery in one element cancause liquidformation. Thecombinedseries-parallelflowsequence,Fi gure15-7,providesforgoodturndown ratioswhileaccommodatinghighfeedrateswithlowresidueflowrates.Areductioninthe number of modulesin parallelaccommodatesthe reductionin flowaspermeateisremoved. Asstatedearlier,separationbymembranepermeationisnotanabsoluteseparation.Each specieshasafinitepermeabilitythroughthemembraneandtheenrichmentisachieveddue torelativepermeabilities,notzeropermeabilityforoneof thespecies.Ashigher puritiesare approached,recoveryofproductdeclinesrapidlyandsingle-stagesystemsbecomeincreas- inglyinefficient.Therefore,single-stageprocessesfrequentlyareapplicabletobulkremoval of aspeciesor the concentrationof feedstoother purificationprocesses. 1 2 5 2 Gas Purification FeedGasPermeateGasResidueGas (Non-Permeate Gas) Figure15-7 . Combined series-parallelflowconfiguration. Asanalternative,multi-stagemembranesystemswithrecompressionarrangementsmay beemployed.Multi-stagearrangements,Fi gure15-8,workwellwhenhighpermeatepurity orimprovedrecoveryratesarerequired.Additionalmulti-stagearrangementsaredescribed bySpillmanet al.(1988). Simulation andDesign Calculations Inconsideringmembranesforvariousapplications,theengineermusthaveameansof estimatingtheperformanceofamembranesystem.Thefundamentallawsofdiffusiondis- cussedearlierinthischapterunder"TransportMechanisms"apply.Mathematicalsolutions forcompositionsof residueandpermeatestreamshavebeendevelopedfortwocomponents andaredescribedintheliterature(e.g.,LaceyandLoeb,1979).Forthreeormorecompo- F eed9=Resid u e mrI Compressor TPermeate Figure15-8. Multi-stageflowconfiguration. Membrane Permeation Processes1253 nentsastrictmathematicalsolutiontothedifferentialequationshasnotbeenfound.Howev- er,anumericalsolutionispossible,andwiththeprevalenceofcomputers,thisapproachis notonlypracticalbutwidelyused. Thebasicequationdefiningtherateofflowofacomponentthroughamembraneisequa- tion15-3whichcanbewrittenintheform: Ji =Ri(Pi,feed-Pi,perm)(15-6) Where:Ji=steadystatefluxof i,moles/(time)(area)=Mi/A M i =flowrateof ithroughagivenarea,moles/ time ki Di Pi R i. . . .l1 Pi =partialpressureof iinthegas A=area k i -solubilitycoefficientfori D i =local diffusivityfori l=membranethickness Pi-permeabilitycoefficientfori Anyconsistentsetofunitsmaybeused.Gasvol ume(e.g.,scf)isoftenusedinsteadof moles. Thepartialpressureofeachindividualcomponentchangesastheseparationiscarriedout becausedifferentcomponentsarebeingremovedfromthefeed(highpressure)sideatdiffer- entrates.Becausethepartialpressureofindividualcomponent sisafunctionofposition alongthemembr anesurface,mat hemat i cal integrationofequat i on15-6overtheentire lengthof amembranesurfaceisaninterestingexercise fortwocomponents,butnotpractical forthreeor morecomponents. Commerci al membranestodayareasymmetricmembranes. Thismeansthattheactive membranesurfaceisonlyaverythinlayerontopofaporoussubstrate.Component swhich permeatethemembranemusttraveloutoftheporousmembranesubstratebeforeentering thebulkpermeatestream.Therefore,theeffectivepartialpressureandmolefractiononthe permeatesideoftheactivemembranesurfaceareonlyafunctionofthematerialpassing throughthemembrane, notafunctionof thebulkpermeatestream. Acomputersimulationof theprocesscanbemadebyconsideringsmallincrementalareas ofthemembraneindividually.Apermeationanalysisandamaterialbalanceareperformed onthefirstincrementalarea.Theresiduegasfromthisareaistreatedasthefeedtothenext areaandtheoperationisrepeated.Theanalysiscontinues,addingareasuntiltheresiduegas meetstheproductpurityrequirementor other processrequirementsareattained. If theassumedareasaresmallenough,thefeedandresiduegascompositionsforeachAA aresimilarenoughthatthedrivingforceforpermeationcanbebasedonthecompositionof thefeedratherthanonanaverageofthefeedandresiduecompositions.Thepermeatecom- positionforeachincrementalareamusttakeintoaccountthepermeationratesofallcompo- nents,butisnotaffectedbypermeatefromtheotherareas. Fi gur e15-9isasimplediagramshowinggasflowsadjacenttoandthroughanincrement of area,AA.Thecomputersimulationisbasedonthisdiagramandthefollowingequations: Forasmallincrementof membranearea,equation15-6becomes Mi =Ri(Pi,feed -- Pi,perm)AA(15-7) 1254Gas Purification ~ - - A A , Smal l I ncr e m(cF eed Ga sRe sid u e Ga s 'NI% l N' v r Z / / / / / / / , ; , . ..k J .=rementof A r e aA ct i v e Me mbr a nePor ou sSu bst r a t e"=Bu l kPe r me a t e I v W '~I r01O= N,=F l owof F eed t oA AN==F l owof Re sid u e f r omA A=F l owof F eed t oNe x t A AO1=Bu l kF l owof Pe r me a t e Pr ior t oA AO2=Bu l kF l owof Pe r me a t e A f t e r A AM=F l owof Pe r me a t e Th r ou g h A AFigure 15-9. Diagram of gas f l ow pattern at small increment of membrane area for computer simulation model. Substitutingforpartialpressures: Mi=Ri(Yi,feed greed-Yi, perm 7tperm)AA (15-8) Themolefractionof anycomponent,i,inthepermeateis Yi,perm =Mi/]~Mi(15-9) Similarly,themolefractionof componentiintheresiduefromincrementareaAAis Yi,residue =N2i/]~N2i(15-10) Combiningequations15-8and15-9yields RiAAy i,feed n feed -(15-11) MiI 1+RiAAzMi~permI Thisisthekeyequationinthecomputersimulationprogram.Otherimportantequations definetheincreaseinbulkpermeateanddecreaseinresiduegasflowateachAA: 02, i --01,i +M i(15-12) N2, i =N1, i -M i(15-13) Membrane Permeation Processes1255 Inequations15-7through15-13, AA =incrementalarea M i -flowof componentithroughAA,moles/ time Mi =f l o wo f allcomponentsthroughAA N1 =flowof feedgastoAA N 2 =flowof residuegasfromAA Ol=bulkflowof permeatefrompriorAA' s O2 =bulkflowof permeateafterAA Ri =permeability ratefori,moles/(time)(area)(partialpressuredifferential) Yi =molefractioniingas 7r =totalpressure Thepermeationrateforeachcomponentimustbeknown,andthetotalfeedpressureand permeatepressuremustalsobeknown. AreasonableZM iisassumedforthefirstincrementofarea,AA.Itisthencorrectedby subsequent iterations,untilasolutionisfoundwhereM icalculatedforeachcomponentequals(withinthespecifiedtolerance)theM i usedinthepreviousiteration.Thesumofall M i fromeachiterationisalsousedforthecalculationof thenew~M i.Atthefirstincrement of membranearea,Yi,feed issimplythemolefractionof componentiinthefeed. AfterthesolutionforM i foreachcomponenthasconverged,theresiduegasmolarflow, N2,i,foreachcomponentiscalculatedbysubtractingMi,thepermeateflow,fromthefeed flowN l,i. Thisresidueflowthenbecomesthefeedflowtothenextsmallincrementof mem- branearea. Theprocedureisrepeateduntilthedesiredproductspecificationsaremet.Aftereach incrementofareaiscalculated,thepermeateflowforeachcomponent,Ol,i,isincreasedby thepermeateflowforthatincrementalarea,M i.Similarly,ateachincrementofarea,the residuegasforeachcomponent,N l,isreducedbythepermeateflowpassingthroughthat incrementofarea,M i.ThesumofallN2,i'sfromthefinalincrementof membraneareaisthe residuegasforthesystem.Eachincrementofareaisalsoaccumulatedsothatthetotalarea requiredisknownaswellasthegascompositionsandquantitiesof theresidueandpermeate streams. Asimplifiedflowsheet fortherequiredcalculationsisshowninFi gur e15-10. Tabl e15-4summarizescomputeroutputfromaprogramusingthecalculationlogicout- linedinFi gur e15-10.Thedataarefromacomputersimulationofaspiral-woundcellulose acetatemembraneunitrecoveringCO2from7.4MMscfd(812mol/hr)of ahighCOzcontent gas,richinhydrocarbons,asmightbefoundinaCO2floodenhancedoilrecovery(EOR) project.Inputtotheprogramconsistedof thefeedcomposition,flowrate,pressure,andtem- perature;thepermeatepressure;andthespecificationthattheresiduegascontain40vol% CO2. Inadditiontotheflowratesandcompositionsoftheresiduegasandpermeate,theoutput indicatestheapproximatemembranesurfacearearequired,andthenumberofincrementsof membraneareausedforthecalculation.At1,000sqftofmembranesurfacearea/cubicfoot ofmembranemodul evol ume, approxi mat el y13.1c u f t ofmembranemodul evol umeis required.Amembranemodule8in.indiameterand36in.longoccupiesapproximately1 cu ftandcontainsapproximately1,000sqftofmembranesurfacearea.Therefore,thesimula- tionindicatesthat13modules(8in.diamx36in.)arerequiredfor theseparation. 1256Gas Purification INPUT SYSTEMDATAAND PRODUCTSPECS INITIALJZEFOR SMALLINCREMENTOF AREA,&A.SET IK FOR EACHCOMPONENTAT A REASONABLEVALUE e. g. ,Mi-Rte~ La.eYi,Feed(~ feed" ~ perm /2) CALCULATEINITIALT..,Mi FROMINITIALM FOR EACHCOMPONENT CALCULATENEW~ FFOREACHCOMPONENTBY EQUATION15-11 NEWM, - OLD M,? ~ _ i. e. ,is...NEW . IK- OLD I~ < 0. 0001 ? . . _NEWM,_ ~ADDNEWM~TO PRODUCENEW ~ -,M REPEATFOR EACHCOMPONENT IS FLAGSET FOR ANY COMPONENT? REPEAl"FOR EACHCOMPONENT NO INCREASEFLOWOF PERMEATEBY FLOWTHROUGH&A, USINGEQUATION15-12 REDUCEFLOWOF PERMEATEBY FLOWTHROUGH&A, USINGEQUATION15-13 REPEAl"FOR EACHCOMPONENT REPEATFOR EACHCOMPONENT CALCULATENEW Y.RESIDUEFROM (EQUALSY,= TO NEXT~ uz.), EQUATION15-10 NO~ - -m m = ~ _ ARE SPECSMET? INCREASE~ BY NEW INCREMENTOF AREA J ~r ou 4 RESULTS J (ENO) I USE NEW Y, FEED(Yi RESIDUEFROMPREVIOUSAREA) AND INITIALIZEWITH fiNAL T..M,FROMPREVIOUSITERATION F ig u re15-10 . Calculation f l owsheetf orcomputersimulationof membranepermeation. Membr anePermeat i onProcesses1257 Table 15-4 Computer Simulation of C02 Recovery Using a Single-Stage, Spiral-Wound Cellulose Acetate Membrane Unit FeedResi duePermeate MOLl MOLMOLl MOLMOLl MOL Component HRFRAC. HRFRAC. HRFRAC.Carbon dioxide735.58.90588744.29.399763691.29.985849 Nitrogen1 . 3 2 .0016261 . 1 2 .010074.20.000291 Methane40.24.04955733.17.2994007.07.010084 Ethane11.74.01445810. 65.0961031 . 0 9 .001559 Hydrogen sulfide.10.000123.00.000006.10.000142 Propane12. 41.01528311. 6 3 .104938.78.001119 Isobutane2.28.0028082.14.019279.14.000206 Butane4 . 7 1 .0058004 . 4 1 .039827.30.000425 Isopentane1 . 4 1 .0017361 . 3 2 .011923.09.000127 Pentane1.34.0016501 . 2 6 .011331.08.000121 Hexane.46.000567.43.003890.03.000041 Heptane.41.000505.38.003467.03.000037 TOTAL812.00110.78701.22 Area =13,110.9Square feet,125Increments used Note: Operating Conditions:Feed at 500.0 psia,140.0 degreesF, permeate100.0 psia. Source:Schendel(1995) Thepermeationrate,R i,isnotnecessarilyaconstant.Forexample,itisknowntovary withtemperatureandpressure,andcorrectionsforthesefactorsarereadilyincludedinthe calculations.Becausechangesincompositionfromincrementtoincrementarenormally small,evencomponent interferenceeffectscanbeincludedin thecalculationsif required. Gascompositioneffectscanbesignificant.Theconcentrationof carbondioxideinnatural gas,forexample,canaffectitsrateofpermeation(perunitofpartialpressuredifferential). DataprovidedbyHogsettandMazur(1983)indicatethatincreasingtheCO2concentration from15to50%in200psianaturalgasincreasestheCO2permeationratefromabout4.0to 5.0(scf)/(ftZ)(hr)(100psi)withGASEPmembranes. HogsettandMazur(1983)alsosuggestasimplifiedapproachforestimatingtheapproxi- matemembranearearequiredforamulticomponent system. Theapproachavoidsthe need to useacomputersimulationmodel,butisreportedlyaccuratetoonlyabout_+ 20%of theactu- alrequiredarea.Themethodisbasedonthefollowingsteps: 1)Sortallcomponentsintotwogroups(fastandslowpermeators),withthefastpermeators typically havingpermeationrates15timesthoseof theslowpermeators. 2)Calculateweightedaveragepermeation ratesforthetwogroups. 3)Useequationsdevelopedfort wo-component systemstocalculatethemembranearea requiredtomeet fastpermeator removal requirements. 1258Gas Purification Therequiredtwo-componentequationsanddetailsof thecalculationprocedurearegiven by Hogsettand Mazur(1983). Ettouneyetal.(1995)comparedtwo-componentandfour-componentmodelsforthe analysisofnaturalgaswellenrichment.Intheflowsystemevaluated,gasfromthewellis recycledthroughapermeationunitandbacktothewell.RemovalofaCO2andHzS-rich permeateresultsinanincreaseinpurityofthewellcontentswithtime.Resultsofthestudy showedthatproperdesignofthissystemrequiresthatthemodelstakeintoconsideration flowpatterns,thepresenceofmorethanonespecies,andpermeabilityratefunctionsthat includetheeffectsof bothcompositionand pressure.Theuseof asimplifiedtwo-component modelwithconstant permeabilitiesgave largedeviationsfrom the moredetailed models. Fields ofApplication Membranesystemsarenowavailablethatareeconomicallyattractiveformanyapplica- tions,andthefieldsof applicationaregrowingsteadily.Themaincurrentcommercialappli- cationsof membrane-basedgaspurificationandseparationare: 1) Hydrogen recovery from nitrogen-bearinggases,e.g.,ammoniasynthesis purgegas 2)Hydrogenremovalandrecoveryfromhydrocarbons(e.g.,methane)andotherslowerper- meatinggases(e.g.,carbonmonoxide) 3)Removalofcarbondioxide,hydrogensulfide,andwatervaporfrommethaneandother hydrocarbongases(e.g.,upgradingnaturalgastomeet pipelinespecifications) 4)Oxygenor nitrogenseparationfromair (e.g.,producingnitrogenfor inert blanketing) 5)Helium removaland recovery from naturalgas 6)Solvent vapor removal from exhaust gases Thesmallmolecularsizeofhydrogenandheliumallowthemtodiffuserapidlythrough membranes.Therefore,theseelementsarereadilyseparablefromlargermoleculegasessuch asmethaneandheavierhydrocarbons.Theacidgases,carbondioxideandhydrogensulfide, andwaterhavelargermoleculesanddiffusemoreslowlythanhydrogenorhelium.Howev- er,theyaremuchmoresolubleinpolymersusedforthemanufactureofmembranes.Since permeabilityistheproductofsolubilityanddiffusivity,highlysolublegasescanhaveper- meationratescomparabletothoseof lightergases;whilegasessuchasmethane,whichhas botha relativelylarge moleculeand lowsolubility,have low permeation rates. Theseparationofoxygenandnitrogenisdifficultbecausethesizeandshape(andhence thediffusivity)ofthemoleculesarequitesimilar.Inaddition,thesolubilityanddiffusivity of thefastergas(oxygen)aregenerallyquitelow,resultinginalowrateof permeationand the requirement for a largemembranearea.However,becauseof the industrialimportanceof thisseparationandthescarcityofsimplealternatives,ithasbeenthesubjectofextensive researchanddevelopmentwork.Inaddition,thefeedstockis"free."Therefore,highrecov- eryisnotarequirementforthisseparation.Lowrecoveryoperationcanbeusedtoimprove theseparationefficiency(i.e.,the permeatepurity). Organicsolventstypicallyexhibitlowdiffusivityratesbutveryhighsolubilitiesinappro- priatemembranes.Asaresult,satisfactorypermeationratescanbeobtainedrelativetothe othercomponentsinexhaustgases.Thepartialpressuredifferentialcanbeimprovedby operatingthepermeatesideundervacuum.Thisisnormallymoreeconomicalthancom- pressinglargevolumesofatmosphericpressureexhaustair.However,theprocessisnot effectiveintheextremelylowpartialpressurerangerequiredforremovingtracesof organic MembranePermeationProcesses1259 solventsfromair,andisusedprimarilyforthebulkremovalofsolventsfromrelativelycon- centratedexhaustairstreams. APPLICATION CASE STUDIES Hydrogen Ammonia Synthesis Purge Ammoni a isproducedbyreactinghydrogenwithnitrogenoveracatalyst.Thehydrogenis usuallyproducedinasteam-methanereformer(SMR)andthenitrogencomesfromtheair suppliedtoasecondaryreformer.Notallthemethane(naturalgas)isconvertedtohydrogen intheSMRnoristheconversionof nitrogenandhydrogentoammoniacompleteinthesyn- thesisreactor.Thereactorproductgases,afterremovaloftheammonia,arerecycledbackto thereactorfeedtoimproveyields.Inertgases,suchasargonfromtheairandmethane,build upinthisclosedloopandreducethenitrogen-hydrogenreactionrate.Therefore,acontinu- ousgaspurgetothefuelgassystemismaintainedtokeeptheseinertsatamanageablelevel. However,thispurgegascontainsvaluablehydrogen.Currentlymanyammoniaplantsrecov- erhydrogenfromthepurgestream.Thisisapopularapplicationformembranesbecauseof thehighpermeationrateof hydrogenandthehigh pressureof thepurgegas. ShirleyandBorzik(1982)reportedontheinstallationofahydrogenrecoverymembrane systemina1,000tonperdayammoniaplantthatresultedina5%overallcapacityincrease. MacLeanetal.(1988)reportedtheinstallationofamembraneunitina600tonperday ammonia plant.Theunitprovidedan89%hydrogenstreamwithan86%hydrogenrecovery, representinga 4%ammonia plantcapacityincrease. Acomparisonofmembraneseparationversuscryogenicseparationinatypicallarge ammoni aplantwasmadebySchendeletal.(1983).Forthecasestudydescribedinthis paper,theoperat i ngcondi t i onsweremodi fi edtoi ncreasethemet hanecont ent atthe entrancetothesynthesislooptoaboutdoublethatallowedwithouthydrogenrecovery.The pertinentprocessvariablesforthe15MMscfdfeedstreamaresummarizedinTabl e15-5. Topreventdensificationofthemembraneorformationofaninsolublephaseinthecryo- genicsystem,theammoniainthefeedtobothsystemsisreducedtoverylowlevelsina waterscrubber.Topreventformationofasolidphaseinthecryogenicunit,molecularsieves areusedtoremove thewaterpickedupinthescrubber. Theinertsconcentrationinthesynthesisloopisheldconstant.Sincethecryogenicsystem producesaslightlypurerrecyclehydrogenstream,thepurgeraterequiredtoholdtheinerts atafixedlevelislessinthissystem.Hydrogenpermeatesthroughthemembraneandis recoveredasalowpressureproduct.Therecompressionrequirementsofthemembranesys- temare,therefore,considerablygreaterthanthoseofthecryogenicunit.Toreducethe recompressioncosts,themembraneunit isoperatedin twostages. Theutility requirementsarealsopresentedinTabl e15-5.Theexternalpowerrequirement forthecompressorhorsepoweristhatrequiredtointegratetherecoveryunitintothetotal plant.Theprimaryadvantageofthecryogenicunitisthel owerrecyclerecompressi on requirements.Thehigheron-skidelectricalcostsforthecryogenicunit reflectstheuseof the molecularsievetodrythefeedstreamtotherecoveryunit.Thecryogenicunitalsorequires asmall purgestreamof nitrogen. Theestimatedcapitalequipmentcostsarevirtuallyidentical,approximately$1.35million (1983dollars).Themembraneunitconsistsoftwoskids,whilethecryogenicunitrequires 1260GasPuri f i cat i on Table 15-5 Comparison of Membrane and Cryogenic Separation Units forHydrogen Recovery in an Ammonia Plant Cryogenic MembraneType, Type,Petrocarbon MonsantoPRISMS-2000 Feed Composition,mole% H 260.860.8 N220.020.0 CH 412.112.2 Ar3.23.1 NH33.93.9 Feed GasQuantity,lbmoles/hr1,7671,503 PressureatSeparatorInlet,psig1,9731,000 Hydrogen Recovery,%95.794.6 Hydrogen Purity,mole%87.892.5 Ammonia Recovery,%99.898.7 RecycleProductto HighStage Compressor,%49.8100 RecycleProductto LowStageCompressor,%50.2 Electricity,kWh~3080 Coolingwater,gpm200225 Steam(600psig),lbs/hr1,9101,760 Nitrogen,scfhStartuponly180 Instrument air,scfh2,1001,800 Turbinecondensate,gpm12Minor make-up Externalpower differential,kWh/h470 Source:Schendelet al.(1983) four.Wheninstallationandmaintenancecostsarefactoredin,thetwoprocessesareconsid- ered to becompetitive. Oxo-alcohol Synthesis Gas Intheproductionof oxo-alcohols,carbonmonoxideisreactedwithhydrogenataone-to- oneratiotoformanaldehyde.Thealdehydeisthenreactedwithpurehydrogentoformthe desiredoxo-alcoholproduct.Thehydrogenandcarbonmonoxidesynthesisgasismadeby steamreformingofnaturalgasorbypartialoxidationofhydrocarbons.Therawsynthesis gashasahydrogentocarbonmonoxideratiorangeof3:1to2:1.Beforethesynthesisgas canbeused,theratioof hydrogen tocarbonmonoxidemustbeadjusted.Cryogenics,molec- ular-sieves,pressureswingadsorption(PSA),andmembranesareusedtomaketheratio adjustment.Thefollowingcasehistoriesindicatehowmembraneshavebeenusedto increasecapacityand/or efficiencyinachievingthecorrectratioof hydrogen/carbonmonox- ideforsynthesis. MembranePermeationProcesses1261 MacLeanandGraham(1980)reportedonamembranesystemusedtodebottleneck Mon- santo' sTexasCityplantwherethecoldboxdidnothaveenoughcapacitytosupplycarbon monoxidefeedstockforaceticacidproduction.Thecoldboxhadthreeproductstreams,a purecarbonmonoxidestreamforaceticacid,a1.3:1hydrogen/carbonmonoxidestreamfor oxo-alcohols,andahydrogenstreamformethanol.Thesolutionwastoinstallamembrane separationsysteminparallelwiththecoldbox.The3.1:1H2/COfeedwassplitbetweenthe membraneunitandthecoldbox.Thecoldboxwasthenoperatedtoproducehydrogenand thepurecarbonmonoxiderequiredforaceticacidproduction.Themembraneunitproduced the1.3:1Hz/COratiostreamrequiredforoxo-alcoholproductionwhilerecovering93%of thecarbonmonoxide.Thepermeatestream,consisting of 96%purehydrogen,wasfedto the methanolplant.Asanindicationofmembraneaging,theseparationefficiencydeclinedless than10%duringthe first three years of operation. AmembraneunitwasalsoinstalledontheTexasCitymethanolplantpurgegasstream (BurmasterandCarter,1983).Themembraneunitrecoveredaboutonehalf of thehydrogen andcarbondioxidethatwaspreviouslysenttothefuelgassystem.Themethanolplant capacitywasincreasedby2.6%,andexceptforcompressorlimitations,couldhavebeen increased by3.9%. Theintegrationofamembranesystemwithpressureswingadsorption(PSA)atanalde- hydesynthesisplantwasstudiedbyDoshietal.(1989).Synthesisgasratioadjustment usingPSAaloneproduceshighpurityhydrogenaswellasthepropersyngasratio.Howev- er,thedrawbackofthesystemistherelativelyexpensivetailgascompressi onthatis required.PSAtailgasisproducedatlowpressureandrequiresalargecompressortoboost thepressuretothealdehydesynthesispressure.Theuseofamembranesystemalonewill producesyngasattheproperhydrogen/carbonmonoxideratio,butthepurityofthehydro- genpermeatestreamislow.Thesyngasproducedisatapproximatelythesamepressureas thefeedgasandcanbeusedforaldehydesynthesiswithoutadditionalcompression.The hydrogenpermeatestreamgenerally hastobeupgraded,whichusuallyresultsinthelossof carbonmonoxide fromthesystem. Anintegratedsystem thatusesbothmembranesandPSAontheproductgasfromapartial oxidation(POX)unit,minimizesthedrawbacksof either system whenit isusedalone.Inthe integratedsystem,Fi gur e15-11,thefeedgasentersthemembraneunitwheretheresidue stream,enrichedincarbonmonoxide,isproducedatapproximately feedpressure.Thelower pressurehydrogenpermeatestreamisfedtoaPSAunitwherethecarbonmonoxideis adsorbedandahighpurityhydrogenstreamisproduced.Thelow-volume,CO-richPSA tailgasstream iscompressedandcombinedwiththeresiduestreamfromthemembraneunit. Thestudyshowstheintegratedsystemtohavelowercapitalandoperatingcoststhaneither oftheindividualsystemsandtoprovide100%recoveryofhydrogenandcarbonmonoxide. Aneconomiccomparisonofastand-alonePSAandanintegratedmembrane-PSAsystemis presented in Tabl e15-6. Catalytic Reforming Catalyticreforming isaprocessinwhichhydrocarbon molecules arestructurally rearranged to higher octane forms.The reforming processisanet producer of hydrogen that,if recovered, canbeusedinhydroprocessing.Thefollowing casehistory of ademonstration membranesys- temfortherecoveryofhydrogenfromacatalyticreformerunit(CRU)off-gaswasdescribed by Yamashiro et al.(1985). 1262GasPuri f i cat i on Natural Gas + Oxygen + Steam CO= TUnoptimizedRecycleH=:CO Ratio ='-. . . . /Syngas Ir (Optional)L r ~H=:CO = 1:1 Enriched H= Enriched CO Product Figure 15-11.Integrated membrane plus PSA system in anatural gaspartial oxidation (POX) plant. (D oshi et al.,1989) Tabl e 15-6 ComparativeEconomics ofPSA A l one vs.aMembrane/PSAIntegratedSystem forProducing Hydrogen and Aldehyde Syngas Basis: RawFeed Rate,MMscfd Synthesis GasRate,MMscfd Hydrogen Product Rate,MMscfd RawH2/COGasPressure, psig Aldehyde Reaction Pressure, psig Hydrogen Use Pressure, psig 20.0 14.4 5.6 420 400 150 COMPOSITIONS,RAW GASH 2SYN. MOL- %H2/COPRODUCTGAS H263.499.99949.1 CO35.410ppm49.1 Ar/N 20.4trace0.6 CH 40.8trace1.2 CO2tracetracetrace H20Sat'dDrySat'd PSAONLY MEMBRANE + PSA Compression Required,BHP Separation Equipment Cost,$MMUS Installation Cost,$MMUS Installed Compressor Cost,$MMUS Compressor Operating Cost,$MMUS (3 yr-8000 hr/yr-5r TotalCapitalCost +3 yrs Operation,($MMUS) 1080 1.425 0.175 0.864 0.966 3.430 415 1.375 0.225 0.332 0.371 2.303 Source." Doshiet al.(1989) MembranePermeationProcesses1263 The850MscfdunitwasinstalledintheNo.2ReformerattheCosmoOilRefineryin Chiba,Japan.TheprocessflowfortheunitisshowninFi gur e15-12.Thereformereffluent iscooled,theliquidandgasseparated,andthegasfedtoanabsorbertoremovetheheavy component s. Gasfromtheabsorber,containingapproxi mat el y80%hydrogenand20% methaneandsaturatedwithabsorptionoilcomponentsatadewpointof95~andapressure of398psia,isfedtoafilterseparatortoremoveanyresidualliquids.Itisthenpreheatedtoa temperatureabovethehydrocarbondewpointoftheresidualgasleavingthemembraneunit andfedtothemembraneseparator.Heatingthegaspreventscondensationoftheheavy hydrocarbonsasthegasdewpointincreaseswithhydrogenremoval, The97%hydrogenpermeategasisofsufficientpurityandpressurethatitcanbefed directlyintothe256psiahydrogensupplysystem.Typicaloperatingconditionsareprovided inTabl e15-7.Theunitwasoperatedatalowpressureratiotodeliver hydrogentotherefin- eryhydrogensupplysystem withoutrecompression.Thislowpressureratioresultedinalow hydrogenrecoveryof about30%. Duringstartup,thetemperatureof thefeedwasvariedfrom104 ~ to180~Athighertem- peraturesthegaspermeationrateincreasedbymorethanafactoroftwo.Acorresponding decreaseofmembraneselectivitywasalsonoted,butwasnotgreatenoughtoalterthesys- temperformancesignificantly.Thepermeategaspressurewasalsovariedwhilemaintaining constantfeedpressure.Itwasshownthatmembraneselectivityandhydrogenpermeation rateareindependentofdifferentialpressureandpressureratiofortherangeofconditions studied(SchellandHouston,1985). Full PI1BIl(I~ 6r ~ F'd~ersepatat~--P,ecyc~R e ~To turth=r F ig u re15-12. Fl owdi agramofref ormersystemwi t hmembr anehydrogenrecoveryu nitonof f g as( Y amashiro et a/.,1985).Reproduced with permissionfrom Hydrocarbon Processing,February 1985 1264GasPurif ication Table 15-7Operating Conditions for Catalytic Reformer Offgas Hydrogen Recovery Membrane System FeedPermeat eResi due Flow,Mscfd850 Composition,vol%: H280 CH 420 DewPoint,~95 Inlet Pressure,psig380 Inlet Temperature,~140 Hydrogen Recovery,%30 9774 326 260 Source:Yamashiro et al., 1985 Another exampleof membranesbeingused for the recovery of hydrogen fromaCRUwas reportedbyLane(1983).TheunitfeedwastheoffgasfromaCRUthatformerlywenttoa hydrogenplant.Thefeedgasflowratewasnominally4MMscfd,butvariedfrom2to10 MMscfd.Themembraneunittypicallyproduced98%hydrogenat250psigwith36%recov- ery.ThesystemoperatedoverfullCRUcycles,fromstart-of-runtoend-of-run,which causedthehydrogencontentofthefeedtorangefrom62to87%.Themembraneunitper- cent hydrogen recovery and product hydrogen purityremained relatively constant. Tolessenscaleuprisks,thesmallestcommercial-sizemembraneunitswereinstalledeven thoughtheywereundersizedfortheflow.Thisresultedinthelow36%hydrogenrecovery. Thepredictedrecoveryforthepropersizemembranesystemis93%withahydrogenpurity of 94%. Hydroprocessing Hydrotreating,hydrodesulfurization,andhydrocrackingareoperationsinwhichhydrogen isusedtosaturatetheolefinsinahydrocarbonstream;removeobjectionableelementssuch assulfur,nitrogen,oxygen,halides,andtracemetals;andcracklargerhydrocarbonmole- culesintosmallerones.Intheseprocesses,freshhydrogenisfedtothereactor,anduncon- sumedhydrogenisseparatedfromthereactoreffluentandrecycledbacktothereactor.A portionof therecycledhydrogenisoftenpurgedfromthesystemtopreventthebuildupof light hydrocarbonsand inertsthat would lower the hydrogen partialpressurein the reactor. H y d r ot r eat i ng . AtConoco' sPoncaCityOklahomarefinery,a71mole%hydrogenhigh- pressure.purgegasstreamfromagas-oilhydrotreaterwassplittofeedthelight-cycleoil hydrodesulfurizerandthecryogenicliquifiedpetroleumgas(LPG)recoveryunit.Thepurge streamwasusedonaonce-throughbasisinbothunitsandthendischargedtothefuelgas system.Theinstallationof a membrane-basedhydrogen recovery unitto producehigh purity hydrogenfromthispurgestreamwasdescribedbyShaveretal.(1991).Aschematicofthe systemispresentedinFigure15-13.Themembraneunitisdesignedtoproduce7MMscfd of95mole%hydrogenfromthe12MMscfdhydrotreaterhighpressurepurgewitha75% MembranePermeat i onProcesses1265 Hy d r og e nr e cy cl eHy d r og e nI~~ - Y_~ J. ~ . . ~ --~ G O H D T-I Ma k e u p ! c omp r e s s orMe mbr a nesy st e m J , oF u e l g a s C r y og e ni c[un.J Figure 15-13. F l ow diagram of membrane system installed topurify gas oil hydrotreater (GOHDT) offgas-to provide hydrogen for light-cycle oilhydrodesulfurizer (LCOHDS) (Shaver et al.,1991). Reproduced with permission from Hydrocarbon Processing, J une 1991 hydrogen recovery. Thehigh-puritypermeate hydrogenissent tothehydrodesulfurizer.The available pressure dropfrom1,050 psigat thehydrotreater to 430psigat thehydrodesulfur- izerprovidesthedrivingforceforthemembraneseparation.Becauseofthehighpurityof thispermeatehydrogen,thehydrodesulfurizeroffgasstreamisstill90pluspercenthydro- gen,andcanbe recycled back to thegas oilhydrotreater. Theresiduestream from themem- braneunitisfedtothecryogenicunitforrecoveryof LPG.Thepretreatment of thefeedto themembraneunitconsistsofaknockoutdrum,afeedpreheater,andadrygasfilter.The economics of thissystem is presented in Table15-8and indicates a1.7-year payback period. Table15-8 Economics ofMembraneHydrogen Recovery Sy stem onHydrotreaterOffgas Investment, $662,000 Debits, $/yr Steam consumption22,000 Lost hydrogen fuel value75,000 Maintenance and overhead29,000 Total126,000 Credits,$/yr Gas oil HDT product upgrade396,000 LCOHDSproduct upgrade74,000 Reduced power consumption42,000 Total512,000 Earnings,$/yr386,000 Simple payback period, yr1.7 Source:Shaver et al.(1991) 1266Gas Purification Operatingdatacollectedovertwoyearsshowconsistentrecoveryratesandpermeatepurity evenwithfeed ratevariation from8 to20MMscfd. Schendeletal.(1983)comparedmembraneandPSAtechnologiesforhydrogenrecovery fromhydrotreaterpurgegas.A7MMscfdpurgestreamat800psigand100~with72% hydrogenwasassumedtobethefeedtothehydrogenrecovery units.Theproducthydrogen couldbereturnedtoeither the250psigmake-upcompressor firststagesuctionortothe450 psiginter-stagesuction.The residuegasisused for fuelgas. For themembraneseparator,a93%hydrogen productisproducedat250psigwithan81% hydrogenrecovery.Theresiduegasisletdowninpressureandfedintothe60psigfuelgas system.Thefeedgastothemembraneunitispreheatedtoavoidcondensation.Fi gure15-14 showstheintegrationofahydrogenrecoveryunitintoatypicalhydrotreaterunitwiththe recovered hydrogen returned to thesuctionof the low-pressure hydrogen makeupcompressor. Pressureswingadsorption(PSA)utilizesmolecular sieves toselectively remove hydrocar- bonsandother impuritiesto produceahighpurityhydrogenstream.Thegreater the pressure swing(betweenthehighpressureofadsorptionandthelowpressureofdesorption),the greater theunit capacityand product recovery. ThePSAsystemevaluationwasbasedontwodifferentoperatingscenarios:productat 450psigandresiduegasat60psig,andproductat250psigandresiduegasat5psig.Inthe 450psigproductcase,thehydrogenisreturnedtothecompressorinterstagesuctionandthe residuegasisfedtothe60psigfuelgassystem.Inthe250psigproductcase,thehydrogen isreturnedtothe250psigcompressorsuctionandtheresiduegasissenttoalowpressure burner.Thehydrogen recovery wasmuchgreater for thelow pressure case. Theoperatingconditionsandtheoperatingandcapitalcostsforthethreecasesarepre- sentedinTabl e15-9.Ascanbeseenfromthecostdata,themembranesystemshowsa Residue ."'"ec~176i H2=,=~ alMake-up""~ ~ ~ ~ ~I Make-upRecycle HydrocarbonCompm~o~Compressor Feed-Hydrotreating "7Reactor1 T . PS.p,to,) I 9=LowPressure vPurge T o '~ Fractionation F ig u re15- 14. Integrationofhy d rog en recoveryunitinto typicalhy d rotreatersystem. ($ c h en d e/ etal . , 1 983 )Membr anePermeat i onProcesses1 2 6 7Table 15-9 Operating Conditions and Economicsfor Hydrogen Recoveryfrom Hydrotreater Purge Gas by Membrane and PSA Systems MembraneP SAP SA WasteGas Pressure,psig(1)605 Feed H2,%727272 Feed Pressure,psig800800800 Feed Temperature,~100100100 Feed Flow Rate,MMscfd777 HighPurityH 2, %9399.599.9 HighPurity H 2, psig250450250 HighPurityH 2, MMscfd4.423.044.05 WasteGas H2,%365134 WasteGas Flow Rate,MMscfd2.583.962.95 H2 Recovery,%816080 CapitalCosts,M$ Equipment5301,050875 Installation100175150 TotalCost6301,2251,025 OperatingCosts,M$/yr (2) Steam for Preheat10w Compression to Reactor Press.14046130 Contrib.to H2 Cost,$/Mscf0.110.050.09 Total H 2 Cost,$/Mscf (3)0.200.290.24 Notes: 1. Membranesystemresiduegas producedat operating pressureand,after pressurereduction, f ed intothe f uel gassystem. 2.Utili~costsbased on5c/kwhand$5/MM Btu. 3.Hecost based on8,000 hr/yr operation f ora f ive-yearspan. Source:Schendelet al.(1983) lower capitalcostand lower total hydrogen cost thaneither of thePSAcases.Themembrane andthelow-pressurePSAcaseincurasignificantcostforrecompressingthehydrogento reactor pressure.Thelow-pressure PSAcasehasthebetter economicsof thetwoPSAcases. However,itmaybedifficulttofindauseforthe5psigwastegas.Also,dependingonthe sensitivityof thehydrotreater tohydrogenpartialpressure,thehigherpurityofthePSAgas mayhavedefiniteadvantages. Tonen Technology K.K.and UBE Industries (1990) reported on the installation of a hydrogen recovery facility at the Wakayama refinery of Tonen Company, Ltd.Polyimide membrane mod- ules,arranged in eight rowsof twotrains each,weredesigned to produce5,153scfmof product hydrogen withaminimumpurityof 95%.Thefeedisfiltered toremove hydrocarbonmistand then preheated to160~176prior to entering the modules. The hydrogen-rich permeate is col- lectedandfedtothehydrogensystem.Theresiduegasiscollected,cooledto150~anddis- charged to the fuel gassystem. Plant operating data are presented in Tabl e15-10. 1268Gas Purification Table 15-10 Operating Data for Hydrogen RecoveryMembrane System at Wakayama Refinery FeedProductOffgas FlowRate,scfm15,49010,0955,395 Pressure, psig343124102 Temperature,~9015599 Composition, mol% H277.698.239.1 CI17.01.645.8 C24.40.212.2 C30.3.09 C4+0.720 H2S, ppm400300590 BTX,ppm300 Mw6.302.313.81 H2 Recovery, %82.5 Source: Tonen Technology and UBE Industries (1990) Hydrocracki ng. Hydrocrackerstypicallyoperateathigherpressuresthanhydrotreatersor hydrodesulfurization(HDS)units.Bollingeretal.(1984)performedastudytooptimize hydrogenrecoveryfromhydrocrackerpurgegasstreams.Variousmembraneseparation operating optionswerestudied.Operating optionsincludedconstant recycle purity,constant purge rate, constant make-up compressor horsepower, and constant hydrogen make-up rate. Theoptimizedsystem,whichincludestherecoveryof hydrogenfromboththehigh-and low-pressure purgestreams, isshowninFigure15-15,whichincludesmaterial balance data for thesystem. Asindicated in the material balance, hydrogen is consumed during the hydro- cracking reaction(chemical hydrogen)andsome of theunreactedhydrogenisdischarged in purgestreams from thehigh-andlow-pressure separators.Theoptimized design depicted in Figure 15-15recovershydrogenfromthetwopurgestreamstherebyminimizinghydrogen losses.Whencomparedtothehydrocrackerwithoutmembraneunits,theoptimizedsystem results in a10%increase in the hydrogen partial pressure of the recycle gas leaving the high- pressureseparator,aslightreductioninmake-uphydrogen,andanincreasedchemical hydrogen consumption(26,7to31.8MMscfd).Assuming thatchemical hydrogen consump- tion per barrel of feed remains constant, the increase in chemical hydrogen consumption cor- respondstoanincreaseinhydrocrackerthroughputof19%. Whilethemake-uphydrogen flowfor the optimized system decreases from 40to38.7MMscfd,the totalflowthroughthe make-up hydrogen compressor increases from 40to59.0MMscfdasbothpermeate streams must be compressed in addition to the make-up hydrogen. " Bu t amer " Of f gas. In the UOP licensed "Butamer" process, normal butane is catalytically isomerizedtoisobutane.Theprocessproducesisobutaneandhydrogenstreams.Evenwith highhydrogenrecycle rates,somefeedstockiscrackedintomethane,ethane,andpropane. To maintain high reactor efficiency, some of the recycle gas is purged,usually to the fuel gas MembranePermeationProcesses1269 MembraneSeparators Reactor High Pressu re MembraneSeparators Stream( ~ ( ~)( ~)~( ~ ( ~ ( ~) Pressu re,psig250180027 0145045022050 Flowrate,MMSCF D38. 7 18.910.24. 414.54. 45.8 Composition,mol e% Ha88.582.061.934. 896.521. 891.7Cl 7 . 012.221.143. 32.642. 35. 3 Ca3.14.610.017.10.7 20. 91.8 C3+1.41.27 . 04. 80.215.01.2 Figure 15-15.Optimized f l ow scheme for recovery of hydrogen f rom hydrocracker purge streams (B ollinger et a/., 1984). Reproduced with permission from Chemical E ngineering Progress, copyright 1984, A merican Institute of Chemical E ngineers system.Thisstreamcontainsnotonlyhydrogenandthecrackingproducts,butalsosome butanes. Membraneshave beenconsidered fortherecovery of hydrogen fromButamer units.How- ever,tomaintaincatalystactivity,smallamountsofanorganicchlorideareintroducedinto thefeedstream.ThechloridesareconvertedtoHC1bythecatalyst,andthepurgegasfrom theButamerunitcontainstracesofHC1. TheeffectofHC1onmembranesisaconcern.A simplifiedprocessschematicfora"Butamer"unitwithhydrogenrecoveryisshowninFig- ure15-16.Themembraneunitislocateddownstreamofacausticwashusedtoremove HC1 fromthepurgegas.SchellandHouston(1982)described theintegrationofamembraneunit witha"Butamer"unit,designed toprocess47,800scfhof feedgas.Theflowratesandpuri- tiesareshowninTable15-11.Theyreport thatunder bone-dryfeed conditions,thecellulose acetatemembrane,whichwaslocatedupstreamofthecausticwashinthisplant,wasnot affectedbyHC1; however,specialmaterialsandadhesiveswereusedtoensureresistanceto theHC1. CooleyandDethloff(1985)reportedonademonstrationunitinstalledinitially upstreamofthecausticwashunit.TheyfoundthatHC1concentrationsof2,000and4,000 ppminthepurgegasimpairedthemembraneperformance.Theunitwasrelocateddown- streamofthecausticwashasshowninFi gure15-16andtheperformanceimproved.The unitwasdesigned fora feed rateof 971scfhat 295psigwith61%hydrogen inthe feedgas. 1270GasPurif ication Table 15-11 Operating Data for Hydrogen Recoveryfrom the Butamer Process FeedResi dual Permeat e GasGasGas Pressure,psia265 FlowRate,Mscfh47.8 Temperature,~110 Composition,mol% H 268.9 C123.7 C2+6.8 HC10.6 24015 16.831.0 100100 17.896.4 63.02.6 19.00.2 0.20.8 Source." Schelland Houston(1982) Make- u p( Hy d r og e nJ" N- Bu t ane Recy cl e Bu t amerUnit Recov er ed Hy d r og en / Cau st ic W ash Isobu t ane Pr od u ct/L oc,. Membr aneRecov er yUnit Figure 15-16. Membrane system for recovering hydrogenfrom Butamer system purge gas. (Cooleyand Oethloff,1985) Other HydrogenApplications MacLeanandNarayan(1982)havedescribedotherapplicationsof membranesystemsfor hydrogenseparation,includingtoluene hydrodealkylationandcoalliquefactionprocesses. CarbonDiox id e, HydrogenSulfide,and W aterRemoval E nhancedOil Recov ery(E OR) WhenCO2isinjectedintoanoilreservoiratsufficientpressure,itdissolvesintheoilpre- sentinthesubstratareducingitsviscosity,allowingittoflowmorefreely,andthereby increasingoilproduction.Whentheoilisbroughttothesurfaceanditspressurereduced,the injectedCO2isreleasedfrom theoilanddischargedwiththeassociatedgas. MembranePermeationProcesses1271 InCO 2floodEORprojects,theCO 2istypicallyrecoveredfromtheassociatedgasand recycled backintotheoilproducingformation.WiththiscontinualrecycleofCO2,boththe volumeandCO2contentoftheassociatedgasprogressivelyincrease.InthedesignofCO2 floodEORsystems,theobjectiveistominimizecapitalexpenditurewhentheassociatedgas volumeandCO2contentarelow,buthaveenoughdesignflexibilitysothatthesystemwill beoperable inthefuturewhentheassociatedgasvolumeandCO2contentarehigh.Inmany ways,themodular natureof membranesmakesthem ideally suited for thisapplication.In the initialphasesoftheEORproject,capitalcostscanbeminimizedbyaddingtheminimum numberofmembranemodulesandCO2re-injectioncompressors.Additionalmembrane modulesandCO2compressioncanbeaddedlaterwhentheassociatedgasvolumeandCO2 content are higher.Other CO2recovery processes donot havethisflexibility. ThefirstcommercialscalemembraneinstallationusedwithCO2floodEORwasthe SacrocprojectinWestTexas.InjectionofCO2intothefieldatvolumesupto200MMscfd beganin1972.TohandletheanticipatedincreaseintheassociatedgasCO2concentration, SacrocinstalledthreeCO2-removalfacilities.Theplantswereinstalledinconjunctionwith threeexistingprocessingfacilitiesoperatedbySunExploration&ProductionCo.,Chevron U.S.A.,andMonsanto.TheSunandChevronfacilitiesusetheBenfieldhotpotassiumcar- bonateprocess,andtheMonsantofacilityusesthemonoethanolamine(MEA)process.The SunhotpotassiumcarbonateplantwasdesignedtoreducetheCO2contentof160MMscfd ofassociatedgasfrom24to0.5%CO2,whiletheChevronplantwasdesignedtoreducethe CO2concentrationof46MMscfdofassociatedgasfrom24to1.0%CO2.TheMonsanto MEAplantwasdesignedtotreat16.5MMscfdofthe24%CO2feedgas,reducingitsCO2 concentration to 0.01%(Parro,1984). Inthelate1970s,SacrocrealizedthattheCO2contentofthegasproducedfromthefield wouldpeakforafewyearsatlevels greater thantheCO2-removal plants' capacity.TheCO2 contentofthefieldgashadgraduallyincreasedfrom0.5%priortoinjectionto40%CO2. SacroccontractedwithTheCynaraCompanytobuildandoperatetwomembraneunits.The newmembraneunitswereinstalledandoperatedintegrallywiththeSunandChevronhot potassium carbonate plants. TheSunmembraneunitwasdesignedtorecover50MMscfdofCO2at520psigwithan allowable pressuredropof 40psig.TheChevronmembraneunitwasdesignedtorecover 20 MMscfdof CO2at 480psigwithanallowable pressuredropof 40psig.Themembranesepa- rationprocessflowisgiveninFi gur e15-17.Thetwoimportantfeaturesofthisdesignare thedehydrationof theinletgasandtheoperationof themembranesatreducedtemperatures. TheinletgasisfirstcooledbycrossexchangewiththeCO2andhydrocarbonproductgas streams.Thisreducesboththemoistureandhydrocarbon contentof thefeedgasandthesize ofthedehydrationequipment.Afterdehydration,thefeedgasiscooledbycrossexchange withtheresiduegasandfinally refrigerated beforebeingdirected tothemembranemodules. Thelow-temperature feedresultsinahigherseparationfactorandareducedvolumeofgas. Themembranesaresinglestageandarearrangedin parallel. Theamountoffeedgasprocesseddependsonthefieldproductionrate,andchangesin flowwerehandledbyincreasingordecreasingthenumberofmembranemodulesonline. Whentheplantwasshutdown,depressurized,andrestarted;thepermeategasflowwas foundtobe5to15%lowerthanpreviouslevels.Thepermeateflowgraduallyimproved overafewdaysof operation,butdidnotrecover entirely.Somelossof fluxremained(with aslightlybetterseparationfactor),andadditionalmoduleswererequiredtomaintainthe same permeate flowrate(Marquezand Hamaker1986). 1272Gas Purification Acid Gas Cross Ex ch a ng e Dehy d r at ionCrossI" ] = ~ e p a ~E x c h a n g e , I" " - Acid. GasIand Refrigeration Hydrocarbons,GastHydrocarbonGas "~(Residue) Figure15-17. Membrane system for treatinggas at Sacroc enhanced oilrecovery (EOR) projectin West Texas.( Cut l er an dJohnson, 1985) Additionalfieldtestdataontheuseofmembranepermeationtoprovideaconcentrated CO2streamforEORwerereportedbyRussell(1984)andMazurandChan(1982).Estimated capitalcostsforprocessinga800psig,20MMscfdfeedgasstreamcontainingvarying amountsof CO2weredevelopedbyCoadyandDavis(1982)andarepresentedin Fi gur e15- 18.Inallcases,the permeatestreamcontains5%hydrocarbonswhile theresiduegascontains 2%CO2.Toobtaina95%CO2permeatestream,atwo-stagemembranesystemwithinter- stagerecompressionisrequiredforfeedgasCO2concentrationslessthan75vol%.Atfeed gasCO2concentrationsgreaterthan75%,thesecond-stagemembraneandinterstagerecom- pressioncanbeeliminated. Lowpermeatepressureresultsintheleastamountofmembranearearequiredtoachieve thedesiredseparation.However,thisisattheexpenseof recompressionhorsepower required forreinjection.Marquez(1991)reportedonaseriesoftestsperformedtodeterminemem- braneperformancewithhigh-permeatebackpressures.Thetestscoveredfeedgaspressures from313to363psig(333psigaverage)andpermeatebackpressuresfrom189to250psig (221psigaverage).Thepermeatecompositionaveragedabout97%CO2and0.13%H2S, representingremovalofabout37%of theCO2and40%of theHzSfromthefeedgas.Itwas concludedthatpermeationintoahighbackpressuresystemisfeasibleandcanresultiscon- siderablerecompressioncostsavings.Newmembranemoduleswereusedforthetests.Per- meationratemeasurementsshowedthatthefluxstabilizedafteramonthof operationat79% of theinitialrate. InworksponsoredbytheU.S.DepartmentofEnergy(1989),thecostwasdevelopedfor amembraneunitinstalledupstreamofanamineunit.Themembraneunitwasdesignedto process170MMscfdoffeedgascontaining17%HzSand45%CO2.Itwasestimatedthat 280, 000ft 2 ofmembranewouldberequiredtoremove70%oftheacidgas.Ataninstalled firstcostof$20perft 2 ofmembrane,thecostoftheunitwouldbe$5.6million.Theesti- matedannualsteamsavingsintheamineplantwere$5millionto$10millionbasedon0.8 to1.6poundsofsteamperpoundofacidgasremovedandacostof$2.28per1,000pounds ofsteam.Thenetannualsavings,includingmembranereplacement,were$4.4millionto $9.4million. Goddin(1982)comparedseveralmethodsforrecoveringCO2fromaCO2-floodproject associatedgasstream.Inthisstudy,theassociatedgashydrocarbonandnitrogenflowrates wereheldconstantwhiletheCO2contentincreasedupto90vol%.Thissimulatesthe changeinassociatedgascompositionovertheEORprojectlife.ThefollowingCO2recovery caseswereevaluated: MembranePermeation Processes1 2 7 35 A o - - 4 m :E v i f )~ 3o O ii ==2 Q. m O 1 IIIIIII Compression El iminatedF eed F l owRat e =20 MMSCF / DI Tot a l F a ci l i t y n~ ~ ' - ~ I ( I ncl u d e sDe h y d r a t l o /Recompressl onITotaIFacil Membrane 10203040506 07 08090 CO= InF eed ,% Figure 15-18. Estimated capital costs for carbon dioxide recovery f rom a20MMscf/d gas stream ( Coady and D avis, 1982). Reproduced with permission from Chemica/ E ngineering Progress, copyright 1982, A merican Institute of Chemica/ E ngineers 1) Conventional Amine Theamineunitisbasedona30%DEAsolution.Thefeedgasiscompressedto285psia andpretreatedtoremoveheavyhydrocarbons.AfterpassingthroughtheDEAabsorber, thesweetoffgasiscompressedto650psiaandsenttotheexistinggasplant.Thesour CO2streamfromtheDEAstripper,at20psig,iscompressedto450psiaandsenttoa SelexolsweeteningprocessdesignedtoreducetheHzScontentoftheCO2productgasto lessthan100ppm. 2)Cryogenic Fractionation(Ryan-Holmes Process) Thefeedtothecryogenicunitiscompressedto625psia,dehydrated,andchilled.The overheadgasfromthefirstcryogeniccolumn,whichcontainsthemethaneandlighter components,issweetpipelinegas.ThebottomproductfromthiscolumnflowstotheCO2 column,wherethesweetCO2leavesintheoverhead.Leanoilisaddedtothetopofthe firstcolumntopreventtheCO2fromfreezing,andalsoisaddedtothetopofthesecond columntobreakupaCO2/ethaneazeotrope.ThebottomsfromtheCO2columnaresent toaleanoilrecoveryunit.ThepropaneandlightercomponentsareprocessedinaDEA unittoremovetheCO2andHzS.TheacidgasesremovedhereareprocessedinaClaus unitforsulfur recovery.Twocaseswereconsidered.CaseAassumesthatallof theethane recoveredhasthevalueofliquidhydrocarbon.CaseBassumesthatonlyaportionofthe ethanehasthehighliquidhydrocarbonvaluesince80%ofthehydrocarbonswouldbe recoveredintheexistinggasolineplant.Creditwasgivenforthedifferentialvalueof ethaneinnaturalgasliquids(NGL)versusfuelgas. 3)TEA BulkC02 Removal IntheTEAbulkremovalprocess,a TEAsolutionisusedtoremove theH2SandCO2 from thefeed.TheCO2andHzSarethenremovedfromtheTEAsolution byflashingit toabout 20psia.TheTEAabsorberoverheadstream,containingabout20%CO2andsomeHzS,is sent toa DEAunit forfinalcleanup.TheacidgasstreamsfromtheDEAunitandthe TEA flashtowerarecompressedto 450psiaandsent toaSelexol unit forHzSremoval. 1274GasPurification 4) MembranePermeation Themembraneunitwasdesignedtoproduceapermeatewithamaximum of 5%hydrocar- bonsandaresiduestreamcontaining20%CO2. TheresiduestreamissenttoaDEAunit for cleanup.Theacidgasfromthe DEAunit andthesour permeatestreamsarecompressed andsenttoaSelexolunitforH2Sremoval.Todeveloparangeofcosts,twocaseswere considered.Thelowcostcaseassumedashortmembranemodulelifeandahighperme- ationrate.Thehigh cost caseassumeda longmodule lifewithalower permeation rate. AsummaryofcapitalcostsversusfeedrateisgiveninTabl e15-12.Fi gures15-19and 15-20presentthecostofremovingCO2fortwodifferentsetsofutilitycosts.Alsoplotted arethecostcurvesforsweeteningtheCO2productstreamusingtheSelexolprocess.The costoftheSelexolprocessisincludedinthecurvesfortheDEA,TEA,andmembrane processes. Forthelowenergycostcase(Fi gure15-19),theleastcostsystem,overmostoftheCO2 concentrationrange,isthecryogenicprocesswithfullcreditforethanerecovery.Whenonly partialcreditistakenforethanerecovery,thecryogenicsystemisnotcompetitive.Mem- branepermeationismoreeconomicalthanDEAandTEAovertheentirerange,andhasan increasingadvantageover DEAatCO2concentrationsabove20%.Theeffectof higher ener- gycost(Fi gure15-20)istomakeTEAandthemembraneprocessmoreeconomicalthan anyof theothersover theentireCO2concentrationrange. Inthemid1980s,severallargeCO2-floodprojectswereinitiatedbasedontheavailability ofnaturallyoccurringCO2broughtinbypipelinetoWestTexas.Atthattimecrudeprices werehighand,significantly,NGLpriceswerealsohigh.Overalleconomicswerequitesimi- lartothescenariopresentedbyGoddininFi gur e15-19.Threeof thelargeprojects:Amera- daHess' sSeminolePlant(2x85MMscfd,77%CO2)(Schaffertetal.,1986;Woodetal., 1986),Shel l ' sWassonDenverunit(275MMscfd,93%CO2)(Flynn,1983;Younetal., 1987),andArco' sWillardunit(72.9MMscfd,86%CO2)(PriceandGregg,1983),anda smaller plant,theMitchellAlvordSouthCO2plant(7.5MMscfd,85%CO2)(McCannetal., 1987)usedcryogenicdistillation(theRyanHolmesprocess).Allof theseprojectsrecovered NGLandcryogenicdistillationwaschosenbecauseeconomicsfavoredNGLproduction. Table 15-12 Summary of C02 Removal Costs for DEA, Cryogenic, TEA, and Membrane Facilities PercentCO 22040608090 Flow,MMscfd18.224.336.573.3148.0 Capital,$MM DEA9.716.425.554.5103.6 CryoA16.220.828.342.673.5 CryoB16.220.828.342.673.5 TEA-DEA- - 15.021.636.965.0 Perm-Low- - 13.818.829.847.0 Perm-High~16.623.437.960.5 Source: Goddin (1982) MembranePermeationProcesses1275 (',,,m 0 = . 30 ~E " 2o., 0 0 FUEL GAS@ $2. 001MM BTU NGL@ $5. 001MM BTU POWER@ $0.051 KWH CRYO-B CR PERMDEA ~ CO2 SWEETINING c o2 iN F E D - ~ LFigure15-19. Estimatedcosts of carbon dioxide removalforaC02-flood EOR pr oj e ct - -lowu til itycost case(Goddin, 1982). Reproduced with permission from Proceedings of the 61st A nnua/ Convention of the GPA, copyright 1982, Gas ProcessorsAssociation 4 u, .O u .3 =E r162 MJ o ."2 \ ~ \ ~ FUEL G A S @ S6.001MM BTU NGL@$7.001MMBTU W HCRYO \-~ TEA ~ l l PERM ~ ~ / C O 2 SWEETENING CO2I NFEED-MOL % Figure15-20 . Estimatedcosts of carbon d iox id e removalforaC02-floodEOR proj ect--- high u til itycost case(Goddin, 1982). Reproduced with permission from the 61st A nnual Convention of the GPA, copyright 1982, Gas ProcessorsAssociation 1276Gas Purification Chevron' sSacrocproject(Parro,1984;SchendelandNolley,1984)andAmoco' sCentral Malletunit(100MMscfd,85%CO2)(Anon.,1985)usedmembranes.However,theSacroc projectpredatestheRyanHolmesprocess,andAmocowasundercontractualobligationto supplygaswithNGL' sstill presenttoadownstreamNGLextractionplant.Thereforeinthe mid-1980s,themarketgenerally sustainedtheconclusionof Goddin(1982)that,whenNGL wasrecovered,economicsfavoredcryogenicdistillation over theuseof either membranesor aminetreating inCO2-flood EORprojects. Interest inCO2-floodEORfellwiththedropinoil pricesthatoccurredintheearly1980s, butseemstobereviving.RecentCO2-flood EORprojectsusingmembranesinclude Mobil' s SaltCreekproject(64-110MMscfd,70%CO2),whichstartedupin1992;Shell' sMcCar- ney,Texasplant,whichbeganoperationin1993(11MMscfd,70%CO2);andAmoco' s Mallet plantinWestTexas(30-100MMscfd,80%CO2),whichstartedupin1994(Cynara, 1995).Oneof the keyfactorsfavoring membranesin these more recent projectsistheability to delay capital expenditures. Schendel(1984)proposedtheintegrationofamembraneprocessandcryogenicdistilla- tion.Feedgasfromthefieldpassesthroughahydrocarbondewpointcontrolunitandthen throughamembraneunit.ThisfirstmembraneunitremovesthebulkoftheCO2.The residueisfedtothecryogenicunit.RatherthansuppresstheCO2/ethaneazeotropewitha hydrocarbonrecyclestream,theCO2-richazeotropeisallowedtogooverheadtoasecond membraneunitwhereCO2isremovedasthepermeate.Theprimaryadvantageofthis schemeisthatthecryogenicunitcanbesizedfortherelativelyconstanthydrocarbonrate, nottheincreasingCO2rateovertime.AstheCO2rateincreases,additionalfront-endmem- branemodulescanbeaddedasrequired.ThedisadvantageofthisconceptisthattheCO2 productisproducedatlowpressure,requiringrecompression.Somecryogenicdistillation processconfigurationscanproducetheCO2productasaliquidwhichcanbeeasily pumped to highpressure(Wood et al.,1986;RyanandSchaffert,1984). Boustany etal.(1983)compared the performanceandeconomicsof hot potassium carbon- ate,cryogenicseparation,andmembranepermeationprocessesforCO2removal.They assumeda100MMscfdfeedstreamcontaining80%CO2availableat25psigand100~ Thegasisassumedtobesweet.Thecryogenicunitissimilar tothat previously described by Goddin(1982)exceptthattheCO2recoveredfromtheDEAunitiscompressedandcom- binedwiththat coming fromthe cryogenic column overhead. Themembraneunitconsistsoftwosectionsinseries.Thepermeatefromthesecondsec- tioniscompressedandrecycledtothefeed.Theresiduefromthesecondstageisfedtoa DEAunit.TheCO2recovered fromtheDEAunitiscombinedwiththatfromthefirstmem- branesectionandcompressedto500psig.Asummaryof operatingconditionsandcostdata ispresentedinTabl e15-13.Itisconcludedthatthemembranepermeationunits(PRISM separation)offer both capitalandoperating costsavings over the other systems. Well Fracturing Inthisprocess,whichisapplicable tocertaintypesof wells,highpressureCO 2 isinjected intoareservoir tofracturetheformation.Aslurryofsandisthenfedintothewelltofillthe fracturesandprovideaporouschannelforgasoroilflow.Theassociatedgasflowafterthe fracturecontainsahighamountof CO2,whichdecreasesto pipeline transmissionlevels over arelatively short period of time.Membranesworkwell for treating the gasresulting fromthe fracturingprocess.Becauseoftheirmodularnatureandportability,theycanbeusedto MembranePermeat i onProcesses1277 Table 15-13 Summary of Data for Comparison of Cryogenic, Hot Potassium Carbonate, andMembrane Processes Hot Potassium CarbonateCryogenicMembrane CO 2 ProductStream CO 2 Recovery,%96.993.096.9 CO2 Purity,%99.895.195.5 Hydrogen Recovery,%99.581.0183.5 CapitalCosts,$MM CO2Removal21.124.216.1 DEATreatingm4.94.9 Feed Compression 25to 250psig11.111.112.1 250to500psig0.96.20.9 Hydrocarbon Liquid Recovery4.0 CO2Compression10 to500psig7.9m5.1 Total Capital45.046.443.1 OperatingCosts,$MM11.788.466.46 Notes: I.AssumesC02columnoperationbased onethanerejection~propanerecovery. Source:Boustanyet al.(1983) maintainpipelinequalitygasuntiltheCO 2levelsreturntonormalandcanthenberemoved andused elsewhere(SchendelandSeymour,1985). Pipeline Natural Gas Typicalspecificationsforpipelinegasar eCO 2lessthan2%,H2Slessthan4ppmv,and waterlessthan7lb/MMscf.Watervaporisaveryfastgasand,therefore,dehydrationto pipelinespecificationscanbeachieved bymembranesystemswhilealso removingtheslow- erCO2orHzSacidgases.MembranepermeationsystemscanremoveCO2totherequired 2%level;however,atlowCO2concentrations,theCO2partialpressuredrivingforceis reducedandasignificantamountofhydrocarbon(primarilymethane)islostwiththeCO2- richpermeate.Theproblembecomesevenmoreseverewhentryingtoremovesubstantial quantitiesofHzStomeeta4ppmvspecification.Normally,withmembranesonlymodest adjustmentsto HzSconcentration in the ppmrangeareeconomically feasible.However,sub- sequent treatment withother processesto meet theH2Sspecification ispossible. TheGasResearchInstitute(GRI)performedanextendedfieldevaluationofamembrane unitoperatingonlowqualitynaturalgas(MeyerandGamez,1995;Leeetal.,1995).The unitwasastandardtwo-tubeGraceMembraneSystems(GMS)testsystemdesignedtotreat 0.5MMscf/dof750psignaturalgascontaining6.0%carbondioxide.Twotypesofasym- metriccelluloseacetatemembrane(standardandhigherdensity)weretested.Duringthe 1278Gas Purification 573-daytestperiodtheunitoperatedsmoothly,reducingtheCO2contentof thegastomeet thepipelinespecificationoflessthan2.0%carbondioxide.Thesystemalsoprovidedgas dehydrationto7lbH20/ MMscfor less. GRIevaluatedprocesseconomicsbasedonthetestresults.Forthecaseof aplanttreating 37MMscf/dat725psig,thestudyindicatedthatthemembraneprocesswouldbecompeti- tivewithDEAorMDEAsystemsifthegascontainedlessthan20%CO2andappreciably lessexpensivethantheamineprocessesifthegascontainedover20%CO2(Meyerand Gamez,1995). BhideandStern(1993)conductedastudytooptimizetheprocessconfigurationandoper- atingconditionsandtoassesstheeconomicsof membrane processesforremovingCO2from naturalgas.Thebasecasewasa35MMscf/dnaturalgasfeedstreamat800psiawithCO2 concentrationsintherangeof5to40mole%.Theoptimumconfigurationwasfoundtobea three-stagesystemconsistingofasinglepermeationstageinserieswithatwo-stageperme- ationcascadewithrecycle. Forthebasecaseoperatingconditions,membraneproperties,andeconomicparameters assumedinthestudy,membraneprocesseswerefoundtobemoreeconomicalthanDEA overtheentirerangeof CO2concentrationsconsidered(withnoH2Sinthefeed).WhenH2S isalsopresent,theresultsshowedthatthecostofmeetingproductgasspecifications(