A COMPARISON OF MEMBRANE SEPARATION AND DISTILLATION

0263±8762/00/$10.00+0.00q Institution of Chemical EngineersTrans IChemE, Vol 78, Part A, March 2000A COMPARISON OF MEMBRANE SEPARATION ANDDISTILLATIONA. B. HINCHLIFFE and K. E. PORTER (FELLOW)Speedwell House, Edgbaston, Birmingham, UKThe cost of separating binary mixtures by distillation is compared with that of separationby membranes. The comparisons are based on design studies for mixtures of varyingmolecular weight and relative volatility. These were chosen to represent a wide range ofconditions from cryogenic to vacuum distillation. The cost estimates for the membraneseparators are based on recently published correlations of costs predicted for commercial scaleplant.The comparisons are biased in favour of membranes and the use of `targeting’ enables thestudy to include hypothetical membranes as well as existing ones.The results explain why polymeric membranes, now used for the separation of hydrogen orwater vapour from gas mixtures and for some air separations, are unlikely ever to be used toseparate those mixtures which are now separated by conventional distillation. The potentialadvantages of molecular sieve membranes of high permeability and selectivity are such that,were they ever made available on a commercial scale, they might replace distillation for someseparations.Keywords: separation; membranes; distillation; economic evaluationINTRODUCTIONAlthough distillation continues to be the most popularmethod of separating ¯uid mixtures it is frequently pointedout that it consumes a considerable amount of energy. Othermethods of separation, including membranes, are suggestedas alternatives. Membranes are used for relatively fewseparations, the separation of hydrogen 14 or water 15from mixtures of gases, the small scale production ofnitrogen and of oxygen-enriched air 20,14 , and more recentlyfor the removal of the heavier hydrocarbons from naturalgas in dew point control 15±17 . The use of polymericmembranes for more than twenty years may be regardedas a mature technology. Nevertheless there is littleinformation on how to design and cost a commercial sizedmembrane separator. Even less is known about the newmolecular sieve membranes which, in laboratory studies,have been shown to have remarkable properties. Thus thereis a lack of comparisons between the cost of separation bymembranes and by distillation. A lingering doubt exists thatmembranes might replace distillation if only more wasknown about cost.As part of a recent study to evaluate the use of membranesfor the separation of carbon monoxide and hydrogen in theacetic acid process 1,2 , it was shown how the costs ofmembrane separators might be estimated from laboratorydata by means of correlations based on estimates of fullscaleplant costs. This provided a starting point for the workreported below. The project was suggested and supported bythe UK Department of Energy, Energy Technology SupportUnit (ETSU), and more details may be found in Porter andHinchliffe 3 .Distillation is a widely used technology. Real systemswere chosen to illustrate the effect of operating conditionson cost comparisons between membrane separation anddistillation. They include examples of vacuum distillation,pressure distillation and cryogenic distillation. The choiceof systems is described below after a description of themethod of estimating the cost of separation by membranes.In comparing the costs of separation, some assumptionswere made to favour the membrane technology.THE IMPORTANCE OF MEMBRANEPERMEABILITY AND SELECTIVITYThe separating properties of a membrane are its `fast gas’permeability and its selectivity. The selectivity is the ratio ofthe fast gas permeability to the slow gas permeability. Thepermeability determines the membrane area required to passthe separated gas through the membrane for a given activity(pressure) drop across the membrane. Thus for a givenselectivity, product speci®cation and pressure drop, thecapital cost of the membrane separator is approximatelyinversely proportional to the permeability. The selectivitydetermines how much of the slow gas passes through themembrane with the fast gas and thus the concentrationchanges achieved. The selectivities of commercially availablepolymeric membranes may vary between 2 and 40. Forthe molecular sieve membrane evaluated here the selectivityis in®nite, i.e. only the fast gas (water) passes through themembrane.In a single stage membrane separator, the concentrationof the slow gas increases on the high pressure side of themembrane as a result of the greater rate of permeation of thefast gas. Indeed, if the membrane area is large enough,255

256 HINCHCLIFFE and PORTERalmost all the fast gas will pass through the membrane andthe slow gas may be produced at a high purity. However,with the ®nite and relatively small selectivities of polymericmembranes, much of the slow gas is lost with the fast gas sothat the amount of the slow gas recovered at a high purity isreduced. It is impossible to produce a high purity fast gas ina single stage with polymeric membranes and for both fastand slow gases, the higher the purity of the product thelower its recovery.In order to achieve high recoveries of high purityproducts with polymeric membranes, it is necessary toadjust the mixture concentration on the high-pressure side ofthe membrane by recycling some of the mixture that haspassed through the membrane back to the membrane feedand use more than one stage. Flow sheets for single and twostage membrane separators are shown in Figures 1 to 3 Theamount of recycle (which must be recompressed) dependson the selectivity which thus determines the running cost.The capital cost of the membrane is approximately inverselyproportional to the permeability and to the pressure dropacross the membrane. The running cost is inverselyproportional to the selectivity. Even with a two stagemembrane separator, the recycle and running cost increaserapidly if high purity products are speci®ed so thatseparation is relatively expensive. The single stage separatoris relatively cheap, but because it can only recover reducedamounts of low purity products, the few applications ofsingle stage separators are for unusual applications.SEPARATION SPECIFICATIONSFor most separations, a high product purity is required toreduce downstream processing costs. A high recovery isrequired to avoid the additional upstream processing costsof producing additional feed. Thus so as to represent themost typical separation, the purity and recovery speci®cationswere each set at 99%, which are easily met bydistillation and by the in®nite selectivity molecular sievemembrane. For the polymeric membranes, the cost ofseparation was also calculated for a reduced purity/recoveryspeci®cation of 95%. This was to allow for the possibilitythat, because of the rapid rise in membrane separation costsat high purities, the total process cost of separation bymembranes, (including small additional upstream anddownstream costs resulting from a reduced speci®cation)might sometimes be less at a slightly lower purity. For anexample of this see Hinchliffe and Porter 1 . Alternatively,this reduction in speci®cation may be considered a means ofcompensating for any reduced ef®ciency resulting frombasing the costs of separation on the cross ¯ow model, i.e.not allowing for any bene®ts of a counter-current ¯owmodule design. This is discussed further below.Estimating the Capital and Running Costs of MembraneSeparatorsCost correlations and ¯ow patternsWhereas the cost of separation by distillation dependsonly on the speci®cation of the separation and on theproperties of the mixture to be separated, the cost ofseparation by membrane technology depends also on theproperties of the membrane. In comparing the twoseparation technologies it is necessary to allow for this.Despite the maturity, in business terms, of membraneseparation, it is reasonable to assume that not all polymericmembranes have yet been developed for commercial use.The evaluation should include those membranes of knownseparating properties, determined in the laboratory, whichmight be developed commercially, and also de®ne whatproperties are required of new, yet to be discovered,hypothetical membranes for them to compete with distillation.Then, after making a judgement about whether it ispossible that such membranes might ever be made andbecome commercially available, it is possible to decidebetween the two technologies.The method described below is based on severalempirical correlations of cost estimates provided bysuppliers of commercial scale membrane plant. Thesewere made for the purpose of identifying the optimumseparation speci®cation during a study of cost comparisonsbetween membrane separators and ¯ash distillation for theseparation of carbon monoxide and hydrogen in a process toproduce acetic acid 1 . The basis of the correlations isSalstonall’s equations for relating concentration changes tomembrane area in terms of permeability and selectivity 4 .These equations were derived for cross ¯ow ¯ow patterns.They were found to correlate the suppliers’ cost estimatesfor both spirally wound membrane modules which are cross¯ow, and hollow ®bre type membrane modules which maybe cross-¯ow, co or counter- current ¯ow. Mass balanceswere estimated to within +/- 2% regardless of the type ofmodule.It has been known for many years that, theoretically, for agiven membrane area, a higher product purity and recoveryis expected if the `fast gas’ permeate is removed so as tocause it to ¯ow countercurrently to the `slow gas’ retentate.See for example Giglia et al. 18 The separation from a cross-¯ow ¯ow pattern is expected to be in between that expectedfor co and counter-current ¯ow. (This result is similar to thatof the theoretical prediction of Lewis for the effect of ¯owpatterns on distillation tray ef®ciency.) The difference inconcentration of the slow gas in the permeate is greatest forpartial separations. In practice it has been found dif®cult toachieve the counter-current ¯ow separation 14 . This has beenexplained by concentration polarization 18,19 and by theeffect on the ¯ow pattern of the presence of the porouspolymer layer provided for most membranes to reduceleakage. The ¯ow pattern is always assumed to be cross-¯ow for assymetric membranes 19 , and it has beenshown 18,19 , that where concentration polarization in thesubstrate is severe, the performance of the module willapproximate cross-¯ow even if the overall ¯ow paths arecounter-current. The cross-¯ow assumption may be themost reasonable one to use for cost estimates and this mayexplain the excellent cost correlations obtained in theprevious work 1 .The extent to which the performance of a moduleapproximates to the idealized counter-current behaviourcan be described using the module ¯ow ef®ciency, g , whereg = (Y p Yê cf )/(Y cc Yê cf ) ´ 100 (1)This was proposed by Giglia et al. 18 who reported a value ofg = 0.51 for their laboratory scale module.Thus if Y cc = 0.99 and Y cf = 0.95 with g= 0.50, thenY p = 0.97. That is, the difference in the product purityspeci®cation between distillation and membranes may inTrans IChemE, Vol 78, Part A, March 2000

A COMPARISON OF MEMBRANE SEPARATION AND DISTILLATION257some circumstances be less than the 99% compared with95% assumed in the evaluation based on the cross-¯ow ¯owpattern. The membranes may be achieving purities nearer tothose achieved in distillation.Membrane process costs calculated from laboratorymeasurementsIt is normal practice to construct membrane separators byjoining together modules. Typically the module consists ofmembranes supported inside a pipe of some 150 to 300 mmdiameter and 3000 millimetres long. Dozens, maybehundreds of modules are required for most of thethroughputs used in this evaluation. The number of modules(and therefore the interconnecting piping, supports, foundationsetc.) is proportional to the membrane area required forthe separation. It is then convenient to introduce a newconcept, the Cost Permeability, which may be calculatedfrom the usual area permeability. The units of thesequantities are:Area permeability, G = nm 3 /hr ±1 bar ±1 m 2Cost permeability, G C = nm 3 /hr ±1 bar ±1 £ ±1As shown in Hinchliffe and Porter 1 , the cost permeabilityfor a given commercially available membrane may becalculated if the cost (or estimated cost) is known for manydifferent sized permeators separating the same system.Using Salstonall’s equation (4) it was found that a linear plotof A s against cost was obtained, where A s is proportional tothe membrane area calculated by Salstonall’s equation (4).A s is that function of concentrations, pressures and ¯owrates which is equal to area multiplied by area permeability.Then assuming that the cost is proportional to area, then alinear plot of A s against cost has a slope of permeabilitymultiplied by cost per unit area, which is the CostPermeability.After calculating the cost permeability in this way forseveral different polymers and module designs, an equationwas proposed for calculating for any polymer, the costpermeability from its area permeability and the costpermeability of a known polymer (2).This is:G c2 = G c1M G 2G 1(2)Where G is the area permeability of the pure fast gas, G c isthe cost permeability and subscript 1 refers to the knownmembrane and subscript 2 to the new membrane. M is themodule design factor which takes account of the effect ofdifferent membrane module designs, hollow ®bre, spirallywound or plate and frame 2 . For modules of the same designwith different membrane materials M = 1.0.It is assumed here that the thickness and cost per unitvolume of polymer of the new membrane are the same asthat of the known membrane. This might be justi®ed bynoting that when a large quantity of a polymer is made, theproduction cost becomes similar to those of other similarpolymers produced in large quantities, and by noting that thecost of the polymer membrane is often a relatively smallpart of the complete plant. The cost permeabilities werecalculated from the low cost of cellulose acetate membranes.Thus the cost of the new membrane may be greaterthan the estimated cost (that is the Cost Permeability may belower and the capital cost of the membrane plant higher thanestimated, thus the comparison may be biased in favour ofmembranes). The errors involved may be allowed for by thefollowing inequality:0.7G C< G C< G C(3)Where G C is the value of the Cost Permeability calculatedby equation (2), and G C * is the value used to make costestimates.It is very unusual to ®nd experimental measurements ofselectivity for binary mixtures. The selectivity valuesquoted in this paper were calculated as the ratio of thepure gas permeabilities, that is, if A is the fast gas and B theslow gas then S * = G A / G B . They were often adjusted toallow for a signi®cant difference in temperature between thelaboratory and the plant. Thus they may be overestimated.Also, the selectivity of the membrane obtained in practicemay be less than that observed in the laboratory due topressure losses in the permeator or diffusion effects. To takeaccount of this, values of the effective selectivity werecalculated from the plant performance data estimated forcellulose acetate membranes 1 and compared with thoseestimated from pure gas permeability data. The followinginequality was established and used to calculate theEffective Selectivity from the ideal selectivity:0.9S > S e > 0.63S (4)Where S is the selectivity based on pure gas permeabilitiesand S e is the effective selectivity which predicts theseparation using Salstonall`s equations.Thus by the use of the cost permeability equation, theeffective selectivity equation and Salstonall’s equation it ispossible to make an estimate of the cost of separating abinary mixture from only knowledge of the laboratoryvalues of fast gas permeability and selectivity.It is also possible to predict the properties required of anew membrane for it to achieve a speci®ed separation cost,for example a cost lower than that achieved by distillation.This has been de®ned as `targeting’ 2 . It is particularly usefulin this study of membrane technology as it avoids theconstraint of an evaluation based only on commerciallyavailable membranes.TargetingThe total annualized cost (TAC) of separation depends onboth the capital investment cost and the running cost. Thepermeability of the membrane sets the capital cost and theselectivity the running cost, thus it follows that there aremany combinations of permeability and selectivity which,for a particular membrane separation process, will result inthe same total cost. A family of isocost lines may be drawnon a ®gure with axes permeability and selectivity, or forconvenience, cost permeability and effective selectivity asde®ned above. One of the isocost lines may be set forexample, at the cost of separation by distillation. Thosemembranes with properties that place them on one side ofthe distillation cost line will result in a membrane separationcost less than distillation, those on the other side will haveseparation cost greater than distillation. All of the iso costlines may be generated for any particular process andseparation, by design and cost equations for imaginarymembranes, each membrane de®ned only by a combinationof permeability and selectivity. Real membranes may thenTrans IChemE, Vol 78, Part A, March 2000

258 HINCHCLIFFE and PORTERbe evaluated by where they appear on the isocost linesdiagram. Note that where the properties of a membrane areevaluated from laboratory data by the inequality relationships(3) and (4) then a rectangle on the targeting gridprovides a range of cost estimates for that membrane.Alternatively, the diagram may be used to guide thedevelopment of new membranes. The isocost line valuesdepend only on the membrane recycle process and on themass transfer driving force (or partial pressure drop) it isproposed to use. A typical targeting diagram is shown asFigure 5. A proposed new membrane is shown as arectangle, the error bands showing an approximate rangeof performances.Membrane process ¯owsheetsExamples of the most commonly used ¯owsheets areshown in Figures 1 to 3. In all cases the feed gas is at a highpressure. Figure 1 shows a single stage separator which isused for the partial recovery of lower purity products.Figures 2 and 3 show two- stage membrane separators withrecycle and recompression of a product stream so as tochange the concentration of the feed. These separators mayin principle produce high purity products at a high recoveryeven for membranes of a low (i.e. ®nite) selectivity. Thecost of production by two- stage separators is signi®cantlylarger than by a single stage separator. Figure 2 showsrecycle to the feed of permeate from the second stage whileFigure 3 shows recycle of retentate. The choice of recycle¯owsheet (permeate or retentate) depends mostly on theconcentration of the feed.The use of membranes to separate condensing mixturesThe proposal to use membranes to separate condensingmixtures (i.e. mixtures below their critical temperature)introduces constraints on operating conditions that do notapply to the use of membranes to separate gases (above thecritical temperature). In gas separation, the pressuredifference across the membrane may, in principle, be setat any value. The mass transfer driving force (or differencein activity) for any component is assumed to be itsdifference in partial pressure across the membrane,isothermal operation (each side of the membrane at thesame temperature) is usual. The separators work at ambienttemperature or above, which is so far above the criticaltemperature that there is no danger of two phase ¯owoccurring. Thus, in the separation of hydrogen and carbonmonoxide the operating pressure was 40 bar and thepressure difference across the membrane about 38 bar.The systems that may be separated by conventionaldistillation must be processed below their critical temperaturesto avoid thermal degradation; thus the activity drivingforces of each of the components depends on temperature.The higher the temperature the higher the vapour pressure.Figure 2. Two unit separator with permeate recycle.There will be for each of the components, some temperatureat which it will decompose. This is why mixtures of largemolecules are distilled under vacuum. For example, theethylbenzene-styrene system at 1008 C would have a totalvapour pressure of only about 0.3 bar. For simplicity it isassumed here that the decomposition temperature of thecomponents of all the hydrocarbon mixtures is 1508 C. It isunlikely that polymeric membranes will withstand atemperature of 1508 C and maintain their mechanicalstrength. It has been assumed that 1008 C is the maximumsafe temperature for polymeric membranes, although thismay be too high for operation over a practical period.The components vapour pressures at 1008 C and 1508 C areshown in Table 2. These are the maximum partial pressuredriving forces for causing permeation through the membrane.A simple rule suggests itself that the available partialpressure driving force is about the same as the systemoperating pressure in conventional distillation. Thus formost systems, the increased membrane area required formass transfer will result in an unacceptably high capitalcost. The systems most likely to compete with conventionaldistillation are those of small molecular weight andincreased operating pressure, such as the propane-butanesystem.The selectivity of polymeric membranesA simpli®ed model for selectivity assumes that thepermeability of each component is proportional to itsdiffusivity (P) and to its solubility (Q) in the membranepolymer. This mode of transport is known as the solutiondiffusionmechanism 13 .Thus the selectivity is given by 13S = P a Q a /P b Q b = [P a /P b ][Q a /Q b ] (5)Polymeric membranes were ®rst used commercially forhydrogen separation but a later application has been forremoving condensables from natural gas (`dew pointFigure 1. Single stage separator.Figure 3. Two stage cascade separator with retentate recycle.Trans IChemE, Vol 78, Part A, March 2000

A COMPARISON OF MEMBRANE SEPARATION AND DISTILLATION259Table 1. Summary of systems selected for analysis.PracticalMixture Average Relative Mean Molecular Operating PressureVolatility Weight (bar)Propane/n-butane 2.7 51 15.9i-Butane/n-butane 1.4 58 6.75Benzene/toluene 2.4 85.5 1.013Ethylbenzene/styrene 1.5 105 0.2Acetone/water 6.0 39.5 1.013Air separation Cryogenic Separation ProcessHydrogen/carbon monoxide Cryogenic Separation Processcontrol’). In hydrogen recovery, the very small hydrogenmolecule is the fast gas, (even although it is expected to bethe least soluble), because of its greater coef®cient ofdiffusion through the membrane polymer. In natural gastreatment, the fast gases are the less volatile, largermolecules which have a larger solubility in the membranepolymer. It is then convenient to refer to the hydrogen typemembranes as diffusion membranes and the gas treatmentmembranes as solubility membranes. There is unlikely to bemuch difference in diffusion coef®cients of mixturecomponents of a similar molecular weight (such as thosefound in the distillation systems evaluated here). Thepolymeric membranes used for distillation systems will thenbe solubility membranes. A simple theory, assumingsolubility to be inversely proportional to the vapour pressureof a component, suggests that for many mixtures, themembrane selectivity will be similar to the mixture relativevolatility in distillation. The values of relative volatility andselectivity for the system/polymer combinations shown inTable 3 con®rm that this is true in several cases.The polymer may introduce deviations from this simpletheory by causing the solubility and transport mechanismsof each of the components to be different so that theselectivity is increased 13 , (this is analogous to the effect ofthe solvent on the vapour-liquid equilibria in the extractivedistillation of a close boiling mixture). For example, Table 7in Prasad et al. 14 , lists some of the polymers that have beendeveloped for air separation, and notes that in general theselectivities are in the range of 4 to 8 and the permeabilitiesin the range of 1 to 20 barrers. Although having selectivitiesgreater than the O 2 /N 2 relative volatility, none of them haveproperties capable of separating oxygen and nitrogen both toa purity of 95% and a recovery of 95% at a cost comparableto cryogenic distillation.A careful search was made of the literature reportingexperimental measurements of the permeability of the testsystem components in various polymers. This yielded a highselectivity polymer, (based on the pure component permeabilityratio), for both the oxygen/nitrogen separation andthe propane/butane separation. Neither of these polymershas ever to the authors’ knowledge been developed as acommercially available membrane. Nevertheless they wereused as part of the evaluation of membrane technology andproduction costs that were estimated by the targetingprocedure outlined above. Most of the laboratory measurementsof permeability were made close to ambienttemperature and, for the propane-butane system, it wasnecessary to estimate the very signi®cant increases inpermeability which might result by using the membranes at1008 C. This was done using available data for the activationenergy for permeation and assuming that the activationenergy was constant over the temperature range to 1008 C 12 .There is an exponential relationship between permeabilityand temperature where:G = G O eê(Ep/R 00 .T)G = Permeability; G O = Constant; E p = Activation energyfor permeation; R’’ = Universal gas constant; T = Absolutetemperature.Thus, it is possible to calculate the permeability atelevated temperatures provided the activation energy forpermeation is constant for the temperature range underconsideration. The change in selectivity of a membranedepends on the relative value of the activation energies forTable 2. Critical temperature of gases and distillation system components.Component MW. T C (8 C) Vapour Pressure (Bara)100 o C 1508 CHydrogen 2 - 240 - -Carbon monoxide 28 - 139 - -Nitrogen 28 - 147 - -Oxygen 32 - 119 - -Propane 44 97 - -n - Butane 58 152 15.5* 33.5i - Butane 58 135 19.9* -Benzene 78 298 1.8 5.8Toluene 92 319 0.74 2.7Styrene 104 374 0.26 1.1Ethyl benzene 106 344 0.35 1.5(Vapour pressure is predicted unless *).Table 3. Relative Volatility and Membrane SelectivityRelativeMembraneSystem Volatility SelectivityPropane/n-Butane 2.7 2.8 (A) 1.4 (B) 58 (C)i-Butane/n-Butane 1.4 2.6 (B)Benzene/toluene 2.4 2.1 (D)Ethylbenzene/styrene 1.5Oxygen/nitrogen 2 3.2 (B) 5.8 (E) 16 (F) 2.2 (C)(A) Commercial hydrocarbon dewpoint control membrane at 408 C.(B) PTFE at 258 C for O 2 /N 2 and 908 C for others.(C) Poly(oxydimethylsilylene) with 10% ®ller scantocel CS (vulcanizedsilicone rubber) at 08 C.(D) Low density polyethylene at 748 C.(E) Cellulose triacetate at 308 C.(F) Cellulose nitrate at 308 C.Trans IChemE, Vol 78, Part A, March 2000

260 HINCHCLIFFE and PORTERpermeation of the two gases permeating through themembrane. In general there will be a reduction in selectivitywith an increase in temperature.The permeability and selectivity of the molecular sievemembranePerhaps the most promising development in membranetechnology has been the development of molecular sievemembranes. These are solids capable of withstanding hightemperatures and pressure gradient without leaking, throughwhich one component may rapidly permeate by molecularmovement in their crystalline structure. The selectivity isvery high and may be in®nite. In this work a zeolitemolecular sieve membrane was evaluated which is availablefor laboratory scale drying of solvents. The Smart ChemicalCompany made performance data available 11 . The sieve hasan in®nite selectivity for the separation of water fromsolvents and is used in the pervaporation mode. In this workit was assumed to be made part of the process shown inFigure 12. It was known that the rate of transport of waterthrough the sieve at any particular concentration andpressure depended on the temperature of the liquid on thehigh-pressure side, that is an activity driving force operated.The rate of mass transfer was calculated from vapour-liquidequilibrium data for the acetone-water system to de®ne awater partial pressure driving force on the high-pressureside of the sieve.The effect of feed concentration on heat requirementsA molecular sieve membrane with in®nite selectivityrequires no recycle to achieve absolute purity. Only lowgradeheat is required to evaporate the transferringcomponent. A simple analysis leads to an understandingof how the feed concentration will in¯uence the runningcost comparison between distillation and such a molecularsieve.The minimum re¯ux ratio for distillation of a binarymixture of constant relative volatility may be estimatedfrom the Fenske-Underwood equation, where x refers to theconcentration of the most volatile component, (m.v.c.).R m = [1/(a ê1)][(x D/x F) êa(1 êx D)/(1 êx F)] (6)Now for the case of a very nearly pure top product, assumex D = 1, and assuming complete recovery of the mostvolatile component as the top product, then if F and D arethe molar ¯ow rates of feed and top product, and x F is theconcentration of the top product in the feed, then x F = D/FandR m = F/D(a ê1) (7)The distillation vapour rate V = (1 + R)DAnd where R = 1.1R mThen the vapour ¯ow rate per mole of feed is:V F = x F+ 1.1/(a ê1) (8)Straight parallel lines for V F against x F are shown on Figure4 for various values of relative volatility, a.For a molecular sieve membrane where the less volatilecomponent (l.v.c) is the fast gas, then:V M = (1 êx F ) (9)If both V Fand V Mare produced by the same low value steamsupply, and assuming the costs of pumping and vacuumFigure 4. Comparison of the vapour load required by molecular sievemembrane with that of distillation.generation to be relatively very small (see Table 7 below),then the running cost of the processes are equal at aconcentration x MFwhen V M = V F.From whichx MF = (0.5)[1 ê1.1/(a ê1)] (10)For x F less than x MF less vapour is needed for distillation.For x Fgreater than x MFless vapour is needed for themembrane.The system used to compare the sieve with distillation isacetone/water. The average relative volatility is 6.0 thus, (asshown in Figure 4); less vapour is needed for the membranefor feed concentrations of acetone more than 0.39 molefaction. (B permeates in Figure 4) For more dif®cultseparations, (a # 2.1),the sieve uses less vapour thandistillation for all concentrations, (V F $ 1.0 at x F = 0).For very easy separations (a very large and 1.1/(a-1)approaches 0) then V F = x F and V MF = 0.5. The sieve has asmaller advantage if the fast gas is the mvc, then V M = x F (Dpermeates in Figure 4).For a very easy distillation separation of this type thesieve has no advantage because the vapour loads ofdistillation and the membrane are the same. The relationshipsderived above are helpful for interpreting the costcomparisons between distillation and solvent drying bymolecular sieve which are presented below.SCOPE OF STUDY AND CHOICE OF SYSTEMSThe operating conditions of distillation columns varyconsiderably depending on the system to be separated. Thusin conventional distillation with condensation of the topproduct by cooling water, the operating pressure rangesfrom 0.1 bar to more than 10 bar. Cryogenic distillation isused to separate small molecular weight materials. Ingeneral, the cost of separation varies with concentration andwith throughput. Table 1 shows the systems chosen torepresent the range of conditions encountered in distillation.Binary systems were used throughout the analysis andexcept where otherwise stated equal molal feeds wereTrans IChemE, Vol 78, Part A, March 2000

A COMPARISON OF MEMBRANE SEPARATION AND DISTILLATION261evaluated. The hydrocarbon systems representing conventionaldistillation were chosen so as to vary relativevolatility and molecular weight.The relative volatility is a measure of the dif®culty of theseparation and determines the re¯ux ratio required and theenergy consumption by the distillation column. Molecularweight is an important variable because, for the conditionsof practical distillation it correlates all physical propertiesand ¯owrates, as shown by Porter and Jenkins, 1979 5 .Large molecular weight mixtures of high normal boilingpoint must be distilled under vacuum to prevent thermaldegradation, and low molecular weight mixtures of lownormal boiling point must be distilled under a pressure suchthat cooling water may condense the top product of thedistillation column. Although the operating pressure forconventional distillation varies signi®cantly, as notedabove, the operating temperature (60 to 1508 C) is alwaysrelatively low (i.e. such that low-grade energy, maybe wastesteam, is used for supplying the heat).The acetone-water system was chosen to enable a costcomparison of distillation with a zeolite molecular sievemembrane. This comparison also demonstrates the effect offeed concentration on the relative cost.Less than 20 materials have a critical temperature of lessthan 508 C and thus may not be condensed by cooling waterhowever much pressure is applied. These are the materialsthat need cryogenic distillation. Cryogenic distillationrequires high-grade expensive electrical energy to drivethe compressors in the cooling cycle and additionalexchangers to retain the cold within the system. Thismeans it is signi®cantly more expensive than conventionaldistillation. Two cryogenic systems were chosen. Thecarbon monoxide-hydrogen system is that used in ourprevious work where it was predicted that membranes mightreplace ¯ash distillation. The methods of cost estimation formembrane separators developed in that work were usedhere. The oxygen-nitrogen air separation was chosenbecause it is the most widely used large scale cryogenicdistillation and cost comparisons included product distributioncosts. The feed concentrations for these two cryogenicsystems were those encountered in practice. Table 1 showsthe systems used to compare the cost of separation bydistillation with that by membranes. Table 2 shows for eachof the system components, the critical temperature and thevapour pressure at 1008 C and 1508 C.ECONOMIC EVALUATION AND THE COST OFENERGYA simple method of amortization of capital costs wasused, i.e. a three-year payback time. The costs are presentedper tonne of distillation column top product so as to permit acomparison of distillation with membrane separation. Thecost of separation, £/ tonne distillate is calculated asCost of Separation =(Operating Cost+ CAPEX/3)/(DxMW) (11)The total capital expenditure (CAPEX) is calculated usingaverage installation factors as recommended by Garrett 6 andother sources 7 and is between 2.0 and for 4.0 times theestimated equipment costs.Most distillation processes require relatively low qualitysteam, i.e. 5 bar. steam at approximately 1508 C is usuallymore than adequate for most distillation processes.Manufacturing sites with well-integrated energy systemsmay have suf®cient 5-bar steam as a by-product fromhigh-pressure steam expanded through turbines. Othersmaller sites may need to generate steam speci®cally fordistillation. Steam costs were calculated to be 0.9 and 2.4p/kW h for these two cases respectively. A recentevaluation for Texaco 8 used a steam cost of 0.6 p/kWhfor four bars steam. A steam cost in the middle range of1.65 p/kW h has been used to estimate separation costs.The cost of the electrical power used to drive compressorsin some of the membrane processes was assumed to be4 p/kWh.DETERMINING THE COST OF SEPARATION BYDISTILLATIONConventional DistillationStandard methods were used to calculate the size and costof the distillation column and its internals, the feed pump,the condenser and the reboiler. The column internals weresieve trays for all systems except the vacuum distillationsystem where structured packing were used. The optimumre¯ux ratio was calculated for each system but with thesteam cost of 0.0165£/KW h the ratio R opt /R m was alwaysless than 1.1. Thus a re¯ux ratio R = 1.1R m was used as thepractical minimum. The results of the design and costingcalculations are summarized in Table 4. The cost ofTable 4. Separation costs using distillation.Ethyl benzine/styrene Benzine/toluene i-Butane/n-Butane Propane/n-ButaneScale of operation(tonnes (dist)/h) 12.5 12.5 12.5 12.5Average MW 105 85.5 58 51Average relative volatility 1.5 2.4 1.4 2.7Operating pressure (atma) 0.2 1.0 6.75 15.9Column diameter (m) 3.30 2.33 2.59 1.98Column height (m) 28.1 25.9 38.6 19Heat load (kW) 14 444 3 243 5 280 2 999Heat load/unit area (kW/m 2 ) 1 690 767 1 152 991Annual capital charge (£k) 385 154 295 196Annual operating cost (£k) 2 205 501 686 404Operating cost/total cost 0.85 0.76 0.76 0.7Cost of separation(£/tonne distillate) 25.90 6.55 10.57 4.25Trans IChemE, Vol 78, Part A, March 2000

262 HINCHCLIFFE and PORTERseparation varies with relative volatility and molecularweight and for conventional distillation is calculated to bebetween 4 and 26 £/tonne distillate.Cryogenic DistillationAir separationThe most commonly used application of cryogenicdistillation is the separation of air. Air separation may beconsidered an example of the most advanced cryogenicdistillation technology and thus provides a realisticcomparison for membrane technology. Here this wassimpli®ed to the separation of oxygen and nitrogen. Thereis a large market for both oxygen and nitrogen. Much of thebusiness of air separation involves gaining a customer mixsuch that the products of separation (including theintermediate product, argon) may be sold and the cost ofseparation shared between them. A large distillation unitsupplying one thousand tonnes/day of oxygen to a methanolmanufacturing plant (say) may simultaneously producelique®ed nitrogen and argon for distribution to othercustomers . Thus the cryogenic distillation plant is wellsuited to satisfying this market for separated air products.Modern plants which use structured packing in the lowpressure columns require the air to be compressed only toabout 6 bar. and much effort is still being made to improveef®ciencies even further.The cost of supplying these gases depends on the scaleof operation. Large-scale production of oxygen andnitrogen (more than 100 tonnes of oxygen a day) justi®esbuilding a plant and supplying the gases by pipeline.Medium scale oxygen and nitrogen supply, in the range0.5-100 tonnes/day is usually made by distributed liquid intankers. Although small scale cryogenic plants areeconomical even as low as 3,000 scfh for nitrogen, thecosts of the products is only marginally less than that ofliquid produced from large scale plants. The supply ofsmall quantities of oxygen and nitrogen is made in highpressuregas cylinders containing about 14.3 kg. of gas andis very much more expensive.Suppliers were approached and agreed to give typical costvalues for these different scales of operation. These were:and Tighe 2 . Costs were estimated on the basis of designcalculations combined with correlations of cost data fromsuppliers. The cost of separation by distillation was£52/tonne (1991) of carbon monoxide.RESULTS OF THE COMPARISON OFMEMBRANES WITH DISTILLATIONPolymeric membranesHydrogen/carbon monoxide.A comparison of membranes and distillation for thissystem was the subject of the previous work 1,2 on a processin which the carbon monoxide from the separator is the feedto a reactor producing acetic acid. The reaction takes placeat a pressure of 40 bar and the carbon monoxide is the slowgas retentate which stays on the high pressure side of themembrane before being fed to the reactor. That is much ofthe compression costs of the membrane separator wouldhave to be incurred anyway by the rest of the process. Thehydrogen rich permeate supplied some of the fuel gas for theprocess. The recycle ¯ow sheet for the membrane separationwas the two unit separator with permeate recycle in Figure2. It was predicted that by optimizing the separation (i.e.reducing the speci®ed concentration of carbon monoxideleaving the membrane from the 97% produced by distillationto 95%), the membrane process annual cost would beless than 80% of the distillation process cost. This was basedon commercially available membranes of cellulose acetate.The result is shown on Figure 5 which is a targeting graphfor membrane separation as part of a process to manufactureacetic acid. Also shown on Figure 5 is the data point for anew polymer, 3(trimethyl)siloxy-c -methacyloxypropylsilane(TRIS). If this could be produced as part of a permeatorthen the cost of separation by membrane is predicted to bereduced to about 60% of the distillation cost. The operatingpressure on the high pressure side of the membrane (set bythe requirements of the process) was 40 bar and that on thelow side 2 bar. The permeabilities and selectivities of themembrane polymers are shown in Table 5.Large scale:Medium scale:Small scale:Oxygen £21-24.5 per tonneNitrogen £24 - 28 per tonneOxygen £44 .5- 49 per tonneNitrogen £51 - 56 per tonneOxygen £758 per tonneNitrogen £1110 per tonneThese different values are helpful in discussing the potentialfor membrane air separation.The separation of hydrogen and carbon monoxideMembrane separators have been used to separatehydrogen from other gases for many years. A comparisonof the cost of separation of this mixture by ¯ash distillationand by membranes led to the methods of estimating thecosts of membrane separators used in this work. It isdescribed in Hinchliffe and Porter 1 and in Porter, HinchliffeFigure 5. Targeting graph for the separation of carbon monoxide andhydrogen.Trans IChemE, Vol 78, Part A, March 2000

A COMPARISON OF MEMBRANE SEPARATION AND DISTILLATION263Table 5. Calculated values of cost permeability and effective selectivity forthree commercial membrane systems, hydrogen-carbon monoxideseparation.Cost permeabilityMembrane Effective selectivity (sm 3 /£k.h.bar)A 31 1.73B 22 2.37C 25 3.05Oxygen and nitrogen (air separation)Single stage membrane separators have been used formany years for some air separations. In this particular case,the relatively low cost of both the feed and that ofcontrolling air pollution by the waste streams make partialseparations acceptable and economically viable on arelatively small scale. High purity nitrogen (the slow gas)for gas purging or inert atmospheres may be economicallyproduced by a single stage separator even at a low recoveryof less than 50%. The low purity permeate from a singlestage membrane may be used to supply oxygen- rich air forcombustion processes. A description of the commercial usesof partial separation of air by single-stage membrane plantsis given by Prasad et al. 14 where the technology is comparedwith pressure- swing adsorption, (PSA), and distillation.Here only high purity, high recovery air separation isconsidered, and assessed whether it is likely that apolymeric membrane might be developed to compete, ona cost basis with distillation.Preliminary calculations were carried out and theseindicated that the most favourable two stage membranecon®guration was the two stage cascade with retentaterecycle shown in Figure 3. The operating conditions wereoptimized for the particular membrane properties. Ingeneral, a high feed pressure was favoured by membranesof lower cost permeability and/or effective selectivity. Thecost of producing gas at the same purity and recovery asdistillation (above 99%) was estimated to be over£1000/tonneproduct. For the reasons given above, a design and targetingevaluation was also made for three different scales ofoperation at 95 % product purity. The two-stage processwith recycle of retentate was assumed for each study.(Figure 3). The results of the optimized cost calculations arepresented as isocost lines in Figures 6, 7 and 8. Productioncosts for three membrane materials are shown on thediagrams. These are cellulose acetate, polyphenyleneoxide(PPO) and the `literature membrane’, cellulose nitrate. Theirmembrane properties are given in Table 3.Figure 6 shows the result of calculations at the large-scalerate of supply of 650 tonnes per day of total product. Thecost of production of 95% purity oxygen and nitrogen usingthe commercially available membrane cellulose acatate isabout £100/tonne, about six times that of the cost ofproduction of 99% purity product by cryogenic distillation.The estimated cost using the cellulose nitrate membrane at£25 /tonne is close to the cryogenic cost of £17 to 24/tonne.This material is not available in the form of a membrane forcommercial scale separation. The targeting isocost curvesshow how membrane properties must be increased tocompete with distillation.Figure 7 shows a similar result for a medium productionrate and Figure 8 for the small production rate. There is aFigure 6. Targeting graph for large-scale separation of oxygen andnitrogen.Figure 7. Targeting graph for medium-scale separation of oxygen andnitrogen.Trans IChemE, Vol 78, Part A, March 2000

264 HINCHCLIFFE and PORTERmembrane will be increased while the costs of distributedliquid will remain the same.Figure 8. Targeting graph for small-scale separation of oxygen andnitrogen.scale effect in that the membrane production costs areestimated to increase as the rate of production decreases. Forthe medium production rate, the cost of production usingthe commercially available membrane is about twice that ofthe distributed liquid cost. A cost of production using thecellulose nitrate membrane is predicted to be about the sameas the distributed liquid cost All membranes are predicted toseparate oxygen and nitrogen (at 95% purity) at a lower costthan that by supply in cylinders. Thus it may be possible todevelop a commercial scale membrane separator to producea reduced (95%) recovery of reduced purity (95%) productsat prices lower than cylinder supply and similar to those ofdistributed liquid. The membrane plant is in effect replacingthe distribution costs rather than the manufacturing costs.High purity oxygen and nitrogen will continue to beproduced for large throughputs by cryogenic distillationand this con®rms current practice. It should be noted that themembrane production costs have been made assuming aconstant rate of supply and that both products may be sold atthe point of manufacture. If the supply is requiredintermittently sometimes at higher rates, or if only oneproduct is to be sold, then the production costs of thePropane and butane.This is a conventional distillation system, and membraneseparation must be below the critical temperature where thepressure and activity levels are set by temperature. For thereasons stated above the membrane evaluation was made fora temperature of a hundred degrees centigrade when thetotal vapour pressure is about 20 bar (depending oncomposition). Two polymers were evaluated. One was acommercially available membrane used for gas dewpointcontrol. The material was not disclosed but process data forpropane/butane were given for 40 degrees centigrade andthis was used to estimate properties at 100 degreescentigrade. The other polymer, vulcanized silicon rubber,was chosen because of the permeability values given in theliterature for pure propane and pure butane at 08 C, whichwere extrapolated to 1008 C. It seems unlikely that thepredicted permeability and selectivity (40) would ever beproduced in a real membrane, but evaluation based on thesecalculated properties serves to illustrate the use of atargeting procedure. The separating properties of the dewpoint control membrane and of silicone rubber are given inTable 6.The evaluation was based on ¯owsheets in Figures 2 and3. Figure 9 shows the cost of separation of propane andbutane using the hydrocarbon dew point control membrane.The cost of separation to produce an inferior product is morethan two orders of magnitude greater than the cost ofseparation by distillation (£4.25/tonne at 99% purity and99% recovery).In order to achieve a dramatic reduction in the estimatedproduction cost by permitting purity and recovery to varywithin reasonable limits, a targeting procedure was developedon the assumption that the separation must be achievedin a single stage (no recycle) and that the purity of theproducts is the same as the recovery. Figure 10 is based onseparating an equal molar mixture of propane and butane.The lower graph shows the separation that may be achievedat various selectivities between 2 and 160 with a feed at onehundred degrees centigrade (5 degrees centigrade superheat).The upper graph then plots the cost of separationagainst selectivity for a range of values of the costpermeability of the membrane between 5-100 nm 3 /£kh.bar.The dotted line on the upper graph marks the cost ofachieving the separation by distillation.From the targeting graph, it can be concluded that a highperformance membrane with the selectivity calculated forsilicon rubber (S = 40) and able to operate safely at 1008 C,would be able to compete with distillation if a costpermeability of more on 5nm 3 /£kh.bar can be realized.The cost of separation would be approximately £3.75 perTable 6. Membrane properties.Gas Dew Point Control MembraneHigh Performance MembraneMaterial: Commercially Con®dential. Vulcanized Silicone Rubber.Fast gas permeability: 0.714 kmol/hm 2 bar 23.5 kmol/hm 2 barSelectivity: 2.76 40Cost permeability:1.01 nm 3 /£k.h.barTrans IChemE, Vol 78, Part A, March 2000

A COMPARISON OF MEMBRANE SEPARATION AND DISTILLATION265tonne of propane at a product purity and recovery of 92%. Itmust be recognized that this increased performance wouldneed a 14-fold increase in selectivity and more than a ®vefold increase in cost permeability compared with thedewpoint control membrane that is commercially available.If the temperature were reduced to reduce the likelihood ofmembrane failure, the cost of separation would rise.The cost of separating the other larger molecular weightsystems (now separated by conventional distillation) wouldbe even higher than those calculated for propane/butanebecause of their lower vapour pressures. The mainconclusion is that it is most unlikely that separation bypolymeric membranes will ever replace conventionaldistillation.Figure 9. Cost of separation of propane and butane using commercial dewpoint control membrane with ¯owsheets as presented in Figure 2 (TURP)and Figure 3 (TSCRR).Figure 10. Targeting graph for the single stage separation of propane andbutane.Evaluation of a molecular sieve membraneHere the zeolite membrane supplied by the SmartChemical Company to remove water from solvents hasbeen evaluated. Some data on this application has beenpublished in the Smart brochure and this was used to makecost estimates suf®cient to illustrate the potential advantagesof the process. Data was available for removing waterfrom propanol, to demonstrate the ability of the membraneto break the propanol-water azeotrope. The system chosenfor comparing the zeolite membrane with distillation wasacetone-water, relative volatility, a =6. At a feed rate of 7.2tonne/hr the cost of separating an equal molar feed intoacetone and water products each of 99 % weight purity was£8.28/tonne acetone. The distillation design is summarizedin Table 7.The water vapour permeation data for the propanol watersystem clearly shows an activity coef®cient effect, in thatthe permeability is higher at low water concentrations. Aline correlating water ¯ux against water concentration forthe propanol water system is shown in Figure 11.Permeabilities were estimated from the shape of the lineto yield the values shown below in Table 8.These may be compared with typical values of permeabilityfor polymeric membranes of between 0.1 to 1.0 nm 3 /m 2 h.bar (noting that the selectivity is in®nite, this illustratesthe great potential of the molecular sieve compared to thepolymeric membrane). These values were used to estimatethe size of the membranes required for the acetone waterseparation. This implicitly assumes that the variation withconcentration of the activity coef®cients of water in acetoneis the same as that of water in propanol, which can only beapproximately true.The membrane process evaluated was pervaporation,with the mixture to be separated as a liquid at a temperatureof 70 degrees centigrade on one side of the membrane, withwater vapour at a low pressure of 8 m.bara, and acondensing temperature of 40 degrees centigrade on theother side. The low pressure side temperature was chosen sothat the vapour product might be pumped from the system asliquid, after condensation by cooling water at 30 degreescentigrade.(i.e. the condenser was assumed to be similar inprinciple to those used for condensing steam from the lowpressure end of a steam turbine) The vacuum system wasassumed to be as described by Mangnall 21 Vacuum/PressureProducing Machines and Associated Equipment 1989, page386, two steam ejectors and a liquid ring pump, and wassized on the basis of assuming a ¯ow of dissolved air andleakage air of 2 kg./hr and the associated water vapourTrans IChemE, Vol 78, Part A, March 2000

266 HINCHCLIFFE and PORTERTable 7. Acetone ± water separation speci®cation.Feed concentration = 50 mol% AcetoneFeed rate = 7.2 tonne/hProduct speci®cation = 99% Product Purity and Recoveryof AcetoneDistillation design dataRelative volatility = 6.0Pressure = 1.013 bara (atmospheric pressure)Column diameter = 1.31 mColumn height = 17mNumber of theoretical plates = 15Re¯ux ratio = 0.41Total heat load = 1925 kWCapital charge = 85 £k/yearOperating cost = 296 £k/yearCost of separation = £8.28/te acetoneMolecular sieve membrane designMode = PervaporationInlet temperature = 708 CFeed side pressure = 1.013 bara (atmospheric pressure)Permeate pressure = 8 m baraCondensing temperature = 408 CHeating/condensing load = 871 kWPower for vacuum system = 1.1 kWCapital charge for vacuum system = 13 £k/yearMembrane area = 105 m 2Operating cost = 132.87 £k/year(which may be an overestimate if no dissolved air passesthrough the membrane.) The capital and operating costs ofthe vacuum system were calculated to be relatively small atabout 5% of the total costs. For cryogenic distillationsystems (de®ned here as those which may not be condensedby cooling water) the permeate gas would have to berecompressed and the costs would be higher.The ¯owsheet used for cost estimates is shown in Figure12. The membrane process design is compared withdistillation in Table 7. It may be noted that for an equalTable 8: Estimated permeability of molecular sieve membrane.molar feed the heat load, and thus the operating cost, for themembrane is less than half that of the distillation column.The cost of separation will depend on the cost per unitarea of the membrane. Membrane process costs are shownin Table 8. For the membrane to achieve the same cost ofseparation as distillation the cost per square metre ofmembrane would be about £1300/m 2 . The variation inseparation cost with membrane cost is shown in Figure 13for 50% water feeds and compared with the cost ofseparation by distillation.At lower water concentrations, for the reasons givenabove, the cost comparisons favour the membrane and it islikely that in separating a mixture at 10% water, the cost ofseparation by the membrane would be signi®cantly lowerthan that of separation by distillation.DISCUSSIONPermeability (nm 3 /m 2 hbar)Above 45 mol% water 33.615 to 45 mol% water 57.00 to 15 mol% water 140.0The estimation of the cost of distillation is well understoodand described in textbooks, for example 9,10 . Distillationcosts may be calculated from the properties of themixture to be separated. There is little or no previous workwhich describes how to estimate the cost of a membraneseparator from the properties of the mixture or whichdescribes how costs depend on membrane properties. Thishas resulted in a situation where engineers are uncertainabout the viability of membrane separators for separationsin general. `Is it likely that membranes separators mightreplace distillation?’ is a frequently asked question. It isbelieved that the design studies reported here have resultedin reasonable cost estimates for both distillation andmembrane separation. The work is based on cost correlationsdeveloped in previous work and described inHinchliffe and Porter 1 and Porter, Hinchliffe and Tighe 2 .The membrane costs have been calculated using the newparameters of Cost Permeability, Effective Selectivity andthe concept of Targeting. This has permitted the evaluationof membrane technology to go beyond commerciallyavailable membranes, to include laboratory studies ofmembranes and to speculate on what properties would berequired of a new membrane if it were ever to compete withdistillation. The results con®rm current practice, which isperhaps not surprising, as membrane separation is a maturetechnology that has been known for more than twenty years.Figure 11. A plot of water ¯ux pervaporating through a 4A molecular sievemembrane vs concentration of water in a water/propanol mixture.Figure 12. Schematic diagram of molecular sieve membrane separator.Trans IChemE, Vol 78, Part A, March 2000

A COMPARISON OF MEMBRANE SEPARATION AND DISTILLATION267Table 9.Cost/m 2 of Module Capital Operating Cost Heat Exchanger. TAC Separation CostMembrane (£k/m 2 ) Charge (£k/yr) (£k/yr) Capital Charge (£k/yr) (£k/yr) (£/te acetone)7.5 1063 134.8 64.2 1262 27.435.0 717 134.8 64.2 916.0 19.912.5 370 134.8 64.2 569.0 12.372.0 301 134.8 64.2 500.0 10.871.5 231 134.8 64.2 430.0 9.351.0 162 134.8 64.2 361.0 7.850.5 92.7 134.8 64.2 291.7 6.340.0 23.3 134.8 64.2 222.3 4.83Polymeric membrane separators may provide an economicallyattractive solution to separating hydrogen fromother gases. This is because of the high membraneselectivity that results from the rapid rate of diffusion ofthe small hydrogen molecule through the membrane, andbecause, either the gas mixture is already at a high pressureor the retentate is to be used at a high pressure, so that thecompression costs may be in part carried by the other partsof the process. They might also provide a cheaper on-sitealternative to the supply of oxygen or nitrogen in cylindersor as liquids in tankers. Here they replace the distributioncosts rather than the manufacturing cost of the separation ofair. The sizeable market for enriched air or inert nitrogenatmospheres is peculiar in that it requires only a lowrecovery, and/or a low purity product from a low cost feedwith pollution free byproducts. For this the membrane costsare low.In these and other (e.g. dew point control membranes)separations of low molecular weight materials above theircritical temperature, the cryogenic distillation alternative isrelatively more costly and there may be advantages formembranes resulting from the need to transport and processthe materials under pressure. Also the membranes will workat ambient temperatures, or not much more, which meansthey are more likely to achieve a reasonable workinglifetime.There are sound and simple reasons why polymericmembranes are unlikely to ever replace conventionaldistillation. To avoid thermal degradation of the membranesand of the materials being separated, the separation processmust be below the critical temperature of the mixturecomponents, thus the driving force for mass transferdepends on temperature and is relatively small. This willresult in a large and costly membrane permeator, which willprobably have a shorter lifetime.Molecular sieve membranes have been shown to offer anattractive alternative to distillation for many but not allseparations. A molecular sieve membrane with in®niteselectivity (such as the one considered in this work) wouldseparate without the need for recycle and the energy usedwould be low cost energy as in distillation. Anotherattractive application of such a molecular sieve separatorwould be as part of a reactor, removing reaction product soas to increase conversion.Very recently a scaled-up version of the Smart ChemicalCompany membrane evaluated here has become availableand further details of it are awaited with interest.Figure 13. Cost of separation of a 50:50 (molar basis) mixture of acetoneand water using a 4A molecular sieve membrane as a function of the costper unit area of the membrane separator.NOMENCLATURED molar ¯ow rate of distillate from the top of the distillation columnF molar feed rate of mixture to be separatedG permeability of pure gas through a membrane, nm 3 /hr bar m 2G C cost permeability as de®ned in text, nm 3 /hr bar £G C* value of cost permeability used in estimating costs for a newmembraneM membrane module design factorP diffusion coef®cient of a gas or vapour in the membrane polymerQ solubility of a gas or vapour in the membrane polymerR re¯ux ratio in distillationR m minimum re¯ux ratioR opt optimum re¯ux ratioS membrane selectivity, ratio of fast gas permeability to slow gaspermeabilityS* membrane selectivity based on pure gas permeabilitiesS e effective selectivity de®ned in equation (4)V vapour rate required by separationV F vapour rate required by distillation per mole of feedV m vapour rate required by molecular sieve membrane separator permole of feedx mole fraction of most volatile componentx MF feed concentration at which V F equals V Ma relative volatility of mixture components in distillationg module ef®ciency (equation 1)SubscriptsD distillateTrans IChemE, Vol 78, Part A, March 2000

268 HINCHCLIFFE and PORTERF feedA, B fast, slow gases permeating through the membrane1,2 known and new membranes in equation (2)REFERENCES1. Hinchliffe, A.B. and Porter K.E. 1997, Gas separation usingmembranes. 1. Optimisation of the separation process using new costparameters, Ind Eng Chem Res, 36: 821±829.2. Porter, K.E., Hinchliffe, A.B. and Tighe B.J., 1997, Gas separationusing membranes. 2. Developing a new membrane for the separation ofhydrogen and carbon monoxide using the targeting approach. Ind EngChem Res, 36: 830±827.3. Porter, K.E. and Hinchliffe, A.B., 1997, Comparison of MembraneSeparation with Distillation. ETSU Agreement Number: E/CA/00125/00/00/4128, Jan. 1997.4. Saltonstall, C.W., 1987, Calculation of the membrane area required forgas separation. J Membr Sci, 32: 185.5. Porter, K.E. and Jenkins, J.D., 1979, Interrelationship betweenindustrial practise and academic research in distillation and absorption,Distillation 1979, IChemE Symp Ser No 56, 1/1.6. Garrett, D.E., 1989, Chemical Engineering Economics, (Van NostrandReinhold).7. Hinchliffe, A.B., 1991, Separation of hydrogen and carbon monoxideusing polymer membranes, PhD Thesis, (Aston University).8. Lee, D., 1994 Trace heating 2000, Presentation to IChemE WestMidlands at Aston University, 29/9/1994.9. Backhurst, J.R. and Harker, J.H., 1979, Process Plant Design,(Heinmann Educational Books).10. Sinnott, R.K., 1993, in Coulson and Richardson’s Chemical Engineering,Volume 6, Second Edition (Pergamon Press).11. Smart Chemical Co. Ltd., 1995, Innovation Brochure, (SmartChemical Co. Ltd., Horsham, West Sussex).12. Brandrup and Immergut (Ed.), 1989, Polymer Handbook 3rd Edition,(Interscience (John Wiley)).13. Koros, W. J. and Fleming, G. K., 1993, Membrane-based gasseparation, J Memb Sci, 83(1): 1.14. R. Prasad, Shaner, R. L. and Doshi, K. J., 1994, Chapter 11 inPolymeric Gas Separation Membranes, Edited by Paul, D. R. andYampol’skii, Y. P. (CRC Press Inc).15. Heinemann, S. B. and Woodward, C., 1996, Natural gas sweetening bythe combination of semi-permeable membranes and ®xed bedabsorbents, 11 th Offshore South East Asia Conf, World Trade Centre,Singapore. 24±27 September 1996.16. Aitken, C., Jones, K. and Tag, A., 1998, PRISM membrane systems forcost ef®cient natural gas dehydration, Gas Producers Association TechMeeting. 18 February 1998, London.17. Cook, P. J. and Losin, M. S., 1995, Membranes provide cost±effectivenatural gas processing, Hydrocarbon Processing, April 1995, 79±84.18. Giglia, S., Bikson, B., Perrin, J. E. and Donatelli, A. A., 1991,Mathematical and experimental analysis of gas separation by hollow®ber membranes, Ind Eng Chem Res, 30: 1239±1248.19. Pan, C. Y., 1983, Gas separation by permeators with high±¯uxasymmetric membranes, AIChE J, 29(4): 545±552.20. Bhide, B. D. and Stern, S.A., 1991, A new evaluation of membraneprocesses for the oxygen-enrichment of air. Effects of economicparameters and membrane properties, J Membr Sci, 62: 37.21. Mangnall, K., 1989, Vacuum/pressure producing machines andassociated equipment, (Hick Hargreaves & Co. Ltd., Bolton, Lancashire,England).ADDRESSCorrespondence concerning this paper should be addressed to ProfessorK. E. Porter, Speedwell House, 1 Speedwell Road, Edgbaston, BirminghamB5 7PR, UK.The manuscript was recieved 5 March 1999 and accepted for publicationafer revision 3 December 1999.Trans IChemE, Vol 78, Part A, March 2000