Separation of Hexane Isomers through Nonzeolite Pores in ZSM-5

Ind. Eng. Chem. Res. 1999, 38, 2775-2781

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Separation of Hexane Isomers through Nonzeolite Pores in ZSM-5 Zeolite Membranes Christopher J. Gump, Richard D. Noble, and John L. Falconer* Department of Chemical Engineering, University of Colorado, Boulder, Colorado 80309-0424

Permeation of n-hexane and 2,2-dimethylbutane (DMB) through two tubular ZSM-5 zeolite membranes was studied as a function of temperature and isomer partial pressures. The maximum permeation selectivity is 650, obtained at high n-hexane pressures. Though the single-gas isomers permeate at similar rates, the presence of n-hexane decreases the DMB permeance dramatically, whereas n-hexane permeation is unaffected by DMB. The nonzeolite pores appear to be different in size or number in the two membranes, whose pores saturate at different n-hexane pressures. Membranes with smaller nonzeolite pores exhibit pore saturation. The dependence of permeance on partial pressure and temperature indicates that most of the DMB permeance is through nonzeolite pores, which can be blocked by preferential n-hexane adsorption. These pores are larger than zeolite pores but are apparently in the nanopore size range. Introduction Zeolites, which have pores that are the size of small molecules and adsorb some molecules strongly, are used in industrial separation processes.1,2 Zeolite membranes, which consist of a continuous layer of zeolite crystals, have the potential to separate mixtures of compounds that are otherwise difficult to separate. In addition, their inorganic crystalline structure gives them mechanical strength and thermal and chemical stability. Separations using these membranes can be accomplished if one molecule in a mixture (1) is larger than the zeolite pores (molecular sieving), (2) diffuses faster through the pores, or (3) preferentially adsorbs within the pores. Preferential adsorption can block the other component from adsorbing inside the zeolite and thus prevent it from diffusing through the pores. High separation selectivities have been reported for a variety of mixtures using different types of zeolite membranes, with the most common being MFI zeolites such as silicalite and ZSM-5. These zeolites have both straight (0.54 × 0.56 nm) and sinusoidal (0.51 × 0.54 nm) pores [diameters measured by X-ray diffraction (XRD)3] running perpendicular to each other. Silicalite is composed of pure silica, whereas ZSM-5 has aluminum substituted into a small fraction of the silicon sites in the crystal lattice. Most zeolite membranes have been prepared on flat disks. Vroon et al.4 studied single-gas permeation and binary mixture separation using silicalite membranes prepared on R-alumina disks. They found that the single-gas flux of methane was a linear function of the partial pressure between 25 and 100 kPa at 298 K but that the fluxes for ethane, propane, and n-butane were not. At 473 K, the fluxes of all four gases were linear functions of partial pressure. This behavior was attributed to the different adsorption strengths of the molecules. They measured high selectivities for the separation of n-butane/i-butane (>10) and n-hexane/2,2dimethylbutane (DMB) (>2000) mixtures at 473 K and concluded that size exclusion controlled transport through the membrane because adsorption is expected to be negligible at 473 K. * To whom correspondence should be addressed. E-mail: [email protected]. Fax: (303) 492-4341.

Membranes have also been prepared on tubular supports.4,5 This geometry is preferable to flat disks for industrial scale-up (more surface area for a given volume). Membranes on alumina supports were studied by Giroir-Fendler et al.6 and exhibited properties similar to those of membranes grown on flat supports for the separation of butane and hexane isomers. They reported permeate compositions of up to 99.5% n-hexane (sweepgas-free basis) from a 45/55 molar ratio n-hexane/DMB feed and high ratios of fluxes up to 473 K. The type of support used affects the properties of the finished membrane. Kusakabe et al.7 studied membranes grown on R- and γ-alumina tubes of varying pore sizes. The various supports exhibited different permeation behavior, but they were unable to find any direct relationship between the support and permeation. As the fraction of n-butane in the feed increased, the permeance of n-butane increased and the permeance of i-butane decreased, leading to a 50% increase in the permselectivity over the 25-75% n-butane range. Funke et al.8 studied single-gas, binary, and ternary permeations for n-octane, i-octane, and n-hexane through tubular, γ-alumina-supported silicalite membranes and found that n-hexane permeated fastest as a pure component but n-octane preferentially permeated in the mixtures. They also found a linear dependence of flux with organic concentration in the feed at 416 K and concluded that the membrane was not saturated with organics, even at partial pressures as high as 18 kPa. Coronas et al.9,10 found that the performance of membranes prepared on R- and γ-alumina tubes by hydrothermal synthesis depended on both the support and the synthesis procedure. Their membranes showed separation selectivities for n-hexane/DMB as high as 2580 at 374 K, but the selectivities decreased dramatically as the temperature increased, so that no separation was obtained at 410 K. Because the ideal selectivity was only 2.1 at 374 K for n-hexane/DMB, separation was concluded to be due to preferential adsorption rather than molecular sieving. The current study uses membranes prepared by the same techniques as Coronas et al.9,10 to study n-hexane/ DMB separations in more detail with two membranes. Single- and mixture-gas permeation rates were com-

10.1021/ie980721v CCC: $18.00 © 1999 American Chemical Society Published on Web 06/11/1999

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pared to understand the transport processes and the types of pores present in the membrane. Mixture-gas permeances were measured as a function of feed composition and temperature to characterize the effects of adsorption and the ability of the linear hydrocarbon to block the branched hydrocarbon through preferential adsorption. Experimental Methods Membrane Preparation. Two ZSM-5 membranes, M1 and M2, were prepared on porous asymmetric γ-alumina supports, using procedures described previously.10 A tubular support from Golden Technologies was used to prepare membrane M1. The 4.7-cm-long support (0.65-cm i.d.) was asymmetric; most of the structure was R-alumina with 200-nm pores, and the inside surface had a 5-µm-thick layer of γ-alumina with 5-nm pores. The support for membrane M2 was obtained from U.S. Filter and had similar dimensions and pore structure. Before synthesis, approximately 1 cm on each end of the support was sealed with a hightemperature glaze (GL 611A, Duncan) to prevent bypass across the membrane. The supports were fired to 1070 K to set the glaze. This high-temperature treatment affects the support structure and pore size of the γ-alumina layer. Permeation measurements using supports on which the γ-alumina layer was deposited before or after the high-temperature glazing showed that the N2 permeance was 140 times larger when the γ-alumina layer was deposited before glazing. The synthesis gel used to grow the ZSM-5 membrane was based on the silicalite gel of Grose and Flanigen;11 the aluminum source was Na2Al2O4‚3H2O, the chargebalancing counterion was sodium, and the Si/Al atomic ratio was 100. The molar component ratio of the modified gel was Na2Al2O4‚3H2O:TPAOH:NaOH:SiO2: H2O ) 0.105:1:3:21:987, where TPAOH is tetrapropylammonium hydroxide. Electron probe microanalysis on membranes prepared by the same procedure9 yielded Si/Al ratios of less than 100, apparently because the basic synthesis gel dissolved alumina from the support and incorporated it into the membrane. The membrane was synthesized on the internal wall of the support by plugging one end of the support with Teflon tape and a Teflon end cap. The inside of the support tube was filled with approximately 2 mL of gel, and then the top end was taped and capped. The capped tube was then placed inside a Teflon-lined, acid digestion bomb (Parr) along with approximately 1 mL of water. The bomb was sealed and placed in an oven at 440 K for 16 h. The membrane was dried at 440 K, and its permeability was tested with N2 at 222 kPa and room temperature. When the nonzeolite pores are sufficiently small, the template molecules fill and block them as they do the zeolite pores and the membrane is impermeable. For membrane M1, a second and third layer was similarly synthesized on the support for 8 h each. After the third layer, membrane M1 was impermeable to N2. Membrane M2 was sealed after two layers. Both membranes were calcined in a muffle furnace at 753 K for 8 h to remove the template molecules from the pores, making the membrane permeable. The oven was heated and cooled at 0.01 and 0.05 K/s, respectively, to minimize thermal stresses during calcination. Membranes synthesized previously by this procedure9 consisted of a 25-30-µm-thick, continuous layer of zeolite crystals on the inside surface of the support. For their mem-

branes, the support did not have a large influence on the permeation properties. Once calcined, the membranes were sealed in a deadend permeation system to test their quality. The ambient-temperature, single-gas permeation rates of N2 and SF6 were measured at a feed pressure of 222 kPa and a permeate pressure of 84 kPa. Because the kinetic diameter of N2 is smaller than that of zeolite pores whereas the kinetic diameter of SF6 is similar, the ratio of the two permeances was used as an indication of the quality of the membrane. Membrane M1 had an N2/SF6 ideal selectivity of 140, and M2 had an ideal selectivity of 250. Separation Apparatus. Hexane isomer mixtures were separated in a continuous-flow system that utilized a cross-flow permeation cell, described in detail elsewhere.8 The vapor feed flowed axially through the membrane tube, and the components permeated radially outward. Silicone O-rings sealed the membrane into the module and prevented bypass over the glazed ends of the support. n-Hexane (Aldrich, 99+%) and DMB (Aldrich, 99+%) were fed to the system as a mixture of liquids via a syringe pump. The feed liquid was vaporized into a preheated helium stream while passing through a heated zone maintained at a minimum temperature of 373 K, which is well above the boiling point of the hexanes. This combined vapor feed then flowed to the membrane module. On the permeate side of the module, a helium sweep gas continuously removed the permeation products from the external surface of the membrane. Helium flow rates for both the feed and the sweep were set at about 40 cm3/min at STP using mass flow controllers. A bypass line allowed the concentration of the hexanes in the feed to be independently measured. To prevent condensation of the organics on the tubing walls, all system lines were wrapped with heating tape and insulated. Thermocouples were placed at various points around the system to allow for the identification and elimination of potential cold spots. The temperatures of the system lines were consistently above 373 K. The highest partial pressure of organics used in the experiments was about 30 kPa, which was well below the lowest dew point of the vapor feed mixtures used (175 kPa at 353 K). Three additional thermocouples were placed in the module feed entrance, sweep entrance, and permeate exit lines. Most experiments were conducted at a membrane temperature of 373 K, which had been found to give high permeation rates, good selectivities, and long periods of separation activity.10 To investigate the temperature dependence of the permeation, some experiments were run at 353 and 398 K. The permeate and retentate streams were analyzed with an HP 6890 gas chromatograph equipped with a flame-ionization detector. A 10-port pneumatic sampling valve was used to switch between the retentate and permeate streams. The chromatographic separation was accomplished using a 6-ft Alltech AT-1200 packed column. The volumetric flow rates of the retentate and permeate streams were measured at atmospheric temperature and pressure (typically 299 K and 84 kPa) using two soap film flowmeters. Procedure. The membranes were calcined for 8 h at 753 K to remove contaminants prior to each experimental run at a given liquid feed composition or temperature. The membranes were purged with helium in the

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Figure 1. Flux of n-hexane versus the mole fraction of organics in helium at 373 K for membrane M1. Feed gas consists of the indicated percentages of n-hexane (balance DMB) in a helium carrier gas.

Figure 3. Separation selectivity for n-hexane/DMB mixtures of indicated percentages of n-hexane (balance DMB) at 373 K and feed concentrations in helium for membrane M1. Table 1. Stream Compositions (Helium-Free Basis) as a Function of the n-Hexane Fraction in the Liquid Feed and the Organic Fraction in the Vapor Feed for Membrane M1 % n-hexane in liquid feed

% organic in feed

80

4.0 23.7 3.9 23.1 4.6 23.9 4.6 27.4 4.9 29.7

60 50 20 10 Figure 2. Flux of DMB versus the mole fraction of organics in helium at 373 K for membrane M1. Feed gas consists of indicated percentages of n-hexane (balance DMB) in a helium carrier gas.

flow system at 373 K before the flow of organics was started, and the system was given approximately 1.5 h to reach steady state. Coronas et al.10 measured the permeances of both isomers in membranes prepared by the same procedure. At 375 K, the permeances reached steady state in less than 0.5 h, and the n-hexane permeance slowly decreased to half its initial value after 17 h. The DMB permeance decreased slower than that of n-hexane. To minimize membrane deactivation, the current experiments were run for less than 8 h. During the course of an experiment, the total organic mole fraction in the feed stream was increased from approximately 0.05 to 0.30 by changing the syringe pump flow rate. Hysteresis effects were checked by decreasing the mole fraction from 0.30 to 0.05 at the end of each experiment but were not seen. Because the module has a cross-flow design, a log mean pressure drop was used to calculate the driving force from the partial pressures. The two pressure drops used in the calculation were between the partial pressures of the feed and the permeate and the partial pressures of the retentate and the sweep (which was pure helium). Results Membrane M1. Figures 1 and 2 show the fluxes of n-hexane and DMB for membrane M1 at 373 K as a function of the total organic fraction in the feed (balance helium) and for various percentages of n-hexane in the liquid organic feed. The flux of n-hexane increases with increased concentrations of n-hexane in the feed, whether

mole fraction of n-hexane retentate permeate 0.702 0.734 0.458 0.474 0.370 0.361 0.126 0.121 0.085 0.063

0.944 0.997 0.818 0.994 0.849 0.983 0.350 0.723 0.173 0.279

the increase is due to a higher percentage of n-hexane in the liquid feed for a given mole fraction of organics in the feed or a higher concentration of total organic at a fixed n-hexane fraction. In contrast, the flux of DMB either increases more slowly or decreases as the mole fraction of organics in the feed increases. At higher concentrations of n-hexane, the DMB flux decreases as the total organic fraction increases. That is, even though the concentration of DMB increases, the DMB flux decreases. However, when n-hexane in the feed is low, such as for the 10% n-hexane feed, the flux of DMB increases with increasing DMB concentration. The separation selectivity, defined as the ratio of permeances, is given in Figure 3 for membrane M1. The selectivity increases linearly with the mole fraction of organics in the feed and with increasing percentage of n-hexane in the feed, with the exception of the selectivity for 80% n-hexane, which is lower than that for 60% n-hexane. For comparison with the selectivities, Table 1 lists representative feed and permeate concentrations on a helium-free basis. Membrane M2. The single-gas fluxes at 373 K for n-hexane and DMB through membrane M2 as a function of the feed partial pressure are given in Figure 4. Both fluxes approach limiting values with increasing partial pressure in the feed, with the n-hexane flux being about twice the DMB flux. Figures 5 and 6 show the fluxes of the two components for equimolar mixtures, which were separated at 353, 373, and 398 K. The flux of n-hexane at various pressures is almost the same for the single gas and the mixtures. The flux of DMB in mixtures is about 2 orders of magnitude lower than the single-gas flux. Table 2 lists representative feed and permeate concentrations on a helium-free basis.

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Figure 4. Single-gas fluxes versus partial pressure for n-hexane and DMB in a helium carrier gas at 373 K for membrane M2.

Figure 5. Flux of n-hexane versus feed partial pressure of n-hexane at 353, 373, and 398 K for membrane M2 (50/50 feed mixture and helium carrier gas).

Figure 6. Flux of DMB versus feed partial pressure of DMB at 353, 373, and 398 K for membrane M2 (50/50 feed mixture and helium carrier gas). Table 2. Stream Compositions (Helium-Free Basis) as a Function of the Temperature and the Organic Fraction in the Vapor Feed for Membrane M2 temperature (K)

% organic in feed

353

3.9 27.5 4.2 1.9 2.4 7.0 16.3 26.9

373 399

mole fraction n-hexane retentate permeate 0.454 0.440 0.405 0.410 0.378 0.383 0.382 0.375

0.997 0.998 0.997 0.996 0.897 0.964 0.996 0.998

The flux of n-hexane in mixtures is relatively independent of temperature at low pressure, but at higher pressures the fluxes at different temperatures are not the same. In contrast, the flux of DMB undergoes a dramatic change with temperature. At low temperature, it increases with increasing organic fraction in the feed, but the flux at 398 K and low partial pressures is more than 2 orders of magnitude larger than that at the lower

Figure 7. Permeance of n-hexane and DMB as a function of n-hexane feed partial pressure at 353, 373, and 398 K for membrane M2 (50/50 feed mixture and helium carrier gas).

Figure 8. Separation selectivity versus n-hexane feed partial pressure at 353, 373, and 398 K for membrane M2 (50/50 feed mixture and helium carrier gas).

temperatures and then decreases with increasing organic fraction in the feed. At high partial pressures, the flux approaches the same value as that seen at lower temperatures. The permeances of n-hexane and DMB at the three temperatures are shown in Figure 7. As with the flux, the n-hexane permeance does not change much with temperature. For the lower temperatures, the DMB permeance is also relatively insensitive to temperature. At 398 K, however, the permeance of DMB is much higher at low partial pressures and decreases by 2 orders of magnitude as the partial pressure increases. This behavior is reflected in the separation selectivity (Figure 8). At 353 and 373 K, the separation selectivity quickly rises to a limiting value of about 650 at low partial pressures, but for 398 K the separation selectivity is only 1/20 of that measured at the lower temperatures in the low-pressure regime. In addition, it does not increase to the same high selectivities of 600-700 until the partial pressure of n-hexane is about 12 kPa. The flux of n-hexane is the same order of magnitude for membranes M1 and M2, but the flux of DMB in mixtures is 1-2 orders of magnitude lower for M2, depending on the partial pressures. Membrane Comparison. Figure 9 shows the nhexane flux versus n-hexane partial pressure. Note that all of the data for membrane M1 in Figure 1 collapse into a single line: the n-hexane flux increases linearly with the n-hexane partial pressure, independent of the DMB partial pressure. The n-hexane flux for membrane M2 is almost the same as that for membrane M1. The DMB permeance decreases with increasing pressure of n-hexane for both membranes, as shown in Figure 10 with permeance on a log scale. The difference in the separation behavior of membranes M1 and M2 is due

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Figure 9. Flux of n-hexane in a mixture with DMB versus n-hexane partial pressure in the feed at partial pressures of DMB from 0.7 to 23 kPa for membranes M1 and M2 at 373 K.

Figure 10. Permeance of DMB in a mixture with n-hexane versus the feed partial pressure of n-hexane for membranes M1 and M2 at 373 K.

to the difference in DMB permeance through each. If flux is plotted instead of permeance, the behavior is similar to that in Figure 10 but with a greater spread in the data. Because the driving force is included in the permeance values, the permeation of DMB depends on both the pressure of n-hexane and the change in partial pressure of DMB from the retentate to permeate sides of the membrane. Discussion Preferential Adsorption. The single-component permeance for n-hexane is only 2-3 times higher than that for DMB at 373 K in membrane M2 (Figure 4), and similar ideal selectivities were reported by Coronas et al.10 for membranes made by the same procedure. Therefore, the high separation selectivities measured for membrane M2 are not due to different diffusion rates or molecular sieving. When the feed consists of significantly more DMB than n-hexane (Figure 2, 10% and 20% curves), the flux of DMB increases with increasing pressure. As the total pressure of n-hexane continues to increase, the rate of increase of DMB flux slows (10% curve) or the flux decreases (20% curve) because nhexane blocks DMB permeance. The permeance of DMB in the mixture increases rapidly with temperature (Figure 6) because n-hexane coverage decreases, and it decreases exponentially as the n-hexane pressure increases (Figure 10). The selectivity increases linearly with n-hexane pressure for membrane M1 over the entire experimental range. However, the selectivity for membrane M2 only increases linearly with n-hexane pressure at 398 K and only after increasing significantly

slower at low pressure. At 358 and 373 K and low n-hexane pressure, DMB permeation through membrane M2 is dramatically inhibited (Figure 6), and selectivity quickly rises with increasing n-hexane pressure to a limiting value. At 398 K, DMB flux through membrane M2 is greater at low pressure of n-hexane, but it decreases more as the n-hexane pressure increases. Thus, the membranes have high separation selectivities because n-hexane preferentially adsorbs and blocks pores. When the n-hexane pressure is sufficiently low, DMB permeates through the membrane. As the nhexane pressure increases, DMB is prevented from adsorbing and therefore cannot permeate through the membrane. For higher n-hexane pressures, the adsorption coverage is higher and n-hexane is even more effective at blocking DMB adsorption. The n-hexane fluxes are similar for both membranes (Figure 9), even though their separation performances are different. Separation performance depends on how well adsorbed n-hexane blocks DMB permeation; n-hexane is more effective in membrane M2 (Figure 10). For similar membranes, Funke et al.8 and Coronas et al.9 concluded preferential adsorption was the reason for separation based only on the temperature dependence of the permeances. The selectivity decreases at higher temperature because the n-hexane coverage is lower and it is thus less effective at blocking DMB. Membrane Pores. The heat of adsorption for nhexane is about 30% higher than that for DMB on silicalite (16.7 versus 13.0 kJ/mol), and the saturation coverage is almost twice as high (10 versus 5.6 wt %),12 so the pure-component adsorption coverages are expected to be higher for n-hexane. The diffusivity for n-hexane in ZSM-5 pores is also much larger than that for DMB because the kinetic diameter of DMB (0.62 nm) is larger than the XRD-measured diameter for ZSM-5 pores, whereas the kinetic diameter for n-hexane (0.43 nm) is much smaller. As measured by both gravimetric and chromatographic sorption studies, the diffusivity of DMB in ZSM-5 is on the order of 10-19 m2/s13 and typical n-hexane diffusivities are 10-12 m2/s.14,15 Indeed, silicalite membranes prepared by Vroon et al.16 and GiroirFendler et al.6 have high ideal and mixture selectivities for n-hexane/DMB, as might be expected for molecular sieving behavior based on the large size difference. Moreover, we have recently prepared H-ZSM-5 membranes that also exhibit both ideal and mixture selectivities greater than 1200.17 Thus, membranes M1 and M2 do not exhibit the large difference in single-gas permeances of these C6 isomers that is expected for zeolite membranes with few defects. Apparently, DMB permeance is large because most of the DMB diffuses through nonzeolite pores. Therefore, the high mixture selectivities appear to be due to preferential adsorption of n-hexane in the nonzeolite pores. The concentration dependence measured for membrane M1 also indicates transport through nonzeolite pores. The linear dependence of the n-hexane flux on the n-hexane partial pressure in membrane M1 indicates that the membrane pores are not saturated at a partial pressure of 16 kPa but are in the Henry’s law region. Sun et al.18 showed that n-hexane adsorption was above the Henry’s law portion of the isotherm on silicalite at 373 K and 0.4 kPa. Because Dunne et al.19,20 observed that silicalite and ZSM-5 zeolite have similar isotherms for smaller alkanes but adsorption is stronger

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on ZSM-5, n-hexane is expected to be above the Henry’s law portion of the adsorption isotherm at lower pressures than 0.4 kPa. Thus, at 16 kPa the zeolite pores are close to saturation. Because the n-hexane flux showed a linear dependence on partial pressure up to 16 kPa for membrane M1 and the selectivity increased as the n-hexane partial pressure increased, the pores used for transport must be in the Henry’s law region at this pressure. That is, both n-hexane and DMB must transport through nonzeolite pores that have different adsorption properties from ZSM-5 pores. These nonzeolite pores have low ideal selectivities (ratio of singlegas permeances) for C6 isomers. The nonzeolite pores are not expected to have uniform sizes and shapes like the zeolite pores; instead, a distribution of sizes is expected. However, their diameters must be similar to those of the zeolite pores. Otherwise, the membranes would not exhibit the high N2/SF6 ideal selectivities measured (140 and 250) and would not separate organic mixtures. Because DMB is slightly larger than the XRD diameter of ZSM-5, the nonzeolite pores would not have to be much larger than the ZSM-5 pores to make the DMB flux much larger and close to that of n-hexane. Molecules move through the nonzeolite pores both by surface diffusion and through the gas phase. As the pore size increases, gas translation becomes more important, and preferential adsorption will be less effective in blocking permeation of the branched isomer. For pores 1.4 nm in diameter, a monolayer of adsorbed n-hexane would reduce the effective pore size so that DMB could not permeate through the gas phase. Simulations of adsorption in silicalite pores at 300 and 362 K indicate that n-hexane effectively blocks 2-methylpentane through adsorption and pore packing.21,22 Even a low pressure of n-hexane prevents the permeation of the branched hydrocarbon through the zeolite pore. Nonzeolite pores may exhibit a similar preferential packing. Preferential adsorption is the mode of separation for nanoporous carbon membranes.23 These membranes separate mixtures of light linear hydrocarbons and hydrogen and recover as much as 98.4% of the hydrogen in the retentate. Apparent pore size distributions for these membranes, as determined by adsorption studies and methane diffusivity, indicate that pores of 0.5-0.6 nm predominate.23,24 Thus, adsorption in zeolite pores does not appear to control the C6 separations for membranes M1 and M2. Coronas et al.9 showed that after exposing their membrane to both isomers at 463 K, the n-hexane permeance decreased to 1% of its original value at 374 K. In contrast, the DMB permeance in the mixture increased an order of magnitude. Because calcination restored the original membrane permeances, the change in membrane properties was likely due to coke deposition that changed the adsorption properties of the nonzeolite pores. The adsorbed n-hexane phase in the nonzeolite pores could be similar to a condensed phase. Although the pressures used were well below the saturation pressure, organics can condense in membrane pores by capillary condensation. The Kelvin equation (eq 1) governs the reduction in vapor pressure of a gas in small pores. In

pred ) p*e(-2γVm/RTr) cos θ

(1)

this equation, p* is the normal vapor pressure, γ is the surface tension of the liquid, Vm is the molar volume of the liquid, R is the gas constant, T is the temperature,

θ is the contact angle for wetting, and r is the radius of the pore. This equation does not accurately predict reduced vapor pressures for pore diameters below 2 nm. Because the surface tension and molar volume lose meaning on this scale, the Kelvin equation overestimates the vapor pressure. The lowest saturation pressure was for the 50/50 mixture of isomers at 353 K for membrane M2. Physical property data were estimated by the ASPEN PLUS process simulator and the PengRobinson equation of state. The reduced vapor pressure in a 0.5-nm-diameter pore with a wetting angle of zero (cos θ ) 1) for n-hexane at 353 K is 16 kPa. For 1-nmdiameter pores, the saturation pressure of n-hexane is 46 kPa. A condensed phase may form under these experimental conditions, though the meaning of a condensed phase for such small pores is unclear because these pore diameters are only a few molecular diameters across. Membranes M1 and M2. Both the size and the number of nonzeolite pores control separation selectivity. In a membrane with smaller nonzeolite pores, n-hexane blocks the pores at lower partial pressures and higher temperatures. Membrane M1 has a lower nhexane/DMB separation selectivity than M2 apparently because M1 has more pores that are not blocked at a given pressure. Because there are fewer of these larger, nonblocked pores in membrane M2, the N2/SF6 ideal selectivity for membrane M2 is about twice that for membrane M1. In addition, the n-hexane flux for M2 is not as linear with n-hexane pressure as the flux through M1 and instead appears to approach an asymptote. The deviation is strongest at lower temperatures, as expected if the nonzeolite pores are filled with n-hexane. This behavior is not seen for membrane M1, even at the highest partial pressures, apparently because its pores are larger and the adsorbed n-hexane cannot completely fill the pores. Thus, membrane M1 is less effective for separations because its pores are larger. Membranes with nonzeolite pores that are larger than the ZSM-5 pores can therefore still be selective for C6 isomer separations. Membrane M2 has a maximum selectivity of 650 and purified a 50/50 mixture of isomers to over 99.8% n-hexane. However, an adsorption-based membrane cannot operate at high temperature. Coronas et al.9,10 reported that separation selectivity quickly dropped to one as the temperature increased to 410 K and that, upon exposure to n-hexane at high temperature, the flux of n-hexane decreased with time because of decomposition in the pores. The loss of selectivity with temperature can be countered by operating at higher partial pressures, as shown in Figure 8. When sufficiently high pressures are used, n-hexane adsorption increases and the selectivity increases. The problem of membrane fouling remains. The sodium form of the ZSM-5 zeolite is not catalytically active for the cracking reaction, although a small fraction of the sites in the zeolite are expected to be in the acid form. The crystal defects of the nonzeolite pores may contain more active sites. This would explain why the membranes of Coronas et al. fouled at higher temperatures, resulting in lower permeances, whereas such fouling has not been reported in molecular sieving membranes.7,17,24 Apparently only the nonzeolite pores contained a significant fraction of active sites. Once the nonzeolite pores are coked, n-hexane probably adsorbs more weakly and is less effective at blocking DMB. To avoid membrane fouling, the separations can be con-

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ducted at lower temperatures. For both our membranes and those reported in the literature,6,25,26 the flux of linear alkanes is a weak function of temperature, so the yield would be similar at lower temperatures. Conclusions Zeolite membranes that contain small-diameter, nonzeolite pores can separate n-hexane/DMB mixtures at 373 K with selectivities as high as 650. Membranes with the smaller nonzeolite pores have these high selectivities. The difference in diffusion rates through the membranes cannot account for these mixture selectivities, because the ideal selectivity is only 2-3. Instead, selectivity is the result of preferential adsorption and pore blocking; n-hexane effectively blocks the permeation of DMB through the nonzeolite pores over a wide range of concentrations. The pores are selectively filled by n-hexane, which then hinders the adsorption and transport of DMB. Acknowledgment is made to the donors of the Petroleum Research Fund, administered by the American Chemical Society, for partial support of this research. We also gratefully acknowledge support by the New Energy and Industrial Technology Development Organization (NEDO) of Japan. We thank David J. Peters and Dr. Xiao Lin for preparing the membranes used. Literature Cited (1) Ruthven, D. M.; Farooq, S.; Knaebel, K. S. Pressure Swing Adsorption; VCH Publishers: New York, 1994. (2) Yang, R. T. Gas Separation by Adsorption Processes; Butterworth: Stoneham, MA, 1987. (3) Flanigen, E. M.; Bennett, J. M.; Grose, R. W.; Cohen, J. P.; Patton, R. L.; Kirchner, R. M. Silicalite, a new hydrophobic crystalline silica molecular sieve. Nature 1978, 271, 512. (4) Bai, C.; Jia, M.-D.; Falconer, J. L.; Noble, R. D. Preparation and separation properties of silicalite composite membranes. J. Membr. Sci. 1995, 105, 79. (5) Kusakabe, K.; Yoneshige, S.; Murata, A.; Morooka, S. Morphology and gas permeance of ZSM-5-type zeolite membrane formed on a porous R-alumina support tube. J. Membr. Sci. 1996, 116, 39. (6) Giroir-Fendler, A.; Peureux, J.; Mozzanega, H.; Dalmon, J. A. Characterization of a Zeolite Membrane for Catalytic Membrane Reactor Application. Stud. Surf. Sci. Catal. 1996, 101, 127. (7) Kusakabe, K.; Murata, A.; Kuroda, T.; Morooka, S. Preparation of MFI-type zeolite membranes and their use in separating n-butane and i-butane. J. Chem. Eng. Jpn. 1997, 30, 72. (8) Funke, H. H.; Kovalchick, M. G.; Falconer, J. L.; Noble, R. D. Separation of Hydrocarbon Isomer Vapors with Silicalite Zeolite Membranes. Ind. Eng. Chem. Res. 1996, 35, 1575. (9) Coronas, J.; Falconer, J. L.; Noble, R. D. Characterization and Permeation Properties of ZSM-5 Tubular Membranes. AIChE J. 1997, 43, 1797.

(10) Coronas, J.; Noble, R. D.; Falconer, J. L. Separations of C4 and C6 Isomers in ZSM-5 Tubular Membranes. Ind. Eng. Chem. Res. 1998, 37, 166. (11) Grose, R. W.; Flanigen, E. M. Crystalline Silica; Grose, R. W., Flanigen, E. M., Eds.; Union Carbide Corp.: Danbury, CT, 1977. (12) Cavalcante, C. L., Jr.; Ruthven, D. M. Adsorption of Branched and Cyclic Paraffins in Silicalite. 1. Equilibrium. Ind. Eng. Chem. Res. 1995, 34, 177. (13) Post, M. F. M.; van Amstel, J.; Kouwenhoven, H. W. Diffusion and Catalytic Reaction of 2,2-Dimethylbutane in ZSM-5 Zeolite. In 6th International Zeolite Conference; Olson, D. H., Bisio, A., Eds.; Butterworths: Washington, DC, 1983: p 517. (14) Choudhary, V. R.; Nayak, V. S.; Mamman, A. S. Diffusion of Straight- and Branched-Chain Liquid compounds in H-ZSM-5 Zeolite. Ind. Eng. Chem. Res. 1992, 31, 624. (15) Talu, O.; Sun, M. S.; Shah, D. B. Diffusivities of n-Alkanes in Silicalite by Steady-State Single-Crystal Membrane Technique. AIChE J. 1998, 44, 681. (16) Vroon, Z. A. E. P.; Keizer, K.; Gilde, M. J.; Verweij, H.; Burggraaf, A. J. Transport properties of alkanes through ceramic thin zeolite MFI membranes. J. Membr. Sci. 1996, 113, 293. (17) Flanders, C. L.; Tuan, V. A.; Noble, R. D.; Falconer, J. L. High-Temperature Separations of Organic Vapors, in preparation. (18) Sun, M.; Talu, O.; Shah, D. B. Adsorption Equilibria of C5-C10 Normal Alkanes in Silicalite Crystals. J. Phys. Chem. 1996, 100, 17276. (19) Dunne, J. A.; Mariwala, R.; Rao, M.; Sircar, S.; Gorte, R. J.; Myers, A. L. Calorimetric Heats of Adsorption and Adsorption Isotherms. 1. O2, N2, Ar, CO2, CH4, C2H6, and SF6 on Silicalite. Langmuir 1996, 12, 5888. (20) Dunne, J. A.; Rao, M.; Sircar, S.; Gorte, R. J.; Myers, A. L. Calorimetric Heats of Adsorption and Adsorption Isotherms. 2. O2, N2, Ar, CO2, CH4, C2H6, and SF6 on NaX, H-ZSM-5, and NaZSM-5 Zeolites. Langmuir 1996, 12, 5896. (21) Krishna, R.; Smit, B.; Vlugt, T. J. H. Sorption-Induced Diffusion-Selective Separation of Hydrocarbon Isomers Using Silicalite. J. Phys. Chem. A 1998, 102, 7727. (22) Vlugt, T. J. H.; Krishna, R.; Smit, B. Molecular Simulations of Adsorption Isotherms for Linear and Branched Alkanes and their Mixtures in Silicalite. J. Phys. Chem. B 1999, 103, 1102. (23) Rao, M. B.; Sircar, S. Performance and pore characterization of nanoporous carbon membranes for gas separation. J. Membr. Sci. 1996, 110, 109. (24) Mariwala, R.; Acharya, M.; Foley, H. Adsorption of Halocarbons on a Carbon Molecular Sieve. Microporous Mater. 1998, 22, 281. (25) Funke, H. H.; Argo, A. M.; Falconer, J. L.; Noble, R. D. Separations of Cyclic, Branched, and Linear Hydrocarbon Mixtures Through Silicalite Membranes. Ind. Eng. Chem. Res. 1997, 36, 137. (26) Vroon, Z. A. E. P. Synthesis and Transport studies of thin Ceramic Supported Zeolite (MFI) Membranes. Ph.D. Dissertation, University of Twente, Twente, The Netherlands, 1995.

Received for review November 17, 1998 Revised manuscript received April 5, 1999 Accepted April 20, 1999 IE980721V