9544
Ind. Eng. Chem. Res. 2008, 47, 9544–9551
Recovery of Petroleum Ether from Solanesol Extracting Solution through Vacuum Hydrophilic Membrane Distillation Xin Y. Qu, Lin Zhang,* De S. Tang, Zhi J. Zhou, and Huan L. Chen College of Mater. Sci. & Chem. Eng., Zhejiang UniVersity, Hangzhou, 310027, China
The objective of this work is to study the feasibility of vacuum membrane distillation (VMD) by hydrophilic membranes to recover petroleum ether from the extracting solution of solanesol. Being different from the traditional membrane distillation (MD) process, hydrophilic membranes were used to avoid the membrane wetting for the nonaqueous solutions. The VMD performances of polyvinylidene fluoride (PVDF) membrane and two kinds of polyacrylonitrile (PAN) membranes with different structures were compared in the process of recovering petroleum ether. The results showed that good solvent flux (>15 kg/m2 h) and solanesol rejection (>98%) for the PAN membranes were obtained, indicating that PAN membranes had considerable potential use in this area, while the PVDF membrane was not appropriate for the nonaqueous solution system because of membrane wetting. The effects of operation conditions on VMD performances of the two PAN membranes were investigated, and it was found that high-feed temperatures, low-downstream pressures enhanced the permeate flux for both membranes, and the flux of PAN1 membrane decreased more obviously than that of PAN2 membrane under high-feed concentration because of its larger pore size. General models including Knudsen and viscous flows were proposed, and a good agreement between the experimental and the theoretical fluxes was obtained. Additionally, temperature and concentration polarizations were proved to have an obvious influence on mass transport of PAN1 membrane. It was advisable to work at the downstream pressure of 20 kPa, the feed temperature of 30 °C, and the feed flow rate of 27.5 mL/s to obtain favorable results for PAN2 membrane. 1. Introduction Generally, membrane distillation (MD) is defined as a thermally driven process in which a microporous hydrophobic membrane acts as a physical support to separate a warm solution, and the temperature difference causes a vapor pressure difference between the membrane sides. There are different ways to make the vapor pressure difference across the membrane to drive a flux: direct contact membrane distillation (DCMD),1 air gap membrane distillation (AGMD),2 sweeping gas membrane distillation (SGMD),3 and vacuum membrane distillation (VMD).4 On the basis of the definition of MD, the main requirement of MD is that the membrane must not be wetted. The relationship between a membrane’s largest allowable pore size and operating conditions is given by the Laplace (Cantor) equation:5 Pliquid - Pvapor ) ∆Pinterface < ∆Pentry )
-2BγL cos θ rmax
(1)
where ∆Pentry is the penetration pressure, γL is the liquid surface tension, θ is the liquid-solid contact angle, rmax is the largest pore radius, and B is a geometric factor determined by pore structure. The wettability is determined by the interaction between the liquid and the polymeric material. There is no wetting occurring at low affinity and the contact angle will have a value greater than 90°, whereas with high affinity the contact angle will be less than 90° and the liquid will wet the surface. As MD is now being observed mostly on the aqueous solutions, hydrophobic membranes have received increasing attention for their use in the MD process. At present, polypropylene (PP),6-9 polytetrafluoroethylene (PTFE),10 and polyvinylidene fluoride (PVDF)11-15 have often been used for the MD process. Among * To whom correspondence should be addressed. Tel: +86-57187952121; e-mail:
[email protected].
the polymers, PVDF is more and more widely used because of its good hydrophobicity, chemical resistance, and thermal stability. Similarly, the definition of hydrophilic membrane distillation is also based on the Laplace (Cantor) equation as illustrated in eq 1. A nonaqueous solution is contacted upstream with a hydrophilic microporous membrane, and the hydrophilicity of the membrane prevents the organic solution from entering the pores resulting in a vapor-liquid interface at each pore entrance. Figure 1a illustrates how the vapor-liquid interfaces are supported at the pore openings of the hydrophilic microporous membrane in the organic solution. In Figure 1a, volatile solvent evaporates from the liquid-vapor interface on the feed side and then diffuses or transfers across the hydrophilic membrane and is either condensed or removed in the permeate side. On the contrary, as Figure 1b shows, the hydrophobic membrane is wetted by the organic solvent because the high affinity between the organic solution and the membrane results in ∆Pentry < 0 < ∆Pintererface. So, the hydrophilic membrane can be used in MD process to treat organic solution. Polyacrylonitrile (PAN) membrane is popular in ultrofiltration process as it combines desired chemical stability with a relatively hydrophilic character, so PAN membrane is chosen in this experiment. Solanesol, a polyisoprenoid alcohol, has gained attention because of its value as a source of isoprene units for the synthesis of metabolically active quinones, for example, ubiquinones and vitamin K analogues.16 Solanesol mostly exists in tobacco leaf, and it is a small amount. Generally, solanesol is extracted by the solvent method and is purified by elution and recrystallization, and in the process, a lot of solvent is used. In our laboratory, a study of extraction and purification of solanesol from ground tobacco leaves has been carried out.17 Petroleum ether is applied as the extractant and a large amount of petroleum ether is consumed in the process, so the recovery of petroleum ether is necessary. Conventional methods of the solvent recovery
10.1021/ie071499+ CCC: $40.75 2008 American Chemical Society Published on Web 11/01/2008
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Figure 1. A schematic presentation of hydrophilic MD process with two kinds of membranes: (a) hydrophilic membrane and (b) hydrophobic membrane.
such as distillation are energy-consuming processes, since the extract solution is very dilute; therefore, new methods should be developed to enhance the efficiency of the recovery process. As is mentioned above, MD is conceived as a process that would operate with a minimum energy requirement and at a minimum cost and with a minimum land requirement for the plant. Recently, VMD has been employed to simply both remove an organic component from wastewater and recover valuable products to a high degree of purity.11,18-24 In this paper, for confirming the feasibility that hydrophilic membranes should be applied to the MD process for organic solvents, VMD experiments by using hydrophilic PAN membranes were carried out to recover petroleum ether from the extraction solution of solanesol. To achieve the optimum operating conditions for petroleum ether recovery by VMD, the effects of feed temperature, permeate pressure, and initial feed solanesol concentration on permeate flux as well as on rejection were investigated. 2. Theory Numerous experimental and theoretical studies have been published in the literature on the molecular transport in the MD process. The transmembrane flux in MD, according to the dusty gas model, can be dominated by the molecular diffusion mechanism, the viscous flow mechanism, or the Knudsen diffusion mechanism.4,25-28 A combination between any of them may also be possible. However, in VMD processes, the mass transfer resistance caused by molecule-molecule collision can be neglected because of the low downstream pressures in the process.11 The Knudsen diffusion mechanism usually dominates the mass transfer through the membrane in VMD. It can be observed that Knudsen type of mass transfer coefficient is mainly a function of membrane morphology and temperature. However, for the fact that membrane pores are not equal (i.e., existence of pore size distribution), the possibility of mass transfer through membrane pores being controlled by more than one mechanism exists. Therefore, the mass flux, NK,P, can be expressed as29 NK,P ) -
(
8 2 εr 3 τδ πRT¯Mw
)
1⁄ 2
∇P-
εr2 P¯ ∇P τδ 8RT¯µ
(3)
where NK,P is the mass flux as function of Knudsen diffusion and viscous flow mechanisms, the ε/τδ accounts for the effective membrane porosity, ε is the membrane porosity, δ is the membrane thickness, τ is the pore tortuosity, ∇P is the pressure j is the average pressure within gradient across the membrane, P j is the average temperature at both the membrane pores, T membrane sides (K), µ is the viscosity of the feed vapor in membrane pores, and r is the pore radius.
With the assumption that all MD processes operate on the principle of vapor-liquid equilibrium (VLE), partial vapor pressure and total pressure can be calculated using VLE data as stated in ref 5. The saturation vapor pressure at the membrane feed side can be estimated by the Antoine equation.5 In fact, the temperature and composition at the vapor-liquid interface differ from the bulk in VMD. The conductive heat transfer across the membrane is negligible because of the low pressure on the permeate side.11,21,26 The heat required for the interfacial evaporation is supplied by the heat flux through the liquid stream, neglecting the heat transfer through the vacuum side, and a simple enthalpy balance is obtained: Nλ ) hf(Tb - TI)
(4)
where hf is the heat transfer coefficient in the liquid phase, and λ is the latent heat of vaporization. Tb is the feed bulk temperature, and TI is the temperature at the liquid-vapor interface. The heat transfer coefficient is given for turbulent liquid flow tube side by the semiempirical correlation30 hfdh
) Nu ) 0.04Re0.75Pr0.33 (5) kT where dh is the effective tube diameter, kT is the thermal conductivity of the liquid, and Nu, Re, and Pr are the Nusselt, Reynolds, and Prandtl numbers, respectively. The temperature at the membrane interface was estimated from the bulk temperature (assumed to be a mean value between inlet and outlet temperatures) using eq 4 for only one vapor permeating. The heat transfer coefficient was estimated using the semiempirical correlation expressed in eq 5. The temperature polarization coefficient was defined as θ)
(TI - Tp) (Tb - Tp)
(6)
where TI, Tp, and Tb are the interfacial, permeate, and bulk feed temperatures, respectively. As θf0 and TpfTI, the resistance in the liquid phase is negligible and the process is controlled by the resistance of the membrane. When θf1 and TIfTb, the resistance in the membrane is negligible and the process is controlled in the liquid phase. Thus, the polarization factor can be used to study the process behavior. The effect of feed concentration on the permeate flux can be obtained according to Raoult’s law PIV ) (1 - x)P*o V
(7)
where PI is the partial pressure of the organic solvent at the interfacial, po* is the saturation pressure of the pure organic
9546 Ind. Eng. Chem. Res., Vol. 47, No. 23, 2008
m (9) A·t where J is the mass of permeate, t is the permeate collection time, and A is the effective area of the membrane. The percent rejection was defined as follows:
Table 1. Membrane Structural Parameters type
PVDF
PAN1
PAN2
average pore size, µm porosity, % thickness, µm contact angle, °
0.3 81 199 92.4
2.0 56 152 55.5
0.5 61 198 61.2
J)
solvent, x is the mole fraction of the solanesol in the solution, and the calculated weight percentage of solanesol is Cmb. Accordingly, the concentration polarization is defined as ς)
Cmb Cb
(8)
where Cb is the solanesol concentration in the feed bulk solution. 3. Materials and Methods 3.1. Materials. The analytical standard of solanesol was obtained from J&K Chemical Ltd. with a purity of >90%; solvents for HPLC mobile phases were of HPLC grade and were obtained from Tianjin Shield Co., China; petroleum ether with boiling point 60∼90 °C was of chemical grade and was obtained from Gaojing Chemical Co., China. The flue-cured tobacco leaves were obtained from Laifeng County, China. Octanol was of analytical grade and was obtained from Shanghai Chemical Co., China. 3.2. Membrane. PVDF and two PAN flat membranes were used. PVDF membrane was laboratory made by the phase inversion method from casting solutions containing 15 wt % of PVDF (Mn ) 850 000, 3F Co., Shanghai, China) and 2 wt % of the nonsolvent additive and the solvent DMAC (>99%, 3S reagent company, Shanghai, China). The PAN membranes were obtained from the Development Center of Water Treatment Technology, Hangzhou, China. The membrane porosity, ε, was defined as the volume of the pores divided by the total volume of the porous membrane. It can usually be determined by gravimetric method. Average pore radius, rf, was investigated by filtration velocity method.31 Contact angle of the membrane surface was measured using a goniometer 14° horizontal beam comparator (22-2000 series model 20-4200, ScherrTumico, St. James, MN). The basic properties of the membranes are listed in Table 1. 3.3. Preparation of Petroleum Ether-Solanesol Solution. Ground tobacco leaf was extracted with petroleum ether for 2 h at 50 °C and then was filtrated; the powder leaf was re-extracted successively with petroleum ether and the extracts were pooled and concentrated to get pasty residue. The residue was saponified and subsequently was loaded onto silica gel column chromatography. The chemical was eluted using a binary solvent mixture of petroleum ether-acetone (98:2, v/v), and the eluent was collected by fraction size of 5 mL. The solanesol concentration of the extracting solution was about 66 mg/L. 3.4. VMD Measurements. In the experiment, the setup assembled is shown in Figure 2. A flat sheet membrane with an effective area of 45 cm2 was installed at the center of the module. The feed was recirculated at a flow rate of 27.5 mL/s. Vacuum in the permeate side was supplied using an oil vacuum pump. Cold traps refrigerated by ice water were used to condense and recover the permeate vapors. Temperature and downstream pressure varied, respectively, in the ranges of 26-46 °C and 10-60 kPa. Feed velocity at the module inlet was 27.5 mL/s corresponding to Re ) 1.2 × 104 at 30 °C. In the experiments, all the experimental conditions, except the specified parameter under investigation, were kept constant. The flux was obtained by the equation
%R )
Cf - Cp × 100% Cf
(10)
where Cp and Cf are the solanesol concentration in the permeate and in the feed, respectively. 3.5. Analysis of Solanesol. The concentration of solanesol in petroleum ether was determined by the Shimadzu HPLC system (Shimadzu SCL-20A, Shimadzu Co., Japan). A Diamonsil TM C18 column (150 mm × 4.6 mm, i.d.) was used at a column temperature of 30 °C, and the injection volume was 5 µL. A binary solvent mixture of acetonitrile-isopropanol (60: 40, v/v) was used as a mobile phase, and measurements were performed at a flow rate of 1.0 mL/min, Detection was done by a Shimadzu SPD ultraviolet detector set at 210 nm. 4. Results and Discussion 4.1. Comparison of the Membrane Performance. To investigate the performance of the VMD process for organic solvent solutions, PVDF, PAN1, and PAN2 membranes were used. The membrane morphologies were characterized with a scanning electron microscope (SEM, FEI SIRION-100), and the results were shown in Figure 3. It was found that all the membranes were porous asymmetric structure, and the surface of PAN1 membrane seemed to be much denser, while the surfaces of the other membranes were more porous. From Figure 4, the porous supporting layer of PAN1 membrane was also different from that of PAN2 membrane for the relatively more bulky and sparse finger pore. In addition, the thickness of the upper dense layer of PAN1 membrane was about 6 µm, and it was only 3 µm for the PAN2 membrane. As illustrated in Figure 1, the liquid-solid contact angle must be greater than 90° for the system to be used in MD to avoid the wettability. The contact angle of PVDF membrane and the organic solution was less than 90°, so from eq 1, the ∆Pentry of the PVDF membrane was under zero. In other words, the ∆Pinterface exceeded ∆Pentry and the organic liquid could penetrate into and through the membrane pores.32 Therefore, although the PVDF membrane with pore size of 0.3 µm, porosity of 81%, and thickness of 199 µm should be proper for MD process from the SEM result, the membrane should be wetted as its hydrophobicity. Otherwise, the contact angles of the PAN1 and PAN2 membranes with water were 55.5° and 61.2°, respectively, which indicated that they were hydrophilic. However, with larger pore size and lower porosity, PAN membranes could be applicable for the organic solvent VMD process.
Figure 2. Scheme of the VMD setup.
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Figure 3. SEM photograph of the surface and cross section of the (a) PVDF membrane, (b) PAN1 membrane, and (c) PAN2 membrane.
Figure 4. SEM photograph of the cross section of the (a) PAN1 membrane and (b) PAN2 membrane.
It can be seen from Table 2 that the PVDF membrane was completely wetted, while the fluxes of PAN membranes were verified to be more than 15 kg/m2 h, which was satisfactory
compared with other separation processes such as distillation. So, PAN membranes were chosen to further verify the feasibility of vacuum hydrophilic membrane distillation.
9548 Ind. Eng. Chem. Res., Vol. 47, No. 23, 2008 Table 2. Membrane Performance membrane 2
J (kg/m h) cp (mg/L) a
PVDF a
66
PAN1
PAN2
15.535 1.82
17.205 1.69
The solanesol extracting solution penetrates the membrane.
Figure 7. Effect of feed temperature on the permeate flux and percent rejection of PAN membranes. VMD operating parameters: P ) 20 kPa; V ) 27.5 mL/s; C0 )66 mg/L.
Figure 5. Effect of downstream pressure on the flux and rejection of PAN membranes. VMD operating parameters: T ) 26 °C; V ) 27.5 mL/s; C0 ) 66 mg/L.
Figure 6. Theoretical permeate flux predicted using eq 3 vs experimental permeate flux for different downstream pressures. VMD operating parameters: T ) 26 °C; V ) 27.5 mL/s; C0 ) 66 mg/L.
4.2. Effect of Vacuum Pressure. In VMD, the driving force was the vapor pressure difference between both sides of the membrane pores so that working at lower downstream pressure would usually result in a higher trans-membrane flux. As it was expected, it can be seen in Figure 5 that the permeate flux increased strikingly as the downstream pressure decreased. However, the rejection of solanesol decreased with increasing downstream pressure. The variations of flux and the solanesol rejection for two PAN membranes with the vacuum pressure were similar, but the rejection of PAN1 membrane was about 3% lower than that of PAN2 membrane. Theoretical fluxes were estimated using eq 3 and were plotted versus the experimental ones in Figure 6 with a relatively good agreement between them. In calculation, the pore size of PAN1 membrane was larger than that of PAN2 membrane, so Knudsen viscous transition should be applicable by eq 3, and the viscous flow mechanism could not be ignored. Contrarily, as mentioned
with the PAN2 membrane, the Knudsen diffusion mechanism dominated the mass transfer through the membrane, and the viscous flow mechanism could be neglected because the mean free path of the diffusing solvent was larger than the pore size of the membrane. It may be observed that the model overestimated the permeate fluxes lightly at lower downstream pressure. In any case, such differences were small (under 8%). One explanation for these discrepancies could be that these experiments did not take into account the mass transfer through the pores of the membrane. Another cause of discrepancies was that the petroleum ether used in our experiments was a mixture of alkyl hydrocarbon such as pentane, hexane, heptane, and so on, and thus, the physical parameters of the petroleum ether were not so exact. Summing up, the downstream pressure had a more considerable effect for the organic solvent systems than aqueous solutions. The flux-pressure curve exhibited an ascending rate behavior with decreasing downstream pressure. However, as illustrated above, with petroleum ether permeate flux increasing, it was more possible for a membrane wetting problem. In the experiments, a downstream pressure of 20 kPa may be considered to be appropriate to obtain a high permeate flux as well as a relatively high solanesol rejection. 4.3. Effect of Temperature. It was usually hypothesized that the flux increased with the increase of temperature through either a reduction in solvent viscosity or an increase in solvent diffusion coefficient in both the liquid phase and the membrane pores. Additionally, the driving force for the permeate increased as the temperature increased because of higher upstream partial pressures. As a result, it can be seen from Figure 7 that the petroleum ether flux increased noticeably with increasing temperature. As Figure 8 shows, the theoretical fluxes that were estimated according to eq 3 had a good agreement with the experimental values, and the analogous results were obtained and compared to Figure 6. The model also overestimated the permeate flux slightly with an error under 5%. The relationship of the permeate flux and feed temperature was analyzed by the Arrhenius equation as follows:33
( )
EJ (11) RT where AJ is a frequency factor, and EJ is the permeation activation energy. The permeate activation energies (EJ) of petroleum ether vapor permeating hydrophilic PAN membranes were obtained J ) AJ exp -
Ind. Eng. Chem. Res., Vol. 47, No. 23, 2008 9549
Figure 8. Theoretical permeate flux predicted using eq 3 vs experimental permeate flux for different feed temperatures. VMD operating parameters: P ) 20 kPa; V ) 27.5 mL/s; C0 ) 66 mg/L.
Figure 10. The temperature polarization coefficient calculated using eq 6 for different feed temperatures. VMD operating parameters: P ) 20 kPa; V ) 27.5 mL/s; C0 ) 66 mg/L.
Figure 9. Arrhenius plot of the effect of temperature on permeate flux.
Figure 11. Effect of initial solanesol concentration on the permeate flux of PAN membranes. VMD operating parameters: T ) 30 °C; V ) 27.5 mL/s; P ) 20 kPa.
from Figure 9 according to the Arrhenius equation, and EJ of PAN1 and PAN2 was 25.38 kJ and 44.16 kJ, respectively. The larger permeate activation energy of PAN2 implied that the permeate flux of PAN2 membrane was more sensitive to temperature than that of PAN1 membrane. The temperature effects were similar to that of hydrophobic MD for aqueous solutions, but the values were a little smaller than the aqueous solution system in MD process.34 To further study the effect of temperature on the membrane permeate flux, the temperature polarization coefficient θ was calculated according to eq 6 shown in Figure 10. It was found that the θ varied approximately from 0.91 to 0.87 for PAN2 membrane and from 0.89 to 0.73 for PAN1 membrane with increasing the temperature from 26 to 46 °C. This indicated that the temperature polarization effects of PAN1 membrane were more significant than that of PAN2 membrane, and the higher the temperature, the more distinct the discrepancy was. As a result, the flux of PAN2 membrane was much higher than that of PAN1 membrane for the same feed temperature. The flux of PAN2 membrane was 53 kg/m2 h at 46 °C, while the flux of PAN1 membrane only was above 20 kg/m2 h. It may be due to the different structures of the two PAN membranes as illustrated in Figures 3 and 4. A survey of the literature revealed that the membrane morphology had a significant impact
on the MD performance,35-37 which resulted most probably from the differences between convection and diffusion of the mass transport. Owing to the high flux at high temperature, the solanesol rejection slightly declined, but at temperatures below 35 °C, more than 98% solanesol percent rejection was achieved. Although the petroleum ether permeate flux increased exponentially with increasing temperature as illustrated in Figure 7, obviously, exceptionally high feed temperature might result in adverse effects of rejection. So, 30 °C could be considered to be optimal in terms of both the flux and the rejection. 4.4. Effect of Feed Concentration. Concentration polarization was generally existent in MD, which would lead to the decline of flux and rejection. It would also greatly reduce the value of γL|cos θ|, which could cause liquid penetration of the membrane. On the other hand, the membrane wettability from pore wetting and capillary condensation would also result in a decrease of the flux. Usually, the application of membrane with the small surface pore could eliminate the internal fouling in the pressure-driven processes.38 Therefore, the presence of large pores on the PAN1 membrane made the concentration polarization and membrane wettability form more easily than the PAN2 membrane when concentration increased. The results are shown in Figure 11, and as anticipated, the flux of PAN1 membrane decreased
9550 Ind. Eng. Chem. Res., Vol. 47, No. 23, 2008
Figure 12. The concentration polarization coefficient for different feed concentration. VMD operating parameters: T ) 30 °C; V ) 27.5 mL/s; P ) 20 kPa.
dramatically; it declined to below 3 kg/m2 h when the solanesol concentration changed to 600 mg/L, and the flux was only 12.5% of that of the pure petroleum ether. However, the flux of the PAN2 membrane diminished slightly; it varied to about 14 kg/ m2 h when solanesol concentration increased to 1000 mg/L, and the flux was 70% of that of the pure petroleum ether. In addition, the solanesol concentration in the permeate for the PAN2 membrane was lower than that of the PAN1 membrane, and the rejection for PAN2 membrane was above 98% at all experimental conditions. Considering the correlation between vapor pressure and mole fraction of solanesol in the solution, the solanesol concentration at the membrane surface was calculated, the concentration polarization coefficient ς was obtained from eq 8, and the results were shown in Figure 12. It can be seen that the concentration at the membrane surface was up to 1-7 times that of the bulk concentration for the PAN1 membrane and 1-3 times that of the bulk concentration for the PAN2 membrane. The PAN1 membrane was more evidently affected by the concentration polarization for its bulky and sparse fingerlike pore. 4.5. VMD Experiments with PAN2. From the above experiments, both relatively high permeate flux and high solanesol rejection were obtained by the PAN2 membrane, so it was verified to be more suitable for the vacuum hydrophilic membrane distillation. At a downstream pressure of 20 kPa, a feed concentration of 66 mg/L, and a feed flow rate of 27.5 mL/s, the durability performances of PAN2 membrane were measured. As illustrated in Figure 13, the steady state was observed after 50 min for the membrane, so the fluxes and rejections of the MD process should be obtained after achieving steady state. From Figure 13, it can be seen that the permeate flux of PAN2 membrane was about 16.5 kg/m2 h and that the rejection stayed above 98% during the entire experiment. When the feed temperature was 26 °C, the flux remained at 14 kg/m2 h, which is a little lower than that of 30 °C. To sum up, the PAN2 membrane gave better results and offered safe operating conditions without pore wetting. Its operating temperature was low enough to be obtained from waste energies in industry plants. 5. Conclusions Two hydrophilic PAN membranes with different asymmetric structures had been validated to be efficient for the recovery of
Figure 13. Permeate flux and percent rejection of the PAN2 membrane as a function of time. VMD operating parameters: V ) 27.5 mL/s; P ) 20 kPa; C0 ) 66 mg/L.
petroleum ether from solanesol solution by VMD, but to the contrary, the hydrophobic PVDF membrane was wetted by organic solvent. The results showed that good solvent flux (>15 kg/m2 h) and solute rejection (>98%) of the PAN membranes were obtained. The performances of two PAN membranes were compared by investigating the membrane structure and operating parameters including vacuum pressure, feed temperature, and initial feed concentration of solanesol; it was been found that the PAN2 membrane with smaller pore size was more suitable for the process. The VMD process with hydrophilic membrane was determined to be sensitive to both the vacuum pressure and the feed temperature, but it was more sensitive to the feed temperature at high vacuum pressure levels. General models taking into account Knudsen and viscous flows were proposed, and a good agreement between the experimental and theoretical fluxes was also obtained. Additionally, temperature and concentration polarizations were proved to have an obvious influence on the VMD process of PAN membranes. Finally, the optimal operation conditions of downstream pressure of 20 kPa, feed temperature of 30 °C, and feed flow rate of 27.5 mL/s were advised for PAN2 membrane. Acknowledgment The authors acknowledge financial support for this work from the National Basic Research Program of China (2003CB615706) and the National Natural Science Foundation of China (no. 20506021). Literature Cited (1) Ferna´ndez-Pineda, C.; Izquierdo-Gil, M. A.; Garcıa-Payo, M. C. Gas permeation and direct contact membrane distillation experiments and their analysis using different models. J. Membr. Sci. 2002, 198, 33. (2) Izquierdo-Gil, M. A.; Garcıa-Payo, M. C.; Ferna´ndez-Pineda, C. Air gap membrane distillation of sucrose aqueous solutions. J. Membr. Sci. 1999, 155, 291. (3) Khayet, M.; Godino, P.; Mengual, J. I. Nature of flow on sweeping gas membrane distillation. J. Membr. Sci. 2000, 170, 243. (4) Bandini, S.; Gostoli, C.; Sarti, G. C. Separation efficiency in vacuum membrane distillation. J. Membr. Sci. 1992, 73, 217. (5) Kevin, W. L.; Douglas, R. L. Membrane distillation. J. Membr. Sci. 1997, 124, 1. (6) Karakulski, K.; Gryta, M. Water demineralisation by NF/MD integrated processes. Desalination 2005, 177, 109. (7) Gryta, M. Osmotic MD and other membrane distillation variants. J. Membr. Sci. 2005, 246, 145.
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ReceiVed for reView November 4, 2007 ReVised manuscript receiVed August 12, 2008 Accepted August 28, 2008 IE071499+
