polyetherimide (PEI) membranes

Journal of Membrane Science 378 (2011) 339–350

Contents lists available at ScienceDirect

Journal of Membrane Science journal homepage: www.elsevier.com/locate/memsci

Processing and engineering of pervaporation dehydration of ethylene glycol via dual-layer polybenzimidazole (PBI)/polyetherimide (PEI) membranes Yan Wang a , Tai Shung Chung a,∗ , Bernard Weijie Neo a , Michael Gruender b a b

Department of Chemical and Biomolecular Engineering, National University of Singapore, 4 Engineering Drive 4, Singapore 117576, Singapore PBI Performance Products, Inc., 9800-D Southern Pine Boulevard, Charlotte, NC 28273, USA

a r t i c l e

i n f o

Article history: Received 12 January 2011 Received in revised form 4 April 2011 Accepted 11 May 2011 Available online 17 May 2011 Keywords: Pervaporation dehydration Polybenzimidazole Ethylene glycol Operation conditions Operation temperature Permeate pressure Feed composition

a b s t r a c t Operating conditions play a significant role in determining the separation performance of a pervaporation process, because they not only manipulate the driving forces to transport permeants but also affect the physicochemical properties of the pervaporation membrane itself. In this study, fundamental governing equations have been derived to correlate separation performance with system operation conditions and intrinsic separation characteristics of the pervaporation membrane. Polybenzimidazole/polyetherimide (PBI/PEI) dual-layer hollow fiber membranes were chosen to study the pervaporation dehydration of ethylene glycol (EG) under different testing protocols. The effects of operational parameters such as operation temperature, permeate pressure, feed composition and operation duration on performance indicators (flux and separation factor, permeance and selectivity) have been investigated. Experimental results show that an increase in operation temperature results in an increase in flux and selectivity, but a decrease in permeance and separation factor. In addition to other factors, decreasing sorption, less EG–water clusters and lower membrane-EG affinity with increasing temperature, play essential roles for the opposite trends. Both flux and permeance decrease with an increase in permeate pressure, while both separation factor and selectivity have an up-and-down trend. An increase in EG composition in the feed from 50 to 90 wt.% results in a lower water flux and permeance, but EG flux and permeance first increase and then decrease. This is due to the combined effect of water-induced membrane swelling and the formation of an EG boundary layer upon the membrane surface. The long-term test up to 33 days proves the membrane durability for EG dehydration. This work may provide useful insights to pervaporation fundamentals, system design and scale up for the EG dehydration. © 2011 Elsevier B.V. All rights reserved.

1. Introduction Ethylene glycol (EG) is a widely used chemical for nonvolatile antifreeze and coolant as well as an intermediate in the manufacture of polyesters. Dewatering of EG is a critical step in its applications and production processes. Currently, EG dehydration is carried out by conventional multi-stage evaporation units followed by distillation columns, characterized by high capital costs. Alternatively pervaporation is an effective technology for the EG/water separation for being highly selective, economical, energy efficient and environmentally benign. Pervaporation has been considered as one of the most promising technologies in the molecular-scale liquid/liquid separations existing in biorefinery, petrochemical, pharmaceutical industries, etc. It has been used to remove trace substances in liquid mixtures, such as removing water from organic solvents to produce

∗ Corresponding author. Tel.: +65 6516 6645; fax: +65 6779 1936. E-mail address: [email protected] (T.S. Chung). 0376-7388/$ – see front matter © 2011 Elsevier B.V. All rights reserved. doi:10.1016/j.memsci.2011.05.020

high-purity solvents. Significant attention has been given to highly hydrophilic polymers, such as poly (acrylic acid) (PAA), poly (vinyl alcohol) (PVA), polyacrylonitrile (PAN), sodium alginate, and chitosan as membrane materials for pervaporation dehydration in the earlier stage. However, these materials lack mechanical strengths and stability in aqueous solutions due to the excessive swelling, leading to a drastic decrease in separation performance. A common method to suppress the swelling is to crosslink the materials while the permeation flux is often compromised. As a result, cross-linking modifications for the development of pervaporation membranes may not be desirable because additional post-treatments incur extra costs and prolong production durations. Another method is the use of composite membranes consisting of a thin active layer upon a microporous substrate [1–11]. The outer surface layer provides the selectivity, while the porous substrate layer offers mechanical strengths and high permeability. Composite membranes with much improved separation performance have been reported extensively in recent years for pervaporation dehydration [1–9]. Co-extruded dual-layer hollow fibers are one example where synergistic separation performance

340

Y. Wang et al. / Journal of Membrane Science 378 (2011) 339–350

can be achieved without the aid of intensive thermal or chemical treatments if inner- and outer-layer materials are properly selected [1–6]. In our previous paper [2], three types of membranes based on polybenzimidazole (PBI) material had been studied for pervaporation dehydration of EG; namely, dense flat-sheet PBI membranes, PBI single-layer and PBI/PEI dual-layer hollow fiber membranes. The PBI/PEI dual-layer hollow fiber membranes show the best separation performance due to the unique combination of physicochemical properties of the PBI outer selective layer and the less swelling characteristics of the PEI supporting layer, as well as the synergistic effect of molecularly designed membrane morphology via dual-layer co-extrusion. The developed PBI/PEI dual-layer hollow fiber membrane has a separation factor up to 4500 and a flux up to 186 g/m2 h, which are better than most other polymeric membranes [2]. As a continued work for EG dehydration, we focus on the engineering and process aspects of separation performance in this study using PBI/PEI dual-layer hollow fiber membranes. According to the solution diffusion model, the transport of each component across a pervaporation membrane is not only dependent on the characteristics of pervaporation membranes but also on the operation conditions [12–19] such as feed composition, process temperature, permeate pressure and other factors, as well as their mutual interactions. For example, if the operation conditions change partial vapor pressures of the components at the feed side or the permeate side, the driving forces for these components across the membrane are also changed. At the same time, membrane morphology and physicochemical properties may vary with operation conditions, causing the changes of intrinsic membrane permeability and selectivity. As a result, both permeation flux and separation factor are determined by many factors. It is therefore necessary to study the effects of operation parameters in order to properly optimize or scale up a pervaporation process. A number of researchers have studied the effects of process conditions on pervaporation systems [12–19], but no deep exploration has been established theoretically and systematically. The objectives of this study are to study the PBI/PEI dual-layer hollow fiber membrane for EG dehydration and to investigate its separation performance as a function of operation conditions. Not only are the conventional performance indicators (flux and separation factor) studied and analyzed in this work, but also permeance and selectivity. By decoupling external process parameters, we obtain a more true characterization of a membrane’s intrinsic performance. We accomplish this by studying the permeance and selectivity since they reflect the real change of membrane properties under varied operating conditions. The influence of key process parameters such as temperature, permeate pressure, feed composition and testing duration on the separation performance of PBI/PEI duallayer hollow fiber membranes will be presented and elucidated using the solution-diffusion model. This work may (1) offer useful criteria for the selection of appropriate operation conditions for EG pervaporation systems to ensure the highest performance and economic viability; and (2) provide insights on the transport mechanism of the pervaporation process through the influence of the external operation conditions. To benchmark our membrane performance against other membranes under similar operating conditions, Table 1 summarizes some recent works for EG dehydration using feed compositions of 70–90 wt.% EG [20–27].

2. Performance evaluation of pervaporation membranes The performance evaluation of a pervaporation membrane is generally based on its capability to separate components from each other. There are two sets of interlinked parameters which have

been widely used to describe this capability, namely: (1) flux and separation factor and (2) permeance and selectivity. Traditionally, the performance of a pervaporation membrane is characterized by flux (J) and separation factor (˛) as defined by the following equations: J=

Q A·t

(1)

yw,1 /yw,2 xw,1 /xw,2

˛1/2 =

(2)

where Q is the total mass transferred over the operation time t, A is the membrane area, subscripts 1 and 2 refer to the two components to be separated in the feed mixture, yw and xw are the weight fractions of the components in the permeate and feed, respectively. For flat-sheet dense membranes, the flux is also generally expressed in normalized flux (JN ), which is defined as the total flux J multiplied by the thickness of membrane selective layer (l). JN = J · l

(3)

Because of the existence of a trade-off relationship between flux and separation factor, that is, the flux and separation factor usually perform in the opposite way, the pervaporation separation index (PSI) [28,29] was therefore proposed to evaluate the overall performance of a membrane, which is defined as follows. PSI = J · (˛ − 1)

(4)

As for the evaluation of intrinsic properties of a specific permeant-membrane system, permeability (or permeance) and selectivity are more representative and accurate since they significantly decouple the effect of process parameters on performance evaluation. Wijmans [30] has elaborated on the importance and differences in using permeability instead of flux to investigate intrinsic membrane properties. The detailed examples of using permeability and selectivity in calculations and applications have been elucidated by our group [1,31–34]. The relationship between permeability or permeance and flux can be expressed as below: Ji =

P  l

i

(xn,i i psat − yn,i pp ) i

(5)

where Pi is the membrane permeability of the component i, a product of diffusivity and solubility coefficients, l is the membrane thickness, xn,i and yn,i are the mole fractions of the component i in the feed and permeate,  i is the activity coefficient, psat is the i saturated vapor pressure, and pp is the permeate pressure. psat and i  i can be calculated by the Wilson equation and Antoine equation, respectively, and obtained with the aid of the AspenTech DISTIL software provided by Hyprotech Ltd., Canada [31]. The term [Pi /l] is also known as permeance (P¯ i ) that is often employed for an anisotropic membrane with an unknown thickness of the dense selective layer. It can be expressed by rearranging Eq. (5) as follows: P¯ i =

Pi Ji = l xn,i i psat − yn,i pp i

(6)

The partial vapor pressure of each component at the feed side can be expressed in terms of fugacity (fi ) as shown below. fi = xn,i i psat i

(7)

The driving force for component i to transport through the membrane is the difference of its partial vapor pressures at the feed side and permeate side and can be written as follows. Driving force = xn,i i psat − yn,i pp i

(8)

The total permeance is therefore defined as the sum total of the permeances of all individual components.

Y. Wang et al. / Journal of Membrane Science 378 (2011) 339–350

341

Table 1 A comparison of pervaporation performance of polymeric membranes for ethylene glycol dehydration. Mass ratio (EG wt.%)

Membrane

Separation factor

Flux (g/m2 h)

Temperature (◦ C)

References

82.5 70 80 80 80 80 80 90 80

Crosslinked PVA–PES composite membrane Commercial NaA zeolite membranes PVA-GPTMS/TEOS hybrid pervaporation membrane Surfaced crosslinked PVA with glutaraldehyde as crosslinking agent Mordenite-filled chitosan–PAA polyelectrolyte complex membranes (PECM) 60 wt.% CS and 40 wt.% PAA polyelectrolyte complex membranes (PECM) 50 wt.% PVA and 50 wt.% ␥-MPTMS polymer inorganic nanocomposite membranes PVA and PS crosslinked with trimesoyl chloride (TMC) PBI/PEI dual-layer hollow fiber membrane

231 1177 714 933 258 105 311 987 1763

383 940 60 211 165 216 67 360 115

80 70 70 70 70 70 70 60 60

[20] [21] [22] [23] [24] [25] [26] [27] This work

The ideal membrane selectivity ˇ is defined as the ratio of permeability coefficients or permeance of the two components. ˇ1/2 =

P1 P¯ 1 or P2 P¯ 2

(9)

The ideal selectivity ˇ can be either weight-based (ˇw, 1/2 ) or mole-based selectivity (ˇn, 1/2 ), depending on the unit of permeability coefficients or permeance. Their relationship is as follows. ˇn,1/2 = ˇw,1/2 ×

M2 M1

(10)

where M1 and M2 are the molecular weights of the two components, respectively. Based on their definitions, the discrepancy of separation factor and selectivity arises fundamentally from the fact that the latter has decoupled the activity coefficient and saturated vapor pressure. The following derivation shows their relationships under different operation conditions. According to Eqs. (6) and (8), the selectivity can be expressed as below: ˇ1/2 =

J1 /(xn,1 1 psat − yn,1 pp ) 1

(11)

J2 /(xn,2 2 psat − yn,2 pp ) 2

Flux (J) is generally weight-based and can be written as follows, Ji = J · yw,i

(12)

Then the calculated selectivity ˇ is also weight-based and can be rewritten as: yw,1 /(xn,1 1 psat − yn,1 pp ) 1

ˇw,1/2 =

yw,2 /(xn,2 2 psat − yn,2 pp ) 2

(13)

In the above equation, the relationships of mole fraction and weight fraction of component i in the feed or permeate are as following: xn,i =

xw,i /Mi



(xw,i /Mi )

,

yn,i =

yw,i /Mi



(yw,i /Mi )

ˇw,1/2 = =

yw,2 · xn,1 1 psat 1

=

yw,1 · ((xw,2 /M2 )/ yw,2 · ((xw,1 /M1 )/

yw,1 /xw,1 M1 2 psat 2 · · yw,2 /xw,2 M2 1 psat 1

 (x /M )) · 2 psat 2  w,i i sat (xw,i /Mi )) · 1 p1

(15)

So the relationship between the weight-based selectivity and separation factor of a pervaporation process under low permeate pressures can be expressed as follows. ˇ =˛·

M1 psat 2 · 2 · M2 psat 1 1

M2 psat 1 · 1 · M1 psat 2 2

(17)

Since the molecular weights of the two components are constants, the saturated pressures are only affected by the operation temperature, the activity coefficients of the components are affected by the operational temperature and feed composition, Eq. (17) indicates that, for a fixed feed system (i.e., operational temperature and feed composition are both constant), the resultant separation factor of a pervaporation membrane is only determined by the membrane intrinsic selectivity, if the downstream permeate pressure is negligible. If the permeate pressure is significant, Eq. (13) can be expressed as Eq. (18) by taking downstream pressure into consideration with the aid of Eqs. (7) and (14). The detailed derivations are given in Appendix A.

ˇw,1/2 ·

f1 yw,1

−

f2 yw,2

 =

ˇw,1/2 M1

1 − M2

 ·

pp



(18)

(yw,i /Mi )

For a binary system with a high separation performance (yw,1 ≈ 1 and yw,2 ≈ 0), Eq. (18) can be further simplified as follows: ˇw,1/2 · f1 −

f2 yw,2

 =

ˇw,1/2 M1

−

1 M2

 · M1 · pp

(19)

Therefore, for a binary feed system with constant feed composition and temperature, the permeate composition of component 2 can be predicted as follows, if membrane selectivity and downstream permeate pressure are known. yw,2 =

f2 ˇw,1/2 · f1 − ((ˇw,1/2 /M1 ) − (1/M2 )) · M1 · pp

(20)

(14)

Since the permeate pressure is kept very low and can be considered negligible in most lab-scale pervaporation tests, Eq. (13) can be simplified as below with the aid of Eq. (14): yw,1 · xn,2 2 psat 2

or ˛ = ˇ ·

(16)

Substituting Eq. (20) into Eq. (2) (i.e., the definition of the separation factor), we can derive separation factor as a function of selectivity and downstream pressure as below: ˛1/2 =





xw,1 1 M1 · · ˇw,1/2 · f1 − ˇw,1/2 − xw,2 f2 M2



· pp − f2

 (21)

Here xw,1 , xw,2 , f1 and f2 are all constants for a fixed binary feed system. Therefore, if the membrane intrinsic selectivity ˇ remains unchanged with testing conditions, separation factor will decrease when the permeate pressure pp increases. However, since membrane morphology and its intrinsic selectivity ˇ may vary with permeate pressure, the resultant separation factor is determined by complicated combined effects of permeate pressure, membrane intrinsic selectivity and its relations with penetrants and permeate pressure.

342

Y. Wang et al. / Journal of Membrane Science 378 (2011) 339–350 Table 2 The spinning parameters of PBI/PEI dual-layer hollow fiber membranes.

Fig. 1. The chemical structures of (a) polybenzimidazole (PBI) and (b) Ultem® 1010 polyetherimide (PEI).

3. Experimental 3.1. Materials The PBI polymer solution was provided by PBI Performance Products, Inc. with the composition of PBI 26.2 wt.%, n,ndimethylacetimide (DMAc) 72.3 wt.%, and lithium chloride (LiCl) 1.5 wt.%. The LiCl serves the function of preventing PBI polymer from phasing out of the solution. Ultem® 1010 polyetherimide were supplied by GE plastics. Fig. 1 shows the chemical structures of PBI and PEI [2]. Polyvinylpyrrolidone (PVP) (Merck, Singapore) with an average Mw of 30 kDa as an additive was employed for hollow fiber membrane spinning. Polymers were dried overnight at 120 ◦ C under vacuum before use. DMAc, employed as the solvent to fabricate hollow fiber membranes, was supplied by Merck with analytical grade and used as received. EG of analytical grade was used to mix with deionized water to prepare binary aqueous solutions as the feed with various compositions. 3.2. Spinning process and modules fabrication The schematic diagram of the dual-layer hollow fiber spinning system has been described elsewhere [4]. The outer-layer dope solution is a PBI® 23 wt.% solution, diluted from the original supplied PBI solution, with a composition of PBI/DMAc/LiCl 23/75.67/1.33 wt.%; while the inner-layer dope solution is a 25.0 wt.% PEI (Ultem 1010) polymer solution, with a composition of PEI/PVP/DMAc 25.0/5.0/70.0 wt.%. The polymer solution was degassed for 24 h before loading into a syringe pump (ISCO 1000). A mixture of 85/15 (w/w) DMAc/water was employed as the bore fluid in order to make a porous inner surface. The dual-layer PBI/PEI hollow fibers were spun by co-extruding the polymer solutions and bore fluid out of the spinneret orifice and subsequent phase inversion in a coagulant bath with a pre-set air gap. Both dope fluid and bore fluid were all filtered through 15 ␮m sintered metal filters before spinning. Tap water was used as the external coagulant at room temperature. The nascent fibers were rolled up by a drum with a free-fall take-up speed, cut into segments, and then rinsed in a clean tapping water bath for at least 3 days to remove the remaining DMAc with water changed daily. 2 batches of hollow fibers were used in this study to confirm the reproducibility. Fibers are abbreviated as PBI-D-A and PBI-D-B with different air gap distances of 2 cm and 5 cm, respectively. Table 2 gives the detailed spinning parameters for the dual-layer co-extrusion process. The as-spun hollow fibers were dried in air naturally after freeze-drying, and then stored in ambient environment. No delamination between PBI outer layer and PEI inner-layer of the resultant hollow fibers was found under microscopy observation. Fig. 2 shows the SEM morphology of the PBI/PEI dual-layer hollow fiber PBI-D-B. The hollow fiber has an outer diameter of about 1250 ␮m, an overall wall thickness of about 240 ␮m, and an outer layer thickness of about 16 ␮m.

Parameters

Range of variables

Outer-layer dope solution composition Inner-layer dope solution composition Bore fluid composition (wt.%) Dimensions of spinneret (mm) External coagulant Temperature Outer-layer dope flow rate (ml/min) Inner-layer dope flow rate (ml/min) Bore fluid flow rate (ml/min) Take-up speed (m/min) Air gap distance (cm) Membrane ID

PBI:DMAc:LiCl (23:75.67:1.33 wt.%) PEI (Ultem 1010):PVP:DMAc (25:5:70 wt.%) DMAc:water (85:15 wt.%) OD1/OD2/ID (1.20/0.97/0.44) Water Ambient (23 ± 2◦ C) 0.5 4 2 4.60 (free fall) 5 PBI-D-A

2 PBI-D-B

More details on the as-fabricated hollow fiber membranes can be found in our previous published work [2]. 3.3. Pervaporation study A laboratory scale pervaporation unit was employed and the details of the apparatus have been described elsewhere [34]. A feed solution of EG/water with a fixed composition was used for pervaporation tests. The feed composition varied less than 0.5 wt.% during the entire experiment and therefore was considered constant because of the large quantity of the feed solution compared to the permeate sample. The feed flow rate was maintained at 0.5 l/min for each hollow fiber module, respectively. Retentate and permeate samples were collected after the membrane being conditioned for about 2 h. The mass of permeate was weighed using a Mettler Toledo balance. The weight fractions of components in the permeate and feed were analyzed through a Hewlett-Packard GC 7890 A with a HP-INNOWAX column (packed with crosslinked polyethylene glycol) and a TCD detector. The pervaporation module was prepared by loading 1 piece of hollow fiber membranes into a perfluoroalkoxy tubing connecting with two Swagelok® stainless steel male run tees with an effective length of around 20 cm. Both ends were sealed by epoxy and cured for more than 24 h at ambient temperature. At least two pervaporation modules were tested for each membrane sample. Unless specified, a binary mixture containing 80/20 wt.% EG/water was chosen as the feed for the study of the effects of operation conditions on pervaporation performance as follows: (1) The effect of operational temperature under a constant permeate pressure of 2 mbar; (2) the effect of permeate pressure under a constant temperature of 60 ◦ C; (3) the stability of long-term performance for 33 days under a constant temperature of 60 ◦ C and permeate pressure of 2 mbar. In addition, we varied feed composition and studied its effects under a constant temperature of 60 ◦ C and permeate pressure of 2 mbar. Both membranes PBI-D-A and PBI-D-B were investigated. Since the obtained results are similar, only the results of the membrane PBI-D-B are presented and discussed in most cases, unless otherwise stated. 3.4. Pervaporation tests under higher permeate pressures The permeate pressure was varied from 2 mbar to 75 mbar to study its effect on separation performance. It was observed that the permeate vapor started to condense along the tubing instead of in the cold trap when permeate pressure was increased to about

Y. Wang et al. / Journal of Membrane Science 378 (2011) 339–350

343

Fig. 2. FESEM images of the membrane morphology of PBI/PEI dual-layer hollow fibers (PBI-D-B).

Table 3 Dew points of water and EG under different pressures. Pressure (mbar)

2 10 15 25 35 50

Dew point (◦ C) Water

EG

−14.53 6.99 13.03 21.08 26.67 32.88

58.57 85.35 92.86 102.8 109.7 117.3

analyzed daily to make sure a constant feed composition. In case of a discrepancy larger than 0.5 wt.% (with the original 80 wt.% EG concentration) was detected, the circulation pump was stopped for a minute to allow the addition of a calculated amount of water (or EG) into the feed tank to adjust the composition back. The samples were collected every 12–24 h generally. The feed circulation was off during the weekend while the permeate vacuum pump was on all the time to avoid severe membrane swelling from the liquid feed.

4. Results and discussion 50 mbar. This was due to the increased dew points of the permeants (water and EG) under the increased vacuum pressure (i.e., lower the absolute vacuum power), as listed in Table 3, which was verified with the aid of Aspen HYSYS. A water/EG mixture with 99.95 wt.% water was found to have a dew point of −14.44 ◦ C under a vacuum pressure of 2 mbar; the dew point rapidly increased to 21.14 ◦ C under 25 mbar; then 32.93 ◦ C under 50 mbar, which was higher than the room temperature and caused the condensation of the permeate in the tubing. Therefore, the condensed permeate was collected for analyses by applying a full vacuum (about 2 mbar) for a very short time (about 2 min) to suck all condensed permeate from the tube to the cold trap. 3.5. The long term stability study of the dual-layer hollow fibers The fabricated membrane module for long-term performance studies is shown in the schematic diagram as in Fig. 3. Small modules with 1 piece of hollow fiber and a much shorter length were prepared. A short sample length was chosen to minimize rapid variation of the feed composition as well as for easy sample collection over a longer period of testing. The feed composition was

Fig. 3. Module fabrication for long term test of PBI/PEI dual-layer hollow fiber membranes.

4.1. Effect of operation temperature The physicochemical properties of polymeric membranes such as membrane morphology, free volume and its distribution vary significantly with operating temperature. In addition, the operating temperature modifies the mutual interactions between permeates, and among permeates and membrane materials. As a result, these effects alter mass transport coefficients of permeants across the membrane. Fig. 4 shows the separation performance of the PBI/PEI dual-layer hollow fiber membrane as a function of operation temperature. The overall flux increases while separation factor decreases with an increase in operation temperature. In addition, both water and EG fluxes increase with increasing temperature, exhibiting a similar trend with total flux. Fig. 5 also plots the separation performance in terms of total permeance and selectivity. Interestingly, the trends of permeance and selectivity vs. temperature are totally opposite with those of flux and separation factor. Based on the solution diffusion model, permeability is a product of diffusivity and solubility. As temperature increases, the diffusion processes becomes faster while the equilibrium solubility decreases. The decreasing permeance trends shown in Fig. 5 indicate that the reduction in solubility is greater than the increment in diffusivity for both water and EG molecules when pervaporation temperature is increased [33]. Similar phenomena have also been reported by previous researchers [33,35–37]. The increased flux of individual components with temperature is thus primarily due to the increase in their driving forces. Since the saturated vapor pressures of permeating components increase with increasing pervaporation temperature, the net driving forces and fluxes across the membrane as written in Eq. (8) increase

344

Y. Wang et al. / Journal of Membrane Science 378 (2011) 339–350 250

250

2000

a

Total Flux Separation Factor

b

Water flux EG Flux

0.7 0.6

150 1800 100 1700

Water flux (g/m 2 ·hr)

Separation Factor

Flux (g/m2·hr)

1900

0.5

150

0.4 0.3

100

EG flux (g/m2 ·hr)

200

200

0.2

50

50

0.1

0 50

60

70

80

0.0

0

1600 40

90

40

50

60

70

80

Temperature (ºC)

Temperature (ºC)

Fig. 4. Effect of operation temperature on separation performance of the PBI/PEI dual-layer hollow fiber membrane.

because the downstream permeate pressure remains the same. Other factors have often been reported to explain the enhanced permeation flux with increasing operation temperature, namely: (1) the increase in thermal motion of polymer chains and free volume inside the polymeric membrane and (2) the enhanced mass transfer coefficients of the permeating components. However, since the permeances of both components still exhibit decreasing trends with increasing operation temperature as shown in Fig. 5 for the current study, these two factors are considered minor compared to the negative impact of the reduced sorption and the positive impact of increased saturated vapor pressure with increasing temperature. Fig. 5(a) shows that selectivity increases with an increase in operation temperature, which is in line with the results in Fig. 5(b), where the percentage of EG permeance decrease with increasing temperature is larger than that of water permeance. The increased selectivity can be attributed to the less EG–water clusters because of the reduced viscosity [38,39] of the feed solution at elevated temperatures and lower membrane affinity towards EG. Fig. 5(b) also shows the calculated driving forces (as defined in Eq. (8)) for EG

1100

a

Total Permeance Selectivity

and water transports. Although EG has a smaller driving force than water, the former has a greater percentage increase with temperature than the later. As a result, the separation factor of water to EG decreases with an increase in operation temperature because of decreasing driving force ratio as described in Eq. (17). Figs. 6 and 7 plot flux and permeance vs. operation temperature, respectively, using Arrhenius equations as follows [40]:



J = J0 exp

EJ − RT



(22)

 E  P

P¯ = P0 exp −

(23)

RT

where J0 and P0 are the pre-exponential factors of flux and permeance, respectively, R is the universal gas constant, T is the operating temperature, and EJ and EP are the apparent activation energies of flux and permeance, respectively. EJ and EP can be calculated from these figures using the least square method and Table 4 summarizes the calculated values of EJ,w , EJ,EG , EP,w , and EP,EG . Good linearity

0.4

900

7

b

0.3

6

5

800

Water Permeance EG Permeance

600

Driving force ratio (Water/EG) Water driving force EG driving force

0.2

450

300 0.1

4 150

700 40

50

60

70

Temperature (ºC)

80

3 90

0

0 40

50

60

70

Temperature (ºC)

Fig. 5. Effect of the operation temperature on permeance and selectivity.

80

Driving force (Bar)

900

Permeance (g/m 2·bar·hr)

1000

Selectivity (Water/EG)

Total Permeance (g/m2·bar·hr)

750

Y. Wang et al. / Journal of Membrane Science 378 (2011) 339–350

140

Total Flux Separation Factor

120

Flux (g/m ·hr)

In(J ) = -4.5953x(1000/T) + 18.487 R² = 0.9979

-0.5

6000

100

5000

80

4000

60

3000

40

2000

20

1000

2

5

7000

-1

4.5

EG In(J ) = -5.0116x(1000/T) + 13.623 R² = 0.9984

4

In(J EG)

In(J w)

Water

-1.5

0 0 3.5 2.8

2.85

2.9

2.95

3

3.05

10

20

30

40

50

60

70

Separation Factor

0

5.5

345

0 80

Permeate Pressure (mbar)

-2 3.1

Fig. 8. Effect of permeate pressure on the total flux and separation factor.

1000/T (1/K) Fig. 6. Arrhenius plots of water flux and EG flux against reciprocal temperature.

7

brane PBI-D-A. This is consistent with the lower permeation flux of membrane PBI-D-A than membrane PBI-D-B for EG dehydration as reported in our previous paper [2].

ln (P ) = 0.5297x (1000/T)+ 5.0422 R = 0.9582

4.2. Effect of permeate pressure

In(Permeance)

6.5

Water

6

5.5 ln (P

) = 1.9677x (1000/T)- 0.8179 R = 0.9794

5

EG

4.5 2.8

2.85

2.9

2.95

3

3.05

3.1

3.15

1000/T (1/K) Fig. 7. Arrhenius plots of water permeance and EG permeance against reciprocal temperature.

exists between logarithmic flux and permeance vs. reciprocal temperature and the values of regression coefficients R2 are close to unity, indicating the experimental data fits the Arrhenius equation well. Compared to EG, water has a smaller kinetic diameter and higher driving force (i.e., higher vapor pressure) to transport through the membrane. Therefore, the former has a higher activation energy of flux than the latter (i.e., EJ,EG > EJ,w ). Interestingly, the activation energies of permeance for both water and EG are negative. These results reconfirm our previous hypothesis that the decrease in sorption with increasing temperature overshadows the increase in thermal motion of polymer chains and the enhanced mass transfer coefficients. In addition, the EJ and EP values are dependent on membrane morphology developed via phase inversion during the spinning process. The EJ and EP values of hollow fiber PBI-D-A are higher than those of PBI-D-B, indicating the former has higher energy barriers than the latter for both water and EG molecules to transport across the membrane. The higher resistance may be resulted from the closer molecular packing and denser morphology because a larger air gap distance was employed during the spinning of memTable 4 Apparent activation energies of flux and permeance for membranes PBI-D-A and PBI-D-B. Membrane

EJ,w (kJ/mol)

EJ,EG (kJ/mol)

EP,w (kJ/mol)

EP,EG (kJ/mol)

PBI-D-A PBI-D-B

40.39 38.21

42.16 41.67

−2.40 −4.40

−15.61 −16.36

The effect of permeate pressure on separation performance of a pervaporation membrane has not received much attention. Most academic researches usually applied full or high vacuum, but it may not be practical or economic from industrial viewpoints. In the solution-diffusion model, the transport of a component across a membrane is driven by the trans-membrane pressure and/or chemical potential difference with the aid of applying vacuum or sweeping gas at the permeate side. As expressed in Eq. (8), an increase in permeate pressure, pp , will lower the driving force and reduce the flux across the membrane. Therefore, a critical permeate p pressure, pcrit,i , described in Eq. (24) exists in pervaporation where the driving force becomes zero and no permeation occurs: p

yi pcrit,i = xi i psat i

(24)

For a feed system of 80/20 wt.% EG/water, the values of xi , ␥i and psat of component water at 60 ◦ C can be calculated as follows: i xwater = 0.4628 mol% water = 1.583 (calculated by Hysis Distil software) ◦ psat water (60 C) = 19.94 kPa (199.4 mbar)

If the water concentration in the permeate, ywater , is very high and close to 1 due to the high separation performance of pervaporation membranes, the critical down stream pressure for water p transport, pcrit,water , can be approximated according to Eq. (24) as below: p

pcrit,water ≈ xwater water psat water = 146 mbar ≈ 0.144 atm

(25)

As a result, the driving force for water molecules to transport across a pervaporation membrane would be zero if the permeate pressure exceeds about 146 mbar for a feed system of 80/20 wt.% EG/water at 60 ◦ C. Figs. 8 and 9 show the effect of permeate pressure on separation performance of the PBI/PEI dual-layer membrane. As expected, both flux and permeance continually decrease with an increase in permeate pressure, while separation factor and selectivity show an up-and-down trend with increasing permeate pressure. Pervaporation could not be carried out successfully if the vacuum pressure was above 75 mbar because of no permeate available to be collected at higher vacuum pressures as stated in Section 3.4. The

Y. Wang et al. / Journal of Membrane Science 378 (2011) 339–350

Water flux

a

EG flux

20

600

15

400

10

200

5

Water Flux (g/m2/hr)

800

250

0.5

200

0.4

150

0.3

100

0.2 0.1

50

EG Flux (g/m2/hr)

25

Water Permeance EG permeance Selectivity

2

y

0

10

20

30

40

50

60

0 80

70

Fig. 9. Effect of permeate pressure on the permeance and selectivity.

performance trends of separation factor and its relationship with selectivity may be explainable using Eq. (21). ˛1/2 =

 1



 M1

xw,1 · · ˇw,1/2 · f1 − ˇw,1/2 − M2 xw,2 f2

· pp − f2



Since most parameters in Eq. (21) for a feed system of 80/20 wt.% EG/water at 60 ◦ C can be found as follows, f1 = fwater = xwater water psat water = 146 mbar f2 = fEG = xEG EG psat EG = 0.5372 mol% × 1.272 × 2.2 mbar = 1.50 mbar

M2 = MEG = 62 g/mol

b

Water permeance EG permeance

350 300

1500

250 200

1000 150 100

500

50 0 50

60

70

80

90

0 100

Feed EG Concentration (wt.%)

is fully swollen and becoming resistance for both water and EG transports.

Eq. (21) can be rewritten as follows



ˇw,1/2





˛1/2

20 1 = · · 146 · ˇw,1/2 − 80 1.5

˛1/2

1 = · [146 · ˇw,1/2 − (ˇw,1/2 − 0.29) · pp − 146] 6

18

1 − 62

p

· 18 · p − 146

(26)

As shown in Fig. 9, the selectivity obtained from experiments is much bigger than 0.29, the item 0.29·pp is negligible. Thus, Eq. (26) can be further simplified as below: ˛1/2 =

2000

Fig. 10. Effect of feed composition on flux and permeance for the PBI/PEI dual-layer hollow fiber membranes of (a) PBI-D-A and (b) PBI-D-B.

M1 = Mwater = 18 g/mol



0.0

0

Downstream Pressure (mbar)

EG permeance (g/m2-mbar-hr)

0

Water Permeance (g/m2-mbar-hr)

Permeance (g/m ·bar·hr)

1000

Selectivity

346

1 · ˇw,1/2 · (146 − pp ) − 24.3 6

(27)

Therefore, separation factor decreases with an increase in permeate pressure pp if ˇ remains unchanged. Since selectivity ˇ actually changes with an increase in permeate pressure pp as illustrated in Fig. 9, the evolution of separation factor with increasing permeate pressure follows the same pattern of selectivity because pp is much smaller than 146 mbar. There are many complicated factors affecting the relationships among selectivity, permeance and permeate pressure. Permeance decreases may be caused by an increase in transport resistance of the membrane porous substrate due to the severer swelling under poor vacuum conditions. The initial selectivity increase is probably owing to two factors; namely, (1) the reduction of the EG permeance since its fugacity at the feed side is very low (1.50 mbar). According to aforementioned discussion, a permeate pressure higher than 1.50 mbar could lead to a sever decline of the EG permeance and (2) the solvent-induced swelling that may cause different degrees of additional transport resistance for water and EG transports. Thus, the selectivity increases initially with increasing permeate pressure because of low EG fugacity and smaller water molecule size. However, a further increase in permeate pressure may lead to the loss of selectivity because the substructure

4.3. Effect of feed composition on pervaporation performance Figs. 10 and 11 show the effects of feed composition on separation performance of PBI/PEI dual-layer hollow fiber membranes. Both water flux and permeance decrease while both EG flux and permeance exhibit an up-and-down trend with an increase in EG concentration. Accordingly, a down-and-up trend is observed for both separation factor and selectivity. There are three factors associated with water permeance decrease with increasing EG concentration in the feed: (1) less swelling of the PBI membrane, (2) the formations of clusters among feed components, and (3) an EG boundary layer near membrane surface. PBI membranes can be swollen by both EG and water molecules because of hydrogen bonding interactions between their hydroxyl groups and the imidazole group of PBI. However, water causes more significant swelling since it has much smaller kinetic diameter and able to penetrate into the polymer chains easily. Water molecules trapped between polymer chains widen the chain–chain distance (d-space) and increase higher free volume of the membrane. The swelling also decreases the required energy for components to transport through membrane and therefore increases their permeation flux. When water concentration decreases, more intermolecular forces form among polymer chains, resulting in a more dense structure [41], and lower diffusivities of the feed components through the membrane. Since the EG/water mixture is a polar–polar system, possible pairs exist among feed components because of their mutual interactions, including water–water, EG–EG, and water–EG clusters [35]. These clusters have larger kinetic diameters and may cause a reduction in water diffusivity through the membrane [35]. A close observation of Fig. 10 shows that the water flux (permeance) drops more drastically at higher EG concentrations. A trendline is added

100,000

a

10,000

1,000 12

b

Selectivity (water/EG)

10

Vapor/permeate EG Composition (wt. %)

Separation factor (Water/EG)

Y. Wang et al. / Journal of Membrane Science 378 (2011) 339–350

347

1.0

0.8

0.6

0.4

0.2 a

0.0 0.0

b

0.2

0.4

0.6

0.8

1.0

Liquid/feed EG Composition (wt. %)

8

Fig. 12. (a) Vapor Liquid Equilibrium (VLE) curve compared with (b) the experimental result of the pervaporation separation for the binary EG–water mixture.

6 4 2 0 50

60

70

80

90

100

Feed EG Concentration (wt.%) Fig. 11. Effect of feed composition on separation factor and selectivity.

to fit to the water flux and permeance curves and a 2nd order polynomial fit gives an R2 value closest to unity. This is probably because that EG molecules forms an EG boundary layer near membrane surface with extensive inter-chain hydrogen bonding at high EG concentrations [42]. The additional resistance from the boundary layer may account for the drastic drops of water and EG fluxes at higher EG concentrations. The up-and-down trends of the EG flux and permeance with increasing EG concentration in the feed may be ascribed to the combined effect of membrane swelling at high water content and EG–EG clusters formed at high EG content. With an increase in EG concentration, more EG–EG clusters could be formed, which leads to a higher EG permeance when water concentration is sufficient so that the PBI membrane is swollen. With less water in the feed, the membrane cannot become swollen enough to allow EG transport. Thus, EG permeation rate decreases drastically when its concentration in the feed is higher than 90 wt.% and reaches nearly zero when the feed contains 98–99 wt.% EG. In addition, as mentioned above, an EG boundary layer forms near the membrane surface at high EG feed concentrations. Therefore, there is a tremendous drop of EG permeation flux to near zero at high EG feed concentrations. Similar EG flux changes with feed composition have also been obtained by previous researchers [25,35]. Fig. 11 shows the trends of separation factor and selectivity with an increase in EG concentration in the feed. The separation factor is plotted in a logarithmic scale against feed concentration because of a large variation of separation factor over the testing range. These performance trends correlate well with the changes of EG flux and permeance shown in Fig. 10. A sudden increase in separation factor is observed when the feed EG concentration is higher than 95 wt.%. A similar phenomenon has also been reported by Du et al. [35], where a separation factor of 32,900 at 99 wt.% feed EG composition was reported. In this study, separation factors of the same magnitude (43,000 and 30,000) are also achieved at 98 and 99 wt.% feed EG compositions. For reader information, separation factors at low and high feed EG concentrations (i.e., 50, 98 and 99 wt.%) might

not be accurate and should only be referred as a rough guide. At a low feed EG concentration of 50 wt.%, the EG concentration in the permeate is very low and probably beyond the detectable level of GC, which introduces uncertainty in the measurement. At high feed EG concentrations of 98 and 99 wt.%, any minor deviation on the measured permeate concentration can lead to a large variation of separation factor. For pervaporation to stay competitive compared in industrial applications, it must show significant advantages or superior performance over other existing separation technologies. For dehydration of concentrated organics, distillation has been the most reliable technology currently. Fig. 12 shows a comparison of separation efficiency between pervaporation and conventional distillation for the EG/water mixture. The Vapor Liquid Equilibrium (VLE) curve as shown in curve a is generated by the Wilson-ideal equation using Aspentech HYSYS with the aid of Wilson-ideal fluid package; while curve b is the permeate concentration from the experimental pervaporation separation process at various feed concentrations from this work. In the VLE curve, coupling effect and other non-idealities of EG and water in the liquid phase are accounted via Wilson equation. The ideal gas equation of state (EOS) is chosen to describe the vapor phase (ideal gas EOS is valid at very low pressures as vacuum is applied at the permeate side of the membrane). From Fig. 12, we observe that the EG content in the permeate from the pervaporation separation is almost zero, while the EG vapor content is much higher in the VLE equilibrium state, especially with high liquid/feed EG compositions. This is because pervaporation is not affected by the VLE curve and the separation is mainly based on solubility and diffusivity selectivity of the membrane. Therefore, pervaporation shows superior separation performance for EG dehydration compared to classic distillation separation. In addition, substantial energy savings can be achieved in pervaporation compared to classic distillation at high EG concentrations. Adsorption could also be a possible alternative technology for separating EG from water, but pervaporation is advantageous at higher EG concentrations, since it can be applied as a continuous process without limitation of a saturation capacity. 4.4. Long-term pervaporation tests For industrial applications, it is necessary that membrane can perform consistently over an extended period of time. A longer operational life of a membrane means reduced operating cost and enhanced economical feasibility of the pervaporation process.

Y. Wang et al. / Journal of Membrane Science 378 (2011) 339–350

150

100

120

99

90

98

60

97 96

30 Flux Water conc. in permeate

0 0

5

10

15

20

25

30

95 35

Water conc. in permeate (wt.%)

Flux (g/m 2hr)

348

Operation Time (day) Fig. 13. Long term pervaporation performance of PBI/PEI dual-layer hollow fiber membranes (total flux and permeate composition).

(4) With the increase in EG feed concentration, water flux and permeance drop while EG flux and permeance show up-and-down trends; accordingly the separation factor and selectivity are of down-and up trends. This can be explained by the combined effects of the driving force, EG–EG clusters, swelling effect and the EG boundary layer. (5) Through the comparison of the VLE curve and the permeate concentration curve from the experimental pervaporation results, pervaporation shows superior separation performance for EG dehydration compared to classic distillation especially at high EG concentrations. (6) The dual-layer hollow fiber membranes exhibit good long-term stability; at least up to 33 days for pervaporation separation of an 80 wt.% EG aqueous solution at 60 ◦ C, as demonstrated by this study. Acknowledgements

Fig. 13 presents the separation performance of PBI/PEI dual-layer hollow fiber membranes for pervaporation dehydration of an 80 wt.% EG/water solution at 60 ◦ C during a 33-day test. The water concentration in permeate and the total flux remain almost constant during the entire testing duration, indicating the long-term stability of the membrane. No obvious swelling of the hollow fiber is observed. Here the separation performance is expressed in terms of permeate water concentration because the separation factor is very high and the exact value of permeate water concentration cannot be obtained accurately. Clearly, this result illustrates that the combination of PBI as the outer selective layer and PEI as the inner supporting layer with the aid of dual-layer membrane fabrication can synergize the stable separation performance effectively. 5. Conclusion In this study, PBI/PEI dual-layer hollow fiber membranes have been studied for pervaporation dehydration of ethylene glycol. Effects of operation temperature, permeate pressure, feed composition and operation time have been investigated. The following conclusions can be made: (1) Governing equations correlating separation performance with system operation conditions and intrinsic separation characteristics of pervaporation membranes have been derived based on the solution-diffusion model. The derived equations suggest (1) the existence of a critical permeate pressure in pervaporation where the driving force becomes zero and no permeation occurs and (2) the variation of separation factor is determined by the combined effect of the permeate pressure, membrane intrinsic selectivity and its relations with penetrants and permeate pressure. (2) Sorption and cluster formation of feed components play important roles on membrane performance. Since an increase in operation temperature results in a decrease in sorption of penetrants and less EG–water clusters, both flux and selectivity increase but both permeance and separation factor decrease with increasing operation temperature. (3) The total flux and permeance decrease, while separation factor and selectivity show an up-and-down trend with increasing permeate pressure. The decreased flux and permeance are mainly due to the reduced driving force and the increased transport resistance. The change of the selectivity is probably caused by low EG fugacity, smaller water molecule size and the change of membrane morphology because of severe swelling under poor vacuum.

The authors thank PBI Performances Products, Inc. (R-279-000279-597) for funding this research. Special thanks are due to Dr. Jing Cai Su for help with hollow fiber spinning. Appendix A. Derivation of the separation factor–selectivity relationship for a pervaporation system For a pervaporation system with a significant permeate pressure, the selectivity of component 1 with respect to 2 is as follows:

ˇw,1/2 =

yw,1 /(xn,1 1 psat − yn,1 pp ) 1 yw,2 /(xn,2 2 psat − yn,2 pp ) 2

The above equation can be rewritten with the aid of Eqs. (7) and (14) and becomes: (f2 /yw,2 ) − (yn,2 · pp /yw,2 )

ˇw,1/2 =

(f1 /yw,1 ) − (yn,1 · pp /yw,1 )

=

(f2 /yw,2 ) − (pp /(M2 · (f1 /yw,1 ) − (pp /(M1 ·

 (yw,i /Mi )))  (yw,i /Mi )))

By multiplying both sides with the denominator, the above equation can be rearranged as:

 ˇw,1/2 ·

ˇw,1/2 ·

ˇw,1/2 ·

f1 yw,1

f1 yw,1 f1 yw,1

−

−

−



pp

M1 · f2

yw,2 f2 yw,2



=

(yw,i /Mi )

= ˇw,1/2 ·

 =

yw,2

−

pp

M1 ·



(yw,i /Mi )

1 − M2

 

·



M2 ·

pp

ˇw,1/2 M1

f2

−

(yw,i /Mi ) pp

M2 ·



(yw,i /Mi )

pp

(yw,i /Mi )

If the pervaporation membrane has very high separation efficiency, which is the case for most high-performance dehydration membranes, then yw,1 can be approximately to be 1 and the above equation can be simplified as: ˇw,1/2 · f1 −

f2 yw,2

 =

ˇw,1/2 M1

1 − M2

 ·



pp

(yw,i /Mi )

Thus the weight-based permeate composition of component 2, yw,2 can be derived. yw,2 =

f2 ˇw,1/2 · f1 − ((ˇw,1/2 /M1 ) − (1/M2 )) · pp /



(yw,i /Mi )

Y. Wang et al. / Journal of Membrane Science 378 (2011) 339–350

Nomenclature DMAc EG LiCl PAA PAN PBI PEI PVA PVP A EJ EP

n,n-dimethylacetimide ethylene glycol lithium chloride poly (acrylic acid) polyacrylonitrile polybenzimidazole polyetherimide poly (vinyl alcohol) polyvinylpyrrolidone membrane area (m2 ) apparent activation energies of flux (kJ/mol) apparent activation energies of permeability (kJ/mol) fi fugacity (partial vapor pressure) of component i at the feed side (mbar) J flux (g/m2 h) JN normalized flux (g ␮m/m2 h) pre-exponential factor for the component flux J0 (g/m2 h) L selectivity layer thickness of the membrane (␮m) Mi molecular weight of the component i (g/mol) pre-exponential factor for the component permeP0 ance (g/m2 bar h) Pi permeability of the component i (g ␮m/m2 bar h) [Pi /l] or P¯ i permeance of the component i (g/m2 bar h) saturated vapor pressure of the component i (mbar) psat i pp permeate pressure (mbar) p pcrit critical permeate pressure (mbar) pervaporation separation index (g/m2 h) PSI Q total mass passing through membrane (g) R universal gas constant (J/K mol) R-squared value of the trendline R2 t operation time (h) operating temperature (◦ C) T xn,i mole fraction of the component i in the feed (mol%) weight fraction of component i in the feed (wt.%) xw,i yn,i mole fraction of the component i in the permeate (mol%) weight fraction of component i in the permeate yw,i (wt.%) ˛ separation factor selectivity ˇ i activity coefficient of the component i Subscripts 1 component 1 with higher or preferred permeability 2 component 2 with lower or less preferred permeability (separation efficiency), to separate component 1 1/2 over component 2 mole-based N i component i in the feed N normalized by membrane thickness J flux permeance P J,w water flux J,EG EG flux P,w water permeance EG permeance P,EG W weight-based

349

Superscripts permeate side p sat saturated vapor pressure

References [1] Y. Wang, S.H. Goh, T.S. Chung, N. Peng, Polyamide-imide/polyetherimide dual-layer hollow fiber membranes for pervaporation dehydration of C1–C4 alcohols, J. Membr. Sci. 326 (2009) 222. [2] Y. Wang, M. Gruender, T.S. Chung, Pervaporation dehydration of ethylene glycol through polybenzimidazole (PBI)-based membranes. 1. Membrane fabrication, J. Membr. Sci. 363 (2010) 149. [3] K.Y. Wang, T.S. Chung, R. Rajagopalan, Dehydration of tetrafluoropropanol (TFP) by pervaporation via novel PBI/BTDA-TDI/MDI co-polyimide (P84) dual-layer hollow fiber membranes, J. Membr. Sci. 287 (2007) 60. [4] R.X. Liu, X.Y. Qiao, T.S. Chung, Dual-layer P84/polyethersulfone hollow fiber for pervaporation dehydration of isopropanol, J. Membr. Sci. 294 (2007) 103. [5] H.M. Guan, T.S. Chung, Z. Huang, M.L. Chng, S. Kulprathipanja, Poly(vinyl alcohol) multilayer mixed matrix membranes for the dehydration of ethanol–water mixture, J. Membr. Sci. 268 (2006) 113. [6] L.Y. Jiang, H. Chen, Y.C. Jean, T.S. Chung, Ultra-thin polymeric interpenetration network with separation performance approaching ceramic membranes for biofuel, AIChE J. 55 (2009) 75. [7] J.H. Kim, K.H. Lee, S.Y. Kim, Pervaporation separation of water from ethanol through polyimide composite membranes, J. Membr. Sci. 169 (2000) 81. [8] H. Yanagishita, D. Kitamoto, K. Haraya, T. Nakane, T. Okada, H. Matsuda, Y. Idemoto, N. Koura, Separation performance of polyimide composite membrane prepared by dip coating process, J. Membr. Sci. 188 (2001) 165. [9] D.M. Sullivan, M.L. Bruening, Ultrathin cross-linked polyimide pervaporation membranes prepared from polyelectrolyte multilayers, J. Membr. Sci. 248 (2005) 161. [10] D.D. Jiang, W. Liu, Polymer Hybrid Membrane Structures, US Patent 20,090,110,873 (2009). [11] D.D. Jiang, W. Liu, P.J. Shustack, J. Wang, Membranes Based on Poly (Vinyl Alcohol-Co-Vinylamine), US Patent 20,090,110,907 (2009). [12] R. Jiraratananon, A. Chanachai, R.Y.M. Huang, D. Uttapap, Pervaporation dehydration of ethanol–water mixtures with chitosan/hydroxyethylcellulose (CS/HEC) composite membranes: I. Effect of operating conditions, J. Membr. Sci. 195 (2002) 143. [13] L. Lin, G. Wang, H. Qu, J. Yang, Y. Wang, D. Shi, Y. Kong, Pervaporation performance of crosslinked polyethylene glycol membranes for deep desulfurization of FCC gasoline, J. Membr. Sci. 280 (2006) 651. [14] E.A. Fouad, X.S. Feng, Use of pervaporation to separate butanol from dilute aqueous solutions: effects of operating conditions and concentration polarization, J. Membr. Sci. 323 (2008) 428. [15] M. Peivasti, A. Madandar, T. Mohammadi, Effect of operating conditions on pervaporation of methanol/methyl tert-butyl ether mixtures, Chem. Eng. Process. 47 (2008) 1069. [16] S. Sommer, T. Melin, Influence of operation parameters on the separation of mixtures by pervaporation and vapor permeation with inorganic membranes. Part 1: dehydration of solvents, Chem. Eng. Sci. 60 (2005) 4509. [17] J.A. Gonzalez-Marcos, C. Lopez-Dehesa, J.R. Gonzalez-Velasco, Effect of operation conditions in the pervaporation of ethanol–water mixtures with poly(1-trimethylsilyl-1-propyne) membranes, J. Appl. Polym. Sci. 94 (2004) 1395. [18] A. Bakhshia, T. Mohammadi, O. Ghaffari Nik, A. Aroujalian, Effect of operating conditions on pervaporation of methanol–water mixtures: part 1, Membr. Technol. 2006 (2006) 7. [19] S.H. Huang, G.J. Jiang, D.J. Liaw, C.L. Li, C.C. Hu, K.R. Lee, J.Y. Lai, Effects of the polymerization and pervaporation operating conditions on the dehydration performance of interfacially polymerized thin-film composite membranes, J. Appl. Polym. Sci. 114 (2009) 1511. [20] F.R. Chen, H.F. Chen, Pervaporation separation of ethylene glycol–water mixtures using crosslinked PVA–PES composite membranes. 1. Effects of membrane preparation conditions on pervaporation performances, J. Membr. Sci. 109 (1996) 247. [21] O.G. Nik, A. Moheb, T. Mohannnnadi, Separation of ethylene glycol/water mixtures using NaA zeolite membranes, Chem. Eng. Technol. 29 (2006) 1340. [22] R.L. Guo, C.L. Hu, F.S. Pan, H. Wu, Z.Y. Jiang, PVA-GPTMS/TEOS hybrid pervaporation membrane for dehydration of ethylene glycol aqueous solution, J. Membr. Sci. 281 (2006) 454. [23] R.L. Guo, C.L. Hu, B. Li, Z.Y. Jiang, Pervaporation separation of ethylene glycol/water mixtures through surface crosslinked PVA membranes: coupling effect and separation performance analysis, J. Membr. Sci. 289 (2007) 191. [24] C.L. Hu, R.L. Guo, B. Li, X.C. Ma, H. Wu, Z.Y. Jiang, Development of novel mordenite-filled chitosan–poly(acrylic acid) polyelectrolyte complex membranes for pervaporation dehydration of ethylene glycol aqueous solution, J. Membr. Sci. 293 (2007) 142. [25] C.L. Hu, B. Li, R.L. Guo, H. Wu, Z.Y. Jiang, Pervaporation performance of chitosan–poly(acrylic acid) polyelectrolyte complex membranes for dehydration of ethylene glycol aqueous solution, Sep. Purif. Technol. 55 (2007) 327.

350

Y. Wang et al. / Journal of Membrane Science 378 (2011) 339–350

[26] R.L. Guo, X.C. Ma, C.L. Hu, Z.Y. Jiang, Novel PVA-silica nanocomposite membrane for pervaporative dehydration of ethylene glycol aqueous solution, Polymer 48 (2007) 2939. [27] M.N. Hyder, R.Y.M. Huang, P. Chen, Composite poly(vinyl alcohol)–poly(sulfone) membranes crosslinked by trimesoyl chloride: characterization and dehydration of ethylene glycol–water mixtures, J. Membr. Sci. 326 (2009) 363. [28] R.Y.M. Huang, X. Feng, Dehydration of isopropanol by pervaporation using aromatic polyetherimide membranes, Sep. Sci. Technol. 28 (1993) 2035. [29] L.Y. Jiang, Y. Wang, T.S. Chung, X.Y. Qiao, J.Y. Lai, A review—polyimides in pervaporation membranes and biofuels separation, Prog. Polym. Sci. 34 (2009) 1135. [30] J.G. Wijmans, Process performance = membrane properties plus operating conditions, J. Membr. Sci. 220 (2003) 1. [31] W.F. Guo, T.S. Chung, T. Matsuura, R. Wang, Y. Liu, Pervaporation study of water and tert-butanol mixtures, J. Appl. Polym. Sci. 91 (2003) 4082. [32] X.Y. Qiao, T.S. Chung, W.F. Guo, T. Matsuura, M.M. Teoh, Dehydration of isopropanol and its comparison with dehydration of butanol isomers from thermodynamic and molecular aspects, J. Membr. Sci. 252 (2005) 37. [33] Y. Wang, L.Y. Jiang, T. Matsuura, T.S. Chung, S.H. Goh, Investigation of the fundamental differences between polyamide-imide (PAI) and polyetherimide (PEI) membranes for isopropanol dehydration via pervaporation, J. Membr. Sci. 318 (2008) 217.

[34] R.X. Liu, X. Qiao, T.S. Chung, The development of high performance P84 copolyimide hollow fibers for pervaporation dehydration of isopropanol, Chem. Eng. Sci. 60 (2005) 6674. [35] J.R. Du, A. Chakma, X. Feng, Dehydration of ethylene glycol by pervaporation using poly(N,N-dimethylaminoethyl methacrylate)/polysulfone composite membranes, Sep. Purif. Technol. 64 (2008) 63. [36] T.S. Chung, W.F. Guo, Y. Liu, Enhanced Matrimid membranes for pervaporation by homogenous blends with polybenzimidazole (PBI), J. Membr. Sci. 271 (2006) 221. [37] W.F. Guo, T.S. Chung, T. Matsuura, Pervaporation study on the dehydration of aqueous butanol solutions: a comparison of flux vs. permeance, separation factor vs. selectivity, J. Membr. Sci. 245 (2004) 199. [38] E. McLaughlin, A.R. Ubbelohde, Cluster formation in relation to the viscosity of melts in the pre-freezing region, Trans. Faraday Soc. 54 (1958) 1804. [39] L.M. Fox, Remington: The Science and Practice of Pharmacy, 21st edition., Lippincott Williams & Wilkins, Philadelphia, PA, 2005. [40] X.S. Feng, R.Y.M. Huang, Pervaporation with chitosan membranes. 1. Separation of water from ethylene glycol by a chitosan/polysulfone composite membrane, J. Membr. Sci. 116 (1996) 67. [41] S.I. Semenova, H. Ohya, K. Soontarapa, Hydrophilic membranes for pervaporation: an analytical review, Desalination 110 (1997) 251. [42] X. Feng, R.Y.M. Huang, Estimation of activation energy for permeation in pervaporation processes, J. Membr. Sci. 118 (1996) 127.