WO1995032051A1 - Pervaporation process with reduced driving force - Google Patents

Abstract

A pervaporation process characterized by a low driving force and low feed/permeate pressure ratio, obtained by operating at permeate pressures above 50 torr (6.7 KPa) or at low feed temperatures. The process is effective in controlling concentration polarization in pervaporation separations of components of high relative volatility, where the evaporative term has a significant effect on the overall separation factor. The process is particularly useful for separating VOCs from water, or for separating organic mixtures.

Description

PERVAPORATION PROCESS WITH REDUCED DRIVING FORCE

FIELD OF THE INVENTION

The invention relates to pervaporation. More particularly, the invention relates to improved pervaporation processes in which the adverse effects of concentration polarization are reduced. BACKGROUND OF THE INVENTION

Pervaporation is a membrane-based process used to separate solutions on the basis of differences in the volatilities or diffusion characteristics of the components. A liquid feed solution contacts one side of a membrane; the permeate is removed as a vapor from the other side. Transport through the membrane is induced by the difference in partial pressure between the liquid feed solution and the permeate vapor. This partial vapor pressure difference can be maintained in several ways, such as drawing a vacuum on the permeate side of the system, sweeping the permeate side with a carrier gas, or simply cooling the permeate vapor, causing it to condense.

Whether carrying out pervaporation experiments in the laboratory or operating larger-scale systems, it is conventional to work with a low pressure on the permeate side, and it is also common to heat the feedstream to some extent, to increase the driving force across the membrane. Table 1 lists some typical conditions identified in a variety of references.

TABLE 1.

Ref Separation Membrane Feed Permeate Pressure temp. pressure ratio

Type Thickness (°Q

1 water/ethanol PVA 10 μm 60 up to 200 mbar (20kPa)

2 benzene/water MTR100 40 15 torr 4 MTR200 (2 kPa)

3 ethanol/water PDMS 70 10 mbar (1 kPa) 30

4 aroma/water PDMS 10 μm 30-60 5xl0- - 50 mb 4-400 ethylacetate/water (<5 kPa)

5 water/IPA TexSep® 70-95 2-189 mmHg (0.3-25 kPa)

6 water/ethanol 100 -12°C

7 TCE/water various 80-150 μm 25 <3 mb 10 toluene/water elastomer (<0.3 kPa) water/ethanol GFT 70-93 -20°C Ref Separation Membrane Feed Permeate Pressure temp. pressure ratio

Type Thickness (°Q

9 ethanol/water PDMS 35 20 mb (2 kPa) 3

10 aromatic he/water PDMS 1-170 μm 0.5-2 mb 13 Pebax 50 1-10 mb (<l kPa)

11 organic/water PDMS 165 μm 23 <l torr (0.1 kPa) 20

12 phenol/water Pebax 42 μm 45 13 Pa 700

13 phenol/water Pebax 1-2 mil 50 5-10 torr 10-20 (0.6-1.3 kPa)

14 water/glycols PVA 80 5 torr (0.6 kPa)

15 water/ketones PVA 65 5 torr (0.6 kPa)

16 water/ PVA 60 2 torr (0.3 kPa) MEK-toluene

17 various various 1-11 mil 25-70 O.l torr 200- organics/water (0.01 kPa) 2,000

18 benzene/water 40 15 torr (2 kPa) >4

19 TCE water MTR 100 30-58 5-40 torr 1-30 ethanol/water MTR 200 (0.6-5.3 kPa) ethylacetate/water

20 various MTR 100 40 10-15 torr >4 organic/water MTR 200 (1.3-2 kPa)

21 ethanol/MEK various 25 9.5-50 torr (1.3-6.7 kPa)

22 ethanol/water PDMS 2 μm 30 up to 6cmHg 1-15 (8 kPa)

23 various butadiene 200-300 30 1-43 mmHg 1-40 organics water copolymer μm (0.1-5.7 kPa)

24 TCE/water PDMS 165 μm 20 3 torr (0.4 kPa) 6

25 ethanol/water PDMS 2.2 μm 30 0.4-4cmHg 1.5-15 ethylacetate/water 1 μm 30 1.5-3 cmHg

(up to 5.3 kPa)

26 water/isopropanol PEI 25 133 Pa

27 water/mineral oil 50 120 mb (12 kPa)

28 TCE/water PDMS 165 μm 25 <1 torr (0.1 kPa) 24

29 TCE water various 5-70 μm 25 <3 mb (0.3 kPa) 10 From the table, it can be seen that typically, permeate pressures are low, such as a few torr, and a typical feed temperature might be 60°C. Typical pressure ratios for workable processes are up to 100 or more. The conventional belief in the art is that the lower the permeate pressure, the better is the separation performance. Many references can be identified that support this belief. For example: 1. A paper by Brun et al. (J.-P. Brun, C. Larchet, R. Melet and G. Bulvestre, "Modelling of the Pervaporation of Binary Mixtures through Moderately Swelling, Non-Reacting Membranes", Journal of Membrane Science, Vol. 23, 1985, pp. 257-283) models the pervaporation process. Figure 2 shows a calculated plot of pervaporation selectivity (separation factor) as a function of permeate pressure. The figure shows a sharp decline in separation factor with increasing permeate pressure. Figure 3 also shows that higher permeate pressures result in poorer separation.

2. A paper by Rautenbach et al. ("The Separation Potential of Pervaporation: Part 1. Discussion of Transport Equations and Comparison with Reverse Osmosis", Journal of Membrane Science, Vol. 25 (1985)) presents a discussion of the pervaporation transport equations and again includes graphs (Figures 8 and 9) that show separation factor falling with increasing permeate pressure. 3. U.S. Patent 4,941,976 recommends a permeate pressure of 10 torr (1.3 kPa), U.S. Patent 4,935,144 recommends 5 torr (0.7 kPa), and U.S. Patent 5,004,861 recommends 2 torr (0.3 kPa). 4. A paper by Lichtenthaler (Reference 1) discusses factors that determine overall mass transport in pervaporation and concludes, amongst other conclusions, that the permeate pressure should be as low as possible (Figure 20). 5. A paper by Lamer et al. (Reference 4) reporting on the pervaporation of aroma compounds shows a decrease of selectivity from over 160 to about 25 as downstream permeate pressure rises from 0.1 mbar to 50 mbar (0.01-5 kPa).

6. A papa^* by Reale et al. (Reference 5) on removal of water or methanol from organic compounds shows in Figure 7 how the permeate concentration of isopropanol (the rejected component) increases with increasing permeate pressure and shows that a significant loss in performance occurs beginning at about 40 torr (5.3 kPa).

7. A paper by Neel et al. ("Influence of Downstream Pressure on the Pervaporation of Water- Tetrahydrofuran Mixtures through a Regenerated Cellulose Membrane (Cuprophan), Journal of Membrane Science, Vol.27 (1985)) examines the influence of downstream pressure on dehydration of tetrahydrofuran through cellulose membranes. In this case, membrane swelling plays a big part in determining separation properties. Again, however, loss of performance as permeate pressure increases is reported (Figure 4).

In all membrane processes, other factors than the membrane properties may influence the separation. As with any fluid flowing across a surface, the velocity profile in a feed solution subjected to pervaporation is not constant across the thickness of the solution layer, because of friction at the solution/membrane interface. The feed solution velocity decreases as the distance from the membrane surface decreases and a stagnant boundary layer is present near the membrane surface. The solution concεntratic-i is uniform outside the stagnant boundary layer, because the flow is turbulent. However, the flow in the boundary layer is laminar, producing a concentration profile across this layer as the faster permeating components are removed preferentially through the membrane. The boundary layer acts as an additional resistance, in series with the membrane, to transport of material from the bulk feed to the permeate side of the membrane. If the membrane has the capacity to transport a component faster than the diffusion rate of that component into the boundary layer, the depleted boundary layer may in fact be the dominant resistance to mass transport, thereby diminishing die driving force across the membrane and reducing membrane separation performance significantly. The effect of concentration polarization is that components that are enriched in the permeate are depleted in the boundary layer, and components that are depleted in the permeate are enriched in the boundary layer. In other words, concentration polarization works against the separation achieved by the membrane, and should be avoided as much as possible.

Boundary layer concentration polarization problems are exacerbated by a low feed flow rate, a low feed concentration of the faster permeating component and a high separation factor between the components to be separated. Unfortunately, many pervaporation separations, such as the removal of small amounts of organic contaminants from water, meet one or several of these criteria. Since the water concentration is high both in the bulk feed and in the boundary layer region, boundary layer effects do not affect water transport adversely. However, die organic/water separation factor is often very high, such as in the hundreds or thousands, and die concentration of the organic in the water is often low, such as in the ppm range, so the effects of concentration polarization on the organic components that are to be removed are often very severe. The standard way of minimizing concentration polarization is to maximize mixing at the membrane surface, that is to minimize the thickness of the boundary layer. This can be done by increasing the feed fluid velocity, thereby promoting turbulence. In practical systems, however, there is a limit to the extent to which this can be done.

In addition, recent references that discuss concentration polarization in pervaporation make the point that, if the feed flow rate cannot be made large enough to overcome concentration polarization problems, as is often the case, another tactic is to use a thicker membrane. For example, "Mass Transfer Characteristics of a New Pervaporation Module for Water Purification", a paper presented by Gooding, Hickey and Crowder at the Fifth International Conference on Pervaporation Processes in the Chemical Industry, lists the factors that affect removal of organics from dilute aqueous streams as: permeate pressure; Henry's Law coefficient; feed concentration; membrane thickness; and liquid phase mass transfer coefficient. It further states that: (a) Permeate pressure must be maintained at a low value,

(b) A large Henry's Law coefficient increases driving force for pervaporation of that component and reduces membrane resistance,

(c) The membrane should be reasonably thin, but, unless the mass transfer coefficient in the liquid boundary layer is large, the boundary layer can easily dominate the total mass transfer resistance, so that making the membrane thinner or more organic-permeable is fruitless.

An earlier paper entitled by Cote and Lipski (Reference 11) expresses similar thoughts. If one is dealing with volatile, hydrophobic compounds (that is, the Henry's Law coefficient is large) the separation is generally boundary-layer-limited and, if it cannot be made membrane-limited by increasing feed flow rate and turbulence, one is better off using a thick membrane than a thin membrane, because the thick membrane will give a better separation factor for the same mass transfer rate. These ideas are reiterated in "The Use of Pervaporation for the Removal of Organic Contaminants from Water" (Reference 28), again by Lipski and Cote; in "A Technico-Econornical Evaluation of Pervaporation for Water Treatment", a paper presented by Cote and Lipski at the Fourth International Conference on Pervaporation Processes in the Chemical Industry; and in U.S. Patent 5,167,825. Each of these teaches that a relatively thick membrane yields better separation performance of organics from water because the thick membrane decreases the water flux to a greater extent than it decreases d e VOC flux, thereby improving the separation factor. As is pointed out in 5,167,825, rubbery membranes are preferred for separating organics from water and very thick rubbery membranes are difficult to make. In '825, the solution to this problem is to impregnate the rubbery selective polymer into the pores of a microporous support, thereby producing a selective layer with an effective thickness of as much as 200 μm.

In these references and other analyses of concentration polarization effects in pervaporation, permeate pressure is generally ignored, or its effect on concentration polarization is not recognized. For example, a poster by H.H. Nijhuis et al. (International Symposium on Synthetic Membranes in Science and Industry, Tubingen, Germany, 1989) gives an analysis of mass transfer resistance on removal of trace organics from aqueous solutions that assumes that the permeate pressure is zero. Likewise, the doctoral thesis of H.H. Nijhuis (Reference 29) lists factors affecting pervaporation performance as feed concentratioQ, boundary layer mass transfer coefficient, membrane permeability and membrane thickness, but does not list or discuss driving force effects. In two of the three chapters that make up the experimental portion of the thesis, permeate pressure is not even mentioned. A paper by Psaume (Reference 24) focuses on the importance of concentration polarization in removing organics from water, but ignores the effects of permeate pressure in the analysis. To our knowledge, an understanding of die important effect of permeate pressure on concentration polarization in certain pervaporation separations has not previously been available to the art. SUMMARY OF THE INVENTION

The invention concerns separation of liquids with high relative volatilities. One example is the separation of volatile organic compounds (VOCs) from water. Other examples include separation of low molecular weight, high volatility organics from high molecular weight, lower volatility organics, dehydration of hydrophobic or low volatility organics, and removal of dissolved gases, particularly air or oxygen, from water or other liquids.

In such processes, we have discovered that reducing the transmembrane driving force, by raising the permeate pressure and/or lowering the feed temperature, actually increases the separation factor. This is a surprising result, which is contrary to conventional teachings about membrane processes, as will be explained in the detailed description that follows.

This discovery provides a useful and beneficial advance in pervaporation technology. These teachings can be adopted in an improved pervaporation process by carrying out the process at higher permeate pressure or lower feed temperature than was previously normal practice. In particular, the invention involves carrying out the pervaporation process at a permeate pressure greater than about 50 torr

(6.7 kPa), and preferably even higher, such as greater than about 60 torr (8 kPa) or 100 torr (13.3 kPa).

In another aspect, the invention involves carrying out the pervaporation process at a feed temperature no greater than about 60°C, and preferably no greater than about 50°C.

In another aspect, the invention involves carrying out the pervaporation process at a lower pressure ratio than was previously normal practice. In particular, the invention involves carrying out the pervaporation process at a pressure ratio lower than about 4, preferably lower, such as below 3, and more preferably even lower, such as below 2. Depending on the vapor pressures of the components to be separated, pressure ratios within these ranges can be obtained by using a permeate pressure greater than about 35 torr (4.7 kPa) in conjunction with a feed temperature lower than about 55 °C, or a permeate pressure greater than about 50 torr (6.7 kPa) with a feed temperature lower than about 60°C, or a permeate pressure greater than about 70 torr (9.3 kPa) with a feed temperature lower than about 50°C, for example. Pervaporation processes in accordance with the invention have the following advantages: 1. Concentration polarization problems can be ameliorated

The invention provides a simple technique to keep boundary layer resistance effects under control in certain pervaporation processes. 2. Energy efficicency

(a) A lower feed temperature may be used

The low driving force required to carry out the process can be achieved by operating at low feed temperature, high permeate pressure or a combination of both. It is not unusual for recommended feed temperatures for pervaporation to be as high as 80° C or even higher. By operating under mild feed temperature conditions, the energy required to heat the feed liquid is correspondingly small.

A second advantage is that many polymer membranes are not resistant to high temperatures. By operating at temperatures below about 60 °C, even the most temperature-sensitive membrane materials can be used.

A third advantage is that feed streams containing labile compounds, such as may be encountered in the food or pharmaceutical industries, for example, may be treated.

(b) A higher permeate pressure mav be used

In the past, pervaporation processes have often been carried out at very low permeate pressures, such as 20 torr (2.7 kPa), 10 torr (1.3 kpa) or below. For laboratory-scale research experiments, low pressure on the permeate side is often obtained by means of a vacuum pump. This pump removes the totality of the permeate stream, including condensable vapors and non-condensable gases, such as air, that may have been dissolved in the feed liquid or have leaked into the apparatus. The need to pass everything through the vacuum pump makes this mode of operation difficult except for small-scale laboratory use.

Largo^* systems often obtain partial vacuum conditions on d e permeate side by using a condenser to liquefy the condensable vapor, so that the pressure on the permeate side is determined by the equilibrium vapor pressure at the condenser temperature. Such systems still need a vacuum pump to remove ncm-condensable gases, which are almost always present to some extent, but the load on the vacuum pump is much reduced. Nevertheless, to reach a pressure as low as a few torr, the condenser temperature often has to be substantially below ambient temperature, so that chilling or even refrigeration may need to be employed. In the case of aqueous solutions, me permeate normally cannot be chilled below 0 ° C. In the case of solutions with lower freezing points, such as ethanol/water mixtures, for example, lower condensation temperatures, such as down to -20 °C, can be and have been employed. For example, a paper by J.L. Rapin, entitled "The Betheniville Pervaporation Unit for the First Large-Scale Productive Plant for die Dehydration of Ethanol" (Reference 8) describes die first GFT ethanol dehydration plant in which three condensers operating at temperatures of 10°C, -5°C and -20°C were used and in which the low temperatures were reached by refrigeration with Freon 22. The lower the cooling temperature that is used, die greater is the energy expenditure. Another issue is the presence of the non-condensable gases. Suppose, for example, that the desired permeate pressure is 20 torr (2.7 kPa). At a particular equilibrium condition determined by temperature, this total permeate pressure may be made up of 19 torr (2.5 kPa) partial pressure of condensables and 1 torr (0.01 kPa) of non-condensable gases. In that case, however, the off-gas vented from the vacuum pump that is removing the non-condensables will contain 19 parts of vapor to every 1 part of non-condensable gas. To prevent this vapor loss, it may be necessary to set the condensation temperature even lower, so ti at the vapor pressure contributes a smaller fraction of the total pressure. This may mean that, to obtain a total permeate pressure of 20 torr (2.7 kPa), the condenser temperature may have to be set to correspond to an equilibrium vapor pressure of only 10 torr (1.3 kPa) or less, for example. In contrast, permeate pressures of 50 torr (6.7 kPa), 60 torr (8 kPa), or above can usually be obtained by condensation at temperatures of 5-25 °C. Besides saving energy, this substantially simplifes the equipment, since large refrigeration units or large vacuum pumps are not needed, and thereby improves operating reliability and decreases costs.

3. The permeate fraction is more concentrated/has a smaller volume The principal product of a pervaporation process may be the residue stream, or the permeate stream, or bodi may be important If the goal of the process is to remove an unwanted contaminant, such as an organic solvent from an industrial wastewater stream, then the clean residue stream is the product, and the solvent-enriched permeate is a secondary waste stream. The more concentrated in solvent is this stream, the easier it often becomes to carry out further treatment, such as to recover usable products. If the stream has to be disposed of, then, generally, the smaller the volume, the lower are the disposal costs. If the product is the permeate stream, then it is usually the case that the higher the purity, the greater is the value of the product, so high concentration in a small volume is again desirable.

4. Very thick membranes need not be used

As discussed in the Background section above, one way to handle concentration polarization problems is to use thick membranes. To separate organic compounds from water, rubbery membranes widi an effective diickness as much as 200 μm are advocated in U.S. Patent 5,167,825. In a program that screened a large number of elastomeric membrane materials for their VOC/water separating properties, H.H. Nijhuis et al. (Reference 7) used membrane stamps with a thickness of 60-80 μm for silicone rubber and 80-150 μm for the other materials.

The processes of the invention provides a technique that relies on driving force, rather than membrane thickness, to circumvent concentration polarization problems. The processes of the invention can, therefore, be carried out using any membrane of any convenient thickness that otherwise fits the requirements of the separation.

It is an object of the invention to provide pervaporation processes in which the adverse effects of concentration polarization are reduced.

It is an object of the invention to provide pervaporation processes in which very low permeate pressures need not be used.

It is an object of the invention to provide pervaporation processes in which very low condenser temperatures need not be used.

It is an object of the invention to provide pervaporation processes in which high feed temperatures need not be used. It is an object of the invention to provide pervaporation processes suitable for removing VOCs from water.

Other objects and advantages of the invention will be apparent from the description of the invention to those of ordinary skill in the art.

It is to be understood that the above summary and die following detailed description are intended to explain and illustrate the invention without restricting its scope.

BRIEF DESCRIPTION OF THE DRAWINGS

Figure 1 is a graph of separation factor as a function of pressure ratio for separation of trichloroethylene

(TCE) from water. Figure 2 is a plot of separation factor as a function of feed temperature for separation of toluene from water at three different feed flow rates.

Figure 3 is a graph of separation factor as a function of feed temperature for separation of toluene from water.

Figure 4 is a graph of TCE reduction factor as a function of stage cut at three different feed temperatures for separation of TCE water.

Figure 5 is a graph of separation factor as a function of feed temperature for separation of TCE from water. Figure 6 is a graph of separation factor as a function of permeate pressure for separation of toluene from water at three different feed flow rates.

Figure 7 is a graph of TCE reduction factor as a function of stage cut at two different driving forces for separation of TCE from water. Figure 8 is a schematic of the process design used for calculations in Examples 18-21.

Figure 9 is a set of gas C-τrorr__tography traces of feed, residue and permeate stream samples from a BTEX removal experiment.

Figure 10 is a graph of permeate concentration as a function of feed concentration for BTEX removal experiments at 10 torr (1.3 kPa) and 70 torr (9.3 kPa) permeate pressure. Figure 11 is a graph of BTEX flux as a function of feed solution concentration for BTEX removal experiments at 10 torr (1.3 kPa) and 70 torr (9.3 kPa) permeate pressure.

Figure 12 is a set of graphs showing feed and permeate concentrations for a set of pervaporation experiments carried out under different driving forces.

DETAILED DESCRIPTION OF THE INVENTION

The invention is a pervaporation process that is applicable to pervaporation separations character-zed by a high relative volatility between the components to be separated, and a high separation factor. The invention is particularly useful in the separation of VOCs from water, and for simplicity, the description of die invention below defines one component as VOC, the other as water. However, the invention is useful for any high relative volatility set of components and is not to be construed as limited to VOC/water separations.

A convenient way to understand pervaporation processes is to assume that the process is divided into two sequential steps. The first is evaporation from the feed liquid to a saturated vapor phase; the second is permeation through the membrane under die driving force of the vapor pressure difference across the membrane. The saturated feed vapor phase is a thermodynamic tool; no such phase actually exists. Nevertheless, this approach is thermodynamically equivalent to conventional treatments, and allows the driving force for permeation to be expressed as a vapor pressure difference across the membrane. The feed liquid is in equilibrium widi the hypothetical feed vapor which, in turn, is in equilibrium with the feed side of the membrane. According to this definition of pervaporation, the overall separation achieved by die process, β -.__-, is given by r pβtVtp r βV-p r BHDI where β„_ is the evaporative separation term, which can be obtained from published vapor-liquid equilibrium data, and β,.,-, is die membrane permeation term derived from the standard solution-diffusion model for gas separation using Fick's Law. If the separation is influenced by the boundary layer, this expression can be written: r pemp ~~ evtp ' bHinαn The relative volatility of a set of components such as VOC/water is defined as: p'_we /p'_valer Relative Volatility = — — — O c voc /c' water where p_TO and p^ are the partial vapor pressures of the components of the feed, and c'^ and c'„__t are the feed liquid concentrations. From the Henry's Law relationships:

P^' _voc ⁼ H^ . c'_n. and p^ - E_wtl . c'^, the relative volatility is also given by:

Relative volatility = H^/H^,, Relative volatility is a convenient approximation for the β„,_ term in the separation factor equation above, because the quantities required for the relative volatility calculation are readily available. Relative volatility and β^,- are identical if the following equations for the partial vapor pressures of the components of the feed liquid (indicated as ' quantities) are valid:

P voc ⁼ P voc ^{- C} voe ' ^C _ιol ^d p -^ = p »„,„ ^• C __„_

In separations of interest in the context of the invention, the relative volatility of die components to be separated is high, so the β term is large and the overall separation factor β .„„_- is large also. By high relative volatility, we mean having a relative volatility greater than about 300, more particularly greater than about 1,000 and especially above 3,000.

We have discovered that, in separations characterized by high relative volatility and high separation factor, reducing the transmembrane driving force, by raising the permeate pressure and/or lowering the feed temperature, actually increases the separation factor. This is an unexpected result, because reducing the driving force lowers both the partial pressure differences and die pressure ratio across the membrane and, conventionally, this would be expected to have a negative effect on both flux and separation performance.

According to Fick's Law, lowering die component partial pressure difference across die membrane should lead to a directly proportional decrease in flux of that component.

As described in U.S. Patent 5,089,033, Figure 1 and columns 11 and 12, for example, there is a relationship between membrane separation α and pressure ratio φ (total feed pressure/total permeate pressure), such that at low pressure ratios, that is when φ « α, the enrichment obtained in a separation is pressure ratio-limited and is essentially independent of die membrane separation capability. Conversely, at high pressure ratios, that is when φ » α, the enrichment obtained is membrane separation-limited and is essentially independent of the pressure ratio. Thus, to take advantage of high membrane separation capability, it is normal to operate at the highest convenient pressure ratios. Yet we have discovered that in some cases, even though the separation factor in pervaporation processes can be very high, such as in the hundreds or thousands, die separation performance is not compromised, but is actually improved, when the pressure ratio is small compared widi the separation factor.

Our strange results appear to arise in situations where concentration polarization, discussed qualitatively in the Background section, is a serious problem and where die relative volatility of die components to be separated is high. We believe that these unusual effects are seen because reducing the driving force reduces d e total permeate flax through the membrane, which in turn alleviates concentration polarization.

Without wishing to be bound by theory, we suggest the following explanation. The equation commonly used to describe the effect of the boundary layer in membrane processes is: 1 1 1

where k is a mass transfer coefficient and the subscripts ov, m and bl refer to overall, membrane and boundary layer. The concentration polarization equation, derived for ultrafiltration and reverse osmosis, but equally valid for pervaporation is:

^Cm ~ ^Cp fv Λ, l = exp v ' c - c

' (2) where c is concentration, the subscripts b, m and p refer to the bulk feed, the membrane feed surface and die permeate, and v_p is the velocity perpendicular to d e membrane surface generated in die boundary layer by the permeate flow. Combining these two basic equations, we obtained the expression: ___ __ ___ , __ « exp^(v'/^{ku )} - I ^ko* ^k _m ^{E V} _P (3) where E = c_C_b is the enrichment achieved in the pervaporation process. The velocity v_p in cm/s is related to the total permeate flux, J_M in mole/cm².s by: v_p = J^p' where p ' is the density in mole/cm³ of the feed solution, so that Equation 3 can also be written as:

___ __ ___ ; L₎ ^eχp^(J"' ^{P ' kh, )} - 1

*.v ^E J_l0/P > (4) The boundary layer resistance term in Equation 4 is, through the term 1-1/E, a function of the separation achieved. If no separation is achieved, E=l and (1-1/E)=0. If E>1, in other words that component is enriched in the permeate, then (1-1/E) is positive and permeation is slowed down by die boundary layer term. Conversely, if E<1, in other words that component is depleted in the permeate, then (1-1 E) is negative and permeation is enhanced by die boundary layer term. Also, it can be seen that the boundary layer resistance term increases with increasing J,-..

Experimentally, we have found that, by increasing the membrane thickness, it is possible to actually reduce the boundary layer resistance. This is a different conclusion from that obtained from die basic resistance model equation. According to this equation, increasing the diickness of die membrane increases the relative contribution of the membrane resistance term but has no absolute effect on the boundary layer term. We have also found that, for enriching separations such as those of interest in the context of the invention, reducing J^ reduces d e slowing effect of the boundary layer term.

Reducing J_ω by reducing die pressure difference across the membrane can be achieved by raising the permeate pressure, by lowering the feed temperature (and hence the feed side vapor pressure) or by a mix of the two.

Our preferred method is to raise the permeate pressure, that is to operate the pervaporation process at a permeate pressure higher than was previously normal practice for high volatility, high separation factor separations. In particular, d e invention involves carrying out the pervaporation process at a permeate pressure greater than about 50 torr (6.7 kPa), and preferably even higher, such as greater than about 60 torr (8 kPa), 80 torr (10.7 kPa), 100 torr (13.3 kPa), 120 torr (16 kPa), or even 150 torr (20 kPa). Furthermore, it is preferred that these pressures be generated, at least substantially, by condensing die permeate at a suitable temperature, and using a vacuum pump only as necessary to remove non-condensed gases. The temperature to which the condenser must be chilled to achieve permeate pressures of this order depends, of course, on the vapor pressures of die permeate components, but will typically be no lower than about 5 °C. It is preferred to operate the condenser at temperatures no lower than about 10°C, more preferably no lower than about 15 °C and most preferably no lower than about 20 °C. A particular advantage of operating in this mode is the low energy requirement compared with conventional pervaporation. For example, if a condenser temperature of 20°C or above is used, this can normally be achieved by means of an air-cooled condenser. If a condenser temperature of 10-20°C is used, this can normally be achieved by means of cooling with available water.

As an alternative or supplementary method to reduce the driving force, the process of the invention can be operated under mild feed temperature conditions. It is not unusual for recommended feed temperatures for pervaporation to be as high as 60 °C, 80 °C or even higher. The preferred feed temperature for operating the processes of the invention is below 60°C, more preferably below 50°C and most preferably below 40°C. Of course, die permeate pressure and die feed temperature should be selected based on the vapor pressures of the components of interest to give an appropriate separation performance taking into account any other relevant considerations, such as energy consumption or costs. In many cases, however, the processes of the invention are characterized by a low pressure ratio, where the pressure ratio is defined as the ratio of the total vapor pressure on the feed side to the total pressure on die permeate side. Processes of the invention may, for example, be characterized by pressure ratios as low as 2 or even lower, such as 1.8, 1.6 or even lower.

Another aspect of the invention is that it provides good opportunities for heat integration. For example, if the incoming feed liquid is cool, the incoming feed stream may be warmed and die permeate vapor cooled by flowing the streams against one another in a heat exchanger. Since only moderate heating and cooling is needed to carry out the processes of the invention, it may be possible to provide a substantial portion of die driving force this way, thereby reducing expenditure on external heating or cooling. As an alternative to direct heat exchange, a heat pump or other indirect heat-exchanging mechanism can be used. As just one specific example, the feed stream can be used to cool an electrical chiller, which in turn provides die coolant for permeate condensation. The feed stream is then fed into the pervaporation system via a heat recovery heat exchanger and a heater.

The membranes used in the invention may take the form of a homogeneous membrane, an asymmetric membrane, a multilayer composite membrane, a matrix incorporating a gel or liquid layer, or any other fcrm known in the art. Airy convenient membrane material may be used. For separating organic compounds from water, rubbery permselective layers are preferred. Suitable rubbery materials are discussed, for example, in "Selection of Elastomeric Membranes for the Removal of Volatile Organic Components from Water" (Reference 7). For organic compound/water separations, we prefer membranes made from silicone rubber or ethylene-propylene copolymers. For other separations, rubbery or glassy membranes as appropriate to the separation may be used.

The membranes can be incorporated into membrane modules of any convenient type, such as spiral-wound, potted hollow-fiber or plate-and-frame. To carry out the process of the invention the feedstream is introduced into an array of one or more membrane modules and flows across the feed surface of the membrane. The non-permeating portion of the feedstream is removed as a liquid residue stream, which is depleted in the faster-permeating component(s). Permeate vapor, enriched in the faster- permeating components), is withdrawn from die permeate side of die membrane. Depending on die mutual solubilities of d e components to be separated, it is frequently the case that the permeate will form a two-phase mixture after condensation. In the case of separation of VOCs from water, for example, an organic phase and an aqueous phase saturated with die organic will typically form. If the organic phase is the desired product, this may be decanted off and die aqueous phase may conveniently be recycled to the pervaporation process, as described in U.S. Patent 5,030,356.

Preferably, the pervaporation process should remove at least 80% of the faster permeating components), more preferably at least 90%, and most preferably 95% or more. The array of membrane modules used to carry out the process may form a single-stage unit, a two- or multi-stage unit, in which the permeate from one stage becomes the feed to the next, a two- or multi- step unit, in which the residue from one step becomes the feed to the next, or combinations thereof.

The invention may be applied to any solution containing components characterized by a high relative volatility. Examples include separation of low molecular weight, high volatility organics from high molecular weight, lower volatility organics, dehydration of hydrophobic or low volatility organics, and removal of dissolved gases, particularly air or oxygen, from water or other liquids. We particularly intend die invention to apply to separations where an organic component is preferentially removed either from water or from another organic component. It is believed that the invention will be found to be especially useful in the removal of VOCs from water, such as as is needed in surface- or groundwater remediation or industrial wastewater treatment, for example.

Representative organic materials that may be separated from water or each other by die process of the invention include, but are not limited to, aliphatic hydrocarbons, such as such as hexane, octane or decane; aromatic hydrocarbons, such as benzene, toluene and xylene; halogenated hydrocarbons, such as perchloroethylene, trichloroethylene, trichloroethane or chlorinated fluorocarbons; esters, such as ethyl acetate or butyl acetate; ketones, such as methyl ethyl ketone; alcohols, such as butanol, hexanol or octanol; naphthas; terpenes; and die like.

If the separation that is to be carried out is the removal of a VOC from water, the relative volatility may be expressed by means of the Henry's Law coefficient. The process of the invention is particularly useful for removing organic components having a Henry's Law coefficient at 25 °C greater than about 2 x 10"" atm.m³/gmol, especially those having a Henry's Law coefficient at 25 °C greater than about 6 x 10"⁴ atπunVgmol, more especially those having a Henry's Law coefficient at 25 °C greater than about 2 x 10³ atm.m³/gmol. The processes of the invention can advantageously be carried out in conjunction with die pervaporation processes described in a co-owned and copending application entided "High Flux Pervaporation Process" (serial number not yet assigned), the specification of which is incorporated herein by reference in its entirety. In that case, the inventions together will provide processes that reduce concentration polarization effects both by working in an optimum membrane thickness range and by using modest transmembrane driving forces. In this case, it is preferred that the membrane thickness be selected to be of such a diickness to yield a transmembrane flux of die more volatile component that is at least 70% of the maximum flux of that component, more preferably at least 80% and most preferably at least 90% of the maximum flux of that component.

The invention is now further illustrated by die following examples, which are intended to show certain aspects of the invention, but are not intended to limit the scope or underlying principles of the invention in any way.

EXAMPLES

EXAMPLE 1: RELATIVE VOLATILITY CALCULATIONS

We calculated die relative volatilities of eight different VOCs representing four classes of organic compounds: aromatic hydrocarbons, chl<_rinated hydrocarbons, esters and ketones. The relative volatilities were calculated from die pure VOC vapor pressure at 40 °C and d e solubUity in water according to the equation: relative volatility = p^ / (p⁰.^ • _„,,) where p^ and p⁰,-*,- are the pure VOC and pure water vapor pressures and c^ is the solubility limit of the VOC in water in mole fraction. The results are given in Table 2.

Table 2

VOC Solubility in wafer Pure vapor pressure 40*C Relative (cfflHg) <kFa) volatility

Toluene 0.07 5.9 (7.9) 8,150

Trichloroethylene 0.11 13.8 (18.4) 16,600

1 , 1 ,2-trichloroethane 0.48 4.8 (6.4) 1,320

Butyl acetate 0.57 2.4 (3.2) 490

Methylene chloride 1.32 76.6 (102.1) 4,950

Ethyl acetate 8.0 18.8 (25.1) 210

Methylethyl-ketone 25 17.7 (23.6) 51

Acetone 100 56.6 (75.5) 10 EXAMPLES 2-10: EFFECT OF FEED TEMPERATURE

EXAMPLE 2

A set of pervaporation experiments to measure die removal of toluene from water was carried out. A spiral-wound membrane πrødde containing 1.1 m² of composite membrane with a 10 μm silicone rubber selective layer was used The feed temperature was 60 °C and die permeate pressure, provided by means of a condenser and a liquid ring vacuum pump, was 50 torr (6.7 kPa). The feed flow rate was 3 gpm (11.417-nin). The experitnαit was repeated for various feed concentrations. The results of the experiment are summarized in Table 3.

Table 3

The separation factors were reasonably constant in the experiment, even though die feed concentration varied by 270-fold.

EXAMPLE 3

Pervaporation experiments were carried out as in Example 2, but this time at a feed temperature of 45 °C. The results are given in Table 4.

Table 4

The separation factor was around 1,500 in both cases. Comparing the data in Tables 3 and 4, it may be seen that the separation factor appears to be improved by a factor of 2-3 by operating at a lower feed temperature. Table 5 repeats the data for a feed concentration of 270 ppm at 60°C and at 45 °C. Table 5

EXAMPLE 4

The experiments of Example 2 to measure the removal of toluene from water were repeated at a flow rate of 5.5 gpm (20.8 IJmin). The feed temperature was 60°C and the permeate pressure was 50 torr (6.7 kPa). The results are summarized in Table 6.

Table 6

The separation factor varied in the range 745-1,280.

EXAMPLE 5

Pervaporation experiments were carried out as in Example 4, but this time at a feed temperature of 45 °C. The results are given in Table 7.

Table 7

The separation factors were again more than doubled by operating at a lower feed temperature. Table 8 repeats the data for a feed concentration of 235 ppm at 60 °C and at 45 °C.

EXAMPLE 6

The experiments of Example 2 to measure die removal of toluene from water were repeated at a flow rate of 8 gpm (30.3 IJmin). The feed temperature was 60°C and die permeate pressure was 50 torr (6.7 kPa). The results are summarized in Table 9.

The separation factor varied in the range 1,320-1,820.

EXAMPLE 7

Pervaporation experiments were carried out as in Example 6, but this time at a feed temperature of 45 °C. The results are given in Table 10.

Table 10

The separation factors were about doubled by operating at a lower feed temperature. Table 11 repeats the data for a feed concentration of 200 ppm at 60°C and at 45°C. Table 11

Concentration (ppm) Permeate flux Stage cut (%) Separation (Feed 200 ppm) (kg/m².h) factor

Residue Permeate (%)

Temp. (°C) (ppm)

45 160 32 0.3 0.02 2,500

60 107 16 0.8 0.05 1,323

EXAMPLE 8

Additional experiments of the type described in Examples 2-7 were carried out. This time the data were plotted graphically. The results are shown in Figure 2. As can be seen, substantially higher separation factors were always obtained at the lower feed temperature.

EXAMPLE 9

A set of pe-vaporation experiments to measure the removal of toluene from a solution of 600 ppm toluene in water was carried out. In this case, four spiral-wound membrane modules with a total membrane area of 4 m² of 20 μm silicone rubber membrane were used. The permeate pressure was 100 ton (13.3 kPa) and die feed temperature was varied from 60-70°C. Representative results of the experiment are summarized in Table 12 and plotted graphically in Figure 3.

Table 12

Concentration (ppm) Permeate flux Separation (kg/m².h) factor

Temp. (°C) Residue (ppm) Permeate (%)

60 13 9.6 0.2 780

65 6.0 5.6 0.3 450

70 5.0 3.6 0.5 300

It can be seen that, over a 10°C temperature change, the separation factor increases over 2.5 fold.

EXAMPLE 10

A set of pervaporation experiments to measure the removal of trichlororethylene (TCE) from solutions containing about 200 ppm TCE in water was carried out, using a spiral-wound membrane module containing 0.2 m² of composite membrane with a 3.5 μm silicone rubber selective layer. The permeate pressure was maintained at 10 ton (1.3 kPa) using a liquid nitrogen trap to remove the VOC and a small vacuum pump to remove non-condensable gases. The parameters varied in the experiment were feed temperature and feed flow rate. Evaluation of the membrane performance was standardized by comparing the stage cut required to reduce the TCE feed concentration by a certain reduction factor. The results for a feed flow rate of 1.8 gpm (6.8 IJmin) are given in Figure 4. As can be seen, the lower the feed temperature, the bettor is die TCE removal. The results are replotted in Figure 5 in terms of the separation factor obtained at different temperatures, using die lowest stage cut data.

EXAMPLES 11-15: EFFECT OF PERMEATE PRESSURE EXAMPLE 11

Experiments of the type described in Example 2 for toluene/water separations were performed, using a feed flowrate of 1 gpm (3.8 IJmin), and this time fixing the feed temperature at 50°C, but varying the permeate pressure. The stage cut varied from 0.1-0.3%. The results are summarized in Table 13.

Table 13

Permeate Concentration Permeate flux Separation pressure (kg/m².h) factor

(ton) (kPa) Feed (ppm) Residue(ppm) Permeate (%)

40 (5.3) 0.7 0.2 0.02 0.6 410

60 (8) 0.6 0.2 0.02 0.5 500

60 (8) 123 38 6.0 0.3 830

80 (10.7) 0.3 0.1 0.01 0.3 800

80 (10.7) 62 20 4.3 0.2 1,200

The separation factor varied in the range 410-1,200. The ratio between the worst separation factor at 40 ton (5.3 kPa) and die best at 80 ton (10.7 kPa) is about 3.

EXAMPLE 12

Toluenewater experiments of the type described in Example 11 were repeated at a feed flow rate of 2 gpm (7.6 IJmin). The results are summarized in Table 14. Table 14

Permeate Concentration Permeate flux Separation pressure (kg m².h) factor

(ton) (kPa) Feed (ppm) Residue (ppm) Permeate (%)

40 (5.3) 0.3 0.1 0.02 0.4 1,190

60 (8) 0.2 0.1 0.02 0.3 1,490

60 (8) 82 31 7.0 0.3 1,440

80 (10.7) 0.2 0.1 0.06 0.2 4,030

80 (10.7) 33 13 4.6 0.2 2,200

EXAMPLE 13

Toluene water experiments of the type described in Example 11 were repeated at a feed flow rate of 3 gpm (11.4 IJmin). The results are summarized in Table 15.

Table 15

Permeate Concentration Permeate Separation pressure flux factor

(ton) (kPa) Feed Residue Permeate (kg/m².h) (ppm) (ppm) (%)

40 (5.3) 0.4 0.2 0.04 0.4 1,400

40 (5.3) 72 30 3.7 0.4 800

60 (8) 0.4 0.2 0.05 0.3 1,700

60 (8) 74 311 6.5 0.4 1,400

80 (10.7) 0.3 0.2 0.06 0.2 2,420

The separation factor varied from a low of 800 to a high of 2,420 at 80 ton (10.7 kPa).

EXAMPLE 14

Additional experiments of the type described in Examples 11-13 were carried out. This time the data were plotted graphically. The results are shown in Figure 6. As can be seen, substantially higher separation factors were always obtained at die higher permeate pressure. EXAMPLE 15

One of the experiments of Example 9 with a 600 ppm toluene feed was repeated using a feed temperature of 60°C and a permeate pressure of 150 ton (20 kPa). The comparison with die previous measurement at 100 ton (13.3 kPa) is listed in Table 16. Table 16

Feed temp. (°C) / Concentration (ppm) Permeate flux Separation

Permeate pressure (kg/m².h) factor

(ton) (kPa) Residue (ppm) Permeate (%)

60 /100 (13.3) 13 9.6 0.2 690

60 / 150 (20) 24 60 0.02 8,320

The process worked, although the very high value for the separation factor at 150 ton (20 kPa) is suspect.

EXAMPLE 16 AND 17: EFFECT OF DRIVING FORCE AND PRESSURE RATIO EXAMPLE 16

The previous examples show the beneficial effect of reducing driving force either by reducing feed temperature at constant permeate pressure or by raising permeate pressure at constant feed temperature.

In this example, two experiments with different feed temperatures and permeate pressures are compared Both experiments measured d e removal of trichlororethylene (TCE) from solutions containing about 200 ppm TCE in water, using a spiral-wound membrane module containing 0.2 m² of composite membrane with a 2 μm ethylene-propylene copolymer selective layer. The feed flow rate was 1.8 gpm (6.8 IJmin) in each case. Evaluation of die membrane performance was standardized by comparing the stage cut required to reduce the TCE feed ccncentraticn by a certain reduction factor. In one case, the feed temperature was 40°C and die permeate pressure was 10 ton (1.3 kPa); in the other, the feed temperature was 30°C and the permeate pressure was 20 ton (2.7 kPa). The results are plotted in Figure 7. As can be seen, the reduced driving force provided by bot raising the permeate pressure and lowering the feed temperature results in better TCE removal. The results are replotted in Figure 8 in terms of the separation factor obtained at different pressure ratios, using the lowest stage cut data. A feed temperature of 40 °C and a permeate pressure of 10 ton (1.3 kPa) provides a feed/permeate pressure ratio of 6; a feed temperature of 30°C and a permeate pressure of 20 ton (2.7 kPa) provides a feed/permeate pressure ratio of 2. The separation factor increases almost fourfold as the pressure ratio drops. EXAMPLE 17

A series of experiments was performed to measure the removal of benzene from water under different driving forces. A bench-scale pervaporation unit containing a single membrane module with a membrane area of 0.18 m² was used to perform the tests. The system could treat about 2-3 gallons

(7.6-11.4 L) of water over a period of 2-3 hours. A small pump was used to circulate feed solution between a feed tank and the test module. A vacuum pump and a dual permeate condenser system provided the appropriate permeate pressure. The permeate stream flow was switched from one condenser to the other from time to time, allowing sampling of the condensed permeate liquid without interrupting operation of d e unit.

The experiments were performed at variable permeate pressures, but at a fixed feed solution temperature of 55°C. A 50-100 ppm benzene solution was used as the test mixture. At 55°C, this solution has an average vapor pressure of 118 ton (15.7 kPa).

The effect of changes in the permeate pressure, and hence driving force, is shown in Figures Ha¬ nd, in which the feed and permeate benzene concentrations are plotted as a function of stage cut. The target in all cases was 95% benzene removal from the feed, so that the feed solution concentration is reduced to 1-5 ppm. As the permeate pressure rises, or the driving force falls, the stage cut required to meet this target falls. The reduced stage cut at high permeate pressures is extremely advantageous, since it reduces the heat load on die permeate condenser. Also die overall enrichment obtained increases from approximately 250 fold at 10 ton (1.3 kPa) to more man 1,000 fold at 95 ton (12.7 kPa). The feed/permeate pressure ratio is 11.8 at 10 ton (1.3 kPa), 3.4 at 35 ton (4.7 kPa), 1.7 at 70 ton (9.3 kPa), and 1.2 at 95 ton (12.7 kPa).

EXAMPLES 18-21: COMPARATIVE CALCULATIONS EXAMPLE 18

A set of calculations was performed using an in-house computer modeling program that simulates the performance of the pervaporation process shown in Figure 8. Referring now to this figure, feedstream, 1, comprising two components to be separated, is brought into contact with membrane unit, 2. The non- permeating portion of the feedstream is removed as a liquid residue stream, 3. The permeate vapor is cooled in condenser, 5. The permeate vapor stream, 4, should contain the components in proportions such that, after condensation, phase separation takes place. The non-condensed fraction, 6, of the permeate vapor, including r-cn-ccndensable gases, is removed by a small vacuum pump, 7. The condensed permeate liquid, 8, passes to a decanter, 9. The product phase, 11, is withdrawn from die decanter. The other phase, 10, is mixed with die incoming feedstream to the system and reprocessed through the pervaporation unit.

The calculations were performed using TCE experimental data for total permeate flux and separation factor at 50 °C. The experimental data had been gathered at different permeate pressures from experiments such as those reported in Examples 11-15. The data were then used to calculate the attributes of a pervaporation system required to reduce die concentration of a TCE-contaminated water stream from 100 ppm to 1 ppm, that is 99% removal of TCE. The feed flow rate was assumed to be 10 gpm (37.9 IJmin). The calculation was repeated for carried out for three permeate pressures, 40, 60 and 80 ton (5.3, 8, and 10.7 kPa). The results are shown in Table 17.

Tabl e l7

Permeate Permeate flux Separation Permeate Membrane Condenser pressure (kg/m².h) factor concentration area (m²⁾ load

(ton) (kPa) (%) (10 Btu/h)

40 (5.3) 0.42 1,140 0.36 21 2.0

60 (8) 0.33 1,450 0.45 21 1.5

80 (10.7) 0.19 2,120 0.66 25 1.0

As can be seen from the table, operating at permeate pressure of 80 ton (10.7 kPa) rather than

40 ton (5.3 kPa) almost doubles die permeate concentration and halves the permeate volume, thereby reducing die load on die condenser to 50% of its value at 40 ton (5.3 kPa). The permeate, stream 8 in Figure 8, is sufficiently concentrated to phase separate, producing a pure TCE product phase and an aqueous phase that is recycled. Although more membrane area is required at lower driving force, the increase in membrane area requirement is modest. EXAMPLE 19

The calculations of Example 18 were repeated using experimental flux and separation factor data obtained for toluene at feed temperatures of 45 ° C and 60 ° C. The feed flow rate was once again assumed to be 10 gpm (37.9 IJmin), the feed concentration 100 ppm and die residue concentration 1 ppb. The permeate pressure was assumed to be 40 ton (5.3 kPa). The results of the calculations are shown in Table 18.

Table 18

The results again show a two-phase permeate. The permeate concentration is nearly doubled and die volume is more than halved by lowering the feed temperature 15 °C. The condenser load is again halved.

EXAMPLE 20

The calculations of Example 18 for TCE woe repeated assuming a feed concentration of 10 ppm and a target residue concentration of 0.1 ppm, that is 99% removal. The feed flow rate was again assumed to be 10 gpm (37.9 IJmin) and the feed temperature 50°C. The results of the calculations are shown in Table 19.

Table 19

Permeate Permeate flux Separation Permeate Membrane Condenser pressure (kg/m².h) factor concentration area (m²⁾ load

(ton) (kPa) (%) (10⁴Btu/h)

40 (5.3) 0.42 1,140 0.34 23 2.1

60 (8) 0.33 1,450 0.41 23 1.6

80 (10.7) 0.19 2,120 0.55 27 1.1

Again, the substantial benefits of operation at relatively high permeate presure are clear.

EXAMPLE 21

The calculations of Example 20 were repeated, die only difference being that the feed concentration was assumed to be 1 ppm and die residue 0.01 ppm. The results are shown in Table 20.

Table 20

Permeate Permeate flux Separation Permeate Membrane Condenser pressure (kg/m2.h) factor concentration area (m²⁾ load

(ton) (kPa) (%) (10⁴Btu/h)

40 (5.3) 0.42 1,140 0.13 30 2.8

60 (8) 0.33 1,450 0.13 29 2.1

80 (10.7) 0.19 2,120 0.15 32 1.3

Once again, the results strongly favor operation at high permeate pressure.

EXAMPLE 22: REMOVAL OF BTEX AROMATICS

A series of tests was performed with a produced water sample from an oilfield. Benzene, toluene and xylenes were the major VOC contaminants, and die concentration of these components was summed and monitored during die course of die experiment. A bench-scale pervaporation unit containing a single membrane module with a membrane area of 0.18 m² was used to perform die tests. The system could treat about 2-3gallons (7.6-11.4 L) of water over a period of 2-3 hours. A small pump was used to circulate feed solution between a feed tank and die test module. A vacuum pump and a dual permeate condenser system provided die appropriate permeate pressure. The permeate stream flow was switched from one condenser to die other from time to time, allowing sampling of the condensed permeate liquid without interrupting operation of die unit.

Two sets of experiments were performed, one at a permeate pressure of 10 ton (1.3 kPa), the other at 70 ton (9.3 kPa), both at 55 °C feed temperature. The feed, residue and permeate stream samples were analyzed by gas chromatography. Figure 9 shows a representative set of traces. Methanol was added to the permeate sample to ensure complete dissolution of die BTEX compounds.

Figure 10 shows the concentration of the permeate solution as a function of feed concentration. The average enrichment, measured as the slope of the graph, obtained at a permeate pressure of 10 ton (1.3 kPa) was 250. When the pressure was raised to 70 ton (9.3 kPa), the enrichment increased to about 600. This significant improvement in separation did not decrease die transmembrane BTEX flux. As Figure 11 shows, the BTEX flux decreases with decreasing feed solution concentration. At the same feed concentration, however, the BTEX fluxes at 10 and 70 ton (1.3 and 9.3 kPa) were the same.

REFERENCES

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Claims

We claim:

1. A pervaporation process for separating a first organic component of a solution from a second component, said first organic component having a relative volatility compared widi said second component of at least about 300, said process comprising operating at a permeate pressure of at least about 50 ton and withdrawing a permeate stream enriched in said first organic component compared with said solution.

2. The process of claim 1, wherein the ratio of vapor pressure on die feed side to said permeate pressure is no greater than about 2.

3. The process of claim 1, wherein said permeate pressure is at least about 80 ton (10.7 kPa).

4. The process of claim 1, wherein said permeate pressure is at least about 100 ton (13.3 kPa).

5. The process of claim 1, wherein said permeate pressure is at least about 150 ton (20 kPa).

6. A pervaporation process for separating a first organic component of a solution from a second component, said first organic component having a relative volatility compared widi said second component of at least about 300, said process comprising operating at such a permeate pressure and feed temperature that the ratio of vapor pressure on the feed side to said permeate pressure is no greater than about 2, and withdrawing a permeate stream enriched in said first organic component compared with said solution.

7. The process of claim 6, wherein said ratio is no greater than about 1.8.

8. The process of claim 6, wherein said ratio is no greater than about 1.6.

9. The process of claim 1 or claim 6, further comprising operating at a feed temperature no greater than about 60 °C.

10. The process of claim 1 or claim 6, wherein said relative volatility is at least about 1,000.

11. The process of claim 1 or claim 6, wherein said relative volatility is at least about 3,000.

12. The process of claim 1 or claim 6, wherein said feed temperature is no greater than about 50°C.

13. The process of claim 1 or claim 6, wherein said feed temperature is no greater than about 40°C.

14. The process of claim 1 or claim 6, wherein at least one of said components is an aliphatic hydrocarbon.

15. The process of claim 1 or claim 6, wherein at least one of said components is an aromatic hydrocarbon.

16. The process of claim 1 or claim 6, wherein at least one of said components is a halogenated hydrocarbon.

17. The process of claim 1 or claim 6, wherein at least one of said components is an ester.

18. The process of claim 1 or claim 6, wherein said components are an organic component and water, and said organic component has a Henry's Law coefficient at 25 °C of at least about 2 x 10^"4 atm.m³/gmol.

19. The process of claim 1 or claim 6, wherein said components are an organic component and water, and said organic component has a Henry's Law coefficient at 25 ° C of at least about 2 x 10³ atm.mVgmol.

20. The process of claim 1 or claim 6, wherein a silicone rubber membrane is used to separate said components.

21. The process of claim 1 or claim 6, wherein an ethylene-propylene copolymer membrane is used to separate said components.

22. The process of claim 1 or claim 6, wherein at least about 90% of said first organic component is removed from said solution.

23. The process of claim 1 or claim 6, wherein said second component is also an organic component.

24. The process of claim 1 or claim 6, wherein said permeate pressure is provided at least in part by condensing at least a portion of said permeate stream, said condensing being carried out at a temperature of at least about 15 °C.

25. The process of claim 1 or claim 6, wherein said solution is heated and said permeate stream is cooled by running said solution and said permeate stream in heat-exchanging relationship against one another.