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Low-Energy Distillation-Membrane Separation Process Yu Huang,*,† Richard W. Baker,† and Leland M. Vane‡ Membrane Technology and Research, Inc., 1360 Willow Road, Suite 103, Menlo Park, California 94025, and EPA-NRMRL, 26 West Martin Luther King DriVe, Cincinnati, Ohio 45268
A low-energy separation process combining distillation and membrane vapor permeation is introduced as an alternative to conventional distillation. The process can be applied to any liquid mixture for which appropriate selective membranes are available. However, in this paper, the discussion is limited to water/organic solvent mixtures using membranes that preferentially permeate water. Such membranes are available. The process is illustrated with two mixtures: ethanol (light component)/water (heavy component) and acetic acid (heavy component)/water (light component) mixtures. In both cases, the combination process reduces the energy consumption of the separation to half that of simple distillation. 1. Introduction There are approximately 40 000 distillation columns of various types in operation in the United States. These columns use about 5 quads (1 quad ≡ 1015 Btu) of energy.1 Distillation is widely used because of its simplicity, low capital cost, and reliability. The key disadvantage of distillation is its low energy efficiency. In the past, the energy cost of distillation was not a serious issue; this is no longer the case. Economical methods of reducing the energy costs of distillation operations are being aggressively pursued. The use of vapor permeation/pervaporation membranes as a low-energy alternative to distillation has been proposed for more than 30 years, yet the current market for this technology is not more than $20 million/year, almost all for the separation of water from ethanol or isopropanol in the pharmaceutical and fine chemicals industry.2 In the early 1990s, major oil companies such as Exxon, Texaco, and Mobil all had research programs focused on developing membrane technology for refinery separations.3-7 Most of these programs have since been abandoned or scaled back. The problem was not the lack of suitably selective membrane materials, but the difficulty of making reliable and economical membranes and membrane modules. Membrane and module components able to operate at temperatures above 100 °C in hydrocarbon liquids were required. Also, early developers of the technology often linked the membrane systems and distillation as a simple series of unit operations. This overlooked the very substantial reductions in the energy of separation achieved when heat integrated process designs are used. Significant progress has been made in solving all of these problems in the last 10 years, especially the last 3-4 years. In this paper, a low-energy separation process combining distillation and membrane vapor permeation is introduced as an alternative to conventional distillation. The membranes and membrane modules used in this new process can be operated at temperatures up to 130 °C and are briefly described. This work was performed at the U.S. Environmental Protection Agency (EPA) Cincinnati Laboratories and at Membrane Research and Technology (MTR) under an EPA/MTR CRADA agreement. This new process is well suited for dehydration of bioethanol and biobutanol to be produced in the next generation of cellulose-to-biofuels plants. More than 500 of these plants * To whom correspondence should be addressed. E-mail: ihuang@ mtrinc.com. † MTR-Membrane Technology and Research, Inc. ‡ EPA-NRMRL, E-mail:
[email protected].
will be needed if the U.S. Department of Energy’s biofuels program is to meet its 2022 targets.8 1.1. Mechanical Vapor Compression. Mechanical vapor recompression is a known method of reducing the energy cost of distillation. The process is illustrated in Figure 1. The overhead vapor from the column is compressed and then condensed in the reboiler at the bottom of the column. This recovers all the latent heat of condensation otherwise lost in the overhead condenser of the column. A fraction of the condensed vapor is sent to the top of the column to become the liquid reflux. The rest is discharged as final product. The energy used by the vapor compressor is only 20% of the energy recovered in the reboiler, so large reductions in total energy consumption are possible. Even though mechanical vapor compression has been known for many years, it was not widely used in the past because the value of the energy saved did not offset the high capital cost of the compressor.9 1.2. Distillation-Vapor Compression-Membrane Separation. We are developing a different approach to mechanical vapor compression. The process combining distillation, vapor compression, and membrane separation is shown in a simplified
Figure 1. Block diagram of a mechanical vapor compression system. After compression, the overhead vapor is condensed in the distillation column reboiler to recover the vapor’s latent heat of condensation.
10.1021/ie901545r 2010 American Chemical Society Published on Web 03/01/2010
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Figure 3. Pervaporation separation of water/ethanol and water/acetic acid mixtures with composite perfluoropolymer membranes. The vapor/liquid equilibrium (VLE) lines are shown for comparison. Figure 2. Block diagram of a distillation-membrane separation hybrid system. In this example, a membrane separation unit is used to separate the compressed overhead vapor from an ethanol (EtOH)/water (H2O) mixture into an ethanol-enriched stream and a water-enriched stream. The waterenriched vapor stream is returned directly to the column, retaining all of its latent heat. The ethanol-enriched stream is condensed in the reboiler, recovering its latent heat.
form in Figure 2. A membrane vapor permeation unit is used to separate the compressed overhead vapor from the distillation column into two streams. In the example shown (the separation of water from ethanol), the permeate vapor enriched in water is returned directly to the distillation column. This recovers all the latent heat of this permeate vapor stream. The residue vapor, depleted in water, is sent to the reboiler heat exchanger where the vapor condenses and its latent heat content is also recovered. The condensed liquid is then split between the product ethanol stream and the liquid reflux stream. As in traditional mechanical vapor compression, the energy consumption of the process is much lower than that of conventional distillation; in addition, a “free” separation step by the membrane unit reduces the size of the column and further reduces energy consumption. Not only is less energy used, but a better separation is achieved. In the process shown in Figure 2, distillation alone cannot produce better than 93 wt % ethanol because of the ethanol-water azeotrope. The combination distillation-membrane process can produce >99 wt % ethanol and consumes at least 50% less energy. The distillation-membrane process illustrated in Figure 2 could, in principle, be applied to any liquid mixture for which appropriate membranes are available. In this paper, however, we will limit our discussion to organic solvent/water mixtures, using membranes that are preferentially permeable to water. This choice allows us to use membranes that are already available or are under development. 1.3. Membranes and Modules. In the process shown in Figure 2, the overhead vapor from the distillation column is compressed from 0.5 bar to 2-3 bar by a one- or two-stage compressor. The dew point of the overhead vapor at a pressure of 3 bar is about 120 °C when the vapor contains about 65 wt % ethanol. This high dew point means that in order to maintain the vapor comfortably above the dew point (to avoid condensation), it is essential to operate the membrane modules with superheated solvent/water vapors at temperatures up to 130 °C. Extensive literature on pervaporation/vapor permeation mem-
branes exists.10-12 Very few of these measurements have been made at temperatures above 60 °C, and only a handful of membranes have been formed into membrane modules. MTR/EPA have been working on ethanol/water separations for the past four years.12-17 Composite membranes have been made from a variety of water-permeable materials ranging from cross-linked hydrophilic polymers (polyvinyl alcohol and cellulose esters) to hydrophobic perfluoropolymers (Teflon AF and Hyflon AD). Producing membranes that are stable in hot ethanol/ water mixtures at temperatures up to 130 °C has been a challenge. These developments are reported in several of our pending patent applications and will be described in detail in a separate publication. Some data showing the membrane separation performance achieved with our membranes for two of the solvent/water mixtures described in this paper are shown in Figure 3. Currently, stable membrane modules (including membranes and other components) that can tolerate ethanol/ water mixtures at 130 °C have also been successfully developed. These modules have been operated under various conditions at both MTR and EPA. MTR’s first field test will come online in the spring of 2010. Membrane performance is quantified in terms of membrane permeance and selectivity using the solution-diffusion equation.18 ji )
Pi (p - piL) L io
(1)
where ji is the molar flux (cm3(STP)/cm2 · s · cm Hg), L is the membrane thickness, pio and FiL are the partial vapor pressures of component i on either side of the membrane, and Pi is the permeability of membrane material, usually written in barrers (1 × 10-10 cm3(STP) · cm/cm2 · s · cm Hg). Because the selective layer in composite membranes is difficult to measure, membrane permeation rates are usually reported as permeances (pressure-normalized fluxes) expressed as Pi ji ) L pio - piL
(2)
where the permeance Pi/L is expressed in gas permeation units or gpu (1 × 10-6 cm3(STP)/cm2 · s · cm Hg).
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Figure 4. Bench-scale test system used for long-term membrane module tests. The feed ethanol concentration is 5 wt %, the liquid feed flow rate to the column is 141 g/min, the feed liquid temperature is 55 °C, and the stripper column pressure is 0.2 bar.
The membrane separation performance is given as the ratio of the permeances of components i and j: Rij )
Pi Pi /L ) Pj /L Pj
(3)
where Rij is the selectivity of the membrane for component i over component j. Permeances are temperature and concentration dependent and can change significantly with the membrane material. In the design calculations that follow, we will use a water permeance of 2000 gpu, an ethanol permeance of 50 gpu, and an acetic acid permeance of 40 gpu. These values are fair representations of the average permeance values of the membranes/membrane modules we have made at MTR. 1.4. Bench-Scale Process Demonstration. Before embarking on a description of how the process will be used in large industrial plants, we will report some proof-of-concept results obtained at EPA with bench-scale membrane modules. The test system is shown in Figure 4. In these experiments, a 6-ft-tall, 3-in.-diameter stripper column was combined with a 2-in.-diameter, spiral-wound membrane module containing about 1.0 m2 of membrane. The overhead vapor was compressed from the column pressure of 0.2 to 1.0 bar using a claw-type vapor compressor. The compressed vapor was then sent to the membrane module. The module was contained in a temperature-controlled oven maintained at about 5-10 °C above the dew point of the incoming vapor, typically at 115-125 °C. During the experiments, three operating modes were used: (1) The membrane permeate and residue streams were condensed independently without any heat integration with the stripping column. (2) The membrane permeate vapor stream was recycled directly to the bottom of the stripper column while the residue stream was condensed separately with no heat integration. (3) The membrane permeate vapor stream was recycled directly to the bottom of the stripper column, and the high-pressure residue stream was condensed in the column reboiler, thus recovering the latent heat content of both membrane streams in the process. In all operating modes, the ethanol-depleted stripper column bottoms liquid and the ethanol-enriched condensed membrane
Figure 5. Demonstration of the energy savings achieved by recycling separated permeate and residue streams to the distillation process. The graph shows the fraction of the feed stream ethanol concentrated into the overhead vapor as a function of the steam energy used in the stripper reboiler.
residue streams were returned to the process feed tank so continuous runs lasting as long as 6-8 weeks were possible. The energy savings achieved by the different operating modes of the bench test system are shown in Figure 5. This figure plots the fraction of the ethanol contained in the liquid feed stream to the column that is stripped into the overhead vapor as a function of the energy supplied in the form of direct steam addition into the base of the column. As the energy supplied to the column is increased, more ethanol is stripped into the overhead vapor stream. The line marked “No recycle of permeate or residue” shows the energy used when these streams are condensed separately and none of their latent heat content is recovered in the stripper. When the process is operated this way, about 675 W of steam are needed in the reboiler to achieve 90% feed ethanol recovery in the overhead vapor. The line labeled “Recycle of permeate” illustrates what happens when heat integration is performed by returning the permeate vapor stream directly to the bottom of the stripper so that its latent heat is recovered. In this mode, 90% recovery of the ethanol
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Figure 6. Flow diagram of a conventional beer still (stripper)/rectification column/molecular sieve bioethanol separation plant. The plant produces 30 million gal/y of pure (>99 wt %) ethanol. Energy usage of the process is approximately 65 million Btu/h.
in the overhead vapor requires about 318 W of steam be delivered to the reboiler. The final line, marked “Recycle of permeate and residue”, shows that when the permeate is recycled to the stripper and the residue is condensed in the reboiler, the steam required to achieve 90% ethanol recovery falls to about 115 W. The electrical energy used to power the overhead compressor must be subtracted from these very large savings in reboiler energy, to determine the overall energy savings. A 75% efficient compressor uses about 90 W of electric power when ethanol recovery is 90%. Even when the electrical power used is multiplied by 3 to take into account the conversion of electric power to heat, the combination distillation-membrane process uses 43% less energy than simple distillation and performs a better separation. 2. Distillation-Vapor Compression-Membrane Separation Process Design When we started our research on distillation-membrane permeation processes, we thought the idea of recovering the latent heat content of the membrane unit’s permeate and residue streams in the distillation column was original with us. Many papers have appeared describing combination distillationmembrane processes, but almost all treat the two processes as a simple combination of unit operations with no heat integration.19-21 However, a search of the patent literature surfaced a number of relevant Japanese references,22-24 and it also appears that other companies are working on similar processes.25,26 There is nothing new under the sun. The heat-integrated process we are developing can be divided into two subgroups, depending on the relative boiling points of the organic component and water: one in which the organic solvent is the lighter (lower-boiling) component and one in which the reverse is the case. We will illustrate the first category of the process using ethanol (light component)/water (heavy component) mixtures. Application of the process to acetic acid (heavy component)/water (light component) will then be used to illustrate the second category of mixtures.
2.1. Designs with Water as the Heavy Component of a Mixture. 2.1.1. Conventional Technology. One of the classic examples of this type of application is the separation of water/ ethanol mixtures. This is required in all bioethanol production plants. The design of a conventional bioethanol separation plant is shown in Figure 6. Approximately 200 of these plants are installed in the United States. A further 500 plants will be built if cellulose-to-ethanol technology takes off. The plant shown produces 30 million gal/y (∼11 500 kg/h) of dry ethanol. Heat integration is used throughout the plant to minimize energy consumption. Even so, the separation plant has a theoretical energy consumption of about 17 000 Btu/gal of ethanol produced, about 20% of the heating value of the ethanol. In the conventional corn-to-ethanol separation process shown in Figure 6, the fermentation broth slurry containing 11.5 wt % ethanol is stripped under moderate vacuum (0.5 bar) in a beer still. This column is run under vacuum to minimize temperaturedependent plating out of dissolved proteins and fats and to allow the reboiler (40 million Btu/h) to use low-grade heat (80 °C boil up) from the multieffect evaporators used to concentrate the wet stillage. Overhead vapor from the beer still (1) is sent to a rectification column that produces an overhead product close to the azeotrope (190 proofsapproximately 93 wt % ethanol) (2) and a bottoms product which is essentially water. This rectification column also runs under vacuum (0.5 bar), and thus, the reboiler (19 million Btu/h) can use low-grade heatsthe trick is to use the heat from condensing dry ethanol product vapor from the molecular sieve dryer (9 million Btu/h) to drive this reboiler, thus requiring only 50% additional energy input to the rectifier reboiler. The molecular sieve dryer is fed with the overhead product of the rectification column (2), evaporated under pressure. This evaporator (16 million Btu/h) is steam heated; typically the steam is generated from a gas-fired boiler. Superheated hot vapor is fed to the adsorbing bed, and the dry ethanol product vapor (3) is condensed. In total, the energy usage of the separation process is approximately 65 million Btu/h (40 million Btu/h in the beer still and 25 million Btu/h in the rectifier and the molecular sieve evaporator).
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Figure 7. Design of a distillation-membrane hybrid process for the separation of a 100 000 kg/h ethanol/water mixture (equivalent to 30 million gal/y of ethanol production). The membrane used has a water permeance of 2000 gpu and an ethanol permeance of 50 gpu. The assumed efficiency of the compressor is 75%. A simple stripper column is used in this design (no rectification section). Table 1. Stream Compositions and Operating Conditions for the Distillation-Membrane Hybrid Process Design in Figure 7 feed 1 ethanol (wt %) water (wt %) pressure (bar) temperature (°C) flow (kg/h)
11.5 88.5 1.0 37 100000
ethanol product 7 99.7 0.3 3.0 91 11450
water product 8 0.1 99.9 0.5 81 88550
2.1.2. New Distillation-Vapor Compression-Membrane Separation Process Design. The design of a distillationmembrane separation process used to perform the same separation in comparison to the above conventional technology (Figure 6) is illustrated in Figure 7. The performance of this design, and others described in this paper, was calculated using a commercial computer process simulator (ChemCad 5.6, Chemstations, Inc., Houston, TX) fitted with code written at MTR for the membrane separation units. The membrane used to separate the compressed ethanol-water overhead vapor has a water permeance of 2000 gpu and an ethanol permeance of 50 gpu. The distillation-membrane process shown in Figure 7 and Table 1 performs an initial separation with a vacuum stripper column as in the conventional bioethanol plant. This stripper produces an ethanol-free bottoms and an overhead vapor (2) at a pressure of 0.5 bar containing 65 wt % ethanol. This vapor is then compressed to 3 bar. Compression increases the temperature of the vapor and a heat exchanger (not shown) integrated with the reboiler is used to cool this vapor to about 120 °C (about 5 °C above the dew point). The compressed gas is then sent to the membrane separation unit. In the block diagrams shown in Figures 2 and 4, a single membrane unit is used to separate the overhead vapor and all of the permeate
vapor is sent to the distillation unit. As explained below, it is often more efficient to separate the membrane unit into two parts. The first membrane unit lowers the water content of the overhead vapor from 35 wt % water (2) to ∼10 wt % water (3). The permeate vapor from this unit (4) has a high water concentration (93 wt % water) and contains the bulk of the water content of the overhead vapor. The remaining water is removed by a second membrane unit. The second unit lowers the water concentration from ∼10 to 0.3 wt % water. Because the vapor being treated by this unit has a lower average water concentration, the permeate (5) contains less water and more ethanol. This stream is condensed and remixed with the feed fermentation broth. The dry ethanol residue vapor stream produced by the second membrane unit (6) is condensed in the stripper column reboiler to recover its latent heat content. 2.1.3. Pressure Ratio of the Membrane Units. Two membrane separation units are used for the process shown in Figure 7 because of the problem of the pressure ratio. The first unit lowers the water concentration from 35 wt % (58 mol %) water to 10 wt % (22 mol %) water. The second unit then lowers the water concentration further to 0.3 wt % (0.76 mol %) water. The importance of the pressure ratio in the separation of gas mixtures can be illustrated by considering the separation of a gas mixture with component concentrations (mol %) nio and njo at a feed pressure of po. A flow of component i across the membrane can only occur if the partial pressure of component i on the feed side of the membrane, niopo, is greater than the partial pressure of component i on the permeate side of the membrane, niLpL. That is niLpL e niopo
(4)
It follows that the maximum separation achieved by the membrane can be expressed as niL nio
e
po pL
(5)
This means that the separation achieved can never exceed the pressure ratio of po/pL, no matter how selective the membrane is. The first membrane unit shown in Figure 7 operates at a feed pressure of 3 bar and a permeate pressure of 0.5 bar, or a pressure ratio of 6. The incoming feed water concentration is 58 mol %, and the exiting (residue) water concentration is 22 mol %. The average feed water concentration is about 40 mol %. At this high water feed concentration, the first membrane unit is not affected by the pressure ratio limitation expressed by eq 5. The second membrane unit shown in Figure 7 also operates at close to 3 bar in the feed. This unit is used to reduce the water concentration in the ethanol product from 22.1 mol % in the feed to 0.76 mol % in the dry ethanol product. The performance of the unit is very much affected by the pressure ratio. If the unit were operated at a pressure ratio of 6, this unit would be completely controlled by the limitation of eq 5. The permeate vapor at any point in the membrane unit would never be more than 6-fold more concentrated than the feed vapor at the same point. The problem of this enrichment limitation is particularly acute at the residue end of the unit where the water concentration falls to the 1-2 mol % water range. This means that the permeate vapor water concentration cannot be more than 6-12 mol % water (88-94 mol % of the permeate has to be ethanol) no matter how selective the membrane is. This result implies a large and wasteful recycle of ethanol back to the beer still.
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Figure 8. Calculation of the total system energy and total membrane area requirement for the 30 million gal ethanol/y plant shown in Figure 7 at various ethanol concentrations of stream 3 in the Figure 7 design. The total energy consumption shown includes the heat used in the column reboiler, and compressor electrical energy converted to heat at 10 000 Btu/ kW (reflecting a conversion efficiency of heat to electricity of 3:1). The efficiency of the compressor is 75%.
The solution to this problem is to increase the pressure ratio across the second membrane unit by reducing the pressure on the permeate side of the membrane to 0.1 bar. This can easily be done by cooling the permeate vapor to 30 °C which completely condenses the vapor and spontaneously creates the vacuum required. At a pressure ratio of 30 (feed pressure 3 bar/ permeate pressure 0.1 bar), the feed end of the second membrane unit is not affected by the pressure ratio limitation of eq 5, and the effect at the residue end of the second membrane unit is much reduced, although not completely eliminated. Increasing the pressure ratio across the second membrane step from 6 to 30 increases the concentration of water in the permeate stream (5) from 19 to 43 wt %. It also reduces the membrane area required to perform the same separation from 14 000 to 2500 m2. However, cooling and condensing the permeate vapor means the latent heat of vaporization of this stream is lost. Fortunately, this stream (5) is much smaller than the main permeate vapor stream (4), so the impact on the total process energy consumption of losing this fraction of the permeate latent heat is not large. A further reduction of pressure on the permeate side of the membrane to below 0.1 bar would be beneficial. However, as a practical matter, the pressure difference required to move the permeate vapor from the permeate side of the membrane to the condenser limits useful permeate pressures to about 0.1 bar. Determining the optimum design of the membrane portion of the process shown in Figure 7 requires calculation of the relative size of the two membrane units. The first unit removes the bulk of the water and recovers the latent heat content of the permeate vapor in the column. The second membrane unit loses the latent heat content of the permeate vapor but has a much reduced membrane area and reduces recycle of ethanol to the column. This balance can be determined by calculating the energy consumption and the membrane area required for the process as a function of the ethanol concentration in the intermediate stream (3) between the two membrane units. These plots are shown in Figure 8. The results show that the best balance between the separations performed by the two membrane separation units occurs when the residue stream (3) contains about 78 mol % (or 90 wt %) ethanol. At this ethanol concentration, the total energy con-
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Figure 9. Trade-off between increased process energy consumption and decreased membrane area as the size of the distillation rectification section increasessfrom a stripper column (no rectification section) to a column producing an overhead vapor at the azeotropic concentration (maximum rectification). The design with no rectification is illustrated in Figure 7. In these energy calculations, the reboiler energy is taken as the theoretical value and the electrical energy used by the compressor is converted to heat energy, assuming 1 kW electricity ) 3 kW heat.
sumption is at a minimum (33 million Btu/h), as is the total membrane area required. The message of Figure 8 is that if the water concentration in the residue stream (3) leaving the first membrane unit is very low, excessive amounts of ethanol in the permeate stream (4) of the first membrane unit will be recycled with the water back to the beer still column. This will increase the energy consumption of the column and the size of the vapor compressor. On the other hand, if the water concentration of stream (3) is too high, the latent heat content of water in the second membrane permeate stream (5) will increase. This latent heat is not beneficially used, so the overall energy efficiency of the process will fall. 2.1.4. Determination of the Balance Between Distillation and Membrane Separation. In designing a distillationmembrane separation hybrid process, the best balance between the separation performed by distillation and the separation performed by the membrane unit must also be calculated. This balance will vary depending on the cost of energy, the capital cost of the separation units, and the separation to be performed. For the case of corn-to-ethanol separation, the overhead vapor concentration can vary from about 65 wt % ethanol (the example illustrated in Figure 7), in which the column has no enriching (rectification) section and just acts as a vapor stripper, to about 93 wt %, the azeotropic value, when the distillation column has a rectification section and performs the maximum separation. To determine the best balance point, a series of calculations for various distillation-membrane separation combinations have been performed. The results are shown in Figure 9. As the ethanol concentration in the overhead vapor increases, that is, as more of the water/ethanol separation is performed by distillation, the membrane area required decreases; however, the energy consumption of the process increases. The design shown in Figure 7 uses a simple stripper column with no rectification section. This turns out to be the lowestenergy design. Condensing a portion of the overhead vapor to produce a liquid reflux stream will increase the concentration of ethanol in the overhead vapor. This decreases the amount of
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Figure 10. Calculation of the optimum stage to recycle the permeate vapor stream (stream 4 in Figure 7) to the distillation column.
membrane area needed to perform the rest of the separation but increases the energy needed for the reboiler. How much rectification is optimum is based on the trade-off between the total energy used in the process (reboiler and compressor) and the membrane area requirement. 2.1.5. Design of the Column. The fermentation beer feed to the stripper column in a corn-to-ethanol plant contains high levels of suspended and dissolved solids, so column designers use many tricks to achieve reliable stable operation. At low ethanol concentrations, the ethanol-water VLE diagram is very favorable to distillation. The reboiler duty decreases and the overhead vapor ethanol concentration increases with the number of theoretical stages in the column. But above 15 stages, a further increase in the number of stages does not significantly improve performance. In the following discussion, a 15-stage stripper column is used for the process shown in Figure 7. In the Figure 7 process, the point at which the permeate vapor stream (4) is recycled to the column is also important. Ideally, the recycle vapor should enter the column at the point where the column vapor phase has the same compositionsno mixing losses. A curve showing the change in overhead vapor ethanol concentration depending on the stage at which the recycle vapor (4) enters is shown in Figure 10. The optimum recycle point occurs between stages 11 and 12. However, the curve shows that there is a significant down side to feeding the permeate at a stage number that is too high (that is, feeding the permeate too close to the bottom), so feeding at stages 8-11 provides a more stable process. 2.1.6. Energy Savings. The energy saving of the process shown in Figure 7 compared to the conventional process (Figure 6) is very substantial. The conventional process uses about 65 million Btu/h of steam (about 17 000 Btu/gal ethanol product). The integrated process shown in Figure 7 uses 1200 kW/h of electricity to power the large compressor, but only 21 million Btu/h of steam. Even when the thermal inefficiency of converting heat to electricity is taken into account, this process uses only about half the amount of heat of the conventional process: 33 million Btu/h (12 million Btu/h of electricity at 10 000 Btu/ kW and 21 million Btu/h of steam, a total of 8600 Btu/gal ethanol product). The overall cost savings will be $3 million/y for a 30 million gal/y ethanol production plant, or about 10¢/gal of ethanol produced. This savings can significantly change the environmental impact of producing the bioethanol. If a cogeneration
Figure 11. Vapor-liquid equilibrium diagrams for various solvent/water mixtures that are potential candidate applications for the proposed new process.
plant is used to produce electricity on site, the savings could be even higher. 2.1.7. Other Applications with Water as the Heavy Component of a Mixture. Many separations in which the heavy component is water could benefit from using the new type of distillation-membrane process. The new technology is likely to be particularly useful with aqueous mixtures that form azeotropes or mixtures that have pinch zones in their vapor-liquid equilibrium diagrams at high solvent concentrations. The vapor-liquid equilibrium diagrams for a number of potential separation candidates are shown in Figure 11. Ethanol/water and isopropyl alcohol (IPA)/water are examples of mixtures with azeotropes; acetone/water and tetrahydrofuran (THF)/water are examples of mixtures with pinch zones at high mole fractions of solvent in the liquid. With all of these mixtures, producing dry solvent from a solvent/water mixture is difficult and energy demands are high, so use of the new process would be particularly attractive. 2.2. Designs with Water as the Light Component of a Mixture. 2.2.1. Conventional Technology. Separating a highboiling solvent from water by distillation is always an energyintensive process. First, for any dilute aqueous feed, much more water than solvent must be evaporated, and second, the latent heat of vaporization of water is about six times larger than that of most organic solvents. Acetic acid/water is a well-known separation problem of this type. The difficulty of separating acetic acid/water mixtures by distillation is illustrated by the vapor-liquid equilibrium curve for acetic acid/water mixtures as shown in Figure 3. Since water and acetic acid (boiling point ) 118 °C) have only small differences in volatility, the VLE curve is narrow. For a typical aqueous acetic acid mixture, the distillation column can economically produce a bottom stream of 90% acetic acid/10% watersgood enough to recycle for use as a process stream. However, producing low levels of acetic acid (
