Water Mixtures via Fractional

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Separation of Vapor-Phase Alcohol/Water Mixtures via Fractional Condensation Using a Pilot-Scale Dephlegmator: Enhancement of the Pervaporation Process Separation Factor Leland M. Vane,*,† Franklin R. Alvarez,† Anurag P. Mairal,‡ and Richard W. Baker‡ National Risk Management Research Laboratory, U.S. Environmental Protection Agency, Cincinnati, Ohio 45268, and Membrane Technology and Research, Inc., Menlo Park, California 94025-1516

In pervaporation, a liquid mixture contacts a membrane surface that preferentially permeates one of the liquid components as a vapor. Our approach to improving pervaporation performance is to replace the one-stage condenser traditionally used to condense the permeate with a fractionating condenser called a dephlegmator. For example, pervaporation of 5 wt % aqueous ethanol yields a vapor containing 35 wt % ethanol. The separation factor for the process is 10. Condensation of this vapor in a dephlegmator yields a vapor product stream containing 90% of the permeating ethanol at a concentration of 85 wt % ethanol. The net result of the combined pervaporation-dephlegmation process is to transform the 5 wt % ethanol feed into an 85 wt % ethanol condensed product. For the overall process, the separation factor increases 11-fold to 108. Pilot-scale parametric experiments were performed with a plate-fin heat exchanger operated as a dephlegmator. The process was modeled with commercial process simulation software; good agreement between the model and the pilot results was obtained. Introduction The cost competitiveness of fermentation processes depends on the efficiency with which the fermentation products are removed from the biological media. The concentration of fermentation products in the fermenter is generally on the order of 10-100 g/L (1-10 wt %). Such low concentrations greatly reduce the cost-effectiveness of traditional separation processes, such as distillation. Furthermore, many fermentation products form azeotropic mixtures with water, thus greatly complicating distillation processes. Pervaporation is an attractive alternative to distillation for this application. In this process, one side of a nonporous membrane is exposed to a liquid feed stream, and a vacuum or sweep gas is applied to the other side. The component or components targeted for removal permeate the membrane and evaporate into the permeate stream. The reduced partial pressure of compounds in the permeate provides the driving force for the separation. Slow-permeating components remain in the liquid residue. In traditional pervaporation systems, the permeate vapor is completely condensed to obtain a liquid that can be further processed. A typical system is shown in Figure 1. The permeate is condensed in a single-temperature heat exchanger (the precondenser). The precondenser feeds into a liquid ring vacuum pump that compresses the condensate and any remaining vapors to atmospheric pressure. The resulting liquid condensate has the same concentration of the desired component as the vapor permeate. Traditional pervaporation systems are generally satisfactory when the permeate concentration is high because of a large membrane/evaporation separation * To whom correspondence should be addressed. Tel.: (513) 569-7799. Fax: (513) 569-7677. E-mail: [email protected]. † U.S. Environmental Protection Agency. ‡ Membrane Technology and Research, Inc.

Figure 1. Schematic diagram of a traditional permeate condensation system.

factor, a high feed concentration, or phase separation of the condensate. For example, in the separation of tetrachloroethylene (perchloroethylene or PCE) from water, the separation factor is several thousand, and the permeate condensate can be separated into an almost-pure PCE phase and a PCE-saturated water phase because of the low water solubility of PCE. In this case, a simple condensation step yields a highly desirable result.1 In contrast, traditional pervaporation systems are generally unsatisfactory when the desired component is present in low concentration and/or exhibits a modest separation factor. (Aroma compounds are examples of such components.2,3) Traditional pervaporation is similarly ineffective in recovering hydrophilic organic compounds, such as ethanol, from water. In these cases, complete condensation of the permeate gives a conden-

10.1021/ie0305667 CCC: $27.50 © 2004 American Chemical Society Published on Web 12/09/2003

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sate requiring additional separation steps to yield a useful product. For example, pervaporation systems for recovering ethanol from aqueous solution typically employ silicone rubber [poly(dimethylsiloxane), PDMS] membranes that achieve an ethanol/water separation factor of 5-10.1,4-10 The separation factor (R) is defined as

REtOH-water )

yEtOH/ywater xEtOH/xwater

(1)

where yi and xi are the permeate and feed mass fractions, respectively, of component i and EtOH is ethanol. Thus, if the feed liquid contains 5 wt % ethanol in water and the PDMS membrane achieves a separation factor of 10, the vapor that permeates will contain 34.5 wt % ethanol. This 34.5 wt % ethanol permeate requires further processing to make it usable as a solvent or fuel additive. (Such applications generally require 99 wt % ethanol.) The additional processing includes distillation to approach the ethanol/water azeotrope (96 wt % ethanol), followed by dehydration with molecular sieves or pervaporation with hydrophilic membranes. In this case, the result of applying pervaporation to the 5 wt % ethanol stream is to increase the concentration of ethanol in the distillation feed stream and reduce the size of the distillation column. The ideal pervaporation system would yield a product that would not require a subsequent distillation step. At present, even the most ethanol-selective membrane materials cannot achieve this goal, especially for low ethanol feed concentrations. Fortunately, modification of the membrane material is not the only avenue for improving the overall separation performance of a pervaporation system. Much research has focused on the impact of feed-side variables, including temperature, concentration, and feed flow rate/concentration polarization. Other studies have evaluated permeate-side properties, such as the effect of permeate pressure on the flux and the separation factor. In nearly all of these studies, complete condensation of the permeate in a single condenser has been the standard operating procedure. Only a few investigators have studied condensation schemes that could alter the overall separation factor of the pervaporation system.2,3,11-14 In these cases, multiple condensers were operated in series or in parallel and at different temperatures and pressures. Because the components of the permeate vapor have different condensation potentials, the condensates differ in concentration, thus yielding an additional separation factor. The most extensive work in this area has been reported by Marin et al.2,3,12 for ethanol/water,12 ethyl acetate/water,2 methyl thiobutanoate/water,3 and 2,3butanedione/water3 binary systems and the ethanol/ ethyl acetate/water ternary system.12 In these papers, the masses and concentrations of condensates produced by two condensers operating in series at different temperatures were compared to those produced by a liquid nitrogen condenser that provided complete permeate condensation. In the ethanol/water system, water condenses at a higher temperature than ethanol. When the first condenser in the series is operated at a higher temperature than the second condenser, an ethanol-depleted condensate is produced at the first condenser, and an ethanol-enriched condensate is generated at the second

condenser. In this manner, Marin et al. were able to fractionate a 38 wt % ethanol (balance water) permeate vapor to obtain an ethanol-enriched condensate with an ethanol concentration as high as 71 wt % at the second condenser and an ethanol-depleted condensate with an ethanol concentration as low as 9 wt % at the first condenser.12 They were not able to achieve these levels of enrichment and depletion simultaneously (see below). The ability to concentrate ethanol from 38 to 71 wt % represents an additional separation factor of 3.9, raising the overall separation factor of the pervaporationcondensation system from 5.5 to 22 (based on a pervaporation feed of 10 wt % ethanol). These results illustrate the tradeoff between the ethanol concentration in the condensate and the fraction of ethanol recovered at each condensation stage. For example, when 30% of the ethanol in the permeate was recovered at the second condenser, the concentration of ethanol in the condensate was 71 wt %. (The remaining 70% of the ethanol in the permeate was recovered at the first condenser, which yielded a condensate containing 28 wt % ethanol.) When ethanol recovery in the second condenser was increased to 96%, the concentration of ethanol in the condensate was only 53 wt %.12 The tradeoff results from the use of single-temperature condensers, each of which delivers only a single vapor-liquid equilibrium (VLE) point. Performance could be improved by increasing the number of condensers in series, with each condenser operating at a different temperature and representing an additional VLE stage. However, the expense and complexity of such a system would be excessive. A reasonable alternative to multiple single-temperature condensers in series is a chilled contactor column. In this high-surface-area column, multiple VLE stages can be created by establishing a temperature profile. The term “dephlegmator” has been used to describe a vapor condensation system consisting of a high-surfacearea contactor and a heat removal system.15-18 In dephlegmation systems, the feed is a hot vapor; cooling generates the downward-flowing liquid. Feed components with a low vapor pressure and high boiling point will condense in the dephlegmator; those with a high vapor pressure and low boiling point will remain in the vapor phase. Thus, for the ethanol/water vapor system, water will be removed as a condensate, and most of the ethanol will remain in the overhead. Three dephlegmator configurations are shown in Figure 2. All three configurations provide sufficient heat and mass transfer for effective separations. In configuration A, a high-surface-area heat exchanger is used as a dephlegmator. In this configuration, heat removal and mass transfer occur on the same surfaces. In contrast, heat removal is achieved in a heat exchanger separate from the mass-transfer device in configurations B and C. In configuration B, a condenser provides the downflowing reflux condensate to a high-surface-area heat exchanger; in configuration C, a condenser sends the condensate to a high-surface-area packed masstransfer column. As in distillation systems, multiple VLE stages can be established by creating a temperature profile in the dephlegmator and providing an adequate mass-transfer surface area. The application of dephlegmation systems to the separation of vapor mixtures can be problematic. Vendors contacted in preparation for this work were not confident about the design methodology for separation

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Figure 2. Dephlegmator configurations: (A) high-surface-area heat exchanger with coolant flow countercurrent to vapor flow, (B) partial reflux condenser above high-surface-area heat exchanger, and (C) partial reflux condenser above column containing high-surface-area packing.

units, such as this one, that operate under vacuum conditions. In addition, unsatisfactory results from dephlegmator separations have been reported in the literature. For example, di Cave et al. reported that “the dephlegmator has been confirmed to be suitable for obtaining additional rectification of the vapor”, but the amount of separation was equivalent to only one theoretical stage.15 In this case, dephlegmator performance might have been limited by the equipment design and operation. A shell-and-tube heat exchanger was used as the dephlegmator; the inner surface of the tube served as the heat- and mass-transfer surface. The coolant flow rate appears to have been too high for the surface area of the exchanger. (This might indicate that the vapor loading rateswhich the article does not provideswas also too high.) Furthermore, the coolant was introduced to the shell at the bottom of the unit, cocurrent to the vapor that entered the bottom of the tube region. Optimal operation of a dephlegmator involves countercurrent coolant flow to generate the necessary temperature profile. The combination of insufficient mass-transfer area and inefficient coolant flow might have caused the poor performance of the dephlegmator in this case. This paper documents chemical process simulations and pilot-scale condensation experiments designed to determine the effectiveness of dephlegmation for the separation of pervaporation permeate vapor. The simulations provide a theoretical basis for assessing dephlegmator efficacy; the pilot-scale experiments provide a practical basis. The model system was ethanol/water vapor. Process Simulations The software program ChemCAD 5.0 (ChemStations, Houston, TX) was used for dephlegmator simulations. The dephlegmator was simulated using a ChemCAD “Tower Plus #1” separation column with user-specified heat removal from each stage of the column (Figure 3). Each simulated stage included a side exchanger for heat removal. No reboiler or reflux condenser was used. Simulations employed the default thermodynamics models selected by ChemCAD based on the ethanol/ water system and the target pressures and temperatures. Specifically, the nonrandom two liquid (NRTL)

Figure 3. Schematic diagram of a ChemCAD Tower Plus #1 column used for dephlegmator simulations. The column is shown with four equilibrium stages.

k-value thermodynamics model, latent heat enthalpy model, and ideal gas heat capacity were applied. In addition, no correction was made to the vapor fugacity, and no vapor-phase association was assumed. The base-case scenario for the simulations involved a 100 kg/h permeate vapor containing 34.5 wt % ethanol (balance water) at 60 °C and 30 Torr absolute pressure introduced into a four-stage dephlegmator with equal amounts of heat removed from each stage. The amount of heat removed in each stage could be adjusted independently to optimize performance of the dephlegmator. The ethanol/water VLE behavior indicates that the first stages will remove the majority of water and, therefore, the majority of the cooling utility will be applied in these stages. Equal heat removal from each stage was selected as the base-case condition to simplify the initial analysis. The permeate vapor properties represent those for a pervaporation system equipped with a silicone rubber membrane (R ) 10) removing ethanol from a 5 wt % ethanol liquid feed stream. The column pressure was the same as the pressure of the vapor feed stream. The simulation parameters that were studied included the number of stages, heat removed per stage, composition

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Figure 4. Schematic diagram of a pilot-scale dephlegmation unit showing the position of pressure and temperature measuring points.

of the feed vapor, pressure of the feed vapor (and column), temperature of the feed vapor, and flow rate of a noncondensing gas (carbon dioxide) introduced with the feed vapor. The simulation results will indicate how these process parameters affect process performance for an equilibrium-based system. The impact of these same parameters on mass- and heat-transfer rates, which will also impact performance, is outside the scope of the simulations performed. Experimental Section The pilot-scale dephlegmator apparatus was assembled in the high-bay area of the U.S. Environmental Protection Agency (USEPA) Test & Evaluation Facility in Cincinnati, OH. Most of the components of the test apparatus are part of the USEPA pervaporation pilot unit. For this phase of dephlegmator testing, pervaporation membranes were not used to create the feed vapor. Instead, an evaporator system generated the feed vapor. A schematic diagram of the evaporator-dephlegmator system is shown in Figure 4. A 1000-gallon epoxy-lined fiberglass-reinforced plastic tank served as the process feed tank. The feed tank was equipped with a recirculation pump and loop to keep the tank well mixed. A stainless steel centrifugal pump transferred fluid from the feed tank to a plate-type heat exchanger (not shown) in which the feed solution was preheated using solution being returned to the feed tank. The feed fluid then passed through a 30-kW electric heater controlled with a thermistor-based controller. After the heater, the feed was filtered using 5- and 1-µm bag filters and exposed to UV radiation to reduce the possibility of microbial growth. A slip stream was taken to feed the dephlegmator column; the rest of the liquid was returned to the feed tank via the heat recovery exchanger. A peristaltic pump metered the slip stream into a tube-in-tube heat exchanger that was warmed with a recirculating heated bath to evaporate the feed liquid. The elevated temperature in the evaporator and the reduced pressure in the dephlegmator resulted in the evaporation of the feed liquid to create the dephlegmator feed vapor. A dephlegmator column consisting of a brazed aluminum plate-fin heat exchanger (Chart Heat Exchangers, La Crosse, WI) was used for the work reported here.

The approximate dimensions of the brazed heat exchanger are 8 in. wide × 9 in. deep × 94 in. high (0.20 m × 0.22 m × 2.4 m). This type of exchanger consists of alternating parallel panels of coolant and process vapor separated by closely spaced thin fins and fused together in a brazing oven. In total, the exchanger contains eight vapor panels and nine coolant panels. The active length of the vapor channels was about 70 in. (1.8 m); the nominal vapor cross-sectional flow area was approximately 0.16 ft2 (0.015 m2), ignoring the area occupied by the fins. The mass-/heat-transfer area was estimated to be 21.6 m2. Pictures of the plate-fin heat exchanger are provided in Figure 5. Multiple VLE stages were established by creating a temperature profile in the dephlegmator. Coolant was introduced at the top of the dephlegmator at a set temperature and flow rate. The warm vapor feed entered at the bottom of the dephlegmator and flowed countercurrent to the coolant. As the vapor was cooled, vapor molecules condensed. The heat of condensation was transferred to the coolant, resulting in a temperature increase in the coolant. Thus, the coolant and the vapor both exhibited a temperature gradient with the lowest temperatures for both coolant and vapor observed at the top of the dephlegmator. A recirculating chiller (Icewagon Industries, model 30 ACXP) provided coolant to the dephlegmator. The coolant temperature was controlled by the chiller; the coolant flow rate was controlled with diaphragm valves and monitored using in-line rotameters. To assess process temperatures, the unit was equipped with eight T-type thermocouples (T1-T7 in Figure 4; T8 not shown) connected to a 12-input digital display unit. The thermocouples were located as follows: T1, vapor feed; T2, dephlegmator bottoms condensate; T3, vapor at one-quarter height; T4, vapor at three-quarters height; T5, overhead vapor; T6, coolant in; T7, coolant out; T8, room temperature (not shown in Figure 4). The accuracy of thermocouples T1-T8 was checked at room temperature using two NIST-certified thermometers. All thermocouples were accurate to within (0.4 °C. Vacuum was supplied to the top of the dephlegmator column by a multistage dry chemical vacuum pump with integrated vapor compression and condensation (Stokes Vacuum, CD-75 dry pump, 75 cfm nominal, 5 hp). A recirculating chiller (Icewagon Industries, model DE5AC) provided -5 to 1 °C coolant to the vacuum pump heat exchangers. The pilot unit was equipped with a vacuum control system that automatically adjusted a butterfly valve on the inlet to the vacuum pump to control the upstream pressure in the vacuum manifold and dephlegmator. Three vacuum gauges (P1-P3 in Figure 4) were used to monitor vacuum level in the system. P1 was a strain gauge (Edwards ASG1000) located at the vapor inlet at the base of the dephlegmator. P2 was a heated electronic manometer (Edwards Barocel pressure sensor model 658 AB) located just upstream of the vacuum control valve. P3 was a digital diaphragm gauge (Vacuubrand DVR2) also located just upstream of the vacuum control valve. A mercury vacuum manometer was attached at locations close to gauges P1-P3 to monitor the accuracy of the gauges on a daily basis. In the vacuum region of interest (30 Torr), all of the gauges were within (3 Torr of the mercury manometer reading. Bottoms condensate from the dephlegmator column was collected in a tank located below the dephlegmator


Figure 5. (a) Picture of the USEPA plate-fin dephlegmator system with letters indicating the (A) plate-fin heat exchanger, (B) evaporator unit, (C) coolant flow control valves and rotameters, (D) vapor feed zone and condensate collection region, and (E) overhead vapor piping. (b) Picture of the lower portion of the plate-fin exchanger prior to assembly, showing (F) the entry region for upflowing feed vapor and exit for downflowing condensate and (G) the coolant exit port. (c) View of the vapor entry region inside the plate-fin exchanger showing alternating coolant and vapor/condensate zones (dark regions with closely spaced fins).

(reservoir 1 in Figure 4). The dephlegmator overhead vapor that condensed in the heat exchangers of the vacuum pump was collected in two tanks located below the vacuum pump (collectively labeled reservoir 2 in Figure 4). During experiments, the condensate tanks were drained at specific intervals to determine the rate of accumulation and the composition of the condensate. System operating conditions, including coolant flow rate, process temperatures, feed flow rate, and vacuum pressure, were maintained and monitored at least twice per condensate collection period. The feed solution was prepared by dissolving 200proof ethanol (AAPER Alcohol, Shelbyville, KY) in the water contained in the feed tank. The target feed concentration of ethanol for most of the experiments (34.5 wt %) matched the conditions chosen for the ChemCAD simulations. The target vapor feed flow rate was 2 kg/h (33 g/min); other feed rates were also studied. The unit was allowed to operate for at least 3 h to reach steady state. One indication of steady state was

constant process temperatures T1-T7. At the end of the start-up period, the condensate reservoirs were drained, and a sample of the feed liquid was obtained. Condensate was collected for periods of 1-4 h between reservoir draining events. The condensates were weighed using an electronic balance with 6-kg capacity and 0.1-g precision. Forty-milliliter samples of each condensate were collected. A valved sampling port located just after the UV system was used to obtain feed samples to determine the composition of the vapor feed to the dephlegmator. A feed sample was acquired each time the condensate reservoirs were emptied. Target conditions were maintained overnight with automatic transfer of condensate. After operating conditions were changed, a period of at least 3 h was allowed to elapse prior to sampling and condensate collection. Ethanol concentrations were determined following EPA Method 8015 using a Tremetrics 9001 gas chromatograph (GC) equipped with a flame ionization detector (FID) and CombiPAL direct injection autosampler

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(LEAP Technologies, Carrboro, NC). Most ethanol concentrations fell outside the calibration range of the GC. As a result, dilution of the samples was necessary. A mass aliquot of each sample was diluted with deionized water in a volumetric flask. Isopropyl alcohol was added to each sample as a surrogate at a concentration of 1000 mg/L to monitor operation of the GC. The diluted samples were transferred to 2-mL autosampler vials sealed with Teflon-lined caps. Each dilution was performed in duplicate. All samples were stored in a refrigerator at 4 °C and analyzed within 24 h. Aliquots of 200-proof ethanol, diluted in the same manner as the dephlegmator samples, served as quality-control (QC) checks. These QC checks were usually within 2% (and always within 4%) of 100 wt % ethanol. The results presented here are based on two to four separate condensate collection periods under steadystate conditions extending over 1-2 days. Average values and, when appropriate, 95% confidence intervals are provided. Although each experiment consisted of replicate condensate collection periods, duplicate experiments were performed periodically to ensure that the same steady-state conditions and performance could be achieved after other experiments had been performed. For each experiment, mass balance closure was calculated as the percentage of ethanol in the feed vapor recovered in either the bottoms or the overhead condensates. Closure averaged 98.2% and ranged from 94.8 to 100.0%. Results

Figure 6. Effect of the heat removed per stage on the overhead ethanol concentration (wt %), bottoms ethanol concentration (wt %), water recovery in the overhead (%), and ethanol recovery in the overhead (%) for a four-stage dephlegmator. Feed vapor: 100 kg/h; 34.5 wt % ethanol, 65.5 wt % water; 60 °C; 30 Torr. Symbols represent simulation results from specific simulation runs; lines are spline curve fits of those results.

Simulations. The manufacturers of the proposed pilot-scale dephlegmator estimated that the height equivalent of a theoretical plate (HETP) was 12-18 in. (0.30-0.46 m). Because the height of the pilot unit was 6 ft (1.8 m), the number of theoretical stages was expected to be 4-6. The four-stage scenario was selected as the simulation base case because this matched the lower end of the expected range. The first variable studied was heat removed per stage. Results of these simulations for the four-stage base case are shown in Figure 6. The figure shows the effect of heat removed per stage on four properties: the concentration of ethanol in the overhead vapor (overhead EtOH concentration), the concentration of ethanol in the bottoms condensate (bottoms EtOH concentration), the percentage of water recovered in the overhead vapor (water recovery in overhead), and the percentage of ethanol recovered in the overhead vapor (EtOH recovery in overhead). Initial increases in the heat removed per stage lead to the selective condensation of water, resulting in an increase in the concentration of ethanol in the overhead vapor and high ethanol recovery. However, further increases in heat removed per stage lead to ethanol condensation, raising the concentration of ethanol in the bottoms condensate and resulting in poor recovery of ethanol in the overhead. According to Figure 6, the transition from selective condensation of water to condensation of both water and ethanol is fairly sharp and occurs at about 11 kW of heat removed per stage. Above this transition point, the concentration of ethanol in the overhead rises slowly while the recovery of ethanol drops rapidly. We selected an operating point that delivered 90% recovery, corresponding to 85.4 wt % ethanol in the overhead vapor, 5.4 wt % ethanol in the bottoms condensate, and 43.6 kW for the total heat

removal rate. Under these conditions, the temperature of the overhead vapor was 14.5 °C, and the temperature of the bottoms condensate was 23.9 °C. Several process variations were studied with the fourstage base-case scenario. First, we examined the effect of carbon dioxide on performance. Ethanolic fermentations produce large amounts of carbon dioxide, some of which dissolves in the fermentation broth as carbonic species. However, the majority of the gas leaves the fermenter via a vent. Silicone rubber pervaporation membranes are permeable to carbon dioxide. As a result, a portion of the dissolved carbon dioxide will be stripped from the fermentation broth and appear in the permeate. From an equilibrium standpoint, the carbon dioxide will dilute the ethanol in the vapor phase, requiring lower temperatures to condense the ethanol. From a heat- and mass-transfer standpoint, the addition of carbon dioxide has at least two competing effectss reducing the residence time in the dephlegmator and increasing the heat-/mass-transfer coefficientsswith the net effect unknown. In the worst-case scenario, all dissolved carbonic species will be removed by the membrane. Water in contact with 1 atm of carbon dioxide at a pH of 7 will contain 0.17 mol/L of carbonic species (H2CO3, HCO3-, and CO32-). The base case 100 kg/h of 34.5 wt % ethanol permeate is assumed to originate from the treatment of 700 kg/h of fermentation broth. Thus, if all carbonic species are removed, the permeate will contain 120 mol/h carbon dioxide (equivalent to 5.3 kg/h or 49 SLPM). Addition of carbon dioxide to the feed in the ChemCAD simulations resulted in slightly higher overhead ethanol concentrations, slightly lower ethanol recoveries, and lower column temperatures when the heat removal rate was held constant. Only when the


Figure 7. Effect of the number of equilibrium stages on the overhead ethanol concentration (wt %), bottoms ethanol concentration (wt %), water recovery in the overhead (%) and ethanol recovery in the overhead (%). Fixed total rate of heat removal (43.6 kW). Feed vapor: 100 kg/h; 35 wt % ethanol, 65.5 wt % water; 60°C; 30 Torr. Symbols represent simulation results from specific simulation runs; lines are spline curve fits of those results.

carbon dioxide load exceeded 20 kg/h did the addition of carbon dioxide have a significant impact on the process parameters. Thus, the presence of reasonable amounts of carbon dioxide or other noncondensable gases is not expected to have a significant impact on dephlegmator performance. Next, we studied the effect of changes in dephlegmator pressure on performance. Small changes in dephlegmator pressure (approximately 10 Torr) caused small changes in process performance, with the main effect being a 6 °C change in process temperatures at the new pressures. (Lower temperatures are needed at lower pressures.) This result indicates that higher coolant temperatures could be used if higher column pressures were employed. For example, 30 °C coolant is required at a column operating pressure of 100 Torr; 9 °C coolant is required at 30 Torr. The latter situation would require a refrigerated system, whereas heat exchange with ambient air might be possible in the former, thereby reducing utility and capital costs. In addition, a vacuum system required to deliver 100 Torr would be less expensive than a system required to deliver 30 Torr. The final variable studied with the four-stage simulation scenario was the effect of ethanol concentration in the feed on overhead ethanol concentration and recovery. For all vapor feed concentrations studied (9.2-34.5 wt % ethanol), 90% ethanol recovery in an overhead product containing 85 wt % ethanol could be achieved. The number of theoretical stages affects the separation efficiency of the dephlegmator. To assess the level of impact, simulations were performed with different numbers of stages while the total amount of heat removed was held fixed at 43.6 kW. The results of these simulations are shown in Figure 7. As expected, performance improved as the number of stages increased,

Figure 8. Comparison of the relationship between ethanol recovery in the overhead and ethanol concentration in the overhead for two simulated dephlegmators and an experimental system composed of two single-temperature condensers in series. Feed vapor for both simulations: 34.5 wt % ethanol, 65.5 wt % water; 60 °C; 30 Torr. Feed vapor for the experimental system (estimated): 38 wt % ethanol, 11 Torr. Literature data were calculated from data provided in Figure 5 of ref 12.

both in terms of the recovery of ethanol in the overhead product and in terms of the concentration of ethanol in that product. Each additional stage provided diminishing gains. The simulation results presented in Figures 6 and 7 suggest that high recoveries and high ethanol concentrations can be achieved with a modestly sized dephlegmator. A closer examination of the simulation results presented in Figure 6 reveals one of the fundamental characteristics of this condenser system: the tradeoff between ethanol recovery in the overhead and ethanol concentration in the overhead. Maximizing the overhead recovery reduces the concentration of ethanol in the overhead product. The tradeoff is illustrated in Figure 8, which shows the expected results for four- and sixstage dephlegmators and the experimental results of Marin et al.12 for two single-temperature condensers operated in series. According to the curves presented in Figure 8, the dephlegmators provide significant improvements in both ethanol recovery and ethanol concentration relative to the condensers-in-series. Further, the transition from the “high-recovery” region (horizontal portion of the dephlegmator operating curve) to the “high-concentration” region (vertical portion of the dephlegmator operating curve) is rather sharp. Thus, the dephlegmators should be able to deliver both high recovery and high ethanol concentration under the appropriate operating conditions. Experiments. The objective of the first set of experiments was to observe the effect of heat removal on dephlegmator performance, as measured by the recovery of ethanol and concentration of ethanol in the overhead vapor. For this set of experiments, the following feed conditions were fixed: 2 kg/h vapor load, 34.5 wt % ethanol, 60 °C, and 30 Torr absolute pressure. The amount of heat removed was controlled via the coolant temperature and flow rate. The operating conditions and results for each dephlegmator experiment are provided in Table SI-1 (see the Supporting Information).

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Figure 9. Comparison of the data from the 2 kg/h base-case experiments (symbols) with the results of four- and six-stage simulations (lines) for equivalent feed conditions (34.5 wt % ethanol, 65.5 wt % water; 60 °C; 30 Torr). Error bars represent 95% confidence intervals.

The results from the first set of experiments (a total of 14 experiments, experiments 4-16 and 19 in Table SI-1), referred to as the base case, are presented in Figure 9 along with the simulation results. As suggested by the simulations, the recovery versus concentration data fell along a single operating curve. In the highconcentration region, the experimental data were consistent with the results from the six-stage simulation. However, in the high-recovery region, the experimental ethanol recoveries were less than predicted by the simulations. Subsequent simulations and experiments with other dephlegmators share this difference between predicted and observed recoveries in the high-recovery region. The cause of the difference is not known, although it might be related to nonideal behavior in the dephlegmator, particularly in the region where the feed vapor interacts with the condensate leaving the bottom of the column. Nevertheless, 90 wt % ethanol at 89% recovery could be attained with the experimental system (see experiments 14 and 19 in Table SI-1). Also, operation of the system was stable in this region. Analysis of dephlegmator experimental results and process parameters showed that recovery was correlated with the bottoms condensate temperature and overhead ethanol concentration was correlated with the overhead vapor temperature. For the 90 wt %/89% recovery experiments, a 16 °C increase in coolant temperature and an 11 °C difference between the temperatures of the bottoms condensate and the overhead vapor were observed. These observations are in agreement with the 10 °C temperature difference predicted by the simulations. The pressure drop through the dephlegmator along the vapor path was estimated to be less than 4 Torr but was too low to be measured accurately. According to the temperature profile within the dephlegmator, most of the heat was removed from the vapor in the bottom 25% of the dephlegmator. As a result, the vapor volumetric flow rate was much lower in the upper 75% of the column than in the lower 25%. To assess the impact of changes in process conditions on dephlegmator performance, we performed experi-

Figure 10. Effect of dephlegmator pressure, feed temperature, inert gas flow rate, and feed loading rate on dephlegmator performance. Error bars are omitted for clarity. All 95% confidence intervals are less than 2% or 2 wt % in absolute terms.

ments in which one process parameter was varied relative to the base-case experiments. These process parameters included 45 °C feed temperature (experiment 17 in Table SI-1), 20 Torr feed pressure (experiment 18), nitrogen added to feed vapor at 0.16 or 0.97 standard liters per minute (SLPM) (experiments 20-22), a 2.4 or 4 kg/h feed rate (experiments 24 and 25-30, respectively), and a feed ethanol concentration of 10 wt % (experiments 31-34). When nitrogen was added to the feed vapor, it was introduced via a mass flow controller located just after the evaporator (see Figure 4). The nitrogen flow rates were selected to mimic the amount of carbon dioxide that might appear in the permeate if an actual fermentation broth was being treated by pervaporation as was discussed earlier. A nitrogen flow rate of 0.97 SLPM was selected to correspond to the worst-case scenario described in the simulation results section. A flow rate of 0.16 SLPM was selected to represent a more likely upper bound of carbon dioxide in the permeate. The results of most of these experiments are presented in Figure 10 with the base-case data and simulation results. (The results of experiments with 10 wt % ethanol feed are shown in Figure 11.) Although changes in the process conditions affected dephlegmator performance, the changes did not appear to alter the operating curve; within experimental error, all experimental data in Figure 10 overlap. In all cases, shifts in performance due to changes in the process parameters could be counteracted by modifying the coolant flow rate and/or coolant temperature. Changing the feed vapor temperature or introducing nitrogen gas had the smallest impact on dephlegmator performance. However, increasing the vapor feed rate required appropriate changes in the coolant flow rate to achieve equivalent performance. Dephlegmator pressure also had a large impact on performance, but adjusting the coolant flow rate returned system performance to desirable levels. In each case, the changes were in line with those predicted by the process simulations. The investigation of the final variable, feed ethanol concentration, yielded surprising results. For the same total mass feed rate (2 kg/h) and ethanol recovery (89%),


Figure 11. Comparison of the data from experiments with the results of four-, six-, and eight-stage simulations for equivalent feed conditions: 10 wt % ethanol, 60 °C, 30 Torr. Error bars represent 95% confidence intervals.

the concentration of ethanol in the overhead product was higher for the 10 wt % ethanol feed (experiment 34) than for the 34.5 wt % ethanol feed (experiments 14 and 19), yielding 93 and 90 wt % ethanol in the overhead, respectively. Statistical analysis of the data sets indicates that the 93 and 90 wt % values are statistically different at the 0.05 level of significance. In addition, for these experiments, the ethanol QC samples were all within 1 wt % of 100 wt %, so the 3 wt % difference between the two sets of data is considered significant. The difference between 90 and 93 wt % might seem small, but the amount of water in the overhead per unit alcohol has been reduced by 32%. Furthermore, this increase in the overhead ethanol concentration was achieved while the feed ethanol concentration was reduced. To explore this observation, dephlegmator simulations were performed for a feed vapor containing 10 wt % ethanol (balance water) to generate operating curves similar to those presented in Figure 9. The results of these simulations for four-, six-, and eight-stage dephlegmators are presented in Figure 11 along with the experimental data obtained with a 10 wt % ethanol feed vapor. Although the 34.5 wt % ethanol experimental results were in closest agreement with a six-stage simulated dephlegmator, the 10 wt % ethanol experimental results were in better agreement with an eightstage simulation, as indicated by the overlap of experimental and simulation results in the high-ethanolconcentration region of the curve. A possible explanation for this result is that the vapor velocity and loading in the upper regions of the dephlegmator are much lower for the 10 wt % ethanol feed vapor than for the 34.5 wt % feed at the same feed vapor mass flow rate. The HETP might be lower in the upper region, thereby increasing the number of effective stages for the 10 wt % feed concentration. In any case, the experiments demonstrate that 93 wt % ethanol can be delivered at high recoveries from a 10 wt % ethanol feed vapor. Discussion The computer simulations predicted that a dephlegmator could effectively separate an ethanol/water vapor

into an ethanol-enriched overhead vapor product and an ethanol-depleted bottoms condensate. The pilot-scale dephlegmator experiments confirmed these predictions. The implications for pervaporation process separation efficiency are significant. For example, the experimental data indicate that a pervaporation system equipped with a modest six-stage dephlegmator can deliver a 90 wt % ethanol product at 90% recovery from a 5 wt % ethanol feed liquid. According to eq 1, the separation factor for this application for a combined pervaporation-dephlegmation system would be 170, which is 17 times higher than the separation factor provided by the pervaporation membrane alone. The same pervaporation-dephlegmation system can deliver a 93 wt % ethanol product from a 1 wt % ethanol feed liquid, which translates into a spectacular process separation factor of 1300. Moreover, dephlegmator size is not limited to the 6-ft (1.8-m) height used in this study. Additional VLE stages could be added (at additional cost) simply by building a taller dephlegmator. For example, a dephlegmator that was 9 ft (2.7 m) tall would generate an overhead product that was 95 wt % ethanol at 90% recovery from a feed vapor of 10 wt %. (This feed vapor concentration corresponds to a pervaporation feed concentration of about 1 wt % ethanol.) The amount of water remaining in the overhead product would be reduced by 28%, from 7 to 5 wt %. At atmospheric pressure, an azeotrope occurs in ethanol/water systems at 95.6 wt % ethanol. At 95 Torr, the azeotropic composition is 99.5 wt % ethanol.19 At pressures below 95 Torr, an ethanol/water azeotrope does not exist. Although the azeotrope disappears, the incremental separation provided by a VLE stage in this region is relatively small. Nonetheless, theoretically, a dephlegmator operating at pressures below 95 Torr could be designed to deliver almost 100 wt % ethanol. However, such a system is unlikely to be practical because of the small VLE driving force as 100 wt % ethanol is approached. Thus, a properly designed dephlegmator of modest height can be expected to deliver 90-95 wt % ethanol product with at least 90% ethanol recovery when operated at pressures below 95 Torr. In this case, the distillation system normally required to refine ethanol/ water pervaporation permeate would be unnecessary. However, the proposed system would need a new column contactorsthe dephlegmator. Because the dephlegmator would operate under vacuum, the gas velocities would be higher than in a distillation column. The dephlegmator might require a larger column diameter to match the number of VLE stages in a comparable distillation column at this lower pressure. However, because a standard distillation column uses a reboiler, the vapor mass flow rate and condensate mass flow rate will be higher than in a comparable dephlegmator, which could lead to a reduction in the efficiency (increase in HETP) for the distillation column. Thus, the comparison between operating a dephlegmator under vacuum with minimum reflux and operating a distillation column with reboil and reflux is not straightforward. Although the dephlegmator used in this work was a high-surface-area heat exchanger (configuration A in Figure 2) requiring a separate coolant source, a simple and less expensive packed column with overhead condensate reflux (configuration C in Figure 2) is expected to be equally effective. Kent and Pigford compared such

182


systems and found that a packed column with reflux (referred to as an “adiabatic distillation column plus ordinary condenser” in the paper) required slightly less area to achieve the same level of separation as a “partial condenser” in which heat was removed from within the column.20 To simplify matters further for a pervaporation system, the final condenser attached to the liquid ring pump could be used to produce the reflux liquid, eliminating the need for a separate reflux condenser. As a result, addition of the dephlegmator separation unit will not require a reboiler or a separate reflux condenser. The next phase of this project will involve testing an integrated pervaporation-dephlegmation pilot system in which the permeate from a pervaporation unit operating on an ethanol/water feed stream will be processed using a dephlegmator. Several styles of dephlegmators (specifically, configurations A and C in Figure 2) will be tested to identify the most cost-effective configuration. Afterward, the relevant engineering cost comparisons between a distillation-only system, a pervaporation-distillation system, and a pervaporationdephlegmation system will be carried out. In this paper, performance of the dephlegmator has been measured by the concentration of ethanol in the overhead and the recovery of ethanol in the feed vapor. The disposition of the ethanol in the bottoms condensate has not been mentioned. In these experiments, 90% of ethanol was recovered in the overhead; the remaining 10% left the dephlegmator in the bottoms condensate. Under these conditions, the concentration of ethanol in the bottoms condensate was found to be approximately the same as that postulated for the feed stream to the pervaporation module. For example, a 5 wt % ethanol pervaporation feed stream should yield a 34.5 wt % ethanol permeate vapor; when this vapor is fed to a dephlegmator that achieves 90% recovery of the ethanol in the overhead, it will yield a bottoms condensate containing about 5 wt % ethanol. The bottoms stream could be recycled back into the pervaporation feed stream without altering the feed concentration, although a slightly larger pervaporation system would be needed because of the higher flow rate. Alternatively, the bottoms stream could be treated with a separate pervaporation module, the permeate of which could be sent to the same dephlegmator used by the first pervaporation module. Thus, the impact of designing the dephlegmator for 90% ethanol recovery should be relatively small. The overall fraction of ethanol recovered depends primarily on the pervaporation unit in a combined pervaporation-dephlegmation system. Finally, although the ethanol/water system has been the focus of this research, the pervaporation separation of other mixtures should also benefit from the addition of a dephlegmator. For example, dephlegmation can be used in pervaporation dehydration systems when the concentration of water is low in the feed and high in the permeate. More specifically, the high-flux, lowselectivity hydrophilic membranes used to dehydrate an alcohol/water stream could produce a permeate containing 1-5 wt % alcohol. A dephlegmator could be used to separate that permeate vapor into an overhead product with a high alcohol concentration (which could be recycled back to the dehydration membrane) and a bottoms condensate with a low alcohol concentration. Thus, the improvements provided by the addition of a

dephlegmator to a pervaporation system potentially apply to a wide range of pervaporation applications. Acknowledgment The authors thank Craig Patterson and Stuart Shealy of Shaw Environmental & Infrastructure, Inc. (formerly IT Corp.), for engineering and fabrication support and Ballard Mullins of the USEPA for analytical support. This project was funded, in part, by a Small Business Innovative Research (SBIR) grant from the National Science Foundation to Membrane Technology & Research. Supporting Information Available: The Supporting Information consists of one table (Table SI-1) in which details of the dephlegmator experiments are provided. This material is available free of charge via the Internet at http://pubs.acs.org. Literature Cited (1) Wijmans, J. G.; Kaschemekat, J. E.; Davidson, J. E.; Baker, R. W. Treatment of Organic-Contaminated Wastewater Streams by Pervaporation. Environ. Prog. 1990, 9, 262. (2) Baudot, A.; Marin, M. Improved Recovery of an Ester Flavor Compound by Pervaporation Coupled with a Flash Condensation. Ind. Eng. Chem. Res. 1999, 38, 4458. (3) Baudot, A.; Marin, M. Dairy Aroma Compounds Recovery by Pervaporation. J. Membr. Sci. 1996, 120, 207. (4) Blume, I.; Wijmans, J. G.; Baker, R. W. The Separation of Dissolved Organics from Water by Pervaporation. J. Membr. Sci. 1990, 49, 253. (5) te Hennepe, H. J. C.; Bargeman, D.; Mulder, M. H. V.; Smolders, C. A. Zeolite-Filled Silicone Rubber Membranes. Part I. Membrane Preparation and Pervaporation Results. J. Membr. Sci. 1987, 35, 39. (6) Shabtai, Y.; Chaimovitz, S.; Freeman, A.; Katchalski-Katzir, E.; Linder, C.; Nemas, M.; Perry, M.; Kedem, O. Continuous Ethanol Production by Immobilized Yeast Reactor Coupled with Membrane Pervaporation Unit. Biotechnol. Bioeng. 1991, 38, 869. (7) Strathmann, H.; Gudernatsch, W. Continuous Removal of Ethanol from Fermentation Broths by Pervaporation. In Separations for Biotechnology; Verrall, M. S., Hudson, M. J., Eds.; Ellis Horwood Limited: Chichester, U.K., 1987; pp 353-359. (8) Slater, C. S.; Hickey, P. J.; Juricic, F. P. Pervaporation of Aqueous Ethanol Mixtures through Poly(dimethyl siloxane) Membranes. Sep. Sci. Technol. 1990, 25, 1063. (9) Hickey, P. J.; Juricic, F. P.; Slater, C. S. The Effect of Process Parameters on the Pervaporation of Alcohols Through Organophilic Membranes. Sep. Sci. Technol. 1992, 27, 843. (10) Molina, J. M.; Vatai, G.; Bekassy-Molnar, E. Comparison of Pervaporation of Different Alcohols from Water on CMG-OM010 and 1060-SULZER Membranes. Desalination 2002, 149, 89. (11) Boddeker, K. W.; Bengtson, G.; Pingel, H.; Siegert, H.; Sufke, T. Treating Synthetic Polymer Dispersions by Organophilic Pervaporation. In Proceedings of the International Conference on Pervaporation Processes in the Chemical Industry, 7th ed.; Bakish Materials Corporation: Englewood, NJ, 1995; pp 333-337. (12) Marin, M.; Hammami, C.; Beaumelle, D. Separation of Volatile Organic Compounds from Aqueous Mixtures by Pervaporation with Multi-Stage Condensation. J. Food Eng. 1996, 28, 225. (13) Escoudier, J. L.; Le Bouar, M.; Moutounet, M.; Jouret, C.; Barillere, J. M. Application and Evaluation of Pervaporation for the Production of Low Alcohol Wines.In Proceedings of the International Conference on Pervaporation Processes in the Chemical Industry, 3rd ed.; Bakish Materials Corporation: Englewood, NJ 1988; pp 387-397. (14) Kaschemekat, J.; Schutt, F.; Wenzlaff, A.; Pervaporation Process of Separating a Liquid Mixture. U.S. Patent 4,900,402, 1990. (15) di Cave, S.; Mazzarotta, B.; Sebastiani, E. Mathematical Model for Process Design and Simulation of Dephlegmators

Ind. Eng. Chem. Res., Vol. 43, No. 1, 2004 183 (Partial Condensers) for Binary Mixtures. Can. J. Chem. Eng. 1987, 65, 559. (16) Rohm, H. J. The Simulation of Steady State Behaviour of the Dephlegmation of Multi-Component Mixed Vapours. Int. J. Heat Mass Transfer 1980, 23, 141. (17) Rohm, H. J. Simulation of the Unsteady State Behaviour of the Dephlegmation of Binary Vapour Mixtures. Lett. Heat Mass Transfer 1978, 5, 307. (18) Jibb, R. J.; Gibbard, I.; Polley, G. T.; Webb, D. R. The Potential for Using Heat Transfer Enhancement in Vent and Reflux Condensers. Unpublished document based on a presentation at Eurotherm Seminar No. 62, Heat Transfer in Condensation

and Evaporation: Application to Industrial and Environmental Processes, Grenoble, France, Nov 17-18, 1998. (19) CRC Handbook of Chemistry and Physics, 61st ed.; Weast, R. C., Ed.; CRC Press: Boca Raton, FL, 1980. (20) Kent, E. R.; Pigford, R. L. Fractionation During Condensation of Vapor Mixtures. AIChE J. 1956, 2, 363.

Received for review July 7, 2003 Revised manuscript received October 28, 2003 Accepted November 6, 2003 IE0305667

Water Mixtures via Fractional

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