Comparing Pervaporation and Vapor Permeation Hybrid Distillation

Feb 2, 2005 - Previous studies have shown that hybrid distillation processes using either pervaporation or vapor permeation can be very attractive for...
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Ind. Eng. Chem. Res. 2005, 44, 5259-5266

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Comparing Pervaporation and Vapor Permeation Hybrid Distillation Processes J. Fontalvo,† P. Cuellar,† J. M. K. Timmer,‡ M. A. G. Vorstman,† J. G. Wijers,† and J. T. F. Keurentjes*,† Process Development Group, Department of Chemical Engineering and Chemistry, Eindhoven University of Technology, P.O. Box 513, 5600 MB Eindhoven, The Netherlands, and ETD&C BV, P.O. Box 8111, 6710 AC Ede, The Netherlands

Previous studies have shown that hybrid distillation processes using either pervaporation or vapor permeation can be very attractive for the separation of mixtures. In this paper, a comparison between these two hybrid processes has been made. A tool has been presented that can assist designers and engineers to decide which process is more convenient for a specific application. Water removal from acetonitrile has been used as an example. A hybrid process with vapor permeation is preferred when the membrane is used either for water removal at high water concentration or just for overcoming the azeotropic composition. When the membrane removes water at water concentrations lower than the azeotropic point, pervaporation is more effective. Recycling part of the product as permeate (product sweep) and applying different pressures in the distillation columns and the membrane unit strongly reduce the required membrane area and the total cost of the process. Relatively low membrane selectivities are required for an economically optimal hybrid membrane-distillation process. Introduction The chemical industry is looking for economical ways to reduce energy demand and to avoid auxiliary agents for the separation of mixtures. Pervaporation (PV) and vapor permeation (VP) are highly suitable to separate close boiling mixtures1,2 and in breaking azeotropic mixtures where distillation is energy intensive or requires the use of an auxiliary agent. Distillation is a well-known operation with lower capital cost than membrane operations. Hybrid membrane-distillation processes exploit the advantages of distillation and membrane operations, while overcoming the disadvantages of both.3 Previous studies have shown the advantages of distillation processes in combination with either pervaporation4-7 or vapor permeation8,9 for various mixtures. Strategies for the optimization of hybrid processes have been developed in the past.1,5,10 However, the comparison between the various types of hybrid processes on the same calculation basis can help designers and engineers to decide which hybrid process is more convenient for a specific application. This work presents an economical and technological comparison between distillationsvapor permeation and distillationspervaporation hybrid processes using silica membranes. Silica membranes are of interest for the dehydration of several solvents because they are resistant to harsh conditions,11,12 high temperatures, and high pressures. As a model case the dewatering of acetonitrile (ACN) is studied. The model assumptions for the simulations are described followed by results for conventional distillation based processes and hybrid processes using pervaporation or vapor permeation. Product sweep, high membrane oper* To whom correspondence should be addressed. E-mail: [email protected]. † Eindhoven University of Technology. ‡ ETD&C BV.

ating pressures combined with low distillation pressures and relatively low selectivities are included to reduce the total cost of a hybrid process. This study shows the influence of these parameters on the required membrane area, total energy consumption, and total cost of the hybrid process. Guidelines for the selection of PV or VP are suggested followed by an economical evaluation. Process Modeling Conventional and hybrid processes were simulated for the separation of 16.000 ton/year of 50 wt % ACN-water mixture at 1 atm and 333 K. The purity of the two product streams is set to 99.9 wt %. ASPEN PLUS 11.1 (AspenTech) was used for the simulations of the distillation tower and other basic equipment in connection with a rigorous subroutine (written by us in MATLAB, The Mathworks Inc.) describing the behavior of the membrane unit including concentration and temperature polarization. The size of the columns, the reboiler and condenser duties, and the water composition in the column and in the outlet streams were calculated using the subroutine RADFRAC in ASPEN PLUS 11.1 (AspenTech). For the membrane unit the outlet temperatures and final water concentrations, retentate total pressure drop and total membrane area were calculated using the MATLAB subroutine. The economic evaluation was performed according to Appendix A. The pervaporation and vapor permeation units consist of a set of membrane modules (and heat exchangers in the case of pervaporation) connected in series. Each membrane module has several ceramic tubes mounted in a shell like a shell and tube heat exchanger. The silica separation layer is located on the inner side of the 7-mm-internal diameter tube. Liquid in pervaporation or vapor in vapor permeation is fed to the inside of the tubes, and reduced pressure is applied at the shell side. The developed MATLAB program allows calculating a series of pervaporation units where the retentate temperature drop or the membrane area per module is

10.1021/ie049225z CCC: $30.25 © 2005 American Chemical Society Published on Web 02/02/2005

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Table 1. Total Flux and Separation Factor for the System ACN-Water Using Silica Membranes water wt % temperature, K permeate pressure, mbar total flux, kg/ m2 h separation factor, γ limiting operating temperature, K limiting pressure (module), bar

9 343 17 3.9 331 448 50

fixed. Cocurrent or countercurrent operation can be simulated, but only results for cocurrent operation between permeate and retentate flows will be presented. The software accounts for local concentration and temperature polarization in the retentate by linking the mass transfer in the retentate (eq 1) with the mass and heat flux through the membrane (eqs 2 and 3). The mass transfer in the retentate (eq 1) was calculated using a Maxwell-Stefan equation expressed in difference form.13 Conventional correlations for flow inside pipes were used for the evaluation of the local mass and heat transfer coefficients. Fluxes through the membrane were computed using eq 2 where the flux for water11 and ACN are proportional to its vapor pressure difference between retentate and permeate. Activity coefficients (γi) were obtained by UNIFAC, and the Poynting factor was taken into account to calculate the vapor pressures of each compound in the retentate

( )

-xi

∆γi γ ji

)

∑ i*j

xjjNi - xjiNj ki,jC h

Ni ) Γi(xIiγIiPvIi - yiPG) q ) hL(TL - TI) + R)

∑i NiHh Li ) ∑i NiHh Gi

(ywI/yACNI) (xwI/xACNI)

Figure 1. Pressure swing process for dehydration of ACN.

(1)

The feed temperature of the PV process was taken as the bubble temperature of the mixture at the feed pressure and composition while for VP it corresponds to the dew temperature. For vapor permeation an overheating of 5 °C is used in the feed stream. The MATLAB subroutine evaluates whether a change of phase occurs in the retentate inside the membrane modules; if so, the temperature is modified accordingly. The retentate temperature drops along a PV unit as a consequence of the vaporization and expansion of the different components at the permeate side. The heat supply necessary to perform an ideal isothermal operation was calculated. The total cost for an actual nonisothermal PV operation, using a series of interstage heat exchangers, is presented in the economic evaluation for comparison with the ideal operation. The retentate temperature drop in VP was neglected because no latent heat is involved in the transport of the components through the membrane.

(2)

Conventional Distillation-Based Processes

(3)

(4)

For VP the concentration and temperature polarization were neglected because of the low mass and heat transfer resistance in the vapor phase. Analogous to pervaporation the flux was calculated proportional to the difference in vapor pressures. The partial vapor pressures in the retentate were calculated by means of fugacity coefficients using the Peng Robinson EOS. The permeate pressure was set at 0.11 atm for both pervaporation and vapor permeation processes. At this pressure the permeate stream can be condensed using cooling water. Values for the water permeance (Γwater) and the membrane separation factor (R) defined in eq 4 were, as a first assumption, taken as constant in the simulations. These values were calculated from typical data supplied by Pervatech BV (Table 1). To reduce the total cost of the PV process the feed flow rate per membrane tube was economically optimized. Concentration and temperature polarization and as a consequence capital cost increase as the feed flow rates decrease, whereas increasing the flow rate will increase the pumping cost. For vapor permeation no optimization of the feed flow was performed because concentration and temperature polarization are much less significant than for pervaporation processes. The feed flow rate per tube was set to obtain a retentate pressure drop lower than 10 mbar.

Distillation-based processes for the separation of ACN-water mixtures have to be designed to overcome the azeotropic composition. This mixture presents a pressure-dependent low-boiling azeotrope, which is close to 30 molar % water at atmospheric pressure. Two conventional processes are described here, i.e., pressure swing distillation and azeotropic distillation. The pressure swing distillation process is presented in Figure 1. This process consists of two distillation columns operating at different pressures. The azeotropic composition is shifted to higher water concentrations as the total pressure increases. As a consequence, the two columns together can operate in the whole range of concentrations between the water-rich stream and the ACN-rich stream. Figure 1 shows that both the feed flow to the second distillation column and the recycle stream flow rate are high. This is due to the high water concentration in the azeotrope at the top of the first and second distillation column. Russell14 patented an azeotropic distillation process using benzene as an entrainer for the separation of ACN-water mixtures; see Figure 2. Water and ACN form a ternary azeotrope with benzene. Upon condensation of the ternary vapor, phase separation occurs into a benzene-free phase and an ACN-rich phase. The ACNrich phase is fed to the second distillation column. The first distillation column produces a water-rich stream and the second distillation column produces an ACNrich stream. The amount of benzene involved in the process is 350 kg/ton of feed. This amount can be calculated by minimizing the reboiler duties as a function of the benzene flow rate. In the economical evaluation section these processes will be compared with the hybrid processes.

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Figure 2. Azeotropic distillation process for dehydration of ACN using benzene as entrainer.

Figure 3. Column-pervaporation-column (CPC) hybrid process for dewatering of ACN.

Membrane-Distillation Hybrid Processes Two different hybrid membrane-distillation processes are of interest for this kind of azeotropic mixture in which either distillation or the membrane unit is used as the final dewatering step.15 The structural and parametric optimization of these hybrid processes is relatively straightforward because the top composition (and feed to the membrane) is limited by the thermodynamic conditions.10 It has been suggested in the literature that the permeate flow should be minimized.16 We have found that in these processes the membrane unit should be operated in such a way that the driving force is as high as possible in order to reduce the required membrane area and thus the capital cost. High driving forces are accomplished by using high feed temperatures or by using a product sweep. The required membrane area and energy consumption are analyzed as a function of the final retentate concentration, membrane feed temperature, and membrane feed pressure when distillation is used as the final dewatering step. A similar analysis is presented for product sweep and the separation factor when the membrane is used as the final dewatering step. Because qualitatively similar results are expected for PV and VP only results for hybrid PV-distillation processes are presented in this section. Distillation as the Final Dewatering Step. Simulations have been performed for processes where pervaporation is used for splitting the azeotropic composition in the hybrid pervaporation-distillation process (Figure 3). The first distillation column produces a water-rich stream and an azeotropic top stream. The liquid top stream is pressurized, heated, and fed to a pervaporation membrane that splits the azeotropic composition. Permeate is condensed, pumped, and recycled to the first distillation column, and the retentate is treated further in a second distillation column. The second distillation column will produce an ACN-rich stream and an azeotropic top stream that is recycled to the pervaporation unit.

Figure 4. Influence of the final retentate water concentration from the pervaporation unit on the membrane area and reboiler duty. Operating pressure of the PV and distillations columns are 5 and 1 atm, respectively. No product sweep applied. Total energy duty considers the energy supplied to the pervaporation unit and reboilers.

Final Retentate Concentration. The final retentate concentration after the pervaporation module is a design variable. Figure 4 shows the effect of the final retentate water concentration on reboiler duty and required membrane area for a pressure of the feed to the pervaporation unit of 5 atm and a pressure in the columns of 1 atm. Low final water concentrations in the retentate will require large membrane areas and low reboiler duties for the second distillation column. The energy duty required for the first distillation column is approximately constant while the total energy consumption is reduced when the final retentate water concentration decreases resulting in an economically optimal final retentate concentration. Liquid Feed Temperature and Pressure. To reduce the required membrane area, the driving force for the water transport through the membrane can be increased using high retentate temperatures. To avoid vaporization of liquid feed, high pressures are required. Under these conditions the required membrane area decreases linearly with pressure when the feed is at the bubble temperature. The main effect of high temperatures is reducing the required membrane area although a small increase in total energy consumption is involved. High feed temperatures to the membrane unit, but lower than the maximum allowed operation temperature as given in Table 1 (448 K), were used in the calculations of all the hybrid membrane-distillation processes. For vapor permeation the maximum operation temperature is obtained at lower pressures than for pervaporation. Membrane as a Final Dewatering Step (Column Membrane CM). In this type of system the membrane unit is used for splitting the azeotropic composition and as a polishing step. Permeate is recycled to the distillation column and retentate is the ACN-rich product stream in the configuration given in Figure 5a. High temperatures (430 and 448 K) and operating pressures (10 and 15 atm) are necessary in the membrane unit in order to reduce the required membrane area. Low operating pressures in the distillation column, however, are advantageous for reducing the water concentration at the column top stream and so the amount of water to be removed by the membrane unit. Table 2 shows temperatures, pressures, and flow rates for the calculations performed for the most effective flow diagram (Figure 5b). The influence of product sweep and membrane separation factor on the energy consumption and required membrane area are analyzed below.

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Figure 5. Column-pervaporation (CP) hybrid process for dewatering of ACN. (a) Schematic representation. (b) Flow diagram for process P2b in Figure 10. See Table 2 for details. Table 2. Stream Properties for the Flow Diagram of Figure 5B, Process P2b in Figure 10 temperature, C pressure, atm vapor fraction mass flow, kg/h mass fractions water ACN

1

2

3

4

5

6

7

8

9

10

11

12

13

14

43.6 0.3 0.0 2000

43.8 0.3 0.0 1208

46.5 0.3 0.0 208

69.2 0.3 0.0 1000

46.0 15.0 0.0 1208

23.5 0.3 0.0 208

23.5 0.1 0.0 208

175.7 15.0 0.0 1208

163.3 0.1 1.0 208

23.1 0.1 0.0 20

175.6 15.0 0.0 1020

23.1 0.1 0.5 20

175.6 15.0 0.0 20

175.6 15.0 0.0 1000

0.500 0.500

0.134 0.866

0.773 0.227

0.999 0.001

0.134 0.866

0.773 0.227

0.773 0.227

0.134 0.866

0.773 0.227

0.001 0.999

0.001 0.999

0.001 0.999

0.001 0.999

0.001 0.999

Product Sweep. When using product sweep (PS) a fraction of the retentate stream is vaporized and recycled at reduced pressure to the permeate side of the pervaporation module (Figure 5). Product sweep decreases the water concentration at the permeate side of the pervaporation membrane. This reduction in water content increases the driving force for the water transport, while the driving force for the ACN transport is approximately constant due to the high ACN concentration in the retentate side. A PS of around 20% of the retentate flow rate results in an important reduction of membrane area while the total energy consumption, including reboiler and pervaporation duties, stays low (Figure 6). At low PS a strong reduction in required membrane area is obtained but further membrane area reductions are almost negligible for high PS (>20%) where the reboiler duty strongly increases. Product sweep has only little influence when the membrane is used just to overcome the azeotropic composition due to the high water driving force available along the membrane. Membrane Separation Factor. The presented results have been calculated with a constant membrane separation factor of 331. However, the separation factor changes with temperature and concentration. For instance, it has been shown17,18 that higher selectivities occur at higher temperatures for different waterorganic mixtures. Simulations using higher and lower

Figure 6. Effect of product sweep on the membrane area required and reboiler duty. Column pressure 0.3 atm. Retentate pressure 10 atm. Total energy duty considers the energy supplied to the pervaporation unit and reboiler.

separation factors than 331 are considered here. A difference to product sweep the effect of selectivity on hybrid processes using either pervaporation or vapor permeation has been shown in the literature8,15 for several mixtures. A low membrane separation factor results in a low water content at the permeate side. This “self-sweeping” increases the permeate stream and the reboiler duty in the distillation column but reduces the required membrane area, but very low separation factors require high membrane areas due to the high recycle flow rate to the distillation column and to the membrane (Figure 7). On

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Figure 7. Effect of membrane separation factor on the required membrane area and reboiler duty. Column pressure 0.3 atm. Retentate pressure 15 atm. Product sweep 2%.

the other hand, high separation factors produce low driving forces, increasing the membrane area but slightly reducing the total energy consumption. At separation factors of around 100, membrane area and total supplied energy (reboiler and pervaporation duties) are low. Fahmy et al.8 also found that membranes with very high selectivities are not desirable when high purity of the retentate is required using VP. For an economically optimal operation, a higher membrane separation factor requires more product sweep. On the other hand, if the permeate stream is used as the product,15 high selectivities are more convenient. Summarizing, product sweep and relatively low separation factors have a positive effect in a hybrid membrane-distillation process reducing the required membrane area in combination with low energy consumption. High operating pressures in the membrane units and low operating pressures in the distillation columns reduce the required membrane area if the membrane is used as the polishing step. Guidelines for Selecting Hybrid Pervaporation-Distillation or Hybrid Vapor Permeation-Distillation Processes The feed composition to the membrane unit in the process presented in Figures 3 and 5 corresponds to the top stream composition from the distillation columns. This composition is the same for hybrid processes using either pervaporation or vapor permeation. Figure 8 presents the driving force available for pervaporation and vapor permeation at 1 and 10 atm at different feedwater concentrations. VP has a higher driving force for water removal than PV in the range of water concentrations higher than the azeotropic point. In the range of water concentrations lower than the azeotropic point the opposite occurs and the difference in driving force increases in favor of PV as the pressure increases. At the azeotropic point the driving forces are equivalent for PV and VP. The difference in water driving force for these membrane processes (Figure 8) is used as a tool for selecting either (or both in some hybrid systems) pervaporation or vapor permeation for a hybrid process. The qualitative concepts introduced in this section are discussed quantitatively in the economic evaluation. Three general cases are discussed below. The membrane unit is used for water removal at high water concentrations and at low water concentrations or for overcoming the azeotropic composition. Vapor permeation is more efficient for water removal at concentrations higher than the azeotropic point

Figure 8. Water driving force for pervaporation and vapor permeation as function of the water concentration at 1 and 10 atm. Driving force calculated as the difference in water partial pressure between permeate and retentate at bubble point (PV) and dew point (VP).

because it requires less membrane area than pervaporation and the energy consumption is also lower due to the availability of vapor from the distillation column.9,19 For water removal at low water concentrations, pervaporation is preferable, especially at high pressures where it requires a smaller membrane and lower energy consumption than VP. In a hybrid VP process three design alternatives can be identified which are all economically less favorable: A partial condenser could be used followed for a compressor (in this case the compression cost is high), the column top stream is condensed, pumped, and vaporized (energetically more expensive than PV because the whole column top stream has to be vaporized), and a high operating pressure could be applied in the distillation column in combination with a partial condenser (under these conditions the water concentration at the top of the distillation column is shifted at high values increasing the amount of water to be removed). In conclusion, when pervaporation is used as the final dewatering step at low water concentrations and high pressures the saving on required membrane area and energy consumption is higher compared to VP. However, a hybrid PV process needs only a total condenser after the distillation tower and a liquid pump to increase the pressure of the column top stream (Figure 5). As a consequence, total and capital costs of the hybrid PV-distillation process are lower than for a hybrid VP-distillation process. The costs of hybrid processes, where VP is used as a final dewatering step, are not included in the economic evaluation. When the membrane is used just for splitting the azeotropic composition, vapor permeation is preferable. In this situation pervaporation and vapor permeation require similar membrane areas where high pressures are applied at the distillation towers and the membranes units. VP is superior to PV because vapor is available at the top of the distillation columns resulting in a lower energy consumption with similar capital costs (see Economical Evaluation). The membrane driving force at several pressures has been presented as a tool for selecting pervaporation or vapor permeation in a hybrid process. This tool has been used for dehydration of ACN, but it can be applied to systems with or without azeotrope depending on which

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Figure 9. Total cost as function of membrane separation factor for the system C (0.3 atm) P (15 atm) and 2% product sweep.

component is preferentially transported through the membrane. For instance, if water is the component with the lower boiling point in a nonazeotropic mixture, vapor permeation using a water-selective membrane is the best option for a hybrid process. Economical Evaluation. Required membrane area and energy consumption have been presented in Figure 7 as a function of the separation factor. Figure 9 shows that the total cost of the hybrid process reduces when the separation factor increases reaching a constant value for a separation factor of 100. In general an optimal separation factor exists with a minimum total cost. Capital, operating, and total costs are presented in Figure 10 for the hybrid processes described above. Capital and total costs are reduced when the operating pressure is increased as shown for hybrid processes at 2, 5, and 10 atm using pervaporation (P1a, P1b, and P1c, respectively) and for hybrid processes at 5 and 10 atm using vapor permeation (V1a and V1b, respectively) where distillation is used as the final dewatering step. In this type of configuration the total and energy costs for VP are slightly lower than for PV at the same pressure, 5 and 10 atm (P1b vs V1a, P1c vs V1b). The cost for a hybrid pervaporation-distillation process using a membrane unit with interstage heating, where the membrane modules and heat exchangers are connected in series (P3), is presented in Figure 10. Similar operating costs are obtained compared to an ideal isothermal membrane operation (P2b). If the number of interstage heat exchangers is optimized, the capital cost of the process P3 is 10% higher than for an ideal operation (P2b). It is found that the membrane

area using interstage heating is around 1.34 times the theoretically minimal membrane area (isothermal operation), which is similar to the value suggested by Bausa et al.4 between 1.2 and 1.3. Hybrid processes show a total cost reduction between 25% and 60% in comparison with conventional processes (C1, C2). Capital costs of hybrid processes are higher than for conventional processes, but the energy consumption and the operating costs are lower. Capital costs for hybrid PV processes are lower (P2a, P2b, P3) than for conventional processes (C1) when pervaporation is used as the final dewatering step (Figure 10). Processes where the membrane is used just to overcome the azeotropic composition are more economical using VP than PV (P1b vs V1a, P1c vs V1b). Hybrid processes using a pervaporation unit as a polishing step (P2b, P3) show the lowest total cost. Discussion and Conclusions The membrane driving force at several pressures has been used as a tool for the economic selection of pervaporation or vapor permeation in hybrid processes. It has been applied for dewatering of ACN where for water removal at high water concentration and just for overcoming the azeotropic composition vapor permeation is more efficient. For water removal at water concentrations lower than the azeotropic point, pervaporation is the best option. The combination of different operating pressures in the membrane unit and in the distillation column has proved to be cost efficient in the design of hybrid processes. A low operating pressure in the distillation column reduces the azeotropic water concentration and a high pressure in the pervaporation unit reduces the required membrane area by increasing the driving force for the water transport through the membrane. Product sweep or the use of a membrane with relatively low separation factor also increases the driving force, reducing the water content in the permeate side of the membrane unit. In general, hybrid processes using either pervaporation or vapor permeation are economically more favorable than conventional process for the separation of ACN-water mixtures with an obtained reduction between 25% and 60% of the total cost of the conventional processes. The lower total cost is mainly due to the lower energy consumption of the hybrid processes.

Figure 10. Comparison of total, capital, and operating costs between conventional distillation-based processes and hybrid membranedistillation processes for dewatering of ACN.

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Acknowledgment The financial support by the Program Office on Economy, Ecology and Technology from the Dutch Ministries of Economic affairs, Education and Environmental affairs under contract EETK20046 is kindly acknowledged. Nomenclature ACOND ) condenser area (m2) BM ) bare module cost estimation (3.57 20 awh ) annual working hours, 8000 (h/year) C ) molar concentration (kmol/m3) C h ) average molar concentration between on the surface and in the liquid bulk (kmol/m3). Cp ) thermal heat capacity (kJ/kmol K) CCAPITAL ) capital cost (Euro) CCOND ) condenser cost (Euro) CD ) distillation column cost, including reboiler and condenser (Euro) CFP ) feed pump cost (Euro) CHE ) interstage heat exchangers cost (Euro) CMEM ) membrane cost (Euro) CMOD ) module cost (Euro) COPERATING ) operating cost (Euro/year) CTOTAL ) treatment cost (Euro/year) CU ) utility cost (Euro/year) CVP ) vacuum pump cost (Euro) Ftotal ) volumetric liquid flow (m3/s) gT )total flow in the permeate (kmol/h) hL ) heat transfer coefficient (kJ/h m2 K) HG ) permeate enthalpy (kJ/h) H h i ) partial molar enthalpy for component i (kJ/kmol) HL ) retentate enthalpy (kJ/h) ∆HG ) enthalpy change from outlet permeate temperature from module to 25 °C (kJ/h) ∆HLj ) enthalpy change in the retentate in the heat exchanger for the stage j (kJ/h) ki,j ) mass transfer coefficient between component i and j in the mixture (m/h) kr ) heat capacity ratio, 1.33 LMTD ) logarithmic media temperature difference (K) mr ) membrane replacement period in Table 3 (years) MSxx ) Marshall and Swift index values for the year 19xx MW ) molecular weight (kg/kmol) Ni ) flux of component i through the pervaporation membrane (kmol/h m2) P ) pressure (bar) PDIS ) discharge pressure, 1.01325 × 102 (kPa) Pvi ) vapor pressure of the component i Po ) pressure at the standard condition, 1.01325 × 102 (kPa) POP ) permeate operating pressure (kPa) ∆P ) pressure drop for the retentate (Pa) pc ) power cost, 0.09 (Euro/kW h) q ) heat flux through the membrane (kJ/h m2) R ) gas law constant, 8.31 m3 (kPa/kmol K) rt ) ratio Euro to Dollar as 1.137 T ) temperature (K) ∆TCOND ) temperature difference for cooling water in the condenser, 10 K Tintake ) absolute temperature of vapor at intake conditions, 298 K To ) temperature at the standard condition, 273 K U ) Overlall heat transfer coefficient in the condenser, 1800 kJ/h m2 K wc ) cooling water cost, 5 × 10-5 (Euro/kg) ws ) steam cost 18 × 10-3 (Euro/kg) WCOND ) condenser power consumption (Euro/h) WFP ) feed pump power consumption (kW) WVP ) vacuum pump consumption (kW)

Table 3. A1 Membrane Replacement Periods Used in the Economic Evaluation membrane replacement period (mr), years

operating temperature, °C

6 4 2

150 160 175

x ) molar fraction in the retentate xj ) average molar fraction between on the surface and in the liquid bulk y ) molar fraction in the permeate WHE ) heat exchangers power consumption (Euro/h) WR ) reboilers power consumption (Euro/h) Greek Variables λw ) heat of vaporization for water (kJ/kg) FP ) pump efficiency, 0.6 Γi ) permeance of component i (kmol/h m2 bar) R ) [(ywI/yACNI)]/[(xwI/xACNI)] separation factor for the system ACN-water γ ) activity coefficient ∆γ ) activity coefficient difference between the surface membrane and the bulk liquid Subscripts I ) interfacial on the membrane L ) liquid bulk phase or retentate G ) gas bulk phase or permeate i ) component i r ) reference w ) water

Appendix A The capital costs of the distillation columns, condensers, and reboilers were calculated using the program Aspen Icarus Process Evaluator 12. The total cost was calculated using the method employed by Oliveira et al.21 and Ji et al.,22 where the total cost (CTOTAL) is based on the following components: capital depreciation and taxes, annual maintenance and labor requirements, module replacement, and energy consumption. Depreciation and taxes have been assumed to be 15% of capital costs (eq A1). Annual maintenance and labor have been taken to be 10% of the capital cost (eq A1).

CTOTAL ) 0.15CCAPITAL + 0.10CCAPITAL + COPERATING (A1) CCAPITAL ) BM(CMEMB + CMOD + CFP + CVP + CCOND + CHE + CD) (A2) COPERATING ) CU + (1/mr)(CMEM)

(A3)

The operating cost (eq A3) takes into account the utility costs (CU) and membrane replacement. The estimated replacement period is presented in Table 3 for several operating temperatures. No supervision, laboratory, or insurance costs have been included.A bare module cost estimation technique20 was used in which the capital cost is estimated to be 3.57 times the cost of the major equipment (membranes, modules, feed pump, vacuum pump, condenser, interstage heat exchangers). This factor (BM) includes piping, valves, instrumentation, and peripheral equipment as well as engineering site preparations and other installation costs.

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The membrane cost was provided by Pervatech as 1400 Euro/m2 and the module cost as 2000 Euro/m2. The feed pump cost was calculated according to ref 21 as follows:

where NS is the number of interstage heat exchangers.

CFP )

Literature Cited

rt

(

∆P MS03 × 15nmod series2.3Ftotal MS68 nmod series

)

0.52

×.93 (A4)

where rt is the ratio Euro to Dollar and MS are the Marshall and Swift index values, 1115 for 2003 and 280 for 1968. The vacuum pump cost and the condenser cost were calculated according to ref 22 and presented in eqs A5 and A6. A similar approach that was used for the condenser was used for evaluating the capital cost of the interstage heat exchangers but using steam at 16 bar (CHE).

(

)

MS03 60gTRT0 CVP ) 4200 rt MS94 3600P0

0.55

(A5)

MS03 CCOND ) rt (1176.7 + 128.1ACOND) MS94

(A6)

ACOND ) ∆HG/ULMTD

(A7)

where

The utility cost (eq A8) was calculated for the feed pump, the vacuum pump, condensers, reboilers, and interstage heat exchangers. The cost for electrical power is 0.09 Euro/kW h, for steam at 16 bar is 18 Euro/ton, and for cooling water is 0.05 Euro/ton with a maximum temperature raise of 10 K for the cooling water.

CU ) awh[pc(WFP +WVP) + WCOND +WHE + WR] (A8) where awh are the annual working hours and pc is the power cost. The power consumption for the feed pump (WFP) was calculated according to21

WFP )

(

1 ∆P F FP total 1000

)

(A9)

The power consumption for the vacuum pump (WVP) and for the condenser (WCOND) was calculated according to ref 22 by means of eq A10. This equation involves the total permeate flow. Because the permeate stream is condensed before the vacuum pump, the power consumption is overpredicted but the influence in the total cost is small.

WVP ) 10%

[( )

gT PDIS kr RT 3600 intakekr - 1 POP

((kr - 1)/kr)

]

-1

(A10)

Water-cooling consumption and steam consumption were calculated as follows:

wc∆HGMWw WCOND ) ∆TCONDCpW

(A11)

WHE ) ws

NS ∆H

∑ j)1

Lj

λw

(A12)

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Received for review August 24, 2004 Revised manuscript received October 25, 2004 Accepted October 26, 2004 IE049225Z