Design and Optimization of Hybrid Separation ... - ACS Publications

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Ind. Eng. Chem. Res. 2004, 43, 5248-5259

Design and Optimization of Hybrid Separation Processes for the Dehydration of 2-Propanol and Other Organics Stefan Sommer* Shell Deutschland Oil GmbH, Rheinland Raffinerie, Werk Wesseling, Ludwigshafener Strasse 1, 50389 Wesseling, Germany

Thomas Melin Institut fu¨ r Verfahrenstechnik, RWTH Aachen,Turmstrasse 46, 52056 Aachen, Germany

The dehydration and recycling of industrial solvents often requires complex processes due to azeotropes and pronounced pinch-points. Combined processes consisting of distillation and pervaporation/vapor permeation might overcome these thermodynamic limitations and offer economically attractive alternatives by simplifying process structure, reducing energy consumption, and avoiding entrainers. The technical and economical feasibility of organic solvent dehydration by pervaporation and vapor permeation with inorganic membranes was evaluated. A survey of the process industry revealed the most attractive applications. The separation of 2-propanol and water was studied in detail. For all process calculations the commercial software Aspen Plus was used. Compatible Fortran routines were developed to incorporate membrane units into the program. Different process configurations to integrate pervaporation and vapor permeation with distillation are compared to conventional separation by extractive distillation. A hybrid process design should take advantage of the specific benefits of both separation methods. Often, savings in energy consumption of up to 85% are possible. Investment and operation costs can be reduced by more than 40% in comparison to conventional processes. The most attractive design is distillation and further purification of the top stream by a membrane unit. Here, the economical benefits increase from adiabatic pervaporation over isothermal pervaporation to vapor permeation. A sensitivity analysis of various parameters affecting the operating costs of the plant was carried out. 1. Introduction Solvents are liquid organic compounds used on a large scale in industry in various functions. The coatings industry consumes nearly 50% of the world’s solvent production. Solvents serve to reduce viscosity, dissolve resins and polymers, and disperse color pigments.1 In the production and recycling of solvents, dehydration plays a major role. Solvents or solvent mixtures very often form azeotropes with water, which cannot be separated by simple distillation. Their treatment requires special thermal processing such as two-pressure, azeotropic, or extractive distillation.2 Separation by pervaporation and vapor permeation is almost independent of the vapor-liquid equilibrium, because the transport resistance of the membrane depends on the mobility and sorption equilibrium. Hence, pervaporation and vapor permeation represent attractive alternatives as effective and energy-efficient techniques for the separation of azeotropic and close-boiling mixtures. Combining these processes with other unit operations, such as chemical reaction or distillation, can especially offer economical advantages. These integrated processes allow lower equipment and energy costs in comparison to stand-alone processes. Since distillation is more economical for the bulk of the separation, the membrane is used only to aid the distillation column or to perform * Author to whom correspondence should be addressed. Phone: +49 2236 79 2201. Fax: +49 2236 79 2968. E-mail: [email protected].

the part of the separation where distillation is difficult or impossible. The separation task is switched among the technologies in such a way that each operates in the region of the composition space where it is most effective.3 Until now, the industrial application of polymeric membranes has been limited to dehydration of few solvents because of their insufficient thermal, mechanical, and chemical resistance. Microporous inorganic membranes are generally superior in terms of stability and performance in molecular separations. This fact opens new fields for membrane technology in chemical processing.4,5 In this paper, we evaluate the technical and economical feasibility of inorganic-membrane-based hybrid systems for the dehydration of organic solvents. A mapping of the process industry with regard to solvent production, utilization, and recycling was carried out to examine the most attractive applications. Separation of a 2-propanol (IPA)/water stream was selected as a representative process example. Several configurations combining pervaporation and vapor permeation with distillation have been developed and optimized. The energy saving potential and the investment and operation costs of new hybrid options are compared to conventional state-of-the-art technology. 2. State-of-the-Art Hybrid Systems including Pervaporation or Vapor Permeation Today, pervaporation and vapor permeation are often used stand-alone or at least decoupled from other unit

10.1021/ie034194d CCC: $27.50 © 2004 American Chemical Society Published on Web 07/10/2004

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operations. It is crucial to identify steps in processes where specific advantages of inorganic membranes are beneficial. The best performance and highest benefit of these membrane processes may be gained by direct integration in high-temperature processes. A few examples for potential applications are given in the following list.5 (a) Spray drying: Cycle gas dehydration without low-temperature condensation leads to energy savings by abolishing the cooling and heating energy requirement.6 (b) Dimethyl carbonate (DMC), methyltert-butyl-ether (MTBE) production: Substitution of the second high-pressure distillation column by vapor permeation in the stripping section results in the smallest membrane area and thus best economics.7-10 (c) Combination with reactions: Equilibrium shift in condensation reactions to form esters, acetals, ketals, and resins results in higher production rates.11-15 (d) Direct placement in the core of the process: Product removal in ammonia synthesis avoids cooling and heating of a large recirculation stream.6 A simplified and novel design in the isomerization reaction of C6 and C8 is possible.16-19 A comprehensive review on pervaporation-based hybrid systems covering aspects of process design and economics in various applications, has been published by Lipnizki et al.20 Generic rules and methods for the design of hybrid separation processes have been presented. Pressly et al.21 developed a general framework for process synthesis based on a classification scheme used in the distillation-membrane hybrid systems for the separation of binary mixtures exhibiting a tangent pinch, azeotrope, or low relative volatility. Their breakeven analysis calculates the minimum membrane performance for which the hybrid is competitive and provides an effective screening tool which examines the process alternatives based on a cost comparison and conceptual analysis. Stephan et al.22 introduced a straightforward extension of the McCabe-Thiele methodology to provide general design criteria, to compare different configurations, and to determine the optimal operating conditions. Pettersen et al.23 generated an explicit algebraic design model for vapor permeation systems based on a black-box representation of the transport across the membrane. This model is limited to binary mixtures. With this model they were able to illustrate some of the tradeoffs in a vapor permeationdistillation hybrid process through parametric studies. Here, a classification of the separation of all kinds of liquid mixtures by a combination of pervaporation/vapor permeation and distillation is given according to the different types of membranes, applications, and system configurations. With hydrophilic membranes, such as poly(vinyl alcohol)/polyacrylonitrile (PVA/PAN), polyetherimide (PEI), A-type zeolite and amorphous silica, water can be separated from aqueous-organic feed mixtures. Organophilic membranes, such as PVA/PAN, Y-type zeolite, and methylated amorphous silica, can be used to remove polar components from organic-organic systems. Hydrophobic membranes, such as poly(dimethylsiloxane)(PDMS),polyorthomethylsiloxane(POMS), polyetherblockamide (PEBA), and ZSM-5- or silicalite1-type zeolites, are applied to separate either nonpolar organic mixtures or organics from water. Pervaporation and vapor permeation are characterized by low fluxes with high purity of the permeating component and are usually applied when small amounts of substance are to be removed. Furthermore, due to the necessary reheating of the feed stream in pervapo-

Figure 1. General hybrid process alternatives for combination of distillation with pervaporation or vapor permeation.

ration and the use of chilling units for condensation of the permeate, these are rather costly membrane processes. Thus, when contemplating possible process designs, four major types of membrane configurations (as shown in Figure 1) can be identified: (1) membrane unit in front of the distillation column to split the azeotrope before distillation; (2) pervaporation/vapor permeation for purification of the top or bottom product of the distillation column often combined with azeotrope breaking; (3) installation of the membrane unit between two columns for coarse splitting of the azeotrope enabling the operation of the distillation column at a considerably lower energy leve; and (4) integration of the membrane by side stream processing to reduce the number of theoretical stages or the reflux ratio. In case of a temperature minimum azeotrope, the membrane unit is best used for the treatment of the distillate stream of the column (2). Several studies refer to the most representative application of the dehydration of ethanol,21,23-30 where the distillate is enriched from around 95 wt % ethanol up to 99.5 wt % or higher. In cases where the feed mixture is close to azeotropic composition, pervaporation can be applied to first overcome the azeotrope and then distillation does the final purification (1). This is an option for solvents such as acetonitrile and methyl ethyl ketone with a high volatility beyond the azeotrope, it is a slight option for 2-propanol, and it is not applicable for ethanol. Through reduction of water from 15 to 7 wt % using hydrophilic pervaporation before a distillation column (1), ethanol, iso-butanol, and isoamyl alcohol could be successfully fractionated from fusel oil, which is a byproduct in ethanol distilleries.31 For the dewatering of 2-propanol, either the hybrid process with the membrane unit between two distillation columns (3) or the purification of the azeotropic distillate stream up to the desired product quality (2) are viable options. This separation problem has been widely discussed in the literature.32-36 Steegs et al.37 presented the de-bottlenecking of an existing distillation process in the production of dimethyl formamide. They increased the throughput by about 100% by adding a hollow-fiber membrane pervaporation operated with sweep-gas to process the bottom stream of the vacuum column (2). Bergdorf32 discussed the dehydration of dimethyl acetal with an azeotrope at 94 wt % organic by a pervaporation unit, which was integrated between two distillation columns (3). The membrane-based hybrid process achieves a final

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water concentration of only 500 ppm and enables a more economic dehydration than two-pressure distillation combined with adsorption using activated carbon as polishing step, due to reduced distillation temperature and reflux ratio. With the development of high-performance membranes, separation problems involving pure organic mixtures can also be solved with hybrid processes. Rautenbach et al.38 analyzed an extractive distillation/ pervaporation process for the separation of a 50% cyclohexane/benzene feed mixture into high purity products with the auxiliary furfural. Cost savings of about 20% can be achieved by using the hybrid process with the membrane installed between the two columns instead of the conventional extractive distillation process (3). Shah et al.7 and Rautenbach et al.9 demonstrated the economical potential of a combined distillation/pervaporation process for the production of dimethyl carbonate (DMC) with the removal of methanol by a pervaporation before the distillation (1) or by processing of the distillate stream (2). Detailed design studies revealed that due to the excellent membrane performance for methanol removal, moderate distillate concentrations could be tolerated. This not only eliminated the second energy-extensive high-pressure column, but also reduced investment and energy costs for the first column. Integrated hybrid processes combining distillation and pervaporation with organophilic membranes can also be used in the production of ethyl-tertbutyl-ether (ETBE). Usually the process mixture of a catalytic reaction of ethanol and isobutene is separated in two distillation columns. Streicher et al.8 and Rautenbach et al.9 showed that by treating the bottom product of the first distillation column (1) the excess ethanol can be recycled to the reactor, while a retentate ETBE with less than 1 wt % ethanol is obtained. Having similar investment costs, the hybrid process saves 5060% of the operation costs. Side stream processing (4) is another possible application of pervaporation and vapor permeation within a hybrid process, as examined in detail for the separation of propane and propylene.21,22,39,40 Due to the low relative volatility and accordingly close boiling points, distillation columns with high tray numbers and reflux ratios are required. A side stream processing through a membrane, located close to the pinch point of the column, can decrease the tray number as well as the reflux ratio. Ho¨mmerich et al.10 and Bausa et al.30 have described how the production of the octane enhancer methyl-tert-butyl-ether (MTBE) could be simplified to bring economic benefits. The new process for the complex separation of the reaction cascade’s effluent, which contains methanol, n-C4, and MTBE would include pervaporation at the side draw in the rectifying section and vapor permeation at the side draw in the stripping section. An exergy analysis based on an energy utilization diagram has been carried out by Ishida et al.41 for a combination of a distillation with a pervaporation module on the column head. The analysis showed distinct advantages in energy utilization for this configuration. Humphrey et al.42 confirmed the opportunity for energy savings in separation technology by improvement and replacement of existing distillation plants with efficient hybrid processes.

Figure 2. Vapor-liquid equilibrium (VLE) of different solvent/ water mixtures at atmospheric pressure.46

3. Solvent Dehydration Applications Dehydration of solvents is an essential step in many processes in the chemical, petrochemical, pharmaceutical, electronics, printing, and coating industries. A survey of the process industry has been carried out to examine the most promising dehydration applications for pervaporation or vapor permeation. A great variety of industrial solvents1,43-45 were compared on the basis of the following criteria: (a) scale of production, plant capacity, and market price; (b) azeotropic or close-boiling mixture with water; (c) low or high boiling component of the mixture; (d) evaporation enthalpy and specific heat capacity of the solvent; and (e) miscibility gaps and mutual solubility of solvent and water. The solvent production amount affects the total volume of energy saving. The worldwide plant utilization is a criterion to assess the need for solvent recycling, while the market price is important to calculate economic advantages of recovery instead of replacement by fresh solvent. Using physical property data, the feasibility of conventional distillation processes can be compared to pervaporation. The vapor-liquid equilibrium (VLE) values for various solvent/water systems, given in Figure 2, were calculated according to ref 46. The separation of close-boiling or azeotropic mixtures with relative volatility around one is especially interesting. The favorable zone for hydrophilic membrane separations is marked in Figure 2. The band covers a relative volatility between 0.7 and 1.5. In the interesting region the mixture contains mainly organic solvent contaminated with water. Due to the minimal volatility differences distillation is difficult, yet the minor component is removed through the membrane without problems. In the zone above the band, energy requirements for distillation are high because the major component evaporates. Methanol and acetone require particularly large amounts of energy because of their high specific enthalpy of vaporization. In the zone below the band one finds high boiling organics such as amides or glycols and pervaporation is only competitive in special cases. Stripping or distillation are usually preferred in this area. The heat capacity of the mixture is significant, since there is an axial temperature drop from inlet to outlet in the pervaporation module, and therefore it influences the operation of the membrane process as discussed elsewhere.47 A miscibilty gap between water and the organic solvent can limit the application of pervaporation. In cases with a low mutual solubility, cooling of the mixture and phase separation can be much more


economic. Aromatics and halogenated hydrocarbons, for instance, can be separated in a settler with a purity above 99.8 wt % at 25 °C, so that pervaporation may be advantageous only if lower water concentrations have to be achieved. Nevertheless, the treatment of multicomponent mixtures such as methyl chlorine/methanol/ water can be again of interest. The survey revealed a number of interesting systems. The pervaporation separation performance of the commercially available inorganic membranes was measured for more than 30 systems48 and for the ones with the greatest application potential, the effect of operation conditions was determined.49 The separation of IPA/ water was selected as a representative dewatering process where polymeric membranes are applicable. From the capacity of the more than 100 worldwide installed pervaporation dehydration plants with polymeric membranes, 40% is for IPA, 20% is for ethanol, and the rest is divided into other solvents such as esters, amines, ethers, and ketones. 2-Propanol (IPA) is the most important alcohol, with a mass production of 555 000 tons in 1992 in Europe.1 The chemical reaction yields an azeotropic mixture which needs to be purified. IPA serves as reactant in the synthesis of many organics such as acetone and hydrogen peroxide. In the coatings and cosmetics industry IPA is applied as a solvent, in medical applications it is used as a disinfectant, and in the food industry it is used for the de-hydrogenation of sugar and gelatin. For its recycling, a dehydration treatment is necessary. 4. Solvent Dehydration of 2-Propanol The evaluation of different process options for the dehydration of 2-propanol (IPA) covers conventional thermal separation and several membrane-based hybrid configurations. The study is based on equal boundary conditions in all designs for a process stream of 1875 kg/h with an initial water concentration of 20 wt % water in the alcohol. Particularly, the same recovery and purity had to be achieved with the integrated novel processes as with the standard technology. The different alternatives were designed according to state-of-the-art knowledge and optimized. All data are normalized to the amount of dehydrated product and given in costs or energy per ton of IPA. The process design for the separation of other azeotropic solvent/water mixtures is similar. Nevertheless, every example is specific and the best configuration depends on the thermodynamic properties of the system and the specific requirements of the industrial environment. 4.1 Conventional Process. In general, simple distillation is applied for dehydration of solvents unless an azeotropic mixture is formed. The VLE for water/IPA at atmospheric pressure is given as a dashed line in Figure 3, showing a temperature minimum azeotrope at 12 wt % water and a close boiling mixture at lower water content.46 The separation characteristic of the inorganic silica membrane, obtained at a feed temperature of 80 °C and a permeate pressure of 20 mbar,48,49 is included for comparison as a solid line in Figure 3. The separation of azeotropic mixtures requires complex and cost intensive special unit operations such as two-pressure-distillation, azeotropic distillation, or extractive distillation. Close boiling mixtures show a low separation factor and as a consequence in distillation a great number of theoretical stages or a high reflux ratio is needed. This makes such conventional processes less

Figure 3. Vapor-liquid equilibrium (dotted line) and silica membrane separation characteristic (solid line) for the system IPA and water.46,48,49

Figure 4. Flow sheet of the extractive distillation for the separation of IPA and water.34

economical. The principle of the two-pressure process is to move the concentration of the azeotrope by a change of pressure. Normally, the reflux ratios are too high to be economically viable. When applying azeotropic distillation, di-isopropyl-ether is added as auxiliary agent. State-of-the-art technology for the dehydration of IPA is extractive distillation with ethylene glycol as the entrainer since it has several advantages in comparison to the azeotropic distillation. With the high boiling auxiliary agent, the activity of water in the azeotropic mixture is reduced, thus leading to an increase of the separation factor.2 The main disadvantages of the conventional process are product impurities due to the use of an entrainer, which are often hazardous to the environment, and the need of a third component to separate the binary IPA/ water mixture, which entails the effort to clean and recycle this auxiliary agent. Hence, investment and energy costs are high because of the operation of three distillation columns with high reflux ratios and recirculation streams. Figure 4 shows the flow sheet of a continuous operating extractive distillation process for the dehydration of IPA. The IPA/water feed mixture is nearly concentrated up to the azeotrope in the first column. The stream removed on the top is slightly below the azeotropic composition. At the bottom, the lower volatile water is withdrawn. In the second column the azeotrope is split. Therefore the auxiliary ethylene glycol (EG) is added almost on the top to extract the water. At the column top, the dehydrated product IPA is removed. At the bottom, a mixture of ethylene glycol and water is withdrawn. In the third column, the separation of EG/ water is performed to recycle the auxiliary to the second column. At the top, almost-pure water is removed and directed together with the water from the first column to a receiving water course. Because of small losses, a


Figure 5. Pervaporation for azeotrope splitting between two distillation columns (A, left), separation up to the product specification (B, center), or as stand-alone process (C, right).

makeup stream is added to the EG gained at the bottom before it is pumped to column two. The product purity of IPA is about 99.91 wt % with small amounts of water (0.09 wt %) and traces of EG. The recovery of IPA from feed to product is 99.98 wt %. The first distillation column has 19 stages and a reflux ratio of 0.637. The extraction column with 58 trays is operated with a reflux ratio of 0.5. The vacuum column has 25 stages and a reflux ratio of 1.5. In the following paragraphs temperature, pressure, and composition of three columns are reported in detail. The IPA/water feed mixture (1875 kg/h, 80 wt % IPA) is preheated in a heat exchanger from 25 to 80 °C. The first column is operated under slight overpressure. The temperature profile is from 103 to 80.25 °C. The mass flow rate of water (99.99 wt %) removed at the bottom is 125 kg/h. The IPA/water mixture withdrawn on top is 1750 kg/h (86 wt % IPA). The evaporator is heated with 1414 kg/h steam at 4 bar and 143.6 °C. The second column is operated under atmospheric pressure. The temperature range is from 169.1 to 82.4 °C. The amount of EG added at tray 6 is almost 5000 kg/h at a temperature of 98.3 °C. The product IPA gained at the top is about 1500 kg/h with a concentration of 99.91 wt %. The 5250 kg/h EG/water mixture (5 wt % water) is pumped from the bottom to the third column. For the heating at1402 kg/h, 16 bar steam at 201 °C is needed. The third column is operated under vacuum with a pressure of 400 mbar to reduce the boiling temperature of EG from 194 to 171.2 °C, which is the bottom temperature. At the top the temperature is 75.8 °C where about 251 kg/h water with a concentration of 99.98 wt % is removed. This water is fed, along with the water from the first column, to a sewage plant. From the bottom, a stream of 5000 kg/h is withdrawn with an amount of 99.97 wt % EG and recompressed to 1.013 bar. The heat exchanger runs with 1181 kg/h, 16 bar steam at 201 °C. 4.2 Alternative Hybrid Process Configurations. Possible hybrid alternatives to the conventional process are combinations of pervaporation and vapor permeation with distillation where the greatest economic potential can be realized. These integrated processes allow lower equipment and energy costs in comparison to stand-alone processes. Due to the low achievable fluxes, which decrease with decreasing feed concentrations, a pervaporation unit should be operated in such a way that the amount being separated by the membrane is as small as possible and withdrawn on the highest concentration level. Pervaporation and vapor permeation offer the chance to improve the performance

of an existing column and to minimize additional separation steps. According to the thermodynamic properties of the system, there are various possible system configurations. For the separation of a binary mixture of IPA and water, three competing alternatives can be identified which are shown in Figure 5. Coarse splitting of the azeotrope with the membrane unit installed between two distillation columns (A, left) and purification of the azeotropic distillate stream up to the desired product quality (B, center) are the two hybrid process options, while the third is a stand-alone pervaporation membrane process (C, right). In case of the dehydration of multicomponent solvent mixtures, there are many more options for system configurations depending on the thermodynamic behavior of the system. For instance, in a ternary mixture with water as a medium boiling component and one high and one low boiling organic solvent, the water can be withdrawn by side-stream processing while the pure solvents are removed at the top and bottom of the distillation column. This has been demonstrated by Kuppinger et al.50 for a process stream of methanol, IPA, and water. They compared an azeotropic distillation with toluene as entrainer to a combination of distillation and side stream processing with a water/methanol selective A-type zeolite membrane. 4.3 Membrane Separation Characteristics. Standard PVA/PAN polymeric membranes are state-of-theart in solvent dehydration today. They exhibit with the system IPA/water a total permeate flux of 0.8 kg/m2h and permeate concentration of 97.3 wt % water at a feed concentration of 10 wt % water, a feed temperature of 80 °C, and a permeate pressure of 20 mbar. This permeate purity corresponds to a selectivity of about 300.14,51-53 Utilizing the same operation parameters for the silica membrane, a total permeate flux of 2.8 kg/ m2h with a water concentration of 98.6 wt % in the permeate and a process selectivity around 800 was observed in pervaporation experiments.48,49 As the specific module costs of inorganic membranes are higher than those of polymeric membranes, the achievable permeate fluxes need to be higher to balance these investment costs by membrane area savings. The measurements showed that the separation characteristics are indeed 3.5 times better under equal boundary conditions. Furthermore, the flux rates increase exponentially with the feed temperature, which is limited to 100 °C53 in the case of polymeric materials while amorphous silica is stable up to at least 200 °C.54 Hence, the higher specific module costs in comparison to those

Ind. Eng. Chem. Res., Vol. 43, No. 17, 2004 5253 Table 1. Investment Costs Blocks of the Main Positions in the Distillation and Membrane Unit distillation unit

membrane unit

packing material collector/distributor unit column jacket heat exchangers pumps

modules and membranes heat exchangers vacuum vessels vacuum and feed pumps chilling unit

unhindered permeate flow (i.e., neglection of flow patterns), neglection of pressure drops, and polarization effects have been made. Figure 6. Interaction of optimization targets and parameters in the process development.

of polymeric membranes can be overcompensated by minimizing the membrane area at elevated temperatures. 5. Process Development with Flow-Sheet Simulation Programs From a design point of view, numerous different optimization targets, shown in Figure 6, can be identified when investigating hybrid processes. This ranges from locally optimized solutions for the single unit (reflux ratio, number of theoretical stages and column pressure for distillation, number of modules, feed inlet temperature, temperature drop per module (in the case of pervaporation), permeate pressure, and membrane choice for pervaporation/vapor permeation) to global optima for the complete process (interface concentrations and flow rates, choice between pervaporation and vapor permeation). The necessary calculation effort is increased by numerous system configurations and recycle streams. Shortcut methods such as the minimum area method developed and employed by Stephan et al.,22 Bausa et al.,30 and Moganti et al.39 can provide an idea about the rough design of a combined process. However, this technique is only applicable to binary systems. For systems with more than two compounds or for determining the tradeoff between plant design (particularly optimized number of modules) and operation costs, suitable computer programs are required. All our results have been calculated with the commercial program Aspen Plus. This modular-based software includes various conventional unit operations (distillation, mixers, reactors, condensers, etc.) which can be combined into any flow scheme. Membrane processes are not included in the standard version, but Aspen Plus offers the possibility of programming userdefined Fortran routines, thus allowing the implementation of any additional unit operation. The pervaporation routine usrpv has been developed at our department facilitating the design and simulation of complete pervaporation units in Aspen Plus.55 The design calculations in usrpv are performed by solving the differential mass and energy balances. The local mass transport in the membrane unit is incorporated by a regression analysis empirical model in the program. This analysis is based on technical scale experiments with real feed mixtures and membrane systems. Permeate flux and composition are calculated according to experimentally determined functions of feed composition and temperature.49 Furthermore, additional assumptions such as decoupled permeate fluxes, 100% module efficiency,

6. Cost Calculation The comparison of the various hybrid process configurations with polymeric and inorganic membranes to the conventional separation by extractive distillation has to be made in subsequent consideration of the resulting specific investment and operation costs. Tables 1 and 2 give an overview of the most important cost data used for the presented calculations. The total costs are obtained by adding operating and capital costs. For the estimation of investment cost at an early stage of process design the proven factor method developed by Lang56,57 was used. The costs for the main items of equipment are determined and the sum is multiplied by a factor to account for other cost items which depend on the complexity of the plant. In the case of distillation this factor is 4.74, whereas for the pervaporation unit it is 2.27. The breakdown of the single investment cost items of the distillation and membrane unit is given in Table 1. The investment costs include membranes and modules (installed costs are 3000.00 euros/m2 for the inorganic shel-and-tube version and 1000.00 euros/m2 for the polymeric plate-and-frame type), heat exchangers (reheating in case of pervaporation), vacuum vessels (for polymeric plate-and-frame modules), vacuum and feed pumps, chilling unit for the membrane unit and packing material, collector/distributor units, column jacket, heat exchangers, pumps, and, if necessary, entrainer for the distillation unit. Based on the investment costs the capital costs are obtained using a capital factor of 0.2638 resulting from a depreciation time of 5 years and an interest rate of 10%. The operation costs have been calculated according to the cost functions given in Table 2 for an equipment availability of 8000 h/a. Yearly maintenance and labor costs have been estimated to be 5% of the capital costs. The operation costs consist of membrane replacement (25% of the installed module costs in case of polymeric flat-sheet membranes and 33% in case of tubular inorganic), permeate condensation, reheating of feed stream, energy consumption for the membrane unit and energy costs for bottom re-boiling and condensation at the column top, electricity consumption of the pumps, makeup for ethylene glycol losses in case of the extraction for the distillation unit, plus COD costs for wastewater treatment in a sewage plant. It should be mentioned that the advantage of a direct cost comparison between all process alternatives and the disadvantage of the inaccuracy of the cost estimate (20-30% in the early project state) have to be weighted carefully. Particularly, the energy costs might vary significantly for different production sites. Nevertheless, cost estimation is a very suitable tool for a preliminary selection of the best process combination. Crucial mis-

5254 Ind. Eng. Chem. Res., Vol. 43, No. 17, 2004 Table 2. Cost Functions for the Determination of Operation Costs in the Distillation and Membrane Unit34,58 energy source

basic price Pb

consumption-specific price Pc

heating steam (6 bar) 12 500 euro/(t/h)a 9 euro/t heating steam (16 bar) 30 000 euro/(t/h)a 10 euro/t cooling water (25 to 35 °C) 70 euro/kWa 0.05 euro/m3 refrigerant (-15 °C) 0.0075 euro/kWh electricity 0.06 euro/kWh ˘ steam‚(Pb + c‚Pc); Kwater ) V˙ water‚Pc‚c; Krefrige ) Q˙ cond,permeate‚(Pb + c‚Pc) Ksteam ) m annual plant availability c 8000 h/yr membrane service lifetime 2 yr membrane replacement costs 25-33% of membrane installation costs

Figure 7. Effect of inter-stage heating on feed temperature and permeate flux over the silica membrane unit in option B. Adiabatic (solid line) and isothermal (dotted line) operation.

Figure 8. Influence of the interface concentration on the distillation (dotted line), pervaporation (dashed line), and total (solid line) separation cost.

takes in basic engineering can be avoided and the influence of important process parameters can be analyzed.

sation costs decrease with increasing permeate pressure as less and less deep cooling energy is required. A sudden drop in condensation costs occurs around 100 mbar due to the fact that cooling water can be used instead of chilling systems. The effects of different process parameters are often contrary. This is shown for the interface concentrations and flow rates of hybrid process B in Figure 8. It can be seen that the closer the distillation gets to the azeotrope (88.2 wt % IPA), the more expensive the distillation will be. On the other hand, at lower IPA concentrations more water has to permeate through the membranes and the costs for the pervaporation process will be higher. The steps in the membrane separation costs occur when the number of necessary modules changes. The operating temperature has always to be chosen as high as possible. In the case of polymeric membranes, the stability of the PVA/PAN material is limited to 100 °C. In the case of inorganic material, silica itself has a higher thermal resistance of up to 240 °C, but the process design pressure level of 10 or 16 bar sets the maximum to 140 °C. A membrane area of 50 m2 is required in option A, 125 m2 is required in option B, and 150 m2 is required in option C with the silica membranes for the dehydration of an IPA/water stream of 1875 kg/h from 80 wt % up to 99.9 wt % product purity with a solvent recovery of 99.98%. In comparison, the area with polymeric membranes is 625 m2 for option B and five times higher than with inorganic ones. As a consequence, the three times higher specific costs of the inorganic membrane modules are more than compensated for by the membrane area reduction due to better performance. By operation with vapor permeation or with the new isothermal Pervap SMS modules for tubular inorganic membranes47,59,60 instead of standard adiabatic pervaporation, the total membrane area in option B is even further reduced to 100 m2. An overview on the optimal design parameters in each of the three

7. Design and Optimization of the Membrane-Based Process Alternatives Essential for an economic use is the detailed optimization both of the individual process steps and the whole hybrid system. One important aspect in pervaporation is the temperature drop per module. The necessary heat for the evaporation of the permeating water is taken from the latent heat of the feed stream, which cools along its way through the module. The total membrane area decreases with the number of modules but costs for piping and intermediate heating increase. The interstage heating effect is shown in Figure 7 for option B with amorphous silica membranes. The total membrane area of 125 m2 is divided into five modules. The temperature decline in the first stages is around 40 K due to the high fluxes. The flux is decreasing exponentially with temperature leading to a deceleration of the effect. Between the exit of the first module and the entrance of the second, the flux is differing with temperature almost by a factor of 4, whereas between module two and three the difference is only a factor of 2. In the later stages at lower water concentration and thus fluxes, the axial temperature drop is negligible, whereas in the beginning additional inter-stage heating would lead to a reasonable reduction of required membrane area and costs. Different sizes of modules have been investigated ranging from 1, 5, and 10 to 25 m2. Due to decreasing specific membrane and module costs with increasing size (0.6 rule), the biggest module is the cheapest process design option.54 Another design parameter is the pressure level for condensation of the permeate vapor. With higher permeate pressure, the membrane area increases as the driving force for transport decreases. However, conden-


Figure 9. Development of the water concentration in the feed (dotted line) and permeate (average, solid line; local, dashed line) with the membrane area in option B with inorganic membranes. Table 3. Design Parameters for the Various Membrane Process Options process configuration [-]

feed concn [%-wt.]

retentate concn [%-wt.]

inlet feed temp [°C]

membrane area [m2]

A Silica PV B Silica PV B Silica VP B Polymeric PV C Silica PV

87.0 84.5 84.5 84.5 80.0

96.0 99.9 99.9 99.9 99.9

140 140 140 100 140

50 125 100 625 150

different membrane process configurations with polymeric or inorganic membranes in pervaporation and vapor permeation is provided in Table 3. Figure 9 shows the development of the feedwater concentration as a function of the membrane area in the inorganic membrane plant option B. More than half of the membrane area of 125 m2 is needed to do the fine dehydration below 1 wt % to the final product specification. At these very low water contents, the permeate flux is rather low and the permeate contains even more alcohol than water. But the average overall water concentration of the permeate is still better than 85 wt %t. The membrane area necessary in option B can be reduced with option A where distillation does the final purification work. It depends on the application whether the larger membrane area or the additional column is the more cost-effective option. If the mixture is close boiling at water contents below the azeotrope, a membrane process operates better than a distillation and option B will show superior performance. Option A becomes more favorable when very low final water concentrations in the ppm region have to be achieved. In this case, there is hardly any driving force for permeation in the end. Hence, the required membrane area increases dramatically making distillation more feasible even at low separation factors. The stand-alone process C becomes more attractive the lower the initial water content is, or when no infrastructure for distillation is available on the production site. 8. Economic Evaluation and Energy Efficiency First of all, the separation costs for the conventional IPA dehydration process are compared with three membrane cases of pervaporation: with ceramic silica membranes between the two distillation columns (A); on top of a distillation column to break the azeotrope and reach the final product specification (B); and as a

Figure 10. Comparison of conventional extractive distillation with the three process options with silica membranes in terms of capital and operation cost.

stand-alone process (C). An overview of specific separation costs in terms of capital and operating costs for the different options is shown in Figure 10. The current market price for the produced IPA is about 500.00 euros/t. The total separation costs of the extractive distillation are 84.06 euros/t. The most efficient membrane configuration, option B with ceramic membranes, is, at 48.18 euros/t, about 42.7% cheaper than the conventional option. The stand-alone process C based upon silica membranes is, at 57.63 euros/, more expensive than hybrid process B, because more membrane area is needed to achieve the product specification, but it is still about 31.4% cheaper than the conventional process. Option A with the membrane unit between the two distillation columns is, at 60.13 euros/t and savings of 28.5%, in the same order of magnitude with a slightly cheaper membrane part and a more expensive distillation due to an extra column. Additionally, the use of polymeric and ceramic membranes in standard adiabatic pervaporation, isothermal pervaporation with Pervap SMS modules, and vapor permeation with silica membranes have been evaluated for the best process configuration B. The economical comparison of the different membranes and operation modes in the same system design is illustrated in Figure 11. The application of polymeric membranes has overall specific costs of 53.68 euros/t and is 11.4% more expensive than the option with inorganic membranes. As already mentioned, this difference is caused by the total membrane area needed in the separation along with its price. The tubular inorganic modules cost 3000 euros/ m2 installed for 125 m2 and for the polymeric plate-andframe design 1000 euros/m2 is charged for the necessary 625 m2. In continuous processes, isothermal operation saves a lot of membrane area compared to adiabatic operation with interstage reheating after each module. Due to the absence of an axial temperature decline and the exponential temperature dependence of the flux, the permeation rates are on a much higher level with constant feed temperatures. There are two ways to realize the isothermal processsby modules with integrated direct heating as the Pervap SMS from Sulzer Chemtech47,59,60 or the use of vapor permeation instead of pervaporation. The isothermal pervaporation has total separation costs of 43.19 euros/t for the dehydration of IPA and is

5256 Ind. Eng. Chem. Res., Vol. 43, No. 17, 2004 Table 5. Energy Consumption for the Distillation Unit in Option B distillation unit

share [%]

annual cost [euros/yr]

top-condensation bottom-reboiling energy consumption pumps

18.7 80.4 0.8

18 292.00 78 549.00 821.00

Table 6. Energy Consumption for the Membrane Unit in Option B

Figure 11. Evaluation of different membranes (PVA/PAN and silica) and operation modes (adiabatic and isothermal pervaporation and vapor permeation) in the same system configuration in terms of investment and operation cost. Table 4. Energy Consumption for the Conventional Separation Process 4-bar steam 16-bar steam electricity cooling water

amount

annual costs [euros/yr]

1.42 t/h 2.58 t/h 20.3 kW 225.6 m3

135 912.00 307 092.00 7 209.00 88 586.00

10.3% cheaper than the standard pervaporation with silica membranes. The reduced costs result from a 20% smaller membrane area and dispensable heat exchangers for the reheating of the feed stream between each stage. In exchange, the specific module costs rise a bit due to costly integration of heat supply by steam into the module construction. The most profitable membrane configuration in combination with distillation and top stream processing of the column is vapor permeation. This option leads to processing costs of 41.79 euros/t, which is even 13.2% below the reference option B with silica membrane pervaporation. The distillation should be operated at the same pressure as the membrane unit to avoid the use of compressors for feed pressure adjustment. The top condenser can be designed smaller as dephlegmator with only partial condensation to create the necessary reflux and transfer of the remaining vapor to the membrane module. In addition to the economical competitiveness, the membrane-based hybrid processes offer some more advantages compared to the conventional technology: (a) purer product, as no auxiliary chemical is needed for azeotrope splitting; (b) simplified process layout by replacement of one or two distillation columns; (c) less organic waste by avoiding the use of additional chemicals in the separation process and by a higher extent of solvent recycling; and (d) minimized energy consumption and therefore lower emissions of CO2, SO2, and NOx. The specific energy requirement for the conventional IPA/water separation costs 45.41 euros per ton of product. The fractions of the total demand are cooling water 16.4%, 4-bar steam 25.2%, 16-bar steam 57%, and electricity 1.3%. The total annual amounts of energy spent and corresponding costs are listed in Table 4. With the three membrane process options, different amounts of energy savings in comparison to extractive

membrane unit

share [%]


permeate condensation reheating of feed stream energy consumption pumps

40.6 52.2 7.3

19 692.00 25 324.00 3 528.00

distillation can be realized. Option A has a lower energy consumption by 50%, option B is lower by 70%, and the stand-alone process option C is lower by 84%. The specific energy requirement for the IPA/water separation in the hybrid case B costs 13.51 euros/t. The consumption of the distillation unit makes up for 9.46 euros/t and of the membrane unit costs 4.05 euros/t. The total annual shares of energy and corresponding costs are summarized in Tables 5 and 6. The conventional process of extractive distillation uses about 3.3 kWh/kg of IPA produced, whereas the hybrid pervaporation and distillation process B requires about 1 kWh/kg for a dehydration from 20 to 0.1 wt % water. Considering the production of 2-propanol in Europe, which is about 555 000 t/year, the energy savings in Europe could be up to 2 PJ/year. These energy savings only account for the production of IPA. In the process industry much IPA is used and recycled regularly to dehydrate alcohol. This recycle stream to be dehydrated is probably much larger than the amount dehydrated in the production process. The size of this recycle stream is not known but a rough guess indicates that the energy savings in Europe in this IPA recycle stream could be as high as 10 PJ/year. 9. Full-Scale Process Design and Cost Breakdown A detailed overview of the design parameters of the membrane-based hybrid alternative B is given in Table 7 for the distillation unit and in Table 8 for the pervaporation unit. The process has been designed for an industrial plant with equal size to compare the novel process to existing technology. The plant capacity is designed for dehydration of an IPA stream with 20 wt % water to produce 1500 kg/h of the pure alcohol. The cost structure of the extractive distillation process is given in Table 9. The entrainer makeup and COD costs, 1227 and 4908 euros, respectively, are almost negligible because 99% of the operating costs are energy expenses, which are listed in Table 4. The cost structure of the hybrid separation process option B with silica membrane is given in Table 10. An analysis of cost structure of the silica-membrane-based hybrid pervaporation process B reveals that the total costs of 48.18 euros/t consists of 28.24 euros/t capital and 19.93 euros/t operation costs. The distillation unit part costs 16.46 euros/t and the membrane unit costs 31.72 euros/t. In the distillation part column and equipment make up for 59 024 euros and heat exchanger and pumps are 24 957 euros, which are 30 and 13% of the overall distillation cost, respectively, whereas energy is almost half of the total cost. In the membrane

Ind. Eng. Chem. Res., Vol. 43, No. 17, 2004 5257 Table 7. Design Parameters and Process Streams of the Distillation Unit in the Full-scale Hybrid Separation Configuration B unit design parameter energy re-boiler kW condenser kW properties pressure bar temperature °C design number of theoretical stages reflux ratio evaporation ratio fraction capacity process streams feed flow rate kg/h water concentration wt % inlet temperature °C back flow rate kg/h water concentration wt % temperature °C bottom flow rate kg/h water concentration wt % temperature °C top flow rate kg/h water concentration wt % temperature °C

Table 10. Cost Structure for the Hybrid Separation Process B with Silica Membrane Pervaporation

number

525 520 1.013-1.016 80-100 8 0.12 2.25 75%

hybrid process capital operation total distillation unit capital operation total membrane unit capital operation total


share total cost [%]

specific cost [euros/tIPA]

338 939.00 239 208.00 578 147.00

58.6 41.4 100.0

28.24 19.93 48.18

83 981.00 113 579.00 197 560.00

14.5 19.6 34.2

7.00 9.46 16.46

254 958.00 125 629.00 380 587.00

44.1 21.7 65.8

21.25 10.47 31.72

1875 20.0 80.1 320 85.8 15.6 375 99.94 99.9 1820 85.5 79.9

Table 8. Design Parameters and Process Streams of the Pervaporation Unit in the Full-Scale Hybrid Separation Configuration B unit

number

design parameter energy reheating condensation modules number area arrangement permeate vessels condensation temperature permeate pressure

kW kW

170 215

m2

5 25

°C mbar

1 15.6 20

process streams feed flow rate kg/h water concentration wt % inlet temperature °C product flow rate kg/h water concentration wt % temperature °C permeate flow rate kg/h water concentration wt % temperature °C

1820 15.5 140 1500 0.1 25 320 85.5 15.6

Table 9. Cost Structure of the Conventional Extractive Distillation Process in the Separation of IPA/Water

capital operation total

annual costs [euros/yr]

specific costs [euros/tIPA]

462 719.00 544 935.00 1 007 654.00

38.56 45.41 83.97

part, investment costs are 67%, maintenance and labor costs account for 20%, and operating costs are only 13% of the total costs. Membranes and modules are the biggest part of these cost blocks. About 41% of the initial

Figure 12. Influence of the cost development of membranes and modules (a) and their long-term stability (b) on the process economics. Dotted line is the conventional process, solid line is the novel hybrid process.

investment and even 63% of the maintenance are for membrane replacement. Vacuum system, pumps and heat exchangers are only 15% and 10%, respectively. 10. Parameter Sensitivity Analysis for the Separation Costs In general, costs are never fixed and always depend on supply and demand on the market. Hence, every cost estimation depends on the future development of prices. Inorganic membranes for pervaporation are still a very recent technology offering large possibilities for price reductions by cheaper production and increasing capacity. On the other hand, the expected stability is higher than with polymeric membranes. The effects of a variation of these factors are shown in Figure 12 for the discussed dehydration of IPA. Despite the high specific costs of inorganic membranes, the process option B is


isothermal pervaporation. Adiabatic pervaporation is the most expensive. A sensitivity analysis of various parameters affecting the operating costs of the plant showed the influence of future price developments. Acknowledgment We gratefully acknowledge financial support by the European Commission under the contracts JOE3-CT970074 and ENK6-CT1999-00015. We are very thankful to ECN, Petten, Netherlands for providing the membranes used in this study. Literature Cited

Figure 13. Effect of variable energy cost with the development of specific prices for cooling (a) and heating (b). Dotted line is the conventional process, solid line is the new hybrid process.

feasible with amorphous silica membranes even if the price would be twice as high. In all previous calculations a membrane lifetime of 2 years was estimated. The break-even is at a membrane stability of only six month. Energy costs also have an influence on the cost efficiency of the separation process. Due to the small amount of permeate and the small fraction of the top condensation, cooling does not have a great impact whereas heating energy has. As specific energy costs are unlikely to decline due to shortage of resources, the membrane process will probably get more economic in the future as can be seen in Figure 13. Conclusions The viability of pervaporation and vapor permeation with inorganic membranes in combined processes for the dehydration of organic solvents could be successfully demonstrated. The most attractive applications were determined in a survey of the process industry and 2-propanol was selected as a representative case. The commercial software package Aspen Plus was combined with a self-developed membrane routine in Fortran providing a flexible and efficient tool for process design calculations. The design and optimization studies indicated that economical benefits can be derived from membrane-based hybrid processes. The best configuration is the purification of the azeotropic distillate stream of the column up to the desired product quality. Investment and operation costs could be reduced by more than 40% and energy savings up to 85% could be demonstrated. Inorganic membranes allow 11% lower total separation costs in comparison to polymeric materials. Among the membrane-based options vapor permeation combined with distillation should be preferred over

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Received for review October 21, 2003 Revised manuscript received May 27, 2004 Accepted June 2, 2004 IE034194D

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