Synthesis of Distillation Processes Using Thermodynamic Models and

ature or pressure, the search for azeotropes begins. For example, the condition for ..... indefinite status (“wildcard character”) and is deter- m...
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Ind. Eng. Chem. Res. 1998, 37, 3112-3123

Synthesis of Distillation Processes Using Thermodynamic Models and the Dortmund Data Bank Ju 1 rgen Gmehling* and Christian Mo1 llmann† Carl von Ossietzky Universita¨ t Oldenburg, Technische Chemie (FB 9), Postfach 2503, D-26111 Oldenburg, Germany

A new and comprehensive procedure for process synthesis for binary and ternary distillation processes with an emphasis on entrainer selection is presented. This procedure consists of two complementary methods for entrainer selection and a variety of graphical tools for the representation of various mixture properties of the given system. The methods for entrainer selection are based on (a) the concept of thermodynamic models, such as group contribution methods (e.g., modified UNIFAC), and (b) the access to a large computerized factual data bank with experimental data on activity coefficients at infinite dilution and azeotropes. Explicit criteria for the selection of solvents for azeotropic and extractive distillation are reported. With the help of residue curves and contour lines it is shown how further information for a comparison and advanced selection of entrainers can be obtained. Introduction Distillation is by far the most important industrial separation process because it offers several advantages. In distillation processes energy is used to create a system of fluid phases (liquid, vapor) that can easily be transported. Especially in continuous distillation, an intensive contact between the liquid and vapor phases as well as a simple separation of the phases due to their different densities can be realized in a countercurrent process. These features facilitate the realization of a multistage process. Finally, distillation is based on vapor-liquid equilibrium (VLE) which can reliably be predicted by modern thermodynamic models, such as gE models (Wilson, NRTL, UNIQUAC (e.g., Prausnitz et al., 1986)), equations of state (e.g., Soave, 1972; Peng and Robinson, 1976), and even by group contribution methods, for example, the UNIFAC methods (Fredenslund et al., 1975, 1977; Hansen et al., 1991; Gmehling et al., 1993). The rank of distillation processes is underlined by the fact that the worldwide throughput of distillation columns (oil refineries, chemical and petrochemical industry, natural gas processing) is approximately 5.3 × 109 tons/year (Porter, 1995) and that, e.g., in the United States nearly 3% of the total energy consumption is needed for running about 40 000 distillation columns (Humphrey and Seibert, 1992). In distillation, only the difference R12 - 1 can be exploited for separation purposes. For systems with separation factors near unity the number of stages required for a separation into the pure components increases enormously. In the case of azeotropic systems (R12 ) 1) a separation by ordinary distillation becomes impossible even with an infinite number of stages (Gmehling and Brehm, 1996). However, to maintain the advantages offered by distillation for azeotropic systems as well, several processes, for example, azeotropic and * Author to whom correspondence is addressed. E-mail: [email protected]. Phone: ++49-4417983831. Fax: ++49-441-7983330. † Present address: AspenTech Europe S.A./N.V., Rue Colonel Bourgstraat 127-129, B-1140 Brussels, Belgium.

extractive distillation, pressure-swing distillation with and without the addition of a separating agent, and hybrid processes (distillation, e.g., combined with crystallization or membrane techniques), have been developed. In azeotropic and extractive distillation a solvent (entrainer) is used besides energy to enable or facilitate the separation of systems with disadvantageous separation factors (0.9 < R12 < 1.1). The selection of suitable solvents has been the subject of numerous investigations (e.g., Laroche et al., 1991, 1992; Simmrock et al., 1991; Stichlmair et al., 1989; Wahnschafft and Westerberg, 1993). The new conceptual design package (Figure 1) presented here specifically addresses the synthesis of distillation processes for systems with disadvantageous separation factors. This means that separation problems which arise from the VLE behavior of a system can be recognized, for example, by calculating all binary, ternary, etc., azeotropic points of multicomponent systems, and strategies to overcome these difficulties can be developed. As a part of the DDBSP program package (Gmehling et al., 1996), the new process synthesis module provides information for the conceptual design of new processes as well as for the analysis of established processes (retrofit problems) and makes extensive use of graphical tools for the synthesis and design of azeotropic and extractive distillation. The entire system has been used successfully in test runs and for consulting purposes. Calculation of Azeotropic Data Process synthesis for distillation processes is based on the reliable knowledge on azeotropic data. Thus, the prediction of azeotropic points using thermodynamic models is a very important step within many procedures, such as the selection of suitable solvents, the calculation of residue curve maps, etc. The calculation of azeotropic data is actually performed by two separate programs: one for the calculation of homogeneous azeotropic points and the other for the calculation of heterogeneous azeotropic points. The input for each of the programs comprises the components, the system pressure for isobaric, and the system temperature for

S0888-5885(97)00782-3 CCC: $15.00 © 1998 American Chemical Society Published on Web 05/23/1998

Ind. Eng. Chem. Res., Vol. 37, No. 8, 1998 3113

Figure 1. Menu for the conceptual design package for the synthesis of distillation processes as a part of DDBSP (Gmehling et al., 1996).

isothermal calculations, respectively. Furthermore, it is the thermodynamic model and the estimation methods for the pure-component vapor pressures and the fugacity coefficients that need to be specified. The output covers data on pressure, temperature, and azeotropic composition (provided that an azeotrope has been detected) and a characterization of the type of azeotrope, for example, homogeneous or heterogeneous, temperature minimum or maximum, saddle point. Homogeneous Azeotropes. The program discussed here is capable of predicting homogeneous azeotropic data for multicomponent mixtures, whereby the real behavior of the vapor phase can be taken into account by the chemical theory (e.g., for carboxylic acids) or the virial equation. The calculations are performed for all binary and higher systems. A binary system exhibits a homogeneous azeotrope, if for any composition the separation factor becomes equal to unity:

R12 )

K1 y1/x1 γ1P1s ) ) )1 K2 y2/x2 γ P s

(1)

2 2

During the azeotropic search, the program actually minimizes the difference between the equilibrium compositions of the vapor and liquid phases. At the azeotropic point, which represents one point of the vaporliquid equilibrium, not only the compositions but also the chemical potentials of the two phases are equal, and thus no further separation can be achieved by either evaporation or condensation as part of an ordinary distillation. Heterogeneous Azeotropes. Binary and ternary heterogeneous azeotropic data are calculated by a separate program. Once the vapor-liquid-liquid equilibrium (VLLE) has been calculated for a given temperature or pressure, the search for azeotropes begins. For example, the condition for binary heteroazeotropic behavior is fulfilled if the calculated equilibrium composition of the vapor phase is situated within the region of liquid-liquid immiscibility, represented by the two liquid compositions xi′ and xi′′. The search for ternary azeotropes is a stepwise procedure in order to investigate the ternary concentration region systematically and usually starts from a binary heterogeneous subsystem. The objective of this search is to find the pressure maximum on the vapor-pressure surface of the ternary system. A detailed description of the current procedure

for the calculation of both homogeneous and heterogeneous azeotropic data is given by Gmehling et al. (1994). An example for azeotropic behavior in multicomponent mixtures is shown in Table 1. The modified UNIFAC (Dortmund) method was applied to calculate the homogeneous quinary system acetone-chloroformmethanol-ethanol-benzene and all its quaternary, ternary, and binary subsystems. Note that this system exhibits six binary and four ternary homogeneous azeotropes, some of which represent temperature maxima or temperature minima while others represent saddle points. All other binary and higher systems show zeotropic behavior. The experimental azeotropic data for each of the systems which are also given in Table 1 enable a direct comparison between experimental and calculated data. The data indicate a good agreement between the UNIFAC predictions and the experimental data from various authors (Gmehling et al., 1994). The ability to directly compare experimental with calculated data is a unique and important tool not only to identify separation problems but also to distinguish correct from erroneous predictions and thus to test the reliability of various thermodynamic models. Residue Curves and Contour Lines Residue Curves. Processes such as azeotropic or extractive distillation are usually based on the addition of a separating agent to at least one binary system with a disadvantageous separation factor (0.9 < R12 < 1.1). For this reason ternary diagrams, i.e., triangular diagrams, are frequently used to represent ternary composition profiles or mass balances around unit operations. Thus, ternary diagrams are well suited for the graphical representation of feasible separations and product composition regions. For “simple” distillation column configurations with one feed, two-outlet (for distillate and bottom product) residue curves can be used to describe internal concentration profiles at infinite reflux ratio. The starting and end points of a residue curve then represent the composition of the distillate and the bottom product, respectively. At the same time both compositions together with the feed composition have to fulfill the mass balance constraint as well, which means that all three compositions are located on the straight material balance line. Residue curves can be calculated by numerical integration of

3114 Ind. Eng. Chem. Res., Vol. 37, No. 8, 1998 Table 1. Predicted (Modified UNIFAC) and Experimental Azeotropic Data (DDB) for the Quinary System Acetone (1)-Chloroform (2)-Methanol (3)-Ethanol (4)-Benzene (5) and All Its Subsystems at Atmospheric Pressure experimentala

predicted system

type of azeotrope

1-2 1-3 1-4 1-5 2-3 2-4 2-5 3-4 3-5 4-5 1-2-3 1-2-4 1-2-5 1-3-4 1-3-5 1-4-5 2-3-4 2-3-5 2-4-5 3-4-5 1-2-3-4 1-2-3-5 1-2-4-5 1-3-4-5 2-3-4-5 1-2-3-4-5

homPmin homPmax none none homPmax homPmax none none homPmax homPmax homPsad homPsad none none none none none homPsad homPsad none none none none none none none

a

t, °C

y1,az

y2,az

64.24 55.31

0.3562 0.7867

0.6438

53.72 59.89 58.10 67.96 57.50 63.14

58.66 68.66

y3,az

y4,az

0.2133 0.6496 0.8636

0.3504 0.1364 0.6150 0.4637

0.3359 0.3448

y5,az

0.2240 0.4820

0.4401

0.1024 0.1244

0.6034

0.3850 0.5363

0.1732

0.4383

0.2942 0.4373

type of azeotrope homPmin homPmax none none homPmax homPmax none none homPmax homPmax homPsad homPsad none none none none none homPsad n.a.b none n.a. n.a. n.a. n.a. n.a. none

t, °C

y1,az

y2,az

64.42 55.56

0.352 0.790

0.648

53.49 59.36 58.00 67.93 57.68 63.30

y3,az

y4,az

0.210 0.651 0.842

0.349 0.158 0.610 0.448

0.319 0.345

y5,az

0.237 0.465

0.390 0.552

0.444 0.190

59.30

Mean values of the data stored in the Dortmund Data Bank. b n.a. ) not available.

dxi(ξ) ) xi(ξ) - yi(x b(ξ)) dξ

i ) 1, ..., n - 1

(2)

where ξ is a dimensionless measure of time. A so-called residue curve map comprises a collection of residue curves within a ternary diagram. As an example, Figure 2 shows the residue curve diagram of the system benzene-2-propanol-water at atmospheric pressure. It can be seen that due to the presence of azeotropes these curves can have different starting points and end points, and the resulting distillation regions are bordered by residue curve boundaries. At infinite reflux ratio a column with a given number of stages is usually expected to have its maximum performance in terms of separation. At the same time, the infinite reflux limiting case can be used to approximate very high but finite reflux ratios. For these reasons residue curves are used within this context to represent column profiles at infinite reflux ratio in distillation processes and to locate product composition regions for a separation. Contour Lines. Within this program system we define contour lines as a series of liquid-phase compositions (xi) that common constant properties can be assigned to. Again, triangular diagrams are useful to display the course of these lines for any ternary system. For the synthesis of distillation processes, contour lines with constant vapor-liquid equilibrium related properties are of particular interest. Table 2 shows the thermodynamic properties that have been defined as objective functions in the algorithm so far. The algorithm itself is based on a spline interpolation among equally distanced points of the VLE for a given ternary mixture. Note that these VLE data points have been calculated rigorously, i.e., by means of gE models or group contribution methods, before. Using the parameters of the spline equation, it is then possible to invert the phase equilibrium calculation, for example,

Figure 2. Residue curve diagram for the system benzene (1)-2propanol (2)-water (3) at atmospheric pressure: thick line, residue curve boundaries; thin line, residual lines; - -, binodal curve; - - -, liquid-liquid equilibrium tie-lines; b, binary and ternary azeotropic points. Table 2. Objective Functions (Constant Properties) To Be Used for the Calculation of Contour Lines P ) constant (isobars) T ) constant (isotherms)

Ki )

yi γiPoyiφisPis ) ) constant xi φ VP

i ) 1, ..., n

i

Rij )

Ki ) constant Kj

i ) 1, ..., n j ) 1, ..., n j*i

F ) |R12 - 1| + |R13 - 1| + |R23 - 1| ) constant

to calculate the liquid mole fractions (provided one xi is fixed in a ternary system) that belong to a given bubble point temperature (objective function). This procedure applies to homogeneous as well as to heterogeneous

Ind. Eng. Chem. Res., Vol. 37, No. 8, 1998 3115

Figure 4. Isothermal curves (t ) 65, 68, 70, and 75 °C) for the system ethanol (1)-water (2)-benzene (3) at atmospheric pressure: b, binary and ternary azeotropic points.

Figure 3. Contour lines for the ternary system acetone (1)schloroform (2)-methanol (3): (a) isotherms (calculated at atmospheric pressure); (b) F ) |R12 - 1| + |R13 - 1| + |R23 - 1| ) constant. b: binary and ternary azeotropic points.

systems. Furthermore, the algorithm is capable of predicting cyclic lines and lines which consist of multiple branches without any problems. In fact, ternary contour line diagrams can reveal a great variety of mixture properties. Figure 3 shows two types of lines out of this variety for the system acetone-chloroform-methanol. From the course of the isothermal lines the character of the ternary azeotrope (saddle point) can be derived. At the same time the plot of constant lines for the function

F)

|Rij - 1| ∑i ∑ j*i

not only helps to locate the ternary azeotropic composition but also allows to study its effect on the relative volatilities over the entire concentration range. An example of the contour line diagram of a ternary heterogeneous system is shown in Figure 4 which displays a selection of four isothermal curves at 66, 68, 70, and 75 °C for the system ethanol-water-benzene together with the liquid-liquid equilibrium tie-lines at atmospheric pressure. Entrainer Selection The majority of processes for the separation of systems with disadvantageous separation factors require

separating agents, i.e., solvents or entrainer, e.g., for azeotropic and extractive distillation or extraction. The solvent is added either to influence the activity coefficients of the components to be separated to a different extent (extractive distillation) or to introduce a new lower boiling azeotrope which for its part can easily be separated (e.g., heterogeneous azeotropes). The question behind the problem of entrainer selection is, how does one find out which solvents enable the separation of a given binary system? The knowledge of simple and at the same time sufficient conditions that have to be fulfilled by entrainer candidates should enable a quick and easy comparison among a high number of solvents. Entrainers that do not meet the requirements could be recognized and discarded immediately. For the development of simple criteria which permit a solvent screening, it takes properties that almost every solvent and every system has. Universal properties that qualify or disqualify a solvent for the separation of a given binary mixture are the boiling points (vapor pressures) of the pure components and the azeotropic points and the activity coefficients (in particular γ∞) in the resulting ternary system. The first step in the design of azeotropic or extractive distillation is then to find the appropriate solvent which fulfills the criteria based on those properties. With the help of either a large and comprehensive factual database or reliable predictive methods used for the calculation of phase equilibrium data, it is possible to provide the necessary purecomponent and mixture information (γ∞, azeotropic data) and thus to find the suitable solvent for the distillation process under consideration. Data Acquisition and Regression The separation task is provided interactively. The user only has to specify the binary system to be separated and the conditions (system pressure, temperature). All other required information (vapor pressures, structural information, interaction parameters, ...) are called from the Dortmund Data Bank (DDB) or will be calculated using a thermodynamic method. Before the solvent selection part actually starts, the program first of all examines the VLE behavior, and the user will be noticed if a separation of the binary system can possibly be achieved without a separating agent. This could be done by either pressure-swing distillation, binary heteroazeotropic distillation, or vacuum (respectively under elevated pressure) distillation. Note that the proposals

3116 Ind. Eng. Chem. Res., Vol. 37, No. 8, 1998 Table 3. Current Status of the DDB Database for Azeotropic Data on Nonelectrolyte Systems (Oct 1997) number of data points

Figure 5. Flowchart for the selection of solvents by access to the Dortmund Data Bank (DDB).

Figure 6. Procedure for the selection of solvents using thermodynamic models (e.g., group contribution methods).

for each alternative are based on the data that has been calculated or retrieved from the database. For pressure-swing distillation a criterion was derived from systems that are actually separated this way, such as water-tetrahydrofuran, etc. The selection of selective solvents is a good example for the importance of having a reliable knowledge of the mixture data (azeotropic data and activity coefficients at infinite dilution) and information on the purecomponent boiling points (vapor pressures). The purecomponent data are calculated using the Antoine equation or estimated with the help of critical data and acentric factors (Gmehling and Kolbe, 1992). The sources of the mixture data are given below. Note that in the case of supercritical conditions and/or missing parameters no calculation of data with thermodynamic models is performed, and thus no recommendation can be provided. Correspondingly, the method using the experimental database fails for systems for which no experimental data are available at all. The complete flowcharts of the two complementary methods for solvent selection are shown in Figures 5 and 6. Experimental Database. For the selection of suitable solvents by access to the Dortmund Data Bank (DDB), two large databases provide the required mixture data. The information about the zeotropic/azeo-

components

2

3

4

∑

zeotropic data azeotropic data total number of systems

20 650 17 727 38 372 18 606

804 1390 2194 1338

51 81 132 97

21 505 19 198 40 703 20 041

tropic behavior is taken from the database for azeotropic data (Gmehling et al., 1994). The approximately 40 700 data points represent a critically evaluated collection of binary, ternary, and a few quaternary systems. The current status of these databases is shown in Table 3. The database on activity coefficients at infinite dilution (Gmehling et al., 1986/1994) which contains approximately 33 200 data points is used to account for information on selectivity and separability according to eq 6. The program starts searching all binary data sets on azeotropic data and activity coefficients at infinite dilution for component 1 and later on for component 2 in the Dortmund Data Bank. Note that at this stage the second component in the binary search has an indefinite status (“wildcard character”) and is determined by the available binary data. A search for ternary data sets is performed for systems with components 1 and 2 and one of the components found during the search for binary data. Since the number of data sets for ternary systems is much lower than the number of binary data sets (see Table 3), the absence of information about the ternary system is interpreted as zeotropic behavior, in order not to discard a solvent because of missing ternary information only. With the selected data it is possible to determine the number of solvents with the required information on selectivity and VLE. For the identified solvents the separation factors at infinite dilution (R∞12) and temperatures at azeotropic composition (Taz, yaz) for a given system pressure will be estimated in the data regression step. Linear regression is used to estimate the values for Taz and R∞12, respectively, from several experimental data sets for the given isobaric (isothermal) condition. Finally, any solvent that fulfills the selection criteria is included in a list with suitable solvents for the separation problem in hand. Thermodynamic Models. With modern and reliable group contribution methods, for example, UNIFAC (Fredenslund et al., 1977; Hansen et al., 1991) and modified UNIFAC methods (Weidlich and Gmehling, 1987; Gmehling et al., 1993; Larsen et al., 1987) and the ASOG method (Kojima and Tochigi, 1979; Tochigi et al., 1990), the prediction of the required information (γ∞i , azeotropic data) about a given mixture is possible from the information about the interaction between the concerned functional groups only. It is worth mentioning that the existing gE models such as the Wilson (Wilson, 1964), the NRTL (Renon and Prausnitz, 1968), or the UNIQUAC (Abrams and Prausnitz, 1975) model can be applied in a similar manner, but as the number of functional groups is much smaller than the number of possible compounds consisting of these groups, the range of application increases and group contribution methods are to be recommended for a comprehensive and systematic search of solvents. Given a binary system to be separated like 2-propanol-water or benzene-cyclohexane for which group interaction parameters for most of the currently defined UNIFAC groups

Ind. Eng. Chem. Res., Vol. 37, No. 8, 1998 3117

to fugacity coefficients and Poynting corrections. For a ternary system all three binary subsystems can be checked this way. In the case of either a high- or lowboiling entrainer which does not introduce any further binary azeotropes, the absence of a ternary azeotrope can be confirmed by the azeotropic rule for ternary systems (Doherty and Perkins, 1979), which reflects the connections between the so-called singular points of a system:

2(N3 - S3) + N2 - S2 + N1 ) 2 Figure 7. Condition for the occurrence of an azeotrope in a system showing positive deviations (ln γ1 > 0) from Raoult’s law (example: ethanol (1)-1,4-dioxane (2)). The temperature dependence (b) indicates the upper temperature limit for the existence of the azeotropic point.

exist, a method such as modified UNIFAC is able to predict interactions for approximately 2500 ternary systems and thus to provide a selection among up to 2500 entrainer candidates. A review on thermodynamic models for the synthesis and design of separation processes is given by Gmehling (1994). The program using thermodynamic models is rather similar to the procedure using experimental data sets. In contrast to the latter, the program illustrated in Figure 6 needs the specification on the solvents to be examined. With the help of data on activity coefficients at infinite dilution (γ∞) and pure-component vapor pressures, it is possible to predict the occurrence of binary azeotropes, and thus it is possible to preselect potential solvents. The underlying principle of this selection is shown in Figure 7. For binary systems showing positive deviation from Raoult’s law the natural logarithm of the ratio of the activity coefficients (ln γ2/ γ1) as well as the ratio of the pure component vapor pressures ln(P1s/P2s) is plotted versus the mole fraction in the liquid phase for a given temperature. A binary azeotrope with a pressure maximum (temperature minimum) occurs when the following condition is fulfilled:

ln γ∞2 > ln

P1s P2s

> -ln γ∞1

(3)

In Figure 7a the azeotropic composition can be identified as the point of intersection of the two curves. Using the above relation the occurrence and disappearance of azeotropes can be calculated provided that reliable information of the activity coefficients at infinite dilution and vapor pressures as a function of temperature are available. In Figure 7b the courses of ln γ∞2 , -ln γ∞1 , and ln(P1s/P2s) for the system ethanol (1)-1,4-dioxane (2) indicate that the azeotrope will disappear for temperatures above 80 °C because for higher temperatures ln γ∞2 is no longer larger than ln(P1s/P2s). This is approved by experimental data. The condition for a binary pressure-minimum azeotrope, which assumes a system showing negative deviation from Raoult’s law, can be derived in a way similar to that of eq 3:

ln γ∞2 < ln

P1s P2s

< -ln γ∞1

(4)

Note that both conditions are simplified with respect

(5)

N1 is the number of nodes among the pure components, N2 is the number of nodes among the binary azeotropes, S2 is the number of saddle points among the binary azeotropes, and N3 and S3 represent the number of ternary nodes and saddles, respectively. This means the singular points are embodied by the pure components and the binary and ternary azeotropic points. As an example Figure 8a shows the residue curve map of the system benzene-cyclohexane-N-methyl-2-pyrrolidone (NMP) at atmospheric pressure. In this system NMP is the high-boiling entrainer for the separation of the binary benzene-cyclohexane temperature-minimum azeotrope. Thus, all residue curves start at the binary azeotrope and approach the highest boiling point (NMP, 202.0 °C). From the data in the residue curve map the following configuration in terms of singular points can be deduced: N1 ) 1, N2 ) 1, S2 ) 0. The azeotropic rule for such a system takes on the following form:

2(N3 - S3) + 1 - 0 + 1 ) 2

(5a)

Equation 5a is fulfilled if the number of ternary nodes is equal to the number of ternary saddle points, a fact that implies multiple solutions (N3 ) S3 ) 0, 1, 2, ...). However, up to now there is no experimental evidence for the existence of more than one ternary azeotropic point in a ternary system (Gmehling et al., 1994), which means that the practical solution to eq 5a is N3 ) S3 ) 0. Thus, for an extractive distillation entrainer it is not necessary to check the existence of ternary azeotropes, meaning that the rigorous calculation of azeotropic data is required for potential azeotropic distillation solvents only. Solvents for extractive distillation can be distinguished from those suitable for azeotropic distillation processes with the help of these insights, and the search for suitable solvents becomes much more effective by a significant reduction of the total time for the selection. However and as indicated earlier, azeotropic systems can also exhibit residue curve boundaries. This is demonstrated in Figure 8b,c, showing the residue curve maps of benzene-cyclohexane-acetone (one boundary) and benzene-cyclohexane-2-butanone (two boundaries). Each residue curve boundary is connecting two singular points, one of which is always a saddle point. For distillation processes these boundaries can limit the product composition regions in such a way that they form compartments within the residue curve map with curves starting and ending at different points. Note that for each of the three examples shown here the number of compartments is equal to the number of stable nodes, which is, however, not a general rule. The behavior can become more complex in the case of positive and negative deviations from Raoult’s law as shown in Figure 3. It has been demonstrated that boundary distillation lines can be crossed (Stichlmair et al., 1989) or moved when the conditions (e.g., pres-

3118 Ind. Eng. Chem. Res., Vol. 37, No. 8, 1998

Figure 8. Singular points in the ternary residue curve map for (a) benzene-cyclohexane-NMP (N1 ) 1, N2 ) 1, S2 ) 0 ), (b) benzenecyclohexane-acetone (N1 ) 2, N2 ) 1, S2 ) 1), and (c) benzene-cyclohexane-2-butanone (N1 ) 3, N2 ) 1, S2 ) 2): b, binary azeotropic point.

sure) are changed (Knapp and Doherty, 1992). Finally, it is once again important to mention that the reliability of the predicted data on the selected solvents can be checked with the help of the large and unique experimental database. Selection Criteria After their retrieval and evaluation, the required data are used to perform a selection procedure. Different distillation processes have different demands. The following paragraphs indicate the most important criteria for the selection of solvents suitable for azeotropic and extractive distillation. Apart from criteria that are related to thermodynamic properties and phase equilibrium behavior, there are other properties that will practically influence the solvent selection. Therefore, it is planned to extend the entrainer selection procedure, taking into account flash points and transport properties, such as viscosity, etc. Experimental data on these properties are already available in a large purecomponent data bank (Rarey et al., 1995). However, despite all the differences the appropriate entrainer for distillation processes should also be thermally stable, nontoxic, and available at a low price. Azeotropic Distillation. The task of a suitable solvent for azeotropic distillation is the formation of one (or more) lower boiling azeotrope(s), which on their part can easily be separated (Gmehling and Brehm, 1996; Stichlmair et al., 1989). This is especially true for heterogeneous azeotropes, and thus the newly introduced azeotrope should favorably be a heteroazeotrope. Suitable solvents can introduce (a) a ternary azeotrope, (b) a binary azeotrope, and (c) two binary azeotropes. The resulting ternary residue curve maps for exemplary systems with one or more azeotropes introduced to a binary mixture together with their corresponding separation sequences are shown in Figure 9a-d. In all cases one azeotropic constituent is separated by the removal of the newly introduced azeotrope. Separation of this component from the entrainer and solvent recovery can be achieved, for example, by extraction (Figure 9a: cyclohexane + acetone with water), exploitation of a liquid-liquid equilibrium (Figure 9b: recycle of entrainer-rich phase S), stripping (Figure 9c: toluenewater separation), and a combination of stripping and a water trap in the separation of ethanol-water using benzene (Figure 9d). For the suitability of an entrainer which introduces a heterogeneous azeotrope, it is also decisive to deter-

Figure 9. Residue curve maps and column sequences for four typical azeotropic distillation processes.

mine the ratio of the amounts of the two liquid phases. If this ratio takes on a value far from unity, the entrainer is probably not useful for the process because

Ind. Eng. Chem. Res., Vol. 37, No. 8, 1998 3119 Table 4. Results (Extract) of the Entrainer Selection for Azeotropic Distillation for the System Acetic Acid (1)-Water (2) by DDB Access components to be separated (1) acetic acid (2) water system pressure ) 101.32 [kPa] Tb(1) ) 391.01 [K] Tb(2) ) 373.15 [K] DDB access

C2H4O2 H2O

64-19-7 7732-18-5

List of Solvents Introducing One Further Binary Azeotrope (with Pressure Maximum) types of azeotropes introduced selective solvent (3)

1-3

2-3

1-2-3

Tb(az,bin.) [K]

Tm(3) [K]

cyclopentanone 3-pentanone ethyl propionate butyl acetate propyl acetate dipropyl ether diethyl ether dibutyl ether 2-pentanone diisopropyl ether ethyl butyl ether 1,2-dichloroethane dichloromethane isopentyl acetate diisobutyl ketone 2,3-butanedione methyl propionate butyl propionate

none none none none none none none none none none none none none none none none none none

hetPmaxMisgap hetPmaxMisgap hetPmaxMisgap hetPmaxMisgap hetPmaxMisgap hetPmaxMisgap hetPmaxMisgap hetPmaxMisgap hetPmaxMisgap hetPmaxMisgap hetPmaxMisgap hetPmaxMisgap hetPmaxMisgap hetPmaxMisgap hetPmaxMisgap hetPmaxMisgap hetPmaxMisgap hetPmaxMisgap

n.a. n.a. n.a. noneMisgap noneMisgap n.a. n.a. n.a. n.a. none n.a. noneMisgap none n.a. n.a. n.a. n.a. n.a.

367.75 355.98 353.15 363.81 355.52 348.55 307.34 366.65 356.15 336.15 349.15 344.80 311.25 366.95 370.15 351.60 344.75 367.95

222.50 234.15 199.25 199.70 178.00 147.05 156.85 175.30 196.25 186.35 149.15 237.65 178.01 195.15 227.17 270.75 185.70 183.65

either the amount of entrainer which can be recycled or the amount of one of the components from the original binary system that is supposed to be gained by decanting will be too small. Note that for the selection of selective solvents using thermodynamic models this ratio is part of the predicted output. For the selection of solvents by access to DDB, the data can be retrieved directly from the liquid-liquid equilibrium (LLE) data bank, which contains more than 10 500 data sets. Extractive Distillation. In extractive distillation the entrainer is supposed to exhibit zeotropic behavior, with all components of the system to be separated. Moreover, its task is to alter the separation factor

R12 )

K1 γ1P1s ) K2 γ P s

(1)

2 2

of the binary system in the desired way, so R12 or 1/R12 takes on values far from unity (Gmehling and Brehm, 1996). A criterion for the selection of selective solvents is the selectivity at infinite dilution S∞12: ∞ R12,S

)

γ∞1,SP1s γ∞2,SP2s

)

P1s ∞ S12,S s P2

(6)

Thus, the entrainer should alter the activity coefficients of the components to be separated to a different extent in order to achieve separation factors far from unity. For a convenient recovery of the entrainer its boiling point must usually be sufficiently higher (e.g., ∆T ) 40 K) than for any component of the mixture to be separated. Recent investigations have demonstrated the use of low-boiling solvents in so-called “reverse extractive distillation” processes (La´ng et al., 1996), but as the amount of solvent has to be increased enormously to ensure a sufficient liquid concentration of the entrainer and the enthalpy of evaporation affects the

energy balance of the process negatively, those kinds of processes will probably only be competitive when no suitable high-boiling solvent is available. Results and Discussion Ester and ether compounds are promising entrainers for the separation of acetic acid and water. This can be seen in Table 4 which gives an extract of the results that were obtained by searching the DDB for azeotropic distillation entrainers. All the entrainers listed here form a binary heterogeneous azeotrope, with one of the components of the binary system to be separated, which is water in this case. To decide which ester or ether is the most suitable one for the process, graphical tools such as residue curve maps can be applied. With the help of residue curves it is possible to decide whether a given solvent yields the desired products by determining the product composition regions for a given feed composition (Stichlmair et al., 1989; Wahnschafft and Westerberg, 1993). The residue curve map and column sequence for each of the solvents mentioned in Table 4 is, in principle, the same as that in Figure 9b, when 3-pentanone is exchanged, for example, by any ester or ether from Table 4. From the residue curve map in Figure 9b it can be seen that the newly introduced binary heterogeneous azeotrope (A) is one product (distillate) of the distillation of a ternary feed composition M, which results from mixing two streams with the compositions F (binary feed of acetic acid and water) and S. The bottom product of this distillation is pure acetic acid. The binary heterogeneous azeotrope (A) splits into two liquid phases after condensation, one of which is the entrainer-rich phase S to be recycled. The other liquid phase contains mostly water that can be purified by stripping it in the second column of the separation sequence. For the separation of vinyl acetate (1) and methanol (2) the selection method using thermodynamic models was applied. The intention in this particular case was

3120 Ind. Eng. Chem. Res., Vol. 37, No. 8, 1998 Table 5. Results (Extract) of the Entrainer Selection for Extractive Distillation for the System Vinyl Acetate (1)-Methanol (2) Using Modified UNIFACa components to be separated (1) vinyl acetate (2) methanol system pressure ) 101.325 [kPa] Tb(1) ) 345.88 [K] Tb(2) ) 337.70 [K] azeotropic data for system 1-2 type of azeotrope: homPmax Tb ) 334.51 [K] model: modified UNIFAC (Dortmund)

C4H6O2 CH4O

108-05-4 67-56-1

List of Solvents Introducing No Further Azeotrope (Extractive Distillation Entrainer) types of azeotropes introduced selective solvent (3)

Tb [K]

R∞1,2

aniline dibutyl ether decane dodecane butyl acetate 1-hexadecene 2-heptanone butylbenzene 4-methyl-2-pentanone 2-methylphenol 1-octadecene 4-methylphenol phenol tetradecane m-xylene p-xylene isopentyl acetate cyclopentanone cyclohexanone 3-hexanone 4-heptanone butyl butyrate

457.83 415.10 447.27 489.43 399.10 558.02 424.20 456.45 389.34 463.97 587.25 474.91 455.08 525.85 412.25 411.50 410.21 403.80 428.15 396.65 419.47 438.15

0.360 (380.47 [K]) 0.194 (366.23 [K]) 0.146 (376.95 [K]) 0.171 (391.00 [K]) 0.270 (360.89 [K]) 0.208 (413.87 [K]) 0.386 (369.26 [K]) 0.142 (380.01 [K]) 0.392 (357.64 [K]) 0.218 (382.52 [K]) 0.230 (423.61 [K]) 0.225 (386.16 [K]) 0.184 (379.55 [K]) 0.196 (403.14 [K]) 0.144 (365.28 [K]) 0.144 (365.03 [K]) 0.266 (380.47 [K]) 0.452 (362.46 [K]) 0.426 (370.58 [K]) 0.350 (360.08 [K]) 0.348 (367.68 [K]) 0.243 (373.91 [K])

1-3

2-3

1-2-3

Tm(3) [K]

none none none none none none none none none none none none none none none none none none none none none none

none none none none none none none none none none none none none none none none none none none none none none

none none none none none none none none none none none none none none none none none none none none none none

266.85 175.30 243.45 263.59 199.70 277.55 237.65 185.25 189.15 304.09 291.15 308.95 314.05 279.02 225.35 286.35 195.15 222.50 242.00 217.65 240.15 181.15

a Minimum boiling point difference (entrainer - binary mixture) ) 40.00 K. Minimum required value for R (1,2 or inverse) at infinite dilution ) 1.500.

to find high-boiling extractive distillation solvents (zeotropic behavior) with a minimum boiling point difference of 40 K. Furthermore, the entrainers are supposed to cause binary separation factors greater than 1.5 or smaller than 1/1.5 by influencing the activity coefficients of the components to be separated to the desired extent. Table 5 summarizes the results that were obtained using the modified UNIFAC (Dortmund) method to calculate the azeotropic data and separation factors for the selection. Obviously, there are some well-known entrainers, such as aniline, dibutyl ether, xylenes, and butyl acetate, included in the selection results. However, there are also some ketones and phenols in the list. Taking phenol as an example, it is important to mention that the melting point of an entrainer candidate is also taken into account. Especially, some highboiling solvents have high melting temperatures as well, which could possibly lead to unwanted crystallization during the actual process. For this reason only entrainers with melting points lower than an upper limiting value defined by the user are selected. Note that the values for the melting points retrieved from the DDB are included in Table 5. From Table 5 it can be seen that all the entrainers lead to separation factors lower than unity. Thus, the volatility of component 1 (vinyl acetate) is smaller in relation to the value for methanol (component 2), which means that using any of these solvents will give methanol as the distillate of the extractive distillation column, whereas vinyl acetate and the entrainer will leave as the bottom product. With the sufficiently high boiling point difference and

zeotropic behavior in all subsystems apart from the binary that is to be separated, it should be rather easy to achieve the separation of vinyl acetate and the entrainer. The temperature value in parentheses right after the value for R12 is the temperature at which the separation factor has been calculated. This is usually an average value calculated from the boiling points of the three components (binary system to be separated plus the entrainer) at the given pressure. For the selection based on data from the DDB this temperature corresponds to a temperature at which experimental activity coefficients at infinite dilution are available in the database. The vapor pressures for the separation factor (see eq 1) are calculated with the help of the temperature- and pressure-explicit Antoine equation. Of course, the Wagner equation shows an improved capability when compared with the Antoine equation. However, the number of available Wagner parameters is much smaller, so that the number of potential solvents considered would be drastically reduced. A third example presented in Table 6 is showing an extract of the results that were obtained using thermodynamic models once again. In contrast to the previous example, this table actually comprises several selection results, such as a list of solvents to be used for extractive distillation, lists of solvents introducing one or two binary azeotropes, a list of solvents introducing a ternary heterogeneous azeotrope, and finally a list of entrainers that can be used in reverse extractive distillation and thus have a lower boiling point than the binary azeotrope to be separated. Note that the assign-

Ind. Eng. Chem. Res., Vol. 37, No. 8, 1998 3121 Table 6. Results (Extract) of the Entrainer Selection for Azeotropic and Extractive Distillation for the System 2-Propanol (1)-Water (2) Using Modified UNIFACa (1) 2-propanol (2) water azeotropic data for system 1-2 type of azeotrope: homPmax Tb ) 353.64 K model: modified UNIFAC (Dortmund)

system pressure ) 101.325 kPa

C3H8O H2O

Tb(1) ) 355.47 K Tb(2) ) 373.15 K

List of Solvents Introducing No Further Azeotrope (Extractive Distillation Entrainer) types of azeotropes introduced selective solvent (3)

Tb [K]

R∞1,2

1,2-ethanediol dimethyl sulfoxide N-methyl-2-pyrrolidone (NMP) N-methyl-2-piperidone N-methyl-6-caprolactam

470.69 466.74 475.13 483.42 510.21

4.234 (399.77 [K]) 5.108 (398.45 [K]) 0.443 (401.25 [K]) 0.401 (404.01 [K]) 0.343 (412.94 [K])

1-3

2-3

1-2-3

Tm(3) [K]

none none none none none

none none none none none

none none none none none

261.65 291.69 248.75 n.a. n.a.

List of Solvents Introducing One Further Binary Azeotrope (with Pressure Maximum) types of azeotropes introduced selective solvent (3)

1-3

butane dichloromethane 2-methylpropane

2-3

none none none

1-2

hetPmaxMisgap hetPmaxMisgap hetPmaxMisgap

Tb(az,bin.) [K]

Tm(3) [K]

272.58 311.34 261.21

134.85 178.01 113.73

none none none

List of Solvents Introducing Two Further Binary Azeotropes (with Pressure Maximum) types of azeotropes introduced selective solvent (3)

1-3

2-3

1-2-3

Tb(1,3) [K]

Tb(2,3) [K]

ethyl acetate methyl butyl ether isopropyl chloride 1-chloropropane ethyl propyl ether

homPmax homPmax homPmax homPmax homPmax

hetPmaxMisgap hetPmaxMisgap hetPmaxMisgap hetPmaxMisgap hetPmaxMisgap

none none none none none

349.18 342.62 309.59 319.47 336.07

344.62 339.36 308.43 317.59 333.02

List of Solvents Introducing a Ternary Heterogeneous Azeotrope solvent (3)

Tb(az) [K]

yaz(1)

yaz(2)

yaz(3)

Tm(3) [K]

benzene methylcyclopentane heptane diisopropyl ether toluene dipropyl ether isopropyl acetate methyl propionate methyl tert-amyl ether 1-chloropentane ethyl tert-butyl ether ethyl tert-amyl ether

339.71 333.48 345.04 335.29 349.16 347.62 349.74 345.12 347.67 349.74 339.18 350.21

0.2031 0.1605 0.3271 0.0886 0.3692 0.2197 0.1733 0.0870 0.1731 0.3608 0.0848 0.3397

0.2527 0.1891 0.3127 0.2122 0.3738 0.3600 0.3958 0.3268 0.3628 0.3847 0.2536 0.3948

0.5442 0.6504 0.3602 0.6992 0.2571 0.4204 0.4309 0.5862 0.4641 0.2545 0.6616 0.2655

278.68 130.73 182.55 186.35 178.16 147.05 199.75 185.70 n.a. 174.15 179.15 n.a.

List of Solvents Introducing No Further Azeotrope (Reverse Extractive Distillation Entrainer) types of azeotropes introduced selective solvent (3)

Tb [K]

R∞1,2

1-butene isobutylene

266.89 266.25

0.231 (331.83 [K]) 0.295 (331.62 [K])

1-3

2-3

1-2-3

Tm(3) [K]

none none

none none

none none

87.85 132.75

a Minimum boiling point difference (entrainer - binary mixture) ) 40.00 K. Minimum required value for R (1,2 or inverse) at infinite dilution ) 1.500.

ment to each class of solvents has been performed automatically, based on the azeotropic data and the activity coefficients at infinite dilution that have been calculated for each component of the DDB for which interaction parameters with groups of the binary system for the modified UNIFAC method exist. Thus, the selection using thermodynamic models is a very systematic approach for the entrainer selection problem with a fairly extended output, occasionally. It is particularly useful for the investigation of certain component classes, such as alcohols, ethers, ..., etc. Finally, Table 7 gives a list of selective solvents (selection by DDB access) which can be applied to the

separation of a binary mixture of benzene and cyclohexane (as an example of the separation of aliphatics from aromatics) by extractive distillation. Note that this list contains solvents that are, on one hand, of practical importance for the industrial process but, on the other hand, cannot yet be described by any group contribution method due to a lack of interaction parameters or group definitions. Thus, any procedure which searches solvents using group contribution methods would fail to notice entrainers like sulfolane, N-formylmorpholine, 6-caprolactam, or N-methyl-6-caprolactam, to name those of this example. (This discussion is based on the capabilities of the modified UNIFAC method with

3122 Ind. Eng. Chem. Res., Vol. 37, No. 8, 1998 Table 7. Extractive Distillation Entrainers (Extract) for the Separation of the Binary System Benzene (1)-Cyclohexane (2) Method: Selection by DDB Accessa components to be separated (1) benzene (2) cyclohexane system pressure ) 101.32 [kPa] azeotropic data for system 1-2 type of azetrope: homPmax Tb ) 351.52 [K]

Tb ) 353.25 [K] Tb ) 353.86 [K]

C6H6 C6H12

types of azeotropes introduced selective solvent (3)

Tb [K]

R∞1,2

N-methyl-6-caprolactam 4-methyl-2-pentanone aniline 1,2-ethanediol N-methyl-2-pyrrolidone (NMP) triethylene glycol sulfolane 6-caprolactam furfural 3-methylphenol cyclohexylamine tetrahydrofurfuryl alcohol nitrobenzene cyclohexanone anisole N-formylmorpholine butyronitrile

510.21 389.34 457.83 470.69 476.15 551.49 559.96 546.32 434.67 475.82 407.14 450.20 483.94 428.15 428.79 519.05 390.77

0.586 (405.77 [K]) 0.420 (293.15 [K]) 0.323 (388.31 [K]) 0.256 (392.60 [K]) 0.299 (394.42 [K]) 0.367 (419.53 [K]) 0.275 (422.36 [K]) 0.267 (348.15 [K]) 0.306 (380.59 [K]) 0.379 (298.15 [K]) 0.646 (303.15 [K]) 0.247 (300.15 [K]) 0.289 (397.02 [K]) 0.293 (293.15 [K]) 0.328 (293.15 [K]) 0.269 (408.72 [K]) 0.195 (298.15 [K])

1-3

2-3

1-2-3

Tm(3) [K]

none none none none none none none none none none none none none none none none none

none none none none none none none none none none none none none none none none none

n.a. none none n.a. n.a. n.a. n.a. n.a. n.a. n.a. n.a. n.a. n.a. n.a. n.a. n.a. n.a.

n.a. 189.15 266.85 261.65 248.75 268.85 301.65 342.45 234.45 284.95 255.15 n.a. 278.85 242.00 235.65 n.a. 161.00

a Minimum boiling point difference (entrainer - binary mixture) ) 35 K. Minimum required value for R (1,2 or inverse) at infinite dilution ) 1.500.

Conclusion

Figure 10. Experimental activity coefficients at infinite dilution as a function of temperature for benzene (b) and cyclohexane (2) in N-formylmorpholine (NFM) from various authors (source: DDB): s, linear regression.

its published group definitions and interaction parameters. In the meantime within the company-supported UNIFAC consortium, the required parameters are already available for the sponsors.) Afterall, the new method of selecting solvents by DDB access gives important additional information and is very innovative and practice-oriented in its results. As mentioned earlier, this method takes advantage of the large and comprehensive factual database. Figure 10 gives as an example the graphical representation of the experimental activity coefficients of benzene and cyclohexane at infinite dilution in N-formylmorpholine (NFM). These data were used to perform a linear regression to give coefficients for the temperature dependence of ln γ∞i . With the help of these coefficients it is then possible to calculate selectivities at infinite dilution at a given temperature, and by using the pure-component vapor pressures, it is finally possible to calculate the corresponding separation factors listed in Table 7.

Group contribution methods and large factual data banks are ideal tools for the synthesis and design of thermal separation processes. New methods for the selection of solvents for azeotropic and extractive distillation have been presented. One method makes extensive use of thermodynamic models, whereas the other one allows a selection of solvents by access to a comprehensive database with experimental data (γ∞i , azeotropic data). The methods work complementarily; i.e., in the limits of thermodynamic models data retrieved from experimental data banks can provide important additional information. Information about the performance of a separation process with a given solvent can be gained from the graphical representation of the mixture properties. This information on, for instance, the minimum entrainer supply, the product composition regions, the feasibility, etc., can be used to decide about the effectiveness and costs and thus to compare solvents for a given separation problem in order to find the “best” entrainer. Further criteria concerning, for example, flash points, viscosity, certain functional groups (to be included or excluded), toxicity, and the extension to other separation processes (e.g., liquid-liquid extraction, absorption, ...), will be included soon. Acknowledgment The authors thank Fonds der Chemischen Industrie for financial support, DDBST GmbH, Oldenburg, Germany, for providing the current version of the Dortmund Data Bank together with the accompanying software package, and Mr. Oliver Noll for his valuable contributions to the present work. Nomenclature Rij ) separation factor of components i and j γi ) activity coefficient of component i

Ind. Eng. Chem. Res., Vol. 37, No. 8, 1998 3123 ξ ) dimensionless measure of time g ) molar Gibbs energy (J/mol) Ki ) K-factor of component i Pis ) saturation vapor pressure of component i (kPa) Sij,k ) selectivity of the solvent k for the binary system of components i and j t ) temperature (°C) T ) absolute temperature (K) Tb ) normal boiling point (K) xi ) liquid-phase mole fraction of component i yi ) vapor-phase mole fraction of component i Superscripts ′, ′′ ) liquid phases E ) excess property ∞ ) at infinite dilution Subscripts az ) at azeotropic composition S ) solvent Abbreviations for Azeotrope Types n.a. ) not available hom ) homogeneous het ) heterogeneous none ) zeotrope Pmax ) pressure maximum Pmin ) pressure minimum Psad ) saddle point Misgap ) miscibility gap (liquid-liquid equilibrium)

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Received for review November 5, 1997 Revised manuscript received March 30, 1998 Accepted April 3, 1998 IE970782D