Liquid Membrane

Ind. Eng. Chem. Res. 2007, 46, 8701-8709

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Propylene/Propane Separation with a Gas/Liquid Membrane Contactor Using a Silver Salt Solution Pavan Chilukuri,† Karlijn Rademakers,† Kitty Nymeijer,‡ Louis van der Ham,† and Henk van den Berg*,† Process Plant Design, UniVersity of Twente, P.O. Box 217, 7500 AE Enschede, The Netherlands, and Membrane Technology Group, UniVersity of Twente, P.O. Box 217, 7500 AE Enschede, The Netherlands

Traditional systems for olefin/paraffin separation, like low-temperature distillation, are expensive and energyconsuming, and therefore, alternative separation methods are desired. The conceptual design of a new propylene/ propane separation process (300 kton/yr at 99.9 wt % propene purity) by means of a gas/liquid membrane contactor and a selective silver salt solution is presented. It involves the selection of an absorption liquid (AgBF4), its concentration (5.5 M, assumed loading of 90%), and the operating conditions of the membrane contactor and the desorber. Subsequently, modeling of the gas/liquid membrane contactor is carried out. The calculated membrane area required for the process is approximately 80 000 m2. The influence of the assumptions made during the conceptual design and modeling stage is evaluated with a sensitivity analysis. Finally, a preliminary design is presented, resulting in a process flow sheet with equipment sized on standard design criteria. A brief economic evaluation shows that the marginal difference between the feed and product prices should be at least 175 $/ton to make the proposed propylene/propane separation process economically feasible. Introduction Olefin/paraffin separation is an important separation in the petrochemical industry.1 Traditional systems, like low-temperature distillation,2 are expensive and energy-consuming, and alternative, less-expensive separation methods are, therefore, required. A significant amount of research has been conducted on olefin/paraffin separation. However, there are currently only a few nondistillation processes being used in the chemical industry.2 This is due to the inherent problems associated with alternative processes. Technologies, such as metal-based adsorption, are prone to deactivation by feed contaminants and not acceptable for practical applications.2 Traditional facilitated transport membranes have similar problems, along with the need for water content control.2 Another alternative method is extractive distillation, but studies2 indicate that this method offers no advantage over traditional distillation for the solvents considered in the study. Recently, Nymeijer and co-workers1,3 reported promising results using a gas-liquid membrane contactor to separate ethylene from ethane. The separation is based on the ability of silver ions to reversibly complexate olefins. In their work, Nymeijer and co-workers1,3 proved the long-term stability of such a process, making it attractive for industrial applications. In the present study, a technical and economical evaluation of such a system is performed for the separation of propylene and propane using an aqueous silver salt solution as the absorption liquid. A structural design method4 is applied to develop, stepby-step, the process flow sheet for the required separation (Figure 1). All calculations are based on mass and heat balances, which are translated into a computational model in Microsoft Excel. The model calculations are solved using an Eulerian * To whom correspondence should be addressed. E-mail: [email protected]. Tel.: +31(0) 53 4894482. Fax: +31(0) 53 4892882. † Process Plant Design, University of Twente. ‡ Membrane Technology Group, University of Twente.

Figure 1. Steps to create process flow sheet.

Figure 2. Scheme of propylene/propane separation using a silver salt solution and a gas/liquid membrane contactor as absorber.5

discretization scheme. Physical solubilities of the hydrocarbons are calculated using Henry’s law. Principle of Separation The separation of propylene from propane is performed in a gas/liquid membrane contactor. Gas and liquid phases are separated by a composite membrane consisting of a dense polymeric top layer on top of a porous support (Figure 2).5 Membrane contactors have a high operational flexibility and

10.1021/ie070556w CCC: $37.00 © 2007 American Chemical Society Published on Web 08/30/2007

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Ind. Eng. Chem. Res., Vol. 46, No. 25, 2007

Table 1. Fixed Variables for Technical and Economical Evaluation separation technology separation method feed composition feed conditions product composition capacity

gas/liquid membrane contactor selective complexation with silver 20% propane 80% propylene 12 bar min. 99.9 wt % propylene and water only in ppm level min. 99.5 wt % propane 300 kton/yr of propylene

allow independent control of the gas and liquid flows, as opposed to state-of-the-art packed-column devices.6 The propylene/propane feed gas enters the membrane absorber at the bottom and flows along the membrane, where it flows countercurrently to a silver salt solution. A membrane is used to bring the gas and liquid phases into contact in a controlled way. Propylene and propane diffuse through the membrane, where propylene selectively reacts with silver ions, resulting in a silver-propylene complex. The reaction is based on the ability of silver ions to reversibly complexate propylene via a combination of a π- and a σ-bond between silver ions and propylene.1 Propane is only physically absorbed in the silver salt solution. The propylene rich silver salt solution leaves the absorber at the bottom. By changing the temperature and/or pressure, the equilibrium of complexation can be influenced, allowing desorption of propylene in a desorber. The lean silver solution is cooled and recycled back to the absorber. The propane rich gas stream leaves the absorber at the top. The process requirements used as input values for the evaluation are listed in Table 1. Conceptual Design A conceptual design of the propylene/propane separation process requires the selection of a silver salt solution and its concentration, the determination of the process temperature and pressure, and the determination of the required functional units. Selection of Silver Salt Solution. The absorption of propylene in silver salt solutions can occur in two different ways, by physical absorption and by chemical absorption. The physical absorption can be described with the distribution coefficient (m) or the Henry coefficient (H). The chemical absorption is based on an equilibrium reaction:1,7,8

Ag+ + C3H6 T Ag-C3H6+

(1)

The equilibrium constant (Kc) describes the equilibrium between the silver complex concentration, the free silver ion concentration, and the propylene concentration in the solution and is dependent on the temperature (eq 2).

Kc(T) )

[Ag-C3H6+] [Ag+][C3H6]

(2)

where Kc(T) is the equilibrium constant of complexation (m3/ mol), [Ag-C3H6+] is the silver complex concentration (mol/ m3), [Ag+] is the free silver ion concentration (mol/m3), and [C3H6] is the propylene concentration (mol/m3) in the solution. The free silver ion concentration in the absorption liquid depends on the degree of dissociation (Kd) of the silver salt in solution, which can be described by eq 3:

Kd(T) )

[Ag+][X-] [AgX]

(3)

Figure 3. Equilibrium phenomena in the system.

where Kd(T) is the equilibrium constant of dissociation (mol/ m3), [X-] is the free salt ion concentration (mol/m3), and [AgX] is the not-dissociated silver salt concentration (mol/m3). In the literature, the two most often investigated silver salt solutions regarding complex formation with propylene are silver nitrate (AgNO3) and silver tetrafluoroborate (AgBF4)8-10 in water. Fogg and Gerrard11 presented the Henry coefficients of propylene and propane in pure water depending on the temperature. A higher temperature results in a lower physical absorption between 0 and 80 °C. Teramoto et al.7 reported the influence of the AgNO3 concentration on the Henry coefficient for ethane and found that a higher concentration also results in a lower physical absorption. Nymeijer et al.1 reported that an increase in AgNO3 concentration resulted in a decrease in degree of dissociation. An increase of the temperature, on the other hand, resulted in an increase in the degree of dissociation of AgNO3. Herberhold9 reported that an AgBF4 solution in water can be considered as fully dissociated, which means that all silver is available for complexation. The relative decrease in the dissociation of AgNO3 results in a relative decrease of the propylene absorption capacity with increasing AgNO3 concentration, as reported by Featherstone and Sorrie.8 This is in contrast to AgBF4, where the absorption capacity increases with increasing salt concentration.8 Baker10 reported the same behavior for the absorption of ethylene in AgNO3 and AgBF4 solutions. The influence of the temperature on the absorption capacity of 1-butene in an AgBF4 solution is reported by Featherstone and Sorrie,8 whereas Nymeijer et al.1 reported the influence of the temperature on the equilibrium constant of complexation for ethylene in an AgNO3 solution, which strongly decreases with increasing temperature. AgBF4 solutions exhibit higher silver efficiencies because of the complete degree of dissociation of AgBF4 compared to that of AgNO3. Higher silver salt concentrations result, in the case of AgBF4, in an increase of absorption capacity, whereas the opposite is true for AgNO3. This resulted in the selection of AgBF4 in water as the selective absorption medium. Determination of Process Temperature, Pressure, and AgBF4 Concentration. Figure 3 presents the different phenomena occurring in the absorption and desorption stages at equilibrium conditions. Mass balances based on these equilibrium conditions are used to determine the optimal temperatures of absorption and desorption and the required silver salt concentration. The equilibrium constant for complexation (Kc) is determined, for 3, 5.5, and 7.5 M AgBF4 solutions, using the absorption capacity reported by Featherstone and Sorrie8 and the Henry coefficients reported by Fogg and Gerrard11 and Teramoto et

Ind. Eng. Chem. Res., Vol. 46, No. 25, 2007 8703

Figure 4. Equilibrium constant for complexation for 1-butene in AgBF4 at various temperatures and concentrations. Table 2. Equilibrium Constants for Complexation at 293 K solution

(mol of C3H6)/(mol of AgBF4)

Kc (m3/mol)

3.0 M AgBF4 5.5 M AgBF4 7.5 M AgBF4

0.45 0.75 0.98

0.2 1.1 26.1

al.7 The results are summarized in Table 2 and show that a higher AgBF4 concentration results in a higher equilibrium constant. The equilibrium constant for complexation is a function of the temperature; hence, changing the temperature makes it possible to influence the amount of propylene absorbed by the silver salt solution and allows desorption. Two similar studies (one about ethylene and one about 1-butene) reported the influence of the temperature on the absorption capacity,1,8 and both assumed that the influence of the temperature on the equilibrium constant for complexation can be described by the van’t Hoff equation,12

d ln Kc d (1/T)

)-

∆Hr R

(4)

where T is the temperature (K), ∆Hr is the enthalpy of reaction (J/mol), and R is the gas constant (J/mol‚K). It is assumed that 1-butene in an AgBF4 solution shows similar behavior as propylene in an AgBF4 solution; hence, the data of Featherstone and Sorrie8 can be applied to calculate the reaction enthalpy of propylene in an AgBF4 solution. The equilibrium constants for complexation at elevated temperatures are calculated using eqs 5 and 6, the Henry coefficients,7,11 and the absorption capacity data of 1-butene in AgBF4 at various temperatures for 1.0, 3.0, 5.5, and 7.5 M AgBF4 solutions.8

[Ag+] + [Ag-C3H6+] ) MAgBF4 × 1000

(5)

[C3H6] + [Ag-C3H6+] ) absorption capacity 1000 × MAgBF4

(6)

where MAgBF4 is the molarity of the silver salt solution (kmol/ m3). The results are plotted in Figure 4, where the slope of the graphs equals -∆Hr /R. Consequently, the average reaction enthalpy for 1.0, 3.0, and 5.5 M absorption liquids is 50 kJ/ mol. Comparing these results with the reaction enthalpy of 27 kJ/mol for ethylene in a 1.0-3.5 M silver nitrate solution1 shows that the current system is more affected by the temperature than the ethylene-silver nitrate solution system.

Figure 5. Functional diagram of propylene/propane separation process based on a gas/liquid membrane absorber using a silver salt solution.

The reaction enthalpy is used to calculate the equilibrium constant for complexation at elevated temperatures, assuming that the influence of the temperature for propylene in an AgBF4 solution equals that for 1-butene in an AgBF4 solution. The vapor pressure of water is calculated with the Antoine equation13 and used to estimate the loss of water to the gas phase in the desorber, because of the temperature increase. Equilibrium calculations are performed at the top and bottom sections of the absorber and desorber in order to select the optimal absorber and desorber temperatures and the silver salt solution concentration. The following assumptions are made: • Absorber and desorber pressures are 12 bar. • There is no pressure drop. • Complexation reactions and phase equilibriums are instantaneous. • A nonselective composite membrane is present between the gas and liquid phases. • The influence of temperature for propylene absorption equals that for 1-butene absorption in an AgBF4 solution. Table 3 summarizes the results of the equilibrium calculations for 3.0, 5.5, and 7.5 M silver salt solutions and for absorber temperatures of 293 and 308 K. Desorption is calculated for two temperatures as well: (i) 353 K, where evaporation of water is still below 4%, and (ii) for the temperature where the required propylene purity of 99.9% still can be achieved, assuming that the evaporated water can be completely separated from the product. Table 3 shows that high product purities can be achieved. On the basis of the equilibrium calculations, several options are promising, and the most promising one needs to be selected. The following issues are important in this respect: • The option must result in an achieved propylene purity of 99.9%. • A lower absorber temperature will result in higher cooling costs. • A higher desorber temperature will result in higher heating costs. • A higher desorber temperature will result in more water evaporation and, hence, in higher drying costs. • A higher silver salt solution flow will result in higher energy, silver salt, and equipment costs. Higher AgBF4 concentrations result in a decrease of the propane absorption and, thus, desorption and in an increase of the propylene absorption, whereas desorption of propylene becomes more difficult because of the more stable complex. Furthermore, silver salts are expensive and the required volume of the silver salt solution has a large influence on the equipment size and energy costs, thus favoring small volumes and lower silver salt concentrations. Therefore, an AgBF4 concentration of 5.5 M seems to be the best choice. This limits the number of viable possibilities in Table 3 to only two: a 5.5 M AgBF4 solution combined with a desorber temperature of 353 K and

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Figure 6. Scheme of a model absorber: (A) overall absorber, (B) staging absorber, and (C) concentration profile membrane surface. Table 3. Results of Equilibrium Mass Balances 3 M AgBF4

5.5 M AgBF4

7.5 M AgBF4

TA (K) TD (K) PH2O (Pa)

293 353 47371

335 21693

353 47371

C3H6 (%) C3H8 (%)

99.91 0.09

99.90 0.10

99.92 0.08

Q (m3/s)

0.10

0.11

required silver salt solution flow (propylene capacity: 300 kton/yr) 0.14 0.17 0.05 0.18 0.06 0.29 0.10

308 336 22704

293 353 47371

308 319 9997

353 47371

320 10521

product composition (assuming complete removal of water) 99.90 99.97 99.90 99.98 99.90 0.10 0.03 0.10 0.02 0.10

an absorber temperature of, respectively, 293 or 308 K. Because an absorber temperature of 293 K requires a higher cooling cost, a 5.5 M AgBF4 solution with an absorber temperature of 308 K and desorption at 353 K is selected for process design purposes. On the basis of the calculated water vapor pressures (Table 3), it can be concluded that a dryer unit, after the separation of propylene from the silver salt solution, is definitely necessary to achieve the required purity of propylene. Furthermore, addition of extra water is required to compensate for this loss of water and to prevent dehydration of the silver salt solution. The outcome of the previous calculations is used for the conceptual design of the process, which is presented in Figure 5.

-

d[Fg[C3H6]g] ) JC3H6 a A dz

(7)

-

d[Fg[C3H8]g] ) JC3H8 a A dz

(8)

where Fg is the gas feed flow (m3/s); [C3H6]g and [C3H8]g are the concentrations of, respectively, propylene and propane in the gas flow (mol/m3); JC3H6 and JC3H8 are, respectively, the

338 24848

353 47084

308 334 20720

99.97 0.03

99.90 0.10

99.98 0.02

99.90 0.10

0.27

0.10

0.42

propylene and propane flux (mol/m2‚s); a is the interfacial area (m2/m3); and A is the cross-sectional area in the gas-phase side (m2). The propylene and propane flux through the membrane can be derived from eqs 9 and 10.14 The inverse values of the mass transfer coefficients for the gas phase (kg), the support layer (ks), the coating layer (kc), and the liquid phase (kl) represent the resistances of the gas flux in the membrane absorber and are schematically presented in Figure 6

JC3H6 )

(

[C3H6]g -

)

[C3H6]l mC3H6

1 1 1 1 + + + kg ks kc mC3H6EC3H6kl

Modeling Gas/Liquid Membrane Contactor The promising results of the equilibrium calculations form the starting point for a rate-based description of the phenomena occurring in the gas/liquid membrane contactor. This model divides the absorber into two sections, the gas phase and the liquid phase, separated by a nonselective composite membrane consisting of a dense top layer on top of a porous support. Each phase is assumed to be plug flow, and the mass balance equations are derived using Figure 6. The propylene and propane concentrations in the gas phase can be expressed with eqs 7 and 8.

293 353 47371

JC3H8 )

(

[C3H8]g -

)

[C3H8]l mC3H8

1 1 1 1 + + + kg ks kc mC3H8kl

(9)

(10)

where [C3H6]l and [C3H8]l are the concentrations of, respectively, propylene and propane in the liquid flow (mol/m3); kg, ks, kc, and kl are the mass transfer coefficients for, respectively, the gas phase, the support layer, the coating layer, and the liquid phase (m/s); and mC3H6 and mC3H8 are the distribution coefficients for, respectively, propylene and propane. Mass transfer coefficients in the gas, the liquid, and the support layer of the membrane are obtained from literature.15 The mass transfer coefficient of the coating layer is derived from data of lab experiments for a 10 µm thick coating layer.15 This value is recalculated for a thinner, but realistic, coating layer thickness of 1 µm (Table 4). In the case of propylene, the flux is enhanced because of complex formation in the liquid. This is expressed as the enhancement factor (EC3H6), which is defined as the ratio of the

Ind. Eng. Chem. Res., Vol. 46, No. 25, 2007 8705 Table 4. Mass Transfer Coefficients of Different Phases in the Membrane Absorber15 mass transfer coefficient (×10-4 m/s)

phase gas phase support layer dense top layer (l ) 1 µm) liquid phase

kg ks kc kl

2.4 4.0 0.5 0.1

flux through the liquid film with reaction relative to the flux through the liquid film without reaction, at the same driving force and mass transfer coefficient kl. The enhancement factor depends on the absorption regime and can be estimated for a reversible reaction with arbitrary kinetics using eq 11,16 which was actually derived for the surface renewal model for a (1,1) irreversible reaction.

EC3H6 )

-HaC3H62 2 × (EC3H6,∞ - 1)

x(

+

HaC3H64

4 × (EC3H6,∞ - 1)2

+

EC3H6,∞ HaC3H62 (EC3H6,∞ - 1)

)

+ 1 (11)

where EC3H6 is the enhancement factor, EC3H6,∞ is the infinite enhancement factor, and HaC3H6 is the Hatta number for propylene. The Hatta number characterizes the absorption regime and is defined as the ratio of the chemical absorption of propylene in the layer to the flux through the layer without reaction, as given in eq 12.16

HaC3H6 )

x

Kf[Ag+]lm DC3H6

EC3H6,∞ ) 1 +

x

(x

[Ag+]bulkKc

Table 5. Diffusion Coefficients of Various Components in Silver Salt Solutions1 component

diffusion coefficient (×10-9 m2/s)

propylene propane silver ion silver-propylene complex

1.08 1.08 1.24 0.90

Because of the similarity of the components, the diffusion coefficients of propane, propylene, propylene-silver complex, and silver ions in a AgBF4 solution are assumed to be equal to the corresponding values of ethylene, ethane, ethylene-complex, and silver ions in a AgNO3 solution, as found in the literature (Table 5).1 Initially it is assumed that the liquid bulk is on its reaction equilibrium, which will be checked afterward. The liquid concentrations in each slice in the absorber (Figure 6) can then be calculated using the equilibrium of complexation and eqs 16 and 17.

(12)

kl2

[C3H6]l,i + [Ag-C3H6+]l,i -

where Kf is the forward reaction rate constant (m3/mol‚s) and DC3H6 is the diffusion coefficient for propylene (m2/s). The forward reaction rate constant available from the literature for ethylene in an AgNO3 solution1 has been used to estimate the Ha number in the present work. The infinite enhancement factor (EC3H6,∞) is the second factor that characterizes the absorption regime and is given by eq 13 for a reversible reaction of type A + B T C16 based on the surface renewal model.17

DAg+ DC3H6

Figure 7. Absorption regimes.1

)

JC3H6 a A h Fl

) [C3H6]l,i+1 +

[Ag-C3H6+]l,i+1 (16) [Ag+]l,i + [Ag-C3H6+]l,i ) MAgBF4 × 1000

where h is the step size (m) and Fl is the liquid flow (m3/s). The gas flow is calculated using eqs 18 and 19 for each slice in the absorber. where nC3H6,g,i and nC3H8,g,i are the mole flows

Fg,i )

(13)

DAg+ + [C3H6]if Kc DAg+-C3H6

(nC3H6,g,i + nC3H8,g,i) R Ti P

nC3H6,g,i+1 ) nC3H6,g,i - JC3H6,i a A h

where DAg+ and DAg+-C3H6 are the diffusion coefficients for, respectively, the free silver ion and the complex (m2/s), [Ag+]bulk is the free silver concentration in the bulk of the liquid (mol/ m3), and [C3H6]if is the propylene concentration at the interface between the membrane and the liquid (mol/m3). Generally, the absorption regions of propylene can be divided into four regimes, but when one of the two following requirements is fulfilled, it can be assumed that the reaction of complexation is infinitely fast and, hence, the liquid bulk is on its reaction equilibrium (Figure 7).

2 < HaC3H6 , EC3H6,∞ (EC3H6 ) HaC3H6)

(14)

HaC3H6 . EC3H6,∞ (EC3H6 ) EC3H6,∞)

(15)

(17)

(18) (19)

of, respectively, propylene and propane in the gas phase in slice i (mol/s). Silver ions are used to complexate or decomplexate propylene, which generates or requires heat. This results in an enthalpy of reaction of -50 kJ/mol (Figure 4). The temperature difference due to reaction is now calculated using eq 20 and assuming that the heat flow across the membrane can be neglected because it is small compared to the enthalpy of reaction.

Ti+1 ) Ti +

([Ag-C3H6]i+1 - [Ag-C3H6]i) ∆Hr Mw CpF × 1000

(20)

where Mw is the mole weight (g/mol), Cp is the heat capacity (J/mol‚K), and F is the density (kg/m3).

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Figure 8. Mole fraction in the gas phase and transmembrane flow of propylene over the length of the absorber.

Figure 11. Effect of support layer mass transfer coefficient (ks) on model results.

Figure 9. Effect of liquid-phase mass transfer coefficient (kl) on model results.

Figure 12. Effect of gas mass transfer coefficient (kg) on model results.

Table 6. Input Values Model Calculations

Figure 10. Effect of coating layer mass transfer coefficient (kc) on model results.

The main input values for the mass balance calculations are the composition of the feed gas and the composition of the outlet liquid in the absorber. The outlet concentrations of the liquid are derived from the results of the mass balance calculations at equilibrium. To do so, a silver loading of 90% is assumed. The loading factor represents the loading of the silver salt solution and influences the different process-sizing parameters. It is the main design variable in the model. The mass balance calculations are performed using the input parameters presented in Table 6 and are based on the following assumptions: • Absorber pressure is constant at 12 bar. • Gas and liquid phases are separated by a nonselective composite membrane. • The process is operated at steady state. • Both phases are plug flow. • The complexation reaction is first order with respect to silver ions as well as propylene.

parameter

value

P T MAgBF4 Fl Fg,in Kf1 A‚a Cpropylene,eq. Cpropane,eq. Csilver,eq. Ccomplex,eq. mpropylene mpropane loading factor step

1.2 × 106 Pa 308 K 5500 mol/m3 0.06 m3/s 0.74 m3/s 200 m3/mol‚s 4000 m4/m3 15 mol/m3 1 mol/m3 796 mol/m3 4704 mol/m3 4.0 × 10-2 9.2 × 10-3 0.9 0.01 m

• Initially, the propane flow through the membrane is neglected. • An annual production of 8000 h is assumed. The results of the mass balance calculations are presented in Figure 8. It shows the propylene mole fraction in the gas phase and the propylene transmembrane flow as a function of the absorber length. At the bottom of the column, the propylene gas concentration is high but the propylene loading in the liquid is also high, resulting in a small driving force for propylene transport and, hence, a small propylene flux. Continuing in the z-direction of the absorber, a decrease of the propylene concentration in the gas phase as well as the propylene loading in the liquid phase is observed, which results at the end of the absorber in low propylene concentrations in the liquid as well as in the gas phase and a maximum propylene flux in the center of the absorber.


Figure 15. Effect of loading factor on model results. Figure 13. Effect of forward reaction rate constant (Kf) on model results.

Figure 16. Effect of propane purity on membrane area. Figure 14. Effect of propylene diffusion coefficient (DC3H6) on model results.

The length of the absorber corresponds with an important design parameter, the required membrane area, and can be calculated using eq 21.On the basis of the model calculations, the needed membrane area to obtain the required purities is 79 640 m2. For a 1 micron tube, the interfacial area a will be 4000 m-1. We can, e.g., assume that the cross-sectional area for the gas phase A ) 1 m2.

Atotal ) a A L

(21)

where Atotal is the required membrane area (m2) and L is the length of the absorber (m). Initially, the propane flux through the membrane is neglected in the model calculations. However, the more propylene is transported through the membrane, the higher is the propane concentration in the gas phase and the higher is the driving force for propane transport, which results in a higher propane flux. Therefore, the propane flux through the membrane in the absorber is calculated afterward applying the following assumptions. • Concentration of propane in the regenerated liquid at the top of the absorber is 0 mol/m3. • Propane gas flow at the top of the absorber is 69 mol/s (maximum propane flow at the top based on feed). • The total gas flow (m3/s) is not influenced by the propane flux through the membrane. The recalculated propane gas flow is compared to the propane gas flow in the situation where no propane flux occurs. The difference shows a maximum of 0.3% in the total propane gas

flow. This allows assuming zero propane flux through the membrane in the absorber. Furthermore, it was assumed that the liquid bulk is at equilibrium, but this needs to be checked using the Hatta number and the infinite enhancement factor. The Hatta number at the bottom of the absorber is 1371 and will only increase in the absorber due to the increase of the silver concentration in the liquid. The infinite enhancement factor in the absorber is 53 and will also increase. Therefore, it can be concluded that the second requirement with respect to the absorption regime (eq 15) is fulfilled. This allows for assuming that the reaction of complexation is infinitely fast and, hence, the liquid bulk is on its reaction equilibrium. Sensitivity Study Model Absorber The final conceptual design depends on the values of the different design parameters: (i) the liquid, coating layer, support layer, and gas mass transfer coefficients; (ii) the forward reaction rate constant; (iii) the silver, propylene, and complex diffusion coefficients; (iv) the loading factor; and (v) the desired propane purity. The influence of the value of each parameter on the results of the gas/liquid membrane contactor design is investigated in a sensitivity study. Base case values are presented in Tables 1, 5, and 6. Figures 9, 10, 11, and 12 show that increasing the mass transfer coefficients will decrease the absorber length. This effect is critical in the range of 1 × 10-6 to 1 × 10-5 m/s for the liquid-phase mass transfer coefficient kl, between 5 × 10-6 and 5 × 10-5 m/s for the coating layer mass transfer coefficient kc, and between 1 × 10-5 and 1 × 10-4 m/s for the gas-phase mass transfer coefficient kg. The support layer mass transfer

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Figure 17. Process flow diagram of propylene/propane separation process based on a gas/liquid membrane absorber using a silver salt solution. Table 7. Results Detailed Factorial Method component

% of FCI

Direct Costs purchased process equipment 25 installation 10 instrumentation 10 piping 8 electrical systems 5 buildings 8 yard improvement 3 service facilities 10 land 2 total direct costs Indirect Costs engineering & supervision 10 construction overhead 1.5 contractors fee 2 contingency 5 total (%) 99.5 total indirect costs silver costs total fixed capital investment

Table 8. Margin Sensitivity Studies

normalized (%)

estimated cost (M$)

25 10 10 8 5 8 3 10 2

10.9 4.3 4.3 3.4 2.1 3.4 1.3 4.3 0.9 34.9

10 2 2 5 100

4.3 0.9 0.9 2.1 8.0 9.5 52.6

coefficient ks does not have a major effect on the absorber length as long as it does not get below a value of 1 × 10-4 m/s. The mass transfer coefficients should be kept above these critical values. Lower values will seriously affect the membrane absorber area and, thus, the overall investment costs. Figure 13 shows the effect of the forward reaction rate constant on the conceptual design. An increase of the forward reaction rate constant will decrease the absorber size. It is worthwhile mentioning that this effect is especially critical, even in the range from 1 to 0.1 m3/mol‚s. A decrease of the rate constant from 1 to 0.1 m3/mol‚s results in a drastic increase of the absorber length. Therefore, it is important to keep the forward reaction rate constant above 1.0 m3/mol‚s. This can be explained by the fact that a lower forward reaction rate constant results in a lower Hatta number and infinite enhancement factor and, consequently, reduced enhancement. An increase of the silver ion, silver complex, and propylene diffusion coefficients results in a decrease of the absorber length, although the effect is not critical. The results are similar for all components and only presented for propylene (Figure 14). Figure 15 shows the effect of the loading factor on the absorber dimensions. From Figure 15, it can be concluded that, with decreasing loading factor, the absorber column will be smaller, which seems logical because of the larger driving force. The total propylene loading of the silver solution, on the other hand, will be lower, resulting in a lower propylene capacity. One of the initial requirements was a propane purity of 99.5%. However, a reduction of the propane purity could result in a significant decrease of the required membrane area. On the other hand, it also results in a higher propylene loss. Figure 16 visualizes the effect of the propane purity on the system dimensions. It indicates that a propane purity above 99.0% shows a significant effect on the absorber length.

margin ($/ton)

payback period (year)

95 100 110 125 130 150 175

17 12 8 6 5 4 3

Process Flow Sheet On the basis of the design methods described by Coulson et al.,13 a preliminary design has been carried out, assuming composite hollow-fiber membranes with an external diameter of 0.5 mm and a fiber length of 1 m (Figure 17). Preliminary calculations show that a simple flash column at 353 K can perform the required desorption of propylene. During desorption, water evaporates, which contaminates the propylene. Cooling the gas flow of the first flash to 308 K will condense the water. Therefore, a second flash unit is selected to separate the majority of the water from the propylene. To achieve the required propylene specification, a flash is not sufficient and an additional adsorber column is selected to reduce the amount of water in the propylene product to the required ppm level. Fire and explosion index (F&EI) as well as HAZOP studies are performed for the absorber unit of the designed propylene/ propane separation. Most problems concern the operating conditions such as flow, temperature, pressure, and concentration and can be solved by adequate control systems and warning alarms. Using the previous criteria to check the hazards of our process, we see that the separation of propylene/propane using a hollow-fiber membrane contactor shows a moderate safety hazard. A detailed HSE study needs to be carried out during the latter stages of the project. Economic Evaluation The conventional method for separating a propylene/propane mixture is low-temperature distillation, which exhibits high operating and investment costs.2 For evaluation purposes, the costs of the newly designed membrane-based separation process are calculated, based on its payback period. The economic evaluation is based on the Peters and Timmerhaus method.18 Preliminary estimates have been used for the total product cost and total capital investment. Membrane costs and membrane life as well as adsorbent cost and its life are inside the battery limit. The program CAPCOST19 is used to estimate the bare equipment cost of the process. A plant cost index (CEPCI) of 51420 has been used. The membrane costs were taken from an informal discussion with industry and were assumed to be 100 $/m2. The total delivered cost of 10.9 M$ is taken as 1.0517 times the total equipment cost of 10.3 M$.


The total fixed capital investment is calculated with the cost of each cost component as a percentage of the total fixed capital investment, as presented in Table 7. The silver costs are taken as a fixed capital investment and are calculated from the holdup silver solution in the process. To correct for possible losses, an additional amount of 5% extra volume of silver is taken into account. The total silver costs are 9.5 M$, and the total fixed capital investment is 52.6 M$. Utility costs amount to ∼2.5 M$/yr. Since the exact costs of feed materials are not readily known, it is decided to calculate the profitability margin between the product and the feed materials for a reasonable payback period of 3 years. Calculations have been carried out according to standard methods18 by varying the price difference between the feed and the products. The calculations show that the difference between the feed and product costs should be at least 175 $/ton for economic operation and in order to have a payback period of 3 years. The results are summarized in Table 8. Conclusion A conceptual design for the separation of propylene and propane based on equilibrium calculations proved that the separation using a gas/liquid membrane contactor and a silver salt solution is technically feasible. Research in the literature showed that an AgBF4 solution exhibits the best properties as the absorption liquid. The conceptual design resulted in the use of a 5.5 M AgBF4 solution at an absorber temperature of 308 K in combination with a desorber temperature of 353 K. This gives the best separation of propylene and propane regarding cost and purity. Modeling of the gas/liquid membrane contactor showed that a membrane area of 80 000 m2 is required to achieve the required propane and propylene purities. Finally an economical evaluation showed that the difference between the feed and product prices should be at least 175 $/ton to obtain a payback period of 3 years. Symbols [ ] ) concentration (mol/m3) A ) cross-sectional area, gas-phase side (m2) a ) interfacial area (m2/m3) Cp ) heat capacity (J/mol·K) D ) diffusion coefficient (m2/s) E ) enhancement factor EA ) activation energy (J/mol) Fg ) gas feed flow (m3/s) Fl ) liquid flow (m3/s) FCI ) fixed capital investment ($) h ) step size (m) Ha ) Hatta number ∆Hr ) heat of reaction (J/mol) J ) flux (mol/m2‚s) k ) mass transfer coefficient (m/s) Kc ) equilibrium constant of complexation (m3/mol) Kd ) equilibrium constant of dissociation (mol/m3) Kf ) forward reaction rate constant (m3/mol‚s) L ) length of absorber (m) m ) distribution coefficient M ) molarity (kmol/m3) Mw ) mol weight (g/mol) n ) mole flow (mol/s) P ) pressure (Pa) Q ) volume flow silver salt solution (m3/s)

F ) density (kg/m3) R ) gas constant (J/mol‚K) T ) temperature (K) x ) mole fraction Subscripts A ) absorber c ) coating layer D ) desorber g ) gas if ) interface l ) liquid s ) support layer ∞ ) infinite Literature Cited (1) Nymeijer, K.; Visser, T.; Brilman, W.; Wessling, M. Analysis of the complexation reaction between Ag+ and ethylene. Ind. Eng. Chem. Res. 2004, 43, 2627. (2) Eldridge, R. B. Olefin/Paraffin Separation Technology: A review. Ind. Eng. Chem. Res. 1993, 32, 2208. (3) Nymeijer, K.; Visser, T.; Assen, R.; Wessling, M. Super selective membranes in gas-liquid membrane contactors for olefin/paraffin separation. J. Membr. Sci. 2004, 232, 107. (4) van den Berg, H. Process Plant Design Course; University of Twente: Enschede, Netherlands, 2004. (5) Nymeijer, D. C.; Visser, T.; Assen, R.; Wessling, M. Composite hollow fiber gas-liquid membrane contactors for olefin/paraffin separation. Sep. Purif. Technol. 2004, 37, 3, 209. (6) Gabelman, A.; Hwang, S. Hollow fiber membrane contactors. J. Membr. Sci. 1999, 159, 61. (7) Teramoto, M.; Matsuyama, H.; Yamashiro, T.; Katayama, Y. Separation of ethylene from ethane by supported liquid membranes containing silver nitrate as a carrier. J. Chem. Eng. Jpn. 1986, 19, 419. (8) Featherstone, W.; Sorrie, A. J. S. Silver-hydrocarbon complexes. J. Chem. Soc. 1964, 5235. (9) Herberhold, M. Complexes with mono-olefinic ligands. In Metal π-complexes; Elsevier publishing company: New York, 1972; Vol. II, Part I. (10) Baker, B. B. The effect of metal fluoroborates on the absorption of ethylene by silver ion. Inorg. Chem. 1964, 3, 200. (11) Fogg, P. G. T.; Gerrard, W. Solubility of gases in liquids; John Wiley & Sons: New York, 1991. (12) Atkins, P. W. Physical chemistry, 5th ed.; Oxford University Press: New York, 1995. (13) Coulson, J. M.; Richardson, J. F.; Sinnot, R. K. Chemical engineering, Vol. 6, Chemical Engineering Design; Butterworth-Heinemann: Woburn, MA, 2005. (14) Seider, W. D.; Seader, J. D.; Lewin, D. R. Product and Process Design Principles, Synthesis Analysis and EValuation; John Wiley & Sons, Inc.: New York, 2003. (15) Nymeijer, K. Gas/liquid membrane contactors for olefin/paraffin separation. Ph.D. Thesis, University of Twente, Enschede, Netherlands, 2003. (16) Versteeg, G. F. Mass transfer accompanied with complex reversible chemical reactions in gas-liquid systems: An overview. Chem. Eng. Sci. 1992, 47, 3181. (17) Chang, C. S.; Rochelle, G. T. Mass transfer enhanced by equilibrium reactions. Ind. Eng. Chem. Fundam. 1982, 21, 379. (18) Peters, M. S.; Timmerhaus, K. D. Plant Design and Economics for Chemical Engineers, fifth ed.; McGraw-Hill: New York, 2003. (19) Turton R.; Bailie R. C.; Whiting W. B.; Shaeiwitz J. A. Analysis, Synthesis and Design of Chemical Processes, second ed.; Prentice Hall: Upper Saddle River, NJ, 2003. (20) http://www.eng-tips.com/viewthread.cfm?qid)120468&page)1 (accessed April 2005).

ReceiVed for reView April 20, 2007 ReVised manuscript receiVed June 30, 2007 Accepted July 12, 2007 IE070556W

Liquid Membrane

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