propylene separation in hybrid membrane distillation systems: Optimization and economic analysis

Separation and Purification Technology 73 (2010) 377–390

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Ethane/ethylene and propane/propylene separation in hybrid membrane distillation systems: Optimization and economic analysis Marzouk Benali ∗ , Bora Aydin 1 Natural Resources Canada, CanmetENERGY, 1615 Lionel-Boulet Blvd., Varennes, Quebec J3X 1S6, Canada

a r t i c l e

i n f o

Article history: Received 16 September 2009 Received in revised form 27 April 2010 Accepted 28 April 2010 Keywords: Multicomponent distillation Ethylene Propylene Hybrid system Membrane separation Economic analysis

a b s t r a c t The combination of membrane and distillation processes to form a hybrid separation system is proposed as an alternative design to replace the current distillation technology. The main objective of this work is to scrutinize the feasibility of numerous hybrid membrane distillation (HMD) schemes through simulation. Silver nitrate is used for modeling/simulation as membrane carriers at a concentration of 6 mol L−1 . Different configurations are suggested to obtain highest product (ethylene or propylene) purity. These technologies are compared in terms of capital, operating and utility costs with the conventional C2- and C3-splitters. For the ethane/ethylene separation, the membrane cascade system resulted in the highest ethylene purity. However the series configuration is more economical than the membrane cascade system. For propane/propylene separation, the top configuration outperformed the conventional C3-splitter and other HMD configurations in terms of propylene purity. Crown Copyright © 2010 Published by Elsevier B.V. All rights reserved.

1. Introduction Ethylene (C2 H4 ) is one of the most important and one of the largest volume petrochemicals produced in the world today. The importance of ethylene is derived from the double bond in its molecular structure making it reactive and industrially convertible to a variety of intermediate and end products. It serves as the principal building block of the petrochemical industry. Worldwide demand for ethylene has grown steadily in the past and is expected to reach 140 million tons per year by 2010. The ethylene production system is very capital intensive; the capital cost is strongly dependent on the nature of the feedstock. In an ethylene plant, the feedstock is subjected to thermal cracking to produce a mixture of hydrocarbons ranging from hydrogen and methane to gasoline and heavier components. In general, the production of ethylene generates liquid, gaseous and solid wastes that have to be managed in a safe manner. The amounts and types of waste streams depend on the type of hydrocarbon feedstock. The major waste

Abbreviations: CEPCI, Chemical Engineering Plant Cost Index; FT, facilitated transport; HIDiC, heat integrated distillation column; ICIS, International Chemical Information Service; M1, membrane 1; M2, membrane 2; PR, Peng–Robinson; SCDS, simultaneous correction distillation column; SRK, Soave–Redlich–Kwong; VR, vapor recompression. ∗ Corresponding author. Tel.: +1 450 652 5533; fax: +1 450 652 5198. E-mail addresses: [email protected] (M. Benali), [email protected] (B. Aydin). 1 Present address: Delft University of Technology, Process and Energy Laboratory, Leeghwaterstraat 44, 2628 CA Delft, The Netherlands.

streams include caustic scrubber effluent, dilution steam condensate, coke and tare at the condensate separators, and system vent gases during start-ups and shutdowns. The waste gas streams are generally burned either in a furnace or in a flare. The flare is one of the most important and expensive parts of the ethylene plant facility, which has to be optimized. The optimization opportunities include design changes, product purity enhancement, energy reduction, production capacity increase, by-products minimizing, revamp economics, and reduction of greenhouse gases. Thermal cracking of propane and/or ethane as feedstocks in the presence of steam remains one of the most important and widely employed processes for ethylene production, which consists basically of four distinct processes, namely: thermal cracking and quenching, compression and acid gas removal (i.e. H2 S, CO2 ), subcooling and product separation, and refrigeration. Fig. 1 illustrates a block flow diagram for a typical ethylene plant. The feedstocks are fed to a bank of parallel pyrolysis furnaces. In the convection zone of the furnace, the feed is preheated to about 600 ◦ C and is then diluted with steam that reduces coking and improves product selectivity. In the radiation zone of the furnace, the feedstock–steam mixture passes through vertical coils where pyrolysis takes place at temperatures above 600 ◦ C. At the exit of the cracking furnace, the outputs are immediately quenched to about 350 ◦ C in a transfer-line exchanger to stop reactions and recover the waste heat for steam generation. The cracked gases are cooled to about 40 ◦ C in a water quench tower to condense the heavy products (e.g. fuel oil) and most of the dilution steam. The cooled gases are then compressed to about 3.5 MPa in four compression stages. Between the 1st and 3rd compression stages, the gases are scrubbed with caustic gas to remove H2 S and

1383-5866/$ – see front matter. Crown Copyright © 2010 Published by Elsevier B.V. All rights reserved. doi:10.1016/j.seppur.2010.04.027

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Nomenclature Ci D F J Keq kH  Mcut n˙ P p Pc RR R S T Tc Tr V xFi xPi xRi z

concentration of component i (mol m−3 ) diffusion coefficient (m2 s−1 ) feed flow rate (mol s−1 ) flux (mol m2 s−1 ) equilibrium constant (m3 mol−1 ) Henry’s constant (mol m−3 Pa−1 ) thickness of membrane (m) membrane module cut rate molar flow rate (mol s−1 ) permeate flow rate (mol s−1 ) pressure (Pa) critical pressure (Pa) retentate in Eq. (7) (mol s−1 ) gas constant in Eqs. (16)–(18) membrane surface area (m2 ) temperature (K) critical temperature (K) reduced temperature volume (m3 ) mole fraction of component i in the feed mole fraction of component i in the permeate mole fraction of component i in the retentate membrane axis

Greek letters ˛i/j separation factor ε porosity  tortuosity ω acentric factor (−log(P/Pc )Tr =0.7 ) − 1 Subscripts c critical value eq equilibrium s–e silver–ethylidene tot total

CO2 . After the 4th compression stage, the gases are cooled to about 15 ◦ C with propylene refrigerant and dehydrated with molecular sieves. The dried gases are cooled to low temperatures in a series of heat exchangers before they enter the separation section. The bottoms of the primary and secondary deethanizers are carried to a depropanizer where C3 components are separated from the heavier components (C4+). The secondary deethanizer overhead is hydrogenated in a catalytic reactor (diameter and volume are 1.818 m and 5.094 m3 , respectively) to convert acetylene into ethylene. As described above, in the purification/separation subsystem of the plant, ethylene is separated from its by-products via a sequence of fractionation steps. Among these steps, several require a cascade of propylene/ethylene refrigeration loops. Propylene refrigerant is used in the chilling–demethanizer subsystem. When the relative volatility between the key components is less than 1.2, the separation process becomes difficult. Subsequently, the distillation in this case requires high-energy consumption, increased refrigeration capacity as well as an increased number of stages, leading obviously to enhanced capital and operating costs. About 70% of the required energy is consumed in the purification/separation subsystem. Therefore, any reduction in the refrigeration load will result in a significant decrease of the operating cost of the propylene refrigeration closed-cycle system as well as the entire plant. As stated in the literature [1], ethane/ethylene, and propane/propylene separations are potential steps for an energy saving of 33% through hybrid technologies involving both membranes and distillation processes. The selection of the best technologies for separation of ethylene from ethane, and propylene from propane is therefore consequential to minimize plant energy consumption and capital cost. Different technologies were investigated to intensify the traditional distillation processes. Ghosh et al. [2] investigated the potential of a hybrid adsorption–distillation system for propane/propylene separation. According to their findings, although there is a reduction in energy consumption there is a need for an innovative adsorption–desorption process or for adsorbents having high selectivity. Schmal et al. [3] compared a heat integrated distillation column (HIDiC) with a vapor recompression (VR) distillation column. They found HIDiC to be 14% more economical than VR in operating cost. Finally, membranes combined with distillation columns were demonstrated in the literature [4–8] as a technolog-

Fig. 1. Block flow diagram of a typical ethylene plant.

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poses and, subsequently, releases C2 H4 in the permeate and the carrier Ag+ , which is then returned to the membrane feed side. The chemical equilibrium can be written as follows: Ag+ + C2 H4  (Ag · C2 H4 )+ complex (aq)

(1)

The equilibrium constant is Keq =

Fig. 2. Schematic diagram of the facilitated transport membrane system.

ical option to optimize the performance of hydrocarbon separation and purification. In this work, membrane separations, which are generally less energy-intensive than distillation processes, have been considered as promising alternatives to enhance the purity of ethylene and propylene as well as to reduce the capital and operating costs of separation processes. In general, such an approach can be also applied for separation of various close-boiling hydrocarbons mixtures as styrene/ethylbenzene, and styrene/xylene isomers. 2. Methods and design of hybrid separation flow diagrams Membrane separation technologies have several advantages over existing mass transfer processes including high selectivity, low energy consumption and simple design [8]. The hybrid membrane processes can be designed in two different configurations: a membrane process combined with a conventional separation process and a membrane process combined with another membrane process. The combination of membrane and distillation technologies to form a hybrid separation system is proposed as an alternative design to replace the current distillation technology. Facilitated transport (FT) membrane technology is used to design several hybrid membrane distillation (HMD) configurations studied in this work. The appropriateness of FT membranes for ethylene and propylene recovery has been proven in [1], and [9–13]. FT membranes involve carrier-mediated transportation in addition to permeate physical dissolution and diffusion. The presence of a carrier that can react reversibly with permeate results in high selectivity and high permeability. There are two types of FT membranes: one is a mobile carrier membrane (liquid membrane) where the carrier can diffuse into the membrane; the other is a fixed carrier membrane where the carrier is immobilized in the membrane [14]. The main objective of the present work is to scrutinize the feasibility of numerous HMD schemes. The hybrid configurations differ in the location of the membrane relative to the distillation column. Top, parallel, bottom, top-bottom and series configurations, along with a membrane cascade system with recycling and without a distillation step are investigated. Agrawal [15] already discussed different possible membrane cascade configurations to separate multicomponent gaseous mixtures. These configurations were drawn based on the analogy with multicomponent distillation schemes.

] [Ag · C2 H+ 4

[Ag+ ][C2 H4 ]

=

Cs–e CAg+ CC2 H4

(2)

The equilibrium constant of the silver–ethylidene ion formation has been evaluated by several authors [16–18]. Obviously, such a constant varies with the concentration of AgNO3 solution. For a given temperature and concentration of AgNO3 , the higher the C2 H4 pressure, the greater the amount of C2 H4 absorbed in the form of the silver–ethylidene complex. On the other hand, for a given concentration of AgNO3 , the lower the temperature, the greater the amount of C2 H4 absorbed in the form of the silver–ethylidene complex. The total carrier concentration within the membrane module is obtained from the following balance: Ctot = CAg+ + Cs–e

(3)

Combining Eqs. (2) and (3), the concentration can be calculated as follows: Cs–e =

Keq CC2 H4 Ctot 1 + (Keq CC2 H4 )

(4)

By assuming that the equilibrium occurs at the feed side and the permeate side, the boundary concentrations of silver–ethylidene ion are determined as a function of known boundary concentrations of ethylene:

⎧ Keq CC2 H4 (z=0) Ctot ⎪ ⎪ ⎨ Cs–e(z=0) = 1 + (Keq CC H ) 2 4(z=0) ⎪ ⎪ ⎩ Cs–e(z=) =

Keq CC2 H4 (z=) Ctot

(5)

1 + (Keq × CC2 H4 (z=) )

where CC2 H4 (z=0) and CC2 H4 (z=) correspond to the solubility of C2 H4 at the feed side and the permeate side, respectively. These boundary concentrations are determined using Henry’s law, expressing the pressure effects on the solubility of a gas in a saturated solution:



CC2 H4 (z=0) = (kH )C2 H4 pC2 H4 (z=0) CC2 H4 (z=) = (kH )C2 H4 pC2 H4 (z=)

(6)

2.2. Membrane model The membrane model is based on the following assumptions: -

Cross-flow along the permeate side of the membrane Plug flow along the feed side of the membrane Steady state process Isothermal operation Instantaneous ethylene/silver reaction. The local permeate-side composition is a function only of the local feed-side composition.

2.1. Properties and equations of FT membrane separation process Fig. 2 shows the general principle of an ethane (C2 H6 )/ethylene (C2 H4 ) FT membrane system consisting of an aqueous solution of silver nitrate (AgNO3 ) transported by capillarity in a microporous membrane. At high pressure, C2 H4 reacts with silver ion (Ag+ ) to form a complex component named silver–ethylidene ion, which diffuses freely across the membrane. Once this complex reaches the opposite low pressure side of the membrane, the complex decom-

The cross-flow pattern is used along all HMD configurations. The permeate composition changes along the membrane length and it is not affected by the composition of any other point along the membrane side. It is only affected by the local flux at that point. The overall mass balance for a membrane module is:



F = P + RR FxFi = PxPi + RR xRi

R

(7)

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The overall mass balance across a differential area of a membrane module is given by

 dn˙ = −J = − Ji dS n

(8)

i=1

The component mass balance is therefore given by ˙ Fi ) d(nx ˙ Fi ndx dS

2.3. Ethane/ethylene hybrid separation

= −Ji

dS +

(9)

xFi dn˙

= −JxPi

dS

(10)

The combination of Eqs. (8) and (10) gives: dxFi dS

=

J i (x − xPi ) n˙ F

(11)

In contrary to distillation, in which separation is based on thermodynamics, the separations are inherently based on rate of transport, i.e. they depend on diffusion. The basic flux across the membrane is equal to the flux across a thin film as proposed by Whitman [19]. For ethylene, the flux ji becomes JC2 H4 , which is proportional to the concentration gradient and obeys to Fick’s law: JC2 H4 =

DC2 H4





CC2 H4 (z=0) − CC2 H4 (z=)



(12)

The quantity C2 H4 / obviously corresponds to a mass transfer coefficient. Due to the microscopic properties of the membrane (i.e. porosity and tortuosity of the pores) and considering the film model [19], this coefficient can be defined as follows: DC2 H4 

≡

DC2 H4 ε  

(13)

In the case of a FT membrane, the flux should include the transport diffusion related to the silver–ethylidene ion. Eq. (12) becomes: JC2 H4 =

DC2 H4 ε  CC2 H4 (z=0) − CC2 H4 (z=)   +

Ds–e ε  Cs–e(z=0) − Cs–e(z=)  

(14)

Considering that the membrane separation feature is strongly influenced by the partition of the solute (i.e. C2 H4 ) between the membrane and the adjacent solution (i.e. AgNO3 ), substitution of Eqs. (5) and (6) in Eq. (14) gives: JC2 H4 =

DC2 H4 ε  





(kH )C2 H4 pC2 H4 (z=0) − pC2 H4 (z=)



+

Ds–e ε  

into CHEMCAD is based on the following sequence: load CHEMCAD, put data into CHEMCAD and run, load data. The Visual Basic server interface allows starting CHEMCAD, loading a simulation case, and transferring information back and forth from Excel in an automated fashion.



Keq Ctot (kH )s–e

The HMD system may be arranged in several configurations. Operating conditions may be varied for each configuration in order to achieve optimum performance. However, the concentration of silver nitrate is maintained constant at 6 mol L−1 . A series configuration is investigated in which two membranes in series are followed by a distillation column. In addition, a membrane cascade system is explored as an alternative to a conventional C2-splitter to recover ethylene from ethane. Fig. 3 illustrates a two-stage membrane cascade based on two single-stage compressors for recycling and a multistage compressor required to compress the ethylene product to the desired pressure, as in the case of a conventional C2-splitter. The fresh stream (208, coming from the demethanizer), composed of 75.66% C2 H4 , 12.61% C2 H6 , 8.58% C3 H6 , 1.34% C3 H8 , 1.80% CH4 , and 0.01% C3 H4 enters at a pressure of 2.61 MPa and a temperature of 28 ◦ C, and then is mixed with the recycled stream (207 and thereafter is fed to the first membrane (M1) at a pressure of 2.61 MPa and temperature of 31 ◦ C. In M1, the recycled ethane is separated as the retentate and directed to the cracking reactor. The recovered ethylene is concentrated as permeate at a pressure of 0.83 MPa and directed, after a compression step, to the second membrane M2 as a feed at a pressure of 1.34 MPa and a temperature of 33 ◦ C. The entire amount of C3 H8 , C3 H6 , CH4 , and C3 H4 is recovered in the retentate (stream 201). Thus, the stream 212 contains only 99.99% C2 H4 and 0.01% C2 H6 . High pure C2 H4 is therefore produced in M2 at a permeate pressure of 0.10 MPa and a temperature of 20 ◦ C. The retentate of M2 is compressed from 1.34 to 2.62 MPa before being mixed with the fresh feed. The surface area of M1 is obviously greater than the surface area of M2 since the highest recovery of C2 H4 occurs in the first separation stage. The ratio of surface area of M2 to surface area of M1 has been varied from 0.10 to 0.80, and the optimal value is 0.40 corresponding to surface areas of 3243 and 8150 m2 , respectively. Fig. 4 illustrates the series configuration of the hybrid system. In this case, the fresh feed (3.50 MPa; −4 ◦ C) is first flashed at a prespC2 H4 (z=0)

1 + [Keq (kH )C2 H4 pC2 H4 (z=0) ]

−

pC2 H4 (z=) 1 + [Keq (kH )C2 H4 pC2 H4 (z=) ]

 (15)

For ethane, which does not react with silver ion, the flux is given by: JC2 H6 =

DC2 H6 ε  





(kH )C2 H6 pC2 H6 (z=0) − pC2 H6 (z=)



(16)

For C3 H6 /C3 H8 separation, similar equations can be written to calculate the local composition of each component in the permeate side. The physicochemical properties of ethylene and propylene in an aqueous silver solution are summarized in Table 1. The membrane porosity, also termed the membrane void volume, is an important parameter affecting the membrane flux. Membranes with higher porosity exhibit greater surface area. The membrane porosity usually lies between 30 and 85%. The membrane tortuosity is the average length of the pores compared to membrane thickness. It is in the order of 2–4. The thickness of the FT membrane is 170 ␮m and the ratio of porosity to the tortuosity used for both C2 H4 /C2 H6 and C3 H6 /C3 H8 separation is 0.258. Eqs. (8) and (11) are solved numerically. The composition of permeate as well as the flux of permeate are calculated from transport equations (14) and (15). The numerical data are incorporated into CHEMCAD [22] by using “Excel Integration” feature. The integration

sure of 2.61 MPa to reach a temperature of −16 ◦ C required to be sent to two membranes in series. For optimal membrane cut rates of 0.99 and 0.90, respectively for M1 and M2, the heat duties of exchangers 102 and 103 are—0.67 and 8.46 MW, respectively. In case of series configuration, the permeate (stream 107) of membrane M1 contains 78.08% C2 H4 , 7.86% C2 H6 , 10.29% C3 H6 , 1.60% C3 H8 , 2.16% CH4 , and 0.01% C3 H4 while the permeate (stream 108) of membrane M2 contains only 99.99% C2 H4 and 0.01% C2 H6 . The resulting permeate is recovered as ethylene product and the retentate leaves the second membrane at a pressure of 0.66 MPa and a temperature of 50 ◦ C, and sent as the feed to the distillation column. A high pressure difference is required to provide an enhanced driving force as well as a high feed concentration to produce high purity ethylene product. The feed to the first membrane stage comes from the bottom of the demethanizer and flashes to a moderate pressure before entering the membrane. The refrigeration load associated with the feed stream is exchanged with other process streams to minimize the overall quantity of propylene refrigerant used in the

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Table 1 Physicochemical properties of components involved in a membrane module. Component C2 H4 C2 H6 (Ag·C2 H4 )+ Ag+ C3 H6 C3 H8 (Ag·C3 H6 )+

Diffusivity (m2 s−1 ) −9

1.87 × 10 – 1.13 × 10−9 1.66 × 10−9 1.63 × 10−9 1.61 × 10−9 1.06 × 10−9

[20] [present work] [20] [21] [21] [21]

Chemical equilibrium constant (m3 mol−1 )

Henry’s constant (mol m−3 Pa−1 )

– – 0.129 [present work] – – – 0.100 [present work]

5.61 × 10−5 2.39 × 10−5 – – 6.17 × 10−5 2.15 × 10−5 –

[present work] [present work]

[present work] [present work]

Fig. 3. Two-stage membrane cascade system for ethane/ethylene separation.

plant. The feed temperature to the membrane system is set to maintain the membrane operation above the freezing point of the silver nitrate solution. The feed from the first stage is concentrated by removing part of the recycled ethane as a retentate. The feed is then directed to the second membrane where the separation yields a

high purity ethylene product. The retentate from the second membrane is directed to a low pressure column while high pressure distillation column is used in the conventional C2-splitter. The key benefits of such a configuration are downsizing the column, reducing the condenser duty, and minimizing the amount of refrigerant.

Fig. 4. Hybrid membrane distillation series configuration for ethane/ethylene separation.

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Table 2 Design details of the conventional C3-splitter (two columns in parallel).

Number of trays Tray efficiency Column diameter (m) Column pressure (MPa) Column reflux ratio

Column 1

Column 2

115 0.85 5.30 1.64 20.33

90 0.85 5.30 1.72 507.51

The diameter of the distillation column and the number of stages are reduced by 67.0% and 78.7%, respectively, in comparison to a conventional C2-splitter system. 2.4. Propane/propylene hybrid separation Four different configurations were tested for propane/propylene separation: top, parallel, bottom and topbottom. The concentration of silver nitrate is 6 mol L−1 . The membrane can be located at the top of the distillation column (Fig. 5). In this case, the membrane performs the final purification. The overhead of distillation column is fed to the membrane at 1.64 MPa and a temperature of 40 ◦ C. The retentate is recycled and fed at the middle of the distillation column at a pressure of 1.66 MPa and a temperature of 48 ◦ C while the permeate pressure is 0.10 MPa. The pressure drop along the feed side of the membrane is assumed negligible. The membrane may also be combined with the distillation column in a parallel configuration (Fig. 6). In such a configuration, the membrane feed stream is withdrawn from one of the intermediate stages of the distillation column, more specifically at the 40th stage. The permeate and retentate streams leaving the membrane are compressed and fed back to the distillation column at 20th and 60th stages, respectively. These two streams enter the distillation column where the stream composition is the same as the tray composition. In the bottom configuration, the membrane performs the final purification of the propane sent from the distillation column (Fig. 7). Refrigerant is required to cool the permeate to its dew point before entering the distillation column. This configuration differs from the top one in that a larger amount of refrigerant is required for the permeate condenser. The feed to the membrane is a saturated liquid. Fig. 8 illustrates the top-bottom configuration in which one membrane is located at the top of the distillation column while another one is located at the bottom of the column. The optimal conditions of top-bottom HMD configuration corresponds to the permeate pressure set at 0.34 MPa for the top membrane and 1.66 MPa for the bottom membrane. The separation within the top membrane is in gas permeation as the membrane feed is a saturated vapor. The bottom membrane is operated in a pervaporation mode since the membrane feed is a saturated liquid. After a compression step, the retentate of top membrane is fed at the 10th stage (1.67 MPa; 44 ◦ C) of the distillation column while the permeate of bottom is fed at the 2nd stage (1.09 MPa; 159 ◦ C). In all HMD systems associated with C3 H6 /C3 H8 separation, the diameter of the distillation column is kept the same (i.e. 5.30 m) as in the conventional C3-splitter for retrofit purpose whereas it is reduced by 25.7% for design purpose of a new plant. For all configurations, the fresh feed stream entering the hybrid separation system is obviously the same as the one of the conventional C3-splitter and is composed of 73.37% C3 H6 , 14.83% C3 H8 , 0.03% C2 H6 , 0.20% C2 H2 , 1.80% CH4 , and 0.01% C3 H4 , 7.89% 1,3C4 H6 , 1.64% 1,-C4 H8 , 0.11% n-C4 H10 , and 1.92% i-C4 H10 . The design details of the conventional C3-splitter are given in Table 2. Each hybrid system has its distinctive minimum reflux ratio. Two main membrane features have a significant effect on the selectivity and the flux rate: silver nitrate concentration and permeate pressure.

For the top configuration, the surface area of membrane was varied from 100 to 1000 m2 whereas the for bottom and parallel configurations it was ranged from 400 to 4000 m2 . The optimal values are given in Table 5 for each HMD system. 2.5. Separation factor of membrane module The parameter to describe the separation efficiency for ethane/ethylene, and propane/propylene mixtures fed to the membrane is the separation factor, ˛i/j , which is also called the selectivity factor of the membrane. It can be formulated as below: ˛i/j =

xiP /xjP

(17)

xiF /xjF

The membrane module cut rate is one of the major independent variables for HMD systems. In practical terms, this variable represents the ratio of how much permeate is recovered per unit feed to the membrane module. There is always a trade-off between the purity and the flow rate of permeate and retentate streams. For a given membrane module cut rate, the flow patterns have a significant effect on the degree of separation as well as on the membrane area required to do so. When a perfect mixing is assumed on both sides of the membrane module, the degree of separation decreases with an increase of the membrane module cut rate. For the crossflow, as considered in the present work, the feed flows through the upstream of membrane surface in plug flow with no longitudinal mixing. There is no flow of permeate along the membrane surface. Thus, the degree of separation increases very slightly (in the order of magnitude of 2–10%) at low membrane module cut rate (typically 0 ≤ ˛i/j ≤ 0.5). As ˛i/j  0.5, the degree of separation increases significantly. 3. Results and discussion The distillation column is simulated with CHEMCAD using the simultaneous correction distillation (SCDS) column type and Soave–Redlich–Kwong (SRK) state equation. SCDS is mainly designed to simulate non-ideal K-value chemical systems. It calculates the derivatives of each equation rigorously, including the derivative of K-value with respect to composition (ıK/ıX) ( ) term, which is significant in chemical system simulation. The equations related to SRK model are given as follows: p=

RT a − 2 V −b V + bV

(18)

where a = 0.42748

R2 Tc2 [1 + f␻ (1 − Tr0.5 )] pc

b = 0.08664

RTc pc

fω = 0.48 + 1.574ω − 0.176ω2

2

(19) (20) (21)

The simulation and optimization of the HMD system use two limiting permeate pressures of 0.10 and 0.34 MPa. The lower value of 0.10 MPa provides high selectivity which minimizes the membrane area required by the HMD system. However, the compression duty was observed to increase when 0.10 MPa was selected as a permeate pressure. Contrarily, the compressor duty load decreased with the upper permeate pressure limit (0.34 MPa). On the other hand, lower selectivity (due to lower compressor load) leads to increased membrane area. The economic performance of various membrane hybrid configurations and membrane cascade system has been assessed and compared to an ethylene production plant with a yearly throughput of 500,000 metric-tons.

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Fig. 5. Top hybrid configuration for propane/propylene separation.

3.1. HMD system analysis 3.1.1. Ethane/ethylene separation HMD systems are characterized by smaller equipment sizing and relatively low operating cost. As seen in Table 3, for the series configuration hybrid system, the number of columns, number of trays and column diameter are lower than in conventional C2-splitters. This results in a lower operating pressure and higher reflux ratio. The optimal configuration could be defined as producing the desired component at the highest purity and lowest operation cost. As explained in detail in the next section,

Table 3 Comparison of the design specifications of the conventional C2-splitter and series configuration hybrid systems. Parameters

Conventional C2-splitter

Series configuration hybrid system

Number of columns Number of trays per column Tray efficiency Column diameter (m) Column pressure (MPa) Column reflux ratio Product purity

2 94 0.85 4.545 2.00 0.09 0.9965

1 40 0.85 1.500 0.65 1.38 0.9979

Fig. 6. Parallel hybrid configuration for propane/propylene separation.

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Fig. 7. Bottom hybrid configuration for propane/propylene separation.

high product purity could be obtained with HMD systems at a lower cost in comparison to the conventional C2-splitter configuration. Fig. 9 shows the effect of membrane module cut rate on the flow rate of C2 H4 in the product and recovery lines (stream 113) for the series configuration. In all HMD and membrane cascade systems, the membrane module cut rate is defined as the ratio of the permeate to the membrane feed flow rate. The flow rate in the product line increases with the increasing module cut rate of the first membrane at constant module cut rate of the second membrane. At constant module cut rate of the first membrane the effect of the module cut rate of the second membrane on the flow rate of C2 H4 in product line is minor. Higher module cut rate signifies higher amount of C2 H4 in the permeate (stream 107) of the first membrane. Thus, higher amount of C2 H4 is sent to the product line via the permeate of the second membrane. In the recovery

line, the flow rate of C2 H4 decreases with the module cut rate of the first and second membrane. A minor effect of the module cut rate on the recovery line is observed for the second membrane at the cut rate values higher than 0.7. However the second membrane should be operated at a membrane module cut rate of 0.9 to achieve maximum throughput of C2 H4 in the product line. Fig. 10 shows the effect of the membrane module cut rate on the separation factor for the series configuration. The membrane module cut rate is expressed as the ratio of the permeate flow rate to the feed flow rate. The mole fractions of C2 H4 and C2 H6 in the permeate and feed were calculated for each membrane module cut rate to determine the separation factor, which is directly proportional to the mole fraction of C2 H4 in the permeate, i.e. enrichment of C2 H4 . The separation factor increases linearly with the membrane module cut rate for both membranes. High separation factor values for the second membrane can be essentially explained with higher mole fraction

Fig. 8. Top-bottom hybrid configuration for propane/propylene separation.

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Fig. 9. Effect of first membrane module cut rate (on the flow rate of C2 H4 in the product and recovery lines for the series configuration.

Fig. 11. Influence of membrane module cut rate of (a) first membrane and (b) second membrane on the flow rate of C2 H4 in the product and recovery lines for the cascade system.

Fig. 10. Membrane module cut rate versus separation factor for the series.

of C2 H4 and lower mole fraction of C2 H6 in the permeate (stream 108) of the second membrane. The effect of membrane module cut rate on the flow rate of C2 H4 in the product and recovery lines (stream 212) for the cascade system is shown in Fig. 11(a) and (b). The flow rate of C2 H4 in the permeate of the first membrane increases with its module cut rate. The permeate of the first membrane (stream 202) is fed to the second membrane and then it is sent to the product line via the permeate of the second membrane (stream 210). At a given module cut rate of the second membrane the flow rate of C2 H4 in the product line increases with the module cut rate of the first membrane (Fig. 11a). Additionally the flow rate of C2 H4 in the product line increases with the module cut rate of the second membrane due to higher flow rate of C2 H4 in the permeate. In the recovery line the flow rate decreases abruptly with the membrane module cut rate for both membranes. For the optimum module cut rate of the second membrane (0.90), an increase in the module cut rate of the first membrane from 0.45 to 0.76 leads to a decrease of 99.7% of C2 H4 in the recovery line (Fig. 11b). Similarly for the second membrane; an increase of the module cut rate from 0.50 to 0.90 results in a decrease of 85.3% of C2 H4 in the recovery line which confirms the separation efficiency of both membranes. Fig. 12 shows the effect of the membrane module cut rate on the separation factor for the cascade system. The separation factor increases linearly with the membrane module cut rate for both membranes. The mole fraction of C2 H4 in the permeate increases while the mole fraction of C2 H6 in the permeate decreases with the increasing membrane module cut rate. Higher separation factor values for the first membrane are

Fig. 12. Membrane module cut rate versus separation factor for the cascade system.

due to lower mole fraction of C2 H4 and higher mole fraction of C2 H6 in the feed (stream 209) of the first membrane. The effect of membrane module cut rate on the total actual compressor power and heat exchanger duty for the series configuration is illustrated in Fig. 13. The heat exchanger duty increases with increasingly membrane module cut rate for both membranes where the actual compressor power decreases with the membrane module cut rate. The effect of membrane module cut rate on the actual compressor power is more pronounced for the second membrane. Fig. 14 shows the influence of the membrane module cut rate on the total actual compressor power and heat exchanger duty for the cascade system. The actual compressor power and heat exchanger duty increase when the membrane module cut rate increases for both membranes where the effect of the membrane module cut

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Fig. 13. Effect of first membrane module cut rate (a) and second membrane module cut rate (b) on the actual compressor power and heat exchanger duty for the series configuration.

Fig. 14. Effect of first membrane module cut rate on the actual compressor power and heat exchanger duty for the cascade system.

Fig. 15. Influence of membrane module cut rate on the flow rate of C3 H6 in the product and recovery lines for top configuration.

Fig. 16. The flow rate of C3 H6 in the product and recovery lines for different membrane module cut rate values of bottom configuration.

rate is more pronounced for the first membrane. The significant increase of the total actual compressor power and heat exchanger duty is essentially due to the increase of 42.1% in the power of the compressor 208 and the duty of the heat exchanger 209 located on the product line (stream 212). Similar analysis for the second membrane shows that the increase in the compressor power and the heat exchanger duty is 24.3%. The corresponding increase of the flow rate of C2 H4 in the permeate is 24.3%. 3.1.2. Propane/propylene separation A systematic analysis as above was performed for the propane/propylene separation. The composition of the product and recovery lines as well as compressor and heat exchanger duties were investigated for all HMD configurations at different membrane module cut rate values. Fig. 15 shows the influence of membrane module cut rate on the flow rate of C3 H6 in the product and recovery lines (streams 109 and 102, respectively) for the top configuration. The flow rate increases with increasing membrane module cut rate. The overhead of the distillation column is the feed of the membrane. The flow rate of C3 H6 depends only on the membrane selectivity which could be given as an explanation for the linear relation between the C3 H6 flow rate and membrane module cut rate. As shown in Fig. 16, the flow rate of C3 H6 in the product line (stream 206) increases whereas C3 H6 flow rate in the recovery line (stream 203) decreases sharply with membrane module cut rate for bottom HMD configuration. The bottoms of the distillation unit is fed to the membrane of which separation efficiency is measured by the amount of C3 H6 recovered and recycled to the distillation

Fig. 17. The flow rate of C3 H6 in the product and recovery lines for different membrane module cut rate values of parallel configuration.

column through the permeate of the membrane (stream 204). As the membrane module cut rate increases, higher flow rate of C3 H6 in the permeate is obtained, which results in the decrease of C3 H6 flow rate in the recovery line by 98.5%. The permeate stream 204 is obviously recycled to a stage in C3-splitter where the composition is the same as that of permeate. Fig. 17 exhibits the effect of membrane module cut rate on the flow rate of C3 H6 in the product and recovery lines (streams 308 and 309 respectively) for the parallel configuration. It is clearly indicated that the membrane module cut rate does not affect sig-

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387

Fig. 19. Membrane module cut rate versus separation factor for top, parallel and top-bottom configurations.

Fig. 18. Effect of first membrane module cut rate (a) and second membrane module cut rate (b) on the flow rate of C3 H6 in the product and recovery lines for top-bottom configuration.

nificantly the amount of C3 H6 in the product and recovery lines: the increase of the membrane module cut rate from 0.49 to 0.78 results in only an increase of 0.34% and 0.87% for the product and recovery lines respectively. The minimum number of stages occurs when the membrane feed composition is close to the composition of C3-splitter feed stage. Fig. 18a and b depicts the effect of membrane module cut rate on the flow rate of C3 H6 in the product and recovery lines (streams 410 and 403 respectively) for the top-bottom HMD configuration. The flow rate in the product and recovery lines increases with increasing membrane module cut rate for the first membrane (Fig. 18a). However the flow rate of C3 H6 in the recovery line is very low and the increase is inconsequential. C3 H6 is sent to the product line via the permeate stream of the first membrane (stream 407). Thus, C3 H6 flow rate in this permeate increases with the module cut rate of the first membrane. The flow rate of C3 H6 in the product line increases and it decreases in the recovery line with increasingly membrane module cut rate for the second membrane (Fig. 18b), which favours the recycle of C3 H6 to the distillation column. The permeate stream 404; whose composition is proportional to the membrane module cut rate, is fed to the 2nd stage of the distillation column. As most C3 H6 is sent to the column via permeate, the flow rate of C3 H6 in the recovery line decreases as membrane module cut rate increases to maximize the yield of the polymer grade propylene. The effect of the membrane module cut rate on the separation factor for top, parallel and top-bottom configurations is shown in Fig. 19. The separation factor values for the top-bottom configuration belongs to the membrane placed at the top of the column. It should be indicated that the similar separation factor values for top and parallel configurations does not result from similar mole fraction values of C2 H4 and C2 H6 in the permeate and feed. The sep-

Fig. 20. Effect of membrane module cut rate on the compressor power and heat exchanger duty for top configuration.

Fig. 21. Compressor power vs. membrane module cut rate for bottom configuration.

aration factor values for the top-bottom configuration are higher than that for the top configuration. This confirms the separation efficiency of the membrane placed at the top of the column for the top-bottom configuration. The effect of membrane module cut rate on the actual compressor power and heat exchanger duty for different HMD configurations is illustrated in Figs. 20–23. For the top configuration, the actual compressor power and heat exchanger duty augment with membrane module cut rate (Fig. 20). Higher membrane module cut rate results in a higher C3 H6 concentration in the

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from the bottoms of the column should be recycled to the column via the permeate of the second membrane; higher membrane module cut rate relatively higher flow rate of C3 H6 requires higher compressor power and heat exchanger duty. 3.1.3. Economic analysis For ethane/ethylene separation with the membrane cascade system, ethylene purity of 0.9999 is obtained and the cost saving is found to be 54.41%. For the series configuration a purity of 0.9979 and a cost saving of 61.45% are found (Table 4). The saving is formulated as the ratio of total investment cost difference between a hybrid configuration and a conventional column to the total investment cost of a conventional column. The expression for the total investment cost is given below: Total investment cost = Capital cost + Operating cost + Utility cost Fig. 22. Effect of membrane module cut rate on the total actual power and heat exchanger duty for parallel configuration.

Fig. 23. Effect of first membrane module cut rate on the actual compressor power and heat exchanger duty for top-bottom configuration.

permeate. The increase in the total actual compressor power is due to increase in the actual power of the compressor located in the product line (equipment 105). For an increase of the membrane module cut rate from 0.54 to 0.85 the compressor actual power increases by 35.21% whereas the heat exchanger duty increases, due to higher C3 H6 flow rate in the permeate (stream 106), by 15.6%. As seen in Fig. 21, the actual compressor power increases with the increase in membrane module cut rate for the bottom configuration. Fig. 22 shows the effect of membrane module cut rate on the total actual power of the compressor and heat exchanger duty for the parallel configuration. The actual power decreases by 79.58% increasing membrane module cut rate in contrast with the heat exchanger duty which increases by 36.0%. The substantial decrease in the total actual power is mainly explained by the tremendous decrease of 88.7% in the actual power of the compressor 305 located on the retentate line as the membrane module cut rate increases from 0.49 to 0.78. The effect of the membrane module cut rate on the total actual compressor power and heat exchanger duty for the top-bottom configuration is shown in Fig. 23. The actual compressor power and heat exchanger duty increase with the membrane module cut rate for the first membrane. The flow rate of C3 H6 in the permeate increases with increasing first membrane module cut rate. Obviously, higher compression power (equipment 405) and heat exchanger duty (equipment 406) is required for higher flow rate of C3 H6 . The power and duty values are increasing function of membrane module cut rate for the second membrane. The permeate of the second membrane is composed of C3 H6 . As all C3 H6 coming

(22)

where Capital cost = 1.82 × Installed cost

(23)

Operating cost = 0.20 × Installed cost

(24)

The installed cost was calculated with CHEMCAD by introducing January 2009 Chemical Engineering Plant Cost Index (CEPCI). The cost estimation calculation is based on the area of the equipment. A rigorous costing routine was defined by the Parser programming language for each equipment separately. Capital and operating costs were calculated based on the definitions given by Douglas [23]. The utility cost consists of electricity and steam prices for each separation system. The cost of ethylene and propylene considered in this work are 1.65 and 1.85 $ kg−1 (2008-basis), respectively. The prices of these two products were taken from the ICIS (International Chemical Information Service) database [24]. The membrane module cost was taken as 215 $ m−2 including all related elements. The membrane replacement cost was assumed to be 95 $ m−2 for a lifetime of 3 years. The comparison of the HMD systems with the conventional C2splitter is given in detail in Table 4. As mentioned above, the main objective of the HMD system is to obtain the desired product at high purity and low cost. For the series configuration, the purity of ethylene is 0.9979 and the cost savings is 61.45%. The membrane module cut rates are of 0.93 and 0.01 for M1 and M2, respectively. For the membrane cascade, the purity of ethylene is enhanced to 0.9999 with membrane module cut rates of 0.75 and 0.90 for M1 and M2, respectively. Even though the total cost savings favours the series configuration, the higher purity of ethylene is obtained with the membrane cascade, showing a cost saving of 54.41%. Thus, for a grass root design, the membrane cascade configuration can be considered as the best alternative configuration to the conventional C2-splitter. On the other hand, it cannot be recommended for a retrofit of an existing plant since its integration will be costly. One of the disadvantages of series and membrane cascade configurations is, however, that more compression duty is required for ethylene product. In comparison to a conventional C2-splitter, the utility cost is 2 times and 4 times higher for membrane cascade and the series configurations respectively. For propane/propylene separation, different HMD configurations were compared with two columns in series. For a given HMD configuration, the effect of the number of stages on the product purity and the capital cost was analyzed. High propylene purity was obtained for the simulation of a column with 90 and 100 stages. The difference in the propylene purity between 90 and 100 is negligible, and 90 was selected as the optimum number of stages for all HMD systems. The top configuration results in a propylene purity of 0.9996 and a saving of 29.84% (Table 5). For the top-bottom configuration, propylene purity and saving were found to be 0.9985 and 36.57%, respectively. For bottom and parallel configurations, the propylene purity is only 0.924. The top-bottom HMD system

Table 4 Ethane/ethylene separation: cost comparison. Ethylene purity

Membrane area (m2 )

Separation factor

Capital cost (US$) ×106

Operating cost (US$) ×106

Utility cost (US$) ×106

Series Membrane cascade Conventional C2-splitter

0.9979 0.9999 0.9950

M1: 8150 M2: 890 M1: 8150 M2: 3243 –

M1: 1.92 M2: 99.92 M1: 998.98M2: 1.80 –

7.88 10.48 24.4

0.87 1.15 2.7

1.86 0.92 0.5

Total savings (%) 61.45% 54.41% –

Table 5 Propane/propylene separation: cost comparison for retrofit purpose. Configuration

Diameter (m) Top Bottom Top-bottom Parallel Conventional C3-splitter

5.30 5.30 5.30 5.30 5.30

Capital cost (US$) ×106

Operating cost (US$) ×106

Utility cost (US$) ×106

182 4.92 790 – Top: 346Bottom: 527 19.70

26.01 25.98 23.11

2.86 2.86 2.54

3.41 3.35 3.53

29.84 30.03 36.57

0.9239 0.9948

2086 3652

4.93 –

26.32 32.58

2.89 3.58

3.52 4.79

28.84 10.98

0.9815

–

–

37.8

4.2

4.00

Propylene purity

Membrane area (m2 )

90 90 90

0.9996 0.9245 0.9985

90 152 205

Distillation column

Separation Factor

Total investment cost savings (%)

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Configuration

Number of trays

–

389

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is the optimum hybrid system for propane/propylene separation to obtain propylene at high purity with a remarkable cost saving at optimal surface areas of 346 and 527 m2 for top and bottom membranes, respectively. In case of designing a new plant integrating top-bottom HMD system, the internal diameter of the 90 stages distillation column can be reduced from 5.30 to 3.94 combined with two membranes having the same surface areas. Based on these specifications, the estimated savings are 53.64%, which makes such HMD very competitive than two-stage distillation system for C3 H6 /C3 H8 separation/purification. As seen in Table 5 parallel configuration is the optimum design in terms of the total investment cost saving for propane/propylene separation. However, the product purity is low for this suggested configuration. For higher propylene purity (≥0.99), a hybrid system with a distillation column of more than 90 trays, should be used. Similar configuration was presented by Kookos [25] in which a distillation column having a diameter of 3.30 m and 250 trays including the condenser and reboiler combined with a membrane of 2245 m2 was suggested to obtain a propylene purity of 0.99. A 17.1% of total cost savings were obtained with the hybrid system in comparison to a single column of 248 trays including the condenser and reboiler and 3.86 m in diameter [25]. The major savings were resulted from the cut in steam cost. In the present work a comparison was done between a hybrid system with a distillation column of 90 trays and a FT membrane of 2086 m2 , and a hybrid system with a column of 152 trays for parallel configuration with a FT membrane of 3652 m2 (Table 5). A propylene purity of 0.9948 was obtained by using a parallel hybrid system with a column of 152 trays yielding total cost savings of 10.98% in comparison to conventional C3-splitter composed of two columns having 115 and 90 stages. The product purity and total cost saving for top-bottom configuration is close to the parallel configuration involving a distillation column of 152 trays. However, the utility cost is increased by 26.3% in comparison to the top-bottom configuration. 4. Concluding remarks For high product purity and incremented both energy and economic savings, the HMD system is proposed as an alternative design to replace current distillation technology. FT membrane technology is clearly a favourable option for the intensification of C2 H4 /C2 H6 and C3 H6 /C3 H8 separation/purification. Different hybrid configurations were compared with conventional C2- and C3-splitters for ethane/ethylene and propane/propylene separation, respectively. To obtain high product purity at relatively low cost, the membrane cascade system is proposed as the optimum configuration for a new ethane/ethylene separation plant. This is achieved at increased membrane requirements corresponding to higher the utility cost by 45.6% in comparison with conventional C2-splitter. However, the separation of C2 H4 from a multicomponent mixture composed of CH4 (1.80%), C2 H4 (75.66%), C2 H6 (12.61%), C3 H8 (1.34%), and C3 H6 (8.58%), can be efficiently achieved in the membrane cascade scheme providing ultra-high purity (99.99%) as well as high yield of C2 H4 . Such two-stage membrane cascade integrated with three compressors can provide a competitive solution than the multistage distillation of such a multicomponent mixture. HMD series configuration is found to be suitable for the retrofit of an ethylene plant. The top-bottom HMD system is proposed as the preferred

design for propane/propylene separation from propylene purity and yield point of view even though the total savings are slightly favourable for parallel HMD system Acknowledgements The authors acknowledge the financial support provided by the Program of Energy Research and Development of Natural Resources Canada. The authors note their gratitude to Chemstations for the discussions with their technical support. References [1] G.E. Keller, R.R. Beebe, R.J. Fruehan, N.N. Li, E.L. Menger, G.P. Pez, P.H. Pfromm, R.W. Rousseau, M.P. Thomas, Separation Technologies for the Industries of the Future, National Academy Press, USA, 1999, p. 128. [2] T.K. Ghosh, H.D. Lin, A.L. Hines, Hybrid adsorption–distillation process for separating propane and propylene, Ind. Eng. Chem. Res. 32 (1993) 2390–2399. [3] J.P. Schmal, H.J. van der Kooi, A. de Rijke, Z. Olujic, P.J. Jansens, Internal versus external heat integration: operational and economic analysis, Chem. Eng. Res. Des. 84 (2006) 374–380. [4] S. Moganti, R.D. Noble, C.A. Koval, Analysis of a membrane/distillation column hybrid process, J. Membr. Sci. 93 (1994) 31–34. [5] T. Pettersen, A. Argo, R.D. Noble, C.A. Koval, Design of combined membrane and distillation processes, Sep. Technol. 6 (1996) 175–187. [6] I.K. Kookos, Optimal design of membrane/distillation column hybrid processes, Ind. Eng. Chem. Res. 42 (2003) 1731–1733. [7] T.G. Pressly, K.M. Ng, A break-even analysis of distillation–membrane hybrids, AIChE J. 44 (2004) 93–105. [8] D.E. Suk, T. Matsuura, Membrane-based hybrid processes: a review, Sep. Sci. Technol. 41 (2006) 595–626. [9] R.J. Valus, R. Eshraghi, A.E. Velikoff, J.C. Davis, High pressure facilitated membranes for selective separation and process for the use thereof, US Patent 5,057,641 (1991). [10] I. Pinnau, L.G. Toy, C. Casillas, Olefin separation membrane and process, US Patent 5,670,051 (1997). [11] J.H. Kim, Y.S. Kang, J. Wong, Silver polymer electrolyte membranes for facilitated olefin transport: carrier properties, transport mechanism and separation performance, Macromol. Res. 12 (2004) 145–155. [12] S.W. Kang, J.H. Kim, K.S. Oh, J. Won, K. Char, H.S. Kim, Y.S. Kang, Highly stabilized silver polymer electrolytes and their application to facilitated olefin transport membranes, J. Membr. Sci. 236 (2004) 163–169. [13] D.E. Gottschlich, D.L. Roberts, Energy minimization of separation processes using conventional/membrane hybrid systems, in: DOE Final Report, 1990, p. 157. [14] L.G. Wua, J.N. Shenb, H.L. Chenb, C.J. Gaoa, CO2 facilitated transport through an acrylamide and maleic anhydride copolymer membrane, Desalination 193 (2006) 313–320. [15] R. Agrawal, Membrane cascade schemes for multicomponent gas separation, Ind. Eng. Chem. Res. 35 (1996) 3607–3617. [16] K.N. Trueblood, H.J. Lucas, Coordination of silver ion with unsaturated compounds. V. Ethylene and propene, J. Am. Chem. Soc. 74 (1952) 1338–1339. [17] M. Teramoto, N. Takeuchi, T. Maki, H. Matsuyama, Ethylene/ethane separation by facilitated transport membrane accompanied by permeation of aqueous silver nitrate solution, Sep. Purif. Technol. 28 (2002) 117–124. [18] M. Teramoto, S. Shimizu, H. Matsuyama, N. Matsumiya, Ethylene/ethane separation and concentration by hollow fiber facilitated transport membrane module with permeation of silver nitrate solution, Sep. Purif. Technol. 44 (2005) 19–29. [19] W.G. Whitman, The two-film theory of gas-absorption, Chem. Met. Eng. 29 (1923) 146–148. [20] M. Terramoto, H. Matsuyama, T. Yamashiro, Y. Katayama, Separation of ethylene from ethane by supported liquid membranes containing silver nitrate as a carrier, J. Chem. Eng. Jpn. 19 (1986) 419–424. [21] R.D. Hughes, J.A. Mahoney, E.F. Steigelmann, Olefin separation by facilitated transport membranes, in: N.N. Li, J.M. Calo (Eds.), Recent Developments in Separation Science, vol. IX, CRC Press, Cleveland, USA, 1982. [22] CHEMCAD V6, Chemstations Inc., Houston, USA, 2008. [23] J.M. Douglas, Conceptual Design of Chemical Processes, McGraw-Hill, USA, 1988, p. 601. [24] ICIS, Chemical Prices and Chemical Industry Trends, 2008, www.icispricing. com. [25] I.K. Kookos, Optimal design of membrane/distillation column hybrid processes, Ind. Eng. Chem. Res. 42 (2003) 1731–1738.