A comparative study of two different membranes applied for auto-thermal methanol synthesis process

Journal of Natural Gas Science and Engineering 7 (2012) 60e74

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Journal of Natural Gas Science and Engineering journal homepage: www.elsevier.com/locate/jngse

A comparative study of two different membranes applied for auto-thermal methanol synthesis process Farhad Rahmani a, b, Mohammad Haghighi a, b, *, Pooya Estifaee a, b, Mohammad Reza Rahimpour c a

Chemical Engineering Faculty, Sahand University of Technology, P.O. Box 51335-1996, Sahand New Town, Tabriz, Iran Reactor and Catalysis Research Center (RCRC), Sahand University of Technology, P.O. Box 51335-1996, Sahand New Town, Tabriz, Iran c Chemical Engineering Department, School of Chemical and Petroleum Engineering, Shiraz University, Shiraz, Iran b

a r t i c l e i n f o

a b s t r a c t

Article history: Received 16 September 2011 Received in revised form 30 March 2012 Accepted 5 April 2012 Available online 10 May 2012

The growing demand for energy and environmental concerns make finding economical and environmentally sustainable solutions for the effective and more utilization of natural gas, as the cleanest fossil fuel, imperative. Among different approaches to reach this goal, chemical conversion is of great interest. Future energy carriers such as hydrogen and methanol can be produced by the natural gas conversion. Recently, the auto-thermal methanol synthesis (AMS) process has become an important alternative for natural gas conversion and monetization. Concerning the equilibrium limitation of methanol synthesis reactions, there are two different configurations for membrane reactor, which may be used in this process. In the first configuration (in situ H2 addition configuration), the Pd/Ag membrane was applied while in the other one (in situ H2O removal configuration) H-SOD membrane was used. Generally, the proposed reactor is composed of three concentric tubes of which the inner tube is separated by a membrane from second one (exothermic side). A steady-state heterogeneous model was developed to investigate the possibility of improving AMS performance by means of two different membranes. The proposed model has been used to compare the performance of two different auto-thermal configurations with respect to the non-membrane one under identical process conditions. It was found that the reactor in the in situ water removal configuration operates with higher methanol yield, higher carbon dioxide removal which causes a lower environmental impact and longer exothermic catalyst life as a result of the more favourable temperature profile as well as reducing H2O promoted catalyst deactivation. However, thermal efficiency of reactions is declined and the dehydrogenation reaction yield in both configurations is not significantly different. Finally, the influence of inlet temperature of sweep gas as one of the key operating variables is investigated on products yield. The results suggest that utilization of this reactor could be feasible and beneficial. Ó 2012 Elsevier B.V. All rights reserved.

Keywords: Auto-thermal membrane methanol synthesis reactor Hydrogen production In situ hydrogen addition configuration In situ water removal configuration Steady-state heterogeneous model

1. Introduction Energy is the very lifeblood of today’s society and economy. However, most of the employed energy, nowadays, comes from fossil fuels as a non-renewable energy source. Concerns about the environmental misdeeds of fossil fuels and the growing gap between increasing demand and shrinking supply of fossil fuels, respectively, have led to significant research into the use of alternative energy carriers. Hydrogen and methanol have been identified as ideal energy carriers to support sustainable energy * Corresponding author. Chemical Engineering Faculty, Sahand University of Technology, P.O. Box 51335-1996, Sahand New Town, Tabriz, Iran. Tel.: þ98 412 3458097/þ98 412 3459152; fax: þ98 412 3444355. E-mail address: [email protected] (M. Haghighi). URL: http://rcrc.sut.ac.ir 1875-5100/$ e see front matter Ó 2012 Elsevier B.V. All rights reserved. doi:10.1016/j.jngse.2012.04.001

development. Widespread use of them, if generated in an advantageous manner, could contribute to alleviation of the growing concerns about the world’s energy supply, security, air pollution and greenhouse gas emissions (Ye et al., 2009). 1.1. Natural gas Natural gas (NG) is one of the major fossil energy sources and an important ingredient in the global energy pool (Economides and Wood, 2009), not the least for power production. NG is also a dominating feedstock for the production of several bulk chemicals such as ammonia, methanol, dimethyl ether (DME) and hydrogen. Its usage fails to provide a solution to deal with huge amount of carbon dioxide emissions during the heating, electricity production, transportation and industrial purposes. For example, NG accounts for 20% of hydrocarbon-related CO2 emissions (Abbas

F. Rahmani et al. / Journal of Natural Gas Science and Engineering 7 (2012) 60e74

and Wan Daud, 2010). On the other hand, a considerable amount of stranded gas reserves are located off shore with low economic incentive to be produced, transported and sold (Bakhtiary-Davijany et al., 2011). In these circumstances, the natural gas that often comes out of the earth along with crude oil is burned off in flares. Flaring is a conventional method of disposing of purge gas in refineries. This method has been the source of much controversial debate as not only does it waste a considerable amount of valuable energy but also results in severe environmental issues in the petroleum and related industries (Tolulope, 2004). The growing demand in the fuels and chemicals market on one hand and increasing environmental concerns and toughening regulations on the other hand make finding economical and environmentally sustainable solutions for effective monetization of NG as well as utilization of the more number of the proven reserves imperative. Among different approaches, chemical conversion is of great interest. Conversion increases the value of the raw material 3e10 times. Future energy carriers such as hydrogen and methanol can be produced by the conversion of natural gas (Løvik, 2011).

61

Natural Gas

Feedstock Purification

Synthesis gas prepration

Steam Reforming

CO2 Reforming

Partial Oxidation

Auto-thermal Reforming

1.2. Natural gas conversion via methanol synthesis process

Methanol Synthesis As a consequence of large natural gas reserves, and recent advances in CH4 chemistry and technology, the future growth of natural gas usage is widely expected: as gaseous fuel, in the production of liquids fuels and/or in manufacturing of petrochemicals (Taniewski, 2008). The conversion of natural gas to methanol is not only a realizable and sustainable alternative for stranded natural gas monetization, but can also be considered as a strategy for medicating environmental problems of NG usage. 1.2.1. Methanol Methanol is considered as an alternative energy source, a medium for the storage and transportation of hydrogen and a starting feedstock for many chemicals as well. It is a clean-burning fuel with versatile applications. Methanol can also be used as a solvent, cleaner or a fuel additive and especially as a building block to produce chemical intermediates such as dimethyl ether (DME) and methyl t-butyl ether (Semelsberger et al., 2006). Commercially, it is produced by the natural gas and specifically by means of syngas (CO, CO2 and H2 mixture) obtained via steam reforming operations (Chinchen et al., 1988). Conversion of remote natural gas to methanol even by conventional methanol technology was reported to be economically competitive compared with shipping LNG. Today, one of the most promising processes for the utilization and fixation of CO2 is the methanol synthesis. 1.2.2. Methanol synthesis process The methanol process consists of four main parts; Feed treatment (purification), synthesis gas preparation, methanol synthesis and product purification and storage. Fig. 1 shows the overall methanol synthesis process. The natural gas is first desulphurized and water impurities are removed. Then the resulting synthesis gas is reformed. The synthesis gas is a mixture of CO, CO2 and H2 which can be produced form different feed stocks: natural gas, higher hydrocarbons and coal. The conventional process for synthesis gas production is steam reforming, but partial oxidation, CO2 reforming, auto-thermal reforming or a combination of them is also used. The main step of methanol process is methanol synthesis. The produced synthesis gas enters to the reactor. Methanol synthesis reactors are designed based on two technologies, high-pressure synthesis operating at 300 bar and using a Cr-based catalyst; and low-pressure synthesis operating between 50 and 100 bar and using a Cu-based catalyst. The methanol synthesis reactor studied

Low-pressure Process High-pressure Process

Product Purification and Storage

Methanol Fig. 1. Overall process for production of methanol.

here is a Lurgi type, which is operated in the low-pressure regime (Kordabadi and Jahanmiri, 2005). Unreacted synthesis gas is separated from crude methanol, then compressed and recycled. Finally, the crude methanol is purified by distillation. 1.3. Hydrogen Hydrogen offers the best energy-to-weight ratio of any fuel. It is not only an ideal energy carrier for the future mainly due to its high conversion efficiency, recyclability and non-polluting nature, but also a fundamental raw material and feedstock in petroleum, chemical engineering, chemical fertilizer and metallurgical industries, for example in the production of ammonia and methanol, upgrading of heavy hydrocarbons, iron ore reduction and food processing (Adris et al., 1997). Meanwhile, hydrogen is often referred to as ‘clean energy’ since its combustion produces only water. However, the greatest difficulties in expanding the use of hydrogen energy are safe storage and transport, because hydrogen is a gas under ordinary temperature and pressure and its liquefaction temperature is around 20 K. Currently, the most favoured feedstock for hydrogen production is natural gas due to its availability and advantageous price. The main drawback of natural gas

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reforming is that this reaction leads to a H2 rich gas mixture containing carbon oxides and other by-products. Therefore, one of the environmentally benign methods to produce hydrogen without CO2-emission is cyclohexane dehydrogenation that provides a good possibility for effective hydrogen storage and supply (Biniwale et al., 2005). 1.4. Auto-thermal methanol synthesis (AMS) process In view of thermodynamic limitation and the high endothermicity of cyclohexane dehydrogenation, chemical equilibrium is favourable for dehydrogenation at high temperature and thereby a large amount of external heat source is required. On the opposite side, methanol synthesis process is exothermic and so the related reactions are thermodynamically favoured at low temperatures. Currently, some alternative processes such as auto-thermal methanol synthesis (AMS) are considered to supply the necessary heat of cyclohexane dehydrogenation reaction through methanol synthesis process and simultaneous produce hydrogen and methanol as environment friendly fuel (Khademi et al., 2010, 2009a,b). Indeed, AMS combines methanol synthesis and cyclohexane dehydrogenation reactions in a single reactor in order to simultaneous hydrogen and methanol production. Auto-thermal methanol synthesis reactor offers the advantage to produce not only hydrogen, but also co-products such as methanol and benzene, used for other scopes, as well as more compact design that can improve the economy of the overall process. Benzene is an important industrial solvent and precursor in the production of drugs, plastics, synthetic rubber, and dyes Besides, possibility of achieving higher degree of in-sit energy integration between the coupled endothermic dehydrogenation reaction and the exothermic methanol synthesis reaction, saving energy and improving the reactions thermal efficiency; and establishment of auto-thermal conditions in both reaction sides are the other clear advantages of this integrated catalytic reactor. However, the auto-thermal processes are not flexible in terms of choosing fuel and their operating variables. 1.5. Membrane assisted to AMS process The AMS process is affected by thermodynamic constraints, which limit reactants conversion. Under these circumstances, membranes may be introduced to auto-thermal methanol synthesis reactors to selectively in situ product removal or reactant addition. This favourably shifts the equilibrium forward by Le ChateliereBrown principle. Separation by membrane is less energy intensive, requiring no phase change in the process and typically provides lowmaintenance operation. Membrane reactors are among the most powerful tools for achieving the targets of process intensification, i.e. to obtain cheaper, safer and more sustainable technologies. 1.5.1. In situ H2 addition Regarding to thermodynamic equilibrium limitations and H2poor synthesis gas produced from the steam reforming reaction, one of the promising candidates in the auto-thermal methanol synthesis reactor (AMSR) is in situ H2 addition via PdeAg membrane, which has also shown increasing importance in membrane reactors in hydrogen production and methanol synthesis processes. The highest hydrogen permeability was observed at an alloy composition of 23 wt% silver (Rahimpour and Ghader, 2003). PdeAg membranes are 100% selective for hydrogen transport (Buxbaum and Kinney, 1996) and utilizing them leads to maximizing the shift effect and producing directly ultra pure hydrogen without using any other purification unit.

Selective H2 addition is a strategy to supply hydrogen on the catalytic methanol synthesis zone in a controlled manner. In this case, the hydrogen feeding through the membrane has been provided by recycling of synthesis gas effluent. However, one of the main problems for the application of this kind of inorganic membrane reactor is the high cost of membranes. Moreover, the recycling and compression of synthesis gas result in huge energy consumption. 1.5.2. In situ H2O removal H2O which is the main by-product of the methanol synthesis reaction accumulates in the gas phase and decreases the partial pressures of the reactants. Besides, water produced during methanol synthesis accelerates the crystallization of Cu and ZnO contained in a Cu/ZnO-based catalyst to lead to the deactivation of the catalyst (Rahimpour et al., 1998; Wu et al., 2001). Therefore, the selective removal of the by-product H2O which cause deactivation of catalyst and may inhibit the reaction rate can increase conversion product yield and catalyst life time. In this study, a new type of membrane on basis of hydroxysodalite (H-SOD) was taken into consideration. The membrane is swept with a sweep gas at low pressure to maintain a high driving force across the membrane. Therefore, no need for an additional compressor is one of the potential benefits for such a configuration. The H-SOD membrane is highly permeable towards H2O and retains H2, CO, CO2 and hydrocarbons almost completely to avoid a costly recovery of the reactants or products from the sweep gas stream. It shows excellently an ideal selectivity (near to absolute) of water to hydrogen. In addition, the membrane demonstrates a high thermal and mechanical stability. The exceptional property of HSOD membrane in comparison with the other widely studied membranes (David et al., 1991; Rohde et al., 2008) is 100% selectivity of water on the basis of molecular sieving in hybrid processes (Khajavi et al., 2009). 1.6. Literature review Up to now, most of the efforts in this field have been focused on the integration of Pd-Ag membranes into hydrocarbon reformers or water-gas shift (WGS) reactors or on the auto-thermal reforming (ATR) carbon compounds (oxygen-containing compounds or hydrocarbons) (Gallucci et al., 2008, 2004, 2007; Gallucci and Basile, 2008; Amandusson et al., 2001; Basile et al., 2006, 2005a,b; Mendes et al., 2010; Rabenstein and Hacker, 2008; Dauenhauer et al., 2006; Wang et al., 2009), but little attention has been devoted to the recuperative auto-thermal membrane methanol synthesis process. Recently, Khademi et al. (Khademi et al., 2009) optimized the operating conditions of a recuperative coupled reactor containing methanol synthesis in the exothermic side and cyclohexane dehydrogenation in the endothermic side. Afterwards, the mentioned reactor has been modified by utilizing a Pd/Ag membrane layer to separate the produced hydrogen in the endothermic side and finally optimal operating conditions have been found (Khademi et al., 2010, 2009). According to this simple literature survey, no paper dealing with feasibility of two different membranes incorporation to methanol synthesis side of AMSR is presented in the literature, so investigation and comparison of auto-thermal membrane methanol synthesis reactors (AMMSRs) performance can be an interesting and novel idea which is the subject of this work. 1.7. Objectives In this work, we decided to study the effect of selective water removal or hydrogen redistribution along an auto-thermal

F. Rahmani et al. / Journal of Natural Gas Science and Engineering 7 (2012) 60e74

63

Fig. 2. Schematic diagram of an in situ H2O removal configuration (IWRC).

membrane methanol synthesis reactor on the simultaneous hydrogen and methanol production. Depending on the chosen strategy (in situ H2 addition and water removal), two types of membranes was considered. In the first strategy (in situ H2 addition), Pd/Ag membrane which is H2 permeable, was applied while in the other strategy (in situ water removal) water permeable H-SOD membrane was integrated to methanol synthesis side of thermally coupled reactor. The motivation is to combine the energy efficient concept of coupling endothermic-exothermic reactions, the membrane-assisted methanol synthesis equilibrium displacement; and enhancement of hydrogen, benzene and methanol production in a single reactor, simultaneously. The steady state, 1-D mathematical model of the thermally coupled membrane multi-tubular reactor is presented to evaluate and compare the performance of both proposed reactor configurations with the non-membrane configuration at the same process conditions such as pressure, temperature, catalyst mass and feed composition. 2. Process description The auto-thermal membrane methanol synthesis reactor (AMMSR) that is developed for simultaneous production of methanol and hydrogen is composed of three concentric tube reactors of which the non-reaction side (inner tube) is separated by a membrane from the adjacent reaction side (exothermic side). Methanol synthesis in the exothermic side supplies the necessary heat for the endothermic dehydrogenation of cyclohexane reaction which assumed to take place in the outer tube. According to utilizing water or hydrogen permeable membrane in the exothermic side, there are two different configurations for the AMMSR. 2.1. In situ H2O removal configuration (IWRC) Fig. 2 shows the schematic diagram of in situ water removal configuration of AMMSR. The wall between first and second tubes is a water selective H-SOD membrane allowing produced water to permeate out of exothermic side and selective removal of water is achieved by co-current flow of sweep gas through the permeation side (inner tube). The characteristics and input data of thermally coupled membrane reactor are listed in Tables 1and 2 (Jeong et al., 2004; Koukou et al., 1997). The operating conditions for exothermic

side were extracted from Rahimpour’s studies (Rahimpour et al., 2005; Rezaie et al., 2005). 2.2. In situ H2 addition configuration (IHAC) The in situ hydrogen addition configuration of the auto-thermal membrane methanol synthesis reactor is shown schematically in Fig. 3. Generally, the process in this configuration is similar to the in situ H2O removal configuration except a few changes. These changes in the new proposed system are as follows: Firstly, the H2 permeable Pd/Ag membrane in exothermic side has been used. Secondly, to supply and add hydrogen into reaction side, the reacting gas leaving exothermic side is recycled, compressed and passed through the inner tube (non-reaction side) in a co-current mode with reacting gas. Hydrogen partial pressure in recycle stream (after it was compressed) is suitable to permeate to exothermic side. After leaving the inner tube (recycling synthesis gas side), the methanol containing gas (product) goes to the separator. The characteristics of IHAC are the same as IWRC, except the inlet pressure of non- reaction side which is 96.98 bar. Table 1 The operating conditions for methanol synthesis process (exothermic side) in the in situ H2O removal configuration. Parameter

Value

Exothermic side Feed composition (mole fraction) CH3OH CO2 CO H2O H2 N2 CH4 Total molar flow rate (mol s1) Length of reactor (m) Inlet pressure (bar) Inlet temperature (K) Density (kg m3) Particle diameter (m) Heat capacity (kJ kg1 K1) Specific surface area (m2 m3) Bed void fraction Density of catalyst bed (kg m3) Wall thermal conductivity (W m1 K1)

0.0050 0.0940 0.0460 0.0004 0.6590 0.0930 0.1026 0.64 7.022 76.98 503 1770 5.47  103 5.0 626.98 0.39 1140 48

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Table 2 The operating conditions for dehydrogenation of cyclohexane to benzene (endothermic side) and permeation side in the in situ H2O removal configuration. Parameter

Value

Endothermic side Feed compositiona (mole fraction) C6H12 C6H6 H2 Ar Total molar flow rate (mol s1) Inlet pressurea (Pa) Inlet temperature (K) Tube diameter (m) Particle diameterb (m) Bed void fraction

0.11 0.0 0.0 0.9 0.1 1.013  105 503 8  102 3.55  103 0.39

Permeation side Feed composition (mole fraction) Ar (sweep gas) H2 Total molar flow rate (mol s1) Inner tube diameter (m) Inlet temperature (K) Inlet pressure (Pa) Thermal conductivity of membrane (Wm1 K1) a b

1.0 0.0 0.64 5.3  102 513 0.1  105 153.95

Obtained from Jeong et al. (2004). Obtained from Koukou et al. (1997).

3.2. Dehydrogenation of cyclohexane The reaction scheme for the dehydrogenation of cyclohexane to benzene is as follows:

C6 H12 4C6 H6 þ 3H2 DH298 ¼ þ206:2 kJ=mol

(1)

The following reaction rate equation of cyclohexane, rc, is used (Itoh, 1987):


3 P k KP PC =PH B 2
 rc ¼ 3 1 þ KB KP PC =PH 2

(2)

where k, KB and KP are the reaction rate constant, the adsorption equilibrium constant for benzene and the reaction equilibrium constant respectively that are tabulated in Table 3. Pi is the partial pressure of component i in Pa. The reaction temperature is in the range of 423e523 K and the total pressure in the reactor is maintained at 101.3 kPa. The catalyst is Pt/Al2O3 (Jeong et al., 2003). 4. Mathematical modelling The following assumptions are considered during the modelling of AMMSR in both configurations as well as AMSR (non-membrane configuration):

3. Reaction scheme and kinetics 3.1. Methanol synthesis In the methanol synthesis, three overall reactions are possible: hydrogenation of carbon monoxide, hydrogenation of carbon dioxide and reverse wateregas shift reaction. In the current work, the rate expressions have been selected from Graaf et al. (1990). The rate equations combined with the equilibrium rate constants (Graaf et al., 1986) provides enough information about kinetics of methanol synthesis over commercial CuO/ZnO/Al2O3 catalysts.

 One-dimensional heterogeneous model (reactions take place in the catalyst particles).  Steady state conditions.  Plug flow pattern is considered in each side.  Axial diffusion of heat and mass are neglected compared with the convection.  No radial heat and mass diffusion in catalyst pellet.  Bed porosity in axial and radial directions is constant.  Gas mixtures considered to be ideal.  Heat loss to surrounding is neglected.

Fig. 3. Schematic diagram of an in situ H2 addition configuration (IHAC).

F. Rahmani et al. / Journal of Natural Gas Science and Engineering 7 (2012) 60e74 Table 3 The reaction rate constant, adsorption equilibrium constant, reaction equilibrium constant for dehydrogenation of cyclohexane. k¼Aexp(B/T)

A

B

k KB KP

0.221 2.03  1010 4.89  1035

4270 6270 3190

65

second term is used for the in situ water removal configuration and f is equal to 1 for water component and 0 for other components. Moreover, the negative sign of second term in these equations is used for the in situ hydrogen addition configuration and f is equal to 1 for hydrogen component and 0 for CO2, CO, H2, CH3OH and inert components. 4.3. Hydrogen permeation in Pd/Ag membrane

4.1. Reaction sides

(3)

The composite membrane in the in situ H2 addition configuration is made of a layer of palladiumesilver alloy. The membrane is deposited as a continuous layer on the outer surface of a thermo stable support. The flux of hydrogen permeating through the Pd/Ag membrane is assumed to follow the halfepower pressure law, Richardson’s law (Richardson, 1904), derived from a combination of Fick’s and Sieverts’ laws and expressed by (Basile et al., 2008; Sammells and Mundschau, 2006):

(4)

Jm ¼ JH2;1 ¼

4.1.1. Solid phase The mass and energy balances for the solid phases of reaction sides of three thermally coupled reactors are analogous and expressed by:


 g av cj kgi;j yi;j  ysi;j þ hri;j rb ¼ 0

av hf




g Tj



Tjs



þ rb

N X






hri;j DHf ;i ¼ 0

i1

where ysi;j and Tjs are the solid-phase mole fraction and temperature in jth side of reactor respectively and h is effectiveness factor of kth reaction in jth side (effectiveness factor is the ratio of the reaction rate observed to the real rate of reaction), which is obtained from a dusty gas model calculations (Rezaie et al., 2005). 4.1.2. Fluid phase The mass and energy balance equations for the fluid phase of reaction sides can be formulate as:


 Fj dyi;j Jm  ¼ 0 þ av cj kgi;j ysi;j  ygi;j þ f Ac;j dz Ac;j g

(5)


 Fj g dTj pDi  g g g U T3  T2   Cpj þ av hf Tjs  Tj Ac;j dz Ac;j g

þf

pDi Ac;j



g U12 T1



g T2

jm þf Ac;j

(6)

ZT2

PH2 is hydrogen partial pressure in Pa. Do and Di stand for the outer and inner diameters of the Pd/Ag layer. The pre-exponential factor P0 above 200  C is reported as 6.33  108 molm2s1Pa0.5 and the activation energy Ep is 15.7 kJmol1 (Rahimpour et al., 2010). 4.4. Water permeation through H-SOD membrane The flux of water permeation through the H-SOD membrane, JH2 O , depends on the water partial pressure difference between two sides of the membrane.

  Jm ¼ JH2 O ¼ Am QH2 O PH2 O;2  PH2 O;1

(10)

where QH2 O is water permeation and it is a specific property of the membrane. QH2 O is constant respect to temperature and the H2O permeance is in a range of 107 and 106 molm2s1Pa1 (Rohde et al., 2008). 4.5. Pressure drop The Ergun momentum balance equation is used to give the pressure drop along the reactor:

! ð1  εÞu2g r ð1  εÞ2 mug dP þ 1:75 ¼  150 dz ε3 d2p ε3 dp

The mass and energy balances for the recycling gas side of IHAC and permeation side of IWRC are expressed by:

where the pressure drop is in Pa.

4.6.1. In situ H2O removal configuration The following boundary conditions are applied:

z ¼ 0;

g

 fJm ¼ 0

(11)

4.6. Boundary conditions

4.2. Non-reaction side

F1

(9)

Cp dT ¼ 0 T1

where Jm is hydrogen permeation rate for Pd/Ag membrane and water permeation rate for H-SOD membrane. In Eq. (6), the positive sign of third term is used for the exothermic side and the negative sign for the endothermic section. In this equation, the fourth and last terms are related to the heat transferred as result of membrane utilization. Therefore, in Eqs. (5) and (6), f is equal to 0 for the nonmembrane configuration while f is equal to 0 for the endothermic and 1 for the exothermic side of membrane configurations.

dyi;1

  Ep
qffiffiffiffiffiffiffiffiffiffiffi qffiffiffiffiffiffiffiffiffiffiffi 2pLP 0  exp PH2 ;1  PH2 ;2 Do RT ln Di

(7)

ygi;j ¼ ygi0;j ;

Tjg ¼ T0g ;

Pjg ¼ P0g

(12)

(8)

where ygi0;j , T0g and P0g are the fluid-phase mole fraction, temperature and pressure at the entrance of jth side of reactor, respectively. The above boundary conditions are also valid for the nonmembrane configuration.

where ygi;1 and T1g are the mole fraction and temperature in the nonreaction side, respectively. In Eqs. (7) and (8), the positive sign of

4.6.2. In situ H2 addition configuration The following boundary conditions for reaction sides of AMMSR are applied:

dz

F1 Cpg1

dT1g  fJm dz

ZT2

  Cp dT þ pDi U12 T2g  T1g ¼ 0

T1

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ygi;j ¼ ygi0;j ;

z ¼ 0; g

g

Tjg ¼ T0g ;

Pjg ¼ P0g

j ¼ 2:3

(13)

g

where yi0;j , T0 and P0 are the fluid-phase mole fraction of ith component, temperature and pressure at the entrance of jth side of reactor, respectively and the boundary conditions for recycling synthesis gas side are as follow:

z ¼ 0;

g

g

g

yi;1 ¼ yif ;2 ; g yif ;2 ,

g Tf ;2

g

T1 ¼ Tf ;2 ;

g

g

P1 ¼ Pf ;2

(14)

g Pf ;2

where and are the fluid-phase mole fraction of ith component, temperature and pressure at the end of exothermic side, respectively.

section. It should be noted that The NAEs which constitute a boundary value problem, was solved using the trial-and-error method in reactant addition configuration. In this approach, the calculation started with initial guesses for the inlet synthesis gas temperature (Tin) and hydrogen mole fraction (yin) to the recycling synthesis gas side, which are unknown (initial conditions). The initial conditions were calculated using the GausseNewton method corrected by the previous values of temperature and hydrogen mole fraction of the synthesis gas in exothermic side outlet in subsequent calculations. Substitution was continued until the convergence criterion was met. 5. Results and discussions

4.7. Auxiliary correlations

5.1. Model validation

Auxiliary correlations should be added to solve the set of differential equations. The correlations used for heat and mass transfer between two phases, physical properties of chemical species and overall heat transfer coefficient between two sides are summarized in Table 4 (Cussler, 2003; Lindsay and Bromley, 1950; Reid et al., 1977; Smith, 1980; Wilke, 1949). The heat transfer coefficient between gas phase and reactor wall is applicable for heat transfer between gas phase and solid catalyst phase. 4.8. Numerical solution The governing equations of the model form a system of coupled equations comprising of partial derivative equations of mass and energy conservation rules for the fluid and solid phases; correlations for the heat and mass transfer coefficients and the physical properties of fluids; as well as nonlinear algebraic equations of the kinetic model. After rewriting the model equations, a set of differential algebraic equations (DAEs) is obtained for both configurations. This set of equations is changed to nonlinear algebraic equations (NAEs) using the backward finite difference approximation. The reactor length is then divided into 100 separate sections and the Gauss-Newton method in MATLAB programming environment is used to solve the non-linear algebraic equations in each

The model of methanol synthesis side was validated against conventional methanol synthesis reactor for a special case of constant coolant temperature under the design specifications. The comparison between simulation and plant data for conventional methanol synthesis reactor is shown in Table 5. It was observed that the model performed satisfactorily well under special case of industrial conditions and the observed plant data were in good agreement with simulation data. In this section, various steady-state behaviours observed in the co-current coupled reactor is analysed and the predicted mole fraction, yield, conversion and temperature profiles are presented. The performance of the thermally coupled reactors is analysed, using different operating variables, for methanol yield, cyclohexane conversion and hydrogen yield as follows:

Hydrogen yield ¼

FH2 ;3 FC6 H12 ;in

(15)

Methanol yield ¼

FCH3 OH;out FCO;in þ FCO2 ;in

(16)

Table 4 Physical properties, mass and heat transfer correlations. Parameter

Equation

Component heat capacity Mixture heat capacity Viscosity of reaction mixtures Mixture thermal conductivity Mass transfer coefficient between gas and solid phases

Cp¼aþbTþcT2þdT2 Based on local compositions Based on local compositions ug  103 kgi ¼ 1.17Re0.42Sc0.67 i Re ¼ Sci ¼

m

m rDim  104

qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi 1:43  107 T 3=2 1=Mi þ 1=Mj Dij ¼ pffiffiffi 1=3 1=3 2Pðvci þ vcj Þ2

Heat transfer coefficient between gas phase and reactor wall

Lindsay and Bromley, 1950 Cussler, 2003

2Rp ug

1y Dim ¼ P y i i i ¼ j Dij

Overall heat transfer coefficient

Reference

Reid et al., 1977

Smith, 1980

1 1 Ai lnðDo =Di Þ Ai 1 ¼ þ þ U hi 2pLKw A o ho   0:407  Cp m 2=3 0:458 rudp ¼ m Cp rm K εB h

Wilke, 1949

F. Rahmani et al. / Journal of Natural Gas Science and Engineering 7 (2012) 60e74

67

Table 5 Comparison between simulation and plant data for conventional methanol synthesis reactor. Reactor inlet

Composition (mol %) CO2 CO H2 CH4 N2 H2O CH3OH Feed flow rate (mols1) Temperature (K)

3.45 4.66 79.55 11.72 0.032 0.08 0.032 0.565 503

Methanol selectivity ¼ 

Reactor outlet Exp.

Calc.

Error%

2.18 1.44 75.71 12.98 0.16 1.74 5.49 0.51 528

2.26 1.5 76.37 12.88 0.15 1.66 5.23 0.5 524.3

3.67 4.167 0.87 0.77 6.66 4.598 4.736 1.96 0.7

FCH3 OH;out    FCO;in þ FCO2 ;in  FCO;out þ FCO2 ;out

Cyclohexane conversion ¼

FC6 H12;in  FC6 H12;out FC6 H12;in (17)

Furthermore, an obvious measure for the performance of the reactor concept is how much heat has to be supplied through the exothermic reaction to maintain the endothermic reaction. The relative heat supply is defined by the fuel ratio J:

J ¼

Available heat of exothermic reaction Maximum required heat of endothermic reaction

(18)

As efficiency of the reactor, we define:

x ¼

Heat actually consumed for endothermic reaction Heat actually released for exothermic reaction

(19)

Optimal conditions imply J / 1þ and x / 1. 5.2. Exothermic side Fig. 4(a)e(b) demonstrates a comparison of temperature, heat transfer and reaction rates profiles along the exothermic side of the membrane and non-membrane configurations. Along the exothermic side of auto-thermal reactors, temperature increases smoothly and hot spots develop as demonstrated in Fig. 4(a) but then decreases. It is clear that by integrated membrane to exothermic side, the temperature in the first part of the reactor is higher mainly due to the lower heat transfer from the exothermic side, and afterwards it reduces mainly due to the higher heat transfer, as shown in Fig. 4(a) and (b). It should be noted that the differences in term of heat transfer from the exothermic side between the auto-thermal membrane and non-membrane configurations is attributed to the simultaneous heat transfer with gas phase in the inner tube and reacting gas in the endothermic side. Anymore, the exothermic temperature control of the auto-thermal membrane reactors is easier due to lower hot spot. The reactor length in the exothermic side can be divided into two sections. The reaction kinetic and equilibrium are controlled in lower and upper sections, respectively. Therefore, implementing a higher temperature at the reactor entrance for a higher reaction rate, and then reducing the temperature gradually towards the end of the reactor for increasing thermodynamic equilibrium conversion is one of the significant issues in methanol synthesis (see Fig. 4(a)e(c)). However, the high temperature in the initiation steps of reaction is more important due to availability of high reactants concentration.

Fig. 4. Variation of (a) exothermic temperature, (b) transferred heat from exothermic side and (c) exothermic reaction rate for the membrane and non-membrane configurations.

Consequently, the most favourable exothermic temperature profile seems to belong to the in situ H2O removal configuration. Fig. 5(a)e(e) shows the comparison of components mole fraction in exothermic side of thermally coupled membrane reactors with coupled reactor. According to this contrast, auto-thermal membrane systems show favourable results relative to non-membrane system.

F. Rahmani et al. / Journal of Natural Gas Science and Engineering 7 (2012) 60e74

0.06

0.089

0.03

0.081

0.015

0.073

0

0.2

0.4

0.6

0.8

5.3. Endothermic side

0.065 1

Dimensionless length

b 0.025

0.05 IWRC

H2O mole fraction

NC

CO2 mole fraction

CH3OH mole fraction

IHAC

0.045

0

recrystallization and longer catalyst life. Therefore, this membrane reactor can be an appropriate reactor type for auto-thermal methanol synthesis process.

0.097 IWRC

IHAC

NC

0.02

0.045

0.015

0.04

0.01

0.035

0.005

0.03

0

0

0.2

0.4

0.6

0.8

1

CO mole fraction

a

0.025

Dimensionless length

c

0.66 IWRC IHAC NC

0.64

H mole fraction 2

0.62

0.6

0.58

Fig. 6(a)e(b) illustrates the axial temperature and reaction rate profiles along the endothermic sides of auto-thermal configurations. The temperature of the endothermic side is always lower than that of the exothermic side in order to make a driving force for heat transfer from the solid wall. At the entrance of endothermic side of auto-thermal reactors, the temperature decreases rapidly and a cold spot form and then the temperature increases (see Fig. 6(a)). The difference of temperature profiles along the reactors can be explained considering again the comments reported for Fig. 4 or, in other words, the difference in term of heat transfer into the endothermic side in the auto-thermal modes. Accordingly, with thermodynamic restrictions and the high endothermicity of cyclohexane dehydrogenation, high temperature results in a higher reaction rate and consequence higher conversion, as shown in Fig. 6(a)e(c). A comparison is reported in Fig. 7(a)e(b) which shows the cyclohexane and benzene; and hydrogen mole fractions along the reactor in the endothermic side of AMS configurations, respectively. As can be seen, the highest reaction yield is achieved in the non-membrane

a Endothermic side temperature (K)

68

512

510

508

506

504

502

500

0.56 498 0

0.54

0.2

0.4

0.6

0.8

1

Dimensionless length 0.52

0

0.2

0.4

0.6

0.8

1

b

1.3

Dimensionless length

Fig. 5. Comparison of (a) methanol and CO2, (b) H2O and CO; and H2 mole fraction along the reactor axis between exothermic sides of the membrane and non-membrane configurations. 1.1

Membrane concept enhances the methanol yield, which is an objective of the proposed design and likely declines carbon dioxide emissions. Based on favourable profile of gas temperature and better performance of H-SOD membrane, the highest reaction yield is achieved in the in situ H2O removal configuration, as can be seen in Fig. 5(a). Moreover, the main difference between the distributions of consumed hydrogen and produced water along the reactor in the two membrane configurations could be related to the different membranes used in them. Consequently, the in situ H2O removal configuration operates with higher carbon dioxide removal and methanol production as well as lower water production which result in reduction of catalyst

0.9

0.7 0

0.2

0.4

0.6

0.8

1

Dimensionless length Fig. 6. Comparison of (a) temperature, (b) reaction rate along the reactor axis between endothermic sides of three types of configurations.

F. Rahmani et al. / Journal of Natural Gas Science and Engineering 7 (2012) 60e74

69

a

b

Fig. 8. Variation of (a) temperature and (b) permeated component mole fraction along the reactor axis in non-reaction side of the membrane configurations (IWRC and IHAC). Fig. 7. Comparison of (a) C6H12 and C6H6; and (b) H2 mole fractions along the reactor axis between endothermic sides of three types of configurations.

configuration. the negligible difference between the configurations performance is attributed to utilizing membrane between exothermic and non-reaction sides, which reduces the reactions’ thermal efficiency and thereby affects the temperature profile so lower yield of reaction is achieved in the membrane configurations. Finally, it could be concluded that the influence of the integrated membrane to exothermic side on the cyclohexane conversion seems to be poor. 5.4. Non-reaction side Fig. 8(a) and (b) demonstrates the variations of permeated components mole fractions and temperature profiles along the permeation and recycling gas sides of the in situ water removal and hydrogen addition configuration systems, respectively. As it can be seen in Fig. 8(a), the temperature profiles pattern in permeation and recycling gas sides is the same as temperature profile pattern in the endothermic side. On the other hand, Fig. 8(b) shows that the hydrogen and water mole fraction profiles trend in the recycling gas and permeation sides are completely different from each other. In fact, the hydrogen mole fraction in the lumen side of the in situ H2 addition configuration diminishes due to the permeation through the membrane into the exothermic side while water mole fraction increases in the in situ water removal configuration due to delivery produced water from the exothermic side via permeation.

A comparison of the performance of NC, IHAC and IWRC is summarized in Table 6. As seen, utilizing HSOD membrane in exothermic side leads to 5.72 and 4.38% enhancement in the methanol yield and 6.64 and 4.61% enhancement in the synthesis gas conversion in comparison to NC and IHAC, respectively. However, the change in the reactor configuration does not significantly affect the performance of endothermic side. Additionally, fuel ratio and reactor efficiency values of non-membrane configuration are closer to the optimal conditions compared with those of membrane configurations. Thereupon, poor auto-thermal coupling is the main shortcoming of incorporating membrane to exothermic side of auto-thermal reactor, slightly suppressing the cyclohexane dehydrogenation. A much more efficient auto-thermal coupling would be possible under proper conditions. In spite of the potential advantages of the in situ water removal configuration, declined reactions’ thermal efficiency limits the cyclohexane dehydrogenation. Overall, the operating and design Table 6 Comparison of reactors performance. Reactor

IWRC IHAC NC

Conversion (%)

Yield

Synthesis gas

C6H12

CH3OH

C6H6

H2

28.63 27.31 26.73

81.64 81.12 82.4

0.3723 0.356 0.351

0.8209 0.816 0.821

2.4538 2.4382 2.4766

Fuel ratio

Thermal efficiency

0.8338 0.7628 0.8552

1.2132 1.4867 1.1608

70

F. Rahmani et al. / Journal of Natural Gas Science and Engineering 7 (2012) 60e74

parameters chosen for the permeation side lead to efficient coupling of the two reactions together with simultaneous methanol and hydrogen production. The efficient coupling of exothermic and endothermic reaction in a single vessel reduces the thermal losses associated with the supply of heat for the energy-intensive endothermic process. A promising route can be using elevated feed temperature for the sweep gas stream. Fig. 9(a) and (b) shows the influence of inlet temperature of sweep gas on the temperature profiles in exothermic and endothermic sides along the reactor length for the in situ water removal configuration. With increasing the temperature of sweep gas, axial temperature variation in both reaction sides becomes higher which is due to higher transferred heat from the solid wall. Fig. 10(a)e(c) illustrates the methanol and hydrogen yield; and cyclohexane conversion behave when the inlet temperature of sweep gas increases from 500 to 531 K. Increasing the feed temperature difference of the exothermic and sweep gas streams leads to enhancing the reactor performance in both reaction sides.

As can be seen, cyclohexane conversion reaches to 100% when the inlet temperature of sweep gas is approximately 531 K. This means optimum feed temperature of the sweep gas is 531 K. In this situation, 0.393 and 3 yields are achievable for methanol and hydrogen, respectively. The results indicate that IWRC in optimum conditions is feasible and beneficial. Fig. 11 presents the comparison of methanol and hydrogen production in NC, IHAC, and IWRC in common and optimum conditions. As demonstrated, the methanol production of IWRC in optimum condition (OIWRC) increases about 9.46, 8.29 and 4.48% relative to that of NC, IHAC and IWRC, respectively. In addition, 40.32, 41.14 and 40.9% enhancement in the hydrogen production rate in comparison with NC, IHAC, and IWRC are seen, respectively. This considerable improvement in the methanol and hydrogen production rate of OIWRC is an obvious consequence of the fact that the elevated inlet temperature of sweep gas remedies the poor thermal efficiency, resulting from utilizing membrane.

Fig. 9. Influence of inlet temperature of sweep gas on the temperature profiles in (a) exothermic and (b) endothermic sides along the reactor length for the in situ H2O removal configuration.

F. Rahmani et al. / Journal of Natural Gas Science and Engineering 7 (2012) 60e74

71

Fig. 10. Influence of inlet temperature of sweep gas on (a) methanol yield, (b) cyclohexane conversion and (c) hydrogen yield along the reactor length for the in situ H2O removal configuration.

72

a

F. Rahmani et al. / Journal of Natural Gas Science and Engineering 7 (2012) 60e74

300

CH3OH production rate (tons/day)

250

200

150

100

50

0

b

OIWRC

IWRC

IHAC

NC

16

H2 production rate (ton/day)

14

12

10

8

6

4

2

0

OIWRC

IWRC

IHAC

NC

Fig. 11. The comparison of (a) methanol and (b) hydrogen production in NC, IHAC, IWRC and OIWRC.

6. Conclusion Improvement in production efficiency of auto-thermal methanol synthesis process by only a few percent can result in significant profit increases, energy supply and conservation, and environmental protection. The development of a membraneassisted auto-thermal methanol synthesis reactor could open the environmentally benign way to increasing the methanol and hydrogen production. In this work, feasibility of in situ H2O removal and H2 addition strategies by means of two different membranes (H-SOD and Pd/Ag) in order to enhancing the performance of AMS process was studied by a one-dimensional model., methanol synthesis process is enhanced by using membrane. However, the cyclohexane dehydrogenation performance slightly decreased because of poor auto-thermal coupling. The simulation results show that the in situ H2O removal configuration (IWRC) gives higher conversion and higher total hydrogen production and operates with lower environmental impact due to lower CO2 emission than the other one. However, the most important result is that in this configuration higher catalyst life of exothermic side is possible because of decreasing H2O promoted catalyst deactivation as well as a more favourable exothermic temperature profile along the reactor. Finally, the effect of inlet temperature of sweep gas as

a key operating variable on the axial temperature profile of exothermic and endothermic sides, methanol yield, cyclohexane conversion and hydrogen yield are shown. Increasing sweep gas temperature results in higher methanol yield, cyclohexane conversion and hydrogen yield due to more efficient auto-thermal coupling .The results show that the optimum inlet temperature of permeation side is approximately 531 K, in the operating conditions of the non-membrane configuration as base case. In these conditions, the cyclohexane conversion is 100%. Besides, 10.55 and 9.28% enhancement in the methanol yield and 17.22 and 18.5% enhancement in the hydrogen yield in comparison with NC and IHAC are seen, respectively. These results suggest that this configuration can be a compelling way to boost methanol and hydrogen production and ensure the other desired results, as well. However, an investigation in relation to the environmental aspects, commercial viability and economic feasibility of the proposed configuration is necessary in order to consider commercialization of the process. Nomenclature specific surface area of catalyst pellet (m2 m3) av cross section area of each tube (m2) Ac inside area of inner tube (m2) Ai outside area of inner tube (m2) Ao C total concentration (mol m3) specific heat of the gas at constant pressure (J mol1) Cp particle diameter (m) dp tube inside diameter (m) Di binary diffusion coefficient of component i in j (m2 s1) Dij diffusion coefficient of component i in the mixture Dim (m2 s1) tube outside diameter (m) Do shell inside diameter (m) Dsh partial fugacity of component i (bar) fi F total molar flow rate (mol s1) G mass velocity (kg m2 s1) gasesolid heat transfer coefficient (W m2 K1) hf heat transfer coefficient between fluid phase and reactor hi wall in exothermic side (W m2 K1) heat transfer coefficient between fluid phase and reactor ho wall in endothermic side (W m2 K1) DHf,i enthalpy of formation of component i (J mol1) K rate constant of dehydrogenation reaction (mol m3 Pa1 s1) rate constant for the 1st rate equation of methanol k1 synthesis reaction (mol kg1 s1 bar1/2) rate constant for the 2st rate equation of methanol k2 synthesis reaction (mol kg1 s1 bar1/2) rate constant for the 3st rate equation of methanol k3 synthesis reaction (mol kg1 s1 bar1/2) mass transfer coefficient for component i (m s1) kg K conductivity of fluid phase (W m1 K1) adsorption equilibrium constant for benzene (Pa1) KB adsorption equilibrium constant for component i in Ki methanol synthesis reaction (bar1) equilibrium constant for dehydrogenation reaction (Pa3) Kp equilibrium constant based on partial pressure for Kpi component i in methanol synthesis reaction thermal conductivity of reactor wall (W m1 K1) Kw L reactor length (m) molecular weight of component i (g mol1) Mi N number of components (N ¼ 6 for methanol synthesis reaction, N ¼ 3 for dehydrogenation reaction) P total pressure (for exothermic side: bar; for endothermic side: Pa)

F. Rahmani et al. / Journal of Natural Gas Science and Engineering 7 (2012) 60e74

Pi r1 r2 r3 r4 ri R Rp Re Sci T U ug U vci yi z

partial pressure of component i (Pa) rate of reaction for hydrogenation of CO (mol kg1 s1) rate of reaction for hydrogenation of CO2 (mol kg1 s1) rate of reversed water-gas shift reaction (mol kg1 s1) rate of reaction for dehydrogenation of cyclohexane (mol m3 s1) reaction rate of component i (for exothermic reaction: mol kg1 s1; for endothermic reaction: mol m3 s1) universal gas constant (J mol1 K1) particle radius (m) Reynolds number Schmidt number of component i temperature (K) superficial velocity of fluid phase (m s1) linear velocity of fluid phase (m s1) overall heat transfer coefficient between exothermic and endothermic sides (W m2 K1) critical volume of component i (cm3 mol1) mole fraction of component i (mol mol1) axial reactor coordinate (m)

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