Use of a micro-porous membrane multi-tubular fixed-bed reactor for tri-reforming of methane to syngas: CO2, H2O or O2 side-feeding

international journal of hydrogen energy xxx (xxxx) xxx

Available online at www.sciencedirect.com

ScienceDirect journal homepage: www.elsevier.com/locate/he

Use of a micro-porous membrane multi-tubular fixed-bed reactor for tri-reforming of methane to syngas: CO2, H2O or O2 side-feeding Afshar Alipour-Dehkordi, Mohammad Hasan Khademi* Department of Chemical Engineering, College of Engineering, University of Isfahan, Isfahan, Iran

highlights
The use of side-feeding strategies could be feasible and beneficial for methane tri-reforming process.
The behavior of membrane tri-reformer with oxygen distributed reactant is different compared to other structures.
The membrane tri-reformers are more environmentally friendly compared to the traditional reactor to reach H2/CO ¼ 1.2.

article info

abstract

Article history:

A one-dimensional heterogeneous model for four configurations of a reactor, three micro-

Received 29 July 2019

porous membrane reactors with O2 (O-MMTR), CO2 (C-MMTR) or H2O (H-MMTR) side-

Received in revised form

feeding strategy and one traditional reactor (i.e., multi-tubular fixed-bed reactor (MTR)),

8 October 2019

was developed to explain tri-reforming of methane to produce syngas. Effect of various

Accepted 12 October 2019

side-feeding strategies on reactor performance containing CH4 and CO2 conversion, H2/CO

Available online xxx

ratio, and H2 yield was investigated under the same condition and then described by chemical species and temperature profiles. It was found that use of side-feeding strategies

Keywords:

could be feasible, beneficial, and flexible in terms of change in membrane thickness and

Methane tri-reforming

shell-side pressure for syngas production with H2/CO ¼ 2 which is proper for methanol and

Side-feeding

Fischer-Tropsch process, and ¼ 1.2 which is suitable for DME direct synthesis. However,

Synthesis gas

the syngas produced by the MTR is only appropriate for the methanol and Fischer-Tropsch

Membrane reactor

synthesis under the base case conditions. Also, the results show that the micro-porous membrane reactors have higher CO2 conversion, based on the H2/CO ¼ 1.2; so these strategies are more environmentally friendly compared to the traditional reactor. © 2019 Hydrogen Energy Publications LLC. Published by Elsevier Ltd. All rights reserved.

Introduction Synthesis gas, including a mixture of H2, CO, and small amounts of CO2 is the most important intermediate product. It may be used in the hydroformylation of olefins (OXO process), methanol, and FischereTropsch synthesis with a

proper H2/CO ratio, in acetic acid synthesis from methanol and carbonylation reactions as a source of pure CO, or in ammonia synthesis and hydrogenation reactions as a source of pure H2 [1]. Syngas can be produced from natural gas, coal, biomass, or virtually any hydrocarbon feedstock such as CH4, by reaction with carbon dioxide (dry

* Corresponding author. E-mail address: [email protected] (M.H. Khademi). https://doi.org/10.1016/j.ijhydene.2019.10.097 0360-3199/© 2019 Hydrogen Energy Publications LLC. Published by Elsevier Ltd. All rights reserved. Please cite this article as: Alipour-Dehkordi A, Khademi MH, Use of a micro-porous membrane multi-tubular fixed-bed reactor for trireforming of methane to syngas: CO2, H2O or O2 side-feeding, International Journal of Hydrogen Energy, https://doi.org/10.1016/ j.ijhydene.2019.10.097

2

international journal of hydrogen energy xxx (xxxx) xxx

reforming), oxygen (partial oxidation), or steam (steam reforming) as follows [2]: CH4 þ CO2 42CO þ 2H2 CH4 þ 1=2O2 4CO þ 2H2 CH4 þ H2 O4CO þ 3H2

DH298 ¼ 247 kJ=mol DH298 ¼ 38 kJ=mol DH298 ¼ 206 kJ=mol

(1) (2) (3)

The syngas produced from the methane dry reforming process has a H2/CO ratio of 1. This process has some disadvantages, such as high-temperature requirement, catalyst deactivation due to the extreme endothermicity of the reaction, and demand for a source of pure CO2 [3]. The exothermic catalytic partial oxidation of methane gives a H2/CO ratio of 2, which is suitable for the production of methanol and hydrocarbons through FischereTropsch synthesis [4]. The methane steam reforming process produces syngas with a high H2/CO ratio, and requires a lot of energy to drive this reaction [5]. As mentioned above, two major operational problems in producing the syngas from natural gas through these three processes are to obtain a suitable H2/CO ratio and to prevent the coke deposition [6]. For the first time, tri-reforming of methane (TRM) was proposed by Song [7] which is a combination of steam reforming of methane, dry reforming of methane, and partial oxidation of methane. The heat released from the exothermic partial oxidation of methane is consumed in other two endothermic reactions, so a combination of these reactions increases the energy efficiency [8]. Carbon formation can be reduced in this process because of the existence of steam and oxygen [9]. Furthermore, the TRM process does not need a source of pure CO2 as a reactant, so effluent gases from the combustion processes of power plants can be used as a CO2 source in reaction. Also, in the TRM process, the H2/CO ratio in the syngas can be controlled by varying feed composition [10]. According to the significant advantages mentioned above for the concept of TRM, this process has been extensively studied by many researchers from modeling and experimental points of view. The performance of the various catalysts for tri-reforming of methane was tested by Izquierdo et al. [11], Majewski and Wood [12], Garcia-Vargas et al. [9,13], and Anchieta et al. [14]. Song and Pan [15] investigated the effect of catalyst type on CO2 and CH4 conversion and conducted both experimental and computational analysis to show that TRM can reduce or even eliminate carbon formation, which is a serious issue in the dry reforming of methane. Aboosadi et al. [16] proposed a single fixed-bed tri-reformer instead of a conventional steam reformer followed by a traditional auto-thermal reformer at Lurgi design. Three years later, Khajeh et al. [17] compared a fluidized-bed tri-reformer with the fixed-bed tri-reformer and demonstrated the advantages of fluidized-bed tri-reformer such as lower pressure drop, a decrease in hot spot temperature in catalytic bed, and an increase in CH4 conversion and hydrogen yield. Furthermore, Aboosadi et al. [16] and Khajeh et al. [18] used a onedimensional heterogeneous model as well as the differential evolution algorithm to optimize the fixed- and fluidized-bed tri-reformer. Inlet feed temperature and reactant compositions were selected as decision variables for maximizing the hydrogen yield subject to 1.5 < H2/CO ratio <2 as a constraint

for methanol production. Farniaei et al. [19] coupled theoretically the tri-reforming and dry reforming of methane in a multi-tubular reactor, in which the TRM provides the required energy to drive the dry reforming of methane. Chein et al. [10] investigated the influence of volumetric flow rate, reactant compositions, pressures, and temperature of feed on TRM performance which characterized by H2 and CO yields, CH4 conversion, and H2/CO ratio using a two-dimensional heterogeneous mathematical model. Manenti et al. [20] presented a modeling, simulation and model validation of a biogas-fed tri-reformer followed by a solid oxide fuel cell for power generation, and a furnace for pre-heating of the stream entering the system. Based on the thermodynamic model, energetic and exergetic analyses were conducted by DiezRamirez [21] to investigate the characteristics of the TRM process; Zhang et al. [22] and Minutillo and Perena [23] evaluated the performance of the TRM process combined with the methanol synthesis reactor; also Wiranarongkorn et al. [24] offered thermally coupled steam and tri-reforming processes integrated with a membrane water-gas shift reactor to produce hydrogen from bio-oil aqueous fraction as feedstock. Several restrictions hinder the accomplishment of the TRM on an industrial scale, such as handling of explosive gas mixtures and forming a hot spot temperature in the catalytic bed. A graceful strategy to overcome these restrictions is to use a membrane reactor. Two main uses of the membrane reactors are to obtain higher conversion in reversible reactions and to achieve higher selectivity of desired product in parallel-series reactions. In the first case, a perm-selective membrane allows one or more components to leave the reaction zone, and in the second case, a micro-porous or permselective membrane is used in a controlled manner to add a reactant to the reaction zone [25]. Two group researchers have recently investigated the concept of using membrane as a feed distributor or removing a chemical component from the reaction zone of TRM reactor. Rahimpour et al. [8] studied a multi-tubular fixed bed trireformer for the production of syngas, in which each tube is contained two concentric pipes assisted with hydrogen and oxygen perm-selective membranes. In this study, the air is fed into the oxygen perm-selective membrane inner tube, and oxygen is permeated into the catalytic side. The hydrogen produced in the reaction zone is taken out through the Pdbased membrane and is swept with inert gas. They showed the use of a membrane reactor could lead to achieve higher methane conversion and hydrogen yield at lower reactant inlet temperature and to reduce the hot spot temperature in the catalytic bed. Rahnama et al. [26] developed a mathematical model to simulate the tri-reforming process as a heat source for the steam reforming of methane in hydrogen permselective two-membrane thermally coupled reactor. The advantages of this reactor configuration are the production of two types of syngas with different H2/CO ratios, getting pure hydrogen, elimination of fired furnace of the conventional steam reformer by substituting it with TRM and reaching higher CH4 conversion. Recently, a micro-porous membrane packed-bed reactor with different side-feeding strategies has been used for direct propylene epoxidation with hydrogen and oxygen to produce propylene oxide (PO) by Lu et al. [27] and Kertalli et al. [28].

Please cite this article as: Alipour-Dehkordi A, Khademi MH, Use of a micro-porous membrane multi-tubular fixed-bed reactor for trireforming of methane to syngas: CO2, H2O or O2 side-feeding, International Journal of Hydrogen Energy, https://doi.org/10.1016/ j.ijhydene.2019.10.097

3

international journal of hydrogen energy xxx (xxxx) xxx

They investigated the effect of O2, H2, or mixture of O2 and H2 side-feeding strategies on the C3H6 conversion, PO selectivity, H2 efficiency, and PO yield; and the results were compared with the traditional packed-bed reactor. The results show that the reactor performance can be improved by controlling the feeding strategy of the reactants. Therefore, controlling the feed composition or in other words adjusting the side-feeding strategy in a micro-porous membrane reactor leads to control the reaction pathway and improve the reactor performance. Also, the side-feeding membrane reactor prevents the direct mixing of explosive gases, which increases the process’s safety. To authors’ best knowledge, this is the first time that the use of a micro-porous membrane with different side-feeding strategies has been conducted for MTR, and it is expected that this strategy will improve the reactor efficiency. So in this study, four types of configuration containing three microporous membrane reactors with O2, CO2 or H2O side-feeding strategy and one traditional multi-tubular fixed-bed reactor is considered to describe tri-reforming of methane for producing synthesis gas. In this regard, a one-dimensional heterogeneous model is developed to investigate the influence of various side-feeding strategies on reactor performance, including CH4 and CO2 conversion, H2/CO ratio, and H2 yield as well as catalyst temperature, molar flow rate of distributed species, and permeation flux through the membrane.

Process description Methane-steam reformer Fig. 1 shows a schematic diagram of the conventional methane-steam reformer (MSR) containing a vertical multitubular furnace-reactor, in which heat is transferred from the combustion gases through tubes by radiation and convection mechanisms. Methane and steam are fed to the tubes packed with Ni catalyst. The specifications of feed composition, operating conditions, and geometry of industrial MSR to produce syngas for methanol synthesis in Zagros Petrochemical Company, Asaluyeh, Iran, are listed in Table 1 [8].

Pre reformed gas/steam Natural gas

Combustion air

Table 1 e Feed composition, operating conditions, and geometry of MSR [8]. Parameter Feed composition (mole %) CO2 CO H2 CH4 N2 H2O Operating conditions Inlet temperature (oC) Inlet pressure (bar) Feed gas flow rate (kmol/h) Heat load on tube (kcal/m2.h) Geometry Number of tubes (in 4 rows) Tube inside diameter (m) Reactor length (m) Catalyst volume loading (m3) Void fraction () Catalyst shape Particle size (mm)

Value 1.72 0.02 5.89 32.59 1.52 58.26 520 40 9129.6 68730 184 0.125 12 27.8 0.4 10-HOLE rings 19  16

Methane tri-reformer A schematic diagram of methane tri-reformer (MTR) is presented in Fig. 2(a). The MTR consists of 184 0.125 m ID tubes and 2 m long, which are placed in a 2.5 m ID shell. The tubes are loaded with NiOeMg/CeeZrO2/Al2O3 catalyst, in which the size and shape of catalysts are the same as MSR catalysts. This catalyst has a good performance for reducing the coke formation on the catalyst surface and reactor wall [19]. Since TRM is an autothermal reaction, no external heat is required to drive the reaction. The molar ratio of feed components and operating conditions of MTR tabulated in Table 2, are similar to those reported by Aboosadi et al. [16].

Membrane methane tri-reformer The membrane methane tri-reformer (MMTR) is a multitubular reactor, in which each tube is covered with a microporous membrane. As shown in Fig. 2(b)e(d), three configurations of MMTR namely, O-MMTR, H-MMTR, and CMMTR are considered in which oxygen, steam, and carbon dioxide play the role of distributed reactants in each configuration, respectively, and permeate through the ceramic membrane. The other feed components except the distributed components are fed to the tube side, where the tri-reforming reaction occurs over the catalyst. To establish a base case, the inlet temperature and pressure of distributed reactants in the shell side are considered 1100 K and 24 bar, respectively; and other operating and geometry parameters are similar to those used for the MTR.

Reaction scheme and kinetics Synthesis gas Combustion products

Fig. 1 e A schematic diagram of conventional MSR.

The Reactions occurring in the MSR include steam reforming of methane (Eqs. (3) and (4)) and the water-gas shift (WGS) reaction (Eq. (5)). Four reactions (3)e(6) are considered for

Please cite this article as: Alipour-Dehkordi A, Khademi MH, Use of a micro-porous membrane multi-tubular fixed-bed reactor for trireforming of methane to syngas: CO2, H2O or O2 side-feeding, International Journal of Hydrogen Energy, https://doi.org/10.1016/ j.ijhydene.2019.10.097

4

international journal of hydrogen energy xxx (xxxx) xxx

Fig. 2 e A schematic diagram of (a) MTR, (b) O-MMTR, (c) H-MMTR and (d) C-MMTR.

CH4 þ 2H2 O4CO2 þ 4H2 Table 2 e Molar ratio of feed components and operating conditions of MTR [16]. Parameter

CO þ H2 O4CO2 þ H2

DH298 ¼ 164:9 kJ=mol

(4)

DH298 ¼ 41:1 kJ=mol

(5)

Value CH4 þ 2O2 4CO2 þ 2H2 O

Molar ratio of feed components CO2/CH4 CO/CH4 H2/CH4 O2/CH4 N2/CH4 H2O/CH4 Operating conditions Inlet temperature (K) Inlet pressure (bar) Feed gas flow rate (kmol/h)

1.33 0.00053 0.082 0.47 0.00053 2.47 1100 20 9129.6

describing the tri-reforming process due to the validation of the kinetic rate expressions of these reactions by Aboosadi et al. [16]. Dry reforming of methane (reaction (1)) is a dependent reaction and can be written as reaction (3) minus reaction (5) [16].

DH298 ¼ 802:7 kJ=mol

(6)

The kinetic model proposed by Xu and Froment [29] over Ni-based catalyst is used for reaction (3)e(5). This model is more general for the steam reforming of methane and has been tested under lab-scale conditions [30]. The kinetic model used for methane combustion (Eq. (6)) is given by Trimm and Lam [31]. Although this model was obtained over the Pt-based catalyst, the parameters of the adsorption model are set for Ni-base catalyst [32]. The rates of reactions (3)e(6) are represented, respectively, as follow: k1 r3 ¼ 2:5 PH2

r4 ¼

k2 P3:5 H2

P3H PCO PCH4 PH2 O  2 KI

PCH4 P2H2 O 

! 

P4H2 PCO2 KІІ

1 ∅2

! 

(7)

1 ∅2

(8)

Table 3 e Kinetic parameters of reaction equilibrium constant, reaction rate constant and adsorption equilibrium constant for reactions (2) and (4)e(6) [16]. Parameter

An

Bn

2

26830 22430 4400

30.114 26.078 4.036

KI (bar ) KII (bar2) KIII k1 k2 k3 k4a k4b KCH4 KCO KH2 KH2 O KCHC 4 K OC 2

k0j (mol/kgcat.s)

1.17 2.83 5.43 8.11 6.82

    

k0i (bar1)

1015 bar0.5 1014 bar0.5 105 bar1 105 bar2 105 bar2 6.65  8.23  6.12  1.77  1.26  7.78 

104 105 109 105 101 107

Ej =DHi (J/mol)

240100 243900 67130 86000 86000 38280 70650 82900 88680 27300 92800

Please cite this article as: Alipour-Dehkordi A, Khademi MH, Use of a micro-porous membrane multi-tubular fixed-bed reactor for trireforming of methane to syngas: CO2, H2O or O2 side-feeding, International Journal of Hydrogen Energy, https://doi.org/10.1016/ j.ijhydene.2019.10.097

international journal of hydrogen energy xxx (xxxx) xxx

5

Fig. 3 e An element of Dz along the axial direction of O-MMTR.

k3 r5 ¼ PH2 r6 ¼



PH PCO2 PCO PH2 O  2 KІІІ



1  2 ∅

(9)

k4a PCH4 PO2 k4b PCH4 PO2 þ 1 þ KCHC PCH4 þ KOC PO2 1 þ KCHC PCH4 þ KOC PO2 4

4

2

∅ ¼ 1 þ KCO PCO þ KH2 PH2 þ KCH4 PCH4 þ KH2 O

(10)

2

PH2 O PH2

(11)

where Pi is the partial pressure of component i, Kn is reaction equilibrium constant, and kj as reaction rate constant and Ki as adsorption equilibrium constants can be defined by Arrhenius and Van’t Hoff relations, respectively: Kn ¼ exp

  An þ Bn Ts

n ¼ I; II; III

(12)

  Ej kj ¼ k0j exp  RT

j ¼ 1; 2; 3; 4a; 4b

(13)

  DHi Ki ¼ k0i exp  RT

i ¼ CH4 ; CO; H2 ; H2 O; CHC4 ; OC2

(14)

The kinetic parameters of reaction equilibrium constant, the reaction rate constant, and adsorption equilibrium constant are given in Table 3.


The multi-tubular reactor is operated at steady-state conditions.
Since L=dp > 30, so mass and heat diffusions in the axial direction are negligible compared to gas bulk movement [33].
Chein et al. [10] showed that the simulation results of the two-dimensional model and the one-dimensional model behavior reported by Aboosadi et al. [16] are close to each other. Also, since Dim L=ðug D2t Þ < 0:00075, the radial diffusion effect can be ignored [34]. Therefore, mass and heat diffusions in the radial direction are negligible (one-dimensional model).
Bed void fraction in radial and axial directions is constant.
The tubes of MTR are wholly insulated, so no heat loss occurs during this reactor.
As the possibility of the gas permeation into the shell side is very low due to the positive pressure difference between the shell and tube sides (~4 bar), it is assumed that no chemical reaction occurs in the shell side. Based on the above assumptions, mass and energy balance equations are derived by considering a differential element along the axial direction of the reactor. A schematic of element Dz inside the O-MMTR is shown, as a sample, in Fig. 3.

Solid-phase The mass and energy balance equations for the solid phase are as follows:

Mathematical model A one-dimensional heterogeneous model has been established to specify the concentration and temperature distributions along the reactor length by considering the mass and heat transfer resistances. The following assumptions are supposed to simplify this model:
Since the gas compressibility factor at the inlet operating conditions is equal to 1.0019 (calculated by Aspen HYSIS V10), the gas phase behavior is assumed to be ideal.

6 X   hj vi;j rj ¼ 0 av kgi Ct yti  yis þ rB

(15)

j¼3

6 X   hj rj ð  DHR Þj ¼ 0 av hf Tt  Ts þ rB

(16)

j¼3

where ni;j is the stoichiometric coefficient of component i in reaction j. Effectiveness factor hj , is considered to account for intra-particle transport limitation of catalysts. This factor is reported for reaction (3)e(6) by De Groote and Froment [35] as: h3 ¼ 0.07, h4 ¼ 0.06, h5 ¼ 0.7, and h6 ¼ 0.05.

Please cite this article as: Alipour-Dehkordi A, Khademi MH, Use of a micro-porous membrane multi-tubular fixed-bed reactor for trireforming of methane to syngas: CO2, H2O or O2 side-feeding, International Journal of Hydrogen Energy, https://doi.org/10.1016/ j.ijhydene.2019.10.097

6

international journal of hydrogen energy xxx (xxxx) xxx

Fluid phase in tube side

Boundary conditions

The mass and energy balance equations for the fluid phase in the tube side are expressed, respectively, as follows:

The Dirichlet boundary conditions have been used to complete the governing equations. At the inlet of each reactor, the gas composition, temperature and pressure in the tube side as well as temperature and molar flow rate of distributed species i in the shell side are specified; therefore:

    1 v Ft yti Ji þ av kgi Ct yis  yti þ 4b ¼ 0 NAc vz Di

(17)

  ZTt   4aU  s  4b 1 v Ft cp;mix Tt þ av hf Ts  Tt þ T  Tt þ Ji cpi dT ¼ 0 NAc Di Di vz

( at

z¼0

Ts

(18) where b ¼ 1 for the compounds which permeate through the membrane in MMTR and b ¼ 0 for conventional MSR and MTR. Also, a is a coefficient which is equal to zero for MTR because of the adiabatic condition and equal to one for MMTR and MSR due to heat transferred between tube- and shell-sides.

yti ¼ yti0 Fsi

¼

Fsi0

Tt ¼ Tt0 T ¼ s

Pt ¼ Pt0

Ts0

in the tube side in the shell side

(27)

where in the shell side, i refers to O2, H2O, and CO2 species for O-MMTR, H-MMTR, and C-MMTR, respectively.

Auxiliary correlations

The following mass and energy balance equations are considered for the fluid phase in the shell side of MMTR:

The auxiliary correlations must be employed in the model to complete the simulation. In this regard, a set of auxiliary correlations for estimation of the heat and mass transfer coefficient between the gas and solid phases, diffusion coefficient, and the heat transfer coefficient between the gas phase and reactor wall are listed in Table 4.

vFsi þ NpDi Ji ¼ 0 vz

(19)

Numerical solution

  ZTt  s  v Fsi cpi Ts t þ NUpDi T  T þ NpDi Ji cpi dT ¼ vz

(20)

Fluid phase in shell side

Ts

where N is the number of tubes in the shell side, and i refers to O2, H2O, or CO2 in each configuration.

Permeation through a micro-porous membrane The permeation rate of gas through the membrane, which is calculated from the dusty gas model [36] is introduced as follows:    B0   1 Dei  t Pi  Psi þ Psi Pt  Ps (21) Ji ¼ RTm d dmi 1 1 1 ¼ þ Dei Dim Dei;k

Dei;k

sffiffiffiffiffiffiffiffiffi 8RT ¼ K0 pMi

2 εm K0 ¼ rp 3 t 1 εm B0 ¼ r2p 8 t

(22)

(23)

The governing equations are a set of simultaneous algebraic equations containing auxiliary correlations, and ordinary differential equations (ODEs) including conservation laws. ODEs are discretized using backward finite difference approximation method to make a set of non-linear algebraic equations. To solve all non-linear algebraic equations, the reactor length is divided into 500 separate sections, and the Gauss-Newton method in MATLAB 2015 programming environment is applied in each section.

Results and discussions In this section, a numerical simulation is performed to predict the mole fraction of each chemical component and temperature profile in the four structures of MTR, O-MMTR, H-MMTR, and C-MMTR. The impact of reactor types on the performance of the reactor is also evaluated. The performance of all reactors is analyzed in terms of methane conversion, hydrogen yield, and H2/CO ratio, which is presented as follows:

(24) CH4 conversion ¼

FCH4 ;in  FCH4 FCH4 ;in

(28)

(25) H2 yield ¼

FH2  FH2;in FCH4 ;in

(29)

The Ergun equation is used to represent the pressure drop along the reactor length.

H2 yield ¼

FH2  FH2;in FCH4 ;in

(30)

ð1  εB Þu2g r ð1  εB Þ2 mug dPt ¼ 150 þ 1:75 2 dz ε3B dp ε3B dp

where FCH4 ; FH2 ; and FCO is the molar flow rate of CH4, H2, and CO, respectively, at a position along the reactor length.

Pressure drop

(26)

Please cite this article as: Alipour-Dehkordi A, Khademi MH, Use of a micro-porous membrane multi-tubular fixed-bed reactor for trireforming of methane to syngas: CO2, H2O or O2 side-feeding, International Journal of Hydrogen Energy, https://doi.org/10.1016/ j.ijhydene.2019.10.097

7

international journal of hydrogen energy xxx (xxxx) xxx

Table 4 e Auxiliary correlations. Parameter

Equation 2=3

kgi ¼ 1:17Re0:42 Sci

Mass transfer coefficient between the gas and solid phases

Ref. ug

[37]

d p ug r Re ¼ m m Sci ¼ rDim Molecular diffusion coefficient in a multicomponent gas mixture

qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi 0:00143T1:75 1=Mi þ 1=Mj Dij ¼ pffiffiffi 1=3 1=3 2 2P½nci þ ncj 

Binary diffusion coefficient

Overall heat transfer coefficient Heat transfer coefficient between the gas phase and reactor wall Heat transfer coefficient between the bulk gas phase and the solid phase

Model validation The proposed mathematical model has been validated by comparing the MSR simulation results with the plant data of Zagros Petrochemical Company, Asaluyeh, Iran, under the geometry and operating conditions listed in Table 1. It is observed in Table 5 that simulation results were in good agreement with the plant data (mean relative error of 3.7 for outlet composition); and the proposed model was an appropriate approximation of the actual plant. Therefore, this model also will be reliable to simulate MTR and MMTR.

Effect of reactor types under the same methane conversion The effect of reactor types on the performance of the reactor is investigated on the basis of the same methane conversion, as shown in Fig. 4. To simplify the benchmarking, simulation is performed for a “base case”. The geometry and operating conditions of four reactor configurations (MTR, O-MMTR, HMMTR, and C-MMTR) for the base case are similar to those listed in Table 2. Also, the parameters of ceramic membrane

Table 5 e Comparison between model prediction and conventional MSR data [8]. parameter

plant

model

REa (%)

Outlet composition (mole %) CO2 CO H2 CH4 N2 H 2O  Outlet temperature ( C) CH4 conversion (%)

5.71 3.15 31.39 20.41 1.29 38.05 710 26.5

5.84 3.42 32.77 19.7 1.28 37 724.7 28.3

2.28 8.57 4.40 3.48 0.77 2.76 2.07 6.79

a

1y Dim ¼ P yi j jsi Dij

RE (%) ¼ ((plant data e model prediction)/plant data)  100.

1 1 D lnðDo =Di Þ Di 1 ¼ þ i þ U hi 2Kw Do ho K 0:8 0:33 g hi ¼ 0:027Re Pr Di K 0:585 0:33 g hf ¼ 1:17Re Pr dp cp m pr ¼ Kg

[38]

[39]

[37] [40]

used in this model are rp ¼ 1.5  106 m, εm =t ¼ 0.112, and d ¼ 5.5  103 m [41]. Moreover, it is assumed that the same methane conversion in membrane reactors is determined over the same reactor length (i.e., z ¼ 0.3 m) by changing the inlet pressure in the shell side. It is observed in Fig. 4(a), that there is not a considerable difference between the behavior of methane conversion along the length of MTR, H-MMTR, and C-MMTR; however, the H-MMTR shows higher methane conversion than the other reactor configurations throughout the reactor length. It is worth noting that this characteristic is difficult to be explained just by the reaction orders of H2O in Eqs. (3) and (4). The inlet pressure in the shell side of O-MMTR, H-MMTR, and C-MMTR is obtained 26.48, 28, and 26 bar, respectively, to reach methane conversion of 94%. Therefore, it should be exerted more pressure to the shell side of HMMTR to reach CH4 conversion of 94% under the same conditions, which leads to an increase in energy consumption. The MTR has achieved this conversion at catalyst length of 0.38 m, so, the MTR needs a longer catalyst length compared to the membrane reactors (i.e., nearly 27%) under the same methane conversion. Since the methane conversion in OMMTR depends strongly on the permeation rate of O2 through the ceramic membrane, the methane conversion increases slowly compared with the other reactor types. To obtain the same methane conversion, the O2/CH4 ratio should be increased from 0.47 to 0.68 at the feed conditions of the OMMTR, which means that the O-MMTR requires more oxygen flow rate (i.e., 44.7%) compared to the other configurations. Fig. 4(b) illustrates the H2/CO ratio along the catalyst length for the four types of the reactor under the same methane conversion (i.e., 94%). At the catalyst bed outlet, the C-MMTR (i.e., 1.96) shows higher H2/CO ratio than the MTR (i.e., 1.63), OMMTR (i.e., 1.74), and H-MMTR (i.e., 1.08). According to the H2/ CO ratio reported by Saad and Williams [42] in the range of 1.5e2.0 as the feed stream for the methanol synthesis plant and 1.7e2.15 for the Fischer-Tropsch process, the results indicate that H2/CO ratio taken from MTR, O-MMTR, and C-

Please cite this article as: Alipour-Dehkordi A, Khademi MH, Use of a micro-porous membrane multi-tubular fixed-bed reactor for trireforming of methane to syngas: CO2, H2O or O2 side-feeding, International Journal of Hydrogen Energy, https://doi.org/10.1016/ j.ijhydene.2019.10.097

8

international journal of hydrogen energy xxx (xxxx) xxx

x 10 4.8

-4

2.2

(e)

3

3.6

2

2.4

1

1.2 H-MMTR C-MMTR O-MMTR

0

0

0.05

0.1

0.15

0.2

0.25

0.3

0.35

0 0.4

Length of reactor (m)

x 10

-6

x 10 7 H-MMTR C-MMTR 6 O-MMTR: Permeation O-MMTR: Consumption

(f) 2

1.8

5

1.6

4

1.4

3

1.2

2

1

1

0.8

0

0.05

0.1

0.15 0.2 0.25 Length of reactor (m)

0.3

0.35

0 0.4

-6

O2 consumed and permeated through membrane (kmol/s)

-3

H2O and CO2 permeated through membrane (kmol/s)

x 10

Molar flow rate of O 2 (kmol/s)

Molar flow rate of H 2O and CO2 (kmol/s)

4

Fig. 4 e Variation of (a) methane conversion, (b) H2/CO ratio, (c) H2 yield, (d) catalyst temperature, (e) molar flow rate of distributed species in the tube side, and (f) permeation rate of distributed species and O2 consumption along the catalyst length of MTR, O-MMTR, H-MMTR, and C-MMTR based on the same methane conversion (i.e., 94%).

MMTR is suitable for these processes. The appropriate range of H2/CO ratio for DME production is 1.2e1.4 [6], so the HMMTR is almost ideal for the DME synthesis process. The variations of H2 yield along the catalyst length of MTR and three membrane reactors are also shown in Fig. 4(c). At the catalyst bed outlet, H2 yield approaches to 1.71, 1.73, 1.41, and 1.83 in the MTR, O-MMTR, H-MMTR, and C-MMTR, respectively. The variation of hydrogen yield in the C-MMTR is higher than other types of reactors. Low concentration of CO2 in C-MMTR leads to a decrease in the rate of reverse WGS reaction, and this subject causes more production of H2 and gives a richer H2/CO ratio and H2 yield than other types of reactors.

Fig. 4(d) shows a comparison between axial catalyst temperature profiles along the MTR, O-MMTR, H-MMTR, and CMMTR. The H-MMTR and MTR have almost a similar behavior in terms of the temperature profile. Also, throughout the catalyst bed, the C-MMTR displays a higher temperature profile than the MTR and H-MMTR. In other words, the MTR (i.e., 1457 K) and H-MMTR (i.e., 1520 K) have a lower hot spot temperature compared to the C-MMTR (i.e., 1666 K). In addition, at the entrance of the O-MMTR, since the released heat by the reaction (6) isn’t enough for endothermic reactions (3) and (4) due to the low oxygen content, thus some of the required heat for endothermic reactions take from sensible heat of gas, and temperature profile decreases rapidly. After

Please cite this article as: Alipour-Dehkordi A, Khademi MH, Use of a micro-porous membrane multi-tubular fixed-bed reactor for trireforming of methane to syngas: CO2, H2O or O2 side-feeding, International Journal of Hydrogen Energy, https://doi.org/10.1016/ j.ijhydene.2019.10.097

international journal of hydrogen energy xxx (xxxx) xxx

that, the temperature begins to increase as a result of oxygen permeation gradually along the reactor length. At the first part of the O-MMTR, the temperature is lower, and at the end part, the temperature is higher than other reactors. Increasing the temperature smoothly in the O-MMTR results in the presence of CH4 conversion profile under other profiles throughout the reactor (see Fig. 4(a)). The profile of molar flow rate of distributed species in the tube side is demonstrated in Fig. 4(e). In the first half of the OMMTR (reactor length ¼ 0e0.1 m), the amount of oxygen permeated through the membrane is more than oxygen consumed in the reaction (6) (see Fig. 4(f)), hence the molar flow rate of oxygen, as a reactant, increases. Afterward, the opposite situation occurs and the oxygen permeation flux becomes lower than its consumption rate, so oxygen content decreases in the tube side. At the entrance of H-MMTR and CMMTR, the molar flow rate of H2O and CO2 increases rapidly as the production of these species increases with temperature at this part of reactors. After that, the molar flow rate of these compounds increases smoothly as the permeation rate of H2O and CO2 through the membrane dominates their consumption rate (see Fig. 4(d)). Fig. 4(f) indicates the variations of permeation rate of distributed species in O-MMTR, H-MMTR, and C-MMTR and the consumption rate of oxygen in O-MMTR. A decrease in the amount of O2 consumption is observed near the reactor entrance, which is associated with the reduction of catalyst temperature in this region, as shown in Fig. 4(d). Subsequently, as the temperature increases, the amount of oxygen consumed increases; finally, the rate of O2 consumption decreases, mainly due to fuel depletion. Since the permeation rate through the membrane is strongly inversely dependent on the gas temperature, the permeation rate of distributed species decreases with the gas temperature and vice versa.

Effect of reactor types under the same H2/CO ratio for methanol production and Fischer-Tropsch process Fig. 5(b) shows a comparison between four types of reactors with different side-feeding strategies under the same H2/CO ratio (i.e., 2), which is proper for methanol production and Fischer-Tropsch process. The results are obtained for all membrane reactors under the “base case” mentioned in Table 2, the shell side pressure of 24 bar and the ceramic membrane parameters of rp ¼ 1.5  106 m and εm =t ¼ 0.112 [41]. As you can see in this figure, the H2/CO ratio in the O-MMTR and CMMTR has a decreasing trend due to the increase in CO2 concentration along the catalyst length, which leads to reacting the reverse WGS reaction. Also, the H2/CO ratio in the H-MMTR decreases rapidly at the entrance of the reactor because of low H2O concentration and after that increases slowly along the reactor with a permeation rate of H2O through the membrane. It should also be noted that according to the “base case”, the H2/CO ratio in the MTR approaches nearly 1.7 at the end of the reactor and this ratio reaches 2 only at the entrance of the reactor (i.e., z ¼ 0.2 m). Moreover, it is considered that the same H2/CO ratio is determined in membrane reactors over the same reactor length (i.e., z ¼ 1.5 m) by changing the thickness of the membrane. To achieve H2/CO ratio of 2 at the reactor outlet, the membrane thickness of O-

9

MMTR (i.e., 1.7 mm), H-MMTR (i.e., 0.16 mm), and C-MMTR (i.e., 0.62 mm) have been obtained. The membrane thickness in the C-MMTR and O-MMTR is approximately 4 and 10 times the membrane thickness in the H-MMTR. So, the O-MMTR requires the maximum membrane thickness to reach H2/CO ¼ 2 in comparison with other membrane reactors; which leads to more cost. Furthermore, to reach the same H2/CO ratio, the H2O/CH4 ratio should be increased from 2.47 to 4 at the feed conditions of H-MMTR, which means that H-MMTR requires more H2O flow rate (i.e., 61.9%) compared to the other sidefeeding strategies. As shown in Fig. 5(a), (c), and (d), to better understand the difference between four reactors, the CH4 conversion, H2 yield, and catalyst temperature profile are also plotted against the reactor length under the same H2/CO ratio (i.e., 2). As can be seen in Fig. 5(a), the H-MMTR (i.e., 99.1%), C-MMTR (i.e., 98.7%), and MTR (i.e., 98.32%) show higher CH4 conversion than OMMTR (i.e., 79.3%). The CH4 conversion in the O-MMTR smoothly increases as a result of the gradual permeation of oxygen from the micro-porous membrane. In Fig. 5(c), the H2 yield profile in the C-MMTR is located above the other profiles at each point of the reactor. The H2 yield at the outlet of the MTR, O-MMTR, H-MMTR, and C-MMTR reaches 1.84, 1.72, 1.98, and 1.96, respectively. The C-MMTR has the highest hot spot temperature (i.e., 1666 K) compared to other reactors, as shown in Fig. 5(d). It should be noted that most ceramic membranes have a high melting point, for example, the melting point of the a-Al2O3 ceramic membrane with a nominal grain size of 5 mm is 2050 ± 4  C [43]. Consequently, there is no temperature restriction for using the ceramic membranes and the hot spot temperature formed in membrane reactors does not damage those. No hot spot forms in the OMMTR, and the temperature remains below 1100 K throughout the reactor. In Fig. 5(e), the trend of molar flow rate of distributed species is similar to that shown in Fig. 4(e). The difference is in the position where the peak of the oxygen molar flow rate curve occurs (i.e., reactor length ¼ 0.5 m), which coincides with equality of the amount of oxygen consumed and permeated through the membrane.

Effect of reactor types under the same H2/CO ratio for DME direct synthesis process To study the effect of reactor types on the performance of TRM, another investigation is also proposed in this work. A comparison between four types of reactors is carried out under the same H2/CO ratio (i.e., 1.2), which is proper for direct production of DME from synthesis gas. As shown in Fig. 6(b), the H2/CO ratio at the reactor outlet is set to 1.2 by changing the membrane thickness over the length of 1.5 m. The operating and geometry parameters are similar to those used for the previous section. Also, the ceramic membrane parameters are rp ¼ 9.3  108 m and εm =t ¼ 0.3 [44]. The membrane thickness is determined 0.5, 0.6, and 0.23 mm for O-MMTR, HMMTR, and C-MMTR, respectively. Also note that the inlet O2/ CH4 ratio in the O-MMTR and CO2/CH4 ratio in the C-MMTR should be increased from 0.47 to 0.83 and 1.33 to 2.48, respectively; so, O-MMTR and C-MMTR need excess O2 (i.e., 77%) and excess CO2 (i.e., 86%) compared to other

Please cite this article as: Alipour-Dehkordi A, Khademi MH, Use of a micro-porous membrane multi-tubular fixed-bed reactor for trireforming of methane to syngas: CO2, H2O or O2 side-feeding, International Journal of Hydrogen Energy, https://doi.org/10.1016/ j.ijhydene.2019.10.097

international journal of hydrogen energy xxx (xxxx) xxx

x 10 1.2

Molar flow rate of H 2O and CO2 (kmol/s)

0.012 H-MMTR C-MMTR O-MMTR

(e) 0.01

1

0.008

0.8

0.006

0.6

0.004

0.4

0.002

0.2

0

0

0.5

1

-4

Molar flow rate of O 2 (kmol/s)

10

0 1.5

Length of reactor (m)

Fig. 5 e Variation of (a) methane conversion, (b) H2/CO ratio, (c) H2 yield, (d) catalyst temperature, and (e) molar flow rate of distributed species in the tube side along the catalyst length of MTR, O-MMTR, H-MMTR, and C-MMTR under the same H2/ CO ratio (i.e., 2).

configurations. This figure displays a decreasing trend in the H2/CO ratio for the MTR, O-MMTR, and C-MMTR and almost a smooth variation in the H2/CO ratio for the H-MMTR. As mentioned above, the C-MMTR requires more amount of CO2 in the feed compared to other reactors to reach H2/CO ratio of 1.2. A more decreasing trend in the H2/CO ratio is found with CO2 concentration for the C-MMTR because of high CO2 concentration increases the rate of reverse WGS reaction. It is noteworthy that the H2/CO ratio in the MTR does not reach 1.2 at any point in the reactor under the base case conditions. The results indicate that membrane reactors exhibit a different behavior under the same H2/CO ratio of 1.2 compared to the same H2/CO ratio of 2. In this regard, Fig. 6(a) shows the variation of CH4 conversion along the reactor

length for four types of the reactor. At the first part of OMMTR, the CH4 conversion increases slowly and then reaches a constant value of 100% at the second part of the reactor. Compared to other reactors, the O-MMTR has the lowest CH4 conversion at the first half of the reactor, but it has the highest value at the reactor outlet due to excess oxygen in the feed stream and generation of more heat to drive reforming reactions. Besides, the MTR, H-MMTR, and C-MMTR have a similar trend in terms of CH4 conversion and reach 98.32%, 95.87%, and 98.62%, respectively. The effect of reactor types on the H2 yield is indicated in Fig. 6(c). The H2 yield has an increasing trend at the first part of the O-MMTR until the CH4 conversion approaches 100% (see Fig. 6(a)); after that, a decreasing trend is found at the second

Please cite this article as: Alipour-Dehkordi A, Khademi MH, Use of a micro-porous membrane multi-tubular fixed-bed reactor for trireforming of methane to syngas: CO2, H2O or O2 side-feeding, International Journal of Hydrogen Energy, https://doi.org/10.1016/ j.ijhydene.2019.10.097

11

-3

x 10 4.5

(e)

H-MMTR C-MMTR O-MMTR

4

3

2

0

1.5

0

0.5

1

-4

0 1.5

Length of reactor (m)

2

x 10

Molar flow rate of O 2 (kmol/s)

Molar flow rate of H 2O and CO2 (kmol/s)

6

O consumed and permeated through membrane (kmol/s)

international journal of hydrogen energy xxx (xxxx) xxx

4

x 10 (f)

3.5

-6

O-MMTR: Consumption O-MMTR: Permeation

3 2.5 2 1.5 1 0.5 0

0

0.5 1 Length of reactor (m)

1.5

Fig. 6 e Variation of (a) methane conversion, (b) H2/CO ratio, (c) H2 yield, (d) catalyst temperature, (e) molar flow rate of distributed species in the tube side, and (f) rate of permeation and consumption of O2 along the catalyst length of MTR, OMMTR, H-MMTR, and C-MMTR under the same H2/CO ratio (i.e., 1.2).

part of the reactor, where the reverse WGS is dominated reaction. Therefore, the WGS reaction plays a decisive role in the H2 yield profile. In Fig. 6(d), the highest temperature is observed at the entrance of the MTR (i.e., 1460 K), H-MMTR (i.e., 1519 K), and CMMTR (i.e., 1665 K) as a result of methane combustion. Since the reactions (3), (4), and (6) do not occur in the second half of the O-MMTR, the temperature remains almost constant along the reactor axis. The results show that in the second half of the reactors, the temperature of the O-MMTR is nearly 230 K more than other reactors. Fig. 6(e) shows how the molar flow rate of distributed species in the tube side changes along the length of O-

MMTR, H-MMTR, and C-MMTR. The alterations of molar flow rate of H2O and CO2 are similar to those shown in Fig. 5(e). However, a different trend is observed for the molar flow rate of O2 in the O-MMTR after the axial position of 0.75 m, where coincides with the lack of methane in the tube side. After this location, the rate of reaction (6) becomes zero at the axial position close to 1 m, and no oxygen is consumed (see Fig. 6(f)). The temperature behavior of the catalyst, as shown in Fig. 6(d), also confirms this content. Therefore, the oxygen molar flow rate increases just due to the permeation of O2 through the membrane. As a result, the use of membrane after the axial position of 1 m

Please cite this article as: Alipour-Dehkordi A, Khademi MH, Use of a micro-porous membrane multi-tubular fixed-bed reactor for trireforming of methane to syngas: CO2, H2O or O2 side-feeding, International Journal of Hydrogen Energy, https://doi.org/10.1016/ j.ijhydene.2019.10.097

12

international journal of hydrogen energy xxx (xxxx) xxx

Fig. 7 e Comparison between MTR, O-MMTR, H-MMTR, and C-MMTR in terms of CO2 conversion under the same H2/ CO ¼ 2 and 1.2.

is not beneficial to the O-MMTR, and eliminating it from the reactor reduces costs.

CO2 utilization Today, one of the energy utilization in the industry is the combustion of carbonaceous fuels, for example, fossil fuelscoal, petroleum, and natural gas. Combustion and complete oxidation of carbon-based compounds can produce CO2; therefore, the concentration of CO2 in the atmosphere increases. Although CO2 is a harmless gas that plays a vital role in the Earth’s carbon cycle, it is also a greenhouse gas which causes increasing concerns for global warming [45]. Methane tri-reforming process is one of the efficient technologies for converting CO2 as a co-reactant to synthesis gas [15]. The purpose of this section is to compare these four types of the reactor in terms of CO2 conversion under the same H2/CO ¼ 2 for methanol production and Fischer-Tropsch process and ¼ 1.2 for DME direct synthesis, as shown in Fig. 7. Under the same H2/CO ¼ 2, the CO2 conversion for O-MMTR, HMMTR, and C-MMTR reach respectively 5, 2.8, and 4.5%, while the MTR has the CO2 conversion of 11.5%. Therefore, the membrane reactors have lower CO2 conversion than MTR. Vice versa, under the same H2/CO ¼ 1.2, the CO2 conversion at the outlet of the MTR, O-MMTR, H-MMTR, and C-MMTR is found 11.5, 25, 29, and 16.7%, respectively, so membrane reactors show higher CO2 conversion than MTR. Generally, the results show that the membrane reactors which produce syngas for DME direct synthesis process have higher CO2 conversion than those used for methanol production and Fischer-Tropsch process.

from tri-reforming of methane was carried out under the same methane conversion (i.e., 94%), the same H2/CO ratio (i.e., 2), and the same H2/CO ratio (i.e., 1.2). A one-dimensional heterogeneous model was developed to investigate the performance of TRM in terms of methane conversion, hydrogen yield, H2/CO ratio, and CO2 conversion at the base case conditions. Finally, the advantages and disadvantages of these strategies are described as follows: (i) The MTR requires a longer catalyst length to achieve the same methane conversion (i.e., 94%) compared to other configurations, (ii) No hot spot forms in the O-MMTR, therefore the catalyst lifetime is expected to increase (iii) In the O-MMTR, the minimum CH4 conversion (i.e., 79.3%), is achieved under the same H2/CO ¼ 2, and the maximum conversion (i.e., 100%) is obtained under the same H2/CO ¼ 1.2 compared to other strategies (iv) In the H-MMTR, the minimum CO2 conversion (i.e., 2.8%), is obtained under the same H2/CO ¼ 2, and the maximum CO2 conversion (i.e., 29%) is achieved under the same H2/CO ¼ 1.2, and (v) The highest and lowest hot spot temperatures form in C-MMTR and MTR, respectively. Generally, the results show that use of O2, CO2 or H2O sidefeeding strategies could be feasible, beneficial, and flexible in terms of change in membrane thickness and shell-side pressure for syngas production with H2/CO ¼ 2 which is proper for methanol and Fischer-Tropsch process, and ¼ 1.2 which is suitable for DME direct synthesis. However, the MTR is only appropriate for producing the synthesis gas with H2 =COy2 under the base case conditions. In addition, the O-MMTR, HMMTR, and C-MMTR have higher CO2 conversion based on the H2/CO ¼ 1.2; so these strategies are more environmentally friendly for DME direct synthesis process compared to the MTR. The authors are looking to apply a proper optimization algorithm for methane tri-reforming based on the sidefeeding strategies in future work.

Nomenclature Ac av cp Ct dp Do Di Dei Dei;k Dim Ej F hf hi ho

Conclusion A comparison between three types of side-feeding strategy and multi-tubular fixed-bed reactor for syngas production

J Kg kg kj

cross-section area of each tube, m2 specific surface area of catalyst pellet, m2 m3 specific heat capacity, J mol1 K1 total concentration, mol m3 particle diameter, m outside diameter of the tube, m inside diameter of the tube, m diffusion coefficient of component i, m2 s1 Knudsen diffusion coefficient of component i, m2 s1 molecular diffusion coefficient of component i in a gas mixture, m2 s1 activation energy of reaction j, j mol1 molar flow rate, mol s1 gas-catalyst heat transfer coefficient, W m2 k1 gas-reactor wall heat transfer coefficient in tube side, W m2 k1 gas-reactor wall heat transfer coefficient in shell side, W m2 k1 permeation rate through the membrane, mol m2 s1 gas thermal conductivity, W m1 k1 Gas-solid mass transfer coefficient, m s1 reaction rate constant, mol kg1 s1

Please cite this article as: Alipour-Dehkordi A, Khademi MH, Use of a micro-porous membrane multi-tubular fixed-bed reactor for trireforming of methane to syngas: CO2, H2O or O2 side-feeding, International Journal of Hydrogen Energy, https://doi.org/10.1016/ j.ijhydene.2019.10.097

international journal of hydrogen energy xxx (xxxx) xxx

Ki Kn Kw M N P rj rp R Re Sc T ug U nc ni;j y z

adsorption equilibrium constant, bar1 reaction equilibrium constant thermal conductivity of reactor wall, W m1 k1 molecular weight, g mol1 number of tubes pressure, bar rate of reaction j, mol kg1 s1 membrane pore radius, m universal gas constant, j mol1 k1 Reynolds number Schmidt number temperature, K velocity of the fluid phase, m s1 overall heat transfer coefficient, W m2 k1 molecular diffusion volume, cm3 mol1 stoichiometric coefficient of component i in reaction j mole fraction, mol mol1 axial coordinate, m

Greek letters adsorption enthalpy, j mol1 DHi DHR enthalpy of reaction, j mol1 r density, kg m3 d membrane thickness, m ε porosity h effectiveness factor m viscosity, kg m1 s1 t tortuosity Superscripts t tube side s shell side Subscripts 0 inlet condition B catalyst bed i chemical species: CO2, CO, H2O, H2, CH4, O2, and N2 in reactor inlet j reaction number m membrane mix gas mixture s solid phase

references

[1] Matar M, Mirbach MJ, Tayim HA. Catalysis in petrochemical processes. Kluwer Academic Publishers; 1988. €  lu S‚. Thermodynamic equilibrium analysis [2] Ozkara-Aydıno g of combined carbon dioxide reforming with steam reforming of methane to synthesis gas. Int J Hydrogen Energy 2010;35:12821e8. [3] Aramouni NAK, Touma JG, Tarboush BA, Zeaiter J, Ahmad MN. Catalyst design for dry reforming of methane: analysis review. Renew Sustain Energy Rev 2018;82:2570e85. € _ Heterogeneous reactor modeling [4] Avci A, Trimm D, Onsan ZI. for simulation of catalytic oxidation and steam reforming of methane. Chem Eng Sci 2001;56:641e9. [5] Sadeghi M, Jafari M, Yari M, Mahmoudi S. Exergoeconomic assessment and optimization of a syngas production system with a desired H2/CO ratio based on methane tri-reforming process. J CO2 Util 2018;25:283e301.

13

[6] Cho W, Song T, Mitsos A, McKinnon JT, Ko GH, Tolsma JE, et al. Optimal design and operation of a natural gas trireforming reactor for DME synthesis. Catal Today 2009;139:261e7. [7] Song C. Tri-reforming: a new process for reducing CO2 emissions. Chem Innovat 2001;31:21e6. [8] Rahimpour M, Aboosadi ZA, Jahanmiri A. Synthesis gas production in a novel hydrogen and oxygen perm-selective membranes tri-reformer for methanol production. J Nat Gas Sci Eng 2012;9:149e59. [9] Garcı´a-Vargas JM, Valverde JL, de Lucas-Consuegra A,  mez-Monedero B, Dorado F, Sa  nchez P. Methane triGo reforming over a Ni/b-SiC-based catalyst: optimizing the feedstock composition. Int J Hydrogen Energy 2013;38:4524e32. [10] Chein R-Y, Wang C-Y, Yu C-T. Parametric study on catalytic tri-reforming of methane for syngas production. Energy 2017;118:1e17. [11] Izquierdo U, Barrio V, Requies J, Cambra J, Gu¨emez M, Arias P. Tri-reforming: a new biogas process for synthesis gas and hydrogen production. Int J Hydrogen Energy 2013;38:7623e31. [12] Majewski AJ, Wood J. Tri-reforming of methane over Ni@ SiO2 catalyst. Int J Hydrogen Energy 2014;39:12578e85.  nchez P. [13] Garcı´a-Vargas JM, Valverde JL, Dı´ez J, Dorado F, Sa Catalytic and kinetic analysis of the methane tri-reforming over a NieMg/b-SiC catalyst. Int J Hydrogen Energy 2015;40:8677e87. [14] Anchieta CG, Assaf EM, Assaf JM. Effect of ionic liquid in Ni/ ZrO2 catalysts applied to syngas production by methane trireforming. Int J Hydrogen Energy 2019;44:9316e27. [15] Song C, Pan W. Tri-reforming of methane: a novel concept for catalytic production of industrially useful synthesis gas with desired H2/CO ratios. Catal Today 2004;98:463e84. [16] Aboosadi ZA, Jahanmiri AH, Rahimpour MR. Optimization of tri-reformer reactor to produce synthesis gas for methanol production using differential evolution (DE) method. Appl Energy 2011;88:2691e701. [17] Khajeh S, Aboosadi ZA, Honarvar B. A comparative study between operability of fluidized-bed and fixed-bed reactors to produce synthesis gas through tri-reforming. J Nat Gas Sci Eng 2014;19:152e60. [18] Khajeh S, Aboosadi ZA, Honarvar B. Optimizing the fluidizedbed reactor for synthesis gas production by tri-reforming. Chem Eng Res Des 2015;94:407e16. [19] Farniaei M, Abbasi M, Rahnama H, Rahimpour MR, Shariati A. Syngas production in a novel methane dry reformer by utilizing of tri-reforming process for energy supplying: modeling and simulation. J Nat Gas Sci Eng 2014;20:132e46. [20] Manenti F, Pelosato R, Vallevi P, Leon-Garzon AR, Dotelli G, Vita A, et al. Biogas-fed solid oxide fuel cell (SOFC) coupled to tri-reforming process: modelling and simulation. Int J Hydrogen Energy 2015;40:14640e50. [21] Dı´ez-Ramı´rez J, Dorado F, Martı´nez-Valiente A, Garcı´a nchez P. Kinetic, energetic and exergetic Vargas J, Sa approach to the methane tri-reforming process. Int J Hydrogen Energy 2016;41:19339e48. [22] Zhang Y, Cruz J, Zhang S, Lou HH, Benson TJ. Process simulation and optimization of methanol production coupled to tri-reforming process. Int J Hydrogen Energy 2013;38:13617e30. [23] Minutillo M, Perna A. A novel approach for treatment of CO2 from fossil fired power plants. Part B: the energy suitability of integrated tri-reforming power plants (ITRPPs) for methanol production. Int J Hydrogen Energy 2010;35:7012e20. [24] Wiranarongkorn K, Authayanun S, Assabumrungrat S, Arpornwichanop A. Analysis of thermally coupling steam

Please cite this article as: Alipour-Dehkordi A, Khademi MH, Use of a micro-porous membrane multi-tubular fixed-bed reactor for trireforming of methane to syngas: CO2, H2O or O2 side-feeding, International Journal of Hydrogen Energy, https://doi.org/10.1016/ j.ijhydene.2019.10.097

14

[25]

[26]

[27]

[28]

[29]

[30]

[31]

[32]

[33]

international journal of hydrogen energy xxx (xxxx) xxx

and tri-reforming processes for the production of hydrogen from bio-oil. Int J Hydrogen Energy 2016;41:18370e9. Lu Y, Dixon AG, Moser WR, Ma YH. Analysis and optimization of cross-flow reactors with distributed reactant feed and product removal. Catal Today 1997;35:443e50. Rahnama H, Farniaei M, Abbasi M, Rahimpour MR. Modeling of synthesis gas and hydrogen production in a thermally coupling of steam and tri-reforming of methane with membranes. J Ind Eng Chem 2014;20:1779e92. Lu M, Wang G, Zhang Z, Duan X, Yuan W, Zhou X. Microporous inert membrane packed-bed reactor for propylene epoxidation with hydrogen and oxygen: modelling and simulation. Chem Eng Process: Process Intensification 2017;122:425e33. Kertalli E, Kosinov N, Schouten J, Nijhuis T. Direct synthesis of propylene oxide in a packed bed membrane reactor. Chem Eng J 2017;307:9e14. Xu J, Froment GF. Methane steam reforming, methanation and water-gas shift: I. Intrinsic kinetics. AIChE J 1989;35:88e96. Xiu G-h, Li P, Rodrigues AE. Sorption-enhanced reaction process with reactive regeneration. Chem Eng Sci 2002;57:3893e908. Trimm DL, Lam C-W. The combustion of methane on platinum-alumina fibre catalysts-I: kinetics and mechanism. Chem Eng Sci 1980;35:1405e13. De Smet C, De Croon M, Berger R, Marin G, Schouten J. Design of adiabatic fixed-bed reactors for the partial oxidation of methane to synthesis gas. Application to production of methanol and hydrogen-for-fuel-cells. Chem Eng Sci 2001;56:4849e61. Iordanidis AA. Mathematical modeling of catalytic fixed bed reactors. The Netherlands: Twente University Press Enschede; 2002.

[34] Nauman EB. Chemical reactor design, optimization, and scaleup. John Wiley & Sons; 2008. [35] De Groote AM, Froment GF. Simulation of the catalytic partial oxidation of methane to synthesis gas. Appl Catal Gen 1996;138:245e64. [36] Rodriguez ML, Ardissone DE, Heracleous E, Lemonidou AA,  pez E, Pedernera MN, et al. Oxidative dehydrogenation of Lo ethane to ethylene in a membrane reactor: a theoretical study. Catal Today 2010;157:303e9. [37] Cussler EL. Diffusion, mass transfer in fluid systems. Cambridge university press; 2009. [38] Fairbanks D, Wilke C. Diffusion coefficients in multicomponent gas mixtures. Ind Eng Chem 1950;42:471e5. [39] Poling BE, Prausnitz JM, O’connell JP. The properties of gases and liquids. New York: Mcgraw-hill; 2001. [40] McCabe WL, Smith JC, Harriott P. Unit operations of chemical engineering. New York: McGraw-hill; 1993. [41] Hussain A, Seidel-Morgenstern A, Tsotsas E. Heat and mass transfer in tubular ceramic membranes for membrane reactors. Int J Heat Mass Transf 2006;49:2239e53. [42] Saad JM, Williams PT. Manipulating the H2/CO ratio from dry reforming of simulated mixed waste plastics by the addition of steam. Fuel Process Technol 2017;156:331e8. [43] MUNRO M. Evaluated material properties for a sintered alpha-alumina. J Am Ceram Soc 1997;80:1919e28. [44] Uchytil P, Schramm O, Seidel-Morgenstern A. Influence of the transport direction on gas permeation in two-layer ceramic membranes. J Membr Sci 2000;170:215e24. [45] Maroto-Valer MM, Song C, Soong Y. Environmental challenges and greenhouse gas control for fossil fuel utilization in the 21st century. Springer Science & Business Media; 2012.

Please cite this article as: Alipour-Dehkordi A, Khademi MH, Use of a micro-porous membrane multi-tubular fixed-bed reactor for trireforming of methane to syngas: CO2, H2O or O2 side-feeding, International Journal of Hydrogen Energy, https://doi.org/10.1016/ j.ijhydene.2019.10.097