Enhancement of Hydrogen Production and Carbon Dioxide Capturing

Article pubs.acs.org/EF

Enhancement of Hydrogen Production and Carbon Dioxide Capturing in a Novel Methane Steam Reformer Coupled with Chemical Looping Combustion and Assisted by Hydrogen PermSelective Membranes Mohsen Abbasi,† Mahdi Farniaei,‡ Mohammad Reza Rahimpour,*,†,§ and Alireza Shariati† †

Department of Chemical Engineering, School of Chemical and Petroleum Engineering, Shiraz University, Shiraz 71345, Iran Department of Chemical Engineering, Shiraz University of Technology, Shiraz 71555-313, Iran § Department of Chemical Engineering and Materials Science, University of California, Davis, 1 Shields Avenue, Davis, California 95616, United States ‡

ABSTRACT: In this novel paper, application of chemical looping combustion (CLC) instead of furnace in a steam reformer assisted by Pd−Ag hydrogen perm-selective membranes (CLC-SRM) for CO2 capture and hydrogen production has been analyzed. NiO18-αAl2O3 particles have been employed as oxygen carriers in CLC-SRM. These particles have shown very high reactivity and allow for working at high temperatures in a CLC process with full methane conversion due to Ni-based oxygen carriers. In the CLC-SRM configuration, the air reactor (AR) and fuel reactor (FR) operate in fast and bubbling fluidization, respectively. In this configuration, reforming tubes are located vertically inside the AR so that methane steam reforming occurs in these fixed bed catalytic tubes that have been covered by the membranes. A steady state one-dimensional heterogeneous catalytic reaction model is applied to analyze the performance of CLC-SRM. Performance of conventional steam reformer (CSR) has been compared with CLC-SRM by investigation of important parameters such as temperature, mole fractions, heat of reaction, rate of reactions, methane conversion, and hydrogen production. The simulation results of CLC-SRM show that by employing CLC-SRM, methane conversion and hydrogen production increase 7.54% and 25.48%, respectively, in comparison with CSR. In addition, results indicated that by increasing feed flow rate of FR from 90 to 180 mol s−1 methane conversion and hydrogen production can increase 16.73% and 40%, respectively. In CLC-SRM, the total amount of methane consumed in the FR and combustion efficiency increases to 1 in the FR, and a huge amount of almost pure carbon dioxide (410 ton day−1) can be captured by removal of water from the FR outlet stream with condensation.

1. INTRODUCTION Because of a fossil energy shortage and global warming, hydrogen is expected to be nominated as an alternative energy source. There are many different methods and various kinds of sources that can produce hydrogen. Industrial production of hydrogen is mainly from the steam reforming of natural gas by production of syngas. Syngas, a mixture of carbon monoxide and hydrogen, plays a very important role in chemical industries and can be converted to bulk chemicals like methanol, dimethyl ether, aldehydes, hydrogen, and ammonia.1−6 A conventional steam reformer consists of vertical tubes packed with Ni-based catalyst located inside a huge furnace. The required heat for endothermic reforming reaction is provided by direct combustion of fossil fuels such as natural gas in the furnace.7−10 The burning of fossil fuels produces huge amount of carbon dioxide per year; carbon dioxide is one of the greenhouse gases that is one of the reasons for global warming and climate change.11−13 Chemical looping combustion (CLC) is a novel combustion technology to decrease the greenhouse gas emissions affecting global warming. In this process, separation of CO2 involves the use of circulating solid oxygen carriers in two interconnected reactors, which transfer oxygen from the air reactor (AR) to the fuel reactor (FR) avoiding the direct contact between air and fuel. The exit gas from the fuel reactor contains only CO2 and H2O, © XXXX American Chemical Society

and after water condensation, almost pure CO2 can be obtained. The oxygen-carrier is regenerated with air in the AR, and it is ready to start a new cycle in FR. The heat involved in the global process is the same as for normal fuel gas combustion.14−17 The concept of CLC was presented for the first time in 1954 by Lewis and Gilliland as a means of CO2 production.18 Ishida et al. recognized the use of CLC as a CO2 capture technology.19 Lyngfelt et al. proposed a CLC reactor design based on circulating fluidized bed (CFB) technology.20 This most common design of a CLC includes a fast fluidized bed riser for the AR and a bubbling fluidized bed for the FR, with the oxygen-carrier in the form of metal oxide particles circulating between them. The research on chemical-looping combustion has been summarized in a number of reviews.21−24 Also CLC with gaseous fuels has been widely investigated.25−28 Few works have been presented in the literature for application of CLC in industrial scales. Lyngfelt et al. calculated basic design parameters of an industrial scale 10 MWth CLC boiler.20 Received: May 13, 2013 Revised: August 15, 2013

A

dx.doi.org/10.1021/ef400880q | Energy Fuels XXXX, XXX, XXX−XXX

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The feasibility of CLC in two interconnected fluidized bed reactors for a capacity of 800 MWth has been analyzed by Wolf et al.29 They compared nickel- and iron-based oxygen carriers in CLC with relation to the system’s electrical and energy efficiencies. Their results indicated that two oxygen carriers had the same system efficiency, but application of CLC with nickelbased oxygen carriers were easier than iron-based oxygen carriers. Rydén and Lyngfelt found that hydrogen production by steam reforming of natural gas with inherent capture of CO2 by CLC should be feasible.30 Marx et al. proposed a basic design for a semicommercial 10 MWth dual circulating fluidized bed type CLC boiler plant for power production based on continuous development of the CLC technology from the laboratory scale to pilot scale.31 Ortiz et al. analyzed the performance of an iron waste material as the oxygen carrier burning a PSA tail gas for a coupling steam reforming of methane in a 500 Wth pilot plant.32 They estimated that a solids inventory of 1600 kg MWth−1 would be necessary to fully convert the PSA off-gas to CO2 and H2O. More recently, the same group found that the solid inventory was considerably lower when using an impregnated Fe2O3/Al2O3 oxygen carrier.33 A new method for production of H2 with CO2 capturing is chemical looping reforming (CLR). CLR is similar to CLC, but complete oxidation of the CH4 is prevented by using a low air to fuel ratio. Hence CLR can be described as a process for partial oxidation of CH4 that it utilizing CLC as a source of undiluted oxygen. In FR of CLR, partial oxidation of CH4 and SR are carried out together and syngas is produced at the outlet of FR.34 CLC-SR is different with CLR. Comparison of CLR with CLC-SR indicates that CLR operates in two sides (AR and FR), but CLC-SR operates in three sides (AR, FR, and SR). Therefore, the cost of materials in CLR is lower than CLC-SR. Another advantage of CLR is that the heat needed for SR is supplied without costly oxygen production. Efficiency of CLR at low pressures is low, and the concept of pressurized CLR has some drawbacks as well. Although in CLC-SR, the CH4 reaction with oxygen carriers in FR can be operated at low pressure while the SR reactions still takes place at elevated pressure in tubes.33 Finally, the change of a conventional steam reformer (CSR) to CLC-SR is easier than CLR, because in CLC-SR, CLC is employed instead of a furnace and operating conditions of SR do not change, but in CLR, operating conditions of SR need to be changed. The concepts of thermally coupled reactors are considered by many researchers.35−39 Rahimpour et al. investigated the application of a thermally coupled reactor in different processes in a review.40 Hydrogen perm-selective membrane systems are applied efficiently in a variety of industrial processes which involve separation of hydrogen from feed streams produced from gasification, reforming, or petrochemical processes. Palladium based membranes have been widely considered as the most suitable membranes for hydrogen production because of their high permeability and good surface properties and because palladium is 100% selective for hydrogen transport.41 The enhancement of the product yield and thermodynamic shifting of the reactions toward the products are the certain superiorities of membranes employed in reactors compared with conventional reactors. A membrane reformer is introduced as a new type of reactor for the purpose of both separation and

production of hydrogen in only one unit.42,43 In this novel study, application of CLC instead of a furnace in a steam reformer assisted by Pd−Ag hydrogen perm-selective membranes (CLC-SRM) for CO2 capture and hydrogen production has been analyzed. NiO18-αAl2O3 particles have been employed as oxygen carriers in CLC-SRM due to Ni-based oxygen carriers have shown very high reactivity and allow for working at high temperatures in a CLC process with full CH4 conversion.44 NiO18-αAl2O3 oxygen carrier has been used in a continuous CLC prototype with a circulating fluidized bed reactor, consisting of two interconnected fluidized bed reactors.28 Additionally, kinetics of this oxygen carrier for reduction with CH4 and oxidation with air have been presented in the literature.45 Rigorous mathematical models have been employed because they are excellent tools for investigation of the basic characteristics of such novel configurations, time saving and reducing the operating costs during the expensive stage of pilot plant development. One-dimensional heterogeneous catalytic reaction mathematical models have been employed for steady state simulation. Performance of CSR is compared with CLC-SRM by investigation of important parameters such as temperature, mole fraction, heat, rate, methane conversion, and hydrogen yield. For better presentation of membranes effect, results of CLC-SRM are compared with coupling of CLC and SR without Pd−Ag hydrogen perm-selective membranes (CLC-SR) with similar reactor design and operating conditions to CLC-SRM.

2. PROCESS DESCRIPTION 2.1. Conventional Steam Reformer (CSR). Figure 1 describes the schematic diagram of CSR for production of syngas in Zagros

Figure 1. Schematic diagram of CSR. Petrochemical Company, Assaloyeh, Iran.46 CSR has a huge top fired furnace with high thermal energy production equal to 69 MWth. In the 184 vertical tubes of CSR, steam reforming of methane occurs over the packed Ni-based catalyst. The properties of CSR have been represented in Table 1. 2.2. Chemical Looping Combustion Coupled with Steam Reformer Assisted by Membranes (CLC-SRM). Figure 2 demonstrates the schematic diagram of the CLC-SRM. Steam reformer characterization is similar to CSR, and only tubular H2 perm-selective membranes have been inserted in tubes. Also CLC has been used instead of a furnace. B


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Table 1. Design Characteristics and Input Data of CSR46

in the SR side, CLC-SRM converts to CLC-SR. The high gas velocity in the AR provides the necessary driving force to obtain a circulation of Ni based oxygen carriers between the AR and FR. Oxidized carriers are collected in the cyclone and led to the FR. Reduced oxygen carriers are transferred from FR back to the AR by using gravity force. Gas leakage between the reactors is prevented by particle-gas locks fluidized with steam.31 Design characteristics of AR and FR such as reactor dimensions, inventory of oxygen carriers, and solid circulation rate are based on experimental results of using this NiO18-αAl2O3 oxygen carrier in the continuous CLC Prototype.28,26,45 Experimental results indicated that for NiO18-αAl2O3 as the oxygen carrier in the continuous CLC Prototype, a solids inventory of about 630 kg MWth−1 was needed for FR. Also on the basis of experiment results, a similar solid inventory is selected for NiO18-αAl2O3 oxygen carrier in AR.28 On the basis of thermal energy production equal to 69 MWth of CLC, solid inventory in AR and FR can be determined as shown in Table 2. Solid circulation rate is calculated based on an oxygen carrier/fuel ratio equal to 2 to reach combustion efficiencies at the maximum values. Limitation of solid circulation rate or the transport capacity of the riser (AR) is the main subject for design of AR and FR. The maximum circulation rate feasible in a CLC plant without increased costs and with commercial experience is 80 kg m−2 s−1.26 AR and FR have been designed based on limitation of the solid circulation rate, sufficient heat transfer from AR to FR, and enough space for loading of NiO18-αAl2O3 oxygen carrier in the reactors. Table 2 shows the specific properties and operational conditions of the CLC-SRM. It must be noted that in the output stream of the SR side of CLC-SRM, there is unconsumed CH4 and the produced gas is not pure synthesis gas. This produced gas can be used for methanol production.

Tube Side inlet temperature [°C] inlet pressure [bar a] total feed gas flow [kmol h−1] number of tubes inside diameter [mm] heated length [m] catalyst volume filled in (total) [m3] design pressure [bar g] design temp [°C] particle size [mm] void fraction [−] heat load on tube (100% design case) [kcal m−2 h−1] reformer duty (100% design case) [GJ h−1] net catalyst shape feed composition (mol %) CO2 CO H2 CH4 N2 H2O Shell Side combustion air

feed gas (fuel)

temperature [°C] pressure [bar] flow rate [sm3 h−1] temperature [°C] pressure [bar] flow rate [sm3 h−1]

520 40 9129.6 184 125 12 27.8 41 790 19 × 16 0.4 68 730 248.2 10-HOLE rings 1.72 0.02 5.89 32.59 1.52 58.26 330 1 114 313 34 3 29 608

3. REACTION SCHEME AND KINETICS 3.1. Steam Reforming of Methane (SR). The SR and water gas shift reactions occur in the steam reformer tubes as follows:

AR and FR operate in fast fluidization and bubbling fluidization regimes. In this novel arrangement, reforming tubes with inserted membranes are located vertically inside the AR and FR is covered in the AR. In the CLC-SRM, heat generated in the AR transfers continuously to the FR and SR sides. By removing Pd−Ag membranes

° = 206.3 kJ mol−1 CH4 + H 2O ↔ CO + 3H 2 ΔH298 (1)

Figure 2. Schematic diagram of CLC-SRM. C


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Table 2. Design Characteristics and Input Data of CLC-SRM

φ = 1 + K COpCO + K H2pH + K CH4pCH + K H2O 2

FR inlet temperature [K] inlet pressure [bar] total feed gas flow [mol s−1] inside diameter [m] heated length [m] particle size [mm] particle density [kg m−3] terminal velocity of particles [m s−1] minimum fluidization velocity [m s−1] void fraction [−] NiO circulation rate [mol s−1] solid inventory [kg] feed velocity [m s−1] AR

Koi [bar CH4 CO H2 H2O

(8)

Also in FR, the endothermic reduction of NiO18-αAl2O3 oxygen carriers with CH4 carries out as follows: ° = 156.5 kJ mol−1 CH4 + 4NiO ΔH298 → 4Ni + CO2 + 2H 2O (9)

The kinetic modeling over Ni-based catalyst in this work is based on the model of Xu and Froment as follows:47

pH p ⎞ ⎛ 2 CO 1 2⎟ ⎜p p − × 2 ⎜ CO H2O ⎟ KIII ⎠ φ ⎝ 2

−38 280 −70 650 −82 900 88 680

10 10−5 10−9 105 bar

° = −479.4 kJ mol−1 O2 + 2Ni → 2NiO ΔH298

(3)

k3 pH

× × × ×

ΔHi [J mol−1]

]

3.2. Oxidation and Reduction of Oxygen Carriers in CLC. For the design of a CLC process, oxidation and reduction rate of oxygen carriers is an important factor to be considered because it is directly related with the solids inventory in the AR and FR. The oxygen carrier particles should have high reactivity to fully convert CH4 in the FR, and to be regenerated in the AR.26 The exothermic oxidation reaction of Ni-based oxygen carriers inside AR is took place as follows:

° = − 41.1 kJ mol−1 CO + H 2O ↔ CO2 + H 2 ΔH298

r3 =

6.65 8.23 6.12 1.77

−1 −4

⎛ −ΔH ⎞ ⎟ K i = K 0i × exp⎜ ⎝ RT ⎠

(2)

⎛ pH 4 pCO ⎞ k2 ⎜ 1 2.5 2 2⎟ p p − × KII ⎟⎠ φ 2 pH 3.5 ⎜⎝ CH4 H2O 2

2

Table 4. Van’t Hoff Parameters for Species Adsorption48

° = 164.9 kJ mol−1 CH4 + 2H 2O ↔ CO2 + 4H 2 ΔH298

r2 =

2

pH

The reaction equilibrium constants and Arrhenius kinetic parameters are listed in Table 3. Also Van’t Hoff parameters for species adsorption are represented in Table 4.48

818 20 1296 20% 2.65 12 2 0.52 864 43 500 7

⎛ pH 3.5 pCO ⎞ k1 ⎜ 2 ⎟× 1 r1 = p p − 2.5 ⎜ CH4 H 2O ⎟ φ2 K pH ⎝ I ⎠ 2

4

(7)

818 10 108 3.75 12 0.2 2400 0.94 0.022 0.52 864 43 500 0.6

inlet temperature [K] inlet pressure [bar] air flow [mol s−1] excess air inside diameter [m] heated length [m] particle size [mm] void fraction [−] NiO circulation rate [mol s−1] solid inventory [kg] feed velocity [m s−1]

pH O

Changing the grain size model is able to predict the reaction kinetics of NiO18-αAl2O3 oxygen carriers. In this model, particles are assumed to be formed by small spherical grains, each one following a shrinking core model during the reaction. To determine oxidation and reduction rate of oxygen carriers the following equations are used:

(4)

rs =

dX s 1 = ; τi dt

for oxidation: Xs = Xox =

m − mred mox − mred (10)

(5)

for reduction: Xs = Xred = 1 − Xox 1 τi = kiCg n

(6)

(11)

Table 3. Reaction Equilibrium Constants and Arrhenius Kinetic Parameters47 reaction, j

koj [mol (kg cat. s)−1]

equilibrium constant, Kj

1

⎞ ⎛ 26830 KI = exp⎜ + 30 114⎟(bar 2) ⎠ ⎝ Ts

1.17 × 10 bar

2

KII = KI × KIII(bar 2)

2.83 × 1014 bar0.5

3

⎛ 4400 ⎞ KIII = exp⎜ + 4.036⎟ ⎝ Ts ⎠

5.43 × 105 bar

15

Ej [J mol−1]

0.5

240 100

243 900

−1

67 130

⎛ − Ej ⎞ kj = k 0j × exp⎜ ⎟ ⎝ RT ⎠ D


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ki = koi e(−Ea / RT )

(12)

Properties and kinetic parameters for the oxidation and reduction of NiO18-αAl2O3 oxygen-carriers are presented in Table 5. Table 5. Properties and Kinetic Parameters for the Oxidation and Reduction of NiO18-αAl2O3 OxygenCarriers45 Properties NiO content (wt %) apparent density (g cm−3) BET surface area (m2 g−1) porosity (%) chemical reaction rate n k0 (mol1−n m3n−2 s−1)

18 2.5 7 43 Ea (kJ mol−1)

0.2 0.7

5 22

reduction with CH4 oxidation with O2

0.2 0.84

4. MATHEMATICAL MODEL The one-dimensional heterogeneous catalytic reaction mathematical model is based on the following assumptions: (1) The model is investigated at steady-state conditions. (2) Gas phases are ideal. (3) Plug flow mode is employed in each side of the reactors. (4) Axial diffusions of heat and mass are negligible. (5) Bed porosity in axial and radial directions is considered to be constant. (6) Heat loss to surroundings is neglected. Figure 3 presents the differential element along the axial direction inside the CLC-SRM. The differential equations describing mole and energy balances in CSR and CLC-SRM are present in sections 4.1 and 4.2. 4.1. Mole and Energy Balance in CSR. 4.1.1. Solid Phase. The mass and energy balance equations for the solid phase in the fixed bed tubes of CSR are expressed by a v cjkgi(yig − yis ) + ηriρb = 0

Figure 3. (a) Differential element along the axial direction inside the sides of CLC-SRM in the dense bed and (b) differential element along the axial direction inside the sides of CLC-SRM in the freeboard zone.

(13)

N g

JH =

s

a v hf (T − T ) + ρb ∑ ηri(ΔHf, i) = 0 i=1

2

(14)

πDj −Ft g dT q =0 Cp , j + a v hf (T js − T jg) + dz Ac w Ac

(16)

− PHTube 2

PHShell ) 2

(17)

E H2 = 15.7 kJ mol−1

The mass and energy balances for a solid and fluid of SR in the CLC-SRM configuration including membranes are as follows: −

(15)

RT

Q 0 = 1.65 × 10−5mol m−1 s−1 kPa−0.5,

where Ts and ysi are temperature and the mole fraction of the icomponent in the solid-phase inside the tubes, respectively. Also η is the effectiveness factor, which is defined as the ratio of the observed reaction rate to the real reaction rate.10 4.1.2. Fluid Phase. The following mass and energy balance equations are used for the fluid phase in CSR: 1 dFt, j − + a v cjkgi , j(yis, j − yig, j ) = 0 A c dz

δm

E H2

( )(

Q 0A exp

JH 1 dFt, j + a v cjkgi , j(yis, j − yig, j ) + ξ 2 = 0 A c dz Ac

(18)

πDj −Ft g dT g ) UAS(TAR − TSR Cp , j + a v hf (T js − T jg) + dz Ac Ac JH T2 πDj g g )+ 2 USM(TSR Cpg dT = 0 − − TM Ac Ac T1 (19)

∫

where T and yi are temperature and the mole fraction of the icomponent in the fluid-phase inside the tubes, respectively. 4.2. Mole and Energy Balance in CLC- SRM. 4.2.1. Mole and Energy Balance in SR Side. For driving mass and energy balance of SR in the CLC-SRM configuration, hydrogen flux with membranes should be included. For estimating the hydrogen flux in membranes, eq 17 can be used, which in this equation the hydrogen flux is a function of diffusivity, membrane thickness, and hydrogen partial pressure.41−43

ξ is equal to 1 for hydrogen and zero for other gases. 4.2.2. Mole and Energy Balance in CLC. AR and FR are operated in the fluidized bed regime, and the fluid phase in these reactors consist of emulsion and bubble phases. The mathematical simulation for modeling of these reactors is developed based on the following assumptions: (1) The dense catalyst bed is considered to be composed of bubble phase and E


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In the dense bed, a gas exchange between the bubbles and emulsion is considered allowing the exchange of products and reactants between these phases by diffusive and bulk flow mechanism. The mole balances for dense bed zone in FR are as follows:

emulsion phase. (2) Because of rapid mixing, the operating condition is assumed to be isothermal, i.e., bubble and emulsion phases have the same temperature. (3) The bubble phase follows the plug flow pattern. (4) The rising velocity of the bubble is constant and equal to the average value. (5) Gas behavior is ideal. (6) Bubbles are spherical with a constant size equal to the average value. (7) The bubble phase contains some catalyst particles, which are involved in reactions but the extent of reaction in the bubble phase is much less than the emulsion phase. (8) In bubbling and fast fluidized beds, because of the good solid mixing, the bed temperature variation in the radial direction is negligible. Therefore the one-dimensional heterogeneous reaction model is reasonable. 4.2.2.1. Mole and Energy Balance in AR. AR operates as a riser with 12 m length for transportation of oxidized oxygen carriers to FR. A valve at the outlet of AR controls the outlet solid circulation rate of oxygen carriers. Because of high velocity in AR, bubbles may be not be recognized clearly. The solid phase exists mainly in the form of disperse particles and clusters. In fact, the AR is divided in the bottom bed (dense bed with bubble and emulsion phases), splash zone (with clusters and dilute phases), and transport zone (with dilute phase, and core annulus structure). However, in this research for simplification of the AR simulation, the plug flow model of gases and bubble and emulsion phases along all AR height have been assumed. According to the above assumptions, the bubble and emulsion phase mole balance equations for an element of length dz are formulated as follows: for emulsion phase: −

for emulsion phase: −

for bubble phase: b −δ dFi + δKbeC tab(yie − yi b ) + aδγηρs (rCH4b) = 0 A c dz

−

−

−

(20)

(27)

1 dFt + ρε r =0 s s CH4 A c dz

(28)

Ft g dT πDi Cp + ρε r ΔHred + UAF(TAR − TFR ) = 0 s s CH4 Ac dz Ac (29)

4.3. Pressure Drop, Boundary Conditions, and Auxiliary Correlations. The Ergun momentum balance equation is used to give the pressure drop along the reactors axes.

(21)

where Kbei is the mass transfer coefficient between the bubble phase and emulsion phases, yie and yib are the emulsion phase and bubble phase mole fractions, respectively, and γ is the volume fraction of the catalyst bed occupied by solid particles in the bubble phase. Fib and Fie are given as follows: Fib = yi b Ft

(22)

Fie = yie Ft

(23)

(1 − ε)ug 2 (1 − ε)2 dP = 150 3 2 + 1.75 dz ε dp ε 3d p2

(30)

The boundary conditions applied to solve the model are described in eq 31. At the entrance of the reactor, the inlet temperature, pressure, and gas compositions of the reactant gas in CSR and CLC-SR are known. Therefore, the following boundary conditions are used: at z = 0:

The energy balance for the fluid phase has been described below:

yi = yi0 ,

T = T0 ,

P = P0

(31)

To complete the simulation, auxiliary correlations should be added to the model. In the heterogeneous model, because of transport phenomena, the correlation estimation of heat and mass transfer between two phases and overall heat transfer coefficient between steam reformer walls and the gas phase in the AR should be considered. The correlations used for physical properties, mass and heat transfer coefficient, and the hydrodynamic concept of fluidization are summarized in Tables 6 and 7. 4.4. Numerical Solution and Model Validation. The developed model contains the set of ordinary differential equations (ODEs) of mass and energy balances and the

Ft g dT Cp + (1 − δ)ρe ηa(ro2e)(ΔHox ) + δρe Ac dz πDi πDi g UAS(TAR − TSR )− ηa(ro2b)(ΔHox ) − Ac Ac UAF(TAR − TFR ) = 0

Ft g dT Cp + (1 − δ)ρe ηa(rCH4e)(ΔHred) + δρe dz Ac πDi UAF(TAR − TFR ) = 0 ηa(rCH4b)(ΔHred) + Ac

Freeboard Zone Above the dense bed starts the freeboard zone, characterized by a strong decay on solids concentration. Mass and energy balance for this zone have been demonstrated by these equations:

for bubble phase:

−

(26)

The following energy balance equation is used for the fluid phase in the dense bed zone of FR:

1−δ + δKbeC tab(yi b − yie ) + a(1 − δ)ηρe (ro2e) A c dz

b −δ dFi + δKbeC tab(yie − yi b ) + aδγηρs (ro2s) = 0 A c dz

(25)

=0

dFie

=0

e 1 − δ dFi + δKbeC tab(yi b − yie ) + a(1 − δ)ηρe (rCH4e) A c dz

(24)

4.2.2.2. Mole and Energy Balance in FR. FR operates in the bubbling fluidized bed regime and consists of dense and freeboard zones.Dense Bed Zone F


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Table 6. Hydrodynamic Parameter49−55 parameter superficial velocity at minimum fluidization

imation is applied to solve the set of ODEs. Therefore, the ODEs are changed into a set of nonlinear algebraic equations. The length of each reactor is divided to 1000 separate segments, and the Gauss−Newton method is used to solve the obtained set of nonlinear algebraic equations in each segment for both sides simultaneously. This procedure is repeated for all nodes in the reactor so that the results of each node are used as the inlet conditions for the next node. Additionally, the validity of the present one-dimensional heterogeneous model is considered by comparing the simulated results with the industrial data of the tubular fixed-bed reactor. Table 8 represents comparison between the model prediction and the plant data. As indicated in this table, the predicted results have a good agreement with the industrial plant data.

equation

umf

⎞ ⎛ μ ⎟ 2 0.5 = ⎜⎜ ⎟[((27.2 + 0.0408Ar ) − 27.2)] ⎝ ρg d p ⎠

Archimedes number

Ar =

d p3ρg (ρp − ρg )g μ2

⎛ − 0.3z ⎞ ⎟ db = dbm − (dbm − db0) exp⎜ ⎝ D ⎠ bubble diameter

⎡ πD 2 ⎤0.4 dbm = 0.65⎢ (u0 − umf )⎥ ⎣ 4 ⎦

db0 = 0.376(u0 − umf )2 mass transfer coefficient (bubble-emulsion phase)

Kbe =

bubble rising velocity

ub = u0 − umf + 0.711 gd p

specific surface area for bubble

δ=

volume fraction of bubble phase to overall bed

6δ ab = db

density for emulsion phase

ρe = ρp (1 − εmf )

terminal velocity of particles

⎛ 4Dj mεmf ub ⎞0.5 umf +⎜ ⎟ 3 ⎝ πd b ⎠

Table 8. Agreement between Model Prediction and Plant Data46

(u0 − umf ) ub

0.5 ⎡ 4gd ⎛ ρ ⎞⎤ p ⎜ p ⎥ ⎢ ⎟ ut = π − 1⎟ ⎥ ⎢ 3C D ⎜ ρ ⎠⎦ ⎝ g ⎣

CD =

methane conversion % composition (mol %) CO2 CO H2 CH4 N2 H2O

26.5

26

5.71 3.15 31.39 20.41 1.29 38.05

5.72 3.19 31.53 20.33 1.30 37.94

FCH4,in

equations

Cp = a + bT + cT 2 + dT −2

component viscosity

C1T

μ=

1+ mixture heat capacity, viscosity of reaction mixtures, and mixture thermal conductivity

C2 T

based on local compositions

ηc =

Re =

Sci = Di m =

Dij =

−0.42

Sci−0.67ug

× 10

3

2

2

(4(yCH )in Fin)

μ

μ ρDim × 10−4 1 − yi

yi

Dij

10−7T 3/2 1/Mi + 1/Mj

ϕ=

P(vci 3/2 + vcj 3/2)

heat transfer coefficient between gas phase and reactor wall

0.407 2/3⎞ ⎛ h ⎜⎛ Cpμ ⎞ ⎟ 0.458 ⎛⎜ μ ⎞⎟ = ⎟ ⎜ Cpρμ ⎜⎝⎝ K ⎠ ⎟⎠ εB ⎜⎝ ρud p ⎟⎠

overall heat transfer coefficient

A ln(Do /Di) A 1 1 1 = + i + i U hi 2πLK w Ao ho

(34)

The yield of hydrogen is expressed as the number of moles of hydrogen formed from 1 mole of methane that flows in the reactor and generally defined from eq 33.10 Fin and Fout are the molar flows at the inlet and outlet of the reactors, respectively, and yi is the mole fraction of gas i. In addition, the ratio of oxygen-carrier circulation flow rate to fuel flow rate (ϕ) is an effective parameter on simulation results and is defined by eq 35:

2R pugρ

∑i ≠ j

(33)

(2yCO + yH O )out Fout 4

kgi = 1.17Re

FCH4,in

C4 T2

(32)

FH2,out − FH2,in

hydrogen yield =

C2

+

CSR

FCH4,in − FCH4,out

methane conversion =

definitions

mass transfer coefficient between gas and solid phases

plant data

To evaluate CLC-SRM performance, some operating variable such as methane conversions, hydrogen yield, and combustion efficiency (ηc) are investigated as follows:

10 Re

Table 7. Auxiliary Correlations49−55 component heat capacity

parameter

FNiO 4FCH4

(35)

where FNiO and FCH4 are the molar flow rates of the NiO and CH4 in the FR, respectively. A value of ϕ = 1 corresponds to the stoichiometric NiO amount needed for a full conversion of the CH4 to CO2 and H2O.14

5. RESULTS AND DISCUSSION 5.1. Thermal Behavior. Figure 4 presents the temperature pattern of CLC-SRM in the four sides. Thermal driving force is created by a temperature difference between exothermic and

nonlinear algebraic equations of the kinetic model, auxiliary and hydrodynamic correlations. Backward finite difference approxG


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Figure 4. Temperature pattern in the four sides of CLC-SRM.

Figure 5. Comparison of SR temperature among CSR, CLC-SR, and CLC-SRM.

endothermic sides. As shown in this figure, AR temperature increases in the reactor entrance due to occurring oxidation of oxygen carriers, but in the second half, temperature decreases due to sensible heat transfer from AR to FR and SR. Important temperature variations along the first half of AR and FR are due to rapid oxidation and reduction of oxygen carriers and a large amount of heat that was generated and consumed, respectively. The outlet temperature of four sides of CLC-SRM have a very low difference because the overall heat transfer coefficient is large due to turbulence of gases and good solid mixing in the fluidized regime of AR and FR. Figure 5 demonstrates the comparison of SR temperature between CSR, CLC-SR, and CLC-SRM. Results show that temperature of SR along the reactor axes in CLC-SR is higher than CSR, and at the end of the reactor they are equal. This subject gives correct design and operating parameters of CLCSR configuration for preparation of enough heat for SR reactions instead of the furnace in CSR. In CLC-SRM, the temperature increase in the first part of the reactor is mainly due to the improved heat supply by the CLC process, and the reduction of temperature in the rest of reactor is due to permeation flux of hydrogen in the membrane side and consuming more heat due to a shift of the endothermic equilibrium reactions.

Figure 6. (a) Heat generation in the AR and heat consumption in the FR and SR and (b) heat transfer from the AR to the FR and SR in CLC-SRM.

For better presentation of thermal behavior, Figure 6 demonstrates heat generation and consumption and also heat transfer from AR to FR and SR in CLC-SRM. As shown in Figure 6a,b, the heat generation of AR is larger than heat consumption in SR and FR and heat transfer from AR to FR and SR. At the reactor entrance, the generated heat is less than H


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Figure 7. (a) Rate of oxidation in the AR side and reduction in the FR side of CLC-SRM. (b) Rate of syngas production in the SR side of CLC-SRM.

Figure 8. (a) Variation of component mole fractions along reactor axes in the AR of CLC-SRM. (b) Variation of solid conversion along reactor axes in the AR of CLC-SRM.

the consumed heat causing the temperature of FR and SR to fall and a cold spot developed after that condition is changed so that an increase along the rest of the reactor length occurs (see Figure 4). 5.2. Rate Behavior. Figure 7 presents oxidation and reduction rates of oxygen carriers in AR and FR plus SR reaction rates in CLC-SRM. As shown in this figure, oxygen carriers have been fully oxidized and reduced in the first half of the reactor because the rate of oxidation and reduction decreases to zero (this subject is clearly presented in next part). In the SR side, the rate of CO and CO2 production (reactions 1 and 2) is higher than the water gas shift reaction. The water gas shift reaction is an exothermic, equilibriumlimited reaction that exhibits decreasing conversion with increasing temperature. This reaction takes place in the temperature range of 583−723 K. For the SR side in CLCSRM with high temperature (800−1000 K), the water gas shift reaction reaches conditions near equilibrium immediately due to the reverse reaction domination.40 However, the equilibrium conversion of the water-gas shift reaction at such high temperatures is quite low. In fact, the high temperature of SR side causes the rates of the direct water gas shift and reverse water gas shift reactions to increase rapidly in the entrance of SR side. Therefore with kinetic parameters that are employed in this modeling, the overall reaction rate becomes very low. This subject is reason of

Figure 9. Variation of component mole fractions along reactor axes in the FR of CLC-SRM.

a very low or near zero value for the overall water gas shift reaction rate in Figure 7b. Reduction of rates in the SR sides is due to reduction of the temperature and cold spot in the reactor entrance. The enhancement in SR reaction rates in the first part of the reactor is mainly due to the improved heat supply by the CLC process, while the enhancement in the rest of the reactor is mainly due to the shift of the equilibrium by removing hydrogen from the reactor with the hydrogen-permeable membrane. I


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Figure 10. (a) Comparison of CH4 conversion along the reactor axes among CSR, CLC-SR, and CLC-SRM. (b) Comparison of H2 yield along the reactor axes among CSR, CLC-SR, and CLC-SRM. (c) Comparison of H2 flow rate along reactor axes between CSR, CLC-SR, and CLC-SRM. (d) Variation of H2 flow rate along reactor axes in SR side of CLC-SRM.

According to Figure 8b in the AR, NiO conversion (Xs) decreases fast from 0.5 to 0 in the entrance of the AR due to the high temperature of AR increases the rate of oxidation. It must be noted that in CLC-SRM, AR dimensions and operating parameters have been designed based on experimental research in the literature for full oxidation of oxygen carriers with φ equal to 2.28,46,47 The variation of component mole fractions along the reactor axes in the dense bed and freeboard zones of the FR has been presented in Figure 9. Results indicate that the total amount of methane is consumed in the dense bed part of FR. Consumption of total methane in the dense bed zone of FR causes the combustion efficiency to increase to 1 in the FR. Comparisons of CH 4 conversion, H 2 yield, and H 2 production along the reactor axes among CSR, CLC-SR, and CLC-SRM have been demonstrated in Figure 10. It is obvious that methane conversion and hydrogen production of SR in the CLC-SRM is higher than CSR and CSR-SR due to better heat and rate effects. Results shows that by employing CLC instead of furnace and assistance with membranes, CH4 conversion increases from 26 to 33.55% and H2 production increases from 18.37 to 23.05 mol s−1 for each tube in CSR compared with CLC-SRM, respectively. The enhancement in methane conversion in the first part of the reactor is mainly due to the improved heat supply by the CLC process, while the enhancement in the rest of the reactor is mainly due to the shift of the equilibrium by removing

Figure 11. Variation of pressure along reactor axes in AR, FR, and SR sides of CLC-SRM.

5.3. Molar Behavior. Figure 8 depicts variation of component mole fractions and oxygen carrier conversion along the reactor axes in the AR. As seen, oxygen is consumed rapidly in the first half of the reactor due to the fast fluidization phenomenon, high gas velocity, and suitable design of the AR. Therefore the oxidation rate of solids becomes zero as shown in Figure 7. The remainder of oxygen at the end of AR is due to excess air that has been employed for combustion and equal to 20%. J


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Figure 12. (a) Effect of FR feed flow rate ratio on the CH4 conversion along the reactor axes in the SR side of CLC-SRM. (b) Effect of FR feed flow rate ratio on the H2 production along the reactor axes in the SR side of CLC-SRM. (c) Effect of FR feed flow rate ratio on the H2 yield along reactor axes in the SR side of CLC-SRM. (d) Effect of the FR feed flow rate ratio on the outlet temperature of the SR side along the reactor axes in CLC-SR. (e) Effect of FR feed flow rate ratio on the combustion efficiency along the reactor axes in the FR side of CLC-SRM.

5.3. Effect of FR Feed Flow Rate on the CLC-SRM Performance. For prediction of the FR feed flow rate on the CLC-SRM performance, 90 mol s−1 CH4 for the FR feed has been considered as the lower band and the feed flow rate ratio for this case is equal to 1. In fact this flow rate is based on CSR results, and for all cases 20% excess air and ϕ equal to 2 have been considered. It must be noted that, as shown in Table 2, the total feed gas flow rate of FR in CLC-SRM is 108 mol s−1. As shown in Figure 12a−c, by increasing the feed flow rate ratio of FR from 1 to 2, CH4 conversion, H2 yield, and H2 production increases 16.73%, 62.03%, and 40%, respectively, due to generation of more heat in AR and receiving more heat from AR in SR. By receiving more heat from AR, the temperature of SR increases and therefore the rate of reactions

hydrogen from the reactor and thus with the hydrogenpermeable membrane. Reduction of the H2 yield in CLC-SRM is due to removing of H2 by membranes and this subject has been shown in Figure 10d. In fact, for determination of the H2 yield, the H2 flow rate in the membrane side has not been considered, and this result only presents the SR side of CLC-SRM. Therefore, the H2 flow rate of CLC-SRM is equal to the H2 flow rate summation of the SR side and membrane side (see Figure 10d). Figure 11 presents pressure along the CLC-SRM reactor axis in SR, AR, and FR sides, respectively, for industrial applications. As shown in this figure, pressure drop in the AR is high due to the high gas velocity for fast fluidization. The pressure drop in the SR, FR, and AR is 5.8, 4.9, and 6.4 bar, respectively. K


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pure carbon dioxide can be captured. Additionally, by considering emission of this large amount of carbon dioxide to the atmosphere in 1 year, it is better to replace CSR with this novel configuration.

6. CONCLUSIONS In this novel research, a concept study for application of CLC instead of a furnace in a steam reformer assisted by hydrogen perm-selective Pd−Ag membranes (CLC-SRM) for CO2 capture and hydrogen production has been proposed. NiO18αAl2O3 particles have been employed as oxygen carriers in CLC-SRM. CLC consist of two interconnected fluidized bed reactors that AR and FR operate in fast fluidization and bubbling fluidization regimes, respectively. Steam reforming of methane occurs over the packed Ni-based catalyst in the vertical tubes located in the AR and Pd−Ag tubular membranes are covered inside the tubes of SR. A steady state one-dimensional heterogeneous catalytic reaction model is applied to analyze the performance of CLC-SRM. Results shows that by employing CLC instead of a furnace and assistance of SR with membranes, CH4 conversion in SR increases from 26 to 33.54% and H2 production increases from 3380 to 4241 mol s−1. The total amount of methane consumed in the dense bed of FR and combustion efficiency increases to 1 in the dense bed part of FR in CLC-SRM. The results show that enhancement in methane conversion and reaction rate of the SR side in the first part of the CLCSRM is mainly due to the improved heat supply by the CLC process, while the enhancement in the rest of the reactor is mainly due to the shift of the equilibrium by removing hydrogen from the reactor with the hydrogen-permeable membrane. By increasing the feed flow rate ratio of FR from 1 to 2 with 20% excess air and ϕ equal to 2, CH4 conversion in the SR side increases from 30.17% to 46.9% and also H2 production increases from 3908 to 5483 mol s−1. The total amount of methane consumed in the dense bed of FR and combustion efficiency increases to 1 in the dense bed part of FR. Therefore, by employing CLC-SRM instead of CSR, 410 ton day−1 of almost pure carbon dioxide can be captured and emission of this large amount to the atmosphere can be reduced.

■

Figure 13. (a) Effect of FR feed flow rate ratio on the AR heat generation in CLC-SRM. (b) Effect of FR feed flow rate ratio on the FR heat consumption in CLC-SRM. (c) Effect of FR feed flow rate ratio on the SR heat consumption in CLC-SRM.

AUTHOR INFORMATION

Corresponding Author

*Phone: +98 711 2303071. Fax: +98 711 6287294. E-mail: [email protected], [email protected]. Notes

The authors declare no competing financial interest.

increase (as shown in Figure 12d). Increasing the FR feed flow rate has a negative effect on the combustion efficiency (η) because AR and FR have been designed for a FR feed flow ratio equal to 1.2. Therefore for the higher flow rate of FR feed flow rate with constant dimensions of AR and FR (see Table 2), some of the CH4 does not convert and combustion efficiency (η) becomes lower than 1 and CO2 cannot be captured at the outlet of FR (see Figure 12e). For better explanation of thermal behavior with increasing feed flow rate ratio of FR, Figure 13 presents variation of heat generation in AR and heat consumption in FR and SR sides along reactor axes in CLC-SRM. 5.4. Carbon Dioxide Capture. One of the advantages of CLC-SRM related to CSR is carbon dioxide capturing. By employing CLC-SRM instead of CSR, 410 ton day−1 of almost

■

L

NOMENCLATURE av = specific surface area of catalyst pellet, m2 m−3 a = activity of oxygen carriers in reduction and oxidation (equal to 1) A = membrane area available for hydrogen permeation (m2) Ac = cross section area, m2 Ai = inside area of inner tubes, m2 Ao = outside area of inner tubes, m2 Ct = total concentration, mol m−3 Cp = specific heat of the gas at constant pressure, J mol−1 dp = particle diameter, m Di = inside diameter, m Do = outside diameter, m dx.doi.org/10.1021/ef400880q | Energy Fuels XXXX, XXX, XXX−XXX

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ΔHf,i = enthalpy of formation of component i, J mol−1 eb = void fraction of catalytic bed δ = bubble phase volume as a fraction of total bed volume δm = membrane thickness of Hydrogen perm-selective membrane (m) γ = volume fraction of catalyst occupied by solid particle in bubble ϕ = oxygen-carrier circulation flow rate to fuel flow rate ηc = combustion efficiency η = effectiveness factor ξ = constant for hydrogen permeation

Dj = outside diameter of tubes in CSR, m Dij = binary diffusion coefficient of component i in j, m2 s−1 Dim = diffusion coefficient of component i in the mixture, m2 s−1 Do = tube outside diameter, m Ea = activation energy, J mol−1 EH2 = activation energy of permeability in membrane, J mol−1 Ft = total molar flow rate, mol s −1 Fi = molar flow rate of component i, mol s −1 Fb = molar flow in bubble phase, mol s −1 Fe = molar flow in emulsion phase, mol s −1 hf = gas−solid heat transfer coefficient, Wm2− K−1 hi and ho = heat transfer coefficient between fluid phase and reactor wall in exothermic and endothermic sides with convection, Wm2− K−1 JH2 = hydrogen permeation rate, mol s−1 k1 = reaction rate constant for the first rate equation, mol kg−1 s−1 k2 = reaction rate constant for the second rate equation, mol kg−1 s−1 k3 = reaction rate constant for the third rate equation, mol kg−1 s−1 ki = rate constant of reaction i, mol kg−1 s−1 bar −1/2 kg,i = mass transfer coefficient for component i, m s−1 K = conductivity of fluid phase, Wm1− K−1 Kw = thermal conductivity of reactor walls, Wm1− K−1 k0i = Pre-exponential factor of chemical reaction rate constant for oxidation and reduction of oxygen-carriers, mol1−n m3n−2 s−1 L = reactor length, m Mi = molecular weight of component i, g mol−1 n = reaction order for oxidation and reduction of oxygencarriers P = total pressure, Pa PTube = total pressure in membrane side, Pa PShell = total pressure is sweep gas side, Pa Pi = partial pressure of component i, Pa Q0 = pre-exponential factor of hydrogen permeability (mol m−1 s−1 kPa−0.5) qw = constant heat flux in CSR (W m−2) R = universal gas constant, J mol−1 K−1 Re = Reynolds number Sci = Schmidt number of component i T = temperature, K t = time, s u = superficial velocity of fluid phase, m s−1 ug = linear velocity of fluid phase, m s−1 UAS = overall heat transfer coefficient between AR side and SR side, Wm2− K−1 UAF = overall heat transfer coefficient between AR side and FR side, Wm2− K−1 USM = overall heat transfer coefficient between SR side and membrane side, Wm2− K−1 XNiO = nickel oxide conversion yi = mole fraction of component i z = axial reactor coordinate, m

Superscripts

g = in bulk gas phase s = at surface catalyst AR = air reactor FR = fuel reactor SR = steam reforming side M = membrane side Subscripts

■

0 = inlet conditions i = chemical species j = reactor side

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Greek Letters

μ = viscosity of fluid phase, kg m−1 s−1 ρ = density of fluid phase, kg m−3 ρ = density of oxygen carriers, kg m−3 ρb = density of catalytic bed, kg m−3 τi = time needed for full conversion, s uci = critical volume of component i, cm3 mol−1 M


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N


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