Production of ultrapure hydrogen from biomass gasification with air

Chemical Engineering Science 64 (2009) 582 -- 592

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Production of ultrapure hydrogen from biomass gasification with air Peijun Ji, Wei Feng ∗ , Biaohua Chen ∗ State Key Laboratory of Chemical Resource Engineering, Beijing University of Chemical Technology, Beijing 100029, China

A R T I C L E

I N F O

Article history: Received 8 May 2008 Received in revised form 30 September 2008 Accepted 16 October 2008 Available online 1 November 2008 Keywords: Biomass gasification Fluidized bed gasifier Integrated process

A B S T R A C T

An integrated process has been proposed for the production of ultrapure hydrogen from biomass gasification with air. The process consists of an air-blown bubbling fluidized bed gasifier, a steam reformer, and a water-gas-shift membrane reactor. A non-isothermal model has been developed to simulate the fluidized bed gasifier, and a one-dimensional model has also been developed to simulate the steam reformer. The simulation results are compared with the experimental data, and good agreement is obtained. Based on the simulation results, the thermodynamic analysis of the integrated process is carried out. The simulation and analysis provide a quantitative tool for gaining insight into the process. © 2008 Elsevier Ltd. All rights reserved.

1. Introduction Biomass is an attractive energy source for hydrogen production. If biomass as a feedstock is derived from dedicated energy crops, then the biomass-based technology can provide a way to recycle CO2 . Through gasification, biomass can be converted into a product gas containing hydrogen, carbon monoxide, carbon dioxide, methane, and C2 gases (Ni et al., 2006). Generally, the product gas from biomass gasifiers contains tar, which is a complex mixture of condensable hydrocarbons. Gasification of biomass offers efficiency, environmental, and operational advantages including the ability to use the product gas in fuel cells. However, there are some limitations for electrochemically converting the product gas in a fuel cell system. Air-blown gasification in fluidized bed reactors typically produces relatively low concentration of hydrogen (Radmanesh et al., 2006), which will lead to a low efficiency of hydrogen conversion in fuel cell; tar will impose serious limitations in the use of the product gas due to fouling of the fuel cell. To overcome the limitations, an integrated process producing ultrapure hydrogen from biomass gasification is proposed as shown in Fig. 1. The process consists of an air-blown bubbling fluidized bed gasifier, a steam reformer, and a water-gas-shift (WGS) membrane reactor. In the gasifier, with air as the gasifying agent, biomass is gasified. The product gas is upgraded in the steam reformer through the conversion of tar, CH4 , and C2 gases. The additional production of H2 is carried out in the WGS membrane reactor; in the meantime the hydrogen is separated through the hydrogen permeable ∗

Corresponding author. Tel.: +86 10 64446249. E-mail addresses: [email protected] (W. Feng), [email protected] (B. Chen). 0009-2509/$ - see front matter © 2008 Elsevier Ltd. All rights reserved. doi:10.1016/j.ces.2008.10.015

membrane. The WGS membrane reactor has a reaction channel and a permeate channel. The hydrogen is permeated through the H2 permeable membrane into the permeate channel. A vacuum pump is used to withdraw the permeated H2 and maintain a low pressure in the permeate channel. The rejected gas containing H2 , CO, and unconverted CH4 is utilized in the catalytic furnace. The separated pure H2 is compressed to 20 bar. The process can make the production of ultrapure hydrogen more efficient. Simulation has been shown as an effective tool to study the complicated phenomena in a fluidized bed reactor (Chen et al., 2007; Radmanesh et al., 2006; Lee et al., 2007; Gungor, 2008; Nikoo and Mahinpey, 2008; Van de Velden et al., 2008). The simulation of the gasifier, steam reformer, and WGS membrane reactor can provide a quantitative tool for gaining insight into and understanding the integrated process. It is very useful for the analysis, evaluation, and design of the process. In this work, non-isothermal models for the air-blown bubbling fluidized gasifier and catalytic steam reformer will be developed. Based on the simulation results for the reactors, the proposed process will be analyzed and evaluated in terms of the pure hydrogen yield (mol/kg biomass) and overall thermodynamic efficiency. 2. Simulation of biomass gasification 2.1. Biomass gasifier model The bubbling fluidized bed gasifier under consideration is schematically shown in Fig. 2a, which has a bubbling fluidized bed of sand particles and a freeboard zone. Biomass is gasified with air as the gasifying agent. An electric heater surrounding the gasifier provides the energy required.

P. Ji et al. / Chemical Engineering Science 64 (2009) 582 -- 592

583

H2

H2O Tar

Biomass Cleaner

Steam Reformer

Gasifier Air

Furnace

H 2-membrane WGS reactor

flue gas rejected fuel gas

Fig. 1. The integrated process for the production of pure hydrogen.

Table 1 Hydrodynamic parameters of the bubbling fluidized bed

biomass

bubble phase

emulsion phase

NF

Gas

Particles

Heat transfer to the emulsion phase

Heat transfer to the bubble phase

bubbling fluidized bed

freeboard zone

product gas

Parameters

dp g umf

Minimum fluidization Reynolds number Bubble diameter

g

0.5

= [27.22 + 0.0408Ar]

0.4

db0 = 0.347[A(ug − umfi )/nd ] Bubble velocity Bubble fraction Bubble-emulsion mass transfer coefficient

− 27.2

db = dbm + (db0 − dbmi ) exp(−0.3z/Dt) dbm = 0.652[A(ug − umfi )]

ub = ug − umfi + 0.711(gdb ) ug − umfi b = ub − umfi

0.4

1/2

1 1 1 = + Kbe Kbc Kce   D1/2 g1/4 umf + 5.85 5/4 db db   Dmf ubr Kce = 6.77 d3b   1/2 (kg g Cpg ) g1/4 umf g Cpg + 5.85 Hbe = 4.5 5/4 db db Kbc = 4.5

interphase mass and heat transfer NF: net flow

gasifying agent (steam)

Equations

gas flow Bubble-emulsion heat transfer coefficient

Fig. 2. Schematic diagram of a bubbling fluidized bed of the gasifier. (a) Gasifier, (b) model for the bubbling fluidized bed. Table 2 Mass and energy conservation equations for the bubbling fluidized bed

A one-dimensional non-isothermal model is developed to simulate the biomass gasifier. A schematic diagram of the model and the two-phase representation of the fluidized bed are presented in Fig. 2b. The model considers that the bubbling fluidized bed contains a particle-free bubble phase and an emulsion phase. The emulsion phase, containing all the particles and a fraction of gases, is maintained at incipient fluidization conditions. The gas in excess of maintaining the incipient fluidization conditions in the emulsion phase is assumed to pass through the bed as bubbles. Mass transfer between the bubble and emulsion phases, due to both molecular diffusion and net flow (Yan et al., 1998), is considered. The heterogeneous reactions and the homogeneous reactions in the emulsion phase are the major contributions to the net flow. These reactions generate gases in excess of what is required to maintain incipient fluidization in the emulsion phase. The transfer of the excess gases from the emulation phase into the bubble phase leads to the net flow. In the developed model, all gases in both bubble and emulsion phases are assumed to be in the plug flow mode, and radial dispersion is not considered because the solids are well mixed within the emulsion phase (Ross et al., 2005). Biomass particles are assumed to be spherical and of uniform size. Biomass pyrolysis is assumed to take place instantaneously after being fed into the gasifier (Radmanesh et al., 2006). The hydrodynamic parameters (Kunii and Levenspiel, 1991) of the bubbling fluidized bed are listed in Table 1. For the bubbling fluidized bed, the equations of mass balance and the equations of energy conservation for the gas phases in the bubble and emulsion

Bubble phase Mass balance   NR  dFbi = A b Kbei (Cei − Cbi ) + fnet,i + b i,j rb,j dz j=1 Energy conservation  NbR dQ dF Hj rb,j ) + b − Ni=1 CPI (Tb − TR ) bi A(b Hbe (Te − Tb ) + fnet Hnet − b j=1 dTb dz dz = N dz i=1 Fbi CPi Emulsion phase Mass balance NHomo NHetero dFei = A(b Kbei (Cbi − Cei ) − fnet,i + (1 − b )(mf j=1 i,j re,j + (1 − mf )( j=1 i,j re,j ))) dz Energy conservation  NHomo NHetero =A(b Hbe (Tb −Te )−fnet Hnet −(1−b )mf j=1 Hj re,j −(1−b )(1−mf )( j=1 Hj re,j ) +b (1 − mf )s hp (Ts − Te )) +  dTe = N dz i=1 Fei CPi

N  dQe dFei − CPi (Te − TR ) dz dz i=1

Boundary conditions z = 0, Cei = Cbi = Ci ,0 , Te ,0 = Tb ,0 = Tinlet

phases are listed in Table 2. The energy equations take into account the heat of reaction, the heat exchanged between the bubble and the emulsion phases due to the temperature difference, the heat transferred due to the net flow, and the heat exchanged between well mixed sand particles and the emulsion phase. So the heat transfer between phases and between material and wall are considered.

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In the model for the freeboard zone, it is assumed that the gas flow is in a plug flow mode. The mass and energy conservation equations are described by Eqs. (1) and (2), respectively. dFfi dz

dTf dz

NHomo



=A

i,j rf ,j

(1)

j=1

=

A

NHomo j=1

Hj rf ,j + dQf /dz − N

N

i=1 Hfi dFfi /dz

(2)

i=1 Ffi CPi

Table 3 Kinetic parameters in Eq. (3) Ei (kJ/mol)

Vi∗ (kg/kg biomass)

Component

Log(k0, i ) (1/s)

Total devolatilization Total gas H2 CH4 C2 CO CO2 H2 O Char Tar

8.30 133.01 0.969 2.88 49.37 0.476 6.17 114.18 0.0016 13.00 251.21 0.0241 9.06 173.85 0.1227 11.75 220.66 0.2164 5.39 97.99 0.0308 6.71 103.01 0.0804 Char = 1−total devolatilization Tar = (total devolatilization )−(total gas)

2.2. Kinetic models After entering the fluidized bed of hot sand particles, the biomass particles undergo fast primary pyrolysis. Biomass pyrolysis generates three different products in different quantities: gas, tar, and char. In the kinetic model it is assumed that the biomass decomposed directly to each product i by a single independent reaction pathway (Hajaligo et al., 1982; Nunn et al., 1985; Radmanesh et al., 2006). The rate of formation of a product i in yield Vi at time t is given by dVi = k0,i e−(Ei /RTP ) (Vi∗ − Vi ) dt

(3)

where k0, i and Ei are the pre-exponential factor and apparent activation energy for component i, respectively. The quantity Vi∗ is the ultimately attainable yield of component i. Table 3 lists the parameter values for each species. The parameters are adopted from the literature (Hajaligo et al., 1982; Nunn et al., 1985). Because the mixing of the biomass particles proceeds more slowly than pyrolysis (Radmanesh et al., 2006), pyrolysis is considered to take place at the feeding zone of the fluidized bed. The thermal cracking of tar, also called secondary pyrolysis, has a significant effect on the final gas composition, because more than half of the primary pyrolysis products accounts for tar. A firstorder kinetic model by Boroson (Boroson et al., 1989) has been widely applied to simulate tar cracking. However, with the model,

Table 4 Reactions for the thermal tar cracking Chemical reaction C6 H6 O→CO+0.4C10 H8 +0.15C6 H6 +0.1CH4 +0.75H2 C6 H6 O+3H2 O→4CO+2CH4 +2.H2 C10 H8 →7.38C+0.275C6 H6 +0.97CH4 +1.235H2 C6 H6 +2H2 O→1.5C+2.5CH4 +2CO C6 H6 O+4O2 →3H2 O+6CO C6 H6 +4.5O2 →6CO+3H2 O C10 H8 +7O2 →4H2 O+10CO

Kinetic equations   105 CC6 H6 O (mol/m3 s) r1 = 107 exp −  RT5  10 r2 = 107 exp − CC6 H6 O (mol/m3 s) RT   3.5 × 105 CC1.6 H CH−0.5 (mol/m3 s) r3 = 1.7 × 1014 exp − 10 8 2 RT   5 4.43 × 10 CC1.3H CH−0.4 CH0.2O (mol/m3 s) r4 = 2.0 × 1016 exp − 2 6 6 2  RT  9650 CC0.5H O CO2 (mol/m3 s) r5 = 0.655 × 103 T exp − 6 6 T   1.2552 × 105 CC−0.1 r6 = 2.4 × 1011 exp − C 1.85 (mol/m3 s) 6 H6 O2 RT   9650 0.5 CC H CO2 (mol/m3 s) r7 = 0.655 × 103 T exp − 10 8 T

Morf et al. (2002) Morf et al. (2002) Jess (1996a) Jess (1996a) Smoot and Smith (1985) Westbrook and Dryer (1984) Smoot and Smith (1985)

Table 5 Heterogeneous and homogeneous reactions in the biomass gasifier Chemical reactions

Kinetic equations

Heterogeneous reaction

  29844 CCO2 (mol/m2 s) r8 = 4364 exp − Tp k1 pH2 O r9 = (1/s) 1 + k2 pH2 O + k3pH2  18522 k1 = 4.93 × 103 exp − T P  3548 k2 = 1.11 × 10 exp − T P   25161 k3 = 1.53 × 10−9 exp TP

C+CO2 →2CO C+H2 O→CO+H2

Homogeneous reaction CH4 +H2 O→3H2 +CO CH4 +2O2 →CO2 +2H2 O 2CO+O2 →2CO2 CO+H2 O→CO2 +H2

2H2 +O2 →2H2 O

  2.4 × 105 1.7 CCH r10 = 3.3 × 1010 exp − C −0.8 (mol/m3 s) 4 H2 RT   24343 0.7 0.8 [CH4 ] [O2 ]2 (mol/m3 s) r11 = 1.58 × 1019 exp − Tg   20129 0.25 0.5 [CO][O2 ]2 [H2 O] [mol/m3 s] r12 = 3.98 × 1020 exp − Tg   1510 2 r13 = 2.78 × 106 exp − (mol/m3 s) (yCO yH2 O − yCO2 yH2 /KEQ )Cmol Tg  3968 KEQ = 0.0265 exp Tg  13127 [O2 ][H2 ] (mol/m3 s) r14 = 2.196 × 1018 exp − Tg

Wurzenberger et al. (2002) Wurzenberger et al. (2002)

Jess (1995) Groppi et al. (2000) Groppi et al. (2000)

De Souza-Santos (1989) Groppi et al. (2000)

P. Ji et al. / Chemical Engineering Science 64 (2009) 582 -- 592

the final gas composition could not be simulated well over wide operation conditions. Reactions and reaction kinetics involved for the tar cracking can greatly affect the modeling results of final gas composition. In this work, phenol is used to represent the tar from primary pyrolysis as discussed in the paper by Gerun et al. (2008). The reactions and reaction kinetics for the tar cracking are presented in Table 4. With phenol as the model compound of the tar, the effect of equivalence ratio (ER) and bed temperature on the gas composition and tar content of the product gas can be described well. The results will be demonstrated and discussed in Section 5. After the primary pyrolysis, in the fluidized bed and freeboard zone, the homogeneous and heterogeneous reactions involved as well as the kinetic equations are listed in Table 5.

The heat capacity CP of the species involved in the reactions is considered to be a function of temperature. Both the thermal conductivity and viscosity of the gas mixtures are a function of the temperature and composition of the local domain. The viscosity of the gas mixtures in the gaseous phase is calculated according to (Bird et al., 2002) N  a=1

x a a  xb ab

(4)

where ab is calculated using     2  1/2  Mb 1/4 Ma −1/2 a 1+ 1+ Mb b Ma 8

1

ab = √

(5)

a and b are calculated using  = 2.6693 × 10−6

√ MT

 

The thermal conductivity is calculated using (Bird et al., 2002)

T/M K = 8.322 × 10−2

2 k

kmix =

N  a=1

xa ka  xb ab

(6)

(7)

(8)

The heat transfer coefficient hp between the sand particle and the gas is calculated according to (Kunii and Levenspiel, 1991) hp dp 1/2 = 2 + 0.6Rep Pr1/3 kmix

(9)

with Rep = Pr =

dp ug g

mix

Cp ug kmix

Table 6 Reactor model for the steam reformer Reaction channel Continuity equation NR  dFSR = ar (ik rk ) dz k=1 Energy equation  NR (−Hk rk ) − q ar k=1 dTSR = 6 dz i=1 Fi CP I

q=

aw kw

w

(TSR − Tfurnace )

Furnace side Continuity equation NR  dFfurnace = afurnace (ik rk ) dz k=1

2.3. Physical properties

mix =

585

Energy equation  NR afurnace k=1 (−Hk rk ) + q dTfurnace  = dz i=1 Ffurnace CPI

volume. At the entrance of the reactor, a uniform distribution of gases is assumed in the two phases. The concentrations of gas species and the temperatures of the emulsion and bubble phases are calculated by solving the differential equations listed in Table 2. The stepwise calculation starts from the bottom of the bed, where flow rates of fluidizing gases are given. If the overall energy conservation of the fluidized bed (the criterion is that the calculated value of the total energy leaving the bed does not exceed 0.5% of the total energy entering the bed) is not attained, a new estimate of the solid temperature is given, and the calculation procedure is repeated as described above. 3. Steam reformer model The steam reformer under consideration includes a reaction channel and a catalytic gas-fired furnace. The catalytic furnace is used to heat the reactants up to a certain temperature. The remaining H2 , CO, and CH4 in the exhaust gas of the membrane reactor are then introduced into the furnace. A one-dimensional steady-state model is developed to simulate the steam reformer of the integrated process as shown in Fig. 1. Table 6 lists the model for the steam reformer, the reactions and the reaction kinetics are listed in Table 7. The reaction of C6 H6 O conversion is not taken into account in the model development, because the phenol has been completely converted in the air-blown gasifier according to the simulation results. Benzene and toluene generated from the cracking of the primary tar are taken as the model compounds of the tar from the gasifier. 4. WGS membrane reactor model

(10)

(11)

2.4. Numerical calculation The program is written in Fortran. The initial value of the solid temperature Tp in the bed is first assumed, and the solid temperature is taken to be constant throughout the bed. The parameters, including the bubble diameter db , the bubble velocity ub , the gas velocity ug , the bubble fraction b , and the net flow fnet , are a function of the height of bed, and they are calculated for each control

The WGS membrane reactor consists of a reaction channel for the conversion of CO formed in the process and a permeate channel for the removal of the permeating H2 through the H2 -membrane. A vacuum pump is used to withdraw the permeated H2 and maintain a low pressure in the permeate channel. As the area of the H2 membrane is set constant during the process analysis, the partial pressure of the permeate channel is the main variable affecting the separation of pure H2 . The continuity and energy equations for the reaction and permeate channels of the WGS membrane reactor are listed in Table 8. The energy equations take into account the heat of reaction, the heat exchanged between the non-reaction side (permeate side) and the reaction side, and the energy carried by diffusing H2 .

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P. Ji et al. / Chemical Engineering Science 64 (2009) 582 -- 592

Table 7 Kinetic models for the steam reformer Kinetic models for the steam reforming Chemical reaction

Kinetic equations   2.50 × 105 CCH4 (mol/m3 s) r1 = 4.5 × 1010 exp − RT   5 1.96 × 10 CC6 H6 (mol/m3 s) r2 = 2.0 × 1011 exp − RT   5 3.32 × 10 0.3 3 C0.2 r3 = 4.3 × 1013 exp − C10 H8 CH2 (mol/m s) RT

CH4 +0.5CO2 +0.5H2 O→1.5CO+2.5H2 C6 H6 +3CO2 +3H2 O→9CO+6H2 C10 H8 +5CO2 +H2 O→15CO+9H2

Kinetic models for the furnace Chemical reaction 2CO+O2 →2CO2

Gas composition (mol %, dry basis)

Table 8 Models for the H2 -membrane WGS reactor Reactor model Reaction side Continuity equation dFH2 dFCO WGS WGS = −ar (1 − B )WGS = ar (1 − B )WGS (−rCO ); (−rCO ) − aH2 NH2 s s dz dz Energy equation dTrWGS 1 WGS =  (HWGS ar (1 − B )(−rCO )WGS − q − aH2 NH2 HH2 ) s Fi CPi dz

sp

WGS (TrWGS − Tnr )

Nonreaction side Continuity equation dGH2 = aH2 NH2 dz Energy equation

Groppi et al. (2000)

a

b c

d

e

c

c

f

g

20 15

CO2

10

CO H2

5

CH4 C2

0.25

(pH2 O )

(pCO2 )

−0.60

The permeability of hydrogen through the H2 permeable membrane is calculated according to

H2



high

( pH

2

0.35

0.40

0.45

Fig. 3. Effect of the equivalence ratio on the gas composition. The bed temperature of gasifier is 1073 K and the temperature of freeboard 823 K. (Ergudenler and Ghaly, 1992; Czernik et al., 1994; Narváez et al., 1996; Maniatis et al., 1988; Chamberland, 1986; Kurkela and Stahlberg, 1992; Rensfelt and Ekstrom, 1989).

(1 − )

 = pCO2 pH2 /(pCO pH2 O Keq ), kWGS = kWGS exp(−ea /RT) 0

Pm exp(−EA /RT)

0.30

Equivalence ratio

ln kWGS = 16.68, ea = 114.6 kJ/mol 0

N H2 =

Groppi et al. (2000)

Exp. data CO2 CO H2 CH4 C2 a--Ergudenler and Ghaly (1992) b--Czernik et al. (1994) c--Narvaez et al. (1996) d--Maniatis et al. (1988) e--Chamberland (1986) f--Kurkela and Stahlberg (1992) g--Rensfelt and Ekstrom (1989) Simulation

0.25

Kinetic model (Bohlbro, 1969) 0.90

Groppi et al. (2000)

0

WGS dTnr 1 =  [q + aH2 NH2 HH2 ] dz j Gj Cpj

WGS = kWGS (pCO ) −rCO

Jess (1996b)

  20129 0.25 0.5 r4 = 3.98 × 1020 exp − [CO][O2 ] [H2 O] (mol/m3 s) Tg Heterogeneous   5533 CCO (mol/m3 s) r5 = 400 exp − Tg Homogeneous   13127 r6 = 2.196 × 1018 exp − [O2 ][H2 ] (mol/m3 s) Tg Heterogeneous   5282 r7 = 791 exp − CH2 (mol/m3 s) T Homogeneous   24343 0.7 0.8 r8 = 1.58 × 1019 exp − [CH4 ] [O2 ] (mol/m3 s) Tg Heterogeneous   10563 CCH4 (mol/m3 s) r9 = 20 exp − T

CH4 +2O2 →CO2 +2H2 O

aH2 km

Jess (1996b)

Kinetic equations Homogeneous

2H2 +O2 →2H2 O

q=

Jess (1996b)

 −

plow ) H 2

(12)

The apparent activation energy EA and the pre-exponential factor Pm of the membrane are 29.73 kJ/mol and 7.71×10−4 mol m/(s m2 bar0.5 ), respectively (Basile et al., 2001). 5. Simulation results 5.1. Air-blown bubbling fluidized bed gasifier Two important parameters affecting the biomass gasification in an air-blown bubbling fluidized biomass gasifier are the ER and the operation temperature of the gasifier. The ER is the air to fuel mass flow ratio used in the gasifier divided by the air to fuel mass flow

ratio required for a complete combustion. Figs. 3–5 present the effect of the ER on the gas composition, tar content, and low heating value of the product gas, in which, the data are illustrated as symbols, and the lines represent the simulation results. As can be seen, the simulated gas compositions, tar content, and low heating value are in good agreement with the experimental data. With the increase of ER, more combustible gases and tar are combusted. This leads to the increase of the concentration of CO2 and the decrease of the low heating value of the product gas as ER increases, as shown in Figs. 3 and 5. The heat released by the combustion reactions promotes the conversion of the tar and hydrocarbons, as described by the models in Tables 4 and 5. Therefore, the tar content decreases with ER, as shown in Fig. 4. Figs. 6–8 show the effect of the bed temperature on the gas composition, tar content, and low heating value of the product gas. The symbols are the experimental data, and the lines represent the simulation results. In the bed temperature range from 973 to 1173 K,

P. Ji et al. / Chemical Engineering Science 64 (2009) 582 -- 592

587

Exp. data (Narvaez, 1996) Simulation

25 Tar content (g/Nm3)

Gas composition (mol%, dry basis)

30

20 15 10 5

25

Exp. data CO2 CO simulation

20

0.30

0.35 Equivalence ratio

0.40

Low heating value (MJ / Nm3)

Fig. 4. Tar content as a function of the equivalence ratio. The bed temperature of gasifier is 1073 K and the temperature of freeboard 823 K.

Exp. data Narvaez et al. (1996) Kurkela and Stahlberg (1992) 8 Maniatis et al. (1994) Gulyurtlu et al. (1994) Ergudenlerand Ghaly (1992) Simulation

6

C2

15 H2 CO2

10 5

950

0.45

CH2

1000

1100 1050 Temperature of bed (K)

1150

Fig. 6. Gas composition as a function of the bed temperature of gasifier. The equivalence ratio is 0.30. The symbols are the experimental data (van der Aarsen et al., 1982; Bilodeau et al., 1993).

20

Tar content (g/ Nm3)

0.25

CH4

CO

C2

0

H2

15

Exp. data Kurkela and Stahlberg (1992) Rensfelt and Ekstrom (1989) Czernik et al. (1994) Narvaez et al. (1996) Simulation

10

5

4

0.25

0.30

0.35

0.40

0.45

Equivalence ratio Fig. 5. Low heating value of the product gas as a function of the equivalence ratio. The bed temperature of gasifier is 1073 K and the temperature of freeboard 823 K. (Narváez et al., 1996; Kurkela and Stahlberg, 1992; Maniatis et al., 1994; Gulyurtlu et al., 1994; Ergudenler and Ghaly, 1992).

the simulated gas compositions are in good agreement with the experimental data. A higher bed temperature favors the conversion of tar and CH4 and the production of H2 and CO. Fig. 8 shows the effect of the bed temperature on the low heating value of the product gas, the data are from different literature. Increasing the bed temperature is favorable for the endothermic reactions as described in Tables 4 and 5. However, some reactions listed in Tables 4 and 5 are exothermic. So the effect of the bed temperature on the low heating value is not so obvious. As can be seen in Fig. 8, the simulated low heating value slightly increases with the bed temperature. Fig. 9 presents the temperature profile along the height of the gasifier. The predicted temperature profile of the gasifier is consistent with the experimental observation (Narváez et al., 1996), which indicated a good isothermicity in the gasifier bed. 5.2. Catalytic steam reformer Fig. 10 shows the profiles of the temperature, gas composition, and tar content along the reformer's axial coordinate, when the temperature of steam reformer is 1073 K. As shown in Fig. 10a, due to the rapid combustion of the fuel gas, the temperature of the furnace is increased rapidly at the entrance. As the heat is transferred from

0 950

1000

1050

1100

1150

1200

Temperature of bed (K) Fig. 7. Tar content as a function of the bed temperature of gasifier. The equivalence ratio is 0.35. (Kurkela and Stahlberg, 1992; Rensfelt and Ekstrom, 1989; Czernik et al., 1994; Narváez et al., 1996).

the furnace to the reaction side, the temperature of the furnace decreases along the reformer's axial coordinate. Because of the dry reforming reaction, the CO2 composition decreases with the reactor's axial co-ordinate, as presented in Fig. 10b. When the temperature of steam reformer is 1073 K, the tar can be completely converted in the steam reformer, as shown in Fig. 10c. Due to the conversion of tar, CH4 , and C2 gases, the compositions of CO and H2 are increased. The simulated exit gas compositions of the steam reformer are compared with the experimental data (Caballero et al., 1997), as illustrated in Fig. 11, which shows that the simulation results are in good agreement with the experimental data. The comparison results indicate that the developed model is appropriate for simulating the catalytic steam reformer, in which the product gas containing tar is upgraded. 5.3. WSG membrane reactor The reactor model for the WGS reactor was verified in the previous work (Feng et al., 2007). Two classes of materials are used almost exclusively in industry as shift catalysts: the iron-based catalysts and the copper-based catalysts. The iron-based catalysts are the so-called high-temperature shift catalysts, the operating temperature is from about 320 to 450 ◦ C. Because the permeability of the H2 -membrane increases with the operation temperature, in this

588

P. Ji et al. / Chemical Engineering Science 64 (2009) 582 -- 592

1300

Temperature (K)

1200

6

4 Narvaez et al. (1996) Rensfelt and Ekstrom (1989) 2 950

1000

1050 1100 Temperature of bed (K)

1100

Reaction tube

1000 900 800

600

1150 40

Fig. 8. Low heating value of the product gas as a function of the bed temperature of gasifier. The equivalence ratio is 0.30. (Narváez et al., 1996; Maniatis et al., 1988; Rensfelt and Ekstrom, 1989).

1200 ER = 0.3 1000 Temperature (K)

Furnace

700

Maniatis et al.(1988) Simulation

Gas composition (% dry basis)

Low heating value (MJ/Nm3)

8

30

CO H2

20

10 CO2 CH4

0 25

600

20 Tar content (g/Nm3)

800

400 0.0

0.4

0.8 Height of gasifier (m)

1.2

Fig. 9. Temperature profile along the air-blown gasifier's axial coordinate.

work, the iron-based catalyst is chosen and the operation temperature of the reactor is set at 720 K.

15

10

5

0 0.0

0.2 0.6 0.4 0.8 Dimensionless steam reformer's axial coordinate (-)

1.0

5.4. Process analysis

Fig. 10. Profiles of temperature, gas composition and tar content as a function of the steam reformer's axial coordinate. The temperature of steam reformer is 1073 K.

The geometric parameters of the air-blown gasifier, steam reformer, and the WGS membrane reactor are listed in Table 9. The three reactors contribute to the H2 production of the integrated process. The overall thermodynamic efficiency of the integrated process is defined as

The pure hydrogen yield, YPure H2 , of the integrated process is defined as

=

ExH2 Ex1 + Ex2

(13)

where ExH2 is the exergy of the pure hydrogen produced, Ex1 is the exergy of the biomass fed into the gasifier, Ex2 is the exergy of the biomass generating the electricity required by the process, which is calculated based on the assumption that the thermodynamic efficiency of producing electricity from biomass is 60%. The exergy of biomass is calculated according to the method in the book (Szargut et al., 1998). The value 21 800 kJ/kg was taken as the chemical exergy per unit mass of biomass (Feng et al., 2004). To avoid carbon deposition in the steam reformer, the ratio of H2 O/C is adjusted to 3.0 (Seo et al., 2002) by mixing steam with the product gas from the air-blown gasifier.

YPure H2 =

pure H2 produced (mol) total biomass consumed (kg)

(14)

The pure H2 yield and the overall thermodynamic efficiency of the process are mainly influenced by three major variables, the ER, the temperature of gasifier, and the temperature of steam reformer. The effect of the temperature of gasifier on the performance of the integrated process is shown in Fig. 12. A higher temperature of the gasifier favors the biomass gasification to produce less tar and CH4 . However, to get a higher temperature of the gasifier, more external energy is required or more biomass is combusted. On the other hand, less tar and CH4 existing in the gas product from the gasifier reduces the energy required by the steam reformer for the conversion of tar and CH4 . Due to the combined effect of the two aspects, there is an

P. Ji et al. / Chemical Engineering Science 64 (2009) 582 -- 592

CH4

Exp.

Exp.

52

Cal.

Cal.

Gas composition

80%

60%

40%

20%

0% Case 1

Case 2

Case 3

Fig. 11. Comparison results between the calculated gas composition and the experimental data for the catalytic steam reformer. Inlet conditions, case 1−TSR = 1053 K, TFurnace = 1148 K, tar 260 (mg/Nm3 ); case 2: TSR = 1048 K, TFurnace = 1143 K, tar 1480 (mg/Nm3 ); case 3: TSR = 1103 K, TFurnace = 1183 K, tar 780 (mg/Nm3 ).

Pure H2 yield (mol/kg biomass)

Cal.

CO

H2

Thermodynamic Efficiency (%)

100%

Exp.

CO2

Table 9 Geometric parameters and conditions for the gasifier, steam reformer and WGS membrane reactor Gasifier Lbed Lfreeboard Dt dp Air

589

0.40 m 1.00 m 0.06 m 250 m 7.6 mol/h

a

TSR = 1023 K

50

48

58

TSR = 1023 K

b 56

54

990

960

1050 1020 1080 Temperature of gasifier (K)

1110

1140

Fig. 12. Effect of the temperature of gasifier on the pure hydrogen yield and on the overall thermodynamic efficiency. The equivalence ratio is 0.20. The temperature of steam reformer is 1023 K.

10

WGS membrane reactor LWGS DWGS Area of the H2 -membrane kw

H2

1.0 m 0.01 m 0.014 m 0.15 J/m s K 3:1

1.00 m 0.07 m 0.125 m2 0.15 J/m s K 5 m

8

Tar (g/Nm3)

Steam reformer LSR DSR Df kw H2 O/C

6 4 2 0

optimum operation temperature of the gasifier, at which the pure hydrogen yield and thermodynamic efficiency are maximized. The operation temperature of the steam reformer, TSR , is another influential parameter. A higher temperature of the steam reformer promotes the conversion of tar and CH4 , and more hydrogen can be produced. However, to obtain a higher temperature of the steam reformer, more H2 and CO have to be recycled to the furnace, and the pure hydrogen yield decreases. In order to get an optimum temperature of the steam reformer, the simulations of the steam reformer, furnace, and WGS membrane reactor are carried out dependently. The simulation results are shown in Figs. 13–15. Fig. 13 indicates that the tar conversion is increased with the temperature of steam reformer. Fig. 14 presents that the amount of hydrogen recycled to the furnace is increased with the temperature of steam reformer. Fig. 15 shows the biomass consumed for electricity generation. At higher steam reformer temperatures, more hydrogen is recycled back to the furnace of the reformer, and hence less electricity is needed to run the pump for maintaining the vacuum pressure, which is

980

1000 1020 1040 1060 Temperature of steam reformer (K)

1080

Fig. 13. Effect of the temperature of steam reformer on the tar conversion. The equivalence ratio is 0.20 and the temperature of gasifier is 1073 K.

required for the hydrogen separation through the H2 membrane. Therefore, the biomass for electricity generation decreases with the temperature of steam reformer. Fig. 16 shows that an optimum temperature of the steam reformer can be found, at which the pure hydrogen yield and thermodynamic efficiency are maximized. Eqs. (13) and (14) define the thermodynamic efficiency and the pure hydrogen yield, respectively. The calculation of the thermodynamic efficiency is based on the exergy of the hydrogen produced, as indicated by Eq. (13). The pure hydrogen yield and the thermodynamic efficiency follow the same trend. Fig. 16 also indicates that the ER influences the pure hydrogen yield and overall thermodynamic efficiency of the

590

P. Ji et al. / Chemical Engineering Science 64 (2009) 582 -- 592

55

0.3

H2 recycled to the SR furnace

0.2

0.1 CH4 recycled

CO recycled

Pure H2 yield (mol/kg biomass)

Molar ratio (mol/mol pure H2)

0.4

ER 0.20

50 0.25

45

0.30

0.35

40

0.40

35

110 Thermodynamic efficiency (%)

LHV of fuel gas recycled (kJ/mol pure H2)

0.0

100 90 80 70

0.20

55

0.25

50

0.30

45

0.35

40

0.40

60 35

980

1060 1000 1020 1040 Temperature of steam reformer (K)

0.006

0.005

0.004

0.003

0.002 980

1000 1020 1040 1060 Temperature of steam reformer (K)

990

1080

Fig. 14. Molar ratio of the recycled CH4 , H2 and CO to the pure H2 produced as a function of the temperature of steam reformer. The equivalence ratio is 0.20 and the temperature of gasifier 1073 K.

Biomass for electricity (kg/mol pure H2)

ER

60

1080

Fig. 15. Biomass consumed for generating electricity as a function of the temperature of steam reformer. The equivalence ratio is 0.20 and the temperature of gasifier 1073 K.

integrated process. A higher ER results in the combustion of more fuel gas and tar in the gasifier, therefore, the pure hydrogen yield and the overall thermodynamic efficiency decrease with the increasing of the ER. 6. Conclusions The non-isothermal models for the air-blown fluidized bed biomass gasifier and catalytic steam reformer have been developed.

1020 1050 Temperature of steam reformer (K)

1080

Fig. 16. Effect of the equivalence ratio and temperature of steam reformer on the pure hydrogen yield and on the overall thermodynamic efficiency. The temperature of gasifier is 1073 K.

The simulated gas composition, tar content, and low heating value of the product gas from the gasifier have been compared with experimental data, and good agreement has been obtained. The modeling of the catalytic steam reformer has been verified with experimental data. The simulation results show that the models developed are appropriate for the modeling of the fluidized bed biomass gasifier and the catalytic steam reformer. The simulation work is the basis for the thermodynamic analysis of the integrated process. The effects of the ER, temperature of the gasifier, and temperature of the steam reformer on the pure hydrogen yield and overall thermodynamic efficiency have been studied and discussed. The results indicate that, about 54 mol of pure hydrogen per kilogram biomass with an overall thermodynamic efficiency of about 58% can be achieved. Notation aH2 ar aw A Ar Ci CPi Cpg d db db0

H2 -membrane area per unit reactor length, m2 /m cross section area per unit reactor length, m2 area of the wall between reaction tube and the parallel furnace of the steam reformer, m2 /m cross section area of the gasifier, m2 Archimedes number concentration of different component, mol/m3 heat capacity of component i, J/mol K heat capacity of gas, J/kg K diameter, m bubble diameter, m initial bubble size, m

P. Ji et al. / Chemical Engineering Science 64 (2009) 582 -- 592

dbm dp D Dt Ex EA Ei fnet,i Fi hp Hbe k k0, i Kbc Kbe Kce L M N NR NHetero NHomo N H2 high

pH

2

plow H 2

Pm Q q

rj T Tinlet TP TR Ts ub ubr ug umf Vi Vi∗ X Ypure H2 z

Hj

Hnet

the limiting size of bubble expected in a bed, m diameter of the particle, m diffusion coefficient of component i in catalyst pellet, m2 /s diameter of the bed, m exergy, J/mol apparent activation energy of H2 membrane, J/mol apparent activation energy for component i, kJ/mol net flow of component i per volume of reactor, mol/(m3 s) molar flow rate of component i in the bubble phase, mol/s heat exchange coefficient between particle and the emulsion phase, J/(m2 s K) heat exchange coefficient between bubble and emulsion phase per unit volume, J/(m3 s K) thermal conductivity of gas, J/(m s K) the pre-exponential factor, 1/s exchange coefficient between bubble and cloud phase per unit volume of bubble phase, 1/s exchange coefficient between bubble and emulsion phase per unit volume of bubble phase, 1/s exchange coefficient between cloud and emulsion phase per unit volume of bubble phase, 1/s length, m molecular weight, g/mol number of components number of reactions number of heterogeneous reactions number of homogeneous reactions H2 permeation rate through the H2 membrane, mol/(m2 s) partial pressure of H2 in the reaction side, bar partial pressure of H2 in the non-reaction side, bar pre-exponential factor of the H2 membrane, mol m/(s m2 bar0.5 ) heat transferred from the gasifier's furnace, J heat flux between the reaction side and the nonreaction side, J/(m s)) reaction rate of j reaction, mol/(m3 s) temperature, K inlet temperature, K temperature of biomass particle, K standard temperature, K temperature of inert particles in the bubbling fluidized bed of the gasifier, K bubble velocity, m/s single bubble rise velocity, m/s gas velocity, m/s minimum fluidization velocity, m/s instantaneous yield of volatile matter for the gaseous component i, kg/kg biomass the ultimately attainable yield of the gaseous component i, kg/kg biomass mole fraction pure H2 yield, mol/(kg biomass) axial coordinate of the reactor, m heat of reaction j, J/mol sensible heat carried by the net flow, J/mol

Greek letters

s

H2 

specific particle surface area, m2 /m3 the thickness of the hydrogen membrane layer, m void fraction of the catalyst bed

mf  mix   i, j ab 

591

void fraction in the dense phase at minimum fluidization conditions viscosity, kg/m s viscosity of gas mixture, kg/m s density, g/m3 characteristic diameter of molecules, Å stoichiometric coefficient of component i of reaction j parameter for calculating the viscosity of gas mixture collision integral

Subscripts b B e f furnace SR WGS

bubble phase catalytic bed emulsion phase freeboard of the gasifier parallel furnace of the steam reformer steam reformer water-gas-shift reactor

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