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Performance evaluation for dual circulating fluidized-bed steam gasifier of biomass using quasi-equilibrium three-stage gasification model
Applied Energy 88 (2011) 5208–5220
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Performance evaluation for dual circulating fluidized-bed steam gasifier of biomass using quasi-equilibrium three-stage gasification model Son Ich Ngo a, Thanh D.B. Nguyen a,e, Young-Il Lim a,⇑, Byung-Ho Song b, Uen-Do Lee c, Young-Tai Choi c, Jae-Hun Song d a
Lab. FACS, RCCT, Department of Chemical Engineering, Hankyong National University, Gyonggi-do, Anseong-si, Jungangno 167, 456-749, Republic of Korea Department of Chemical Engineering, Kunsan National University, Gunsan, Jeonbuk 573-701, Republic of Korea High Temperature Processing R&D Department, Korea Institute of Industrial Technology (KITECH), Cheonan 331-825, Republic of Korea d R&D Center, SeenTec Co., Ltd., #608 Office Plazza, 7-2 Yongho-dong, Changwon, Gyeongnam 641-969, Republic of Korea e School of Chemical Engineering, Hanoi University of Science and Technology, No. 1 Dai Co Viet, Hanoi, Viet Nam b c
a r t i c l e
i n f o
Article history: Received 21 April 2011 Received in revised form 8 July 2011 Accepted 27 July 2011 Available online 20 August 2011 Keywords: Steam gasification Woodchips Dual circulating fluidized-bed (DFB) Solid circulation ratio Quasi-equilibrium three-stage gasification (qETG) model Parametric study
a b s t r a c t The effects of gasification temperature (TG) and steam to fuel ratio (c) on product gas composition and yield were experimentally investigated for steam gasification of pine woodchips in a bench-scale circulating fluidized-bed (CFB) gasifier with external heat supplier. To evaluate process performance in a dual circulating fluidized-bed (DFB) with heat carrier (silica sand), a quasi-equilibrium three-stage gasification (qETG) model was developed and validated with experimental data of biomass steam gasification. The model was divided into three stages including biomass pyrolysis, char–gas reactions, and gas-phase reactions. Carbonic and methane formation ratios were considered at the pyrolysis stage under the assumption of spontaneous decomposition. At the second and third stages, char–gas and gas-phase equilibrium reactions were corrected by two empirical equations concerning the steam participation ratio and the non-equilibrium factor, respectively. Using the qETG model, parametric study on TG and c was performed to predict final gas composition, carbon conversion, char residue, gas yield, lower heating value, additional fuel ratio, solid circulation ratio, heat recovery and H2 to CO molar ratio. Focusing on the solid circulation ratio and H2/CO molar ratio, several effective operating conditions were suggested from the contour of performance criteria. Ó 2011 Elsevier Ltd. All rights reserved.
1. Introduction As we know, there are nine general sources of energy on the Earth: Solar, biomass, wind, wave, hydro, tidal, geothermal, nuclear, and fossil. Except for the last three, the remaining six are generally called renewable sources of energy, as they are not depleted with time [1]. Biomass sources in the plant form store solar energy as chemical energy during their growth by photosynthesis. This storage of chemical energy can be released and converted to many other useful forms of energy such as heat, electricity, light, and automotive power through some upgrading and conversion systems [2]. In recent years, much attention has been paid to the thermochemical conversion from biomass since: (1) the use of biomass for energy has the potential to control the net emission of CO2 to the atmosphere [2] and (2) it can reduce the dependence of fossil sources. Among the thermochemical conversion processes such as gasification, pyrolysis and combustion, gasification has been ⇑ Corresponding author. Tel.: +82 31 670 5207; fax: +82 31 670 5445. E-mail address: [email protected] (Y.-I. Lim). 0306-2619/$ - see front matter Ó 2011 Elsevier Ltd. All rights reserved. doi:10.1016/j.apenergy.2011.07.046
considered to be the most cost effective process since it converts biomass into clean combustible gaseous products [3]. Steam gasification has become an area of growing interest, since it produces a gaseous fuel with relatively higher hydrogen content, which could be used in fuel cells as a new recognizable technology for the future to produce power in a much cleaner manner. It also has other advantages in comparison with air or air–steam gasification such as, (1) producing a gas with higher heating value, (2) reducing the dilution effect of nitrogen from air, and (3) eliminating the need for an expensive oxygen plant when both air and oxygen are used as gasification reagents [4]. However, steam gasification is a more complex process, because the heat necessary to the process is not directly supplied by the partial combustion of the feedstock during the gasification process [4]. Among several types of gasifiers such as fixed bed, down- and updraft, and fluidized-bed, the fluidized-bed has been known as a promising alternative , thereby maximizing the gas product yield due to the uniform temperature distribution and high mass and heat transfer rates [5]. The dual circulating fluidized-bed (DFB) system has many advantages that are capable of scaling up and
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Nomenclatures A empirical parameters (–) B empirical parameters (–) cp molar heat capacity (kJ/kmol/K) Cp heat capacity (kJ/kg/K) DG0T G Gibbs free energy change of reaction at 25 °C (kJ/kmol) DG0f standard Gibbs energy of formation at 25 °C (kJ/kmol) h molar enthalpy (kJ/kmol) Hbiomass sensible heat of biomass fed to gasifier (kJ/h) Hsens sensible heat of unreacted char leaving gasifier (kJ/h) char Hdemand heat demand (kJ/h) heat entering into gasifier (kJ/h) Hg,in Hg,out heat leaving gasifier (kJ/h) Hloss heat loss through surrounding environment (kJ/h) Hcomb combustion heat of syngas (kJ/h) product sensible heat of syngas leaving gasifier (kJ/h) Hsens product Hsteam heat of steam fed to gasifier (kJ/h) DH0f standard enthalpy of formation at 25 °C (kJ/kmol) K equilibrium constant (–) K corrected equilibrium constant (–) Lwater latent heat of water (kJ/kg/h) LHVbiomass lower heating value of biomass (kJ/kg) LHVproduct lower heating value of producer gas (MJ/Nm3) LHVchar lower heating value of char (kJ/kg) Madd amount of additional fuel to combustor (kg/h) Mbiomass biomass feed rate (kg/h) amount of carbon in biomass (kg/h) MC MC,in amount of fixed carbon in biomass (kg/h) Mchar amount of the unreacted char (kg/h) MC,convert amount of carbon converted to producer gas (kg/h) MC,unreact amount of unreacted char (kg/h) Msand amount of recirculated sand (kg/h) n number of moles (kmol/h)
development. However, it is also known as a complex system in construction and installation [6]. Numerous studies for biomass gasification by means of modeling and simulation include (1) thermodynamic equilibrium models [7– 10], (2) kinetic rate models [1,11–13], and (3) neural network models [14,15]. In the kinetic rate models, initial conditions and kinetic parameters are not well known because of a variety of feedstock [11]. The neural network models as a kind of black-box models have achieved high prediction accuracy. However, it is hard to obtain physical meaning from these models, and the scale-up and adaption abilities of the neural network models are restricted. The kinetic models predict the progress of product composition with respect to the residence time in a gasifier, whereas the equilibrium models provide the maximum yield of a desired product which is achievable from a gasification system [16]. Although kinetic rate models are considered as a rigorous approach, equilibrium models are valuable because they can predict thermodynamic limits which are used to design, evaluate and improve the process [17]. The equilibrium models have been used for preliminary study on the influence of the most important process parameters. Several equilibrium models have been proposed for coal gasification [3,18], but these models cannot directly be applied to biomass due to significant differences in chemical and physical properties between coal and biomass. In biomass gasification, the pyrolysis step is one of the most important steps for gasification, since it strongly affects the final gas product composition [19–22]. The knowledge of devolatilization in the pyrolysis step is crucial for a precise prediction. Therefore, the more detailed the pyrolysis is taken into account, the better the prediction will be.
p Vproduct radd rcir R Rt Tambient Tfeed TG Tpyrol Triser T in steam Vyield x y z
empirical polynomial parameter (–) volumetric flow rate of product gas (Nm3/h, dry basis) additional fuel ratio (kg/kg) solid circulation ratio (kg/kg) gas constant (kJ/kmol/K) retention time (min) ambient temperature (K or °C) temperature of biomass fed to gasifier (K or °C) gasification temperature (K or °C) pyrolysis temperature (K or °C) riser temperature (K or °C) inlet temperature of steam (K or °C) gas yield (Nm3/kg biomass) mass fraction (–) mole fraction (–) empirical variable (1/K)
Greek letters c steam to fuel ratio (kg/kg) e reaction coordinate (kmol) j non-equilibrium factor (–) ge heat recovery (–) / ratio of molar fractions(–) /C carbon conversion (–) char residue (–) /CR m stoichiometric number (–) a species involved reactions v total stoichiometric number (–) Subscripts i species index j reaction index
In this study, we investigate the biomass gasification with the steam agent in a bench-scale circulating fluidized-bed (CFB) gasifier and develop a quasi-equilibrium three-stage gasification (qETG) model for the prediction of process performance in dual circulating fluidized-bed (DFB). The qETG model is divided into three main stages: (1) pyrolysis of volatiles in biomass, (2) solid–gas reactions between biomass char and gasifying reagents (carbon dioxide or steam) in the fluidized bed, and (3) gas-phase reactions among the gaseous species in the free board of the gasifier. At each stage, empirical models are established based on the experimental data to calculate the gaseous components. Especially, the deviation from equilibrium reaction is taken into account in the third stage by a non-equilibrium factor. The model is first validated by the experiment data conducted in the bench-scale CFB gasifier with pine woodchips, and the data taken from the literature [17,23]. Then the contour of performance criteria (i.e., the solid circulation ratio, the H2/CO molar ratio, the additional fuel ratio, and the heat recovery) is obtained from the parametric study of TG and c using the qETG model. Focusing on the solid circulation ratio and H2/CO molar ratio for Fisher–Tropsch (FT) synthesis, several effective operating conditions were proposed for a 100 MW DFB biomass gasifier.
2. Experiment Experiments were performed in a bench-scale CFB gasifier (0.15 m I.D. 3.0 m height) operating at atmospheric pressure. The schematic diagram of the experimental set-up is shown in Fig. 1. The gasification system consists of a CFB gasifier with
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Fig. 1. Schematic diagram of experimental set-up for bench-scale CFB gasifier.
Table 1 Proximate and ultimate analyses of pine woodchips used in this study. Proximate analysis (wt.%, wet basis)
Ultimate analysis (wt.%, dry basis)
H2O Volatile Fixed carbon Ash
6.4 75.9 17.4 0.3
C H O N S Cl
Total
100.0
50.8 5.4 43.6 Not detected 6.2103 Not detected 99.8
electrical heaters, a screw feeder for continuous supply of the feedstock, two cyclones, and an impinger box for collecting fly ash, a tar trap, and a gas analyzer. Pine woodchips of sizes under 10 mm were used as a feedstock. The proximate and ultimate analyses of the woodchips are listed in Table 1. It is addressed that the biomass material contained large volatiles (76%) and low ash content (0.3%) in comparison with common coal materials [24,25], and a little sulfur was detected from the ultimate analysis. The screw feeder with a feed rate of about 5 kg/h continuously supplied the feedstock. The gasifier temperature was controlled by electrical heaters, which could be heated up to 1300 °C. Steam was used as a fluidized and gasifying agent. Superheated steam at 400 °C was introduced to the gasifier to remove condensable components. The on-line gas analyzer (Model: AO2020, ND-IR analyzer, ABB Co. Ltd., Germany) monitored the composition of syngas. Gas chromatography (Model: 7890A, Agilent Technologies, USA) was employed for the quantitative analysis of sulfur compounds, where highly sensitive Flame Photometric Detector (FPD) was generating a flame with hydrogen and oxygen. The flow rates of hydrogen and oxygen were 100, and 60 ml/min, respectively. The inlet and detector temperatures were 250 °C. The PLOT capillary column
(GC-gaspro, Agilent, I.D. = 0.32 mm, L = 5 m) was used since it is suitable for light molecular weight hydrocarbon compounds or sulfur compounds. For the identification of the sample gas, the area and retention time (Rt) was calibrated by standard gases of COS (49.6 ppm) and H2S (177.6 ppm) with N2 balanced. Helium was used as a carrier gas, and the oven temperature was maintained at 50 °C. The COS was detected after 7.0 min and the H2S was detected after 8.1 min. Temperature was measured at the six locations of which the third one from the bottom (T3 in Fig. 1) was considered as the gasification temperature (TG). Steam was injected from the bottom of the gasifier and the steam to fuel ratio (c) was adjusted by the steam flow rate. The two operating conditions ranged 700 6 TG 6 900 °C and 0.3 6 c 6 1.0 kg/kg. At TG = 800 °C and c = 0.3 (kg/kg), Fig. 2a shows temperature profiles of the six measuring points with respect to the running time of biomass feeding. The temperatures slightly dropped down just after feeding. The temperatures and the product gas composition were stabilized after 5 min (see Fig. 2). The gasification temperature (TG) measured at T3 was controlled by manipulating the electrical heater. The syngas composition measured after 30 min of the running time was used for the final one. 3. Quasi-equilibrium three-stage gasification (qETG) model and process performance evaluation A quasi-equilibrium three-stage gasification (qETG) model involving four empirical equations is developed to investigate the influence of two operating conditions, gasification temperature (TG) and steam to fuel ratio (c), on process performance of a 100 MW (20 tons/h) DFB gasifier. The parameters of the four empirical equations are estimated from experimental data published in the literature [18,23,26,27]. The empirical equations are
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(a)
900
o
Temperature ( C)
1000
800 700 T1 T2 T3 T4 T5 T6
600 0
Gas composition (vol%, dry basis)
50 H2 CO 2 CH 4
(b) 40
CO
30
20
10
0 0
5
10
15
20
25
30
Time (min) Fig. 2. Experimental data in the bench-scale CFB performed at TG = 800 °C and c = 0.3 kg/kg: (a) temperature profile, and (b) syngas composition.
used to calculate the gaseous components at each stage. Table 2 shows main reactions and products in the three-stage gasification. Three char–gas reactions and a water–gas shift reaction are taken into account in this model. The qETG model is validated with experimental data obtained from the bench-scale CFB as well as the other data from the sources. Parametric study on TG and c is performed to predict final gas composition (yi), carbon conversion (/C), char residue (/CR), lower heating value (LHVproduct), heat recovery (e), gas yield (Vyield), additional fuel ratio (radd), and solid circulation ratio (rcir). Fig. 3 summarizes the structure of qETG model and the scope of parametric study. 3.1. Quasi-equilibrium three-stage gasification (qETG) model The three stages of gasification are divided into biomass pyrolysis, char–gas reactions, and gas-phase reactions. At the pyrolysis stage under the assumption that volatiles and tar in biomass are
completely and spontaneously converted into five gaseous components (CO, CO2, CH4, H2, and H2O), the carbonic formation ratio (/CO=CO2 ) and the methane formation ratio (/CH4 =H2 ) determine gas composition after pyrolysis by the elemental balance (see Appendix A.1). The two empirical equations in the first stage (/CO=CO2 and /CH4 =H2 ) were derived on the basis of experimental data [26] conducted at a wide range of temperatures (400–900 °C), using three kinds of biomass (woodchip, coconut shell and straw). Fig. 4 illustrates the dependency of the two ratios on pyrolysis temperature (Tpyrol). As shown in Fig. 2b, the temperatures are distributed inside the gasifier. Since pyrolysis and decomposition of biomass depend on its pore size distribution, heating rate [28] and residence time, it is not easy to determine an appropriate Tpyrol in the range of 700 6 TG 6 900 °C. Thus, Tpyrol was assumed to be the same as TG in our study. Char–gas and gas-phase reactions are supposed to reach an equilibrium state. At the second stage, the steam participation ratio (b) is used to determine the amount of steam involved in the char– gas reactions with respect to TG [3,18]. At the third stage, a nonequilibrium factor (j) is employed to account for the deviation from thermodynamic equilibrium of the water–gas shift reaction [23,27,29]. Due to b and j, the qETG model compensates for a deviation from equilibrium reaction according to the gasification temperature (see Appendices A.2 and A.3). The steam reforming reaction (CH4 + H2O CO + 3H2) is assumed not to take place in the third stage (see Table 2). This reaction is not favorable at temperatures over 800 °C, since the Gibbs free energy of formation has a positive value [30]. Moreover, Dupont et al. concluded from a modeling approach that the steam reforming of CH4 was kinetically limited [22] in those temperatures. Therefore, only one reaction for water–gas shift is considered in this stage. Fig. 5 shows the temperature dependency of the thermodynamic equilibrium constant (K4, see Appendix A.4), and the experimental equilibrium constant (K4,exp = (yCO2 yH2 Þ=ðyCO yH2 O )) for the water–gas shift reaction. The molar fraction ratios (K4,exp) between reactants and products obtained from experimental value of Wei et al. [27] and Herguido et al. [23] are far from the thermodynamic equilibrium constants (K4). The former [27] was measured in a down-draft gasifier and the latter [23] was in a fluidized bed which is closer to K4 than the former. To compensate this discrepancy, the non-equilibrium factor (j) has been used elsewhere [29,31]. Applying the non-equilibrium factor (j) to K4, the modified equilibrium constant (K 4 ¼ j K 4 ) follows the trends of experimental data. The non-equilibrium factor (j) may depend on feedstock, operating conditions, and the type of gasifiers. In this study, j was determined so that our experimental syngas compositions measured for CO, CO2, and H2 were predicted as best as possible by the qETG model. In Table 3, the fifth-order polynomial of j is suggested as the best fitting
Table 2 Reactions and products assumed in the qETG model. Stage
Reactions
Products
Components assumed in qETG
Refs.
Pyrolysis
First step: thermal decomposition Second step: tar cracking
CO, CO2, CH4, H2 and H2O
CO, CO2, CH4, H2 and H2O
[19,20,45,48]
Char unreacted, CO, H2, H2O-residue
[3,18]
CO, CO2, H2, H2O
[23,27,30,38,47]
Solid–gas reactions
Water–gas shift reaction
C(s) + CO2(g) M 2CO(g)
CO, CO2, H2, heavier hydrocarbon (e. g., C2H6, C2H4, and C3H6), and inert tar. (Char unreacted) CO, CO2, H2, (H2O-residue)
C(s) + H2O(g) M CO(g) + H2(g) C(s) + 2H2O(g) M CO2(g) + 2H2(g) CO(g) + H2O(g) M CO2(g) + H2(g)
CO, CO2, H2, H2O
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β
κ
Fig. 3. Structure of the qETG model and scope of parametric study.
15
2
(a)
Woodchip [26] Straw [26]
1.6
Empirical model
1.4
5
K 4*
κ
1.2
value
φ
CO/CO2
10
exp. [23] exp. [27] K4
1.8
Coconut shell [26]
0 400
500
600
700
T
800
0.8
900
o
pyrol
1
( C)
0.6 0.4
15
(b)
0.2
Woodchip [26]
0 700
Coconut shell [26] 10
φ
CH4/H2
Straw [26]
800
850
900
o
TG ( C)
Empirical model
Fig. 5. Temperature dependency of experimental (K4,exp), thermodynamic (K4), and corrected (K 4 ) equilibrium constants, and non-equilibrium factor (j).
5
0 400
750
500
600
T
700 pyrol
800
900
(oC)
Fig. 4. Ratios of CO/CO2 (/CO=CO2 ) and CH4/H2 (/CH4 =H2 ) versus pyrolysis temperature (Tpyrol): (a) /CO=CO2 , and (b) /CH4 =H2 .
function, and it converges to the thermodynamic equilibrium constant as the temperature increases. It is assumed that these four empirical equations only depend on the gasification temperature (TG), as indicated in Table 3. A detailed description of the qETG model is given in the Appendix.
3.2. Performance criteria
gas(LHVproduct), (5) additional fuel ratio (radd), (6) solid circulation ratio (rcir), (7) heat recovery (ge), and (8) H2 to CO molar ratio (/H2 =CO ). Carbon conversion (/C, dimensionless) is defined as the mass fraction of carbon content (MC,convert, kg/h) converted from the product gas by pyrolysis and gasification to the mass of carbon in biomass (MC, kg/h).
/C ¼
ð1Þ
Char residue (/CR, dimensionless) is calculated by the mass ratio of unreacted carbon after char–gas reactions (MC,unreact, kg/h) to the amount of fixed carbon at the initial condition (MC,in, kg/h). /CR ¼ ¼
The main purpose of this study is to evaluate process performance, when the targeting feedstock (i.e., pine woodchips) with a constant feed rate (Mbiomass = 20,000 kg/h) will be used for biomass steam gasification in a DFB gasifier. The following criteria for process performance are investigated: (1) Carbon conversion(/C), (2) char residue (/CR), (3) gas yield (Vyield), (4) lower heating value of product
M C;conv ert carbon in CO; CO2 and CH4 after 2nd stage ¼ carbon in biomass MC
M C;unreact MC;in ðM C;in þ carbon in gas after 1st stageÞ ðcarbon in gas after 2nd stageÞ M C;in ð2Þ
The heat released by combustion of product gas calculated as follows [32]:
Hcomb product ¼ 285:63nH2 þ 282:99nCO þ 890:3nCH4
(Hcomb product ,
MJ/h) is
ð3Þ
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Carbonic formation ratio
B1 /CO=CO2 ¼ nnCOCO ¼ A1 exp T pyrol
A1 = 4.70 103
[48]
2
3
Methane formation ratio Char–gas reactions
Steam participation
/CH4 =H2 ¼
nCH4 nH2
B2 ¼ A2 exp T pyrol
B1 = 7.16 10 K A2 = 2.30 103 B2 = 5.40 103 K A3 = 5.14 101
nH O;0 b ¼ nH 2O;tot ¼ A3 exp TBG3
[26] [3]
2
B3 = 7.54 103 K Water–gas shift reaction
Non-equilibrium factor
j = p1 z5 + p2 z4 + p3 z3 + p4 z2 + p5 z + p6
[23]
p1 = - 3.2 103; p2 = - 7.8 103; p3 = 2.3 103; p4 = 7.4 102; p5 = 3.0 101; p6 = 4.0 101;
[27]
1070 z ¼ T G 62:0
where nH2 ; nCO and nCH4 (kmol/h) are the mole numbers of combustible product gases. The lower heating value of the final product gas, LHVproduct (MJ/m3), is then calculated as:
LHV product ¼
Hcomb product V product
ð4Þ
V yield
ð5Þ
The lower heating value of biomass (LHVbiomas, kJ/kg) is calculated by the Boie formula [10]:
LHV biomass ¼ 34; 835xC þ 93; 870xH 10; 800xO þ 6280xN þ 10; 465xS
ð6Þ
xC, xH, xO, xN, and xs are the mass fraction of the elements C, H, O, N, and S in biomass, respectively.In the DFB process, the heat for the gasification reaction taking place in the gasifier (Hdemand) is provided by the combustion heat of the unconverted char and additional fuel (if needed) in the riser. The silica sands play a key role as a heating carrier from the combustor to the gasifier. It is assumed in this study that the sand particles absorb all the heat released from the combustion. The heat demand of the gasifier may be expressed as:
Hdemand ¼ Hg;out Hg;in
ð7Þ
where Hg,out (kJ/h) is the heat leaving the gasifier and Hg,in (kJ/h) is the heat of the feed entering into the gasifier. The total inlet heat (Hg,in, kJ/h) is calculated as:
Hg;in ¼ Mbiomass LHV biomass þ Hsens biomass þ H steam Hsens biomas
where conditions:
ð9Þ
C p;biomass was assumed as 1.0 kJ/kg/K for woodchips [33]. The heat of steam (Hsteam, kJ/h) includes sensible heat and latent heat caused by the vaporization of water:
ð10Þ
The total outlet heat (Hg,out, kJ/h) is obtained as follows: sens sens Hg;out ¼ 1000ðHcomb product Þ þ H product þ M char LHV char þ H char
þ Hloss
ni cpi ðT G Þ ðT G T ambient Þ þ nH2 O cp;H2 O ðT G Þ
T ambient ÞÞ
ð11Þ
where LHVchar (kJ/kg) and Mchar (kg/h) are the lower heating value and the mass flow rate of the unreacted char, respectively. Hsens product
ð12Þ
where ni (kmol/h) from the final gas product and cpi ðT G Þ (kJ/kmol/K) are the mole number and the heat capacities at TG (K) of CO, CO2, H2, and CH4, respectively. nH2 O ; cp;v apor and Cp,water are the mole number of H2O, the molar heat capacity of vapor and the heat capacity of water, respectively. Sensible heat of char (considered as pure carbon) is expressed as:
Hsens char ¼
M C;unreact cp;char ðT G Þ ðT G T ambient Þ 12
ð13Þ
where the unreacted char (MC,unreact) is given in the numerator of Eq. (2) and cp;char ðT G Þ is the molar heat capacity of char at TG (K). The molar heat capacity of the pure components in high temperatures is expressed in terms of fourth order polynomials (NIST webbook [34]) in Table 4. With low temperatures, the thermodynamic property was estimated as a constant. Since the ambient temperature (Tambient) is equal to the biomass feed temperature (Tfeed), the sensible heat of biomass (Hsens biomas ) is zero in Eq. (9). The heat loss of DFB gasifiers reported in the literature [35] was about 0–30% of the heat value of biomass fed to the gasifier. In the qETG model, the heat loss (Hloss) from the gasifier is assumed to be 10% (Hloss = 0.1 Mbiomass LHVbiomass). The circulation rate of silica sand (Msand, kg/h) is obtained from the heat consumption of the gasification.
Msand ¼
1:1 Hdemand C p;sand ðT riser T G Þ
ð14Þ
where Cp,sand is the heat capacity of silica sand (Cp,sand = 0.64 kJ/kg/K [36]). Triser is the temperature in the riser, which is assumed to be 950 °C in DFB. The solid circulation ratio (rcir) is given as the ratio of the circulation rate of the sand particles to the fuel feed rate (Mbiomass):
rcir ¼
Hsteam ¼ M biomass c ðC p;water ð373 T ambient Þ þ Lwater þ C p;v apor ðT in steam 373ÞÞ
X
ðT G 373Þ þ 18nH2 O ðLwater þ C p;water ð373
ð8Þ
(kJ/h) is the sensible heats of biomass at the inlet
Hsens biomas ¼ M biomass C p;biomass ðT feed T ambient Þ
Hsens product ¼
i
where Vproduct (Nm3/h) is the volume of final product gas on the dry basis. The yield of the final gas product (Vyield, Nm3/kg biomass) is determined as the volume of the product gas (Vproduct) obtained when 1 kg of biomass is gasified.
V product ¼ M biomass
and Hsens char (kJ/h) are the sensible heats of the product gas and the unreacted char, respectively. The sensible heat of CO, CO2, H2, CH4 and H2O at the out stream was computed by the enthalpy formulations:
M sand Mbiomass
ð15Þ
The larger the rcir, the more the energy consumption for the bed particles circulation between the combustor and gasifier, as well as the higher the breakage and attrition of bed particles in the combustor [9]. However, Seo et al. [37] suggested to maintain the solid circulation rate over a certain amount for stable transportation in the loop-seal of the DFB gasifier. The additional fuel (Madd) would be supplied if an additional heat were required for endothermic gasification reaction:
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Table 4 Operating conditions and thermodynamic properties for pure components used in heat balance. Operating conditions Tambient T in steam Tfeed
Constants 298 K (25 °C) 673 K (400 °C)
Lwater C p;water
2260 (kJ/kg/h) 4.1813 (kJ/kg/K)
298 K (25 °C)
C p;v apor
2.080 (kJ/kg/K)
973–1173 K (700–900 °C)
TG
Fourth order polynomial equations of molar heat capacity (kJ/kmol/K) cp,CO 29:7 6:50 103 T G þ 1:83 105 T 2G 9:39 109 T 3G þ 1:08 1012 T 4G cp;CO2 29:3 2:24 102 T G þ 2:65 104 T 2G 4:15 107 T 3G þ 2:01 1010 T 4G cp;H2 27:0 þ 1:19 102 T G 2:41 105 T 2G þ 2:15 108 T 3G 6:15 1012 T 4G cp;CH4 36:2 5:11 102 T G þ 2:21 104 T 2 1:82 107 T 3 þ 4:90 1011 T 4 G
G
G
cp;H2 O
33:8 5:95 103 T G þ 2:24 105 T 2G 9:96 109 T 3G þ 1:10 1012 T 4G
cp,char
27:0 þ 4:58 104 T G 4:5310 2:1810 þ 8:0010 7:2110 TG T2 T3 T4
3
6
8
G
M add ¼
Hdemand 1:12 Hsens char M char LHV char LHV biomass
G
50
where the heat loss due to the transportation of sand particles from the riser to the gasifier is supposed to be 12% of Hdemand. The additional fuel ratio (radd, dimensionless) is defined as the mass ratio of the additional fuel (Madd) to the feed rate (Mbiomass):
45
Madd Mbiomass
ð17Þ
The heat recovery (ge, dimensionless) is expressed as the ratio of the heat of the product gas (LHVproduct) to that of the total biomass fed to the system:
ge ¼
V yield LHV product ð1 þ r add Þ LHV biomass
ð18Þ
Note that the units of LHVproduct and LHVbiomas are different and Vyield multiplied in Eq. (18) makes ge dimensionless. 4. Results and discussion In this section, the proposed model (qETG model) is validated with the experimental data measured in the bench-scale CFB gasifier as well as the data taken from the published works [17,23]. The effects of the gasification temperature (TG) and the steam to fuel ratio (c) on the DFB gasifier performances are examined by means of their parametric study. Effective operating conditions are proposed for the production of syngas suitable for Fisher–Tropsch (FT) synthesis. 4.1. Model verification In Fig. 6, three experiment sets operated at c = 0.3 kg/kg and TG = 700, 800, and 900 °C are compared to the qETG model results. The H2 and CO2 concentrations increase with the rise in temperature, while the concentrations of CH4 and CO show an opposite trend. Since the water–gas shift reaction (CO + H2O M CO2 + H2) proceeds forward at the temperatures above 700 °C in the presence of steam [4,38], the increase in H2 and the decrease in CO are expected from the reaction when TG increases. However, for a higher temperature range, the water–gas shift reaction has been thought to be less important and the Boudouard reaction (C + CO2 M 2CO) and the primary and secondary water gas reactions (C + H2O M CO + H2, and C + 2H2O M CO2 + 2H2) have a more prevailing role [4]. The CO2 formed by the secondary water gas reaction could be consumed by the Boudouard reaction, which could explain a slight increase observed in CO2 concentration when the gasification temperature rises. The concentration of CH4 produced mainly in the pyrolysis stage decreases with TG (see Fig. 4). The model results of
final gas composition (%vol)
ð16Þ
r add ¼
10
G
40 35 30 25 20 15 10 5
700
750
800
850
900
o
TG ( C) CO-Model value H2 -Model value
CO-Exp.value H2 -Exp.value
CO2 -Model value
CO2 -Exp.value
CH4 -Model value
CH4 -Exp.value
Fig. 6. Effect of gasification temperature (TG) on final gas composition obtained from qETG model in comparison with experiment data at c = 0.3 (kg/kg).
the final gas composition show a similar tendency to the experimental data, as seen in Fig. 6. Table 5 gives the comparison of H2/CO molar ratio (/H2 =CO ), gas yield (Vyield), and lower heating value (LHVproduct) between the experiment and the qETG model at c = 0.3 kg/kg. /H2 =CO increases with the increase of temperature, as expected in Fig. 6. The error of /H2 =CO between the experiment and the model increases with temperature but is limited to 10%. It is experimentally observed that Vyield increases with an increase of TG from 700 °C to 800 °C, since the endothermic reactions of tar cracking and char gasification (char–CO2 and char–steam) proceed forward when TG increases [27]. However Vyield at TG = 900 °C is slightly less than that obtained at TG = 800 °C. It might be attributed to the tar formation at elevated temperatures [39]. Furthermore, the gasification with pure steam has been known to produce a large content of tar (30–80 g/m3), compared to the gasification with steam–O2 mixture and air agents (4–30 and 2–20 g/ m3, respectively) [40]. In the qETG model where tar formation is not considered, the syngas yield increases with temperature. The error of Vyield is relatively big at 700 °C where the gasification reactions are far from equilibrium [23,27,29]. The lower heating value (LHVproduct, MJ/Nm3) decreases a little with the increase of temperature because the volume of product gas (Vproduct, Nm3/h) increases
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S.I. Ngo et al. / Applied Energy 88 (2011) 5208–5220 Table 5 Comparison of H2/CO molar ratio (/H2 =CO ), gas yield (Vyield), and lower heating value (LHVproduct) between experiment and qETG model at c = 0.3 kg/kg. TG (°C)
700 Exp. 0.72 0.35
/H2 =CO ðÞ 3
V yield
Nm kg
LHV product
MJ m3
12.4
Model 0.73 0.80 15.9
Error (%) 1.4 128 28
Exp. 1.16 1.36 12.0
900 a
Model
Error (%)
1.25 1.12
Exp.
1.50 1.3
11.7
11.7
Errora (%)
Model
1.66 1.1
7.8 17.6
13.4
9.6 18.2
12.4
6.0
valueModel prediction Error ð%Þ ¼ Experiment 100. Experiment value
relatively much with the rise of TG (see Vyield in Table 5), even though the methane production rate (nCH4 ) released from pyrolysis increases slightly with TG (see Eqs. (3) and (4)). The error of LHVproduct between the experiment and model is restricted to about 30%. To validate the qETG model in the wide range of c, experiment data taken from Karmakar and Datta [17] and Herguido et al. [23] are utilized. Figs. 7 and 8 show the effect of c on the final gas composition, comparing the qETG model values with the experimental data. Here, the two empirical equations of the pyrolysis stage are slightly modified because different feed materials are employed. The averages of root mean square errors (RMS) for the two comparisons between the model and experiment are 3.97% and 5.18%, respectively. The RMS was calculated by the following equation:
sffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi PN 2 k ðexperiment k Modelk Þ RMS ¼ N
60
final gas composition (%vol)
a
800 a
50
40
30
20
10
ð19Þ 0
where N is the number of data points. As observed from the two figures, an increase in c leads to the increase of H2 and CO2 concentrations, while CO and CH4 decrease [1,4,7,27]. These behaviors were the same as in our experimental observation. The final gas product composition in biomass gasification with respect to c is mainly influenced by the water–gas shift reaction [8,10,23,38]. As shown in Figs. 6–8, the qETG model predicts the tendency of the syngas composition with respect to TG and c reasonably well. Using the qETG model, the process performance of a 100 MW DFB gasifier is evaluated by parametric study on TG and c. 4.2. Parametric study Fig. 9 shows the effects of TG and c on (a) carbon conversion, (b) char residue, (c) gas yield and (d) lower heating value of the final dry gas product. The increase of TG promotes the pyrolysis and the char–gas reactions, then the carbon conversion (/C) increases, and consequently the gas yield (Vyield) increases. However, since CO2 increases, CH4 decreases (see Fig. 6), and Vyield increases (see Fig. 9c) as TG increases, the lower heating value of product (LHVproduct) decreases (see Eq. (4)), as is depicted in Fig 9d. Contrasting with the variation of carbon conversion, the char residue (/CR) decreases with an increase in T (see Eq. (2)). As the steam to fuel ratio (c) increases, the gasification reactions are also promoted (see Table 2). That results in the increase of carbon conversion (/C), gas yield (Vyield) and the decrease of char residue (/CR). The carbon conversion reaches to 1.0 at TG = 900 °C and c = 3.0, as shown in Fig. 9a, which means that all the carbon in biomass is converted into gas species through the first and second stages. By contrast, the char residue (Fig. 9b) tends to be zero. When both TG and c are low, the char residue shows a maximum value (=1.0). In this area, the char–gas reactions do not take place and only pyrolysis and a water–gas shift reaction occur to a limited extent. Fig. 10 illustrates the parametric study of TG and c for (a) additional fuel ratio (radd), (b) solid circulation ratio (rcir), (c) heat
0.6
0.8
1
1.2
1.4
1.6
1.8
γ (kg/kg) CO-Model value
CO-Exp.value [17]
H2 -Model value
H 2 -Exp.value [17]
CO2 -Model value
CO2 -Exp.value [17]
CH4 -Model value
CH4 -Exp.value [17]
Fig. 7. Effect of steam to fuel ratio (c) on final gas composition obtained from qETG model in comparison with experiment data taken from Karmakar and Datta [17] at TG = 750 °C.
recovery (ge), and (d) H2/CO molar ratio (/H2 =CO ). Since the rising of TG and c requires a higher heat demand (see Eqs. (11)–(13)) and leads to a smaller char residue (see Fig. 9b), the additional fuel ratio (radd) and solid circulation ratio (rcir) have to increase, as shown in Fig. 10a and b. When TG and c are over 780 °C and over 0.6, respectively, the additional fuel is required. The solid circulation ratio (rcir) is unrealistically too big at high values of TG and c. At a higher gasification temperature, the temperature difference between the riser and gasification becomes lower and the solid circulation (Msand) increases (see Eq. (14)). Furthermore, the higher steam amount and temperature need the higher energy demand (Hdemand) which is provided by hot silica sand to maintain the endothermic gasification reactions. Consequently, the solid circulation ratio (rcir) increases. It is desired for rcir to keep as low as possible, as mentioned above. The circulation ratio of bed particles was reported at about 15–20 [9] and about 30–80 times of the biomass dry feed rate [10,35,41]. Therefore the rcir over 100 may be avoided for practical operations, for which the gasification temperature is desirable to be below 850 °C, as shown in Fig. 10b. As seen in Eq. (18), the heat recovery (ge) was defined as a function of Vyield, LHVproduct, and radd. As the rises of both TG and c makes Vyield and radd increase whereas LHVproduct decreases, the conflicting behavior displays a maximum area of heat recovery (ge = 0.775) in the range of 0.0 6 c 6 0.1 and 850 6 TG 6 870 °C, as illustrated in Fig. 10c. The heat recovery tends to decrease with the increase of
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S.I. Ngo et al. / Applied Energy 88 (2011) 5208–5220
4.3. Effective operations for DFB steam gasifier of pine woodchips
final gas composition (%vol)
60
For FT synthesis, the common H2/CO ratio (/H2 =CO ) range is 1.8– 2.7, which depends on catalysts used [42,43]. The additional fuel (radd) is desirable to be as low as possible and the heat recovery (ge) should be as high as possible for energy efficiency. As discussed earlier, the solid circulation ratio (rcir) is concerned in the range from 35 to 80 [10,35,41,44] in this study. The contours of the four performance criteria are plotted in Fig. 11 with respect to TG and c, where the four criteria were chosen within 1.8 6 /H2 =CO 6 2.7, 0 6 radd 6 0.3, 35 6 rcir 6 80, and 0.73 6 ge. Here, the region between two solid bold lines of /H2 =CO is considered as an operating region suitable for FT synthesis. Since there is no optimum point simultaneously satisfying the above criteria, five points (A, B, C, D and E in Fig. 11) on the two bold lines of /H2 =CO are marked, which give an interesting feature, respectively. At point A, the heat recovery (ge) is maximized as well as the additional fuel ratio (radd) keeps lower. However, the solid circulation ratio is too high (rcir = 80). Point B indicates the highest heat recovery at /H2 =CO = 2.7 and rcir = 80, where the additional fuel ratio is about 0.29. At points A and B, heat efficiency of the product gas is high but operational ability is low because of the high solid circulation ratio. Point E is an alternative so that the solid circulation ratio is lowered (rcir = 55), while the heat recovery keeps high and the additional fuel ratio reaches a minimum. Points A, B and E are located at relatively high TG (758 6 TG 6 814°C) and low c (0.69 6 c 6 1.82) for high heat efficiency (0.751 6 ge 6 0.768) at the cost of the high solid circulation ratio (55 6 rcir 6 80). To increase /H2 =CO up to 2.7 and lower the solid circulation ratio (rcir = 64), point C is chosen where ge = 0.73 and radd 0.32. Point D located on the line of /H2 =CO = 1.8 has such a feature that the solid circulation ratio (rcir = 47) and the heat recovery (ge = 0.728) are relatively low. Points C and D are operated at relatively low TG (700 6 TG 6 741 °C) and high c (2.68 6 c 6 3.0), showing low solid circulation ratios (47 6 rcir 6 64). Since both points C and D are
50
40
30
20
10
0
0.5
1
1.5
2
2.5
γ (kg/kg) CO-Model value
CO-Exp.value [23]
H2 -Model value
H2 -Exp.value [23]
CO2 -Model value
CO2 -Exp.value [23]
CH4 -Model value
CH4 -Exp.value [23]
Fig. 8. Effect of steam to fuel ratio (c) on final gas composition obtained from qETG model in comparison with experiment data taken from Herguido et al. [23] at TG = 750 °C.
c except for in the lower region of c, because an excess amount of steam needs additional fuel and deteriorates the heat recovery. Fig. 10d depicts the H2/CO molar ratio (/H2 =CO ) with respect to TG and c. As shown in Figs. 6–8, the /H2 =CO increases obviously when both TG and c rise. The /H2 =CO computed by the model has about a maximum of 5 in the range investigated.
(b) 1 φCR (-)
1
0.8
0.4 900
0 700
3 2
800
800
1 700
o
TG ( C)
0
900
γ (kg/kg)
(d) 18 16
3
2
0
γ (kg/kg)
o
2.5
1
2
3
TG ( C)
(c) Vyield (Nm3 /kg, dry basis)
0.5
0.6
LHVproduct (MJ/Nm , dry basis)
φC (-)
(a)
1.5 1 0.5 900 800 o
TG ( C)
700
0
1
2
γ (kg/kg)
3
14 12 10 700 800 o
TG ( C)
900
3
2
1
0
γ (kg/kg)
Fig. 9. Parametric studies of TG and c: (a) carbon conversion (/C), (b) char residue (/CR), (c) gas yield (Vyield), and (d) lower heating value (LHVproduct).
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(a)
(b)
0.8
rcir (kg/kg)
r add (kg/kg)
0.6 0.4
200
0.2 3
0 900 0
TG ( C)
γ (kg/kg)
o
TG ( C)
(d)
0.75 0.7 0.65
1 700
0
γ (kg/kg)
6
φH2/CO (-)
0.8
4 2
0 900 900
0
3 2
800
1 700
o
(c)
0 900
2 800
ηe (-)
400
1
800 2
γ (kg/kg)
3
700
3 2
800 1
o
TG ( C)
o
TG ( C)
700 0
γ (kg/kg)
Fig. 10. Parametric studies of TG and c: (a) additional fuel ratio (radd), (b) solid circulation to fuel ratio (rcir), (c) heat recovery (ge), and (d) H2/CO molar ratio (/H2 =CO ).
Fig. 11. Contours of process performance criteria for a 100 MW DFB gasifier calculated by the qETG model with respect to TG and c.
situated at the border of the operating conditions, the two points are dependent on the operating boundary of TG and c. Process performance for the 100 MW DFB gasifier is evaluated at the five interesting points in Table 6. Since each point has a particular value for the eight performance criteria, one can choose an
operating condition suitable for a purpose of the producer gas. However, it is worth noting that the process performance evaluation is valid only if the parameters of the empirical equations used in the qETG model are well estimated for a given feedstock and gasifier.
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S.I. Ngo et al. / Applied Energy 88 (2011) 5208–5220
Table 6 Performance evaluation for steam gasification of pine woodchips in a 100 MW DFB gasifier at five interesting points of Fig. 11. Points
TG (°C)
c (kg/kg)
/C
/CR
Vyield (Nm3/kg)
LHVproduct (MJ/Nm3)
radd
rcir
ge
/H2 =CO
A B C D E
814 791 741 700 758
0.69 1.82 3.0 2.68 1.12
0.73 0.80 0.81 0.74 0.73
0.78 0.59 0.56 0.75 0.80
1.29 1.50 1.47 1.10 1.22
12.68 11.88 12.14 13.39 13.02
0.173 0.308 0.349 0.219 0.159
80 80 64 47 55
0.768 0.751 0.730 0.728 0.763
1.8 2.7 2.7 1.8 1.8
5. Conclusions In the experimental work, the pine woodchips of particle sizes of under 10 mm were used in a CFB (circulating fluidized-bed) steam gasifier as a feedstock. The result showed the final gas compositions about 26–42% of H2, 25–37% of CO, 16–19% of CO2 and 8– 11% of CH4, when the steam to fuel ratio (c) was 0.3 kg/kg and the gasification temperature (TG) was from 700 °C to 900 °C. The quasi-equilibrium three-stage gasification (qETG) model was proposed to investigate the effects of TG and c on DFB (dual circulating fluidized-bed) process performances such as (1) Carbon conversion (/C), (2) char residue (/CR), (3) gas yield (Vyield), (4) lower heating value (LHVproduct), (5) additional fuel ratio (radd), (6) solid circulation ratio (rcir), (7) heat recovery (ge), and (8) H2 to CO molar ratio (/H2 =CO ). The qETG model includes three stages: biomass pyrolysis, char–gas reaction, and gas phase reaction, where empirical equations are involved to take into account the deviation from theoretical equilibrium reactions. Final gas compositions predicted by the model had a reasonable agreement with our experimental data and previous studies from the literature. The prediction accuracy of the qETG model would be enhanced, if tar formation and cracking in the pyrolysis stage are considered. By parametric study, the eight process performances were evaluated for a 100 MW DFB gasifier with respect to TG and c. As both TG and c increase, the carbon conversion, the product gas yield, additional fuel ratio, solid circulation ratio, and H2 to CO molar ratio increase, while the lower heating value of product gas decreases. The heat recovery shows a maximum point because of a conflicting behavior among the product gas yield, the additional fuel ratio, and the lower heating value of product gas. Considering the syngas production suitable for FT synthesis, the contour of the four performance criteria (/H2 =CO , radd, ge, and rcir) was plotted as a function of TG and c. Since there was no general optimum point simultaneously satisfying the criteria, five operating conditions showing an interesting feature were suggested. In the region of high TG (758 6 TG 6 814 °C) and low c (0.69 6 c 6 1.82), high heat efficiency is found but the solid circulation ratio is too high. The low heat efficiency and low solid circulation ratio are observed in low TG (700 6 TG 6 741 °C) and high c (2.68 6 c 6 3). The contour of performance criteria provides a guideline for the selection of effective operating conditions. Acknowledgement This work is supported by Bilateral International Collaborative R&D program under the Ministry of Knowledge Economy, Korea. Appendix A
consist of CO, CO2, CH4, H2, and H2O. In the secondary pyrolysis (tar cracking step), the thermal cracking of tar takes place. Tar cracking in the gas phase has significant impact on the accurate prediction of pyrolysis gas composition. The products of tar cracking are composed of CO, CO2, H2, hydrocarbons (e.g., C2H6, C2H4, and C3H6), and inert tar [19,20,45]. Prasad and Kuester [7] proposed an empirical model obtained from a linear regression to predict the composition of pyrolysis gas, in which the gas components were expressed as a function of pyrolysis temperature and steam to fuel ratio. A simplified approach was reported by Sharma, and Gao and Li [46,47], herein the product gas of pyrolysis was assumed to be only CO, CH4, and H2O. The atomic balances determined the composition of these components. Sadaka et al. [48] calculated the pyrolysis gas composition using the elemental balance under the assumption that the product consists of CO, CO2, CH4, H2, and H2O. The approach proposed by Sadaka et al. [48] is used in this study to predict the composition of the gaseous product released from biomass pyrolysis. In this approach, volatiles and tar in biomass are considered to be completely converted to the five gaseous components mentioned above without any inert tar remaining. The elemental balances are written as:
nc ¼ nCO þ nCO2 þ nCH4
ðA:1Þ
nH ¼ 4nCH4 þ 2nH2 þ 2nH2 O
ðA:2Þ
nO ¼ nCO þ 2nCO2 þ nH2 O
ðA:3Þ
where n is the number of moles. The nitrogen content in biomass (if any) is assumed to form inert nitrogen gas. Thus the elemental balance for nitrogen is
nN ¼ 2nN2
ðA:4Þ
In order to close the elemental balance equations, two terms defined as follows are required [48]: The fraction of CO formation,
/CO=CO2 ¼
nCO nCO2
ðA:5Þ
and the fraction of CH4 formation,
/CH4 =H2 ¼
nCH4 nH2
ðA:6Þ
The two fractions are determined by pyrolysis experiments for a biomass used. The experimental data show that with an increase in temperature /CO=CO2 increases, while /CH4 =H2 decreases. It is supposed that the reaction between a certain part of char and CO2 by Boudouard reaction (C + CO2 ? 2CO) leads to the increase of /H2 =CO when the pyrolysis temperature increases [26]. Increase in the H2 concentration by the dehydrogenation of saturated hydrocarbons such as C2H6 and C3H8 might be the explanation for the decrease in /CH4 =H2 .
A.1. Pyrolysis A.2. Char–gas reaction The non-oxidative decomposition of biomass called pyrolysis may be divided into two main steps: (1) primary and (2) secondary pyrolysis. In the first step, biomass thermally decomposes into gases, tar, and char. The gases released from the primary pyrolysis
The second stage of the qETG model is the char–gas reactions, which are composed of a series of complex and competing reactions between biomass char and gasification reagents. However,
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S.I. Ngo et al. / Applied Energy 88 (2011) 5208–5220
the Boudouard reaction, primary char–steam reaction and secondary char–steam reaction are assumed to take place and to reach an equilibrium state in this study:
CðsÞ þ CO2 ðgÞ $ 2COðgÞ
ðIÞ
CðsÞ þ H2 OðgÞ $ COðgÞ þ H2 ðgÞ
ðIIÞ
CðsÞ þ 2H2 OðgÞ $ CO2 ðgÞ þ 2H2 ðgÞ
ðIIIÞ
reactant. Each reaction proceeds forward or backward to a certain degree ejej, called the reaction coordinate, in order to reach an equilibrium from the initial stage [50]. At the equilibrium state, the mole fraction of species ai is expressed as [50]:
yi ¼
P ni;0 þ j mi;j ej P n0 þ j mj ej
ðA:9Þ
Since the steam amount involved in the char–gas equilibrium reaction increases with the increase of gasification temperature, it is needed to determine the ratio of the steam amount involved in reactions (II) and (III) to the total water amount in this stage [3,18]. The steam participation ratio (b) is estimated from the char–gasification experimental data taken from Yoshida et al. [18]. In Table 3, b is expressed as a function of gasification temperature.
where ni,0 and n0 are the moles of species ai and the total moles of all species in the system at the initial state, respectively. vj is the total stoichiometric number of the jth reaction which is defined as:
A.3. Gas-phase reaction
yCO ¼
It was reported that the water–gas shift reaction dominates in biomass gasification at the temperature higher than 670 °C [23,27,38]:
COðgÞ þ H2 OðgÞ $ CO2 ðgÞ þ H2 ðgÞ
ðIVÞ
The methanation reaction C(s) + 2H2(g) ? CH4(g) in the second stage and the steam reforming reaction CH4(g) + H2O(g) ? CO(g) + 3H2(g) in the third stage of the qETG model are not supposed to take place. A.4. Calculation procedures
nC
3
2
1 1
2
3
0
7 6 2 07 7 6 nCO2 7 7 6 7 1 0 7 6 nCH4 7 76 7 7 6 0 27 7 6 nH2 7 7 6 7 0 0 5 4 nH2 O 5 0 0 nN2
0 2
4 2 0 0
0
0
0
/CO=CO2
0
0
0
0
1 /CH4 =H2
0
3
1 0
6n 7 60 6 H7 6 6 7 6 6 nO 7 6 1 6 7¼6 6n 7 60 6 N7 6 6 7 6 40 5 41 0
nCO
mi;j aj ¼ 0
mi;j
ðA:10Þ
i
By applying Eq. (A.10) to the reactions of the second stage (reactions I, II, III), the mole fraction of the gas species is expressed as:
nCO;0 þ 2e1 þ e2 n0 þ e1 þ e2 þ e3 nCO2;0 e1 þ e3 yCO2 ¼ n0 þ e1 þ e2 þ e3 nH2 ;0 þ e2 þ 2e3 yH2 ¼ n0 þ e1 þ e2 þ e3 nH O;0 e2 2e3 yH2 O ¼ 2 n0 þ e1 þ e2 þ e3 nCH4 ;0 yCH4 ¼ n0 þ e1 þ e2 þ e3 nN2 ;0 yN2 ¼ n0 þ e1 þ e2 þ e3
ðA:11Þ ðA:12Þ ðA:13Þ ðA:14Þ ðA:15Þ ðA:16Þ
2 K 1 ¼ y1 CO2 yCO
ðA:17Þ
y1 H2 O
ðA:18Þ
K2 ¼
yCO yH2
2 K 3 ¼ y2 H2 O yCO2 yH2
ðA:7Þ
DG0T G ;j K j ¼ exp RT G
! ðA:20Þ
where D0T G is calculated from standard Gibbs energy of formation:
X
DG0T G ;f ;i ¼ DG0T G ¼ ðDH0T G T G DS0T G Þ
ðA:8Þ
where vi,j is the stoichiometric number of species ai involved in the jth reaction. It is positive if ai is the product and negative if ai is the
ðA:21Þ
i
The molar enthalpy hi of a component i in the gas phase is defined by the enthalpy of formation D0fi and the molar heat capacity Cpi according to the following equation [50].
DH0T G ¼
X
m i hi
i
Z
ðA:22Þ
TG
cpi ðT G ÞdT G
ðA:23Þ
298:15
In the third stage, the thermodynamic equilibrium constant of water–gas shift reaction is first calculated by the same way as the second stage, then the non-equilibrium factor (j) is applied:
K 4 ¼ j K 4
j
ðA:19Þ
From the Gibbs free energy equation, we have:
hi ¼ DH0f ;i þ
In the second stage, the molar fraction of the gas species is calculated. For a stoichiometric multi-reaction system, a set of i species ai involved in j independent reactions can be written as:
X
X
The equilibrium constant of each reaction is required to determine the reaction coordinates:
Because the hydrogen and oxygen content in char decreases sharply as temperature increases [26,49], the char remaining after biomass pyrolysis is assumed to be pure carbon. The moles of carbon in volatiles are determined by the difference between fixed carbon content in the proximate analysis and carbon content in the ultimate analysis. The moles of H, O, and N are then calculated from the ultimate analysis. The composition of each species in pyrolysis gas is obtained by solving element-based balances. These compositions are used as the initial conditions of the equilibrium model in the second stage for char gasification. At the second and third stages of the qETG model, instead of using element-based balances, the compositions of gaseous components are calculated by using species-based balances. The equilibrium constant for each reaction is calculated from the Gibbs free energy change of the reaction [50]. At the third stage, the equilibrium constant (K4) of the water–gas shift reaction (reaction (IV)) is corrected by the non-equilibrium factor (j). In the first stage, six linear equations are established to calculate the compositions of pyrolysis gas:
2
vj ¼
where
K 4
ðA:24Þ
is the corrected equilibrium constant.
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Applied Energy journal homepage: www.elsevier.com/locate/apenergy
Performance evaluation for dual circulating fluidized-bed steam gasifier of biomass using quasi-equilibrium three-stage gasification model Son Ich Ngo a, Thanh D.B. Nguyen a,e, Young-Il Lim a,⇑, Byung-Ho Song b, Uen-Do Lee c, Young-Tai Choi c, Jae-Hun Song d a
Lab. FACS, RCCT, Department of Chemical Engineering, Hankyong National University, Gyonggi-do, Anseong-si, Jungangno 167, 456-749, Republic of Korea Department of Chemical Engineering, Kunsan National University, Gunsan, Jeonbuk 573-701, Republic of Korea High Temperature Processing R&D Department, Korea Institute of Industrial Technology (KITECH), Cheonan 331-825, Republic of Korea d R&D Center, SeenTec Co., Ltd., #608 Office Plazza, 7-2 Yongho-dong, Changwon, Gyeongnam 641-969, Republic of Korea e School of Chemical Engineering, Hanoi University of Science and Technology, No. 1 Dai Co Viet, Hanoi, Viet Nam b c
a r t i c l e
i n f o
Article history: Received 21 April 2011 Received in revised form 8 July 2011 Accepted 27 July 2011 Available online 20 August 2011 Keywords: Steam gasification Woodchips Dual circulating fluidized-bed (DFB) Solid circulation ratio Quasi-equilibrium three-stage gasification (qETG) model Parametric study
a b s t r a c t The effects of gasification temperature (TG) and steam to fuel ratio (c) on product gas composition and yield were experimentally investigated for steam gasification of pine woodchips in a bench-scale circulating fluidized-bed (CFB) gasifier with external heat supplier. To evaluate process performance in a dual circulating fluidized-bed (DFB) with heat carrier (silica sand), a quasi-equilibrium three-stage gasification (qETG) model was developed and validated with experimental data of biomass steam gasification. The model was divided into three stages including biomass pyrolysis, char–gas reactions, and gas-phase reactions. Carbonic and methane formation ratios were considered at the pyrolysis stage under the assumption of spontaneous decomposition. At the second and third stages, char–gas and gas-phase equilibrium reactions were corrected by two empirical equations concerning the steam participation ratio and the non-equilibrium factor, respectively. Using the qETG model, parametric study on TG and c was performed to predict final gas composition, carbon conversion, char residue, gas yield, lower heating value, additional fuel ratio, solid circulation ratio, heat recovery and H2 to CO molar ratio. Focusing on the solid circulation ratio and H2/CO molar ratio, several effective operating conditions were suggested from the contour of performance criteria. Ó 2011 Elsevier Ltd. All rights reserved.
1. Introduction As we know, there are nine general sources of energy on the Earth: Solar, biomass, wind, wave, hydro, tidal, geothermal, nuclear, and fossil. Except for the last three, the remaining six are generally called renewable sources of energy, as they are not depleted with time [1]. Biomass sources in the plant form store solar energy as chemical energy during their growth by photosynthesis. This storage of chemical energy can be released and converted to many other useful forms of energy such as heat, electricity, light, and automotive power through some upgrading and conversion systems [2]. In recent years, much attention has been paid to the thermochemical conversion from biomass since: (1) the use of biomass for energy has the potential to control the net emission of CO2 to the atmosphere [2] and (2) it can reduce the dependence of fossil sources. Among the thermochemical conversion processes such as gasification, pyrolysis and combustion, gasification has been ⇑ Corresponding author. Tel.: +82 31 670 5207; fax: +82 31 670 5445. E-mail address: [email protected] (Y.-I. Lim). 0306-2619/$ - see front matter Ó 2011 Elsevier Ltd. All rights reserved. doi:10.1016/j.apenergy.2011.07.046
considered to be the most cost effective process since it converts biomass into clean combustible gaseous products [3]. Steam gasification has become an area of growing interest, since it produces a gaseous fuel with relatively higher hydrogen content, which could be used in fuel cells as a new recognizable technology for the future to produce power in a much cleaner manner. It also has other advantages in comparison with air or air–steam gasification such as, (1) producing a gas with higher heating value, (2) reducing the dilution effect of nitrogen from air, and (3) eliminating the need for an expensive oxygen plant when both air and oxygen are used as gasification reagents [4]. However, steam gasification is a more complex process, because the heat necessary to the process is not directly supplied by the partial combustion of the feedstock during the gasification process [4]. Among several types of gasifiers such as fixed bed, down- and updraft, and fluidized-bed, the fluidized-bed has been known as a promising alternative , thereby maximizing the gas product yield due to the uniform temperature distribution and high mass and heat transfer rates [5]. The dual circulating fluidized-bed (DFB) system has many advantages that are capable of scaling up and
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Nomenclatures A empirical parameters (–) B empirical parameters (–) cp molar heat capacity (kJ/kmol/K) Cp heat capacity (kJ/kg/K) DG0T G Gibbs free energy change of reaction at 25 °C (kJ/kmol) DG0f standard Gibbs energy of formation at 25 °C (kJ/kmol) h molar enthalpy (kJ/kmol) Hbiomass sensible heat of biomass fed to gasifier (kJ/h) Hsens sensible heat of unreacted char leaving gasifier (kJ/h) char Hdemand heat demand (kJ/h) heat entering into gasifier (kJ/h) Hg,in Hg,out heat leaving gasifier (kJ/h) Hloss heat loss through surrounding environment (kJ/h) Hcomb combustion heat of syngas (kJ/h) product sensible heat of syngas leaving gasifier (kJ/h) Hsens product Hsteam heat of steam fed to gasifier (kJ/h) DH0f standard enthalpy of formation at 25 °C (kJ/kmol) K equilibrium constant (–) K corrected equilibrium constant (–) Lwater latent heat of water (kJ/kg/h) LHVbiomass lower heating value of biomass (kJ/kg) LHVproduct lower heating value of producer gas (MJ/Nm3) LHVchar lower heating value of char (kJ/kg) Madd amount of additional fuel to combustor (kg/h) Mbiomass biomass feed rate (kg/h) amount of carbon in biomass (kg/h) MC MC,in amount of fixed carbon in biomass (kg/h) Mchar amount of the unreacted char (kg/h) MC,convert amount of carbon converted to producer gas (kg/h) MC,unreact amount of unreacted char (kg/h) Msand amount of recirculated sand (kg/h) n number of moles (kmol/h)
development. However, it is also known as a complex system in construction and installation [6]. Numerous studies for biomass gasification by means of modeling and simulation include (1) thermodynamic equilibrium models [7– 10], (2) kinetic rate models [1,11–13], and (3) neural network models [14,15]. In the kinetic rate models, initial conditions and kinetic parameters are not well known because of a variety of feedstock [11]. The neural network models as a kind of black-box models have achieved high prediction accuracy. However, it is hard to obtain physical meaning from these models, and the scale-up and adaption abilities of the neural network models are restricted. The kinetic models predict the progress of product composition with respect to the residence time in a gasifier, whereas the equilibrium models provide the maximum yield of a desired product which is achievable from a gasification system [16]. Although kinetic rate models are considered as a rigorous approach, equilibrium models are valuable because they can predict thermodynamic limits which are used to design, evaluate and improve the process [17]. The equilibrium models have been used for preliminary study on the influence of the most important process parameters. Several equilibrium models have been proposed for coal gasification [3,18], but these models cannot directly be applied to biomass due to significant differences in chemical and physical properties between coal and biomass. In biomass gasification, the pyrolysis step is one of the most important steps for gasification, since it strongly affects the final gas product composition [19–22]. The knowledge of devolatilization in the pyrolysis step is crucial for a precise prediction. Therefore, the more detailed the pyrolysis is taken into account, the better the prediction will be.
p Vproduct radd rcir R Rt Tambient Tfeed TG Tpyrol Triser T in steam Vyield x y z
empirical polynomial parameter (–) volumetric flow rate of product gas (Nm3/h, dry basis) additional fuel ratio (kg/kg) solid circulation ratio (kg/kg) gas constant (kJ/kmol/K) retention time (min) ambient temperature (K or °C) temperature of biomass fed to gasifier (K or °C) gasification temperature (K or °C) pyrolysis temperature (K or °C) riser temperature (K or °C) inlet temperature of steam (K or °C) gas yield (Nm3/kg biomass) mass fraction (–) mole fraction (–) empirical variable (1/K)
Greek letters c steam to fuel ratio (kg/kg) e reaction coordinate (kmol) j non-equilibrium factor (–) ge heat recovery (–) / ratio of molar fractions(–) /C carbon conversion (–) char residue (–) /CR m stoichiometric number (–) a species involved reactions v total stoichiometric number (–) Subscripts i species index j reaction index
In this study, we investigate the biomass gasification with the steam agent in a bench-scale circulating fluidized-bed (CFB) gasifier and develop a quasi-equilibrium three-stage gasification (qETG) model for the prediction of process performance in dual circulating fluidized-bed (DFB). The qETG model is divided into three main stages: (1) pyrolysis of volatiles in biomass, (2) solid–gas reactions between biomass char and gasifying reagents (carbon dioxide or steam) in the fluidized bed, and (3) gas-phase reactions among the gaseous species in the free board of the gasifier. At each stage, empirical models are established based on the experimental data to calculate the gaseous components. Especially, the deviation from equilibrium reaction is taken into account in the third stage by a non-equilibrium factor. The model is first validated by the experiment data conducted in the bench-scale CFB gasifier with pine woodchips, and the data taken from the literature [17,23]. Then the contour of performance criteria (i.e., the solid circulation ratio, the H2/CO molar ratio, the additional fuel ratio, and the heat recovery) is obtained from the parametric study of TG and c using the qETG model. Focusing on the solid circulation ratio and H2/CO molar ratio for Fisher–Tropsch (FT) synthesis, several effective operating conditions were proposed for a 100 MW DFB biomass gasifier.
2. Experiment Experiments were performed in a bench-scale CFB gasifier (0.15 m I.D. 3.0 m height) operating at atmospheric pressure. The schematic diagram of the experimental set-up is shown in Fig. 1. The gasification system consists of a CFB gasifier with
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Fig. 1. Schematic diagram of experimental set-up for bench-scale CFB gasifier.
Table 1 Proximate and ultimate analyses of pine woodchips used in this study. Proximate analysis (wt.%, wet basis)
Ultimate analysis (wt.%, dry basis)
H2O Volatile Fixed carbon Ash
6.4 75.9 17.4 0.3
C H O N S Cl
Total
100.0
50.8 5.4 43.6 Not detected 6.2103 Not detected 99.8
electrical heaters, a screw feeder for continuous supply of the feedstock, two cyclones, and an impinger box for collecting fly ash, a tar trap, and a gas analyzer. Pine woodchips of sizes under 10 mm were used as a feedstock. The proximate and ultimate analyses of the woodchips are listed in Table 1. It is addressed that the biomass material contained large volatiles (76%) and low ash content (0.3%) in comparison with common coal materials [24,25], and a little sulfur was detected from the ultimate analysis. The screw feeder with a feed rate of about 5 kg/h continuously supplied the feedstock. The gasifier temperature was controlled by electrical heaters, which could be heated up to 1300 °C. Steam was used as a fluidized and gasifying agent. Superheated steam at 400 °C was introduced to the gasifier to remove condensable components. The on-line gas analyzer (Model: AO2020, ND-IR analyzer, ABB Co. Ltd., Germany) monitored the composition of syngas. Gas chromatography (Model: 7890A, Agilent Technologies, USA) was employed for the quantitative analysis of sulfur compounds, where highly sensitive Flame Photometric Detector (FPD) was generating a flame with hydrogen and oxygen. The flow rates of hydrogen and oxygen were 100, and 60 ml/min, respectively. The inlet and detector temperatures were 250 °C. The PLOT capillary column
(GC-gaspro, Agilent, I.D. = 0.32 mm, L = 5 m) was used since it is suitable for light molecular weight hydrocarbon compounds or sulfur compounds. For the identification of the sample gas, the area and retention time (Rt) was calibrated by standard gases of COS (49.6 ppm) and H2S (177.6 ppm) with N2 balanced. Helium was used as a carrier gas, and the oven temperature was maintained at 50 °C. The COS was detected after 7.0 min and the H2S was detected after 8.1 min. Temperature was measured at the six locations of which the third one from the bottom (T3 in Fig. 1) was considered as the gasification temperature (TG). Steam was injected from the bottom of the gasifier and the steam to fuel ratio (c) was adjusted by the steam flow rate. The two operating conditions ranged 700 6 TG 6 900 °C and 0.3 6 c 6 1.0 kg/kg. At TG = 800 °C and c = 0.3 (kg/kg), Fig. 2a shows temperature profiles of the six measuring points with respect to the running time of biomass feeding. The temperatures slightly dropped down just after feeding. The temperatures and the product gas composition were stabilized after 5 min (see Fig. 2). The gasification temperature (TG) measured at T3 was controlled by manipulating the electrical heater. The syngas composition measured after 30 min of the running time was used for the final one. 3. Quasi-equilibrium three-stage gasification (qETG) model and process performance evaluation A quasi-equilibrium three-stage gasification (qETG) model involving four empirical equations is developed to investigate the influence of two operating conditions, gasification temperature (TG) and steam to fuel ratio (c), on process performance of a 100 MW (20 tons/h) DFB gasifier. The parameters of the four empirical equations are estimated from experimental data published in the literature [18,23,26,27]. The empirical equations are
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(a)
900
o
Temperature ( C)
1000
800 700 T1 T2 T3 T4 T5 T6
600 0
Gas composition (vol%, dry basis)
50 H2 CO 2 CH 4
(b) 40
CO
30
20
10
0 0
5
10
15
20
25
30
Time (min) Fig. 2. Experimental data in the bench-scale CFB performed at TG = 800 °C and c = 0.3 kg/kg: (a) temperature profile, and (b) syngas composition.
used to calculate the gaseous components at each stage. Table 2 shows main reactions and products in the three-stage gasification. Three char–gas reactions and a water–gas shift reaction are taken into account in this model. The qETG model is validated with experimental data obtained from the bench-scale CFB as well as the other data from the sources. Parametric study on TG and c is performed to predict final gas composition (yi), carbon conversion (/C), char residue (/CR), lower heating value (LHVproduct), heat recovery (e), gas yield (Vyield), additional fuel ratio (radd), and solid circulation ratio (rcir). Fig. 3 summarizes the structure of qETG model and the scope of parametric study. 3.1. Quasi-equilibrium three-stage gasification (qETG) model The three stages of gasification are divided into biomass pyrolysis, char–gas reactions, and gas-phase reactions. At the pyrolysis stage under the assumption that volatiles and tar in biomass are
completely and spontaneously converted into five gaseous components (CO, CO2, CH4, H2, and H2O), the carbonic formation ratio (/CO=CO2 ) and the methane formation ratio (/CH4 =H2 ) determine gas composition after pyrolysis by the elemental balance (see Appendix A.1). The two empirical equations in the first stage (/CO=CO2 and /CH4 =H2 ) were derived on the basis of experimental data [26] conducted at a wide range of temperatures (400–900 °C), using three kinds of biomass (woodchip, coconut shell and straw). Fig. 4 illustrates the dependency of the two ratios on pyrolysis temperature (Tpyrol). As shown in Fig. 2b, the temperatures are distributed inside the gasifier. Since pyrolysis and decomposition of biomass depend on its pore size distribution, heating rate [28] and residence time, it is not easy to determine an appropriate Tpyrol in the range of 700 6 TG 6 900 °C. Thus, Tpyrol was assumed to be the same as TG in our study. Char–gas and gas-phase reactions are supposed to reach an equilibrium state. At the second stage, the steam participation ratio (b) is used to determine the amount of steam involved in the char– gas reactions with respect to TG [3,18]. At the third stage, a nonequilibrium factor (j) is employed to account for the deviation from thermodynamic equilibrium of the water–gas shift reaction [23,27,29]. Due to b and j, the qETG model compensates for a deviation from equilibrium reaction according to the gasification temperature (see Appendices A.2 and A.3). The steam reforming reaction (CH4 + H2O CO + 3H2) is assumed not to take place in the third stage (see Table 2). This reaction is not favorable at temperatures over 800 °C, since the Gibbs free energy of formation has a positive value [30]. Moreover, Dupont et al. concluded from a modeling approach that the steam reforming of CH4 was kinetically limited [22] in those temperatures. Therefore, only one reaction for water–gas shift is considered in this stage. Fig. 5 shows the temperature dependency of the thermodynamic equilibrium constant (K4, see Appendix A.4), and the experimental equilibrium constant (K4,exp = (yCO2 yH2 Þ=ðyCO yH2 O )) for the water–gas shift reaction. The molar fraction ratios (K4,exp) between reactants and products obtained from experimental value of Wei et al. [27] and Herguido et al. [23] are far from the thermodynamic equilibrium constants (K4). The former [27] was measured in a down-draft gasifier and the latter [23] was in a fluidized bed which is closer to K4 than the former. To compensate this discrepancy, the non-equilibrium factor (j) has been used elsewhere [29,31]. Applying the non-equilibrium factor (j) to K4, the modified equilibrium constant (K 4 ¼ j K 4 ) follows the trends of experimental data. The non-equilibrium factor (j) may depend on feedstock, operating conditions, and the type of gasifiers. In this study, j was determined so that our experimental syngas compositions measured for CO, CO2, and H2 were predicted as best as possible by the qETG model. In Table 3, the fifth-order polynomial of j is suggested as the best fitting
Table 2 Reactions and products assumed in the qETG model. Stage
Reactions
Products
Components assumed in qETG
Refs.
Pyrolysis
First step: thermal decomposition Second step: tar cracking
CO, CO2, CH4, H2 and H2O
CO, CO2, CH4, H2 and H2O
[19,20,45,48]
Char unreacted, CO, H2, H2O-residue
[3,18]
CO, CO2, H2, H2O
[23,27,30,38,47]
Solid–gas reactions
Water–gas shift reaction
C(s) + CO2(g) M 2CO(g)
CO, CO2, H2, heavier hydrocarbon (e. g., C2H6, C2H4, and C3H6), and inert tar. (Char unreacted) CO, CO2, H2, (H2O-residue)
C(s) + H2O(g) M CO(g) + H2(g) C(s) + 2H2O(g) M CO2(g) + 2H2(g) CO(g) + H2O(g) M CO2(g) + H2(g)
CO, CO2, H2, H2O
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β
κ
Fig. 3. Structure of the qETG model and scope of parametric study.
15
2
(a)
Woodchip [26] Straw [26]
1.6
Empirical model
1.4
5
K 4*
κ
1.2
value
φ
CO/CO2
10
exp. [23] exp. [27] K4
1.8
Coconut shell [26]
0 400
500
600
700
T
800
0.8
900
o
pyrol
1
( C)
0.6 0.4
15
(b)
0.2
Woodchip [26]
0 700
Coconut shell [26] 10
φ
CH4/H2
Straw [26]
800
850
900
o
TG ( C)
Empirical model
Fig. 5. Temperature dependency of experimental (K4,exp), thermodynamic (K4), and corrected (K 4 ) equilibrium constants, and non-equilibrium factor (j).
5
0 400
750
500
600
T
700 pyrol
800
900
(oC)
Fig. 4. Ratios of CO/CO2 (/CO=CO2 ) and CH4/H2 (/CH4 =H2 ) versus pyrolysis temperature (Tpyrol): (a) /CO=CO2 , and (b) /CH4 =H2 .
function, and it converges to the thermodynamic equilibrium constant as the temperature increases. It is assumed that these four empirical equations only depend on the gasification temperature (TG), as indicated in Table 3. A detailed description of the qETG model is given in the Appendix.
3.2. Performance criteria
gas(LHVproduct), (5) additional fuel ratio (radd), (6) solid circulation ratio (rcir), (7) heat recovery (ge), and (8) H2 to CO molar ratio (/H2 =CO ). Carbon conversion (/C, dimensionless) is defined as the mass fraction of carbon content (MC,convert, kg/h) converted from the product gas by pyrolysis and gasification to the mass of carbon in biomass (MC, kg/h).
/C ¼
ð1Þ
Char residue (/CR, dimensionless) is calculated by the mass ratio of unreacted carbon after char–gas reactions (MC,unreact, kg/h) to the amount of fixed carbon at the initial condition (MC,in, kg/h). /CR ¼ ¼
The main purpose of this study is to evaluate process performance, when the targeting feedstock (i.e., pine woodchips) with a constant feed rate (Mbiomass = 20,000 kg/h) will be used for biomass steam gasification in a DFB gasifier. The following criteria for process performance are investigated: (1) Carbon conversion(/C), (2) char residue (/CR), (3) gas yield (Vyield), (4) lower heating value of product
M C;conv ert carbon in CO; CO2 and CH4 after 2nd stage ¼ carbon in biomass MC
M C;unreact MC;in ðM C;in þ carbon in gas after 1st stageÞ ðcarbon in gas after 2nd stageÞ M C;in ð2Þ
The heat released by combustion of product gas calculated as follows [32]:
Hcomb product ¼ 285:63nH2 þ 282:99nCO þ 890:3nCH4
(Hcomb product ,
MJ/h) is
ð3Þ
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S.I. Ngo et al. / Applied Energy 88 (2011) 5208–5220 Table 3 Empirical equations used in this study. Pyrolysis
Carbonic formation ratio
B1 /CO=CO2 ¼ nnCOCO ¼ A1 exp T pyrol
A1 = 4.70 103
[48]
2
3
Methane formation ratio Char–gas reactions
Steam participation
/CH4 =H2 ¼
nCH4 nH2
B2 ¼ A2 exp T pyrol
B1 = 7.16 10 K A2 = 2.30 103 B2 = 5.40 103 K A3 = 5.14 101
nH O;0 b ¼ nH 2O;tot ¼ A3 exp TBG3
[26] [3]
2
B3 = 7.54 103 K Water–gas shift reaction
Non-equilibrium factor
j = p1 z5 + p2 z4 + p3 z3 + p4 z2 + p5 z + p6
[23]
p1 = - 3.2 103; p2 = - 7.8 103; p3 = 2.3 103; p4 = 7.4 102; p5 = 3.0 101; p6 = 4.0 101;
[27]
1070 z ¼ T G 62:0
where nH2 ; nCO and nCH4 (kmol/h) are the mole numbers of combustible product gases. The lower heating value of the final product gas, LHVproduct (MJ/m3), is then calculated as:
LHV product ¼
Hcomb product V product
ð4Þ
V yield
ð5Þ
The lower heating value of biomass (LHVbiomas, kJ/kg) is calculated by the Boie formula [10]:
LHV biomass ¼ 34; 835xC þ 93; 870xH 10; 800xO þ 6280xN þ 10; 465xS
ð6Þ
xC, xH, xO, xN, and xs are the mass fraction of the elements C, H, O, N, and S in biomass, respectively.In the DFB process, the heat for the gasification reaction taking place in the gasifier (Hdemand) is provided by the combustion heat of the unconverted char and additional fuel (if needed) in the riser. The silica sands play a key role as a heating carrier from the combustor to the gasifier. It is assumed in this study that the sand particles absorb all the heat released from the combustion. The heat demand of the gasifier may be expressed as:
Hdemand ¼ Hg;out Hg;in
ð7Þ
where Hg,out (kJ/h) is the heat leaving the gasifier and Hg,in (kJ/h) is the heat of the feed entering into the gasifier. The total inlet heat (Hg,in, kJ/h) is calculated as:
Hg;in ¼ Mbiomass LHV biomass þ Hsens biomass þ H steam Hsens biomas
where conditions:
ð9Þ
C p;biomass was assumed as 1.0 kJ/kg/K for woodchips [33]. The heat of steam (Hsteam, kJ/h) includes sensible heat and latent heat caused by the vaporization of water:
ð10Þ
The total outlet heat (Hg,out, kJ/h) is obtained as follows: sens sens Hg;out ¼ 1000ðHcomb product Þ þ H product þ M char LHV char þ H char
þ Hloss
ni cpi ðT G Þ ðT G T ambient Þ þ nH2 O cp;H2 O ðT G Þ
T ambient ÞÞ
ð11Þ
where LHVchar (kJ/kg) and Mchar (kg/h) are the lower heating value and the mass flow rate of the unreacted char, respectively. Hsens product
ð12Þ
where ni (kmol/h) from the final gas product and cpi ðT G Þ (kJ/kmol/K) are the mole number and the heat capacities at TG (K) of CO, CO2, H2, and CH4, respectively. nH2 O ; cp;v apor and Cp,water are the mole number of H2O, the molar heat capacity of vapor and the heat capacity of water, respectively. Sensible heat of char (considered as pure carbon) is expressed as:
Hsens char ¼
M C;unreact cp;char ðT G Þ ðT G T ambient Þ 12
ð13Þ
where the unreacted char (MC,unreact) is given in the numerator of Eq. (2) and cp;char ðT G Þ is the molar heat capacity of char at TG (K). The molar heat capacity of the pure components in high temperatures is expressed in terms of fourth order polynomials (NIST webbook [34]) in Table 4. With low temperatures, the thermodynamic property was estimated as a constant. Since the ambient temperature (Tambient) is equal to the biomass feed temperature (Tfeed), the sensible heat of biomass (Hsens biomas ) is zero in Eq. (9). The heat loss of DFB gasifiers reported in the literature [35] was about 0–30% of the heat value of biomass fed to the gasifier. In the qETG model, the heat loss (Hloss) from the gasifier is assumed to be 10% (Hloss = 0.1 Mbiomass LHVbiomass). The circulation rate of silica sand (Msand, kg/h) is obtained from the heat consumption of the gasification.
Msand ¼
1:1 Hdemand C p;sand ðT riser T G Þ
ð14Þ
where Cp,sand is the heat capacity of silica sand (Cp,sand = 0.64 kJ/kg/K [36]). Triser is the temperature in the riser, which is assumed to be 950 °C in DFB. The solid circulation ratio (rcir) is given as the ratio of the circulation rate of the sand particles to the fuel feed rate (Mbiomass):
rcir ¼
Hsteam ¼ M biomass c ðC p;water ð373 T ambient Þ þ Lwater þ C p;v apor ðT in steam 373ÞÞ
X
ðT G 373Þ þ 18nH2 O ðLwater þ C p;water ð373
ð8Þ
(kJ/h) is the sensible heats of biomass at the inlet
Hsens biomas ¼ M biomass C p;biomass ðT feed T ambient Þ
Hsens product ¼
i
where Vproduct (Nm3/h) is the volume of final product gas on the dry basis. The yield of the final gas product (Vyield, Nm3/kg biomass) is determined as the volume of the product gas (Vproduct) obtained when 1 kg of biomass is gasified.
V product ¼ M biomass
and Hsens char (kJ/h) are the sensible heats of the product gas and the unreacted char, respectively. The sensible heat of CO, CO2, H2, CH4 and H2O at the out stream was computed by the enthalpy formulations:
M sand Mbiomass
ð15Þ
The larger the rcir, the more the energy consumption for the bed particles circulation between the combustor and gasifier, as well as the higher the breakage and attrition of bed particles in the combustor [9]. However, Seo et al. [37] suggested to maintain the solid circulation rate over a certain amount for stable transportation in the loop-seal of the DFB gasifier. The additional fuel (Madd) would be supplied if an additional heat were required for endothermic gasification reaction:
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S.I. Ngo et al. / Applied Energy 88 (2011) 5208–5220
Table 4 Operating conditions and thermodynamic properties for pure components used in heat balance. Operating conditions Tambient T in steam Tfeed
Constants 298 K (25 °C) 673 K (400 °C)
Lwater C p;water
2260 (kJ/kg/h) 4.1813 (kJ/kg/K)
298 K (25 °C)
C p;v apor
2.080 (kJ/kg/K)
973–1173 K (700–900 °C)
TG
Fourth order polynomial equations of molar heat capacity (kJ/kmol/K) cp,CO 29:7 6:50 103 T G þ 1:83 105 T 2G 9:39 109 T 3G þ 1:08 1012 T 4G cp;CO2 29:3 2:24 102 T G þ 2:65 104 T 2G 4:15 107 T 3G þ 2:01 1010 T 4G cp;H2 27:0 þ 1:19 102 T G 2:41 105 T 2G þ 2:15 108 T 3G 6:15 1012 T 4G cp;CH4 36:2 5:11 102 T G þ 2:21 104 T 2 1:82 107 T 3 þ 4:90 1011 T 4 G
G
G
cp;H2 O
33:8 5:95 103 T G þ 2:24 105 T 2G 9:96 109 T 3G þ 1:10 1012 T 4G
cp,char
27:0 þ 4:58 104 T G 4:5310 2:1810 þ 8:0010 7:2110 TG T2 T3 T4
3
6
8
G
M add ¼
Hdemand 1:12 Hsens char M char LHV char LHV biomass
G
50
where the heat loss due to the transportation of sand particles from the riser to the gasifier is supposed to be 12% of Hdemand. The additional fuel ratio (radd, dimensionless) is defined as the mass ratio of the additional fuel (Madd) to the feed rate (Mbiomass):
45
Madd Mbiomass
ð17Þ
The heat recovery (ge, dimensionless) is expressed as the ratio of the heat of the product gas (LHVproduct) to that of the total biomass fed to the system:
ge ¼
V yield LHV product ð1 þ r add Þ LHV biomass
ð18Þ
Note that the units of LHVproduct and LHVbiomas are different and Vyield multiplied in Eq. (18) makes ge dimensionless. 4. Results and discussion In this section, the proposed model (qETG model) is validated with the experimental data measured in the bench-scale CFB gasifier as well as the data taken from the published works [17,23]. The effects of the gasification temperature (TG) and the steam to fuel ratio (c) on the DFB gasifier performances are examined by means of their parametric study. Effective operating conditions are proposed for the production of syngas suitable for Fisher–Tropsch (FT) synthesis. 4.1. Model verification In Fig. 6, three experiment sets operated at c = 0.3 kg/kg and TG = 700, 800, and 900 °C are compared to the qETG model results. The H2 and CO2 concentrations increase with the rise in temperature, while the concentrations of CH4 and CO show an opposite trend. Since the water–gas shift reaction (CO + H2O M CO2 + H2) proceeds forward at the temperatures above 700 °C in the presence of steam [4,38], the increase in H2 and the decrease in CO are expected from the reaction when TG increases. However, for a higher temperature range, the water–gas shift reaction has been thought to be less important and the Boudouard reaction (C + CO2 M 2CO) and the primary and secondary water gas reactions (C + H2O M CO + H2, and C + 2H2O M CO2 + 2H2) have a more prevailing role [4]. The CO2 formed by the secondary water gas reaction could be consumed by the Boudouard reaction, which could explain a slight increase observed in CO2 concentration when the gasification temperature rises. The concentration of CH4 produced mainly in the pyrolysis stage decreases with TG (see Fig. 4). The model results of
final gas composition (%vol)
ð16Þ
r add ¼
10
G
40 35 30 25 20 15 10 5
700
750
800
850
900
o
TG ( C) CO-Model value H2 -Model value
CO-Exp.value H2 -Exp.value
CO2 -Model value
CO2 -Exp.value
CH4 -Model value
CH4 -Exp.value
Fig. 6. Effect of gasification temperature (TG) on final gas composition obtained from qETG model in comparison with experiment data at c = 0.3 (kg/kg).
the final gas composition show a similar tendency to the experimental data, as seen in Fig. 6. Table 5 gives the comparison of H2/CO molar ratio (/H2 =CO ), gas yield (Vyield), and lower heating value (LHVproduct) between the experiment and the qETG model at c = 0.3 kg/kg. /H2 =CO increases with the increase of temperature, as expected in Fig. 6. The error of /H2 =CO between the experiment and the model increases with temperature but is limited to 10%. It is experimentally observed that Vyield increases with an increase of TG from 700 °C to 800 °C, since the endothermic reactions of tar cracking and char gasification (char–CO2 and char–steam) proceed forward when TG increases [27]. However Vyield at TG = 900 °C is slightly less than that obtained at TG = 800 °C. It might be attributed to the tar formation at elevated temperatures [39]. Furthermore, the gasification with pure steam has been known to produce a large content of tar (30–80 g/m3), compared to the gasification with steam–O2 mixture and air agents (4–30 and 2–20 g/ m3, respectively) [40]. In the qETG model where tar formation is not considered, the syngas yield increases with temperature. The error of Vyield is relatively big at 700 °C where the gasification reactions are far from equilibrium [23,27,29]. The lower heating value (LHVproduct, MJ/Nm3) decreases a little with the increase of temperature because the volume of product gas (Vproduct, Nm3/h) increases
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S.I. Ngo et al. / Applied Energy 88 (2011) 5208–5220 Table 5 Comparison of H2/CO molar ratio (/H2 =CO ), gas yield (Vyield), and lower heating value (LHVproduct) between experiment and qETG model at c = 0.3 kg/kg. TG (°C)
700 Exp. 0.72 0.35
/H2 =CO ðÞ 3
V yield
Nm kg
LHV product
MJ m3
12.4
Model 0.73 0.80 15.9
Error (%) 1.4 128 28
Exp. 1.16 1.36 12.0
900 a
Model
Error (%)
1.25 1.12
Exp.
1.50 1.3
11.7
11.7
Errora (%)
Model
1.66 1.1
7.8 17.6
13.4
9.6 18.2
12.4
6.0
valueModel prediction Error ð%Þ ¼ Experiment 100. Experiment value
relatively much with the rise of TG (see Vyield in Table 5), even though the methane production rate (nCH4 ) released from pyrolysis increases slightly with TG (see Eqs. (3) and (4)). The error of LHVproduct between the experiment and model is restricted to about 30%. To validate the qETG model in the wide range of c, experiment data taken from Karmakar and Datta [17] and Herguido et al. [23] are utilized. Figs. 7 and 8 show the effect of c on the final gas composition, comparing the qETG model values with the experimental data. Here, the two empirical equations of the pyrolysis stage are slightly modified because different feed materials are employed. The averages of root mean square errors (RMS) for the two comparisons between the model and experiment are 3.97% and 5.18%, respectively. The RMS was calculated by the following equation:
sffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi PN 2 k ðexperiment k Modelk Þ RMS ¼ N
60
final gas composition (%vol)
a
800 a
50
40
30
20
10
ð19Þ 0
where N is the number of data points. As observed from the two figures, an increase in c leads to the increase of H2 and CO2 concentrations, while CO and CH4 decrease [1,4,7,27]. These behaviors were the same as in our experimental observation. The final gas product composition in biomass gasification with respect to c is mainly influenced by the water–gas shift reaction [8,10,23,38]. As shown in Figs. 6–8, the qETG model predicts the tendency of the syngas composition with respect to TG and c reasonably well. Using the qETG model, the process performance of a 100 MW DFB gasifier is evaluated by parametric study on TG and c. 4.2. Parametric study Fig. 9 shows the effects of TG and c on (a) carbon conversion, (b) char residue, (c) gas yield and (d) lower heating value of the final dry gas product. The increase of TG promotes the pyrolysis and the char–gas reactions, then the carbon conversion (/C) increases, and consequently the gas yield (Vyield) increases. However, since CO2 increases, CH4 decreases (see Fig. 6), and Vyield increases (see Fig. 9c) as TG increases, the lower heating value of product (LHVproduct) decreases (see Eq. (4)), as is depicted in Fig 9d. Contrasting with the variation of carbon conversion, the char residue (/CR) decreases with an increase in T (see Eq. (2)). As the steam to fuel ratio (c) increases, the gasification reactions are also promoted (see Table 2). That results in the increase of carbon conversion (/C), gas yield (Vyield) and the decrease of char residue (/CR). The carbon conversion reaches to 1.0 at TG = 900 °C and c = 3.0, as shown in Fig. 9a, which means that all the carbon in biomass is converted into gas species through the first and second stages. By contrast, the char residue (Fig. 9b) tends to be zero. When both TG and c are low, the char residue shows a maximum value (=1.0). In this area, the char–gas reactions do not take place and only pyrolysis and a water–gas shift reaction occur to a limited extent. Fig. 10 illustrates the parametric study of TG and c for (a) additional fuel ratio (radd), (b) solid circulation ratio (rcir), (c) heat
0.6
0.8
1
1.2
1.4
1.6
1.8
γ (kg/kg) CO-Model value
CO-Exp.value [17]
H2 -Model value
H 2 -Exp.value [17]
CO2 -Model value
CO2 -Exp.value [17]
CH4 -Model value
CH4 -Exp.value [17]
Fig. 7. Effect of steam to fuel ratio (c) on final gas composition obtained from qETG model in comparison with experiment data taken from Karmakar and Datta [17] at TG = 750 °C.
recovery (ge), and (d) H2/CO molar ratio (/H2 =CO ). Since the rising of TG and c requires a higher heat demand (see Eqs. (11)–(13)) and leads to a smaller char residue (see Fig. 9b), the additional fuel ratio (radd) and solid circulation ratio (rcir) have to increase, as shown in Fig. 10a and b. When TG and c are over 780 °C and over 0.6, respectively, the additional fuel is required. The solid circulation ratio (rcir) is unrealistically too big at high values of TG and c. At a higher gasification temperature, the temperature difference between the riser and gasification becomes lower and the solid circulation (Msand) increases (see Eq. (14)). Furthermore, the higher steam amount and temperature need the higher energy demand (Hdemand) which is provided by hot silica sand to maintain the endothermic gasification reactions. Consequently, the solid circulation ratio (rcir) increases. It is desired for rcir to keep as low as possible, as mentioned above. The circulation ratio of bed particles was reported at about 15–20 [9] and about 30–80 times of the biomass dry feed rate [10,35,41]. Therefore the rcir over 100 may be avoided for practical operations, for which the gasification temperature is desirable to be below 850 °C, as shown in Fig. 10b. As seen in Eq. (18), the heat recovery (ge) was defined as a function of Vyield, LHVproduct, and radd. As the rises of both TG and c makes Vyield and radd increase whereas LHVproduct decreases, the conflicting behavior displays a maximum area of heat recovery (ge = 0.775) in the range of 0.0 6 c 6 0.1 and 850 6 TG 6 870 °C, as illustrated in Fig. 10c. The heat recovery tends to decrease with the increase of
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4.3. Effective operations for DFB steam gasifier of pine woodchips
final gas composition (%vol)
60
For FT synthesis, the common H2/CO ratio (/H2 =CO ) range is 1.8– 2.7, which depends on catalysts used [42,43]. The additional fuel (radd) is desirable to be as low as possible and the heat recovery (ge) should be as high as possible for energy efficiency. As discussed earlier, the solid circulation ratio (rcir) is concerned in the range from 35 to 80 [10,35,41,44] in this study. The contours of the four performance criteria are plotted in Fig. 11 with respect to TG and c, where the four criteria were chosen within 1.8 6 /H2 =CO 6 2.7, 0 6 radd 6 0.3, 35 6 rcir 6 80, and 0.73 6 ge. Here, the region between two solid bold lines of /H2 =CO is considered as an operating region suitable for FT synthesis. Since there is no optimum point simultaneously satisfying the above criteria, five points (A, B, C, D and E in Fig. 11) on the two bold lines of /H2 =CO are marked, which give an interesting feature, respectively. At point A, the heat recovery (ge) is maximized as well as the additional fuel ratio (radd) keeps lower. However, the solid circulation ratio is too high (rcir = 80). Point B indicates the highest heat recovery at /H2 =CO = 2.7 and rcir = 80, where the additional fuel ratio is about 0.29. At points A and B, heat efficiency of the product gas is high but operational ability is low because of the high solid circulation ratio. Point E is an alternative so that the solid circulation ratio is lowered (rcir = 55), while the heat recovery keeps high and the additional fuel ratio reaches a minimum. Points A, B and E are located at relatively high TG (758 6 TG 6 814°C) and low c (0.69 6 c 6 1.82) for high heat efficiency (0.751 6 ge 6 0.768) at the cost of the high solid circulation ratio (55 6 rcir 6 80). To increase /H2 =CO up to 2.7 and lower the solid circulation ratio (rcir = 64), point C is chosen where ge = 0.73 and radd 0.32. Point D located on the line of /H2 =CO = 1.8 has such a feature that the solid circulation ratio (rcir = 47) and the heat recovery (ge = 0.728) are relatively low. Points C and D are operated at relatively low TG (700 6 TG 6 741 °C) and high c (2.68 6 c 6 3.0), showing low solid circulation ratios (47 6 rcir 6 64). Since both points C and D are
50
40
30
20
10
0
0.5
1
1.5
2
2.5
γ (kg/kg) CO-Model value
CO-Exp.value [23]
H2 -Model value
H2 -Exp.value [23]
CO2 -Model value
CO2 -Exp.value [23]
CH4 -Model value
CH4 -Exp.value [23]
Fig. 8. Effect of steam to fuel ratio (c) on final gas composition obtained from qETG model in comparison with experiment data taken from Herguido et al. [23] at TG = 750 °C.
c except for in the lower region of c, because an excess amount of steam needs additional fuel and deteriorates the heat recovery. Fig. 10d depicts the H2/CO molar ratio (/H2 =CO ) with respect to TG and c. As shown in Figs. 6–8, the /H2 =CO increases obviously when both TG and c rise. The /H2 =CO computed by the model has about a maximum of 5 in the range investigated.
(b) 1 φCR (-)
1
0.8
0.4 900
0 700
3 2
800
800
1 700
o
TG ( C)
0
900
γ (kg/kg)
(d) 18 16
3
2
0
γ (kg/kg)
o
2.5
1
2
3
TG ( C)
(c) Vyield (Nm3 /kg, dry basis)
0.5
0.6
LHVproduct (MJ/Nm , dry basis)
φC (-)
(a)
1.5 1 0.5 900 800 o
TG ( C)
700
0
1
2
γ (kg/kg)
3
14 12 10 700 800 o
TG ( C)
900
3
2
1
0
γ (kg/kg)
Fig. 9. Parametric studies of TG and c: (a) carbon conversion (/C), (b) char residue (/CR), (c) gas yield (Vyield), and (d) lower heating value (LHVproduct).
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(a)
(b)
0.8
rcir (kg/kg)
r add (kg/kg)
0.6 0.4
200
0.2 3
0 900 0
TG ( C)
γ (kg/kg)
o
TG ( C)
(d)
0.75 0.7 0.65
1 700
0
γ (kg/kg)
6
φH2/CO (-)
0.8
4 2
0 900 900
0
3 2
800
1 700
o
(c)
0 900
2 800
ηe (-)
400
1
800 2
γ (kg/kg)
3
700
3 2
800 1
o
TG ( C)
o
TG ( C)
700 0
γ (kg/kg)
Fig. 10. Parametric studies of TG and c: (a) additional fuel ratio (radd), (b) solid circulation to fuel ratio (rcir), (c) heat recovery (ge), and (d) H2/CO molar ratio (/H2 =CO ).
Fig. 11. Contours of process performance criteria for a 100 MW DFB gasifier calculated by the qETG model with respect to TG and c.
situated at the border of the operating conditions, the two points are dependent on the operating boundary of TG and c. Process performance for the 100 MW DFB gasifier is evaluated at the five interesting points in Table 6. Since each point has a particular value for the eight performance criteria, one can choose an
operating condition suitable for a purpose of the producer gas. However, it is worth noting that the process performance evaluation is valid only if the parameters of the empirical equations used in the qETG model are well estimated for a given feedstock and gasifier.
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Table 6 Performance evaluation for steam gasification of pine woodchips in a 100 MW DFB gasifier at five interesting points of Fig. 11. Points
TG (°C)
c (kg/kg)
/C
/CR
Vyield (Nm3/kg)
LHVproduct (MJ/Nm3)
radd
rcir
ge
/H2 =CO
A B C D E
814 791 741 700 758
0.69 1.82 3.0 2.68 1.12
0.73 0.80 0.81 0.74 0.73
0.78 0.59 0.56 0.75 0.80
1.29 1.50 1.47 1.10 1.22
12.68 11.88 12.14 13.39 13.02
0.173 0.308 0.349 0.219 0.159
80 80 64 47 55
0.768 0.751 0.730 0.728 0.763
1.8 2.7 2.7 1.8 1.8
5. Conclusions In the experimental work, the pine woodchips of particle sizes of under 10 mm were used in a CFB (circulating fluidized-bed) steam gasifier as a feedstock. The result showed the final gas compositions about 26–42% of H2, 25–37% of CO, 16–19% of CO2 and 8– 11% of CH4, when the steam to fuel ratio (c) was 0.3 kg/kg and the gasification temperature (TG) was from 700 °C to 900 °C. The quasi-equilibrium three-stage gasification (qETG) model was proposed to investigate the effects of TG and c on DFB (dual circulating fluidized-bed) process performances such as (1) Carbon conversion (/C), (2) char residue (/CR), (3) gas yield (Vyield), (4) lower heating value (LHVproduct), (5) additional fuel ratio (radd), (6) solid circulation ratio (rcir), (7) heat recovery (ge), and (8) H2 to CO molar ratio (/H2 =CO ). The qETG model includes three stages: biomass pyrolysis, char–gas reaction, and gas phase reaction, where empirical equations are involved to take into account the deviation from theoretical equilibrium reactions. Final gas compositions predicted by the model had a reasonable agreement with our experimental data and previous studies from the literature. The prediction accuracy of the qETG model would be enhanced, if tar formation and cracking in the pyrolysis stage are considered. By parametric study, the eight process performances were evaluated for a 100 MW DFB gasifier with respect to TG and c. As both TG and c increase, the carbon conversion, the product gas yield, additional fuel ratio, solid circulation ratio, and H2 to CO molar ratio increase, while the lower heating value of product gas decreases. The heat recovery shows a maximum point because of a conflicting behavior among the product gas yield, the additional fuel ratio, and the lower heating value of product gas. Considering the syngas production suitable for FT synthesis, the contour of the four performance criteria (/H2 =CO , radd, ge, and rcir) was plotted as a function of TG and c. Since there was no general optimum point simultaneously satisfying the criteria, five operating conditions showing an interesting feature were suggested. In the region of high TG (758 6 TG 6 814 °C) and low c (0.69 6 c 6 1.82), high heat efficiency is found but the solid circulation ratio is too high. The low heat efficiency and low solid circulation ratio are observed in low TG (700 6 TG 6 741 °C) and high c (2.68 6 c 6 3). The contour of performance criteria provides a guideline for the selection of effective operating conditions. Acknowledgement This work is supported by Bilateral International Collaborative R&D program under the Ministry of Knowledge Economy, Korea. Appendix A
consist of CO, CO2, CH4, H2, and H2O. In the secondary pyrolysis (tar cracking step), the thermal cracking of tar takes place. Tar cracking in the gas phase has significant impact on the accurate prediction of pyrolysis gas composition. The products of tar cracking are composed of CO, CO2, H2, hydrocarbons (e.g., C2H6, C2H4, and C3H6), and inert tar [19,20,45]. Prasad and Kuester [7] proposed an empirical model obtained from a linear regression to predict the composition of pyrolysis gas, in which the gas components were expressed as a function of pyrolysis temperature and steam to fuel ratio. A simplified approach was reported by Sharma, and Gao and Li [46,47], herein the product gas of pyrolysis was assumed to be only CO, CH4, and H2O. The atomic balances determined the composition of these components. Sadaka et al. [48] calculated the pyrolysis gas composition using the elemental balance under the assumption that the product consists of CO, CO2, CH4, H2, and H2O. The approach proposed by Sadaka et al. [48] is used in this study to predict the composition of the gaseous product released from biomass pyrolysis. In this approach, volatiles and tar in biomass are considered to be completely converted to the five gaseous components mentioned above without any inert tar remaining. The elemental balances are written as:
nc ¼ nCO þ nCO2 þ nCH4
ðA:1Þ
nH ¼ 4nCH4 þ 2nH2 þ 2nH2 O
ðA:2Þ
nO ¼ nCO þ 2nCO2 þ nH2 O
ðA:3Þ
where n is the number of moles. The nitrogen content in biomass (if any) is assumed to form inert nitrogen gas. Thus the elemental balance for nitrogen is
nN ¼ 2nN2
ðA:4Þ
In order to close the elemental balance equations, two terms defined as follows are required [48]: The fraction of CO formation,
/CO=CO2 ¼
nCO nCO2
ðA:5Þ
and the fraction of CH4 formation,
/CH4 =H2 ¼
nCH4 nH2
ðA:6Þ
The two fractions are determined by pyrolysis experiments for a biomass used. The experimental data show that with an increase in temperature /CO=CO2 increases, while /CH4 =H2 decreases. It is supposed that the reaction between a certain part of char and CO2 by Boudouard reaction (C + CO2 ? 2CO) leads to the increase of /H2 =CO when the pyrolysis temperature increases [26]. Increase in the H2 concentration by the dehydrogenation of saturated hydrocarbons such as C2H6 and C3H8 might be the explanation for the decrease in /CH4 =H2 .
A.1. Pyrolysis A.2. Char–gas reaction The non-oxidative decomposition of biomass called pyrolysis may be divided into two main steps: (1) primary and (2) secondary pyrolysis. In the first step, biomass thermally decomposes into gases, tar, and char. The gases released from the primary pyrolysis
The second stage of the qETG model is the char–gas reactions, which are composed of a series of complex and competing reactions between biomass char and gasification reagents. However,
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the Boudouard reaction, primary char–steam reaction and secondary char–steam reaction are assumed to take place and to reach an equilibrium state in this study:
CðsÞ þ CO2 ðgÞ $ 2COðgÞ
ðIÞ
CðsÞ þ H2 OðgÞ $ COðgÞ þ H2 ðgÞ
ðIIÞ
CðsÞ þ 2H2 OðgÞ $ CO2 ðgÞ þ 2H2 ðgÞ
ðIIIÞ
reactant. Each reaction proceeds forward or backward to a certain degree ejej, called the reaction coordinate, in order to reach an equilibrium from the initial stage [50]. At the equilibrium state, the mole fraction of species ai is expressed as [50]:
yi ¼
P ni;0 þ j mi;j ej P n0 þ j mj ej
ðA:9Þ
Since the steam amount involved in the char–gas equilibrium reaction increases with the increase of gasification temperature, it is needed to determine the ratio of the steam amount involved in reactions (II) and (III) to the total water amount in this stage [3,18]. The steam participation ratio (b) is estimated from the char–gasification experimental data taken from Yoshida et al. [18]. In Table 3, b is expressed as a function of gasification temperature.
where ni,0 and n0 are the moles of species ai and the total moles of all species in the system at the initial state, respectively. vj is the total stoichiometric number of the jth reaction which is defined as:
A.3. Gas-phase reaction
yCO ¼
It was reported that the water–gas shift reaction dominates in biomass gasification at the temperature higher than 670 °C [23,27,38]:
COðgÞ þ H2 OðgÞ $ CO2 ðgÞ þ H2 ðgÞ
ðIVÞ
The methanation reaction C(s) + 2H2(g) ? CH4(g) in the second stage and the steam reforming reaction CH4(g) + H2O(g) ? CO(g) + 3H2(g) in the third stage of the qETG model are not supposed to take place. A.4. Calculation procedures
nC
3
2
1 1
2
3
0
7 6 2 07 7 6 nCO2 7 7 6 7 1 0 7 6 nCH4 7 76 7 7 6 0 27 7 6 nH2 7 7 6 7 0 0 5 4 nH2 O 5 0 0 nN2
0 2
4 2 0 0
0
0
0
/CO=CO2
0
0
0
0
1 /CH4 =H2
0
3
1 0
6n 7 60 6 H7 6 6 7 6 6 nO 7 6 1 6 7¼6 6n 7 60 6 N7 6 6 7 6 40 5 41 0
nCO
mi;j aj ¼ 0
mi;j
ðA:10Þ
i
By applying Eq. (A.10) to the reactions of the second stage (reactions I, II, III), the mole fraction of the gas species is expressed as:
nCO;0 þ 2e1 þ e2 n0 þ e1 þ e2 þ e3 nCO2;0 e1 þ e3 yCO2 ¼ n0 þ e1 þ e2 þ e3 nH2 ;0 þ e2 þ 2e3 yH2 ¼ n0 þ e1 þ e2 þ e3 nH O;0 e2 2e3 yH2 O ¼ 2 n0 þ e1 þ e2 þ e3 nCH4 ;0 yCH4 ¼ n0 þ e1 þ e2 þ e3 nN2 ;0 yN2 ¼ n0 þ e1 þ e2 þ e3
ðA:11Þ ðA:12Þ ðA:13Þ ðA:14Þ ðA:15Þ ðA:16Þ
2 K 1 ¼ y1 CO2 yCO
ðA:17Þ
y1 H2 O
ðA:18Þ
K2 ¼
yCO yH2
2 K 3 ¼ y2 H2 O yCO2 yH2
ðA:7Þ
DG0T G ;j K j ¼ exp RT G
! ðA:20Þ
where D0T G is calculated from standard Gibbs energy of formation:
X
DG0T G ;f ;i ¼ DG0T G ¼ ðDH0T G T G DS0T G Þ
ðA:8Þ
where vi,j is the stoichiometric number of species ai involved in the jth reaction. It is positive if ai is the product and negative if ai is the
ðA:21Þ
i
The molar enthalpy hi of a component i in the gas phase is defined by the enthalpy of formation D0fi and the molar heat capacity Cpi according to the following equation [50].
DH0T G ¼
X
m i hi
i
Z
ðA:22Þ
TG
cpi ðT G ÞdT G
ðA:23Þ
298:15
In the third stage, the thermodynamic equilibrium constant of water–gas shift reaction is first calculated by the same way as the second stage, then the non-equilibrium factor (j) is applied:
K 4 ¼ j K 4
j
ðA:19Þ
From the Gibbs free energy equation, we have:
hi ¼ DH0f ;i þ
In the second stage, the molar fraction of the gas species is calculated. For a stoichiometric multi-reaction system, a set of i species ai involved in j independent reactions can be written as:
X
X
The equilibrium constant of each reaction is required to determine the reaction coordinates:
Because the hydrogen and oxygen content in char decreases sharply as temperature increases [26,49], the char remaining after biomass pyrolysis is assumed to be pure carbon. The moles of carbon in volatiles are determined by the difference between fixed carbon content in the proximate analysis and carbon content in the ultimate analysis. The moles of H, O, and N are then calculated from the ultimate analysis. The composition of each species in pyrolysis gas is obtained by solving element-based balances. These compositions are used as the initial conditions of the equilibrium model in the second stage for char gasification. At the second and third stages of the qETG model, instead of using element-based balances, the compositions of gaseous components are calculated by using species-based balances. The equilibrium constant for each reaction is calculated from the Gibbs free energy change of the reaction [50]. At the third stage, the equilibrium constant (K4) of the water–gas shift reaction (reaction (IV)) is corrected by the non-equilibrium factor (j). In the first stage, six linear equations are established to calculate the compositions of pyrolysis gas:
2
vj ¼
where
K 4
ðA:24Þ
is the corrected equilibrium constant.
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