Thermodynamic equilibrium analysis of H2-rich syngas production via sorption-enhanced chemical looping biomass gasification

Journal Pre-proof Thermodynamic Equilibrium Analysis of H2-rich Syngas Production via Sorptionenhanced Chemical Looping Biomass Gasification Rei-Yu Chein, Wen-Huai Hsu PII:

S0960-1481(19)31588-5

DOI:

https://doi.org/10.1016/j.renene.2019.10.097

Reference:

RENE 12466

To appear in:

Renewable Energy

Received Date: 2 July 2019 Revised Date:

15 October 2019

Accepted Date: 17 October 2019

Please cite this article as: Rei-Yu Chein, Wen-Huai Hsu, Thermodynamic Equilibrium Analysis of H2-rich Syngas Production via Sorption-enhanced Chemical Looping Biomass Gasification, Renewable Energy (2019), doi: 10.1016/j.renene.2019.10.097 This is a PDF file of an article that has undergone enhancements after acceptance, such as the addition of a cover page and metadata, and formatting for readability, but it is not yet the definitive version of record. This version will undergo additional copyediting, typesetting and review before it is published in its final form, but we are providing this version to give early visibility of the article. Please note that, during the production process, errors may be discovered which could affect the content, and all legal disclaimers that apply to the journal pertain. © 2019 Published by Elsevier Ltd.

Highlights ►Sorption-enhanced chemical looping biomass gasification (SE-CL-BG) was studied. ►Syngas selectivity decreases with increase in oxygen carrier amount introduced. ►Higher H2 production can be obtained from SE-CL-BG with temperatures lower than 750°C. ►H2 production can be enhanced using more CO2 sorbent in SE-CL-BG.

1 2 3 4

Thermodynamic Equilibrium Analysis of H2-rich Syngas Production via Sorption-enhanced Chemical Looping Biomass Gasification Rei-Yu Chein, Wen-Huai Hsu

5 6 7

Department of Mechanical Engineering, National Chung Hsing University, Taichung City, Taiwan 40227

8 9

Abstract

10 11

In this study, thermodynamic equilibrium analysis of the sorption-enhanced chemical

12

looping biomass gasification (SE-CL-BG) using Fe2O3 as the oxygen carrier, CaO as the

13

CO2 sorbent, and CO2 or H2O as the gasifying agent for producing H2 rich syngas was

14

conducted. Based on the amount of OC introduced, highly selective syngas can only result

15

when a small amount of oxygen carrier is introduced. Due to carbon and hydrogen

16

oxidations, the yields of CO and H2, cold gas efficiency, and the second-law efficiency of

17

SE-CL-BG case were found to be lower than the conventional biomass gasification case in

18

which no oxygen carrier and CO2 sorbent were introduced. Compared with conventional

19

biomass gasification, the advantage of SE-CL-BG is that biomass gasification can be

20

operated at lower temperatures (500~750°C) with higher H2 yield due to the enhanced

21

water-gas shift reaction and lower heat duty due to heat release from the CO2 absorption

22

reaction. The computed results indicated that CaO loses the ability to absorb CO2 as the

23

temperature becomes higher than 800°C.

24 25 26

Keywords: sorption-enhanced chemical looping biomass gasification (SE-CL-BG),

27

thermodynamic analysis, gasifying agent, oxygen carrier, and CO2 sorbent.

28 29 30

1

31

Nomenclature

32

CaO/C

mole of CO2 sorbent/mole of carbon in biomass

33

CO2/C

mole of gasifying agent CO2/mole of carbon in biomass

34

Ex

exergy, kJ/kg

35




molar flow rate, kmol/hr

36

H2O/C

mole of gasifying agent H2O/mole of carbon in biomass

37

H2/CO

molar ratio of H2 to CO in syngas

38



mass flow rate, kg/s

39

OC/C

mole of oxygen carrier/mole of carbon in the feedstock

40

OC/biomass

mass of OC/mass of biomass

41

P

pressure, kPa

42

Q

specific heat input or output, kJ/kg

43

T

gasification temperature, °C

44

X

species conversion

45

y

component mole fraction in syngas

46

Y

species yield

47 48

Greek symbol

49

β

correction factor for biomass chemical exergy

50 51

Subscript

52

0

standard condition

53

agent

agent

54

bio

biomass

55

calc

calcination 2

56

ch

chemical

57

fr

fuel reactor

58

i

species i in flow stream

59

in

inlet

60

mix

gaseous mixture

61

out

outlet

62

ph

physical

63

syngas

syngas

64 65

Superscript

66

0

standard condition

67 68 69

Abbreviation

70

BG

biomass gasification

71

CL-BG

chemical looping biomass gasification

72

SE-CL-BG

sorption-enhanced chemical looping biomass gasification

73

CGE

cold gas efficiency

74

HHV

higher heating value, kJ/kg

75

LHV

lower heating value, kJ/kg

76

OC

oxygen carrier

77

RWGS

reverse water gas shift

78

WGS

water gas shift

79 80

3

81

1. Introduction

82

With the global increase in energy demand, limited amounts of conventional fossil fuels,

83

and the growing concerns for global warming, searching for alternative energy has received

84

much attention. Among the various alternative energy forms, hydrogen is considered as an

85

important future energy carrier [1]. Hydrogen is also an important raw material in the

86

chemical industry that can be used as a fuel in fuel cells to produce electrical energy. For

87

reasons of sustainability, the use of renewable fuel sources such as biomass for hydrogen

88

production has received considerable attention [2].

89 90

Biomass conversion into syngas is the first step for H2 production from biomass. Biomass

91

gasification (BG) is viewed as one of the most important and promising technologies for

92

converting biomass into syngas [3]. Generally, gasifying agents such as air, oxygen, steam or

93

their mixtures are required during the biomass gasification process to obtain high quality

94

syngas [4-6]. However, problems such as low calorific value, high operating cost, and high

95

heat requirement occur when these gasifying agents are used. Hence, exploring biomass

96

gasification technology aiming for low operating cost, system self-heating, and high quality

97

syngas are essential. Chemical looping biomass gasification (CL-BG) is proposed because it

98

can realize the above-mentioned targets [7]. CL-BG is originated from chemical looping

99

combustion (CLC) and it involves using an oxygen carrier that circulates between the air and

100

fuel reactors to provide the oxygen and heat needed for gasification [8]. In contrast to

101

traditional biomass gasification technologies, CL-BG has several potential benefits. Oxygen

102

carrier recycling can provide the oxygen needed for gasification, saving the cost of making

103

pure oxygen. The oxygen carrier oxidation reaction in the air reactor is exothermic while the

104

fuel reduction reaction in the reactor is endothermic. The heat for the endothermic reduction

105

reaction can thus be provided by the circulating oxidized oxygen carrier coming from the air 4

106

reactor at higher temperatures. Moreover, the oxygen carrier can also be used for catalyzing

107

tar cracking that reduces the tar content in the biomass case [9]. Hence, CL-BG seems to be a

108

promising alternative compared with other proposed approaches.

109 110

H2 production from CL-BG has received much attention because of the advantages stated

111

above. The study by Lou et al. [10] provided an extensive review of chemical looping H2

112

production technology. As pointed out by them, high process efficiency and CO2 capture can

113

be provided by the chemical looping technology. Udomsirichakorn and Salam [11] reviewed

114

research work on CaO-based CL-BG. They pointed out that H2O is usually adopted as the

115

gasifying agent in order to gain H2-enriched syngas. Yan et al. [12] proposed a three biomass-

116

based chemical looping hydrogen generation system. They reported that the highest cold gas

117

efficiency resulted from using the CaO-looping process. Kuo et al. [13] numerically simulated

118

a biomass gasification process using steam as the gasifying agent and integrated with an iron-

119

based chemical looping hydrogen production system. Using torrefaction to overcome the

120

inherent problems of raw biomass such as high moisture content, low energy and bulk

121

density, poor grindability, and hydroscopicity [14], they indicated that better system

122

performance can be obtained when torrefied wood was used as the feedstock for gasification

123

as compared with raw wood.

124 125

In obtaining high purity H2 from the reforming reaction, research efforts have focused on

126

an alternative concept that combines the reforming reaction with in situ CO2 separation,

127

namely sorption enhanced reforming [15]. CO2 capture in situ drives the reaction towards

128

higher conversion (Le Chatelier’s principle) and higher H2 yield. Recently, a sorption-

129

enhanced chemical looping reforming process that combines chemical looping and CO2

130

absorption into one process was proposed using CH4 and steam as the feedstock [16-18]. 5

131

Three interconnected fluidized bed reactors (i.e., fuel reactor, calcination reactor and air

132

reactor) are used to produce high-purity hydrogen without the need for CO2 and N2 separation.

133

The same principle can be extended to H2 production from biomass gasification. The

134

sorption-enhanced reaction for CO2 removal and chemical looping biomass gasification can

135

be combined for H2 production, termed as SE-CL-BG. In SE-CL-BG, CO2 sorbent is

136

circulated along with the oxygen carrier, adsorbs CO2 in the fuel reactor, and desorbs CO2 in a

137

calcination reactor. In the study by Udomchoke et al. [19], sorption enhanced chemical-

138

looping reforming for hydrogen production from biomass was carried out using corn stover as

139

the feedstock. They pointed out that biomass conversion increases when the oxygen carrier

140

and CO2 sorbent amounts fed into the reformer increased. In the study by Detchusananard et

141

al. [20], integration and optimization of a sorption enhanced steam biomass gasification with a

142

solid oxide fuel cells and gas turbine system for the power and heat productions from biomass

143

were investigated using multi-objective optimization approach. They found that the maximum

144

exergy efficiency of 61.2% can be achieved for the system proposed. In the study by Fuch et

145

al. [21], sorption enhanced dual fluidized bed steam gasification of four different fuels was

146

studied. In addition to reaction temperature, they pointed out that the sorbent cycle rate is one

147

of the important factors affecting the H2 production. In the study by Schweitzer et al. [22], H2

148

production from biomass gasification was studied in terms of efficiency and feasibility. Based

149

on the techno-economic analysis, they concluded that an economic H2 production via sorption

150

enhanced gasification is possible.

151 152

Although extensive studies on CL-BG for syngas production can be found in the literature,

153

relatively little information is available on SE-CL-BG performance with various gasifying

154

agents. The present study carries out detailed thermodynamic SE-CL-BG analysis using CaO

155

and Fe2O3 as CO2 sorbent and oxygen carrier. Using the fuel reactor operating temperature as 6

156

the primary parameter, the oxygen carrier/biomass molar ratio, type of gasifying agent, and

157

amount of CO2 sorbent effects on the carbon conversion, H2 yield, CO yield, cold gas

158

efficiency, and second-law efficiency were investigated.

159 160

2. Model development

161

Biomass gasification for syngas production based on thermodynamic equilibrium theory is

162

modeled in this study. The Aspen Plus process simulator (v10.1) is used to establish this

163

model. The following assumptions are made to simplify the model:

164

(1) The process is operated under steady-state and steady-flow conditions.

165

(2) The biomass is considered made up of carbon, hydrogen, oxygen and nitrogen only.

166

(3) All chemical reactions are assumed to be at equilibrium.

167

(4) Pressure drops and heat losses are neglected.

168

(5) Char is assumed as carbon and tar formation is neglected [23].

169

(6) Ash is an inert species and is discharged into the environment at ambient temperature.

170

(7) CO2 sorbent catalytic activity is not considered.

171

(8) The product stream contains H2, CO, CO2, CH4, N2, H2O and solid carbon only and all

172

gases are ideal gases.

173

In general, biomass gasification mainly produces different products: gases, condensable tars,

174

and solids (char and ash). Tars are condensable organic compounds formed in

175

thermochemical processes such as gasification and pyrolysis [24]. Although tar formation in

176

the biomass gasification process is inevitable, gases are the predominant product in

177

gasification. Therefore, tar formation is usually neglected in the theoretical analysis of

178

biomass gasification [6,25]. Moreover, gaseous hydrocarbon species such as C3H6, C3H8, and

179

C4H10 are neglected in the product stream since their amounts are relatively small as

180

compared with CH4. 7

181 182

Three approaches are used to evaluate H2 production performance from biomass

183

gasification. In the first approach, conventional biomass gasification process is carried out,

184

termed as BG. In BG, no oxygen carrier or CO2 sorbent is introduced. The results from BG is

185

used as a comparison basis in this study. In the second approach, chemical looping biomass

186

gasification is carried out, denoted as CL-BG. In this approach, oxygen carrier is introduced

187

for providing oxygen. The third approach carries out sorption-enhanced chemical looping

188

biomass gasification, denoted as SE-CL-BG. The CO2 sorbent is also introduced in addition

189

to the oxygen carrier. The Aspen Plus simulation flow diagrams and chemical reactions that

190

take place for each approach are described in the following.

191 192

2.1 BG

193

Figure 1(a) shows the BG flow process established in Aspen Plus. Biomass is regarded as a

194

non-conventional component and defined using the ultimate and proximate analyses. The

195

biomass feedstock (BIOM) is decomposed into its elements (C, H, O, N, S, etc) through the

196

RYield reactor (DECOM). The elemental stream (IN-GASF) enters the fuel reactor (FUEL-

197

REA) modeled as an RGibbs reactor. In conventional biomass gasification the fuel reactor is

198

known as the gasifier. The gasifying agent (AGENT) is fed to the fuel reactor along with the

199

biomass. In an RGibbs reactor the chemical equilibrium biomass gasification is modeled by

200

minimizing the Gibbs free energy and the product at the fuel reactor outlet is the syngas

201

denoted as SYN. To simplify the analysis, the syngas is assumed to be mainly composed of

202

CO, H2, CH4, CO2, H2O, and solid carbon. The Peng-Robinson equation of state is used to

203

model the chemical species thermophysical properties involved in the computation.

204 205

8

206 207 208

(a)

AGENT SYN2

209 210

TEFSYN

TEFAGENT

211 212

AGENT2

DECOM

SYN

213

QDECOM

214

FUEL-REA

BIOM

215 IN-GASF

216 217

QGASI

218

Q

219 220 221

AGENT

222

AIR

SYN2

(b)

FG2

TEFSYN

TEFAGENT

TEFAIR

TEFFG

223 224

AGENT2

225 226

DECOM

QDECOM

AIR2 FUEL-REA

AIR-REA

SYN

FG SEP2

SEP1 BIOM

ME+SYN MEO+N2

227

IN-GASF

228

ME QGASI Q

229 230

MEO

231 232

AGENT

234

(c)

TEFSYN

TEFAGENT

AIR

CO22

SYN2

233

FG2

TEFAIR

CALC

TEFFG

TEFCO2

235 237 238

SYN

QDECOM

236

CO21 DECOM

AIR2

CO2+SOL

SEP1

AGENT2

SEP2

FUEL-REA IN-GASF

ME+SYN

ME+CACO3 MEO+N2

QCALC

240 242

FG SEP3

BIOM

239 241

AIR-REA

QAIRR ME+CAO QGASI

Q

Q

Q MEO+CAO

243 244 245 246 247 248

Figure 1 Process flowsheet for the biomass gasification reaction based on Aspen Plus. (a) conventional biomass gasification (BG), (b) chemical looping biomass gasification (CL-BG), and (c) sorption-enhanced chemical-looping biomass gasification (SE-CL-BG).

249 250 251

In Figure 1(a), TEFAGENT and TEFSYN are blocks for evaluating the total exergy for

252

AGENT and SYN streams, respectively. These blocks can be defined through the ACM 9

253

provided in Aspen Plus. For a flow stream (mixture of gas and liquid in general), the total

254

exergy is the sum of physical exergy Exph and chemical exergy Exch [26],

255

Ex = Ex + Ex 

256

where the physical exergy is calculated by,

257

Ex = h  − h   − T s  − s  

258

and the chemical exergy is calculated by,

259

   # Ex  = v ∑ y, Ex ,, + 1 − v  ∑ y,!" Ex ,,!"

260

In these equations h is the enthalpy, s is the entropy, v is the gas mass fraction, n is the

261

number of gaseous species, n! is the number of liquid species, yi is mole fraction of species i,

262

 Ex  is the standard chemical exergy. The superscript and subscript ‘0’ stand for standard

263

conditions (T0=298 K and p0= 1 atm), the subscript mix is the liquid and gas species mixture,

264

gas is the gas phase, and liq is the liquid phase.

(1)

(2a)





(2b)

265 266

The biomass exergy can be evaluated as [27],

267

Ex% = β LHV%

268

β is a correction factor for chemical exergy of biomass. Based on statistical correlations

269

developed by Szargut [28] for solid biofuels, β can be evaluated as,

270

β=

271

where C, H, N, and O are the mass fraction of carbon, hydrogen, nitrogen, and oxygen

272

obtained from ultimate analysis. LHV%; is the lower heating value of biomass which can be

273

evaluated as [29],

274

LHV% = HHV% − 137.58228

275

where higher heating value of biomass HHV%; is provided from proximate analysis of

276

biomass.

(3)

++,..//12 .3+435/1[,.73//1], .+439/1 2.+:+5/1

(4)

(5)

10

277 278

In the gasifier, biomass is firstly devolatilized to produce volatile matter, tar, char, and

279

syngas during a pyrolytic process [30]. The general reaction can be written as [31],

280

C H: O → char + tar + syngas

281

As mentioned above, the tar formation is neglected in this study. The gasification reactions

282

depend on the gasification agent used. Using steam and CO2 as gasification agents, the main

283

reactions are [25],

284

Boudouard reaction:

285

 C + CO: ↔ 2CO, ∆H:4KL = 172.4 kJ/mol

286

Carbon partial oxidation:

287

 C + 0.5O: ↔ CO, ∆H:4KL = −112 kJ/mol

288

Carbon combustion:

289

 C + O: ↔ CO: , ∆H:4KL = −393 kJ/mol

290

Water-gas reaction:

291

 C + H: O ↔ CO + H: , ∆H:4KL = 131 kJ/mol

292

Water-gas shift reaction:

293

 CO + H: O ↔ CO: + H: , ∆H:4KL = −41 kJ/mol

294

CH4 formation:

295

 C + 2H: ↔ CH+ , ∆H:4KL = −74.9 kJ/mol

296

Hydrogen combustion:

297

 H: + 0.5O: ↔ H: O, ∆H:4KL = −242 kJ/mol

298

Methane-steam reforming:

299

 CH+ + H: O ↔ CO + 3H: , ∆H:4KL = 206 kJ/mol

(14)

300

 CH+ + 2H: O ↔ CO: + 4H: ,∆H:4KL = 165 kJ/mol

(15)

(6)

(7)

(8)

(9)

(10)

(11)

(12)

301

11

(13)

302

2.2 CL-BG

303

In Figure 1(b) the CL-BG flow process is presented. An air reactor (AIR-REA) is

304

introduced in addition to the fuel reactor as compared with Figure 1(a). Based on the CLC

305

concept, the metal oxide oxygen carrier (MEO) is introduced to provide oxygen for reactions

306

taking place in the fuel reactor. The reduced metal oxide oxygen carrier (ME) circulates into

307

the air reactor and re-oxidized through the combustion with air. It is assumed that the

308

equilibrium oxidation reactions also occur in the air reactor operated adiabatically. As shown

309

in Figure 1(b), the product from the air reactor is sent to a separator (SEP2) at which the re-

310

oxidized oxygen carrier (MEO) and N2 are separated. Since the oxidation reaction in the air

311

reactor is a highly exothermic reaction, the re-oxidized oxygen carrier has high temperature

312

and can provide the energy required for endothermic reactions as it circulates back to the fuel

313

reactor.

314 315

In this study, Fe2O3 is used as the oxygen carrier because of its high reactivity and

316

environmental friendliness [32]. addition to the reactions described in Eqs. (6)~(15), possible

317

chemical reactions between oxygen carrier and pyrolytic products in the fuel reactor are [32],

318

 CO + 3Fe: O3 → CO: + 2Fe3 O+ , ∆H:4KL = −37.67 kJ/mol

(16)

319

 CO + Fe: O3 → CO: + 2FeO, ∆H:4KL = −3.217 kJ/mol

(17)

320

 H: + 3Fe: O3 → H: O + 2Fe3 O+ , ∆H:4KL = −2.011 kJ/mol

(18)

321

 H: + Fe: O3 → H: O + 2FeO, ∆H:4KL = 32.438 kJ/mol

(19)

322

 3CH+ + 4Fe: O3 → 8Fe + 3CO: + 6H: O, ∆H:4KL = 824.852 kJ/mol

(20)

323

 CH+ + 4Fe: O3 → 2H: O + CO: + 8FeO, ∆H:4KL = 317.87 kJ/mol

(21)

324

 CH+ + 3Fe: O3 → 2H: + CO + 2Fe3 O+ , ∆H:4KL = 221.76 kJ/mol

(22)

325

 C + 3Fe: O3 → CO + 2Fe3 O+ , ∆H:4KL = 133.65 kJ/mol

(23)

326

 3C + 2Fe: O3 → 3CO: + 4Fe, ∆H:4KL = 467.97 kJ/mol

(24)

12

327 328

 C + 2Fe: O3 → CO: + 4FeO, ∆H:4KL = 164.877 kJ/mol

(25)

In the air reactor, the reduced oxygen carrier is re-oxidized. The possible reactions are

329

[33],

330

 3O: + 4Fe → 2Fe: O3 , ∆H:4KL = −826.0 kJ/mol

(26)

331

 O: + 4FeO → 2Fe: O3 , ∆H:4KL = −554.68 kJ/mol

(27)

332

 O: + 4Fe3 O+ → 6Fe: O3 , ∆H:4KL = −483.07 kJ/mol

(28)

333 334

As shown in Figure 1(b), two separators are used to separate the gas mixture and solid

335

particles. In SEP1, solids (Fe, FeO, Fe2O3, Fe3O4, and carbon) are separated from the gas

336

mixture and enter the air reactor where air is introduced to oxidize Fe, FeO, Fe3O4 and carbon.

337

The air amount introduced to the air reactor is such that all of the carbon is burned into CO2

338

and all of the Fe and Fe oxides are oxidized into Fe2O3. This can be accomplished using the

339

design specs flow sheeting option in Aspen Plus. In SEP2, Fe2O3 is separated from the gas

340

mixture and circulates into the fuel reactor. In addition to TEFAGENT and TEFSYN,

341

TEFAIR and TEFFG are added into the CL-BG to evaluate the exergies of air (AIR) and flue

342

gas (FG). From the above analysis, CL-BG is a complex process. Multiple reactions occur in

343

the fuel reactor that competes with each other, leading to the final gasified product at the fuel

344

reactor outlet.

345 346

2.3 SE-CL-BG

347

According to Le Chatelier’s principle, chemical reaction can be shifted to the product side

348

if one of the product species is selectively removed [34]. For biomass gasification, Florin et

349

al. [35] and Masnadi et al. [36] reported that the in situ CO2 removal using CaO could

350

significantly enhance the H2 production and a high H2 concentration can be obtained. This is

351

known as the sorption-enhanced biomass gasification (SE-BG). In Figure 1(c), the SE-BG is 13

352

combined with CL-BG and becomes the sorption-enhanced chemical looping biomass

353

gasification (SE-CL-BG). In addition to Fe2O3, the CO2 sorbent CaO is also introduced into

354

the fuel reactor for CO2 adsorption. A calcination reactor (CALC) and a separator (SEP3) are

355

introduced as compared with the CL-BG case shown in Figure 1(b). In the fuel reactor CaO

356

absorbs CO2 and becomes CaCO3. The calcination reactor is used to regenerate CaO from

357

CaCO3 by heat addition according to the sorption reaction [37],

358

 CaO + CO: ↔ CaCO3 , ∆H:4KL = −178 kJ/mol

359

The CO2 absorption is an exothermic reaction and the heat released from this reaction can

360

reduce the heat duty for biomass gasification in fuel reactor. At SEP1, solid particles are

361

separated from the gas mixture and sent into the calcination reactor. At the calcination reactor,

362

CO2 is desorbed from CaCO3 and CaO is regenerated. Since the regeneration reaction is an

363

endothermic reaction, heat supply (QCALC) to the calcination reactor is required. In many

364

applications, the calcination reactor is operated at atmospheric pressure in a rotary- or

365

fluidized-bed system. From equilibrium curve of the reaction shown in Eq. (29), calcination is

366

found to occur at temperatures above 890°C in the presence of CO2 with a pressure of 1 atm

367

[37]. For the SE-CL-BG considered, the CO2 partial pressure is usually less than 1 atm. Under

368

such case, calcination reactor would permit the calcination reaction to be conducted at

369

temperatures lower than 890°C. In this study, the calcination reactor operating temperature is

370

set at 880°C [16-18].

(29)

371 372

The mixture at the calcination reactor outlet is sent to a separator (SEP2) at which CO2 is

373

separated from solid particles which can be collected for further sequestration or reuse. The

374

solid particles are sent to the air reactor (AIR-REA) at which the reduced oxygen carrier is re-

375

oxidized and carbon is burned with the air introduced into the air reactor. The product from

376

the air reactor is sent to a separator (SEP3) at which the re-oxidized oxygen carrier and CaO 14

377

are separated and circulated into the fuel reactor.

378 379

It is assumed that the air reactor is operated adiabatically and equilibrium oxidation

380

reactions occur. Since the oxidation reactions in the air reactor are highly exothermic

381

reactions, the re-oxidized oxygen carrier has high temperature and can provide the energy

382

required for endothermic reactions as it circulates back to the fuel reactor. In Figure 1(c),

383

TEFCO2 is added for evaluating exergy from CO2 separated from SEP2.

384 385

3 Results and discussion

386

3.1 Model verification and numerical parameters

387

The study of Renganathan et al. [38] is used to verify the model established in this study.

388

According to Renganathan et al. [38], the biomass is composed of carbon (C) 46%, hydrogen

389

(H) 6%, oxygen (O) 48% and without moisture content. The mass flow rate of

390

biomass m%; = 24 kg/s is used. Using the Ryield reactor the corresponding carbon molar

391

flow rate is 3310 kmol/hr. For the fuel reactor operated under isothermal condition, biomass

392

with inlet temperature of 25°C, and CO2 as gasifying agent with molar ratio of CO2/C=0.5

393

and inlet temperature of 25°C, Figure 2(a) shows the effect of gasification reaction

394

temperature (T) on biomass gasification performance using the flow process shown in Figure

395

1(a). In Figure 2(a), yi is the mole fraction of gas species at fuel reactor outlet, XC is the

396

carbon conversion, CGE is the cold gas efficiency, and XCO2 is the CO2 conversion. XC, CGE,

397

and XCO2 are defined as, YZ,[\ 2YZ,]^_`] YZ,[\

398

X1 =

399

CGE =

400

X15d =

(30)

]^_`] b/c]^_`]

(31)

[\ b/c[\ YZed,`f_g 2YZed,]^_`] YZed,`f_g

(32) 15

401

402

When H2O is used as the gasifying agent, H2O conversion is defined as, X/de =

Yhde,`f_g 2Yhde,]^_`] Yhde,`f_g

(33)

403

In Eqs. (30)~(33), F is the molar flow rate, m is the mass flow rate, and LHV is the lower

404

heating value. Good agreement is obtained with Figure 3 presented in the study by

405

Renganathan et al. [38]. Figure 2(b) presents the gasifying agent amount effect on the

406

gasification performance using the model built in Figure 1(a). Good agreement is also

407

obtained compared with Figure 6 in the study by Renganathan et al. [38]. Based on these

408

comparisons, the model developed in this study is correct and can be extended further for the

409

studies on CL-BG and SE-CL-BG. The numerical parameters used in computing Figure 2 are

410

used for the BC-LG and SE-CL-BG cases for comparison convenience.

411 412 413 414 415 416 417 418 419 420 421 422 423 424 425

16

426 427 428 429 430 431 432 433 1.6

436

+

1.2

+

439 440

1-XC yCO yH2 H2/CO yCO2 yCH4 CGE XCO2

1.4

437 438

4.5

(a)

435

X

4 3.5 3

1 2.5

441 442

0.8

443 444

X

+

0.6

X

X

X

X

X 2 1.5

445 446

0.4

448

X

+

447

1

+

0.2

+

+

+

+

+

0.5

+

449 450

0 500

600

700

800

o

900

1000

1100

0 1200

T ( C)

451 452 453

1.6

(b)

1-XC yCO yH2 H2/CO yCO2 yCH4 CGE XCO2

454 455

1.4

+

456 457 458 459

1.2

x 1+

+

460 461 462 463 464 465 466 467 468 469 470

+

0.8

+ x x

0.6

x

x +

+

x +

x +

x +

x

x

x

+

+

+

x

0.4 0.2 0

0

0.1

0.2

0.3

CO2/C 17

0.4

0.5

H2/CO

434

471 472 473 474 475

Figure 2. Model verification by comparing gasification performance with the results reported by Renganathan et al. [38]. (a) gasification temperature effect, and (b) effect of gasifying agent CO2 amount.

476 477 478 479 480 481

3.2 Oxygen carrier amount effect on CL-BG and SE-CL-BG performance

482

In CL-BG and SE-CL-BG, the oxygen carrier is circulated between the fuel and air

483

reactors. In Figure 3 the amount of oxygen carrier effect on CL-BG is shown for the fuel

484

reactor operation temperature which varies from 700°C to 900°C using CO2 as the gasifying

485

agent with the molar CO2/C ratio of 0.5. In Figure 3 the oxygen carrier mass flow rate is

486

normalized by the biomass mass flow rate, denoted as OC/biomass. In Figure 3(a) the mole

487

fractions are presented for CO, H2, CO2, H2O and CH4 at the fuel reactor outlet as a function

488

of OC/biomass. It is seen that yH2 and yCO decrease while yH2O and yCO2 increase as the

489

OC/biomass increases. From the variations in yH2, yCO, yH2O, and yCO2, it is expected that a

490

complex interaction between the water-gas reaction (Eq. (10)), water-gas shift reaction (Eq.

491

(11)), carbon combustion (Eq. (9)), hydrogen combustion (Eq. (13)), and methane-steam

492

reforming reaction (Eqs. (14) and (15)) occur as more oxygen is provided from the oxygen

493

carrier. With more oxygen is provided in this range, it is also expected that carbon and

494

hydrogen combustions (Eq. (9) and Eq. (13)) will gradually dominate. For OC/biomass>12.2,

495

all H and C are converted into H2O and CO2. That is, complete carbon and hydrogen

496

oxidations are achieved. In this OC/biomass range the gasification becomes the combustion.

497 498

In the fuel reactor reactions between the oxygen carrier and biomass also take place. Figure

499

3(b) presents the yields for reduced Fe-based oxides (Fe, FeO, Fe2O3, and Fe3O4), denoted as

500

Yi, at the fuel reactor. The yields for these products are defined as the ratio of mass flow rate 18

501

of Fe-based oxide to the mass flow rate of biomass. When the OC/biomass is small, Fe2O3 is

502

mainly reduced into Fe as indicated in Eq. (24) because of insufficient oxygen provided from

503

the OC. With a higher OC/biomass ratio, Fe2O3 is reduced into Fe and FeO according to Eqs.

504

(24) and Eq. (25). In this range the reduction of Fe2O3 into FeO is gradually dominated as the

505

OC/biomass increases. When the OC/biomass>12.2, FeO remains constant while Fe2O3

506

increases as the OC/biomass increases. This indicates that the gasification becomes

507

combustion and excessive oxygen carrier is provided. Note that no Fe2O3 reduced to Fe3O4 is

508

found for the OC/biomass range studied. In addition to the gaseous species reaction with the

509

oxygen carrier described in Eqs. (16)~(22); the oxygen carrier may also react with solid

510

carbon as shown in Eqs. (23)~(25). That is, the result shown in Figure 3(b) is an overall yield

511

contributed from the reactions shown in Eqs. (16)~(25).

512

Using CO2 as the gasifying agent, Figure 3(a) shows that yCO increases while H2 decreases

513

as T increases for 0
514

performance will be addressed in detail in the next section. The reaction temperature effect on

515

the Fe-oxide yields at the fuel reactor outlet is insignificant as shown in Figure 3(b).

516

Figure 4 presents the OC/biomass effect on SE-CL-BG performance with the fuel reactor

517

operating temperature varied from 700 to 900°C. In addition to introducing CO2 as the

518

gasifying agent, CaO is introduced as the CO2 sorbent. For the purpose of comparing the

519

results with those shown in Figure 3, the same amount of CO2 is used for computing the

520

results shown in Figure 4. The amount of CaO introduced is equal to that of CO2, ie, CaO/C=

521

CO2/C=0.5. From Figure 4(a), higher yH2 can be obtained compared with the CL-BG case

522

shown in Figure 3(a) for the T=700°C case. For the T=800 and 900°C cases, the variations in

523

yH2, yCO, yH2O, and yCO2 are about the same as those shown in Figure 3(a). This indicates that

524

CaO loses it is capability for CO2 removal when the reaction temperature is higher than

525

800°C. More detailed discussion on the temperature effect on SE-CL-BG will be addressed in 19

526

the next section. In Figure 4(b) the yields of reduced Fe-based oxides corresponding to Figure

527

4(a) are shown. It can be seen that the addition of CaO produced more pronounced effect on

528

Fe2O3 reduction for 0
529

4(b).

530 531

Note that the OC/biomass effect on LC-BG and SE-CL-BG using H2O as the gasifying

532

agent can also be performed by replacing CO2 as described in Figures 3 and 4. Similar results

533

as those shown in Figures 3 and 4 can be obtained when H2O is used as the gasifying agent.

534

That is, complete biomass oxidation can be found when the OC/biomass is higher than 12.2

535

and the Fe oxide reductions are independent of the gasifying agent.

536 537 538 539 540 541 542 543 544 545 546 547 548 549 550 551 552 553 554 555 556 557 558 559

20

560 561 562 563 564 565 566 567 568 569 570 571 572 573 574 575 576 577 578 579 580 581 582 583 584 585 586 587 588 589 590 591 592 593 594 595 596 597 598 599 600 601 602 603 604 605 606 607 608 609

21

610 611 612 613

Figure 3 OC flow rate effect on CL-BG using CO2 as the gasifying agent with CO2/C=0.5. Fuel reactor operating temperature varies from 700°C to 900°C. (a) gas species mole fractions at the fuel reactor outlet, (b) yields of Fe oxides at reactor outlet.

614 615 616 617 618 619 620 621 622 623 624 625 626 627 628 629 630 631 632 633 634 635 636 637 638 639 640 641 642 643 644 645 646 647 648 649 650 651 652 653 654 655 656 657 658 659

22

660 661 662 663 664

Figure 4 OC flow rate effect on SE-CL-BG using CO2 as the gasifying agent and CaO as CO2 sorbent. CO2/C=CaO/C=0.5. Fuel reactor operation temperature varies from 700°C to 900°C. (a) gas species mole fractions at the fuel reactor outlet, (b) yields of Fe oxides at reactor outlet.

665 666 667 668 669 670 671 672 673

3.3 Fuel reactor temperature effect on biomass gasification performance

674

From the results shown in Figures 3 and 4, syngas production with high H2 and CO mole

675

fractions can only be obtained when the OC/biomass ratio is small. To examine the fuel

676

reactor operating temperature effect on gasification performance, the amounts of OC,

677

gasifying agent CO2, and CO2 sorbent are fixed as OC/C= CO2/C=CaO/C=0.5. Figure 5

678

presents comparisons of yH2, yCO, XC, and XCO2 between BG, CL-BG, and SE-CL-BG. Figure

679

5(a) shows that higher yCO from BG results compared with CL-BG and SE-CL-BG cases for

680

the T>700°C. This result indicates that with the introduction of oxygen carrier in CL-BG and

681

SE-CL-BG, carbon oxidation producing CO2 instead of CO is more favored due to the oxygen

682

supply when gasification temperature is high. With CaO introduced into SE-CL-BG, a slight

683

decrease in yCO from SE-CL-BG is obtained compared with that from BG and CL-BG cases

684

for T<700°C. However, higher yCO results from SE-CL-BG compared with CL-BG for the

685

700°C800°C, both CL-BG and SE-CL-BG have the same yCO. Since CO2

686

absorption by CaO is an exothermic reaction, the result indicates that CaO loses the CO2

687

absorption capability as the gasification temperature becomes higher than 800°C.

688 689

Figure 5(b) shows that the H2 production from CL-BG and SE-CL-BG is lower than that

690

from BG for the T>800°C case. For this temperature range, identical yH2 is obtained from CL-

691

BG and SE-CL-BG because CaO loses the ability to absorb CO2. This implies that H2O 23

692

formation is more favored when oxygen is supplied in gasification. With the introduction of

693

CO2 sorbent in the SE-CL-BG case, H2 production can be enhanced for T<800°C and has a

694

maximum value at T=725°C. For the T<750°C case, yH2 from SE-CL-BG is higher than that

695

from BG and CL-BG cases. This result agrees with the results reported using classical

696

sorption enhanced H2 production reported in the literature [15].

697 698

Figure 5(c) shows the carbon conversion as a function of gasification temperature. The

699

carbon conversion can be enhanced in CL-BG and SE-CL-BG for the T<800°C case. That is,

700

CO or CO2 production from carbon oxidation is more favored with oxygen provided from the

701

oxygen carrier. As shown in Figure 5(c), the CL-BG case has the lowest fuel reactor operating

702

temperature for complete carbon conversion. Complete carbon conversion results when the

703

temperature is high. In Figure 5(d), negative CO2 conversion for both BG and CL-BG results

704

with lower T regime. For the BG case, CO2 is produced rather than consumed when T is low.

705

As T>650°C, positive CO2 conversion is obtained from the BG case. With increased T, the

706

endothermic Boudouard reaction (Eq. (7)) is favored, converting CO2 into CO. Moreover, the

707

reverse water-gas shift reaction also contributes to increase the CO2 conversion and CO

708

concentration as T is high. The CO2 conversion variation trend from CL-BG is about the same

709

as BG. Therefore, both the Boudouard and reverse water-gas shift reactions also take place in

710

CL-BG as in BG. However, CO2 produced due to the introduction of oxygen from the oxygen

711

carrier lowers the CO2 conversion compared with the BG case. For the SE-CL-BG case,

712

positive CO2 conversion can be obtained for T<800°C due to CO2 absorption. As CaO loses

713

the CO2 adsorption ability for T>800°C, CO2 conversion becomes identical to the CL-BG

714

case.

715 716

The H2/CO ratio, CGE, η2nd, and heat duty of the fuel reactor corresponding to the results 24

717

discussed in Figure 5 are presented in Figure 6. Because of enhanced H2 production, a higher

718

H2/CO ratio can be obtained from the SE-CL-BG case for T<800°C as shown in Figure 6(a).

719

Using CO2 as the gasifying agent an H2/CO ratio lower than unity is obtained for T>800°C.

720

Because of lower CO and H2 production for T>800°C, as shown in Figures 5(a) and Figures

721

5(b), CGE for both CL-BG and SE-CL-BG is lower than that of the BG case as shown in

722

Figure 6(b).

723 724

In Figure 6(c), the second law efficiency is compared for the three approaches. The second-

725

law efficiency is defined as,

726

η:j =

727

In Eq. (34), Ex and Ex;op are the exergy input to and output from the system, respectively.

728

Based on Figure 1, Ex and Ex;op for each approach are described as follows.

729

BG:

730

Ex = Ex%; + Exqr + Exsp , Ex;op = Ext

731

CL-BG:

732

Ex = Ex%; + Exqr + Exsp + Exr , Ex;op = Ext + ExYu

733

SE-CL-BG:

734

Ex = Ex%; + Exqr + Ex ! + Exsp + Exr , Ex;op = Ext + Ex15: + ExYu

735

In Eqs. (35)~(37), Exqr and Ex ! are the exergies due to heat transfer to fuel and calcination

736

reactor, respectively. These two terms are defined as,

737

Exqr = Qqr 1 − T /T, Ex ! = Q ! 1 − T /T ! 

738

Where Qfr and T are the heat duty and fuel reactor operating temperature, respectively. Qcalc

739

and Tcalc are the heat supply and calcination reactor operating temperature, respectively. Note

740

that heat production may occur in some cases instead of heat supply for the fuel reactor. In

klmg

(34)

kn_

25

(35)

(36)

(37)

(38)

741

these cases the Exfr becomes the exergy output.

742 743

As shown in Figure 6(c), BG has the highest η:j among the three approaches for

744

T>800°C. This is due to a greater exergy contribution from the syngas for the BG case. In the

745

T<800°C case, BG has lower η:j compared with the CL-BG and SE-CL-BG cases. By

746

referring to the fuel reactor heat duty shown in Figure 6(d), higher η:j in CL-BG and SE-

747

CL-BG is due to less heat supplied to the fuel reactor in this temperature range. Figure 6(d)

748

shows that fuel reactor thermo-neutral operation is possible at T=580°C and 700°C for the

749

CL-BG and SE-CL-BG cases, respectively. For temperatures lower than these values, the fuel

750

reactor energy output is possible due to exothermic reactions in the fuel reactor. While for the

751

BG case, the heat duty is always positive for the temperature range studied. The energy output

752

also contributes to higher η:j for the CL-BG and SE-CL-BG cases shown in Figure 6(c).

753 754 755 756 757 758 759 760 761 762 763 764 765

26

766 767 768 769 770 771 772 773 774

0.6

(a)

1 (c)

775 776

0.9 0.5 0.8

777 778

0.7

0.4

0.6

XC

780

yCO

779 0.3

0.5

781 782

0.4 0.2 0.3

783 784 785 786

0 500

787

0.2

BG CL-BG SE-CL-BG

0.1

600

700

800

900

T (oC)

1000

1100

BG CL-BG SE-CL-BG

0.1 0 500

1200

600

700

800

900

T (oC)

1000

1100

1200

788 789

0.6

1

(b)

790 791

0.75 0.5 0.5

792 793

0.4 0.25

yH2

XCO2

794 795

(d)

0.3

0

796 797

-0.25 0.2

798 799

-0.5 BG CL-BG SE-CL-BG

0.1

800 801 802

0 500

600

700

800

900

T (oC)

1000

1100

BG CL-BG SE-CL-BG

-0.75 -1 500

1200

600

700

800

900

T (oC)

1000

1100

1200

803 804 805 806

Figure 5 Biomass gasification performance comparisons as a function of the fuel reactor temperature using CO2 as the gasifying agent. CO2/C=OC/C=CaO/C=0.5. (a) yCO, (b) yH2, (c) XC, and (d) XCO2.

807 808

27

809 810 811 812 813 814 815 816 817 818 819 820 821 822 823

5

824

1

(a) BG CL-BG SE-CL-BG

825 4

826

0.8

827

0.7

828

0.6

η2nd

830

3

H2/CO

829

(c)

0.9

2

831

0.5 0.4

832

0.3

833

1

0.2

834 0.1

835 0 500

836

600

700

837 838

800

900

T (oC)

1000

1100

0 500

1200

10

1.2 (b)

600

700

800

900

T (oC)

1000

1100

1200

(d)

839 840

1

841 842

heat duty (MJ/kg)

5 0.8

844 845 846 847

CGE

843 0.6

0.4

849

-5

BG CL-BG SE-CL-BG

848 0.2

0

BG CL-BG SE-CL-BG

850 851

0 500

600

700

800

900

T (oC)

1000

1100

-10 500

1200

600

700

800

900

T (oC)

1000

1100

1200

852 853 854 855 856

Figure 6 Biomass gasification performance comparisons as a function of the fuel reactor temperature using CO2 as the gasifying agent. CO2/C=OC/C=CaO/C=0.5. (a) H2/CO ratio, (b) CGE, (c) second-law efficiency, (d) fuel reactor heat duty.

857 858

28

859 860 861 862 863 864 865 866 867 868 869 870 871 872 873

With the gasifying agent replaced by H2O, Figures 7 and 8 show the fuel reactor

874

temperature effect on the gasification performance from the three approaches. For these

875

results, the amounts of gasifying agent, oxygen carrier, and CO2 sorbent are chosen as

876

H2O/C=OC/C=CaO/C=0.5. As H2O is used as the gasifying agent, the forward or reverse

877

water-gas shift reaction (Eq.(11)) can be enhanced depending on the reaction temperature.

878

This results in a decrease in CO and increase in H2 amounts. Comparing Figures 7(a) and

879

7(b), with Figures 5(a) and 5(b) obtained by using CO2 as gasifying agent, lower yCO and

880

higher yH2 are obtained for the case with H2O as the gasifying agent. However, lower yCO and

881

yH2 are obtained in the CL-BG and SE-CL-BG cases as compared with BG case because of

882

carbon and hydrogen oxidations with oxygen provided from the oxygen carrier for T>800°C.

883

Due to CO2 absorption, slightly lower yCO results from SE-CL-BG for the T<800°C case

884

shown in Figure 7(a). Figure 7(b) shows yH2 is enhanced in the SE-CL-BG case and has a

885

maximum value at T=700°C. Because of the reverse water-gas shift reaction, a slight decrease

886

in yH2 occurs as T increases. As shown in Figure 7(c), carbon conversion can be enhanced

887

using H2O as the gasifying agent [25, 32]. Complete carbon conversion can be achieved at

888

T=650°C for the CL-BG and SE-CL-BG cases which is lower than that for BG case. As

889

shown in Figure 7(d), H2O is produced when T is low and this leads to negative H2O

890

conversion. The oxygen carrier introduction in the CL-BG case enhances H2O formation with 29

891

more negative H2O conversion resulting. Because of the water-gas shift reaction, more H2O

892

can be converted into H2 when CO2 is absorbed at T<800°C, positive H2O conversion can be

893

obtained from SE-CL-BG case. Due to H2O formation in the CL-BG and SE-CL-BG cases,

894

lower XH2O results compared with the BG case. Moreover, H2O conversion decreases and

895

may become negative as T increases for the T>800°C case because of the reverse water-gas

896

shift reaction.

897 898

Figure 8 presents the H2/CO ratio, CGE, η:j , and fuel reactor heat duty corresponding to

899

the results shown in Figure 7. Because of higher H2 and lower CO productions as H2O is used

900

as the gasifying agent, higher H2/CO is obtained as shown in Figure 8(a). When T is high,

901

H2/CO ratio with a value of unity can be obtained from the three approaches studied. Figures

902

7(a) and 7(b) show lower syngas content is obtained from the CL-BG and SE-CL-BG cases

903

for T>800°C compared with the BG case. This leads to lower CGE and η:j from the CL-BG

904

and SE-CL-BG cases as shown in Figures 8(b) and 8(c). However, higher η:j can be

905

obtained from the CL-BG and SE-CL-BG cases for T<700°C as shown in Figure 8(c). This is

906

due to the exergy contribution from negative heat duty shown in Figure 8(d). For the SE-CL-

907

BG case fuel reactor thermo-neutral operation results at T=650°C.

908 909 910 911 912 913 914 915

30

916 917 918 919 920 921 922 923 924 925

0.6

926

1 (c)

(a)

0.9

927

0.5 0.8

928 929

XC

932

0.6

yCO

931

0.7

0.4

930

0.3

0.5 0.4

933 0.2

934

0.3

935

BG CL-BG SE-CL-BG

0.1

936

0.1

937 0 500

938

600

700

939 940

0.6

941 942

BG CL-BG SE-CL-BG

0.2

800

900

T (OC)

1000

1100

0 500

1200

1

(b)

600

700

800

900

T (oC)

1000

1100

1200

(d)

0.75 0.5

943

0.5

944

947

0.25

XH2O

946

0.4

yH2

945

0.3

948 949

-0.25 0.2

950 951

0

-0.5

BG CL-BG SE-CL-BG

0.1

BG CL-BG SE-CL-BG

-0.75

952 953 954

0 500

600

700

800

900

T (oC)

1000

1100

1200

955

31

-1 500

600

700

800

900

T (oC)

1000

1100

1200

956 957 958

Figure 7 Biomass gasification performance comparisons as a function of the fuel reactor temperature using H2O as the gasifying agent. CO2/C=OC/C=CaO/C=0.5. (a) yCO, (b) yH2, (c) XC, and (d) XCO2.

959 960 961 962 963 964 965 966 967 968 969 970 971 972 973 974 975 976 977

14

1

(a) BG CL-BG SE-CL-BG

978 979

12

(c)

0.8

980 10

981

984

8

η2nd

983

0.6

H2/CO

982

6

0.4

985 4

986

BG CL-BG SE-CL-BG

0.2

987

2

988 989

0 500

600

700

990

800

900

T (oC)

1000

1100

0 500

1200

600

700

800

900

T (oC)

1000

1100

1200

991 992

10

1.2 (b)

8

994

1

7

996

heat duty (MJ/kg)

995 0.8

CGE

997 998

0.6

999 1000

0.4

1001 1002

1005

0 500

600

700

800

900

T (oC)

1000

1100

6 5 4 3 BG CL-BG SE-CL-BG

2 1 0

BG CL-BG SE-CL-BG

0.2

1003 1004

(d)

9

993

-1 -2 -3 500

1200

32

600

700

800

900

T (oC)

1000

1100

1200

1006 1007 1008 1009 1010

Figure 8 Biomass gasification performance comparisons as a function of fuel reactor temperature using H2O as the gasifying agent. CO2/C=OC/C=CaO/C=0.5. (a) H2/CO ratio, (b) CGE, (c) second-law efficiency, (d) fuel reactor heat duty.

1011 1012 1013 1014 1015 1016 1017 1018 1019 1020 1021 1022 1023 1024

3.4 CaO amount effect in biomass gasification.

1025

From the results shown in Figures 6~8, the benefit of using SE-CL-BG is that better

1026

performance can be obtained for T<800°C. Based on this result, the CaO amount effect on

1027

SE-CL-BG is examined in greater detail using CO2/C=0.5 as the gasifying agent. Figure 9

1028

shows that CaO/C varies from zero to 1 and its effect on H2 mole fraction, H2/CO ratio, CGE

1029

and heat duty. With an increased CaO/C ratio, the H2 concentration can be enhanced further

1030

as shown in Figure 9(a). With increased H2 amount, the H2/CO ratio is also increased as

1031

shown in Figure 9(b). Figure 9(c) shows that CGE is decreased with increased CaO/C ratio

1032

for T<800°C. As CO2 is removed in this temperature regime, less CO production is expected.

1033

Therefore, CGE decreases due to less CO contribution as CaO/C increases. Due to the

1034

exothermic CaO carbonation reaction, the fuel reactor heat duty can be reduced with

1035

increased CaO/C ratio as shown in Figure 9(d). The simulation results show that for a CaO/C

1036

ratio higher than 1.5, no further SE-CL-BG enhancement can be achieved.

1037 1038 1039

33

1040 1041 1042 1043 1044 1045 1046 1047 1048 1049 1050 1051 1052 1053 1054 1055 1056 1057 1058 1059 1060 1061 1062 1063 1064 1065 1066 1067 1068 1069 1070 1071 1072 1073 1074 1075 1076 1077 1078 1079 1080

34

1081 1082 1083 1084 1085 1086

Figure 9 CaO/C ratio effect on SE-CL-BG performance using CO2/C=0.5 as the gasifying agent. (a) yH2, (b) H2/CO ratio, (c) CGE, and (d) fuel reactor heat duty.

1087 1088 1089 1090 1091 1092 1093 1094 1095

4. Conclusion This study examined chemical looping biomass gasification based on the thermodynamic

1096

equilibrium theory. The following conclusions can be made based on the obtained results:

1097

(1) In the CL-BG and SE-CL-BG cases, syngas selectivity decreases with the increase in

1098

oxygen carrier amount introduced due to more complete carbon and hydrogen oxidations

1099

contained in the biomass.

1100

(2) The CO2 sorbent CaO loses its capability as the reaction temperature is higher than 800°C.

1101

Higher H2 production can be obtained from SE-CL-BG with temperatures lower than

1102

750°C.

1103

(3) The CO and H2 yields, cold gas efficiency and second-law efficiency in the CL-BG and

1104

SE-CL-BG cases are lower than those for conventional biomass gasification due to carbon

1105

and hydrogen oxidation as oxygen is supplied from the oxygen carrier.

1106

(4) H2 production can be enhanced using more CO2 sorbent in SE-CL-BG. With more CaO

1107

introduced, less heat duty to fuel reactor is required due to more heat release from

1108

exothermic CaO carbonation reaction.

1109

35

1110 1111 1112 1113 1114 1115 1116 1117 1118 1119 1120

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