Hydrogen production from steam methane

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Hydrogen production from two-step steam methane reforming in a fluidized bed reactor

Kang Seok Go

^a

, Sung Real Son

^a

, Sang Done Kim

^a,

*, Kyoung Soo Kang

^b

, Chu Sik Park

^b

aDepartment of Chemical and Biomolecular Engineering, Energy and Environment Research Center, Korea Advanced Institute of Science and Technology (KAIST), 335 Gwahangno, 373-1 Guseong-dong, Yuseong-gu, Daejeon 305-701, Republic of Korea

bHydrogen Energy Research Center, Korea Institute of Energy Research, 71-2 Jang-dong, Yuseong-gu, Daejeon, 305-343, Republic of Korea

a r t i c l e i n f o

Article history:

Received 27 August 2008 Received in revised form 24 November 2008

Accepted 25 November 2008 Available online 30 December 2008 Keywords:

Hydrogen production Steam methane reforming Chemical looping Iron oxides Fluidization

a b s t r a c t s

Based on the chemical looping (CL) concept, two-step steam methane reforming (SMR) system with a reduction/oxidation (redox) reaction of iron oxides to produce pure hydrogen is proposed. This system consists of fuel reactor (FR) and steam reactor (SR). The feasibility of producing pure hydrogen without any purifying steps in SR and the synthetic gas (syngas, COþH2) to the Fisher–Tropsch reaction in FR was evaluated. The effect of reaction temperature on the redox reactivity of iron oxides was determined in thermo- gravimetric analyzer (TGA) and the effects of gas velocity and reactant concentration on the reactivity at the isothermal condition in a fluidized bed reactor were determined. The conversion range of the optimum reaction (FeO/Fe3O4) in the two-step SMR system is found to be from 0 to 0.5 based on the iron oxides. In this condition, hydrogen (H2) with CO- free and synthetic gas (syngas) having the H2/(2COþ3CO2) molar ratio of 0.65 can be obtained in SR and FR, respectively. The solid inventory ratio of FR to SR to compensate the difference of reaction rates in the redox system of FeO/Fe3O4is determined to be 2.9.

1. Introduction

Research activities on H2production increase rapidly since it is an environmental friendly fuel that does not emit carbon dioxide which is known as one of the global warming gases[1].

In particular, H2as a fuel for the mobile and stationary power generation by using fuel cell has a great potential to increase H2demand in the near future[2–4].

Steam reforming of natural gas (i.e., mainly methane) is the most common and well developed technology used for H2

production at large scales[5,6]. This process mainly has three parts and the simplified flow diagram of this process is shown inFig. 1. Desulfurized methane is catalytically reformed at the temperature range of 970–1100 K and pressure up to 3.5 Mpa to produce synthetic gas (syngas) mixture of CO and H2[6]:

CH4þH2O/COþ3H2; DH⁺_{298 K}¼ þ241 kJ=mol (1) Exothermic reaction of CO with H2O (water–gas shift reaction) is then carried out to produce H2at 470–820 K[7–10].

COþH2O/CO2þH2; DH⁺_{298 K}¼ 41:1 kJ=mol (2) As a final stage, pressure swing adsorption (PSA) is commonly used to remove CO2(mainly), water, methane, and CO from the offgas. In this manner, production of H2from conventional SMR process has many complex reaction steps to purify H2so that the capital investment would be high and the process efficiency is reduced[11,12]. Cost of WGS and PSA practically accounts for about 30% of the total cost for producing H2as reported by Myers et al[13]. In this respect, many researches are underway to produce syngas and H₂by chemical looping

*Corresponding author. Tel.:þ82 42 350 3953; fax:þ82 42 350 3910.

E-mail address:[email protected](S.D. Kim).

A v a i l a b l e a t w w w . s c i e n c e d i r e c t . c o m

j o u r n a l h o m e p a g e : w w w . e l s e v i e r . c o m / l o c a t e / h e

doi:10.1016/j.ijhydene.2008.11.062

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(CL) system to improve conventional SMR process [14–23].

Chemical looping concept is similar to the chemical looping combustion (CLC) where oxygen needed to combust the fuel can be taken from the metal oxides in a fuel reactor (FR) and then the reduced metal oxides can be regenerated by air in air reactor (AR). Therefore, the metal oxides in this reaction are used as an oxygen carrier between FR and AR. As a result, CO2

can be inherently separated when the fuel is combusted [24,25].

For production of syngas or H2 from the CL system, the reported main studies can be summarized as: (A) partial oxidation of methane with oxygen carrier in FR and the recovery of heat from the product gases in AR to generate the electric power and steam or transfer of heat from AR to FR thorough the oxygen carrier [14–18], (B) carbonation and calcination with carbon carrier such as dolomite to improve the selectivity of H2 among the offgas in FR as well as the syngas production with the oxygen carriers in (A)[14,18–20], (C) generation of the reduced oxygen carrier from complete oxidation of methane with the inherent separation of CO2in FR and the reduced oxygen carrier then reacts with steam to produce pure H2in SR and regeneration of itself[21]. These reported studies on SMR with CL concept are summarized in Table 1.

Among the above three cases, (C) has rather simple concept for the process design since homogeneous solid material is circulating inside the reactor. With this concept, conventional SMR process for the production of purified hydrogen is divided into methane oxidation and steam reduction steps with oxygen carriers. If carbon deposition does not take place on the surface of the solids, production of

purified H2from the steam reduction is possible without any post-process such as WGS and PSA. Therefore, it can reduce the purification cost significantly compared to conventional processes.

However, researches of this concept have been sparse based on the experimental studies but only for the conceptual consideration[21]. Also, studies on mass transfer between gas and solid for the mass production of H2have not been conducted. Therefore, in the present study, iron oxide was selected as an oxygen carrier since it is environmental friendly having relatively low cost[26]and proposed the two-step SMR with concept (C) for production of pure H2in SR without post- process. In addition, gaseous products in FR can be changed into two ways which are full or partial oxidation of methane by iron oxides and these reactions are presented inEqs. (3)–(5).

Complete methane oxidation

4Fe3O4þCH4/12FeOþCO2þ2H2O;DH⁺_{298 K}¼463:7 kJ=mol (3) Partial methane oxidation

Fe3O4þCH4/3FeOþCOþ2H2; DH⁺_{298 K}¼280:9 kJ=mol (4) Steam reduction

3FeOþH2O/Fe3O4þH2; DH⁺_{298 K}¼ 74:7 kJ=mol (5) In the partial oxidation of methane (Eq. (4)), it is noticed that the produced syngas (COþH2) can be utilized as a feedstock to the Fisher–Tropsch reaction. Although the product gas contains some CO2, all the product gases (H2, CO, CO2) can be converted to FT products with a H2/(2COþ3CO₂) ratio equal to about 1.05[27]. In that case, the two-step SMR process can reduce CO2emission and produce high purity hydrogen and

Steam reformer Natural gas

feedstock HDS/ZnO Shift reactors

(470-520K) steam

Pre-refomer

HTS LTS (620-820K)

PSA

800K 1100K

H₂product

Heat exchanger

Heat exchanger Fig. 1 – Simplified flow diagram of conventional SMR process.

Nomenclature AR Air reactor CL Chemical looping

CLC Chemical looping combustion Ci Concentration of componenti, % FR Fuel reactor

MFR Solid mass weight in the fuel reactor, kg MSR Solid mass weight in the steam reactor, kg n__i Molar rate of componenti, mol/s

ns Theoretical mole of oxygen corresponds to the oxidation from FeO to Fe₃O₄, mol

PSA Pressure swing adsorption SR Steam reactor

SMR Steam methane reforming

WGS Water–gas shift reaction TGA Thermo-gravimetric analyzer T Temperature, K

t time, min

Ug Superficial gas velocity, m/s

Umf Minimum fluidizing gas velocity, m/s W Weight of the measured sample, mg wox Weight of the fully oxidized sample, mg wred Weight of the fully reduced sample, mg X Fractional conversion of metal oxides

xox Fractional oxidation conversion of metal oxides x_red Fractional reduction conversion of metal oxides Greek symbols

6H298 K+ Enthalpy of reaction at 298 K, kJ/mol

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the feedstock to the FT synthesis. Simplified flow diagram of two-step SMR system is shown inFig. 2.

For two-step SMR system, the effect of reaction temperature on the redox reactivity of iron oxides was determined in

thermo-gravimetric analyzer (TGA) and the effects of gas velocity and reactant concentration on the reactivity at the isothermal condition in a fluidized bed reactor were determined. Based on composition of the product gases at the

HDS preheater

FR

SR Natural gas

feedstock HDS/ZnO

Steam

H₂product Purification

Heat exchanger

Heat exchanger Syngas

product 470-520K

Iron oxides circulation 1070-1170 K

FT process Heat

exchanger CO₂

sequestration

Eq(3)

Eq(4)

Eq(5) Some H₂moves to FR to utilize in FT process

Fig. 2 – Simplified flow diagram of two-step SMR process.

Table 1 – Chemical looping researches for the improvement of SMR.

Research group

Class/

Research type

Process concept Process characteristics

Wolf and Yan[14]

A, B

Extended CLC: oxygen carrier circulates along with a CO2acceptor between three fluidized bed reactor (Fuel/CO2desorption/Air reactor)

Inherent H2production

Conceptual

Addition of CO2acceptor (CaO) in FR

Zafar et al.[15] A

Integrated hydrogen and power production with CO2capture

Inherent CO2separation

Experimental

Necessary for additional processes: WGS, PSA

Ryden and Lyngfelt[16]

A Conceptual

Energy for the endothermic reforming reactions is provided by indirect combustion that takes place in two separate reactors: one for air and one for fuel

The offgas from the PSA unit is used as fuel in the fuel reactor.

The tubes located inside the CL fuel reactor.

Improved selectivity towards H2and higher reformer efficiency with CO2capture Ryden

et al.[17]

A

Two interconnected fluidized beds for CL reforming

Complete conversion of natural gas

Experimental

High selectivity towards H2and CO

Implementation of the continuous reaction Duponta

et al.[18]

A, B

Unmixed steam reforming: steam reforming that uses separate air and fuel–steam feeds

90% Selectivity of H2in FR using dolomite Experimental

Johnsen et al.[20]

B

Ni-based catalyst and dolomite circulate between reformer and regenerator

98–99 Vol.% hydrogen after 4 cycles at 873 K

Experimental

^{Loss of CO}²up-take capacity of the dolomite

Chiesa et al.[21]

C

Three-reactors CL system, where iron oxide particles are circulated to: (i) oxidize natural gas (ii) reduce steam, to produce hydrogen as the final product of the process, (iii) consume oxygen from an air stream

Direct production of the H2and CO2from natural gas, by means of a process that simpler than the conventional technologies with CO2

capture capabilities Conceptual

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optimum condition in the redox reaction, the concept and the chemical reactions for the production of pure H2from SR and syngas from FR were analyzed with variation of iron oxide conversion and the guideline for designing the continuous process of two-step SMR system is presented.

2. Experimental

2.1. Reaction characteristics of iron oxides in thermo-gravimetric analyzer

The redox reactions of Fe/Fe₂O₃(Bayer Co.) were carried out in thermo-gravimetric analyzer (TGA, SETARAM 92). The sample weight of 7–8 mg (0.1–2.5 micron) and the gas velocity of 0.08 m/s were selected to eliminate the diffusion effect on the reaction. The maximum experimental temperature was limited up to 1173 K which is similar to the operating condition of conventional SMR reaction. Air was injected into the analyzer to prevent the pre-reduction before attaining the isothermal condition. When the desired temperature was reached, methane with 10 Vol.% nitrogen gas was supplied into TGA and then variation of sample weight was measured with time. For oxidation of Fe, a sample was prepared in advance by using hydrogen having the same conversion.

Inert gas was injected to prevent pre-oxidation before attaining the isothermal condition. Saturated steam was then supplied into the analyzer through a steam bubbler at 333 K.

2.2. Experimental apparatus

A fluidized bed reactor (83 mm-ID1350 mm-high) made of inconel 600 pipe is schematically shown inFig. 3. The reactor was heated with an electric kantal wire and insulated with ceramic fiber wool to prevent heat loss to surroundings. The

fluidizing gas was injected into the bed via a plenum chamber to a sintered plate distributor having 40% opening area. Before entering the gas into the bed, a pre-heater was used in heating the reducing gas and generating steam from water supplied by a water pump. To measure the bed pressure, two pressure taps were mounted flush with the wall of the reactor at 35 mm and 145 mm above the distributor plate and pressure trans- mitters (Sensys) were used.

2.3. Experimental procedure

To obtain stable fluidization, zirconia–silica beads (Saint- Gobain Co. Ltd.) were added as an inert solid medium in the fluidized bed reactor. Initially, 10 wt.% hematite (total bed weight: 3.415 kg) was premixed with zirconia–silica beads and then it was loaded into the bed to make a static bed height having an aspect ratio (length/diameter) of 4. To maintain the initial oxidation state of iron oxides to Fe3O4, first redox cycle was carried out. To prevent the carbon deposition from the catalytic cracking of methane[28], the reaction in this study is limited from FeO to Fe3O4. For the reduction of iron oxides, methane with dilution by nitrogen gas was supplied through a mass flow controller (MKS Co. Ltd.). After finishing the reduction step, the residue gases were purged by nitrogen gas and then the oxidation reaction was conducted. In the oxidation step, the injected water was evaporated in a preheater at 573 K and then steam was supplied by nitrogen gas into the reactor. Physical properties of the particles used in this study are presented inTable 2.

2.4. Analysis of product gases and solids

The fractional conversion by the weight change in TGA is defined as:

Fractional reduction conversion from Fe2O3to Fe is

Fig. 3 – Schematic diagram of the fluidized bed reactor system.

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xred¼ ðwoxwÞ=ðwoxwredÞ (6) Fractional oxidation conversion from Fe to Fe2O3is

xox¼ ðwwredÞ=ðwoxwredÞ (7)

wherexis the fractional conversion,wis the weight of sample at any time andwox,wredare the weights of the sample after fully oxidized and reduced, respectively.

Concentrations of product gases from the redox reaction in the fluidized bed was measured by infrared gas analyzer (Fuji Co., Ltd.) to detect CH4, CO, CO2, by thermal conductivity analyzer (Fuji Co., Ltd.) to detect H2and H2O was calculated from the mass balance. Composition of product gases was also confirmed by gas chromatograph (HP 5890 series II equipped with a Molsieve 5A and PoraPak Q column).

The fractional conversion of the product gases is defined as:

xred¼ Rt1

t0

n_COþ2n_CO2þn_H2O

dt ns

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xox¼ Rt1

t0

n_H2

dt ns

(9) where the fractional conversions of reduction and oxidation are determined by integrating the molar evolution rate of each components (n_i) with time fromt0tot1andnsis the theoretical oxygen mole corresponds to oxidation from FeO to Fe3O4.

The phase of iron oxides before and after reaction was analyzed by X-ray diffractometer (XRD, D/MAX-IIIC (3 kW), RIGAKU).

3. Results and discussion

3.1. Effect of the temperature on the redox reactivity of iron oxides in TGA

The effect of temperature on the reactivity of reduction and oxidation of iron oxides in two-step SMR process was determined in TGA as shown in Fig. 4. As can be seen, the conversions of both reduction and oxidation increase with increasing temperature. Also, it is found that the effect of temperature on the reduction reactivity of iron oxides is more pronounced than that of the oxidation due to the large endothermic reduction reaction ofEqs. (3) and (4). The reaction rate is shown to be slow in the range of 0.2–0.4 and it is

more sensitive to temperature in the reduction of Fe3O4. In that region, the reduction rate decreases with increasing FeO (xred¼0.33) phase since FeO reduction by H2and CO is unfa- vorable in thermodynamic aspects[28,29]. As can be seen in Fig. 4(a), the initial stage of reaction, Fe2O3(atxred¼0)/Fe3O4

(atxred¼0.11), is very fast regardless of the reaction temperature. However, the reduction temperature should be limited above 1073 K since oxidation of iron oxides by water vapor can reach up to Fe3O4[30,31]. After almost complete consumption of oxygen at 1148 K and 1173 K, carbon deposition is presented by the catalytic methane decomposition with increasing the sample weight as report by Cho et al.[32].

For oxidation of Fe inFig. 4(b), the conversion increases with increasing temperature and the reaction occurs Table 2 – Physical properties of iron oxides and inert

thermal medium.

Properties Symbol Dimension Fe2O3

(hematite)

Zirconia–

silica bead Particle size

range/mean size

dp mm 0.1–2.5/0.63 63–125/90

Density of particle

r kg/m³ 5180 3850

Minimum fluidizing gas velocity

Umf m/s N/A 0.0067

Fig. 4 – Fractional (a) reduction conversion of Fe2O3and (b) oxidation conversion of Fe at a given temperature in TGA.

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comparatively at lower temperatures compared to the reduction. However, the oxidation rate is lower than the reduction rate. At 1173 K, the reaction rate in the initial stage is fast and levels off at the later stage. Based on the data from TGA, the optimum working temperature in this study is determined to be 1173 K based on the reaction kinetics and similar operating temperature in conventional process.

3.2. Reduction characteristics of iron oxides in the fluidized bed reactor (FR)

Chemical looping system can be achieved in a two-step SMR system in the fluidized bed reactor. The effect of gas velocity on the reduction of Fe3O4in the bed is shown inFig. 5. The reactivity exhibits its maximum value at gas velocity of 0.0287 m/s (4.28Umf) where gas–solid mixing would be highest due to the gross solids circulation caused by drift and wake transport of bubbles and small scale local-mixing in bubble wakes [33,34]. At the optimum gas velocity, the effect of methane concentration on reactivity of the reduction is shown inFig. 6. As can be seen, the maximum conversion of around 0.32 can be attained at methane concentration above 25%. The reactivity in TGA is higher than that in the fluidized bed reactor because the gas–solid contacting or mixing is limited due to the reactant gas encapsulated in bubbles in the reactor.

3.3. Oxidation characteristics of iron oxides in the fluidized bed reactor (SR)

The performance of oxygen carrier in terms of the maximum H2evolution rate in SR with 50% steam injection as a function of gas velocity is shown in Fig. 7. As can be seen, the maximum evolution rate of H2can be obtained at gas velocity of 0.0402 m/s (6U_mf) that is a little higher velocity than that in

FR. This difference of the optimum gas velocity can be attributed to gas volume change from the reduction of Fe3O4

by methane accompanied by volume expansion in the reac- tionEq. (3)but oxidation of FeO has no volume change in the reactionEq. (4).

The effect of steam concentration on the maximum evolution rate of H2is shown inFig. 8where the evolution rate of H2is improved with increasing steam concentration while the increasing rate becomes lower over 50% steam.

Fig. 5 – Effect of gas velocity on the reduction conversion of iron oxides (C_CH₄[25%;T[1173 K;t[20 min).

Fig. 6 – Effect of methane concentration on the reduction conversion of iron oxides (Ug[0.0287 m/s;T[1173 K;

t[20 min).

Fig. 7 – Hydrogen evolution rate as a function of gas velocity in oxidation of iron oxides (C_H₂_O[50%;

T[1173 K).

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3.4. Feasibility study of two-step SMR system

At the optimum condition (Ug: 0.0287 m/s, CCH4: 25%) composition of the product gases from FR with time and the conversion was determined. The reactivity is changed at 40 min (x_Fe₃_O₄_/FeOy0:55) after reaction started as can be seen inFig. 9(a). In the former part, it is shown that concentrations of H2O and CO2of the product gases are relatively high due to oxygen in Fe3O4were consumed. On the other hand, it can be seen that concentration of H2 and CO increases after the conversion of 0.55. This result can be elucidated by the reaction of catalytic decomposition of methane with phase change of iron oxides and this mechanism is shown inFig. 10. As can be seen, methane is slowly decomposed at the initial reaction (in the absence of catalyst [32]) of Fe3O4 with methane to produce solid carbons and H2which are fully oxidized to H2O and CO2by reduction of Fe3O4. Also CO gas can be generated through the further reaction of CO₂with solid carbons (Bou- douard reverse reaction[32]). However, the rate of catalytic methane decomposition increases with increasing FeO phase in the middle of the reaction so that production rates of solid carbon and H2rapidly increase[32,35]. After all, direction of the reaction is favorable to the partial oxidation of methane under the fixed amount of oxygen carriers in the batch reaction system. In particular, the reason why the reduction rate of iron oxides is not improved despite the fast increase of methane consumption in the latter stage is also explained by the partial oxidation of methane. From the data in the figure, the average molar ratio of H2/(2COþ3CO2) in the latter part is found to be about 0.65. The phases of Fe3O4(magnetite) and FeO (wuestite) can be confirmed from the XRD patterns of sampling solids after the reduction (Fig. 11b).

The carbon deposition can be confirmed by the mass balance of carbon at a fixed methane supply (25%) in the batch

reaction. From the gas composition obtained with time in Fig. 9(a), carbon deposition is mainly caused by the partial oxidation of methane (i.e., after 40 min). However, any carbon component was not observed on the surface of the sampling particles after the reduction (Fig. 11b) because the deposited carbon is small and almost emitted in the form of CO2. After finishing the reduction reaction, residual gases were purged by nitrogen gas for 20 min, where small amount of CO2is still detected in the reactor. Subsequently, composition of the product gases with time and the oxidation conversion of FeO Fig. 8 – Hydrogen evolution rate in oxidation of iron oxides

as a function of steam concentration (Ug[0.0402 m/s;

T[1173 K).

Fig. 9 – Gas concentration from the fluidized bed reactor (a) reducing with methane (CCH4[25%;Ug[0.0287 m/s), (b) oxidizing with steam (C_H₂_O[50%;Ug[0.0402 m/s) of iron oxides at 1173 K,6: CO,

^{: CO}2,-: CH4,,: H2,;: H2O.

Fig. 10 – Mechanism of the methane oxidation with iron oxides.

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were determined with 50% steam introduced in the reactor as shown inFig. 9(b). With steam injection to the reactor, FeO is oxidized fast to produce H2under the excess oxygen present.

About 10 min after starting reaction (xox: 0.45–0.5), the abrupt change of the reaction rate occur and this point is similar with the result in FR based on the conversion from FeO to Fe3O4. At the oxidation conversion of FeO over 0.9, the oxidation rate sharply decreases. Meanwhile, it has been expected that solid carbon deposited on particles in FR can appear to be in the form of CO, CO2, and CH4in SR. However, it is shown that CO and CH4are not present in the product gases except small amount of CO2 because the equilibrium composition is favorable to the direction of CO2production. Based on XRD patterns of the sampling solids after the oxidation (Fig. 11c), it can be claimed that most of iron oxides are re-oxidized to the solid phase of Fe3O4(magnetite) by steam.

In the present study, H2without CO can be generated from the two-step SMR reaction with the redox of FeO/Fe3O4. The CO-free gas in SR [36,37] can directly feed to the polymer electrolyte membrane (PEM) fuel cell. The results from FR are divided into following two cases: the full oxidation of methane is favorable below 50% conversion of Fe3O4/FeO at oxygen carrier rich condition while the syngas production for the feedstock to the Fischer–Tropsch reaction and the reduction of Fe3O4prevail over 50% conversion of Fe3O4/FeO at oxygen carrier lean condition. In particular, the production of both purified H2and syngas mixture for Fischer–Tropsch reaction can be obtained from H2generated from SR that supplied to FR to satisfy the proper ratio of product gas composition as shown inFig. 2.

3.5. Consideration of designing a continuous two-step SMR system with redox of FeO/Fe₃O₄

Based on the present study, the design base for the continuous process of two-step SMR with the redox of FeO/Fe3O4could be obtained. The reaction conversion range could be 0–0.5 from FeO to Fe3O4since H2evolution rate in SR is fast at the lower oxidation state between FeO and Fe3O4(Fig. 9b) and the syngas production and reactivity of solids in FR is also favorable at the lower oxidation state of FeO (Fig. 9a). To compensate the

difference of the redox reaction rates, the solid inventory ratio of two-reaction regions should be determined from the present experimental data. This is defined as the ratio of solid mass weight (MFR/MSR) that corresponds to an average resi- dence time of solids in each reactor. Based on the data inFig. 9, the solid inventory ratio can be determined to be about 2.9. As maintaining the mass balances of the reaction, the amount of H2production depends on the contact time between gas and solids since the increase of contact time improves the conversion of methane and steam. In this respect, bubbling fluidized bed reactors would be a best choice since it has long reaction time with good solid mixing. Therefore, two bubbling fluidized beds system which has the double pressure loops as a driving force for solid circulation with two gas streams (fuel and steam) [38]is proposed for two-step SMR reaction. In addition, a pneumatic transfer line for upward solids transport and non-mechanical valves for the control of solid flow are also needed for the continuous reaction system.

4. Conclusions

The two-step steam methane reforming system with the redox of FeO/Fe3O4 for H2 production using the chemical looping process is confirmed in the present study and the guideline for the construction of continuous process is proposed. In TGA, the range of reaction temperature is found to be over 1073 K for FR and the reaction rate increases with increasing temperature in both FR and SR. The optimum operating conditions (Ug¼0.0287 m/s and C_CH₄>25% in FR and Ug¼0.0402 m/s and C_H₂_O>50% in SR) were determined for the mass transfer between gas and solid in the fluidized bed. It is found that pure hydrogen except for CO2 can be obtained from the reaction of FeO/Fe3O4in SR at 1173 K.

With respect to the syngas production for the Fischer–Tropsch reaction and the oxidation reactivity of FeO, the conversion range of the redox reaction is found to be 0–0.5 from FeO to Fe3O4. Based on the present experimental data, the design basis for the continuous two-step SMR is proposed as the double-loop solids circulation system with fuel and steam streams. The solid inventory ratio of FR to SR to compensating the reaction rate is determined to be 2.9.

Acknowledgements

This Research (Paper) was performed for the Hydrogen Energy R&D Center, one of the 21st Century Frontier R&D Program, funded by the Ministry of Education, Science and Technology of Korea.

r e f e r e n c e s

[1] Schultz MG, Diehl T, Brasseur GP, Zittel W. Air pollution and climate forcing impacts of a global hydrogen economy.

Science 2003;302:624–7.

[2] Bak YC. Status for the technology of hydrogen production from coals. Energy R&D 1993;15(2):191–201.

Fig. 11 – XRD patterns of solid particles: (a) raw material, (b) sample after reduction (c) sample after oxidation.

(9)

[3] Dixon RK. Advancing towards a hydrogen energy economy:

status, opportunities and barriers. Mitigation Adapt Strateg Glob Change 2007;12:325–41.

[4] National hydrogen energy roadmap. Washington, DC: U.S.

Department of Energy; 2002.

[5] Sperling D, Cannon JS. The hydrogen energy transition:

moving toward the post-petroleum age in transportation.

London: Elsevier Academic Press; 2004.

[6] Kothari R, Buddhi D, Sawhney RL. Comparison of environmental and economic aspects of various hydrogen production methods. Renew Sust Energ Rev 2008;12:

553–63.

[7] Fu Q, Saltsburg H, Flytzani-Stephanopoulos M. Active nonmetallic Au and Pt species on ceria-based water–gas shift catalysts. Science 2003;301:935–8.

[8] Hardacre C, Ormerod RM, Lambert RM. Platinum-promoted catalysis by ceria: a study of carbon monoxide oxidation over Pt(11 l)/CeO2. J Phys Chem 1994;98:10901–5.

[9] Ghenciu AF. Review of fuel processing catalysts for hydrogen production in PEM fuel cell systems. Curr Opin Solid State Mater Sci 2002;6:389–99.

[10] Yeung CMY, Yu KMK, Fu QJ, Thompsett D, Petch MI, Tsang SC. Engineering Pt in ceria for a maximum metal–

support interaction in catalysis. J Am Chem Soc 2005;127:

18010–1.

[11] A National Vision of America’s Transition to A Hydrogen Economy - To 2030 and Beyond. Available from:www.eren.

doe.gov.

[12] Peng Z. A novel hydrogen and oxygen generation system.

MS. Thesis. Louisiana State University; 2003.

[13] Myers DB, Ariff GD, James BD, Lettow JS, Thomas CE, Kuhn RC. Cost and performance comparison of stationary hydrogen fueling appliances 2002. Task 2 report, Grant No.

DE-FG01–99EE35099.

[14] Wolf J, Yan J. Parametric study of chemical looping combustion for tri-generation of hydrogen, heat, and electrical power with CO2capture. Int J Energy Res 2005;29:

739–53.

[15] Zafar Q, Mattisson T, Gevert B. Integrated hydrogen and power production with CO2capture using chemical-looping reforming redox reactivity of particles of CuO, Mn2O3, NiO, and Fe2O3using SiO2as a support. Ind Eng Chem Res 2005;44:

3485–96.

[16] Ryden M, Lyngfelt A. Using steam reforming to produce hydrogen with carbon dioxide capture by chemical-looping combustion. Int J Hydrogen Energy 2006;31:1271–83.

[17] Ryden M, Lyngfelt A, Mattisson T. Synthesis gas generation by chemical-looping reforming in a continuously operating laboratory reactor. Fuel 2006;85:1631–41.

[18] Duponta V, Rossa AB, Knighta E, Hanleya I, Twiggb MV.

Production of hydrogen by unmixed steam reforming of methane. Chem Eng Sci 2008;63:2966–79.

[19] Barelli L, Bidini G, Gallorini F, Servili S. Hydrogen production through sorption-enhanced steam methane reforming and membrane technology: a review. Energy 2008;33:554–70.

[20] Johnsen K, Ryu HJ, Grace JR, Lim CJ. Sorption-enhanced steam reforming of methane in a fluidized bed reactor with dolomite as CO2-acceptor. Chem Eng Sci 2006;61:1195–202.

[21] Chiesa P, Lozza G, Malandrino A, Romano M, Piccolo V.

Three-reactors chemical looping process for hydrogen production. Int J Hydrogen Energy 2008;33:2233–45.

[22] Zeman F, Castaldi M. An investigation of synthetic fuel production via chemical looping. Environ Sci Technol 2008;

42:2723–7.

[23] Jin GT, Ryu HJ, Jo SH, Lee SY, Son SR, Kim SD. Hydrogen production in fluidized bed by chemical-looping cycle.

Korean J Chem Eng 2007;24(3):542–6.

[24] Richter HJ, Knoche KF. Reversibility of combustion processes.

ACS symposium series. Washington DC: American Chemical Society; 1983. pp. 71–85.

[25] Lyngfelt A, Leckner B, Mattisson T. A fluidized-bed combustion process with inherent CO2separation;

application of chemical-looping combustion. Chem Eng Sci 2001;56:3101–13.

[26] Abad A, Adańez J, Garcıá-Labiano F, de Diego LF, Gayań P, Celaya J. Mapping of the range of operational conditions for Cu-, Fe-, and Ni-based oxygen carriers in chemical-looping combustion. Chem Eng Sci 2007;62:533–49.

[27] Dry ME. The Fischer–Tropsch process: 1950–2000. Catal Today 2002;71:227–41.

[28] Steinfeld A, Frei A, Kuhn P. Thermoanalysis of the combined Fe3O4-reduction and CH4-reforming processes. Metall Mater Trans B 1995;26(B):509–15.

[29] Steinfeld A, Kuhn P. High-temperature solar

thermochemistry: production of iron and synthesis gas by Fe3O4-reduction with methane. Energy 1993;18(3):239–49.

[30] Charvin P, Abanades S, Flamant G, Lemort F. Two-step water splitting thermochemical cycle based on iron oxide redox pair for solar hydrogen production. Energy 2007;32:1124–33.

[31] Lorente E, Pena JA, Herguido J. Kinetic study of the redox process for separating and storing hydrogen: oxidation stage and ageing of solid. Int J Hydrogen Energy 2008;33:615–26.

[32] Cho P, Mattisson T, Lyngfelt A. Carbon formation on nickel and iron oxide-containing oxygen carrier for chemical- looping combustion. Ind Eng Chem Res 2005;44:668–76.

[33] Lim KS, Zhu JX, Grace JR. Hydrodynamics of gas–solid fluidization. Int J Multipahse Flow 1995;21:141–93.

[34] Abrahami SN, Resnick W. Fluidised bed behavior near incipient fluidization in a three-dimensional bed. Trans Inst Chem Engrs 1974;80:80–7.

[35] Bak YC, Cho KJ. Status for the technology of hydrogen production from natural gas. Korean Chem Eng Res 2005;

43(3):344–51.

[36] Gottesfeld S. Preventing CO poisoning in fuel cells. U.S.

Patent 4,910,099, 1990.

[37] Gottesfeld S, Pafford J. A new approach to the problem of carbon monoxide poisoning in fuel cells operating at low temperatures. J Electrochem Soc 1988;135:2651–2.

[38] Kunii D, Levenspiel O. Fluidization engineering. 2nd ed.

London: Butterworth-Heinemann; 1991.