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Syngas production from oxidative methane reforming and CO cleaning with water gas shift reaction
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 3 8 ( 2 0 1 3 ) 1 0 8 2 6 e1 0 8 3 2
Available online at www.sciencedirect.com
journal homepage: www.elsevier.com/locate/he
Syngas production from oxidative methane reforming and CO cleaning with water gas shift reaction Donglai Xie*, Jie Zhao, Ziliang Wang, Yajun Zhang The Key Laboratory of Fuel Cell Technology of Guangdong Province, South China University of Technology, Guangzhou 510640, China
article info
abstract
Article history:
Fuel cell based heat and power cogeneration is considered to be well qualified for a dis-
Received 5 November 2012
tributed energy system for residential and small business applications. A fuel processing
Received in revised form
unit including an oxidative steam methane reformer, a high temperature shift reactor and
27 December 2012
a low temperature shift reactor is under development in South China University of Tech-
Accepted 1 January 2013
nology. Performance of the unit is experimentally investigated in a bench-scale exper-
Available online 29 January 2013
imental setup. Processor performance under typical operating conditions is tested. The influence of reaction temperature, methane space velocity in the oxidative steam methane
Keywords:
reformer, and air to carbon molar ratio on unit performances is experimentally studied. It
Hydrogen
is found that under the typical operating conditions, the total energy efficiency reaches
Reforming
88.3%. The efficiency can further be improved by utilizing the sensible heat of the refor-
Water gas shift reaction
mate gas. The current study has been focused on the chemical performances such as
Combined heat and power
methane conversion of the reformer and CO concentration in the synthesis gas down-
cogeneration
stream water gas shift reactors. Heat integration of the unit will be further implemented in future to improve energy efficiency. Copyright ª 2013, Hydrogen Energy Publications, LLC. Published by Elsevier Ltd. All rights reserved.
1.
Introduction and background
Fuel cell technologies have been of special interest in recent years due to their high efficiency in energy conversion and less/no harmful emissions compared with other energy conversion systems [1]. Fuel cell based heat and power cogeneration (CHP) is considered to be well qualified for a distributed energy system for residential and small business applications, which generally incorporates the proton exchange membrane fuel cells (PEMFCs) with hydrogen production from natural gas or other fuels [2e5]. These applications, also known as microCHP systems, offer benefits of on site and real time generation of electricity with a high efficiency and utilization of the heat
produced during operation of fuel cells and fuel processing for hot water generation and residential heating [6]. When introducing fuel cell based micro-CHP systems into the market, small-scale hydrogen generation systems from various hydrocarbon feed stocks are required due to the lack of hydrogen distribution infrastructure. A kW-scale hydrogen production unit is under development in the New Energy Group in South China University of Technology. Natural gas is selected as feedstock for hydrogen production, as its infrastructure is well established in most countries. The fuel processing unit constitutes of an oxidative steam methane reformer (OSR), a high temperature shift reactor (HTS) and a low temperature shift reactor (LTS). The
* Corresponding author. Tel./fax: þ86 20 22236985. E-mail addresses: [email protected], [email protected] (D. Xie). 0360-3199/$ e see front matter Copyright ª 2013, Hydrogen Energy Publications, LLC. Published by Elsevier Ltd. All rights reserved. http://dx.doi.org/10.1016/j.ijhydene.2013.01.012
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 3 8 ( 2 0 1 3 ) 1 0 8 2 6 e1 0 8 3 2
targeted CO concentration in the reformate gas out of the unit is 0.3%e0.6%. The OSR converts natural gas, steam and air into a hydrogen rich gas. The outlet gas is then fed to the CO shift reactors where CO concentration is expected to be reduced to 2.0%e4.0% in the HTS and to 0.3%e0.6% in the LTS. The unit can directly supply feed gas to a high temperature PEMFC [7]. If a regular low-temperature PEMFC is employed, the CO concentration in the product gas from the unit needs to be further reduced to be less than 10 ppm through selective CO oxidation or CO methanation. The current study focuses on the design and chemical performance of theses reactors under various operating conditions, while the heat integration will be implemented later in the system assemble stage. Fuel cell based micro-CHP system is a novel energysupplying technology represented the main developing trends for the future’s residential energy supply. However, the technology is far beyond maturity. Very few testing results have been published. The experimental results presented could help to understand the mechanisms and performances of the fuel processing unit in the CHP system.
2. The reactors design and experimental setup 2.1.
The reactors design
The sizing of the OSR, HTS and LTS are based on the simulation result of a 1 kW scale fuel cell based micro-CHP system with Aspen Plus [8]. To produce H2 rich reformate gas from natural gas, water and air, a reformer integrated with steam generation is designed and patented [9]. The integrated reforming system constitutes of a reforming reactor, a combustion chamber, a burner, an external methane heating coil, an external water coil, an internal mixture gas coil, a flue gas outlet nozzle, and a reformate gas outlet nozzle. Methane and water enter the external methane and external water coils outside the combustion chamber, respectively. They then join together and enter the internal mixture gas coil around the reforming reactor. The reactor is essentially a stainless steel tube with an inner diameter of 80 mm and a length of 200 mm, where reforming catalyst is filled. An air distributor is inserted inside the reactor for oxidative steam methane reforming. The air distributor is made of a 6 mm outer diameter stainless steel tube with four pairs of 2 mm diameter orifices (8 holes in total) drilled through the tube wall. 800 g of commercial nickel catalyst Z413Q supported on gamma alumina are loaded in the reactor. The catalyst is supplied by Shandong Qilu Keli Chemical Institute Co. (China) and has been widely used in industrial natural gas reforming systems in China. The details of the reformer and catalyst can be found elsewhere [10]. The HTS, as well as the LTS, is made of a 200 mm long stainless steel tube with inner diameter of 60 mm. Feed gas enters the vessel from a nozzle on its top flange and the reactant gas leaves the reactor from a nozzle on its bottom flange. Commercial catalyst KLB-101/9 with active component of Fe2O3eCr2O3 is loaded in the HTS, while commercial catalyst KLB-201 with active component of CuOeZnO is loaded in the LTS. Both catalysts are supplied by Shandong Qilu Keli Chemical Institute Co. (China). Both the HTS and LTS are
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equipped with electrical heaters for bed heat up during startup. Insulation ribbons are wrapped outside of these reactors. The number of the layers of the insulation ribbon can be manually changed during the experiments in order to maintain the bed temperature at a desired value, as the water gas shift reaction is exothermal.
2.2.
Experimental setup
Performance of the hydrogen production unit is tested in a bench-scale experimental setup. Schematic layout of the setup is shown in Fig. 1. Due to the lack of natural gas distribution infrastructure in the laboratory, bottled methane (industrial pure, supplied by Guangzhou Zhuozheng Industrial Gas Ltd., China) rather than piped natural gas is employed as process feed. Bottled liquefied petroleum gas (LPG, supplied by Guangzhou Gas Ltd., China) is fed into the burner for combustion. The volumetric composition of the LPG is 1.0% C2H6, 9.0% C3H6, 4.5% C3H8, 54.0% C4H8, 26.2% C4H10, and 5.3% Cþ 5. The flow rate of process CH4 is controlled by a mass flow controller (MFC). The flow rates of process air and water are monitored by rotameters. The flow rate of LPG for combustion is measured by a wet gas flow meter. The process methane and water under room temperature are preheated by hot flue gas issued from the burner to ensure that the water is in vapor phase before being injected into the reactor. The methane and water are heated separately in the external methane and water coils outside the combustion chamber. They are further heated to a temperature around 500 C in the internal mixture gas coil. The preheated methane and steam mixture then proceed to the catalyst reactor to produce a hydrogen rich reformate gas containing largely H2, CO2, CO, and steam. Air is introduced through the eight orifices of the air distributor. The reactor off gas from the OSR is directed to HTS and LTS in sequence. The reformate gas compositions downstream of each reactors are analyzed by a gas chromatography (GC910, Shanghai Kechuang Chromatograph Instrument Ltd., China) equipped with a thermal conductivity detector. The connection tubes between reactors are well insulated.
2.3.
Unit performance characterization
The OSR performance is characterized by reformate gas composition (mainly H2 and CO concentration) and methane conversion. Methane conversion is defined as: XCH4 ¼
cCO þ cCO2 in the product gas cCH4 þ cCO þ cCO2 in the product gas
(1)
The HTS and LTS performances are characterized by CO concentration in the outlet gas and CO conversion. The CO conversion is defined as: XCO ¼ 1
cCO in the product gas cCO in the feed gas
(2)
The influence of following operating conditions on reactor performance is studied: (1) reaction temperature; (2) air to carbon molar ratio in the feed (AC); and (3) methane space velocity in OSR. OSR methane space velocity is defined as the
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Fig. 1 e Schematic layout of the experimental setup.
ratio between the volumetric flow rates of CH4 to OSR under standard conditions (273.15 K and 0.1 MPa) and the catalyst volume in OSR.
3.
Experimental results and discussion
The hydrogen production unit has experienced intensive tests after being built. More than 600 h’ operation has been carried out aperiodically in a six-month period with different operating conditions. The unit has shown stable functionality and no obvious deactivation of the catalysts is observed.
3.1.
Performance of a typical case
Table 1 shows the performance of the OSR, HTS and LTS under typical operating conditions of Fm ¼ 14.0 mol h1, FL ¼ 4.1 mol h1, Fa ¼ 33.6 mol h1 (AC ¼ 2.4), and Fw ¼ 56.0 mol h1 (SC ¼ 4.0). Under such conditions, the hydrogen concentrations downstream of the OSR, HTS and LTS are 40.61%, 43.45% and 45.56% (dry basis), respectively. The hydrogen concentration of 45.56% downstream of LTS is higher than 30.5% reported by Heinzel et al. [11], 40% by Lee et al. [12], and lower than 51.3% reported by Lin et al. [13] and 48% by Adachi et al. [14]. The methane conversion in the OSR is
92.56%, while the CO conversion is 59.62% in the HTS and 97.47% in the LTS. The CO concentrations downstream of these reactors are 5.56%, 2.25% and 0.06% (dry basis), respectively. It is obvious that the CO concentration downstream of the LTS meet the design target of 0.3%e0.6%. The CO concentration of 0.06% downstream of the LTS is much lower than 8.9% reported by Heinzel et al. [11]. The oxidative steam methane
Table 1 e Performance of the OSR, HTS and LTS under typical operating conditions. (Fm [ 14.0 mol hL1, FL [ 4.1 mol hL1, Fa [ 33.6 mol hL1, and Fw [ 56.0 mol hL1.) Reactor Average bed temperature, C Downstream gas composition, % H2 N2 CO CH4 CO2 Key component conversion, %
OSR
HTS
LTS
630
420
300
40.61 37.63 5.56 3.62 12.58 92.56 (CH4)
43.45 37.12 2.25 3.27 13.91 59.62 (CO)
45.56 36.11 0.06 3.18 15.10 97.47 (CO)
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reforming reactions are limited by thermodynamic constraints. Experimental results from the same reactor under steam methane reforming and oxidative steam methane reforming operating conditions have been compared with the predicted values by Gibbs free energy minimization models in a previous study [10]. As stated previously, in the experimental test rig, the flow rates of process feeds (methane, water and air) and combustion LPG are all measured. Flue gas composition is analyzed by a flue gas analyzer. The combustion air flow rate can be calculated from the oxygen concentration in the flue gas and LPG flow rate [15]. Syngas composition downstream of each reactor is measured by a gas chromatography. Gas temperatures of all streams are intensively monitored by thermocouples as shown in Fig. 1. Hence, enthalpy of these streams can be calculated from the stream flow rate, composition and temperature. Fig. 2 shows the flow of enthalpy in each unit operation of the setup from these calculations. The enthalpy loss in each reactor and the connection tube is shown as DE in the figure. The total energy efficiency, which is defined as the total enthalpy carried by the reformate gas and flue gas divided by the enthalpy carried by feed methane and combustion LPG, is 88.3% under the typical operating conditions. Here the energy carried by flue gas is considered as useful energy, as its energy can be further employed in the downstream process to produce hot water. A major heat loss takes place in the connecting tubes. When the temperature of the reformate gas decreases from 630 C in the OSR to 420 C in the HTS and further to 300 C in the LTS, the heat released by the reformate gas temperature drop has not been recovered in the current configuration. When a micro-CHP system is built, the sensible heat of the reformate gas can be employed to produce hot water, and the system energy efficiency can be improved. Another significant energy loss is caused by the burner, which has a total heat loss of 1849.7 kJ h1, corresponding to 7.3% of the total energy input. The current burner is a domestic boiling table. In the experiments it is observed that some of the flue gas has escaped from the OSR heat transfer surfaces to the environment. A special designed atmospheric burner is under construction to improve flue gas flow pattern.
3.2.
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Effect of HTS and LTS bed temperatures
The effect of bed temperature on reformer performance has been previously studied extensively [8,10,16,17]. Hence it is not studied here. The water gas shift reaction is significantly affected by bed temperature. As the water gas shift reaction is not controlled by thermodynamic equilibrium, higher bed temperature helps the reaction kinetics. Fig. 3 shows that the CO concentration decreases from 2.8% to 2.0% in the HTS when the bed temperature increases from 220 C to 460 C and the feed gas molar composition is maintained at cH2 ¼ 44:27%, cN2 ¼ 35:48%, cCO ¼ 5.78%, cCH4 ¼ 3:25%, cCO2 ¼ 11:22%. For the LTS, when the bed temperature increases from 120 C to 290 C and the feed gas molar composition maintains at cH2 ¼ 443:70%, cN2 ¼ 36:87%, cCO ¼ 2.07%, cCH4 ¼ 3:26%, cCO2 ¼ 14:00%, the CO concentration decreases from 2.0% to 0.09%. Obviously, higher temperature improves the CO conversion in both HTS and LTS reactors for the temperature range investigated. Due to the limitation of the electrical heater capacity and the insulation conditions, higher bed temperatures for the HTS and LTS have not been tested.
3.3.
Effect of OSR methane space velocity
Gas space velocity is a crucial parameter influencing catalyst performance. It shows the amount of feed that the reactor can handle per unit volume of catalyst. When the space velocity is low, the contact time between the reactant and catalyst is long. Hence it is easy for the reaction to reach equilibrium. In the current study, the OSR, HTS and LTS performances are all experimentally investigated against the OSR methane space velocity. In the study, the process methane feed changes from 3.5 mol h1 to 21 mol h1, while the steam to carbon molar ratio maintains at 3.5 and the air to carbon molar ration maintains at 2.4, corresponding to an OSR methane space velocity from 82.5 h1e495.2 h1. For such a small scale chemical plant, heat dissipation to the environment is a serious issue, as the heat carried by gas stream is limited due to its very low flow rate. In the experiments, gas space velocity is changed by varying the feed flow rate while maintaining the catalyst volume unchanged. Hence,
Fig. 2 e Energy flow within the fuel processing unit under typical operating conditions (Fm [ 14.0 mol hL1, FL [ 4.1 mol hL1, AC [ 2.4, SC [ 4.0).
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Fig. 3 e Influence of bed temperature on CO concentration in HTS (Fm [ 14 mol hL1, SC [ 4.0, AC [ 2.4, inlet gas composition: cH2 [44:27%, cN2 [35:48%, cCO [ 5.78%, cCH4 [3:25%, cCO2 [11:22%) and LTS. (Fm [ 14 mol hL1, SC [ 4.0, AC [ 2.4, inlet gas composition: cH2 [443:70%, cN2 [36:87%, cCO [ 2.07%, cCH4 [3:26%, cCO2 [14:00%.)
low space velocity corresponds to low feed gas and reformate gas flow rates. When the reformate gas flow rate is low, it is difficult to maintain a bed temperature much higher than the environment. The top part of Fig. 4 shows the influence of methane space velocity on bed temperatures of OSR, HTS and LTS. It is obvious that bed temperatures increase with increasing OSR methane space velocity. Hence, the influence of space velocity on reactors’ chemical performance is a combination of both bed temperature and gas catalyst contact time. As shown in Fig. 4, the methane conversion of the OSR is not significantly affected by the space velocity, as well as the CO concentration in the reformate gas. It indicates that the OSR catalyst works well in the current space velocity range. For the HTS, it can be seen that when the OSR methane space velocity increases from 82.5 h1e495.2 h1, CO conversion decreases from 70% to 50%, while CO concentration in the reformate gas increases slightly. It indicates that the OSR space velocity has notable influences on HTS performance. For the LTS, when the OSR space velocity increases from 82.5 h1e495.2 h1, CO conversion decreases slightly from 99% to 98%, while CO concentration in the reformate gas increases slightly from 0.02% to 0.07%, which is far below the development requirement of 0.3%e0.6%. Hence the fuel processor works well under the OSR methane space velocity range of 82.5 h1 to 495.2 h1
3.4.
Effect of air to carbon molar ratio
Air to carbon ratio is an important factor for oxidative steam reforming. It has dominant impact on reaction temperature and product gas composition. As the ratio increases, the rate of the oxidation reaction increases and consequently heats up the reactor. This leads to a higher methane conversion. However, excess oxygen will also consume hydrogen and CO in the product gas, leading to a decline in H2 and CO yields in the OSR. Fig. 5 shows the influence of air to carbon molar ratio on bed
Fig. 4 e Effect of OSR methane space velocity on bed temperatures, methane conversion in OSR, CO conversions in HTS and LTS, and CO concentrations downstream of OSR, HTS, and LTS. (SC [ 3.5, AC [ 2.4.)
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temperature, H2 and CO concentrations in reformate gas of the OSR, HTS and LTS. It can be seen that when AC changes from 1.4 to 3.4 under the operating conditions of Fm ¼ 14 mol h1 and SC ¼ 4.0, OSR bed temperature increases significantly from 560 C to 650 C. Similar results are thermodynamically predicted and have been reported by others [12].
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The HTS and LTS bed temperatures also increase significantly while maintaining the same insulation conditions. Due to the excess nitrogen in the feed air, H2 concentration in the reformate gas is significantly diluted. Hence the H2 concentration in the reformate gas decreases with increasing the air to carbon ratio. Similar experimental results have been reported by Lee et al. [6]. The influence of air to carbon ration on CO concentration is quite complex. Higher bed temperature in the reformer helps to produce more CO due to the thermal dynamical oxidative steam methane reforming equilibriums. However, higher temperatures of the HTS and LTS beds help the conversion of CO to CO2 and H2. Excess nitrogen in the air also dilutes the CO concentration in the reformate gas. The measured variation of CO concentration with the air to carbon molar ratio is also plotted in Fig. 5. It can be seen that although the CO concentrations downstream of OSR and HTS are significantly affected by the air to carbon molar ratio, the variation of CO concentration downstream of the LTS is only between 0.02% and 0.05%，far below the development requirement of 0.3%e0.6%. It indicates that the fuel processing unit works well under the air to carbon molar ratio range of 1.4e3.4 in terms of CO concentration.
4.
Concluding remarks and future work
A fuel processor constituting of an oxidative steam methane reformer, a high temperature water gas shift reactor and a low temperature water gas shift reactor is developed and tested. Processor performances under typical operating conditions are tested. The influence of reaction temperature, OSR methane space velocity, and air to carbon molar ratio on performances of the OSR, HTS and LTS is experimentally investigated. It is found that for the current experimental setup: 1. Under the typical operating conditions of Fm ¼ 14.0 mol h1, FL ¼ 4.1 mol h1, AC ¼ 2.4, SC ¼ 4.0, the hydrogen and CO concentrations downstream of LTS is 45.56%, and 0.06% (dry basis), respectively. The total energy efficiency reaches 88.3%. The efficiency can further be improved by utilizing the sensible heat of the reformate gas. 2. Higher temperature improves the CO conversion in both HTS and LTS reactors for the temperature range investigated. 3. The influence of OSR methane space velocity on reactors’ chemical performance is a combination of both bed temperature and gas catalyst contact time. When the OSR space velocity increases from 82.5 h1e495.2 h1, CO concentration downstream of LTS increases slightly from 0.02% to 0.07%, which meets the development target of 0.3%e0.6%. 4. The fuel processing unit works well under the air to carbon molar ratio range of 1.4e3.4, as the variation of CO concentration downstream of the LTS is only between 0.02% and 0.05%, far below the development requirement of 0.3%e0.6%. Fig. 5 e Effect of air to carbon molar ratio on bed temperatures, H2 and CO concentrations downstream of OSR, HTS, and LTS. (Fm [ 14 mol hL1, FL [ 4.1 mol hL1, SC [ 3.5.)
The current study has been focused on the chemical performances of the reformer and water gas shift reactors. Heat integration of the unit will be further implemented in future to improve energy efficiency. The current fuel processing unit
10832
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will be further integrated with a CO selective oxidation reactor and a PEMFC to build a micro-CHP system.
Acknowledgment Financial supports from the Guangdong Nature Science Foundation (project # S20122010010314) and Guangzhou Scientific Research and Development Program (project # 2010JD00021) are gratefully acknowledged.
Nomenclature
AC c Fa FL Fm Fw SC X
air to carbon molar ratio in the feed molar species concentration in the reformate gas mixture, % flow rate of process air to OSR, mol h1 flow rate of LPG in gas phase to burner, mol h1 flow rate of process methane to OSR, mol h1 flow rate of process water to OSR, mol h1 steam to carbon molar ratio in the feed methane or CO conversion defined in Equation (1) or (2)
Acronyms HTS high temperature shift reactor LPG liquefied petroleum gas LTS low temperature shift reactor CHP combined heat and power cogeneration MFC mass flow controller OSR oxidative steam methane reformer PEMFC proton exchange membrane fuel cell
references
[1] Donitz W. Fuel cells for mobile applications, status, requirements and future application potential. Int J Hydrogen Energy 1998;23:611e5. [2] Okada O, Yokoyama K. Development of polymer electrolyte fuel cell co-generation systems for residential applications. Fuel Cells 2001;1:72e7.
[3] Gigliucci G, Petruzzi L, Cerelli E, Garzisi A, Mendola AL. Demonstration of a residential CHP System based on PEM fuel cells. J Power Sources 2004;131:62e8. [4] Hubert CE, Achard P, Metkemeijer R. Study of a small heat and power PEM fuel cell system generator. J Power Sources 2006;156:64e70. [5] Gencoglu MT, Ural Z. Design of a PEM fuel cell system for residential application. Int J Hydrogen Energy 2009;34: 5242e8. [6] Lee D, Lee HC, Lee KH, Kim S. A compact and highly efficient natural gas fuel processor for 1 kW residential polymer electrolyte membrane fuel cells. J Power Sources 2007;165: 337e41. [7] Andreasen SJ, Vang JR, Kær SK. High temperature PEM fuel cell performance characterisation with CO and CO2 using electrochemical impedance spectroscopy. Int J Hydrogen Energy 2011;36:9815e30. [8] Xie DL, Wang ZL, Jin L, Zhang YJ. Energy and exergy analysis of a fuel cell based micro combined heat and power cogeneration system. Energy Buildings 2012;50:266e72. [9] Xie DL, Peng A, inventor; South China University of Technology, assignee. A hydrocarbon reformer integrated with steam generation. 23 May 2012, Chinese patent ZL201010298945.8. [10] Xie DL, Peng A, Wang ZL, Zhang YJ. Hydrogen production for fuel cells by a methane reformer integrated with steam generation. Int J Chem React Eng 2012;10(1). http://dx.doi.org/ 10.1515/1542-6580.2980. [11] Heinzel A, Vogel B, Hubner P. Reforming of natural gashydrogen generation for small scale stationary fuel cell systems. J Power Sources 2002;105:202e7. [12] Lee SHD, Applegate DV, Ahmed S, Calderone SG, Harvey TL. Hydrogen from natural gas: part I: autothermal reforming in an integrated fuel processor. Int J Hydrogen Energy 2005;30: 829e42. [13] Lin ST, Chen YH, Yu CC, Liu YC, Lee CH. Dynamic modeling and control structure design of an experimental fuel processor. Int J Hydrogen Energy 2006;31:413e26. [14] Adachi H, Ahmed S, Lee SHD, Papadias D, Ahluwalia RK. A natural-gas fuel processor for a residential fuel cell system. J Power Sources 2009;188:244e55. [15] Cornforth JR. Combustion engineering and gas utilisation. 3rd ed. London: E & FN SPON; 1992. [16] Xie DL, Qiao WY, Wang WX, Yu H, Peng F. Reaction/ separation coupled equilibrium modeling of steam methane reforming in fluidized bed membrane reactors. Int J Hydrogen Energy 2010;35:11798e809. [17] Ye GY, Xie DL, Qiao WY, Grace JR, Lim JC. Modeling of fluidized bed membrane reactors for hydrogen production from steam methane reforming with Aspen Plus. Int J Hydrogen Energy 2009;34:4755e62.
Available online at www.sciencedirect.com
journal homepage: www.elsevier.com/locate/he
Syngas production from oxidative methane reforming and CO cleaning with water gas shift reaction Donglai Xie*, Jie Zhao, Ziliang Wang, Yajun Zhang The Key Laboratory of Fuel Cell Technology of Guangdong Province, South China University of Technology, Guangzhou 510640, China
article info
abstract
Article history:
Fuel cell based heat and power cogeneration is considered to be well qualified for a dis-
Received 5 November 2012
tributed energy system for residential and small business applications. A fuel processing
Received in revised form
unit including an oxidative steam methane reformer, a high temperature shift reactor and
27 December 2012
a low temperature shift reactor is under development in South China University of Tech-
Accepted 1 January 2013
nology. Performance of the unit is experimentally investigated in a bench-scale exper-
Available online 29 January 2013
imental setup. Processor performance under typical operating conditions is tested. The influence of reaction temperature, methane space velocity in the oxidative steam methane
Keywords:
reformer, and air to carbon molar ratio on unit performances is experimentally studied. It
Hydrogen
is found that under the typical operating conditions, the total energy efficiency reaches
Reforming
88.3%. The efficiency can further be improved by utilizing the sensible heat of the refor-
Water gas shift reaction
mate gas. The current study has been focused on the chemical performances such as
Combined heat and power
methane conversion of the reformer and CO concentration in the synthesis gas down-
cogeneration
stream water gas shift reactors. Heat integration of the unit will be further implemented in future to improve energy efficiency. Copyright ª 2013, Hydrogen Energy Publications, LLC. Published by Elsevier Ltd. All rights reserved.
1.
Introduction and background
Fuel cell technologies have been of special interest in recent years due to their high efficiency in energy conversion and less/no harmful emissions compared with other energy conversion systems [1]. Fuel cell based heat and power cogeneration (CHP) is considered to be well qualified for a distributed energy system for residential and small business applications, which generally incorporates the proton exchange membrane fuel cells (PEMFCs) with hydrogen production from natural gas or other fuels [2e5]. These applications, also known as microCHP systems, offer benefits of on site and real time generation of electricity with a high efficiency and utilization of the heat
produced during operation of fuel cells and fuel processing for hot water generation and residential heating [6]. When introducing fuel cell based micro-CHP systems into the market, small-scale hydrogen generation systems from various hydrocarbon feed stocks are required due to the lack of hydrogen distribution infrastructure. A kW-scale hydrogen production unit is under development in the New Energy Group in South China University of Technology. Natural gas is selected as feedstock for hydrogen production, as its infrastructure is well established in most countries. The fuel processing unit constitutes of an oxidative steam methane reformer (OSR), a high temperature shift reactor (HTS) and a low temperature shift reactor (LTS). The
* Corresponding author. Tel./fax: þ86 20 22236985. E-mail addresses: [email protected], [email protected] (D. Xie). 0360-3199/$ e see front matter Copyright ª 2013, Hydrogen Energy Publications, LLC. Published by Elsevier Ltd. All rights reserved. http://dx.doi.org/10.1016/j.ijhydene.2013.01.012
i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 3 8 ( 2 0 1 3 ) 1 0 8 2 6 e1 0 8 3 2
targeted CO concentration in the reformate gas out of the unit is 0.3%e0.6%. The OSR converts natural gas, steam and air into a hydrogen rich gas. The outlet gas is then fed to the CO shift reactors where CO concentration is expected to be reduced to 2.0%e4.0% in the HTS and to 0.3%e0.6% in the LTS. The unit can directly supply feed gas to a high temperature PEMFC [7]. If a regular low-temperature PEMFC is employed, the CO concentration in the product gas from the unit needs to be further reduced to be less than 10 ppm through selective CO oxidation or CO methanation. The current study focuses on the design and chemical performance of theses reactors under various operating conditions, while the heat integration will be implemented later in the system assemble stage. Fuel cell based micro-CHP system is a novel energysupplying technology represented the main developing trends for the future’s residential energy supply. However, the technology is far beyond maturity. Very few testing results have been published. The experimental results presented could help to understand the mechanisms and performances of the fuel processing unit in the CHP system.
2. The reactors design and experimental setup 2.1.
The reactors design
The sizing of the OSR, HTS and LTS are based on the simulation result of a 1 kW scale fuel cell based micro-CHP system with Aspen Plus [8]. To produce H2 rich reformate gas from natural gas, water and air, a reformer integrated with steam generation is designed and patented [9]. The integrated reforming system constitutes of a reforming reactor, a combustion chamber, a burner, an external methane heating coil, an external water coil, an internal mixture gas coil, a flue gas outlet nozzle, and a reformate gas outlet nozzle. Methane and water enter the external methane and external water coils outside the combustion chamber, respectively. They then join together and enter the internal mixture gas coil around the reforming reactor. The reactor is essentially a stainless steel tube with an inner diameter of 80 mm and a length of 200 mm, where reforming catalyst is filled. An air distributor is inserted inside the reactor for oxidative steam methane reforming. The air distributor is made of a 6 mm outer diameter stainless steel tube with four pairs of 2 mm diameter orifices (8 holes in total) drilled through the tube wall. 800 g of commercial nickel catalyst Z413Q supported on gamma alumina are loaded in the reactor. The catalyst is supplied by Shandong Qilu Keli Chemical Institute Co. (China) and has been widely used in industrial natural gas reforming systems in China. The details of the reformer and catalyst can be found elsewhere [10]. The HTS, as well as the LTS, is made of a 200 mm long stainless steel tube with inner diameter of 60 mm. Feed gas enters the vessel from a nozzle on its top flange and the reactant gas leaves the reactor from a nozzle on its bottom flange. Commercial catalyst KLB-101/9 with active component of Fe2O3eCr2O3 is loaded in the HTS, while commercial catalyst KLB-201 with active component of CuOeZnO is loaded in the LTS. Both catalysts are supplied by Shandong Qilu Keli Chemical Institute Co. (China). Both the HTS and LTS are
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equipped with electrical heaters for bed heat up during startup. Insulation ribbons are wrapped outside of these reactors. The number of the layers of the insulation ribbon can be manually changed during the experiments in order to maintain the bed temperature at a desired value, as the water gas shift reaction is exothermal.
2.2.
Experimental setup
Performance of the hydrogen production unit is tested in a bench-scale experimental setup. Schematic layout of the setup is shown in Fig. 1. Due to the lack of natural gas distribution infrastructure in the laboratory, bottled methane (industrial pure, supplied by Guangzhou Zhuozheng Industrial Gas Ltd., China) rather than piped natural gas is employed as process feed. Bottled liquefied petroleum gas (LPG, supplied by Guangzhou Gas Ltd., China) is fed into the burner for combustion. The volumetric composition of the LPG is 1.0% C2H6, 9.0% C3H6, 4.5% C3H8, 54.0% C4H8, 26.2% C4H10, and 5.3% Cþ 5. The flow rate of process CH4 is controlled by a mass flow controller (MFC). The flow rates of process air and water are monitored by rotameters. The flow rate of LPG for combustion is measured by a wet gas flow meter. The process methane and water under room temperature are preheated by hot flue gas issued from the burner to ensure that the water is in vapor phase before being injected into the reactor. The methane and water are heated separately in the external methane and water coils outside the combustion chamber. They are further heated to a temperature around 500 C in the internal mixture gas coil. The preheated methane and steam mixture then proceed to the catalyst reactor to produce a hydrogen rich reformate gas containing largely H2, CO2, CO, and steam. Air is introduced through the eight orifices of the air distributor. The reactor off gas from the OSR is directed to HTS and LTS in sequence. The reformate gas compositions downstream of each reactors are analyzed by a gas chromatography (GC910, Shanghai Kechuang Chromatograph Instrument Ltd., China) equipped with a thermal conductivity detector. The connection tubes between reactors are well insulated.
2.3.
Unit performance characterization
The OSR performance is characterized by reformate gas composition (mainly H2 and CO concentration) and methane conversion. Methane conversion is defined as: XCH4 ¼
cCO þ cCO2 in the product gas cCH4 þ cCO þ cCO2 in the product gas
(1)
The HTS and LTS performances are characterized by CO concentration in the outlet gas and CO conversion. The CO conversion is defined as: XCO ¼ 1
cCO in the product gas cCO in the feed gas
(2)
The influence of following operating conditions on reactor performance is studied: (1) reaction temperature; (2) air to carbon molar ratio in the feed (AC); and (3) methane space velocity in OSR. OSR methane space velocity is defined as the
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Fig. 1 e Schematic layout of the experimental setup.
ratio between the volumetric flow rates of CH4 to OSR under standard conditions (273.15 K and 0.1 MPa) and the catalyst volume in OSR.
3.
Experimental results and discussion
The hydrogen production unit has experienced intensive tests after being built. More than 600 h’ operation has been carried out aperiodically in a six-month period with different operating conditions. The unit has shown stable functionality and no obvious deactivation of the catalysts is observed.
3.1.
Performance of a typical case
Table 1 shows the performance of the OSR, HTS and LTS under typical operating conditions of Fm ¼ 14.0 mol h1, FL ¼ 4.1 mol h1, Fa ¼ 33.6 mol h1 (AC ¼ 2.4), and Fw ¼ 56.0 mol h1 (SC ¼ 4.0). Under such conditions, the hydrogen concentrations downstream of the OSR, HTS and LTS are 40.61%, 43.45% and 45.56% (dry basis), respectively. The hydrogen concentration of 45.56% downstream of LTS is higher than 30.5% reported by Heinzel et al. [11], 40% by Lee et al. [12], and lower than 51.3% reported by Lin et al. [13] and 48% by Adachi et al. [14]. The methane conversion in the OSR is
92.56%, while the CO conversion is 59.62% in the HTS and 97.47% in the LTS. The CO concentrations downstream of these reactors are 5.56%, 2.25% and 0.06% (dry basis), respectively. It is obvious that the CO concentration downstream of the LTS meet the design target of 0.3%e0.6%. The CO concentration of 0.06% downstream of the LTS is much lower than 8.9% reported by Heinzel et al. [11]. The oxidative steam methane
Table 1 e Performance of the OSR, HTS and LTS under typical operating conditions. (Fm [ 14.0 mol hL1, FL [ 4.1 mol hL1, Fa [ 33.6 mol hL1, and Fw [ 56.0 mol hL1.) Reactor Average bed temperature, C Downstream gas composition, % H2 N2 CO CH4 CO2 Key component conversion, %
OSR
HTS
LTS
630
420
300
40.61 37.63 5.56 3.62 12.58 92.56 (CH4)
43.45 37.12 2.25 3.27 13.91 59.62 (CO)
45.56 36.11 0.06 3.18 15.10 97.47 (CO)
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reforming reactions are limited by thermodynamic constraints. Experimental results from the same reactor under steam methane reforming and oxidative steam methane reforming operating conditions have been compared with the predicted values by Gibbs free energy minimization models in a previous study [10]. As stated previously, in the experimental test rig, the flow rates of process feeds (methane, water and air) and combustion LPG are all measured. Flue gas composition is analyzed by a flue gas analyzer. The combustion air flow rate can be calculated from the oxygen concentration in the flue gas and LPG flow rate [15]. Syngas composition downstream of each reactor is measured by a gas chromatography. Gas temperatures of all streams are intensively monitored by thermocouples as shown in Fig. 1. Hence, enthalpy of these streams can be calculated from the stream flow rate, composition and temperature. Fig. 2 shows the flow of enthalpy in each unit operation of the setup from these calculations. The enthalpy loss in each reactor and the connection tube is shown as DE in the figure. The total energy efficiency, which is defined as the total enthalpy carried by the reformate gas and flue gas divided by the enthalpy carried by feed methane and combustion LPG, is 88.3% under the typical operating conditions. Here the energy carried by flue gas is considered as useful energy, as its energy can be further employed in the downstream process to produce hot water. A major heat loss takes place in the connecting tubes. When the temperature of the reformate gas decreases from 630 C in the OSR to 420 C in the HTS and further to 300 C in the LTS, the heat released by the reformate gas temperature drop has not been recovered in the current configuration. When a micro-CHP system is built, the sensible heat of the reformate gas can be employed to produce hot water, and the system energy efficiency can be improved. Another significant energy loss is caused by the burner, which has a total heat loss of 1849.7 kJ h1, corresponding to 7.3% of the total energy input. The current burner is a domestic boiling table. In the experiments it is observed that some of the flue gas has escaped from the OSR heat transfer surfaces to the environment. A special designed atmospheric burner is under construction to improve flue gas flow pattern.
3.2.
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Effect of HTS and LTS bed temperatures
The effect of bed temperature on reformer performance has been previously studied extensively [8,10,16,17]. Hence it is not studied here. The water gas shift reaction is significantly affected by bed temperature. As the water gas shift reaction is not controlled by thermodynamic equilibrium, higher bed temperature helps the reaction kinetics. Fig. 3 shows that the CO concentration decreases from 2.8% to 2.0% in the HTS when the bed temperature increases from 220 C to 460 C and the feed gas molar composition is maintained at cH2 ¼ 44:27%, cN2 ¼ 35:48%, cCO ¼ 5.78%, cCH4 ¼ 3:25%, cCO2 ¼ 11:22%. For the LTS, when the bed temperature increases from 120 C to 290 C and the feed gas molar composition maintains at cH2 ¼ 443:70%, cN2 ¼ 36:87%, cCO ¼ 2.07%, cCH4 ¼ 3:26%, cCO2 ¼ 14:00%, the CO concentration decreases from 2.0% to 0.09%. Obviously, higher temperature improves the CO conversion in both HTS and LTS reactors for the temperature range investigated. Due to the limitation of the electrical heater capacity and the insulation conditions, higher bed temperatures for the HTS and LTS have not been tested.
3.3.
Effect of OSR methane space velocity
Gas space velocity is a crucial parameter influencing catalyst performance. It shows the amount of feed that the reactor can handle per unit volume of catalyst. When the space velocity is low, the contact time between the reactant and catalyst is long. Hence it is easy for the reaction to reach equilibrium. In the current study, the OSR, HTS and LTS performances are all experimentally investigated against the OSR methane space velocity. In the study, the process methane feed changes from 3.5 mol h1 to 21 mol h1, while the steam to carbon molar ratio maintains at 3.5 and the air to carbon molar ration maintains at 2.4, corresponding to an OSR methane space velocity from 82.5 h1e495.2 h1. For such a small scale chemical plant, heat dissipation to the environment is a serious issue, as the heat carried by gas stream is limited due to its very low flow rate. In the experiments, gas space velocity is changed by varying the feed flow rate while maintaining the catalyst volume unchanged. Hence,
Fig. 2 e Energy flow within the fuel processing unit under typical operating conditions (Fm [ 14.0 mol hL1, FL [ 4.1 mol hL1, AC [ 2.4, SC [ 4.0).
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Fig. 3 e Influence of bed temperature on CO concentration in HTS (Fm [ 14 mol hL1, SC [ 4.0, AC [ 2.4, inlet gas composition: cH2 [44:27%, cN2 [35:48%, cCO [ 5.78%, cCH4 [3:25%, cCO2 [11:22%) and LTS. (Fm [ 14 mol hL1, SC [ 4.0, AC [ 2.4, inlet gas composition: cH2 [443:70%, cN2 [36:87%, cCO [ 2.07%, cCH4 [3:26%, cCO2 [14:00%.)
low space velocity corresponds to low feed gas and reformate gas flow rates. When the reformate gas flow rate is low, it is difficult to maintain a bed temperature much higher than the environment. The top part of Fig. 4 shows the influence of methane space velocity on bed temperatures of OSR, HTS and LTS. It is obvious that bed temperatures increase with increasing OSR methane space velocity. Hence, the influence of space velocity on reactors’ chemical performance is a combination of both bed temperature and gas catalyst contact time. As shown in Fig. 4, the methane conversion of the OSR is not significantly affected by the space velocity, as well as the CO concentration in the reformate gas. It indicates that the OSR catalyst works well in the current space velocity range. For the HTS, it can be seen that when the OSR methane space velocity increases from 82.5 h1e495.2 h1, CO conversion decreases from 70% to 50%, while CO concentration in the reformate gas increases slightly. It indicates that the OSR space velocity has notable influences on HTS performance. For the LTS, when the OSR space velocity increases from 82.5 h1e495.2 h1, CO conversion decreases slightly from 99% to 98%, while CO concentration in the reformate gas increases slightly from 0.02% to 0.07%, which is far below the development requirement of 0.3%e0.6%. Hence the fuel processor works well under the OSR methane space velocity range of 82.5 h1 to 495.2 h1
3.4.
Effect of air to carbon molar ratio
Air to carbon ratio is an important factor for oxidative steam reforming. It has dominant impact on reaction temperature and product gas composition. As the ratio increases, the rate of the oxidation reaction increases and consequently heats up the reactor. This leads to a higher methane conversion. However, excess oxygen will also consume hydrogen and CO in the product gas, leading to a decline in H2 and CO yields in the OSR. Fig. 5 shows the influence of air to carbon molar ratio on bed
Fig. 4 e Effect of OSR methane space velocity on bed temperatures, methane conversion in OSR, CO conversions in HTS and LTS, and CO concentrations downstream of OSR, HTS, and LTS. (SC [ 3.5, AC [ 2.4.)
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temperature, H2 and CO concentrations in reformate gas of the OSR, HTS and LTS. It can be seen that when AC changes from 1.4 to 3.4 under the operating conditions of Fm ¼ 14 mol h1 and SC ¼ 4.0, OSR bed temperature increases significantly from 560 C to 650 C. Similar results are thermodynamically predicted and have been reported by others [12].
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The HTS and LTS bed temperatures also increase significantly while maintaining the same insulation conditions. Due to the excess nitrogen in the feed air, H2 concentration in the reformate gas is significantly diluted. Hence the H2 concentration in the reformate gas decreases with increasing the air to carbon ratio. Similar experimental results have been reported by Lee et al. [6]. The influence of air to carbon ration on CO concentration is quite complex. Higher bed temperature in the reformer helps to produce more CO due to the thermal dynamical oxidative steam methane reforming equilibriums. However, higher temperatures of the HTS and LTS beds help the conversion of CO to CO2 and H2. Excess nitrogen in the air also dilutes the CO concentration in the reformate gas. The measured variation of CO concentration with the air to carbon molar ratio is also plotted in Fig. 5. It can be seen that although the CO concentrations downstream of OSR and HTS are significantly affected by the air to carbon molar ratio, the variation of CO concentration downstream of the LTS is only between 0.02% and 0.05%，far below the development requirement of 0.3%e0.6%. It indicates that the fuel processing unit works well under the air to carbon molar ratio range of 1.4e3.4 in terms of CO concentration.
4.
Concluding remarks and future work
A fuel processor constituting of an oxidative steam methane reformer, a high temperature water gas shift reactor and a low temperature water gas shift reactor is developed and tested. Processor performances under typical operating conditions are tested. The influence of reaction temperature, OSR methane space velocity, and air to carbon molar ratio on performances of the OSR, HTS and LTS is experimentally investigated. It is found that for the current experimental setup: 1. Under the typical operating conditions of Fm ¼ 14.0 mol h1, FL ¼ 4.1 mol h1, AC ¼ 2.4, SC ¼ 4.0, the hydrogen and CO concentrations downstream of LTS is 45.56%, and 0.06% (dry basis), respectively. The total energy efficiency reaches 88.3%. The efficiency can further be improved by utilizing the sensible heat of the reformate gas. 2. Higher temperature improves the CO conversion in both HTS and LTS reactors for the temperature range investigated. 3. The influence of OSR methane space velocity on reactors’ chemical performance is a combination of both bed temperature and gas catalyst contact time. When the OSR space velocity increases from 82.5 h1e495.2 h1, CO concentration downstream of LTS increases slightly from 0.02% to 0.07%, which meets the development target of 0.3%e0.6%. 4. The fuel processing unit works well under the air to carbon molar ratio range of 1.4e3.4, as the variation of CO concentration downstream of the LTS is only between 0.02% and 0.05%, far below the development requirement of 0.3%e0.6%. Fig. 5 e Effect of air to carbon molar ratio on bed temperatures, H2 and CO concentrations downstream of OSR, HTS, and LTS. (Fm [ 14 mol hL1, FL [ 4.1 mol hL1, SC [ 3.5.)
The current study has been focused on the chemical performances of the reformer and water gas shift reactors. Heat integration of the unit will be further implemented in future to improve energy efficiency. The current fuel processing unit
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will be further integrated with a CO selective oxidation reactor and a PEMFC to build a micro-CHP system.
Acknowledgment Financial supports from the Guangdong Nature Science Foundation (project # S20122010010314) and Guangzhou Scientific Research and Development Program (project # 2010JD00021) are gratefully acknowledged.
Nomenclature
AC c Fa FL Fm Fw SC X
air to carbon molar ratio in the feed molar species concentration in the reformate gas mixture, % flow rate of process air to OSR, mol h1 flow rate of LPG in gas phase to burner, mol h1 flow rate of process methane to OSR, mol h1 flow rate of process water to OSR, mol h1 steam to carbon molar ratio in the feed methane or CO conversion defined in Equation (1) or (2)
Acronyms HTS high temperature shift reactor LPG liquefied petroleum gas LTS low temperature shift reactor CHP combined heat and power cogeneration MFC mass flow controller OSR oxidative steam methane reformer PEMFC proton exchange membrane fuel cell
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