INTERNATIONAL JOURNAL OF ENERGY RESEARCH Int. J. Energy Res. 2013; 37:49–58 Published online 6 July 2011 in Wiley Online Library (wileyonlinelibrary.com). DOI: 10.1002/er.1882
RESEARCH ARTICLE
200-MW chemical looping combustion based thermal power plant for clean power generation Raman Sharma1,*,†, Munish Kumar Chandel2, Arnaud Delebarre3 and Babu Alappat1 1
Department of Civil Engineering, Indian Institute of Technology Delhi, Hauz Khas, New Delhi 110016, India Center on Global Change, Climate Change Policy Partnership, Duke University Durham, Box 90658, Durham NC-27708, USA 3 École Superieure des Sciences et Technologies de L’Ingénieur de Nancy, 2 rue Jean Lamour, Vandoeuvre Lès Nancy Cedex, F-54519, France 2
SUMMARY The present study demonstrates a possible configuration of a 200MW chemical looping combustion (CLC) system with methane (CH4) as fuel. Iron oxide-based oxygen carriers were used because of its non-toxic nature, low-cost, and wide availability. We analyzed the effects of different variables on the design of the system. For the air reactor (oxidizer), bed mass is independent, and for the fuel reactor (reducer), it decreases with increase in the conversion difference between the air and fuel reactors. On the other hand, the pressure drop in the air reactor is unchanged, whereas for the fuel reactor, it decreases with the same increase of conversion difference between air and fuel reactors. Also, entrained solid mass flow rate from the air to fuel reactor shows a decreasing trend. Bed mass, bed height, pressure drop, and residence time of the bed materials decrease with increase in the conversion rates in the air and fuel reactors. Residence time of bed material in the air and fuel reactor reduces with increase in the temperature of the air reactor. Copyright © 2011 John Wiley & Sons, Ltd. KEY WORDS chemical looping combustion; energy; CO2 capture; global warming; thermal power plant Correspondence *Raman Sharma, Department of Civil Engineering, IIT, Delhi, 110016, India. † E-mail:
[email protected] Received 7 February 2011; Revised 30 May 2011; Accepted 1 June 2011
1. INTRODUCTION Scientific consensus attributes the cause of global warming to the increase in the concentration of atmospheric greenhouse gases because of human activities [1]. Although the discovery of green house effect in 1824, the green house gasses have never shown any sign of decreasing. The current level is roughly 430ppm CO2 equivalent and is increasing at an average concentration of 2ppm per year [2]. To keep pace with the rapid economic growth, nations across the globe require fossil fuels. Globally, prime energy needs are satisfied by burning fossil fuels in thermal power plants. CO2 is an inevitable product of the burning of carbon-containing fuels and is the main cause behind the global warming. Hence, greenhouse gas emission reduction, especially CO2 emission reduction, is required. Several techniques and alternative strategies are being investigated these days to minimize the use of fossil fuels. However, fossil fuels would continue to be in use in the foreseeable future. In this situation, carbon capture and storage may
Copyright © 2011 John Wiley & Sons, Ltd.
be a solution to the fossil fuel-related global warming, where CO2 is separated from the industrial sources and transported to the storage locations for long term isolation from the atmosphere. This fact encourages the development of CO2 capture techniques like chemical looping combustion (CLC). CLC is a technology to capture CO2 at the power plants. Chemical looping combustion is a promising technology with inherent separation of CO2. No energy is wasted in the separation of CO2 from the flue gas; however, the technology is not yet demonstrated at commercial scale. CLC involves solid carriers to supply oxygen to the fuel for combustion [3]. Metal oxides are recycled between an air reactor, where these are oxidized by the air, and a fuel reactor, where these are reduced by fuel. With this arrangement, air is never mixed with the fuel, and the generated CO2 does not dilute with the nitrogen of air [4]. The outgoing gas from the oxidation step will contain nitrogen and unreacted oxygen. The reduction step gives CO2 and H2O from the fuel reactor. Condensation will remove H2O, and nearly pure CO2 can be collected.
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1.1. Global energy use pattern and carbon capture
2. CHEMICAL LOOPING COMBUSTION (CLC)
Natural gas has a share of 22% of world energy mix, which makes it one of the most important fuels after petroleum (39%) and coal (25%) [4]. Public electricity and heat production contribute 34.8% of the greenhouse gas emissions globally [5]. By the year 2050, renewables could supply up to 30% of the world’s energy needs. Hence, fossil fuels will continue to be used in the near future [6]. Carbon capture and storage has been presented as an approach for fossil fuel based CO2 emission reduction. There are different precombustion and post-combustion CO2 capture technologies under different stages of the development. The problem with existing CO2 capture techniques is that they are expensive and have a large energy penalty [7]. An example, a coal-based integrated gasification combined cycle (IGCC) plant employing CO2 capture may increase the cost of electricity generated by 25%, and an oxy fuel combined process may consume 25–28% of electricity generated [8]. Similarly, a natural gas-based power plant employing the post combustion amine-based CO2 capture may reduce the efficiency of plant by 8–10% [4]. On the other hand, if the CO2 capture takes place using CLC, the efficiency will reduce by only 2–3% points [4]. Also, a syngas chemical looping (SCL) process co-producing hydrogen and electricity has an overall efficiency of 64% with 100% carbon capture [8]. The design of 200-MW CLC system is presented in this paper. This could be the size of an initially built industrial scale CLC system. CH4 is used fuel as the CLC is better poised for the gaseous fuels than for solid fuels. An excel-based program developed by the TOTAL (www. total.com), based on the equations presented by Lyngfelt et al. [3] and Chandel et al. [9], and mass and energy balance equations are used in the design. Heat and mass flow sheet is presented for the system. We also evaluated the effect of different variables on design of a CLC system.
Chemical looping combustion has been explained in the literature as the combustion that may be carried out in two successive reactions [3,10]. The CLC system consists of two interconnected fluidized bed reactors, an air reactor, and a fuel reactor (Figure 1). Fuel is introduced in gaseous form into the fuel reactor containing a metal oxide, MeO. The metal oxides react with the fuel according to the following equation: ð2n þ mÞ MeO þ Cn H2m ! ð2n þ mÞ Me þ m H2 O þ n CO2
(1) The outgoing gas stream from the fuel reactor contains CO2 and H2O. Nearly pure CO2 can be collected after the condensation of water vapor. The reduced oxygen carrier, Me, is circulated back to the air reactor, where it is oxidized to MeO. Me þ 1=2 O2 ! MeO
(2)
The flue gas from the air reactor will contain N2 and any unreacted O2. The extent of the reaction depends on the metal oxides and the reaction conditions. Reaction (1) is often endothermic and reaction (2) is exothermic with respect to metal oxides. Heat of the reaction from (1) and (2) is same as that of conventional combustion of fuel with oxygen [3]. The advantage of CLC as compared with conventional combustion is that the N2 present in the air does not dilute the CO2. CO2 is inherently separated with no extra energy required. Chemical looping combustion system could have a similar configuration as that of the circulating fluidized bed (CFB) combustion system where the main riser of CFB is an air reactor and loop–seal is a fuel reactor. Several
N2 and unadsorbed O2
CO2 and H2O
MeO (+Me)
Fuel Reactor
Air Reactor Me (+MeO)
Air
Fuel
Figure 1. Reactors arrangement in Chemical looping combustion, MeO/Me denotes metal oxide[3].
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researchers have advocated this kind of configuration. The air reactor will be a high-velocity reactor operating in a fast-fluidized bed regime and the fuel reactor will be a low velocity reactor [3,11–13]. The concept of dual circulating fluidized beds has been presented by Kolbitsch et al. [14]. CLC has been proposed as two interconnected pressurized fast fluidized bed reactors by Wolf et al. [4], and two high velocity beds with a j-type loop have been presented by Mattisson et al. [15]. Most of the researchers have discussed gaseous fuels for the CLC process using iron, nickel, copper, and manganese-based oxygen carriers. There are some studies discussing the adaptability of the CLC for solid fuels. The main problem with burning solid fuel in CLC is the low solid–solid reaction rate between the solid fuel and the metal oxide. Instead, the solid fuel needs to be gasified, and the generated gases then can be burned using the metal oxides in a CLC system [16]. 2.1. Gaseous fuels and metal oxide combinations for chemical looping combustion The chemical looping combustion process requires high heat and mass transfer between the gas and solid oxygen carriers. Hence, two interconnected fluidized beds have an advantage over other designs [11]. Metal oxides used in CLC must possess high rates of oxidation and reduction, enough strength to tolerate particle breakage and attrition. Low cost and environmental friendliness add on to the advantages of the metal oxide. Oxides of Cd, Co, Cu, Fe, Mn, and Ni have been enlisted in the literature for the application in CLC [17,18]. Gaseous fuels like liquefied natural gas (LNG), CH4, syngas, and hydrogen have widely been tested for the CLC process using Fe-based oxygen carrier having Yttria-stabilized zirconia (YSZ), Al2O3, ZrO2, TiO2, and MgAl2O4 as support materials [11,12,15,18,19]. Fe2O3/ MgAl2O4 metal oxide system has shown maximum reactivity in the study of Mattisson et al. [19]. In the study of Cho et al. [18], Cu-, Ni-, and Fe-based oxygen carriers have shown high reactivity. However, the copper-based oxygen carriers suffer from the problem of agglomeration [4]. Fe-, Mn-, and Cu-based carriers have been found to be very reactive, environmental friendly, and of low cost in the study of Mattisson et al. [15], whereas a nickelbased carrier is slightly expensive. In his study of Abad et al. [12], Fe-based oxygen carriers have shown low agglomeration, low carbon deposits, and no mass losses. High-crushing strength, low attrition, high durability, and high reactivity have been observed for the iron based metal oxides. All these gave the authors a reason to use iron as an oxygen carrier. Moreover, ironbased oxygen carriers are easily available, environmental friendly, and are of low cost. Table I presents the characteristic properties of the oxygen carrier used in this study. The aging and the variation in the carrier characteristics are not considered. It was assumed that the flue gases from the fuel reactor do not have any fuel content, and there is Int. J. Energy Res. 2013; 37:49–58 © 2011 John Wiley & Sons, Ltd. DOI: 10.1002/er
no carbon deposition in the fuel reactor. As described by [10,20], there is no thermal NOx formation at the temperature of 1000 C and, hence, it is assumed that the N2 of the air remains the same in the flue gases.
3. GAS–SOLID HYDRODYNAMICS 3.1. Continuity equations The gas continuity equations for the gas phase and solid phase [21] are given by Equations (3) and (4), respectively Xn @ g eg rg þ r eg rg Vg ¼ r n¼1 gn @t
(3)
There are M solid-phase continuity equations of type XnSm @ eSm rSm Þ þ r eSm rSm VSm Þ ¼ r n¼1 Smn @t
(4)
where e, r, and V are the volume fraction (dimensionless), the density (kg/m3), and the velocity (m/s). rn is the rate of production of the nth chemical species (kg/m3 s). n is the total number of chemical species. The first term on the left in Equation (3) and (4) accounts for the rate of mass accumulation per unit volume, and the second term is the net rate of convective mass flux. The term on the right accounts for interparticle mass transfer because of the chemical reactions. 3.2. Energy equations The internal energy balance [22] for the fluid phase is written in terms of the fluid temperature eg rg CPg
@Tg þ Vg rTg @t
¼ r qg Hg1 Hg2 ΔHrg þ Hwall Twall Tg
(5) where qg is the fluid-phase conductive heat flux (J/m2 s), Hg1, and Hg2 describe fluid–solids interphase heat transfer (J/m3 s), Hrg is the heat of reaction (J/m3 s), and the last
Table I. Numerical values for the parameters used in the design. Fluidizing velocity in air reactor (m/s) Fluidizing velocity in fuel reactor (m/s) Air reactor temperature ( C) Outlet gas temperature ( C) Conversion of oxygen carriers in air reactor Conversion difference in oxidizer and reducer (ΔX) Conversion rate in the oxidizer (%/min) Conversion rate in the reducer (%/min) Oxidation heat reaction (kJ/mol O2) Apparent particle density (kg/m3)
9.6 0.16 1000.0 450.0 0.98 0.02 Fe2O3/Al2O3 13 4 481.04 3250
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term accounts for the heat loss to the wall. The thermal energy balance for the M=1 solid phase is given by @TS1 eS1 rS1 CPS1 þ VS1 rTS1 ¼ r qS1 þ Hg1 ΔHrS1 @T (6) where qS1 is the solid-phase-1 conductive heat transfer (J/m2 s), HrS1 is the heat of reaction (J/m3 s). All other solids phases are assumed to be in thermal equilibrium to simplify the numerical solution of the energy equations. The thermal energy balance for all the other solids phases (m=2 to M) in terms of an average temperature TS2 (K), is XM @TS2 e r C rT (7) þ V Sm PSm Sm S2 Sm m¼2 @t ¼ r qS2 þ Hg2 ΔHrS2 where, qS2 is the average solid-phase conductive heatflux (J/m2 s), Hg2 is fluid–solid interphase heat transfer, and HrS2 is the heat of reaction (J/m3 s).
mbed;ox ¼
m fuel ¼
Pfuel Hi
hbed;ox
hbed;red ¼
m fuel Sr Mfuel
(9)
The air ratio l is given by the volume fraction of oxygen in the air from the oxidizer, XO2 ;ex 0:21 1 XO2 ;ex l¼ (10) 0:21 XO2 ;ex 0.21 is the volume fraction of oxygen in air. The air mass flow m air is calculated by m l (11) m air ¼ o 0:233 where 0.233 is the mass fraction of oxygen in air. The fluidizing velocities in oxidizer uox and reducer ured are m air Va;air Tox Aox Ta
(12)
m fuel Va; fuel Rr Tred Ared Ta
(13)
uox ¼
ured ¼
Rr is the recirculation ratio of the total flow of gas supplied to reducer and the fuel flow. Bed mass and bed height in oxidizer and reducer are calculated by 52
mbed;red rsð1 eÞAred
(14) (15) (16)
(17)
the mass fraction of oxygen in the carrier Ro (in its fully oxidized state) Ro ¼
Mf :ox Mf :red Mf :ox
(18)
The degrees of oxidation Xox, reduction Xred, and the difference in conversionΔX are given by Xox ¼
Mactual:ox Mf :red Mf :ox Mf :red
(19)
Xred ¼
Mf :ox Mactual:red Mf :ox Mf :red
(20)
ΔX ¼ Xox Xred
(21)
(8)
The amount of oxygen needed, m o , for oxidation of the fuel is m o ¼ MO2
ΔPred Ared g mbed;ox ¼ rsð1 eÞAox
mbed;red ¼
4. DESIGN METHODOLOGY The design is based on the equations described by [3,23,24] with the assumption of complete combustion, the fuel consumption m fuel is calculated using
ΔPox Aox g
The presented system is a circulating one, with two interconnected fluidized beds, a fast fluidized bed air reactor and a low velocity fuel reactor. To capture the outgoing solid particles from the air reactor, a cyclone is used. The size of the metal oxides particles is assumed to be 200 microns with an apparent density of 3250kg/m3. The superficial gas velocity in the air and fuel reactors are taken as 9.6m/s and 0.16m/s, respectively based on the particle size used and the regimes of fluidization to be maintained in the reactors. The inlet gas velocity in the fuel reactor was in the range of 10Umf . The air flow rate in the air reactor was 70m3/s. A residence time of 9s is selected in the air reactor for the oxidation of the oxygen carrier. CH4 (Lower Calorific Value (LCV) of 50MJ/kg, specific heat (Cp) of 2.09kJ/kg C and density of 0.717kg/m3) is used as a fuel in this study. Temperature of the air reactor is selected as 1000 C. The heating of the system is carried out using air as fluidizing agent in the air and fuel reactors. When the temperatures in respective reactor are high enough to begin fuel operation, air is replaced with nitrogen in the fuel reactor, and after a few minutes nitrogen is switched to fuel. The incoming air to the air reactor can be preheated to the desired temperature. The air reactor temperature is kept at 1000 C. For time saving purposes, it is desirable to have some external means of control over the fuel reactor temperature at start-up. The reactivity and durability of metal oxides can be enhanced by doping with Al2O3, TiO2, MgO, or YSZ [3]. Int. J. Energy Res. 2013; 37:49–58 © 2011 John Wiley & Sons, Ltd. DOI: 10.1002/er
200-MW chemical looping combustion based thermal power plant
Fe2O3/Al2O3 (60% and 40% mass) has been used in the design calculations for this paper. In the Fe2O3/Fe3O4 metal oxide system, conversion of fuel to CO2 and H2O is complete [11]. Hence, Fe2O3 to Fe3O4 conversion is assumed in this paper (Equation (22)). The equations of the reaction phase in the fuel reactor and air reactor are CH4 þ 12 Fe2 O3 ! CO2 þ 2H2 O þ 8Fe3 O4 ΔH ¼ þ141:6kJ=mol O2 þ 4 Fe3 O4 ! 6 Fe2 O3
(22)
ΔH ¼ 481:04kJ=mol (23)
Fractional reduction rate for Fe-based metal oxide is assumed to be 4%/min. However, this value was reported to be 5.2%/min for Fe-based oxygen carrier [18]. Higher rate may not be possible as it is a function of many factors: manufacturing process, thermal conditions, and fuel characteristics. Reduction rate is assumed to be one-third of the oxidation rate. Assuming a 2% difference in the conversion between oxidizer and reducer as reported earlier [11], 98% conversion of metal oxides is assumed in the oxidizer and 96% in the reducer (Table I). A smaller conversion difference will ensure a large solid circulation rate and less temperature difference between the air and the fuel reactors. Air is assumed to be preheated to 374 C. Not all oxygen of air is assumed to react in the air reactor, and the flue gases from air reactor is assumed to have 4% oxygen in it. CH4 entered in the fuel reactor at ambient temperature. Variables considered for the analysis are bed diameter, bed height, entrained solid mass flow rate from the air to the fuel reactor, fluidization velocity, conversion difference, temperature, and conversion rate in the air and fuel reactors. A sensitivity analysis also presented here to predict the effect of change of one design variable on the other parameters.
5. RESULTS AND DISCUSSION The 200-MW CLC-based system burning CH4 using Fe2O3/Al2O3 oxygen carrier is presented in Figure 2 and detailed in Table II. The air reactor is a high velocity bed with a solid mass flow rate of 3780kg/s to the fuel reactor. The pressure drop is 11kPa and the heat of reaction is 233 kW corresponding to CH4 and iron oxygen carrier. Ambient air is preheated to 374 C before entering into the air reactor and the nitrogen gas, and some unreacted oxygen leaves the air reactor at 1000 C. On the other hand, the fuel reactor is a low velocity bed with a solid mass flow rate of 3763kg/s to the air reactor. The inlet gas velocity in the fuel reactor is in the range of 10Umf. Oxygen carriers entered in the fuel reactor at 1000 C. The pressure drop is 7kPa, and the heat of reaction is 33MW. Flue gas leaves the fuel reactor at about Int. J. Energy Res. 2013; 37:49–58 © 2011 John Wiley & Sons, Ltd. DOI: 10.1002/er
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980 C. The sensible heat is extracted from the air reactor and from the flue gases of the fuel reactor. Various heat recovery units with their heat transfer values are also presented in the Figure 2. The air reactor has a set temperature of 1000˚C. The exhaust gas from the air reactor consists of 4% (by volume) of O2 and 96 % (by volume) of N2. The bed particles have a residence time of 9.2s in the reactor. The air flow rate in the air reactor was maintained to ensure the circulation of bed material with the interconnected reactors’ system. The diameter of the air reactor is half of that of the fuel reactor so as to maintain the desired fluidization gas velocity in the fuel reactor. The fuel reactor has a temperature of 979˚C. Residence time of the bed material in the fuel reactor is more than that of the air reactor because of the endothermic nature of the reaction that requires more time to complete. The height of the air reactor is more as compare with the fuel reactor; this is in accordance with the regimes of fluidization to be maintained in the reactors. 5.1. Sensitivity analysis of design parameters The impact of various parameters, such as the conversion difference between air reactor and the fuel reactor (ΔX), temperature in the air reactor, and conversion rate of oxygen carrier in oxidation and reduction, on the design of CLC were analyzed. The analysis has been performed by assuming five sets of values for each variable described in previous texts and by keeping all other variables the same as that of the base case. The fluidization velocities in the air and fuel reactor are selected such that the air reactor maintains a fast fluidization regime, whereas the fuel reactor remains at a low-velocity bed. (a) Changing conversion difference between air and fuel reactors (ΔX). Five sets of conversion difference ΔX=0.02, 0.04, 0.06, 0.08, and 0.10, as quoted by Mattisson and Lyngfelt [11], have been assumed. Bed mass in the air reactor is independent of the conversion difference between the air and the fuel reactor; however, it decreases with increase in the ΔX for fuel reactor (Figure 3). The conversion rate in the air reactor has been fixed at 0.98, and in the fuel reactor, it is decreasing to accommodate this increase in the conversion difference between the two reactors. Decrease in the conversion rate in the fuel reactor decreases the amount of oxygen required and to accommodate this decrease in the oxygen requirement, the bed material in the fuel reactor is also decreasing. Pressure drop in the air reactor remains unchanged as bed mass is not changing here; however, for fuel reactor, it decreases as bed mass is decreasing with the increase in the conversion difference between air and fuel reactors. Entrained solid mass flow from air reactor to fuel reactor is also decreasing as we are considering reduced 53
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Temperature (°C) 400.0 Pressure (kPa) 101.32
Heat Extracted (MW) 40.36 Heat Extracted (MW) 20.98
Oxygen depleted air O2 4.0 N2 96.0
Flow rate (kg/s) 65.86 Temperature (°C) 1000.0 Sensible heat (MW) 71.55
Flue gas Flow rate (kg/s) 19.94 Temperature (°C) 979.7 Sensible heat (MW) 30.38
H2O 66.67 CO2 33.30 Temperature (°C) 400 Pressure (kPa) 101.32
MeO Diameter (m) 6.27 Bed Material (ton) 34.89 Bed height (m) 20.54 Residence time (s) 9.23 Fluidizing velocity (m/s) 9.6 Pressure drop (kPa) 11.08 Reaction heat (MW) 233.31 (Air Reactor)
Heat extracted (MW) 107.64
Flow rate (kg/s) 3779.23 Temperature (°C) 1000.00 Sensible heat (MW) 3316.27 O2 transferred (kg/s) 15.94
Diameter (m) 14.27 Bed Material (ton) 112.90 Bed height (m) 1.18 Residence Time (s) 30.0 Fluidizing velocity (m/s) 0.16 Pressure drop (kPa) 6.92 Reaction Heat (MW) -33.31 (Fuel Reactor)
Me Flow rate (kg/s) 3763.29 Temperature (°C) 979.00 Sensible heat (MW) 3231.16
Temperature (°C) 373.62 Sensible Heat (MW) 31.00
Flow rate (kg/s) 4.00 Temperature (°C) 25.39 Sensible heat (MW) 0.00 Pressure (kPa) 108.24
Temperature (°C) 25.48 Sensible heat (MW) 0.00 Pressure (kPa) 101.32
Temperature (°C) 450.0 Sensible Heat (MW) 31.19
Compressor power (MW) 0.041
Fuel Flow Flow rate (kg/s) 81.81 Temperature (°C) 25.00 Sensible heat (MW) 0.00 Pressure (kPa) 108.25 Air
flow
Compressor power (MW) 0.742 Temperature (°C) 25.00 Pressure (kPa) 101.32
Temperature (°C) 25.48 Pressure (kPa) 101.32
Temperature (°C) 25.00 Pressure (kPa) 101.32
Oxygen depleted air
Figure 2. Chemical-looping combustion system for 200 MW capacity with Fe2O3/Al2O3 as oxygen carrier and CH4 as fuel.
conversion rates in the fuel reactor for the sensitivity analysis. Figure 3 demonstrates the effect of the change of conversion difference between the air and fuel reactors (ΔX) on residence time, bed material, bed height, pressure drop, and entrained solid mass flow rate from the air to the fuel reactor at conversion rate in the air reactor=7%/min, conversion rate in the fuel reactor=2%/min, fluidization 54
velocity in the air reactor=9.6m/s and fluidization velocity in the fuel reactor=0.16m/s. (b) Changing conversion rates in the air and fuel reactors. Five different sets of conversion rates in the air and fuel reactors have been used here as quoted by Chandel et al. Int. J. Energy Res. 2013; 37:49–58 © 2011 John Wiley & Sons, Ltd. DOI: 10.1002/er
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Table II. Details of the chemical looping combustion system with CH4 for 200-MW power plant. Metal oxides Fuel Fluidized beds
Fe2O3/Al2O3 Methane (CH4) Air reactor
Diameter (m) Bed Height (m) Bed Mass (ton) Temperature ( C) Reaction Heat (MW) Pressure Drop (kPa)
6.27 20.54 34.89 1000 233.31 11.08 Fuel Reactor Diameter (m) 14.27 Bed height (m) 1.18 Bed mass (ton) 112.90 Temperature ( C) 979.70 Reaction heat (MW) 33.31 Pressure drop (kPa) 6.92 Air Reactor outlet O2 % (Vol) 4.0 N2 % (Vol) 96.0 Air reactor outlet flow rate (kg/s) 65.86 Air reactor inlet flow rate (kg/s) 81.81 Fuel Reactor outlet CO2 (%) 33.3 H2O (%) 66.7 N2 (%) 0 Flue gas flow rate (kg/s) 19.93 Fuel flow rate (kg/s) 3.999 Mass Transfer Oxygen transfer (kg/s) 15.94 Entrained solid mass flow from oxidizer (kg/s) 3779.22
[9] and Abad et al. [12] and presented in Table III. The amount of bed material and residence time of bed material in the air and fuel reactors are decreasing with increase in the conversion rate in the air and fuel reactors (Figure 4). This is in accordance with the increase in the reactivity of the metal oxides used. Similarly, bed height and, hence,
Table III. Assumed conversion rates in the air and fuel reactors used in this study for sensitivity analysis. Series Fractional conversion rate Fractional conversion rate number in air reactor (%/min) in fuel reactor ( %/min) 1 2 3 4 5
7 9 13 15 18
2 3 4 5 6
pressure drop in the air and fuel reactor decrease with the increase in the conversion rate as bed material in the reactors is decreasing. However, solid mass flow rates in between air reactor and fuel reactor are coming out to be independent of the conversion rates. Figure 4 demonstrates the effect of the change of conversion rate in the air and fuel reactors on the other design parameter like residence time, bed material, bed height, and pressure drop. For the fixed values of fluidization velocity in the air reactor=9.6m/s, fluidization velocity in the fuel reactor=0.16m/s, conversion difference between air and fuel reactor (ΔX)=0.02. (c) Changing temperature in the air reactor Five different sets of air reactor temperatures have been used here as quoted by Lyngfelt et al. [3] and are presented in Table IV. The height of the air reactor and the residence time of the bed material in the air reactor decreases with the increase in the temperature of the air reactor. An increase in the temperature increases the rate of reaction. Because of this increase in the activity of the metal oxides, the quantity of bed material in the air reactor and fuel reactors decreases and so do the heights of the air and fuel reactor (Figure 5). However, temperature can only be increased within the permissible limits, decided by the melting point of the metal oxide used, otherwise, there may be sintering of metal oxides and a reduction in the activity of the metal oxide and a reverse trend may occur.
Figure 3. Effect of change of conversion difference between air and fuel reactors on various design parameters. Int. J. Energy Res. 2013; 37:49–58 © 2011 John Wiley & Sons, Ltd. DOI: 10.1002/er
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Figure 4. Effect of change of conversion rate in air and fuel reactors on various design parameters.
Table IV. Assumed temperature in the air reactor. Series number
Temperature in the air reactor ( C)
1 2 3 4 5
700 800 900 950 1000
Figure 5 demonstrates the effect of change of the temperature in the air reactor on the other design parameter such as residence time, bed material, bed height, and pressure drop. For the fixed values of the conversion rate in the air reactor=7%/min, in the fuel reactor=2%/min, conversion difference between air and fuel reactor (ΔX)=0.02, fluidization velocity in the air reactor=9.6m/s and fluidization velocity in the fuel reactor =0.16m/s.
6. CONCLUSIONS Chemical looping combustion-based 200-MW power plant using a Fe2O3/Al2O3 oxygen carrier is designed for the
CH4 combustion. The present study demonstrates that the size of the CLC system for Fe-based carriers is very high because of the lower conversion rate of Fe-based carriers. Conversion of the oxygen carrier decides the size of the reactor. High conversion rate ensures smaller reactor size. The size of the reactor can be reduced using the oxygen carriers with higher conversion rate. An increase in the conversion rate in the air and fuel reactors decreases the bed mass in both reactors because of the increase in the reactivity of the metal oxide in the reactors. Also, increasing conversion difference in the air and fuel reactors reduces the bed mass in the fuel reactor. This is because of the decrease in the conversion rate in the fuel reactor, which is ultimately reducing the oxygen requirement and so does the amount of inventory in the reactor. As bed mass is decreasing in the reactor, the pressure drop also shows a decreasing trend in the particular reactor. On the other hand, a temperature rise in the air reactor decreases the amount of the bed material in the air and fuel reactors and height of the reactors, which is because of the increase in the rate of reaction in the air reactor. However, temperature in a reactor can only be increased based on the melting point of the metal used as the metal oxide.
Figure 5. Effect of change of air reactor temperature on various design parameters.
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NOMENCLATURE
Va,fuel
Aox
Xox Xred ΔX
Ared e Hg Hi Hr hbed, ox hbed, red g M Mactual, ox Mactual, Mfuel Mf .ox Mf
.red
MO2 mbed,ox mbed,red mfuel mo mair n
Pfuel ΔPox ΔPred q
Ro Rr
rn rox rred Sr Ta Tox Tred TS Umf uox ured V
Va,air
red
= cross sectional area of air reactor (oxidizer) (m2) = cross sectional area of fuel reactor (reducer) (m2) = voidage decide the bed height (dimensionless) = fluid-solids inter-phase heat transfer (Jm-3 s) = heating value of the fuel (MJkg-1) = heat of reaction (Jm-3 s) = bed height in the oxidizer (m) = bed height in the reducer (m) = acceleration due to gravity (ms-2) = number of solid phase (dimensionless) = actual molar mass of the carrier in its partially oxidized state (kgkmol-1) = actual molar mass of the carrier in its partially reduced state (kgkmol-1) = molar mass fuel (kgkmol-1) = molar mass of the fully oxidized carrier (kg kmol-1) = molar mass of the fully reduced carrier (kg kmol-1) = molar mass of oxygen (kgkmol-1) = bed mass in oxidizer (kg) = bed mass in reducer (kg) = fuel consumption (kgs-1) = oxygen transferred (kgs-1) = air mass flow rate (kgs-1) = total number of chemical species = power to be generated (MW) = pressure drop oxidizer (Pa) = pressure drop reducer (Pa) = conductive heat flux (Jm-2 s) = mass fraction of oxygen in the carrier (dimensionless) = recirculation ratio of the total flow of gas supplied to reducer and the fuel flow (dimensionless). = the rate of production of the nth chemical species (kg m-3 s) = fractional conversion rates in the oxidized (% sec-1) = fractional conversion rates in the reducer (% sec-1) = stoichiometric ratio for the reaction between fuel and oxygen (dimensionless) = ambient air temperature (K) = oxidizer temperature (K) = reducer temperature (K) = system temperature (K) = minimum fluidization velocity (ms-1) = fluidizing velocity in the oxidizer (ms-1) = fluidizing velocity in the reducer (ms-1) = velocity (ms-1) = specific volume of air at ambient temperature (m3 kg-1)
Int. J. Energy Res. 2013; 37:49–58 © 2011 John Wiley & Sons, Ltd. DOI: 10.1002/er
XO2 ;ex
= specific volume of fuel at ambient temperature (m3 kg-1) = degree of oxidation (conversion) (dimensionless) = conversion in the reducer (dimensionless) = conversion difference in oxidizer and reducer (dimensionless) = volume fraction of oxygen in the air from the oxidizer (dimensionless)
Subscript g ox red s
= = = =
Greek e l r
= volume fraction (dimensionless) = air ratio (dimensionless) = density of solid material (kgm-3)
gas phase oxidizer reducer solid phase
ACKNOWLEDGEMENT Authors are thankful to French Energy Company TOTAL.
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Int. J. Energy Res. 2013; 37:49–58 © 2011 John Wiley & Sons, Ltd. DOI: 10.1002/er
