CO2 Capture with Chemical Looping Combustion of Gaseous Fuels

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CO2 capture with chemical looping combustion of gaseous fuels: An Overview Jing Li, Hedong Zhang, Zuopeng Gao, Jie Fu, Wenya Ao, and Jianjun Dai Energy Fuels, Just Accepted Manuscript • DOI: 10.1021/acs.energyfuels.6b03204 • Publication Date (Web): 28 Feb 2017 Downloaded from http://pubs.acs.org on March 1, 2017

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CO2 capture with chemical looping combustion of gaseous fuels: An Overview Jing Lia,b, Hedong Zhanga, Zuopeng Gaoa, Jie Fua, Wenya Aoa, Jianjun Dai∗a a College of Chemical Engineering, Beijing University of Chemical Technology 15 Beisanhua East Road, Chaoyang District, Beijing, China 100029 b GreenGenTech Energy Inc., 18 Gardengate Way, Ottawa, ON K2H 5Z2 Canada ______________________________________________________________________________________

Abstract Nowadays the world consumes more and more natural gas (NG) and other fossil fuels (e.g. coal and crude oil). However, most of CO2 is released to atmosphere without capture for either NG combustion or other processes. Chemical-looping combustion (CLC) is a two-step combustion technology for power and heat generation with inherent CO2 capture, using either gaseous fuels or solid and liquid fuels. Previous review focused on CLC of solid fuels or CLC of all types of fuels without in-depth and specific discussions for gaseous fuel CLC systems. China is one of the largest NG, coal and crude oil consumers in the world and it is essential to develop an alternative technology to gaseous fuel (e.g. NG) combustion. This paper summarized recent research and development work on CLC using gaseous fuels, including the technological and economic assessment, types of oxygen carriers (OCs), reactor types, coke formation and OCs poisoning, efficiency and exergy analyses, model development based on a literature survey. The plant efficiency of NG-CLC can be up to 52-60% (LHV) including CO2 compression based on calculations and simulations, which is about 3-5% more efficient than a NGCC with CO2 capture. Ni-based materials have been widely developed and applied for NG because of its fast kinetics for methane conversion. CuO-Cu2O/Cu, Mn3O4-MnO and Fe2O3-Fe3O4 are typical OCs with high selectivity towards CO2 and H2O. The operating conditions are closely dependent on reactor configurations, hydrodynamics, mass and heat balances and characteristics of the OCs in the system. CLC and other CO2 capture technologies were also compared in the present study, which has rarely been investigated in previous review. From simulation and process analysis, a conceptual design of a NG-CLC power plant of thermal input 655 MWth was conducted to clarify its technological advantages and economic benefits compared to other power generation processes. The air reactor, fuel reactor and OCs do not impose significant economic barriers for scale-up and commercialization of CLC. Keywords: Chemical looping combustion; CO2 capture; fluidized beds; power generation. ∗

Corresponding author: Tel: 86-010-64452091; e-mail: [email protected], [email protected]

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Table of Contents CO2 capture with chemical looping combustion of gaseous fuels: An Overview ........................1 Abstract ..............................................................................................................................................1 1.

Introduction................................................................................................................................3

2.

CLC concept and competing CO2 capture technologies ........................................................7

3.

Process descriptions and characteristics of CLC with gaseous fuels ..................................10

3.1 Process description ................................................................................................................................. 10 3.2 CLC of gaseous fuels vs. CLC of solid and liquid fuels ........................................................................ 14 3.3 Reactor types and characteristics ........................................................................................................... 18 3.4 Equilibrium analysis for oxygen carriers ............................................................................................... 23 3.5 Efficiency and exergy analysis ............................................................................................................... 25 3.6 Design and scale up of CLC................................................................................................................... 31

4.

Oxygen carriers ........................................................................................................................35

4.1 Ni-based oxygen carriers........................................................................................................................ 36 4.2 Cu-based oxygen carriers ....................................................................................................................... 39 4.3 Fe-based oxygen carriers........................................................................................................................ 42 4.4 Mn-based oxygen carriers ...................................................................................................................... 48 4.5 Co-based oxygen carriers ....................................................................................................................... 49 4.6 Mixed oxide oxygen carriers .................................................................................................................. 50 4.7 Naturally occurring and low cost materials as OCs ............................................................................... 56 4.8 Coke formation and oxygen carriers poisoning ..................................................................................... 61

5.

Modeling development of CLC...............................................................................................65

5.1 Fluid dynamics models .......................................................................................................................... 66 5.2 Reaction kinetics model ......................................................................................................................... 69 5.3 Process simulations and mathematical models for CLC systems .......................................................... 77

6.

Comparison of COE and CO2 cost for different CO2 capture technologies.......................85

7.

Conclusions and future work ..................................................................................................88

Nomenclature: .................................................................................................................................91 Abbreviations: .................................................................................................................................94 References:.......................................................................................................................................97 Figure Captions .............................................................................................................................134

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1. Introduction Emission of greenhouse gases (in particular CO2, CH4 and N2O) is the main contributor to global warming with CO2 being the most prevalent of these gases. CO2 emissions resulting from human activity have led to an increase of the atmospheric CO2 concentration from a preindustrial level of 280 to 360-396 ppmv. Release of CO2 from fossil fuel combustion is the most important source of these emissions [1–4]. Post-combustion capture, pre-combustion decarbonization, and oxy-fuel combustion are three major routes to limit CO2 emissions (see Table 1). However, commercially available CO2 capture technologies are not cost effective and considerably increase the cost of electricity. Chemical looping combustion (CLC), initially proposed by Ritcher and Knoche [5], is an indirect two-step combustion technology with inherent separation of CO2 from N2 (see Figure 1). CLC is a precommercialization technology, but experimental tests and simulation studies on its integration with power plants indicate that it has the potential to be more efficient than all other CO2 capture technologies [6–27]. CLC usually consists of two reactors, a fuel reactor (FR) and air reactor (AR), with the oxygen carrier (OC) transferring oxygen from the air to the fuel. Its main advantage is the ability to inherently separate both CO2 and H2O from the N2 gas stream, while maintaining high efficiency of heat and power generation. After condensation of water, pure CO2, free of nitrogen, can be compressed for sequestration or enhanced oil recovery (EOR) [28–33]. Moreover, CLC minimizes NOx formation since both the AR and the FR operate at temperatures < 1200oC so that there is only possible NOx formation from fuel nitrogen [34]. The CLC concept can also be used for H2 production [35–42]. The fuel could be gaseous fuels (such as syngas, NG and propane), liquid fuels (e.g. diesel, bitumen and heavy oils) or solid fuels (e.g. coal, biomass and coke).

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Previous resreach on CLC has focused on the development of suitable OC materials. Transition metal oxides such as nickel (Ni), copper (Cu), cobalt (Co), iron (Fe) and manganese (Mn) are good candidates, given their favorable reductive and oxidative (redox) thermodynamic properties [17,20,43–55]. Besides high reactivity toward the fuel gas and air during hundreds (or even thousands) of reduction-oxidation cycles, the OCs must fulfill other characteristics such as high resistance to attrition, lack of agglomeration, and minimum carbon deposition when interacting with carbon-based fuels. Other environmental and economic aspects must also be considered when selecting the OC. High CLC performance requires good gas-solid contacting between OC and gases. The reactor design is very important since gas-solid contacting is strongly related to reactor configuration. Interconnected fluidized beds are one promising technology for CLC with advantages over competing configurations, such as good gas-solid contacting, uniform temperature, good mass and heat transfer, high reaction rate and throughput, flexible solids circulation and small footprint. The OC particles are circulated between AR and FR, transferring oxygen and heat from the AR to the FR. The solid circulation rate is dictated by the heat of reaction, the fuel gas flow rate and the oxygen transfer capacity of the OC [28,56,57]. Loop seals or other non-mechanical valves and devices can be used to prevent gas mixing between the two reactors. The volumetric gas flow in the AR is approximately 2.5 and 10 times larger than that of syngas and CH4, respectively, because there is a large amount of nitrogen in the air. To keep a reasonable size of the reactors, a high velocity riser has been proposed for the AR in a CLC plant when CH4 is used as the fuel gas [28]. CLC reactor design, modeling, mass and energy balance and hydrodynamics have been explored in previous study. Lyngfelt et al. [28] presented a design of an atmospheric CLC reactor including a high-velocity riser for the AR and a low-velocity fluidized bed (FB) for the FR. A 4 ACS Paragon Plus Environment

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conceptual design of a 10 kWth CLC operating at atmospheric pressure was provided by Kronberger et al.[57]. The NG-CLC process was demonstrated in two continuous operation CLC units of 10kWth thermal input during 100 h operation using a Ni-based OC [58–60] and a Cubased OC [61–63]. Whereas Ryu et al. [64] presented the results of a 3.5 h continuous run in a 50 kWth CLC system. Wolf et al. [65] assessed a 800MWth CLC system consisting of two interconnected pressurized FBs with NG as the fuel. Johansson et al. [66] analyzed the gas leakage between the reactors, and Kronberger et al. [67] and Xu [68] used a cold model to find critical design parameters of the CLC system. Fernandez and Alarcon [69] anazlyed a process scheme based on fixed-bed reactors and presented it as a possible alternative to carrying out the CLC of NG at high pressure with ilmenite as the OC. The operation at high pressure permits the use of highly efficient power cycles. However, complex heat management strategies and switching valves able to function at high temperatures are required. A conceptual design of 500MWth capacity using NG as the fuel and ilmenite as the OC has been provided. Spallina et al. [70] demonstrated the design and operation strategies of dynamically operated packed-bed reactors of a CLC system included in an integrated gasification combined cycle (IGCC) for electric power generation. The CLC reactors employed ilmenite as the OC and operated sequentially across the following phases: oxidation, purge, reduction and heat removal. The results indicated that 14–16 units with 5.5 m of internal diameter and 11 m of length are required for continuous operation of a 350–400 MWe coal-fired power plant. Both thermodynamic analysis and Aspen Plus® simulations have been performed for CLC systems with NG and syngas as the fuels. An exergy analysis conducted by Anheden and Syedberg [71] indicated that when a CLC system with a Fe-based OC is used to a retrofit an IGCC (Integrated Gasification Combine Cycle) plant, an increase in exergetic efficiency can be achieved. However, the energy for CO2 compression was not considered. The Aspen Plus® 5 ACS Paragon Plus Environment

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simulation showed that the integrated gasification-CLC (IG-CLC) system has the potential to achieve 43.2% (LHV) efficiency for electricity generation with 99% of the CO2 captured [72]. Several process configurations have been proposed that integrate CLC reactors with combined cycle (CC) power plants utilizing NG and synthesis gas fuels. Wolf et al. [73] reported a thermal efficiency as high as 52-53% in a CLC combined cycle (CLC-CC) plant operating at 1200ºC and 13 bar in the AR, which represents a 3-5% more efficient process than a natural gas combined cycle (NGCC) system with CO2 capture, and about 3-5% higher than that of post-, oxy-fuel or pre-combustion CO2 capture methods [74–76]. The efficiency would be slightly lower in an atmospheric CLC steam cycle (CLC-SC). Preliminary economic assessments have suggested that CLC is a promising combustion process with high efficiency and low cost CO2 capture. However, commercialization of CLC mainly depends on the availability of excellent OCs [43,49,50,73,77– 80]. Almost all experimental data now available are from specific reactor types with specific OCs and operating time may not be sufficient to ensure the quality of the OCs and the long-term performance of the CLC system over many cycles. Design optimization, continuous operation, detailed engineering and cost analysis at larger scales are still needed. However, the technical risk of the CLC system is low because all CLC processes were carried out in standard process equipment and used elements of proven technology [8,28,81]. Nowadays the world consumes large amounts of NG and other fossile fuels (e.g. coal and crude oil). However, most of CO2 is released to atmosphere without capture for either NG combustion or other prococesses. The objective of this work is to present recent research and development work on CO2 capture using gaseous fuel CLC system from a technological and economics point of view, including different types of OCs, different reactor types, conversion, reaction rates, gas yields, reaction kinetics, solids circulation rates, bed inventory, hydrodynamics and performance of the CLC system. A brief comparison is provided of CLC and 6 ACS Paragon Plus Environment

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other CO2 capture technologies. The process parameters are closely related to each other, and the operating conditions depend heavily on the reactivity of the OCs. The solids circulation rate is determined by the hydrodynamics, mass and heat balances and characteristics of the OCs in the system. From simulation and process analysis, a conceptual design of a NG-CLC power plant of thermal input 655 MWth is outlined in order to clarify the technical advantages and economic benefits from a CLC process compared to other CO2 capture technologies.

2. CLC concept and competing CO2 capture technologies Ritcher and Knoche initially proposed CLC concept in 1980s [5], which is based on the transport of oxygen from air to the fuel by means of an OC. One of the most important aspects for CLC concept is inherent CO2 capture. Hence, the tradional CO2 capture technologies are briefly introduced in this section for comparison with CLC, showing advantages of the CLC system over other technologies. Various CO2 capture methods have been used to remove CO2 from impure NG, power plants, food processing and chemical industries (Table 1 and Table 2). A variety of methods is used to separate CO2 from gas mixtures during production of hydrogen for petroleum refining, ammonia production and in other industries [2,31,81–87]. Selection of a CO2 capture technology depends on many factors, e.g. partial pressure of CO2, extent of CO2 recovery required, impurities in the gas stream (e.g. particulate and acid gases), purity requirements of the CO2 product, capital and operating costs, and cost of additives to overcome fouling and corrosion. Nowadays the most commonly used CO2 capture methods include absorption by physical and/or chemical solvents, adsorption (e.g. Pressure Swing Adsorption-PSA), membrane, cryogenic and cyclic calcium looping technologies.

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Conventional CO2 capture technologies for large scale power plants (e.g. post-combustion, pre-combustion and oxy-fuel combustion) are energy intensive resulting in a significant decrease of the overall efficiency, resulting in a price increase of the produced electricity. Both pre- and post- combustion capture are based on separation by membrane or other costly conventional separation technologies of high energy consumptions. Pre-combustion produces CO2 rich gas stream, leading to easier separation of CO2 and a slightly lower operation cost. However, its cost is still high. Oxy-combustion would seem a good alternative, given that high-concentration CO2 is obtained in the gases. However, the oxygen required for the process must be produced in an air-separation unit, also highly energy-demanding [88]. CLC reaction occurs with a lower irreversibility compared to a conventional combustion, leading to a somewhat higher overall thermal efficiency [88]. It is estimated that CO2 capture may lead to an energy efficiency penalty that can reach 1015% points depending on the technology adopted [89–91]. Romano [92] and Hawthorne et al. [93] have done thermodynamic analysis and simulation using the carbonator for CO2 capture and the calciner for sorbent (i.e. CaO) regeneration, a net efficiency of 37.4% (LHV) was predicted for the selected reference case, with 97% CO2 capture from his simulation [92]. Dean et al. [82] reported the current status of development for cyclic calcium looping for CO2 capture from power generation, cement manufacture and hydrogen production. The sorbents in this case are derived from cheap, abundant and environmentally-benign limestone and dolomite, and the process impose a relatively small efficiency penalty on the power/industrial process (i.e., estimated to be 6–8%, compared to 9.5–12.5% from amine-based post-combustion capture). CLC was proposed to have a relatively high energetic efficiency in integrated power generation with inherent CO2 separation [94–96]. Alvaro et al. [88] carried out simulations to evaluate the energetic efficiency of the CLC based power plant under diverse working conditions. A 8 ACS Paragon Plus Environment

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comparison of a conventional integrated gasification power plant with pre-combustion capture of CO2 has been made. Two different synthesis gas compositions have been tried to check its influence on the results. The energy saved in carbon capture is found to be significant and even notable, inducing an improvement of the overall power plant thermal efficiency of around 7% in some cases. For CLC at atmospheric pressure, a steam cycle could achieve about 40-42% (LHV) efficiency when using gaseous or solid fuels including energy consumptions for CO2 compression before transport and sequestration [97,98]. CLC combined cycle is another option, but CLC reactor should operate at higher pressure and temperature to increase energy efficiency. A pressure in a CLC system above 1.3 MPa is not recommended because theoretical calculations show little effect beyond this pressure on the overall efficiency [73,99]. The CO2 turbine downstream the reduction reactor was reported no substantial impact on the energetic efficiency of the process, mainly due to the higher compression work needed to re-pressurize the CO2 stream to sequestration conditions after decompression by the CO2 turbine. However, there is controversy in CO2 turbine application in CLC processes, the compression of CO2 was reported to reduce the efficiency by about 2% based on previous study [73]. CO2 as a commodity can be used in firefighting, food industry, fish farms, agricultural greenhouse and other chemical industries. However, the overall demand for CO2 is very small compared to the total CO2 emitted annually, and sequestration of CO2 is therefore essential, with considerable efforts being focused on enhance oil recovery (EOR), sequestration in depleted oil and gas reservoirs, mineral carbonation, and saline aquifers sequestration [32,100–102]. For EOR, the market price of CO2 varies widely, depending on both the price of crude oil and the amount of CO2 required to produce a barrel (Bbl) of oil. The cost of CO2 avoided ($/tonne) is calculated from the following equation (see Table 3):

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C avoided =

COEcapture − COE reference ( M CO2 ) reference − ( M CO2 ) capture

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(1)

The cost of CO2 avoided can be used to compare the ecomonics for different CO2 capture technologies. More information is provided in Section 6. For CLC, thermo-gravimetric analyzers (TGA), packed/fixed beds, moving beds and FBs have all been investigated for different OCs, determining reaction kinetics, conversion, yield and longevity (see Section 1). Reaction kinetics models have been developed for both reduction and oxidation. Solids circulation rate, bed inventory, hydrodynamic performance of the CLC system have been studied for FB configurations. Reactor design, modeling and hydrodynamics still require investigation. CLC has the potential for achieving efficient and low-cost CO2 capture. Scale-up to commercial CLC systems is promising, and mainly depends on the properties of the OCs prepared by different methods.

3. Process descriptions and characteristics of CLC with gaseous fuels 3.1 Process description Previous works provided the basic foundations of CLC and its potential application to gas turbine or combined cycle (CC) power generation. Further developments consider the use of alternative fuels for the chemical looping combustor [103]. A coal-direct chemical looping (CDCL) process has been recently patented. Such a system converts pulverized coal feedstock to fuel in one integrated system without additional gasification, and allows both electricity and/or hydrogen production [104–106]. The chemical looping concept has been extended to consider reforming (Chemical Looping Reforming, CLR) and gasification (Chemical Looping Gasification, CLG), where complete oxidation of the fuel is prevented by using low air to fuel ratio [107]. Further, the simple CLC configuration has been extended to include a third reactor 10 ACS Paragon Plus Environment

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that favors the simultaneous generation of hydrogen [108]. The oxidative coupling of methane (OCM) is an attractive alternative concept for ethylene production from methane based feedstocks using chemical looping concept [109]. The basic principle of OCM is the activation of methane by a catalyst material (e.g. Na2WO4/Mn/SiO2) which leads to methyl radical formation by C-H bond cleavage. The methyl radicals can couple to ethane in the gas phase, close to the catalyst surface. More information is available in [110]. In a typical CLC system, the OC particles circulate between two reactors present in the process and provide two-step combustion of the fuel. The process of CLC avoids direct contact between fuel and air, and achieves inherent separation of CO2 from N2 [111,112](see Figure 2). As mentioned above, CLC primarily consists of two main reactors: an air reactor (AR) and a fuel reactor (FR). In the AR, OC particles are oxidized by oxygen from air as shown in Reactions 2-6. The oxidized form of the OC is then transported to the FR, transferring heat and releasing oxygen in the FR without any direct efficiency loss.. If the fuel is in gaseous form, no additional steam and/or CO2 are needed for the FR. The gaseous fuel then reacts with oxygen available in the OC particles. In the FR, the OC is reduced. The reduced metal oxide particles are returned to the AR, where they are again oxidized by air. In CLC, the total heat evolved from the combined oxidation and reduction remain the same as in conventional combustion which is in accordance with the Hess's law.i.e the total enthalpy change during the complete course of a chemical reaction is the same whether the reaction is made in one step or in several steps. The product gas from FR and the flue gas from AR can both be used for heat and/or power generation with almost pure CO2 stream produced after condensation of water. CLC realizes inherent separation of both CO2 and H2O from flue gases, and therefore, the energy penalty for the capture of CO2 is minimized. Furthermore, very low NOx emissions have been reported due to the low combustion temperature and flameless 11 ACS Paragon Plus Environment

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combustion [113]. NOx formation usually occurs well above 1200 °C, which is hardly the highest temperature attained in CLC [7]. The main reactions in the FR (with CH4 and H2 as fuels) and AR are ( 2n + m ) Me x O y + C n H 2 m → (2n + m ) Me x O y −1 + nCO2 + mH 2O

(2)

Me x O y −1 + 1 / 2O2 → Me x O y

(3)

where MexO y denotes a metal oxide and Me x O y −1 represents the reduced form of the metal oxide. The oxidation of OC is exothermic, whereas the reaction in the FR may be either endothermic or exothermic, depending on the OC and the fuel types. The total amount of heat evolved from these reactions is the same as for normal combustion of the fuel. The following reactions occur in the AR and FR for CLC with NG as fuel and Ni-based metal oxide as the OC: (O2 + 3.76 N 2 ) + 2( Ni + Al 2 O3 ) → 2( NiO + Al 2 O3 ) + 3.76 N 2

(4)

2( NiO + Al 2 O3 ) + 0.5CH 4 → 0.5CO2 + H 2 O + 2( Ni + Al 2 O3 )

(5)

The net reaction combining these two reactions is

(O2 + 3.76 N 2 ) + 0.5CH 4 → 0.5CO2 + H 2O + 3.76 N 2

(6)

A makeup flow of new OCs is required to compensate for the natural decay of activity and solid losses due to attrition/fragmentation during operation. For simplicity, the makeup flow rate can be neglected since it should be small relative to the circulation flow of solids. The oxygen balance in the AR and FR is then (7)

FO2 ∆X O2 = dF f ∆X f

The circulation rate, expressed as mass flow of completely oxidized OCs, mainly depends on the type of OC and the fuels used, as well as on the solids conversion difference in the AR and FR

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o

m oc = 0.001br M NiO Ff ∆X f /(x NiO ∆X s ,FR )

(8)

Although Equation 8 gives the circulation rate necessary to satisfy the oxygen mass balance in the CLC system, it is also necessary to consider other aspects related to the hydrodynamic behavior and heat balance in the system, addressed in the next sections. The flow rates and diameters of the two reactors must be such that the reduction of the oxidized metal oxides (e.g. NiO) balances the oxidation of reduced metal oxides or metal (e.g. Ni). This will occur if the flow rates correspond to 2 moles of O2 for every mole of CH4 per unit time. As a result, assuming ideal gas behavior, we can have 0.21(πD12 / 4)(U1 / T1 ) = 2(πD22 / 4)(U 2 / T2 )

(9)

where D denotes reactor diameter, U superficial velocity, T the temperature, while the subscripts 1 and 2 refer to the air and fuel reactors, respectively. The corresponding diameter ratio is

D2 / D1 = 0.21U 1T2 / 2U 2T1

(10)

For oxidation of Ni, the heat release is 17% greater than for conventional combustion of NG. This leads to a drop in temperature in the FR since the total heat of reaction in the AR and FR is the same as for conventional combustion. Most OCs give such a temperature fall. For a CLC process, heat input is essential for start-up of the AR because the heat release only occurs when the oxidation occurs. The temperature in the air and fuel reactors is related to the fuel type, OC type and operating conditions (e.g. gas velocity, solids circulation rate and bed mass), as well as the process control strategy. Both reaction heat and preheating of air and CH4 need to be considered when carrying out the energy balance calculations. In general, the energy balance for the AR and FR can be written

∑F

in

input

.h + QOX =

∑F

out

.h + Q remove, AR

(11)

output

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∑F

in

input

.h =

∑F

out

.h + Qre + Qremove, FR

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(12)

output

More information abut mass and heat balance of CLC is available in previous study [68].

3.2 CLC of gaseous fuels vs. CLC of solid and liquid fuels CLC can use fuels in the gas, liquid or solid forms to achieve indirect combustion. There were extensive investigations for different types of fuels using different types of reactors and OCs in the past decades. However, previous study mainly focused on specific reactor types and configurations, specific OC types and properties. Hence, these results are not comparable and may not be applicable to large scale commercial CLC processes. A 300W CLC pilot plant of continuous operation with NG and syngas as fuel was reported [27,114]. It showed that unburned methane was detected, and CO and H2 were present at low concentrations in the exit flue gases when NG was used as the fuel. However, the combustion of syngas was complete for all experimental conditions with no CO or H2 present. Garcia-Labiano et al. [30] used syngas from an IGCC (integrated gasification combined cycle) power plant as fuel for the CLC study. Tian et al. [31] employed the simulated syngas as fuel for the CLC research, which showed stable reactivity for production of CO2 from fuel gas at 800 and 900oC and full consumption of hydrogen during the reaction. Foreroetal. [35] used CH4 and H2S as the fuel. The influence of H2S concentration on the product gas distribution, combustion efficiency, sulphur splitting between the FR and the AR, and agglomeration tendency were investigated. Iliutaetal. [36] used CH4 as fuel in two different reactors: one was a fixed bed micro-reactor and the other one was a fluidized be d reactor. In the FB reactor, complete conversion of the fuel was achieved within several minutes. However, in the fixed bed micro-reactor, the fuel conversion tends to decrease after some time due to imperfect gas-solid contact and slow reaction rate of fuel with the OCs. Johansson et al. [37] demonstrated CLC of methane as gaseous fuel with Nibased and Fe-based OCs. Kol- bitsch et al. [38] used NG, which mainly consisted of CH4 with 14 ACS Paragon Plus Environment

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only traces of other gas species (C2H6, other CxHy, CO2 and NO2), as fuel in their CLC study. The CLC experiments were carried out mostly with gaseous fuels. However, with advance in time, the growing interest on CLC of solid and liquid fuels has increased. Lots of different OCs have been developed using different metal oxides, support materials and production methods. Four different OCs were tested with focus on reactor temperature, total solids inventory, specific fuel reactor inventory, solids circulation rate and air to fuel ratio to investigate the interchangeability of OCs in a 120 kWth pilot plant [15]. Different reactor configurations and potentially limiting factors for each type of OC were identified. Changing OCs in an existing chemical looping plant seems to be possible without big changes of the reactor configuration. Previous studies of low-cost OCs for use with solid fuels covered iron ore [115–117], manganese ore [118], ilmenite, industrial waste materials [41,119], comparisons of materials of different sources were also investigated [120,121]. Aspen Plus® simulation indicated that the energy efficiency of the CDCL (coal direct chemical looping) process could exceed 80% (HHV) for H2 production and >50% for electricity generation with 100% of the CO2 captured [81]. Tests for a coal direct chemical looping (CDCL) process have been carried out in a 2.5 kWth bench scale moving bed unit at Ohio State University. Different feedstocks (e.g. simulated coal volatiles, lignite coal char, bituminous coal char, and anthracite coal) have been tested. It was found that coal/coal char conversion as high as 95.5% has been obtained. The CO2 concentration in the flue gas was > 97% (dry basis) in all cases. In addition, the reactivity of the OC was maintained after three redox cycles [55]. OC research for solid fuels has focused mainly on oxides of Ni, Fe, Mn and Cu. Nickel oxide is expensive and easily deactivated by sulfur, being less suitable for the solid fuels than iron and manganese oxides. Metal oxides may also be combined forming new compounds with new properties although these materials have not yet been tested successfully in operation with exception of calcium manganates[122,123]. Ilmenite 15 ACS Paragon Plus Environment

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(FeTiO3), a naturally occurring mineral and combined oxide, is often used with solid fuels due to its low-cost, reasonably high reactivity towards syngas and good fluidization behaviour [124– 128]. Mixed oxides build on synergies of mixing OC materials with different properties, for example, addition of limestone to ilmenite in CLC using solid fuels [129,130]. CDCL doesn’t need any gasifier, and there is no corresponding oxygen requirement.The solid fuel is mixed with the OC in the FR. The OC reacts with the in-situ gasification products formed inside the FR. The FR is fluidized by H2O and/or CO2. Chemical-looping with oxygen uncoupling (CLOU) combustion systems have the potential to assist in the capture of CO2 from power plants and have been investigated in previous studies [8,55,131–137].

CLOU processes require OC

materials to be able to release oxygen in the FR and to regenerate by re-oxidation in oxygen-rich atmosphere in the AR at elevated temperature. Hence, the OC is segregated to liberate gaseous oxygen in the FR to promote the combustion of the fuel [137,138]. This process has recently received great interests to overcome the slow char gasification step in the CDCL process. Copper oxide is a CLOU material although it has a higher cost than iron and manganese oxides. The largest experimental setup (1MWth) of CDCL has been tested with ilmenite as the OC. However, the plant didn’t demonstrate autothermal operations. The oxygen transfer capacity of ilmenite is low, which is the principle impediment to commercializing a process with this ore (huge solid inventory) [139]. CLC of liquid fuels used non-sulphurous and sulphurous kerosene, as well as heavy fuel oils and heavy vacuum residues generally produced during refining of crude oil, as the fuel [140]. Bitumen and asphalt, domestic fuel oil and waste cooking oil were reported in the past years for CLC study, using either packed bed or FB reactor [141–143]. A continuously operating reactor of capacity 300W has been designed, constructed, and success- fully operated using non-

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sulphurous and sulphurous kerosene as fuel. The conversion of fuel carbon to CO2 was as high as 99% [140]. Very limited studies have been reported so far on CLC of liquid fuels. CLC of gaseous fuels means the primary fuel is in the gas form (e.g. NG, syngas and refinery gas, etc.), which reacts directly with the OCs and the fundamental chemistry behind it is the same as discussed above. If the gasification of solid fuels (e.g.coal, petcoke, solid wastes or biomass) is first carried out and the produced syngas is introduced in the CLC, it is usually termed as the Syngas-CLC. In the CLC of gaseous fuels, the reaction rate is dependent on gassolid contact, interphase mass and heat transfer, OC particle properties and reaction kinetics. For the fluidzed bed CLC system, hydrodynamics is crucial for both oxidation and reduction reactions. The solid phase in the CLC of gaseous fuels is mainly OC particles. Whereas, the solids in the CLC of solid fuels include OC particles, ash and some carbonaceous particles (e.g. char), there are also complicated interactions between gas/ vapor and solids, which add complexity to the system analysis. For CLC of gaseous fuels, the gaseous fuel is not only the fuel combusted, but also a fluidization gas for FB configurations. There is no ash in the reactor, different from the case of CDCL. CLC of solid fuels involves char gasification, as the ratelimiting step in the solid fuels conversion, different from CLC of gaseous fuels [144,145]. The feeding systems of liquid and solid fuels to the FR is also completely different from CLC of gaseous fuels. CLC of gaseous fuels are applicable for almost any type of reactor. However, OC properties and reactor performance still need further research. CO2 concentration in the flue gas from NG combustion is in the range of 3-8v%, lower than that from coal combustion (~15v%), which indicates that post-combustion CO2 capture for NG combustion is even more difficult compared to coal combustion processes. Hence, CLC of NG is very promising as an alternative to NG combustion with inherent CO2 capture. Previous study and review discussed CLC of different types of fuels together, and ignored the specific 17 ACS Paragon Plus Environment

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characteristics for CLC of different fuels. The present study overcame this problem and mainly focused on CLC of gaseous fuels to try to clarify the important aspects for CLC burning combustible gas. However, experience in CLC of solid and liquid fuels is useful for development of CLC with gaseous fuels with proper modification and caution.

3.3 Reactor types and characteristics Various reactor configurations were investigated for CLC systems (e.g. interconnected moving beds and interconnected fluidized beds), such as coated monolith reactors [146], packed beds [147] or FBs [148]. Interconnected fluidized-beds are usually considered the most promising CLC concept , especially for CLC with gaseous fuels. However, other reactor types have also been tested and analyzed. Hydrodynamic study provides some key information for the design of reactors, especially for selection of the flow regime, system configuration and operating conditions of FB reactors.

For FB applications, minimum fluidization velocity, minimum bubbling velocity, minimum slugging velocity and terminal velocity should been calculated and analyzed based on the properties of OCs (e.g. particle density, size) and gases. There are several equations used to calculate minimum fluidization velocity. One of the commonly used equations is 2

Re mf = C1 + C 2 Ar − C1

(13)

where C1 =27.2 and C2 =0.0408 based on Grace [149]. Terminal velocity can be calculated from equations recommended by DallaValle [150]. The gas velocity in the AR provides the driving force for the circulation of particles [28]. As mentioned above, the flow regimes in the CLC system could be chosen as fast fluidization for the AR and bubbling fluidization for the FR in order to obtain a reasonable reactor size and performance. Hydrodynamic models developed for fast FBs, bubbling fluidized beds (BFB) and other fluidization regimes [149,151–154] can be used to simulate CLC performance. For a CLC 18 ACS Paragon Plus Environment

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system, the superficial gas velocity in the AR is normally in the (5~20) U t range, whereas the superficial gas velocity in the FR is in the range of (5~30) U mf , and the loop-seal and downcomer superficial gas velocities are of order of magnitude ~ U mf . For a packed bed of particles, if the fluidization velocity is less than the minimum fluidization velocity, the particles remain fixed. Pressure profiles and relationships between superficial gas velocities, total solids inventory and net solids circulation rate for FBs have been reported [57,67,155–158]. However, most hydrodynamic studies have been conducted in a narrow range of operating conditions in coldflow models consisting of relatively small-scale fluidization columns. A detailed operating mapping of a CLC system consisting of one AR operating in the fast fluidization regime and one FR in the bubbling fluidization regime has been prepared at UBC, including superficial gas velocities, solids circulation flux, pressure profile, solids hold-up, gas leakage, aeration gas velocity and their relationships [68] (See Figure 2). Gas leakage between reactors can adversely affect CO2 capture efficiency and CO2 purity. The pressure in the two reactors should be approximately equal in order to minimize gas leakage between them. Johansson et al. [66], Kronberger et al. [57,67], Min [68] and Saayman et al. [159] have measured gas leakage in cold-flow models and investigated the effect of operating conditions. Gas leakage is generally not a major problem during the operation of systems with loop-seals since it can be avoided by using steam or other gases. CFD modeling for CLC study is presented in Section 4. One of the important features that an OC must possess is to have high resistance to attrition, especially for two interconnected fluidized bed CLC system. In a continuously operated CLC unit for gaseous fuels, several particle characteristics related to attrition resistance were analysed on 23 natural and synthetic OCs, prepared by different methods based on different metal oxides 19 ACS Paragon Plus Environment

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[160]. The OCs tested were based on Mn, Ni, Cu and Fe, which were mainly prepared by the incipient impregnation methods. However, the Mn based OCs were prepared by spray-drying, and there is also a natural Fe-based ore as the OC. The main supports are γ-Al2O3, α-Al2O3, CaAl2O4, MgAl2O4 and ZrO2. The particle size used was 100-500µm, depending on different OCs. Particle crushing strength and Air Jet Index (AJI) were determined for the fresh materials, as well as the attrition rate and the corresponding particle lifetime during multi-cycle redox reactions. A comparison was made of the different methods used to evaluate attrition behaviour. Schwebeletal. [85] tested a parallel arrangement of a moving bed FR and a FB AR. This arrangement reduced the power requirement for fluidization, avoided fuel segregation with less char at the reactor exit. Thon et al. reported a 25 kW CLC system of coupled FBs with solid fuels and ilmentile as the OC [86]. The CLC system consisted of a circulating fluidized bed (CFB) AR coupled with a two-stage BFB FR. Carbon capture efficiency of ~90% was achieved in this system. However, the presence of unconverted combustible gas in the FR was reported. Penthoretal. [87] reported a coupled operation of two dual fluidized bed (DFB) pilot plants.The first one is a 100kW steam gasifier, the syngas produced was introduced to a 120kW CLC of continuous operation. The coupling of the two was done by a hot product gas fan. Chiu et al. [57] used an annular dual-tube moving bed reactor to study the CLC process with polyurethane and polypropylene as the fuel and Fe-based metal oxides as the OC. Strohle et al. reported a pilot plant CLC of capacity 1MWth based on two interconnected CFB reactor with hard coal as the fuel and ilmentite as the OC [88]. Abad etal. [89] developed interconnected CFB CLC unit where they achieved a carbon capture of 88% at 991oC and a total oxygen demand of 8.5% with a solid inventory in the FR of 470kg/MWth. FBs can be employed for CLC with either gaseous fuels or solid and liquid fuels due to the fuel flexibility and excellent gas and solid contacting.

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An alternative to interconnected fluidized bed CLC configurations is packed bed design [161–164]. The packed bed reactor shows high tunability in heat management and reactor configuration, because the air and fuel reactors are completely independent [165]. In a packed bed reactor configuration for CLC, the OC is statically contained and alternatively exposed to oxidizing and reducing gases. Solid circulation is not required for this case. Hence, high-pressure operation can be achieved without much difficulty. The packed bed reactor is compact and offers potential for better utilization of the OC and low capital cost. Moreover, a two packed-bed in series could help the CLC system to obtain the desired temperature rise (450-1200oC) for hot air and avoid fuel slip, making the OC selection become less critical [166,167]. The packed bed CLC system has been tested at up to 7 bar using syngas as fuel and ilmenite as OC [168]. It has been found that wet syngas has to be used as reducing fuel to avoid excessive carbon deposition that reduces the CO2 capture efficiency of the system. A maximum temperature rise of 340oC has been obtained in the reactor, which is in good agreement with the theoretical prediction when accounting for the heat losses. The main disadvantages of the packed bed reactor are its dynamic and batch operation, and the necessity of high temperature valves. It has been shown that the maximum temperature is independent of the gas mass flow rate or oxidation kinetics of the OC in a packed-bed reactor CLC system, which results in high flexibility to changes in production capacity, with little disturbance due to changes in reaction kinetics [147]. However, the overall process efficiency of FB CLC and packed bed CLC systems using Ni-based oxides as OC and syngas as fuel is similar [169]. Dahl et al. [170,171] proposed a rotating reactor system, with the OC materials being rotated between different gas streams flowing radially outwards through the metal oxide bed, avoiding the periodic switching valves operating with gases [172,173]. However, pressure drop and temperature control when the processes are scaled up will be challenging because of the 21 ACS Paragon Plus Environment

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large fraction of OCs in the fixed bed reactors (i.e., low bed void fraction) and the high pressure conditions [174,175]. Although an inert gas is introduced to avoid mixing of the two reacting gases, gas interchange between the AR and FR is a challenge and may be unavoidable based on previous research. At IFP-France, a 10 kWth unit with three interconnected bubbling beds (one FR and two AR) has been designed and constructed. Pneumatic L-valves were used to control the solid circulation rates independently of the gas flow and solids inventory in each reactor [7,176]. Ryu et al. [177] reported a 50 kWth CLC unit usig solid injection nozzles inside each reactor to control the solids flow. Aeration gas fed to the two loop seals was used to adjust the solids circulation rate for a CLC unit consisting of two interconnected fluidized beds at UBC [68], with the AR operating in the fast fluidization flow regime, while the FR is a BFB). A prototype countercurrent moving bed bench scale unit was initially studied using iron oxide to convert syngas and methane [35,178]. Zeng et al. employed a moving bed reducer with iron oxide as the OCs and NG as the fuel. A bench-scale countercurrent moving bed demonstration unit was built and tested, and the reducer operating conditions were studied using thermodynamic models. The thermodynamic equilibrium model also established a system baseline performance. An optimized set of operating criteria were determined from the Ellingham Diagram analysis. A thermodynamic criterion for selecting iron oxide based OC material and designing the reaction system was developed. Solids analysis determined the iron oxide conversion and verified the lack of carbon deposition or iron carbide formation. An optimal set of operating conditions (e.g solids to gas ratio and temperature) is identified for further testing on a larger scale [178]. The Korean Institute of Energy Research has also considered using a moving bed for the Three Reactor Chemical Looping (TRCL) system [179]. Countercurrent moving beds have also been studied by Wu et al. [180] and Chiu [181] at The 22 ACS Paragon Plus Environment

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National Taiwan University of Science and Technology to determine the proper conditions for complete methane conversion in a moving bed consisting of iron and aluminum oxides. They have tested gasified polyurethane and polypropylene and have shown the feasibility of treating plastic wastes in a moving bed reducer [178,181]. The moving bed reactor with a counter-flow pattern can be used to enhance the utilization of iron-based OCs and implement the coproduction of electricity and hydrogen via step oxidation of reduced OCs [166,167,182–185]. In general, smaller OC particles lead to larger pressure drop, but larger particles may lower the utilization of the OCs due to the intra particle diffusion limitation. Moreover, a highly exothermic reaction (e.g., the oxidation of the reduced oxides) in the fixed bed system may result in hot spots and thermal runaway, which is detrimental for the stability of the OC and for the safety of the operation [166,167,182–185]. The size and geometry of OCs are one of the key factors to determine the efficiency of a large-scale CLC system in fixed bed reactors, because they strongly affect the dynamic conditions of the gas–solid reactions (e.g. the intra particle mass transfer limitation for reactants, the pressure drop and the flow distribution) [165]. The monolithic OCs (4.5 cm long, 6.0 cm in diameter, square cell size of 2.0 mm, and wall thickness of 0.9 mm) used in the CLC of methane show high activity in a high gas hourly space velocity (GHSV, 6000/h). This can be attributed to the special geometric structure and layered microstructure [165]. For the large scale utilization, multiple points of the gas feed along the reactor may be employed for the monolith to improve the gas-solid reaction. The powder and monolithic OCs show similar reduction behaviors either in hydrogen or in methane atmosphere [165].

3.4 Equilibrium analysis for oxygen carriers CLC with gaseous fuels, e.g. NG, has been widely investigated [49,186] and it has been shown that high gas conversion can be achieved with acceptable OC properties. The main 23 ACS Paragon Plus Environment

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differences among Cu-, Fe-, and Ni-based OCs in relation with the design of a CLC system are available from previous studies [43]. Metal oxides with the capacity to transfer oxygen in a CLC system are essential for successful CLC operation. The metal oxides should show a favorable tendency toward high conversion of fuel gas to CO2 and H2O, or in the case of chemical looping reforming (CLR), to CO and H2. Jerndal et al. [187] identified oxides of Cu, Ni, Co, Fe and Mn with favorable thermodynamics for CH4 and syngas conversion. At temperatures and pressures relevant to CLC, CH4 is not thermodynamically stable in the FR and variable amounts of CO2, H2O, CO, and H2 appear, depending on the operating conditions. A higher equilibrium constant means a higher conversion of the reducing gas. CuO-Cu2O/Cu, Mn3O4-MnO and Fe2O3-Fe3O4 are typical OCs with high selectivity towards CO2 and H2O (see Table 4 and Table 5). For Ni-based OCs, an equilibrium conversion of 99-99.5 % for H2 and 98-99.4% for CO can be obtained at 800-1000oC. When Al2O3 is used as a supporting material, the formation of NiAl2O4 is favorable at CLC conditions, leading to a lower conversion of these gases (93-98 %). For a CoO/Co system, the maximum conversion at equilibrium conditions is 95-97% for H2 and 92-97% for CO at 800-1000oC. Full conversion could be obtained in the redox system Co3O4CoO, but oxidation to Co3O4 is not favoured in air at temperatures above 880oC [7]. The CaSO4-CaS system for CLC application has also been reported [7,188–190], with thermodynamic limitations similar to NiO for maximum conversion of H2 and CO (See Table 5). Fe3O4 usually appears as an intermediate product in Fe2O3 reduction, with further reduction of Fe3O4 to FeO being a slow reaction. Formation of Fe(II)Al2O4 or Fe(II)TiO3 changes the thermodynamics when Al2O3 or TiO2 is used as the support, the reduction of Fe(III) to Fe(II) provides more oxygen to the fuel leading to almost complete conversion of CO and H2 to CO2 and H2O, respectively. The equilibrium constant of Fe3O4-FeO with H2 and CO is 1.5-2.5 and 24 ACS Paragon Plus Environment

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even 1300oC with stable redox cyclic performance. NiO loading on Yttria stabilized Zirconium (YSZ) provides high solid diffusivity for the nickel oxide ion and helps to improve the composite material reactivity, as well as regenerability [49]. NiO on bentonite also has been investigated, showing promising activity and stability under repeated redox cycles [249,261–263]. However, the oxidation curve showed that the oxidation reaction rate was quite small. Moreover, the thermal

stability of NiO/bentonite OCs is limited at higher temperature. TiO2 is also a common support material for catalysts, which has been tested for OC development. It was found that NiO was more prone to interact with TiO2, forming NiTiO3, which is less reducible than NiO. The loaded nickel is completely converted to NiTiO3 after several reduction/oxidation cycles, leading to lower reactivity of the OC. Moreover, coke formation was greater than for NiO on other common support materials [264–266].

NiO

supported on SiO2 and ZrO2 has also been investigated. It was shown that the reactivity of both NiO/SiO2 and NiO/ZrO2 decreased with repeated reduction–oxidation cycles above 900oC. Formation of nickel complexes is the main reason for the decreasing reactivity over time [241,245,246]. Different preparation methods, such as spray drying, incipient impregnation and mechanical mixing, have been used to improve performance of the OCs and reduce the production cost. OC properties and quality (e.g. reactivity and crushing strength) rely heavily on the preparation methods [267]. High reactivity and low NiAl2O4 formation were found in some cases to result from mechanical mixing or impregnation methods [246,267–270]. Ni-based OCs prepared by impregnation on α-Al2O3 had very high reactivity, with low attrition rates and low 38 ACS Paragon Plus Environment

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agglomeration during FBs operation [269]. Dueso et al. [271] concluded that about 80% of the Ni reduced in the FR was oxidized to free NiO, while the remaining Ni oxidized to NiAl2O4, indicating that the formation of the spinel compound (i.e. NiAl2O4) cannot be completely avoided. Despite these limitations, NiO on Al2O3-based support holds significant promise as a potential OC for large-scale CLC applications [174,213,271]. Crushing strength and attrition resistance are important properties for OCs, especially for uses in moving-bed and/or fluidizedbed processes. After about 200 cycles in the temperature range of 500-800oC, there is a small decrease in the solid conversion of Ni-CaAl2O4 with H2, CO and CH4 as reducing agents in the TGA and fixed bed tests, mainly due to agglomeration of the NiO grains [272]. Nevertheless, the redox kinetics is still sufficiently fast for low temperature applications if the OC is pre-activated. The kinetics rates for the gas-solid reactions and gas-phase catalytic reactions have been determined, which can be used to predict the performance of the activated NiO/CaAl2O4 OC for low temperature CLC applications. More information of Ni-based OCs is available in previous literatures [254,273].

4.2 Cu-based oxygen carriers Cu-based OCs have several advantages relative to other OCs, such as: 1) Highly reactive in both reduction and oxidation; 2) Reduction and oxidation reactions are both exothermic; 3) CuO reduction is favored thermodynamically to reach complete conversion using gaseous fuels, such as methane; 4) Higher oxygen transport capacity than Fe; 4) Less costly than Ni; 5) Low toxicity. However, CuO has not received significant attention due to its tendency to undergo agglomeration and decomposition at relatively high temperatures [274–277]. Different support materials (e.g. Al2O3, bentonite, CuAl2O4, MgO, MgAl2O4, SiO2, TiO2 and ZrO2) and different preparation methods (impregnation, co-precipitation, spray drying, freeze granulation and mechanical mixing) have been used to make Cu-based OCs. Impregnation 39 ACS Paragon Plus Environment

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on α-Al2O3, γ-Al2O3, MgAl2O4 or NiAl2O4-Al2O3 has been found to be optimum preparation method to avoid agglomeration and improve performance of Cu-based OCs [276,277]. Impregnation on SiO2, TiO2 or γ -Al2O3 [236,276–278] or co-precipitation with Al2O3 [279] produced Cu-based OCs with excellent chemical stability and good mechanical strength. If Al2O3 is used, CuAl2O4 is formed, becoming the dominant in the OC after a few cycles. However, CuAl2O4 is highly reducible, showing very high reduction reaction rates, similar to those of CuO [236,277–279]. SiO2 is quite inert to Cu, even at high temperatures and did not form any Cu-SiO2 complex [277,279]. Particle reactivity and stability were reasonable, with reactivity inferior to that of Ni supported on SiO2. However, the copper based OC suffered from CuO decomposition to Cu2O [241,268,274–276,280,281]. TiO2, as a Cu support, displays limitations for CLC due to its tendency to form CuTiO4 [277,279,280]. It was shown that composite particles consisting of CuO (28-37wt%) and inert support materials such as ZrO2, YSZ, CeO2, and MgAl2O4 provided full conversion of CH4 at 900−925°C, and were also found to release gas-phase O2 into inert atmosphere at these temperatures when fluidized with N2, whereas OCs using semi-active support such as Fe2O3, Mn2O3, and Al2O3 formed combined spinel structures with CuO. Materials with semi-active support had less reactivity with CH4 [282]. The rate of CuO oxygen uncoupling becomes significantly higher when reaction temperature reaches 900oC, whereas the increase in the reaction rates between CuO and H2/CO/CH4 is less profound [283]. Low CuO content (e.g. 50 cycles, Segregation of CuO from Al2O3 in the CuAl2O4 was observed during gaseous fuels combustion, which produced more available oxygen for CLOU than the initial material. It was reported that ZrO2 support could increase the chemical stability, mechanical strength and the reactivity of Cu-based OC [285,286]. Wang et al. [285] showed that the reduction and oxidation rates of Cu-based OCs with different inert supports increased in the order of ZrO2> TiO2> SiO2. Gayán et al. [286] suggested that the attrition rate of the Cu-based OCs supported on ZrO2 was low and stable (0.045%/h). However, the reaction mechanisms between CuO supported on ZrO2 and reaction gas (CO, O2) have not been investigated, and the detailed electronic property and the sintering inhibition mechanism of CuO/ZrO2 are still not well understood [287]. Adánez et al. [61] and de Diego et al. [81,275] tested 15 wt% CuO impregnated on γ-Al2O3 in 500 Wth and 10 kWth CLC units with syngas and CH4 as fuels. 120 h of operation were achieved in the continuous 10kWth CLC unit, and the effects of the operating conditions on fuel conversion and performance of the OC were tested. The OC to fuel ratio, ϕ, was found the most important parameter affecting fuel conversion. For a FR temperature of 800oC and ϕ>1.4, CH4 could be 100% converted to CO2 and H2O without CO and H2 emissions, no observable agglomeration and carbon deposition, attrition rate almost constant and as low as 0.04 wt%/h during 50 h of operation. Forero et al [288] analyzed the behavior of CuO impregnated on γAl2O3 and found that only 60 h stable operation could be reached for the FR at 800oC and AR at 900oC due to agglomeration. Gayán et al. [277] reported the Cu-based OCs with γ-Al2O3 41 ACS Paragon Plus Environment

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modified with 3 wt% NiO as support achieves more than 67 h of stable operation with FR and AR temperatures of 900oC and 950oC, respectively. In addition, Forero et al. [289] discovered that sulfur impurities present in the feed gas did not affect the reactivity of the OC and that full CH4 conversion was reached in the FR. Penthor et al.[290] tested a 120 kW chemical looping pilot plant with NG as the fuel and Cu-based oxides as the OC, the Cu-based OC showed good performance regarding conversion of CO and H2 (~100%). However, only moderate conversion (up to 80%) was achieved for CH4. Cao et al. [291,292] studied a FB CLC process at 600oC using solid fuels and Cu-based OCs. No particle agglomeration was observed. However, after the test, Cu2O was found, attributed to the decomposition of CuO. It was shown that CuO decomposition reactions occurred in the AR at a low oxygen concentration (e.g. < 20 wt%). An excess of air is suggested in the AR to minimize CuO decomposition [257,280]. Previous study indicated that agglomeration occurred at 950oC for the Cu-based OCs, mainly depending on CuO content and preparation methods [280].

4.3 Fe-based oxygen carriers Fe-based OCs are considered to be an attractive option for CLC applications due to their low cost and toxicity [43,293]. However, they have low oxygen transport capacity, weak redox characteristics and low methane conversion. Different oxidation states (Fe3O4, FeO, or Fe) can be found when hematite (Fe2O3) is reduced. Only the transformation from hematite to magnetite (Fe2O3-Fe3O4) may be applicable for industrial CLC systems due to thermodynamic limitations. Further reduction to wustite (FeO) or Fe is slow and would produce more CO and H2, reducing CO2 purity produced by the FR [187]. Preparation methods included mechanical mixing, freeze granulation, impregnation and coprecipitation. The active metal oxide content has ranged from 20 to 100 wt%, most studies have included metal contents exceeding 60 wt% due to the low oxygen transport capability of the Fe42 ACS Paragon Plus Environment

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based OC. Some of the works have used pure Fe2O3. The original hematite (Fe2O3) is a nonporous and smooth textural material of low specific surface area. It changes to a coarser texture with cracks and fissures when exposed to alternating reduction and oxidation cycles. The oxidation rates of these natural ores are adequate in CLC. Some researchers [257,294] have found agglomeration problems in the bed associated with phase change from wustite to magnetite when oxidized in air. Abad et al. [43] noted that Fe2O3 metal contents below 10 wt% are unlikely to be suitable for CLC operation due the limited solid circulation rates between interconnected fluidized-bed reactors. To improve reactivity and/or overcome agglomeration, a variety of materials have been used as supports for Fe-based OCs, e.g. Al2O3, MgAl2O4, SiO2, TiO2, ZrO2, with alumina being the most common. The solid-state reaction between Fe and Al2O3 is considered to be the main cause of the loss in particle reactivity [49]. However, with Al2O3 or TiO2 as the support, iron aluminate (FeAl2O4) or iron titanate (FeTiO3) can be formed as the reduced compounds corresponding to Fe(II). The reduction of Fe(III) in Fe2O3 to Fe(II) in FeAl2O4 and FeTiO3 increases the oxygen transport capacity from the AR to FR compared with hematite (Fe2O3) being reduced and transformed into Fe+2.67 in magnetite (Fe3O4). This is beneficial to fully convert the fuel gas to CO2 and H2O [125,295,296]. Note that Fe(II) in FeAl2O4 and FeTiO3 to Fe are very slow reactions, perhaps too slow to use these reactions for CLC application. Excess metal loadings may not increase oxygen carrying capacity much due to interactions between the metal and support. MgAl2O4 has been considered as a support for Fe2O3 to avoid aluminate formation, and increase the reactivity and stability up to 1100◦C [241,251,297,298]. When SiO2 is used as the support, unreactive iron silicates are formed, which can drastically reduce the reactivity of the OC after a number of cycles [241,245]. Fe2O3 supported on YSZ showed stable activity over 43 ACS Paragon Plus Environment

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cyclic reductions and oxidations without metal-support interactions confirmed by XRD analysis of the oxidized samples. Reduction rates of Fe2O3/YSZ were slower than for Fe2O3 supported on Al2O3, at least partly because of complete reduction of Fe2O3 to Fe for Fe2O3/YSZ OCs, and formed iron carbide (Fe3C) contributing to the inferior Fe2O3/YSZ reactivity [264,299] . Several group of authors have shown that Fe-based OCs have enough reactivity both at atmospheric [43,251,268,297] and pressurized conditions [300] for H2 and CO fuel gases, while being lower for CH4. Fe-based OCs also bring other benefits to CLC systems: low tendency for carbon deposition [300] and no risk of sulphide or sulphate formation at any sulfur concentration or operating temperature [187]. Most previous works were conducted in laboratory reactors in a batch-wise mode (TGA or fluidized-bed reactors) with gaseous fuels (usually methane) for CLC application. Several Febased materials have been investigated for CLR application. There are a few pilot studies ranging from 300 Wth to 10 kWth for continuous circulation units using gaseous (e.g. syngas and CH4) and solid fuels (e.g. biomass) with Fe-based OCs (20-100 wt% Fe2O3) [26,60,261,301– 303]. A test has been performed in a 300Wth continuous unit at temperatures from 800 to 950oC for a total of 40 h operation with NG or syngas as fuel, using an OC of 60 wt% Fe2O3 on an Al2O3 support prepared by freeze granulation [26]. There was no sign of deactivation, agglomeration, carbon deposition, and very little attrition. The combustion efficiency of syngas was about 99% for all experimental conditions, while methane was detected in the flue gas from the FR, and combustion efficiencies ranged up to 94% for NG combustion. Hence, Fe-based OC was better suited for syngas than for methane combustion. Lyngfelt and Thunman [60] conducted research on Fe-based OCs in a 10kWth pilot plant using CH4 as fuel at high concentrations. Up to 2-8 v% CO was observed in the flue gas from the FR during 17 h of continuous operation, even at low fuel flow rates, high circulation rates or high FR temperatures. 44 ACS Paragon Plus Environment

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Similar conclusions were reached by other researchers [261,302]. Ortiz et al. [302] found that 100% combustion efficiency of PSA off gas (CH4, CO, CO, H2) could be obtained at a FR temperature of 880oC and OC to fuel ratio > 4 during 50 h of operation with no carbon deposition, agglomeration or de-fluidization problems. Combined oxides of iron, manganese and silicon have been used as OCs for CLC combustion in a FB reactor with continuous circulation of solids designed for a thermal power of 300 W [304], full conversion of syngas and above 95% conversion of NG above 900 oC have been achieved. The study showed that it was possible to achieve very high fuel conversion with combined oxides of iron, manganese and silicon as the OC. However, the mechanical stability of the particles is rather poor and need to be improved. Galinsky et al. [305]investigated the effect of support on the cyclic redox performanceof iron oxides as well as the underlying mechanisms. The results indicated that the redox properties of the OC, e.g. activity and long-term stability, were significantly affected by interactions of the support and iron oxides. The perovskite supported OC exhibited high activity and stability compared to OC with ceria support. The high stability of perovskite supported OC is attributable to its high mixed ionic–electronic conductivity. Deactivation of ceria-supported samples may result from solid-state migration of iron cations and subsequent enrichment on the OC surface. Application of CLC to solid fuels involves char gasification, as the rate-limiting step in the coal conversion [144,145]. Pure Fe2O3 as OC was the focus using TGA or batch reactors for testing of solid fuels in past years. Low reactivity was found for pure Fe2O3 powders, while Fe2O3/MgAl2O4 exhibited good performance in batch FBs. Bimetallic OCs consisting of Fe2O3 and a small amount CuO (5 wt%) is more effective for coal char conversion than an OC without copper addition, mainly because gaseous oxygen is released at a higher temperature from CuO

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[67,306–309]. Further research in continuous operations is needed to gain a more adequate understanding of the performance of Fe-based OCs. Lin et al. studied the detailed adsorption, oxidation and dissociation mechanism of CO on the perfect and reduced Fe2O3 surfaces using periodic density functional theory [310]. The correlation between the interaction mechanism and reduction degree of iron oxide was also investigated. The results indicated reduction of iron oxide favors adsorption step of CO. After CO adsorption on surface, CO oxidation and decomposition occur subsequently. The surface plays two roles, namely, serving as reactant to oxidize CO and as catalyst to catalyze CO decomposition. Partial rather than complete reduction of Fe2O3 contributes to more efficient CO oxidation and less carbon deposition. The results provided a detailed understanding of behavior of the perfect and reduced iron oxide surfaces during CLC process, which will favor the optimization of CLC process [310]. Fe-based OCs have been extensively studied for the CLC process, which have shown sufficient activity among transition metal oxides. However, reactivity improvements are still necessary to achieve high conversion in short residence times to minimize the size of commercial CLC process [7,8,311,312]. Several pilot scale tests have been conducted using Fe-based natural mineral OCs and NiO-based OCs, the latter showed promising results in pilot scale tests[43,239,240]. However, Ni-base OCs require special handling because of the potential health effects of nickel dusts [313,314]. A bench scale countercurrent moving bed unit was studied, which ran for 15 h with syngas conversions higher than 99.5% and an iron oxide conversion of 50% (mol O consumed per total mol O in Fe2O3) [35,178]. Similarly, Tong et al. have performed parametric studies on a 25 kWth moving-bed reducer reactor by changing the gas hourly space velocity (GHSV) of the NG injected [185]. Their results showed that high conversion of methane (98%) could be achieved at 46 ACS Paragon Plus Environment

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a number of conditions in the FR temperature of 930-975oC, methane flows from 2 to 5 slpm, and OC flow rate of 150 g/min. These results showed high conversion of OC, with the maximum OC conversion being 43.6% at 98.9% methane conversion [185]. MgO-promoted natural mineral hematite (Fe2O3) and synthetic mixed metal CuO – Fe2O3/Al2O3 have shown excellent reactivity and stable performance during cyclic CLC tests conducted with CH4/air at 700–850oC in an atmospheric FB reactor [313]. The presence of MgO on the hematite OC significantly improved the oxygen utilization of hematite for methane CLC. Cyclic CLC tests conducted in the FB with MgO promoted hematite showed better performance than that with hematite. The CuO–Fe2O3/ Al2O3 OC exhited excellent performance during the 25-cycle FB CLC tests conducted at 800oC with methane and air. Full conversion of CH4 to CO2 and stable oxygen transfer capacity were observed during all the cycles. The fluidization of the OC was good, with no particle agglomeration as is traditionally observed with CuO-containing materials. Attrition resistance of both OCs with particle size of 100–150 µm was better than that of standard FB cracking catalysts [313]. Reactivity and stability of the MgO-promoted hematite were also tested in the thermogravimetric analyzer (TGA) and bench scale reactors. The incorporation of 5 wt% MgO led to an increased reaction rate and an increase in oxygen utilized as compared to the pure hematite OC. These studies revealed that the best performing OC was the 5 wt% MgO/Fe2O3 which exhibited no observed degradation in the kinetics and conversion performance in the methane step over 15 reduction and oxidation cycles. The Mg-promoted OC also showed reduced coke formation as compared to the pure hematite carrier [314]. OCs were made from Fe2O3 powder by granulation and impregnation with Ca(NO3)2.4H2O [315]. The reactivity and cyclic stability of the OCs were tested by alternating reducing (using H2 or CO) and oxidizing environments (using air and CO2), in a TGA and a FB reactor at 850-950°C. It was found that 47 ACS Paragon Plus Environment

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the one containing 20 mol% CaO showed optimal performances. However, others showed improvements in oxygen carrying capacity, cyclic stability and reactivity over unmodified iron oxide but to a lesser extent [315].

4.4 Mn-based oxygen carriers Mn-based materials are considered non-toxic and inexpensive. Moreover, their oxygen transport capacity is higher than for iron compounds. However, relatively few studies have dealt with Mn-based materials as the OCs for CLC. The highest oxidized manganese compound, MnO2, decomposes at 500oC. Mn2O3 is more thermodynamically stable than MnO2 in air at relatively high temperatures [316]. However, only Mn3O4 is present at temperatures above 800oC [260]. Therefore, only the transformation between Mn3O4 and MnO is considered for CLC applications. Pure manganese oxides (e.g. Mn2O3, Mn3O4) have shown low reactivity with CH4 and coal [317,318]. Manganese oxides react with inert materials, e.g. SiO2, TiO2, Al2O3 or MgAl2O4, to form highly irreversible and unreactive phases which reduce the reactivity of OCs [236,245,257,268]. The mechanical strength of the Mn-based OC was low when sepiolite (a complex magnesium silicate) was the support [268]. Manganese oxides with bentonite (aluminium phyllosilicate) as binder has shown good reactivity to a mixture of H2 and CO. However, the reactivity was very sensitive to the presence of H2S in the gas mixture [319]. OCs with ZrO2 as the support showed good reactivity and stability through consecutive redox cycles [186]. During heat treatment and reactivity testing, Mn-based OCs with ZrO2 showed agglomeration or underwent a phase transformation, producing cracks in the structure [195,320,321]. To avoid these problems, new OCs were prepared with ZrO2 stabilized by addition of MgO, CaO or CeO2 [195]. Mn-based OCs supported on ZrO2 stabilized with MgO showed good reactivity with syngas [322], but lower reactivity was found for CH4 [186]. These particles were better suited for syngas than for methane combustion in terms of 48 ACS Paragon Plus Environment

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reactivity, which was demonstrated in a continuously operated 300Wth CLC unit [28]. Very high efficiencies (>99.9 %) were obtained at temperatures of 800-950oC for syngas combustion. For NG combustion, some methane was detected in the gas outlet from the FR and combustion efficiencies ranged from 88 to 99%. The perovskite-based OCs were synthesized. Ca(OH)2 and MnCO3 powders with an average particle size of ~46µm (325 mesh) were mixed in a certain mass ratio with or without La2O3, Fe2O3, ZrO2 and SrCO3 to make the OC particles. This new modified calcium manganese perovskite structures applicable in CLC were investigated [323]. All prepared samples showed a porous surface and the perovskites phase formation confirmed by XRD results. Reactivity and oxygen uncoupling behaviors of the prepared OCs were also evaluated using a FB CLC reactor using methane as the fuel at four different temperatures (800, 850, 900, 950oC). All OCs showed acceptable quantity. Oxygen uncoupling properties and reactivities for methane combustion of 12 OC particles, produced from mixtures of Mn and Mg oxides with optional addition of TiO2 or Ca(OH)2 were investigated in a quartz batch reactor at 810oC, 850oC, 900oC and 950oC. The addition of Ca(OH)2 facilitated oxygen release and combustion of methane. However, addition of TiO2 did not have a considerable effect on either oxygen uncoupling or reactivity of the OC. OCs with greater extent of oxygen release generally have better methane combustion properties [137].

4.5 Co-based oxygen carriers The cobalt oxide has been considered as a possible OC because of its high oxygen transport capacity, but subject to high cost and environmental concerns. Several oxidation states can be involved in the redox reactions cycles with cobalt, Co3O4 is unstable above 900oC, being converted into CoO. Hence, only the transformation between CoO and Co is considered for CLC

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applications, although it is less thermodynamically favourable, with maximum conversion of 9597% for H2 and 87-97% for CO in the temperature range of 800-1200oC [7]. With Al2O3, TiO2 and MgO as inert supports, the metal oxide and support materials suffered strong interactions, forming unreactive compounds such as CoAl2O4, CoTiO3 and Mg0.4Co0.6O. Mattisson et al. [236] reached a similar conclusion with an OC prepared by impregnation using Al2O3 as support, indicating that these materials were not suitable for CLC. Jin et al. [259,299] observed that CoO/YSZ OCs exhibited good reactivity and low carbon deposition in their TGA studies. Ryu et al. [177] reported 25 h of continuous operation in a 50 kWth CLC unit with a Co-based material supported on CoAl2O4. They reported a 99.6% of CH4 conversion, although the attrition resistance of the OC required improvement to accomplish long-term CLC operation.

4.6 Mixed oxide oxygen carriers Mixed metal oxides sometimes provide better features than individual metal oxides. However, only limited studies have focused on mixed metal OCs for CLC application. The investigations have been performed either by mixing different active metal oxides in the same particle or combining different OCs each composed of single metal-oxides. The purpose is to increase reactivity, thermal stability and mechanical strength, improve reaction rate, conversion and efficiency, decrease carbon deposition and poisoning of the OC, and minimize cost and toxicity.

4.6.1 Cu–Ni: Adanez et al. [47] and Johansson et al. [298] reported a stable bimetallic Cu–Ni/Al2O3 OC. It was claimed that Cu and Ni stabilized each other, providing improved performance at high reaction temperatures and high oxygen carrying capacity. Cu-based OC (13 wt% CuO) prepared using γ-Al2O3 as support modified with a small NiO addition (3 wt%) was tested with CH4 as 50 ACS Paragon Plus Environment

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fuel in a 500 Wth CLC of continuous operation [277]. These particles showed high metal oxide utilization, complete CH4 combustion, and low and stable attrition rate without agglomeration during 67 h of operation at high temperature (AR at 950oC and FR at 900oC). A particle lifetime of 2700 h was estimated for these particles.

4.6.2 Co-Ni and Co-Fe Jin et al. [264,299] reported synergetic effects of NiO and CoO supported on YSZ. The OC was prepared with equal amounts of NiO and CoO, showing excellent regenerability under repeated oxidation and reduction. Jin et al. [299,324] found that the reduction and oxidation reaction rates of Co-NiO supported on YSZ were slightly lower than those of the individual metal oxides because a solid solution (NiCoO2) between NiO and CoO was formed. The bimetallic oxides provided excellent performance, with good reactivity, complete avoidance of carbon deposition, and significant regenerability for repeated cycles of reduction and oxidation. Despite these encouraging results, YSZ supported bimetallic Co-Ni materials have not received further attention, possibly due to low thermal stability and the high cost of the YSZ zeolite. Some researchers [325–327] have demonstrated excellent reactivity and stability of a modified Co– Ni/Al2O3 OC. It was observed that Co helped to increase the dispersion of nickel and support the uniform and unchanged Ni metal particle size by enhancing the CoAl2O4 abundance and minimizing the NiAl2O4, thereby improving the reactivity and thermal stability of the OC, especially at high temperature. Co–Ni/Al2O3 increases CH4 conversion by 15% compared to unpromoted Ni-based OC in their tests. It was reported that mixed Fe-Co and Fe-Mn oxides loaded with 60% manganese transfer 3.6% of its oxygen [9].

4.6.3 Fe-Ni: A test was conducted in a TGA for a bimetallic Fe-Ni/Al2O3 [328], Addition of NiO to Fe2O3/Al2O3 particles was observed to improve the activity, but to decrease the mechanical 51 ACS Paragon Plus Environment

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strength. Son and Kim [261] carried out experiments in a continuous CFB using Fe-Ni/bentonite OCs. It was found that the reactivity of the OC particles increased with increasing NiO content. The optimum mass ratio of NiO/Fe2O3 was found to be 3 (NiO/Fe2O3= 75:25). The mechanical mixing of Ni-based and Fe-based particles has been investigated, it was found that a bed of Fe2O3 with only 3 wt% nickel oxides was sufficient to give very high CH4 conversion [298]. In addition, these researchers showed that the mixed-oxide system produced significantly more CO2 than expected based on the sum of the metal oxides operated separately, giving evidence of synergy when NiO is combined with Fe2O3. Rydén et al. [329] observed similar findings by mixing NiO-MgAl2O4 either in a bed of Fe2O3-MgZrO2 [294], in a bed of ilmenite [329] or in waste products from the steel industry [330]. The positive effects of the Ni addition have also benn found in continuous units. Ortiz et al. [302] reported an increase in combustion efficiency in a continuous 500 Wth CLC prototype in which PSA off-gas was the fuel.

4.6.4 Fe-Cu: OCs containing Cu and Fe with the spinel structure prepared by different methods were tested in a TGA apparatus and in a batch fluidized-bed reactor [58,331–333]. The best working spinel formulation was Cu0.95Fe1.05AlO4 [334], with high oxygen transfer capacity, high oxidation rate, but a relatively low reduction rate compared with a reference Ni-based OC (NiONiAl2O4). Lambert et al. [333] found that impregnating NiO on this spinel material increased both its oxygen capacity and the reactivity of the resulting material. However, addition of CuO on the spinel led to agglomeration and defluidization of the bed during the reduction and oxidation. The CuO content should be < 20 wt% to avoid these problems. Combined Cu-Fe and Mn-Ni OCs were investigated as bed materials for CLC [111]. The aim was to identify material combinations with a high reactivity towards gaseous fuels, such as CO, H2 and CH4, as well as sufficient mechanical durability. For this purpose, 18 different OCs 52 ACS Paragon Plus Environment

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were spray-dried and calcined at defined temperatures. Gas conversion as well as release of gaseous oxygen was investigated in a batch FB reactor setup at temperatures between 850 and 1050°C. For mechanical durability, a jet-cup attrition rig was used. Moreover, properties like specific surface area, oxygen transfer capacity and crystalline phase composition were examined to physically and chemically characterize the OC particles. For both the Cu-Fe and Mn-Ni-based materials, OCs could be produced showing a high reactivity with gaseous fuels like CO or CH4 while having a sufficiently high mechanical strength. These properties make them interesting candidates for application in CLC.

4.6.5 Mn-Fe, Mn-Ni, Mn-Cu and Mn-Ca Fe-Mn mixed oxides were studied to check for synergetic effects between the two [335]. OCs consisting of Fe2O3-MnO2 supported on ZrO2 and sepiolite were tested [336]. Both carriers exhibited excellent reaction performance and thermal stability for a CLC process at 800oC. The support had a substantial effect on the reaction rate. Sepiolite appeared to be a better support than ZrO2. Moreover, Fe-Mn-O mixed oxides, as well as Ni-Mn-O materials, have oxygen uncoupling properties (releasing O2). From evaluation for CH4 conversion, only one Mn-Fe-O material showed enough high reactivity and mechanical strength to have potential as an OC [337]. Particles discussed in Section 4.6.4 [111], Section 4.3 and 4.4 [304,323] involved Fe-Mn and NiMn based OCs. Naturally occurring manganese oxide minerals were usually disregarded as the OC owing to its low crushing strength. Adding iron oxides (i.e. Mn-Fe) and calcium (i.e. Mn-Ca) improved its mechanical resistance [338,339] and showed better performance than ilmenite for gaseous fuels although its reactivity may be lower than nickel oxide [45]. Xu et al. [340] increased a Chinese manganese ore (40% manganese) activity by impregnating it with 10% copper but the carrier agglomerated. Larring et al. [341] evaluated Fe and Mn-based OC minerals, with Mn and Fe 53 ACS Paragon Plus Environment

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content ranging from 0% to 39%, and from 24% to 82%, respectively. It showed that the oxygen transfer capacity of pure iron ores was greater than that of mixed iron and manganese ores. Manganese oxides outperform the other ores only in the case of chemical looping with a solid fuel [342], indicating manganese oxides are inferior to iron ore for CLC of gaseous fuels. Mn3O4 releases 7% of its oxygen (by mass), which makes it attractive for CLOU process. However, due to equilibrium limitation, CLOU system with manganese oxides would need to be operated at temperatures below 800oC [343]. A Fe-Mn mixed oxide spray dried with alumina as a support was among the best carriers in terms of oxygen transfer (2.5% by mass) [9]. Impregnating Fe-Mn mixed oxides with cerium inhibits formation of iron-manganese spinel structure, thereby increasing its stability [344]. Mixed copper-manganese oxides prepared by coprecipitation lost as much as 12% of its weight as oxygen [345], indicating excellent oxygen releasing ability and good CLOU performance. Synthesizing manganese oxides is an alternative to the natural ores, more information is available in previous study. CaMn0.875Ti0.125O3 of the perovskite structure has been examined as OC for CLOU of NG, by 70 h of experiments in a CFB reactor system [122]. For CLC, combustion efficiency of 99.8% can be achieved when 1900 kg OC/MW NG was used. The OCs retained their physical properties, reactivity with CH4 and ability to release O2 throughout the testing period and there were no problems with the fluidization or formation of solid carbon in the reactor.

4.6.6 Other complex mixed oxides OCs: A 10 kWth pilot reactor of continuous operation was used to examine a calcium manganite based OC with NG as fuel during 99 h operation. The composition of the OC can be expressed by the formula CaMn0.775Ti0.125Mg0.1O3. The main part of the OC forms a perovskite crystal structure which has oxygen releasing properties. The fuel conversion was generally above 95% and full conversion was reached at certain operating conditions. The elutriation of fines (110 h continuous operation. The OC exhibited sufficient reactivity to completely convert syngas to CO2 and H2O at 880oC, with required OC of 400kg/MWth. However, the combustion efficiency with CH4-containing fuel gases ranged only from 75% to 80% because of the lower reactivity of CH4 compared to CO and H2. Some researchers [121] found a Mn ore to be very promising for CLC application based on its reactivity, while others considered the Mn ore unsuitable due to poor mechanical strength and unsatisfactory fluidizing properties [120]. Five OCs based on three manganese ores originating from Slovakia, Brazil and South Africa, respectively, were investigated in a batch-type fluidized reactor [369]. Besides the natural manganese ores, two more materials were prepared where 38.4 wt% of Ca(OH)2 was mixed with 61.6 wt % of the Brazilian and the South African Mn-ore, respectively. Spherical OC particles were produced using the spray-drying methods. The study 58 ACS Paragon Plus Environment

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showed that manganese ores are highly promising as OCs for CLOU combustion. Furthermore, it was possible to produce a low-cost, reactive and mechanically stable OC through spray-drying of a mixture of manganese ore and Ca(OH)2. A comparison between the pure ores and the combined materials show that the addition of Ca(OH)2 had generally a beneficial impact on the mechanical stability of the OCs. The produced OCs by mixing Brazilian ore and Ca(OH)2 showed a decrease in reactivity towards methane, but a higher reactivity towards syngas in comparison to the pure ore. While combined OCs of South African ore and Ca(OH)2 showed a major improvement in reactivity as compared to the pure ore towards both syngas and methane, which may be attributed to formation of a perovskite material with significant CLOU properties.

4.7.3 CaSO4 from natural anhydrite CaSO4 from natural anhydrite is a low cost material with much higher oxygen transport capacity than other materials. Interest on the use of CaSO4 as OC has been shown by several research groups in China [189,370]. The reduction reactions rate is low for CaSO4 [190,371– 374]. Small fractions of CO and H2 cannot be converted to CO2 and H2O due to thermodynamic equilibrium limitations [371,375]. ALSTOM is developing a Hybrid Combustion-Gasification Chemical Looping Process with CaSO4 as the OC for heat generation, syngas production or hydrogen generation [376,377]. Complete combustion of solid fuels is needed for heat generation, but partial oxidation is required if syngas is the desired product. Production of hydrogen can be accomplished by coupling a CaCO3/CaO cycle to the CaSO4/CaS cycle with the use of an additional calciner. Tian and Guo [378] investigated the reduction behaviour of CaSO4 by CO and the results showed that CaSO4 could be converted to CaS completely on the condition that the partial pressure of CO in the atmosphere was big enough. Deng et al. [379] investigated the reaction kinetics model of the FR (CaSO4 + H2) to mimic the behaviour of fuel gas flow in the reactors and a multiphase hydrodynamics model was developed. One of the concerns using 59 ACS Paragon Plus Environment

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CaSO4 is the possible formation of CaO by side-reactions evolving SO2 to the reacting gases [189]. More research is required for CaSO4 as the OC, especially for continuous CLC operations. Experiments were conducted in a laboratory scale fixed-bed reactor under CO–H2–N2 atmosphere for Fe2O3/CaSO4 composite OCs prepared by impregnation method with natural anhydrite as active support and Fe2O3 as additive [380]. It showed that Fe2O3 could enhance the reduction of CaSO4 to CaS and increase CaS yield, over 85% SO2 released from the decomposition of CaSO4 alone was inhibited by Fe2O3/CaSO4 OCs with the Fe2O3 loading contents higher than 5 wt% at 950◦C. The higher the Fe2O3 loading content, the higher the reactivity and sulfur inhibition ability Fe2O3/CaSO4 OCs could possess. Higher conversion of CO and H2 was achieved at higher temperature and Fe2O3 loading content. Redox cyclic tests indicated that Fe2O3/CaSO4 OC could keep a better recyclability and stability than CaSO4. A solar-hybrid trigeneration system based on methane CLC was proposed to produce electricity, chilled water for cooling and hot water [112]. CaS and CaSO4 are the cycle materials of the CLC, and the reduction reaction in the CLC is driven by solar thermal energy. The thermodynamic performances of the new CLC trigeneration system, including energy and exergy efficiencies, were analyzed and compared based on design conditions and variable parameters. The results indicated that the optimal solar heat collection temperature was approximately 900oC, the pressure ratio of the air compressor is 20, and the energy and exergy efficiencies reach 67% and 55%, respectively. The output ratios of the three products vary with the solar collection temperature and pressure ratio. Meanwhile, the net exergy efficiency and the saving rate of the solar collection area are expected to be 24% and 63%, respectively.

4.7.4 Pyrolusite Pyrolusite, a manganese oxide containing more than 85% amorphous MnO2 in its fresh form, can transfer more than 3 times oxygen compared to ilmenite-as much as 7% of its own 60 ACS Paragon Plus Environment

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mass. In addition, the oxygen transfer capacity can be increased by preheating pyrolusite under an oxidizing atmosphere. It can then transfer almost double its experimental oxygen transfer capacity versus pyrolusite pre-treated in an inert environment. Together with excellent oxygen carrying capacity, it has adequate attrition resistance and costs fraction of ilmenite. The solid inventory in the reduction reactor would represent 8% of that required for ilmenite and the solid circulation would drop by 30%, but the fresh particle makeup would be 4 times higher based on measurements from a laboratory-scale submerged jet attrition mill [381]. There are still arguments whether the pyrolusite or modified OCs based on pyrolusite are suitable for CLC processes. More research needs to be conducted in the future.

4.8 Coke formation and oxygen carriers poisoning There are several possible ways of carbon formation in CLC depending on different fuels in the FR and operating conditions, pyrolysis (or methane cracking if CH4 is the fuel gas) and Boudouard reactions are amongst the most probable reactions leading to coke formation [382]. The exothermic Boudouard reaction is more likely to take place at lower temperatures, whereas pyrolysis or methane cracking is an endothermic reaction thermodynamically favored at high temperatures. Both pyrolysis (or cracking) and Boudouard reactions are known to have a limited importance based on their kinetics if there is not a catalyst. Transition metals, such as Ni and Fe, can catalyze coke formation depending on reaction conditions (e.g. the availability of oxygen, fuel conversion, reaction temperature and pressure) [249,258,265,383,384]. Carbon can be gasified by steam or oxidized by oxygen. WGS is usually considered to reach equilibrium under CLC operating temperatures. Hence, WGS is very important for product gas composition from the FR. More CO2 make the Boudouard reaction move to CO direction reducing carbon deposition tendency on the particle surface. It is important not only to understand the possible carbon formation mechanisms but also to identify the operating parameters that need to be 61 ACS Paragon Plus Environment

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implemented to reduce or avoid carbon deposition, which can cause decaying reactivity of the OC. Ishida et al. [264] reported carbon deposition on YSZ supported Ni- and Fe-based OCs and concluded that carbon formation is mainly caused by the Boudouard reaction, with the amount of carbon deposited on the Ni-based OCs being greater than that found on Fe-based OCs under the same conditions. Moreover, higher temperature decreases carbon formation, and carbon deposition can be completely removed at 900oC for NiO/YSZ OCs. However, care should be taken when selecting the maximum operating temperatures since higher temperatures also cause metal sintering and phase transformation. Ryu et al. [249,383] reported similar findings for Ni-based OCs. Other researchers suggested that pyrolysis or cracking is the main reaction by which carbon deposition occurs on NiO/YSZ or NiAl2O3 carriers [258]. Carbon deposition by pyrolysis can be controlled by steam, steam from the reduction reaction should also be accounted in the required steam calculations [103,249,250,259,264] Ishida et al. proposed a correlation between the carbon deposition boundary and the input molar ratio of water vapor to fuel gas. In addition, carbon formation on a Ni-based OCs is strongly dependent on the oxygen availability, high oxygen availability from the OCs reduce carbon formation [258]. Jin et al. [103,259] considered a bimetallic CoO–NiO/YSZ system for a CLC study, their study suggested that carbon deposition can be completely eliminated by employing cobalt as a promoter along with NiO being the main active species on the YSZ support. It is speculated that formation of a solid solution between nickel and cobalt might change the free energy of the carbon formation reactions, giving rise to reduced coke deposition. It was reported that the Mg-promoted hematite OCs showed reduced coke formation as compared to the pure hematite carrier [314]. Fuel gases (e.g. NG, refinery gas, syngas, biogas), solid fuels (e.g. coal and biomass) and liquid fuels (e.g. heavy oil) are considered as potential fuels for CLC. These fuels contain some 62 ACS Paragon Plus Environment

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impurities, one of them is sulfur. NG usually contains a few of H2S about 20 ppmv. H2S concentrations in the refinery gas vary depending on the site and may reach up to 800 ppmv in some sites, while this value may even rise substantially to 8000 ppmv for the raw syngas [385]. The sulfur compounds may react with both the oxygen and supported metal in the FR, which can contribute to not only SOx formation but also the formation of metal sulfides. Ni-based OCs are very sensitive to sulfur. It was found that the conversion of H2S to SO2 is enhanced at high temperatures and low pressures of the FR. The formation of SO2 resulted from the oxidation of H2S with H2O, CO2 and even with NiO. Both SO2 and H2S could react with the active metal (NiO or Ni) to form sulfides or sulfates (NiS, NiS2, Ni3S2 or NiSO4). Ni3S2 was found to be the most thermodynamically stable sulfide [187,375,382]. The experimental investigation with Ni/αAl2O3 found a Ni3S2 rich component in the FR, while NiSO4 was a prevalent species for the OC present in the AR [252,386]. Dong et al. [387] exposed Ni-YSZ to hydrogen containing 100ppmv H2S at 727oC, Raman spectroscopy detect vibration mode of Ni3S2 although thermodynamic calculations indicate that Ni-YSZ would be stable. A desulfurization step before fuel combustion for fuels with higher sulfur content is suggested, and sulfur contents < 100 ppmv H2S was recommended in an industrial CLC plant for Ni-based OCs [252,386]. Ni is more prone to react with H2S or SO2 forming nickel sulfides. However, agglomeration problems have never been observed from previous investigations although the high sulfides formation was found in some extreme cases (melting point of Ni3S2 is 789oC). Cu2S was identified as the most thermodynamically favorable species under an oxygen deficient condition for Cu-based OCs [187,375]. Forero et al. [289] reported that 95% of the sulfur in the FR was released as SO2 under normal operating conditions and that Cu2S was observed only during operation at low oxygen to fuel ratios. No agglomeration was detected during the operation. Note that melting point of Cu2S is ~1130oC, higher than that of Ni3S2. 63 ACS Paragon Plus Environment

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Wang et al. conducted experiments with a sulfur-containing synthesis gas (4000 ppm H2S) as fuel and copper ore as the OC [388]. A weight gain was observed in all TGA experiments with 4000 ppm H2S synthesis gas as fuel, due to the sulfidation of the copper ore OC. It was reported tha the main metal sulfide products were Cu2S and FeS, while the sulfur species in the gas phase were mainly SO2, COS, and CS2. H2S was easier to react with copper oxides than iron oxides. A laboratory scale FB reactor was then used to investigate sulfidation of copper ore at 900oC. The results showed that the sulfidation of copper ore degraded its oxygen transport capicity and reactivity to some extent [388]. Gu et al. [389] observed that H2S was converted into COS, CS2 and SO2 at the exit of the FR and FeS was formed for the Fe-based OC in the FR. The higher the FR pressure is, the more FeS is found in the OC. Thermodynamic studies also reported Fe0.84S as the most present solid sulfur compound for Fe-based OCs [187,375]. CoS0.89 and MnSO4 were also identified as the favorable states of sulfur compounds in case of Co- and Mn-based OCs [187,375]. Based on TGA of several bentonite-supported metal oxides, Tian [390] reported that the redox rates decreased in the presence of H2S. The highest decrease was observed for NiO, while the lowest was for Mn2O3. Jerndal et al. [187] and Mattisson et al. [382] recently discussed the effects of sulfur components in CLC processes. It was shown that CuO, FeO, MnO-based OCs can convert H2S completely to SO2 between 600 and 1200oC. At similar reaction conditions, it is predicted that the H2S conversion is somewhat lower for Ni-based OCs. This reaction is significantly influenced by temperature and pressure, with the oxidation rate of H2S being increased at higher temperatures and lower pressures [391]. SOx and/or other sulfur compounds may be observed in the FR, metal sulfides formed in the FR may lead to release of SOx from the AR. The emission in the FR may adversely affect the quality of CO2 and its compression, transportation and 64 ACS Paragon Plus Environment

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sequestration. In general, sulfur results in decaying reactivity of the OC and enhance agglomeration because of the low melting points of metal sulfides. Pretreating of the fuel to remove sulfur upstream the FR or desulfurization downstream FR is necessary to meet the requirements of the environmental regulations. A TGA test was conducted using coal as fuel and MnFe2O4 as the OC. Itt was reported that sulfur present in coal was converted to solid MnS, which was further oxidized to MnSO4. The formation of both MnS and such manganese silicates as Mn2SiO4 and MnSiO3 should be addressed to ensure the full regeneration of the reduced MnFe2O4[392]. Ilmenite was used as the OC in CFB of CLC. The sulfur balance showeds that the conversion in the FR of inbound sulfur is around 72%. Furthermore, it was found that 75% of the S-containing gas is SO2, and only 25% is H2S. The nitrogen analysis indicates that 62% of the nitrogen fed with the coal is converted to gas, and that the nitrogen in this gas is distributed as 1wt% HCN, 11wt% NO, 26 wt% NH3, with the balance probably being N2 [393]. The nitrogen transfer of fuel-N in the coal is investigated with a NiO/Al2O3 OC under a continuous operation in a 1 kWth interconnected fluidized bed prototype [34]. Experiments by Sundqvist et al. [31] using 50% CH4 diluted with 0.5% SO2 in N2 as fuel, suggested that CaMn0.875Ti0.125O3-δ reacts with SO2 during CLC and forms CaSO4. It can be concluded that the effect of sulfur and other fuel impurities on such OCs needs to be carefully considered and experimentally examined.

5. Modeling development of CLC Modelling of the fuel and air reactors is beneficial for the design, scale-up and optimization of the CLC process. Most of the modeling works were developed for the two interconnected fluidized bed reactors concept, the most widely used configurations for CLC systems. The AR is designed as a high-velocity riser and the FR as low velocity BFB [68,394]. Nevertheless, the FR 65 ACS Paragon Plus Environment

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can also be operated in the fast fluidization regime to increase the fuel load [395]. Modelling of fluidized-bed reactors is generally divided into three main areas, which are closely related to each other: fluid dynamics, reaction kinetics, and mass and heat balance. Fluid dynamics, mass balances and heat balances in the reactor must be solved simultaneously because of the variation of reaction rates and gas properties. The suspension density and actual reaction rate at each position inside the reactor, the possible gas expansion as fuel is converted (e.g. when methane or coal is used), the size and growth of the bubbles, or relevance of reactions in the freeboard are other factors to be considered. The mathematical modelling of each reactor will improve the understanding of the fluid dynamics coupled with the complex chemistry happening in the reactors. In addition, the solids circulation rate and the solids inventory in the CLC reactors should be evaluated. Some modeling work for optimization of the FR operation has also been done [396–398].

5.1 Fluid dynamics models Fluid dynamics describe types of gas-solid contacting and interactions. The simulation approaches can be categorized as particle based methods, macroscopic fluid dynamics models and multi-phase computational fluid dynamics models (CFD) (see Table 9) [7,399].

5.1.1 Particle based methods Particle based models offer a high level of details but are computationally demanding and limited in the number of particles they can address. It is usually impracticable to model even a laboratory scale reactor with realistic particle size based on currently available computer systems, limiting the use of particle based model [275,400].

5.1.2 Macroscopic fluid dynamics models There are some empirical equations to predict hydrodynamic performance for FBs based on previous research. The macroscopic models estimate distribution of the gas flow among 66 ACS Paragon Plus Environment

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emulsion and bubbles as well as the distribution of solids concentration in the bed by using empirical equations. A more complete model also includes the solids distribution in the freeboard region. Kolbitsch et al. [401] considered the diffusion resistance between bubbles and emulsion by using a model parameter, which simulates an effective amount of solids exposed to gas phase. The actual value of this parameter in a FB could be different in different zones of the reactor, and it is difficult to determine. Models based on the two-phase theory for BFBs [402] or for FBs in the turbulent and fast fluidization regimes [403] have been used for CLC simulation. These models were developed to predict the fluid dynamics of large FB reactors. Pallarés and Johnsson [403] predicted the vertical profile of solids, showing good agreement with the experimental data for units as large as 226MWth CFB. Combustion efficiency for a 12 MWth biomass CFB boiler was also adequately predicted [404]. These models integrate the complex chemistry, where a fuel gas (e.g. NG) reacts with a continuously circulated OC, with the complex fluid dynamics of large fluidized-bed reactors, consuming reasonable computing time (order of minutes). In this way, modelling and simulation of the FR for CH4 as fuel gas has been developed for a 10 kWth BFB and a 120 kWth fast FB [194,398]. These models have been validated against experimental results obtained in the CLC units built at Spanish National Research Council (ICB-CSIC) and Vienna University of Technology (TUWIEN), respectively [7]. The macroscopic models have a great potential for the simulation, design and optimization of large FB reactors in CLC systems.

5.1.3 Multi-phase computational fluid dynamics models The computational fluid dynamic (CFD) codes are based on momentum, heat and mass transfer and do not required detailed assumptions in the modelling procedure. To date few CFD simulations have been performed for a full CFB due to the complexities in geometry and the flow physics, requiring a large computational effort. The task of simulating a full scale CFB is 67 ACS Paragon Plus Environment

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very challenging, and improvement of CFD methods for modelling full scale FBs is in development [405]. The fluid dynamics for AR, cyclone and FR in a small scale cold-flow CLC unit has been successfully modelled by using CFD codes [406,407]. With the improvement of numerical methods and more advanced hardware technologies, the use of CFD codes is becoming more affordable. CFD models for commercial scale BFBs has been recently presented [408,409]. Therefore, the use of CFD models can be of interest for the development of the CLC technology in the near future. Some researchers [379,410] applied multi-phase CFD modeling for CLC processes, especially for mall scale CLC systems (300-1000 Wth). However, most of them only focused on the FR operated as a bubbling bed in batch mode without solids circulation, and left the AR and the complex interaction between both reactors unconsidered. Important progress has been done validating the CFD models with experimental results obtained in small-scale facilities using gaseous fuels [411,412], or coal [413]. Up to date, the more complete modelling of a CLC system using CFD codes is the simulation of a BFB for the FR coupled to a riser for the AR using CH4 as fuel and Mn- or Ni-based OCs [399,414]. Kruggel-Emden et al. [399] developed 2D modeling for both air and fuel reactors using a multi-phase fluid dynamics framework. The CFD model was developed separately for air and fuel reactors and coupled through inlet and outlet conditions. Spherical shrinking core reaction kinetics model was employed. The developed CFD model was tested for chemical looping operated with CH4 as fuel gas and Mn3O4 as OC, including the initial degree of reduction of the solid, the kinetic model, the solid circulation rate and the air reactor length. Longer simulations, the stability of operation points, appropriate control strategies, and part load analysis has yet to be tested for the CFD model. Model validation by experimental data has to be done in the future.

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The relatively complex processes affecting the reaction of fuel gas with the OC such as full fluid dynamics, reactivity of the OC, the reaction pathway and the effect of solids circulation rate has not yet been modelled using CFD in the range of the CLC technology (10-1000kWth). The macroscopic models are effective tools for the simulation, design and optimization of CLC technologies until CFD codes for CLC process are more robust and affordable.

5.2 Reaction kinetics model The reaction rate of an OC is one of the most important properties in the CLC system as it dictates the solids circulation flux and bed inventory required in the system. Hence, kinetics studies are considered important for CLC reactor designs since both the fuel and air reactors involve gas solid reactions. The kinetics study of the gas-solid reaction, especially for the supported metal/metal oxide, is very complex due to the heterogeneity of the active species on the support surfaces. A particle reaction model is supposed to provide valuable information about the mass and heat transport processes inside the carrier particles during reduction and oxidation cycles. However, careful considerations are needed to include oxidation and reduction kinetics. Kinetics is also significantly influenced by the oxygen transport capacity and OCs’ physical and chemical properties. All the reactions involved in the reaction scheme should be included, and related kinetic parameters need to be determined [7,415]. Most of the experimental studies have been done in a TGA. In these experiments, the absence of external mass transfer control was checked. The temperature programmed reduction or oxidation (TPR, TPO) technique has been used for kinetics determination. Fluidized-bed and fixed-bed reactors have aslo been used, taking measures to reduce mass transfer limitations [416] or using a reactor model accounting for the mass transfer processes [417]. The activation energy, pre-exponential factor and the reaction order must be determined for each reaction involved in the reaction scheme. In general, the reaction order was found to be in the range 0.8-1.0. The 69 ACS Paragon Plus Environment

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activation energy seems to follow the tendency CH4 > H2 > CO>O2 [56]. Nevertheless, an important dispersion in the values for each type of OCs was observed. The interaction between the metal oxides and the support affects the activation energy. Reducing the affinity of the metal oxide with the support reduces the activation energy and makes reduction of the metal oxide easier [50,77,260]. Thus, it can be concluded that the kinetic parameters for every OC should be determined specifically. Extrapolation of kinetic models to various OCs is not yet viable [415]. The solid-state kinetics of the reduction and oxidation are mainly developed using nucleation and nuclei growth model (NNGM) and unreacted shrinking core model (USCM) [415]. Ryu et al. [418,419] assumed pseudo first-order reaction for the reduction of NiO by CH4, the rate constant, k for NiO/bentonite particle based on the Arrhenius equation is as follow:

k re = 73.63 exp(−

16349 ) RT

(14)

The reaction rate of OCs is the most important property of particles, because the solid circulation rate between AR and FR, and the amount of bed materials necessary in the two reactors is inversely proportional to the reaction rate of the OC [115]. The reaction rate varies widely depending on the particle size, temperature and pressure, the composition and type of the metal oxides and the support, and gas composition [280,299,420]. In order to describe the reaction rate, the degree of oxidation, X , called solid conversion, is introduced. X =

m a − m red m ox − m red

(15)

then the reaction rate could be presented by dX . The conversion rates of OCs can be increased dt

further by reducing the particle size or by increasing the reaction temperature. The OC must have enough reactivity to fully convert the fuel gas in the FR, and to be regenerated in the AR. Many studies of CLC focus on the OC development. Different OCs were testified in the TGA, fixed 70 ACS Paragon Plus Environment

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beds, FBs, and the particle properties including reactivity were measured, as mentioned in above sections. However, the reactivity data are usually obtained for one single operation condition, and limited information can be extracted for design and optimization purpose. Reaction rates under different operating conditions (e.g. different temperature, pressure and gas concentrations) should be determined [49,50,252,253,257,415,421,422]. The reactions with the OC in the air and fuel reactors can be considered as non-catalytic gas-solid reactions. The most frequently used models for predicting the time dependence of the solids conversion and the effect of operating conditions on the reaction rate are Changing Grain Size Model (CGSM), Shrinking Core Model (SCM), Nucleation and Nuclei Growth Models (NNGM) [7,68,394,423]. Gas-solid reaction kinetics becomes more complicated considering supported metals and metal oxides with added promoters and additives into the OC. Preparation methods of the OC may also affect physical properties, chemical composition and structure of the OC, which is closely related to reaction mechanism and chemistry. The rate-controlling step of the gas–solid reactions in CLC process is defined by observing the reaction rates at different set of experimental conditions as well as applying theoretical calculations. Kinetics parameters are mainly assessed using TGA, TPR, TPO and validated employing reaction data obtained from experimental study (e.g. fixed beds and FBs). Reaction kinetics is crucial for design, optimization and scale-up of the CLC process.

5.2.1 Changing Grain Size Model (CGSM) The Changing Grain Size Model (CGSM) [424,425] takes most of steps involved in gassolid reactions into account: gas film transfer, diffusion through the interstices among the grains (i.e. pores), diffusion through the product layer around the grain and chemical reaction on the interface in the grain. The grain size changes and the unreacted core size shrinks as reaction proceeds (See Figure 5) [426]. 71 ACS Paragon Plus Environment

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The local reaction rate in terms of the gas concentration in the pores can be expressed as

S 0 (rc / r0 ) 2 − rg = kC k s rc rc s g 1+ (1 − ) Ds r1

(16)

Effectiveness factor was considered to take into account the gas diffusion in the pores for the reaction rate, which is referred to as the diffusion-reaction model [426]. Szekely et al. [427] used specific conversion function for each time accounting for the specific resistance followed by summing the time of each one. Erri and Varma [428,429] observed that the reduction reaction of Ni-based OC particles supported on NiAl2O4 (40 wt% NiO, particle size up to 425 µm) were not limited by diffusion effects. The rate of diffusion thought the film, pores and product layer presents a negligible resistance based on most of studies related to OC particles, especially at high temperature (e.g. >800oC). The resistance to heat and mass transfer in the gas film and inside the particle together with the chemical reaction on the particle surfaces were also considered by some researchers for both oxidation and reduction reactions with different fuel gases (CH4, H2 and CO) and metal oxides (Ni, Cu, Fe, Mn and Co) for the CLC application, a diffusion coefficient dependent on solid conversion was used [56,240,430]. The rate controlling step can change in the course of the reaction [426,430]. If we ignore particle size and mass transport resistance and assume the solids conversion is uniform throughout the solid, the model corresponds to a Shrinking Core Model (SCM) in the grains considering only the chemical reaction term, which was used to calculate the kinetic parameters of the reduction and oxidation reactions for Cu, Ni, and Fe based OCs [56,190,295,430]. Particle properties (e.g. particle size, porosity and specific surface area, pore size and structure), active metal oxide content, support materials and promoters/additives, gas composition are all related to resistance to the mass or heat transfer, as well as to reaction rates.

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5.2.2 Shrinking core model When the resistance to gas diffusion in the unreacted particle is very high, the Shrinking Core Model (SCM) in the particle should be considered. Thus, a layer of the solid product is formed around an unreacted core inside the particle. The time required to reach a certain conversion is calculated in a similar way as that in the grain model, but replacing the radius of the grain by the radius of the particle [7] (see Figure ). Xiao and Song [431] investigated the kinetics of CaSO4 reacting with CO in more detail using the shrinking core model, kinetic parameters with high precision were obtained. In general, the unreacted shrinking-core models are classified as (1) the grain model, which describe the solid reactant phase as a juxtaposition of dense objects and (2) the pore model, which consider the porous solid as a collection of hollow objects [50,432]. The models of the above classification are different in calculating the surface area of the active sites. However, they are not totally independent and each of these models can possibly be derived from a more generalized mathematical form depending upon the appropriate structure of the solid material [50]. The unreacted shrinking core model incorporates particle size and pore structure of the solid reactant particles. According to this model, as the reaction progresses the metal–metal oxide interface moves towards the center of the grain, leaving behind a porous metallic/metal oxide product layer through which gaseous reactants and products diffuse [424,433]. The SCM has been used to calculate the kinetic parameters for reduction with CH4 and H2 and oxidation of millimetre-sized Ni-based particles. The particles are considered as a matrix of non-porous individual grains of uniform size [43,240,264,295,418,434]. The conversion vs. time curves indicate that reduction was controlled by chemical reaction, but oxidation was in the intermediate regime between chemical reaction control and product layer diffusion control [418]. 73 ACS Paragon Plus Environment

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According to the SCM, the reaction rate is first order in R p for chemical reaction control, second order for product layer diffusion control, and in the interval 1.5-2.5 order in R p for gas film diffusion. This fact was properly analyzed by Ishida et al. [264] for particle sizes from 1 to 3.2 mm. The SCM has been able to predict experimental data for smaller particles [261,418]. However, uncertainty arises about the suitability of the SCM in these cases because of the high porosity and small size of the particles, in the order of 100 µm. Assume a hypothetical first order gas-solid reaction shown by Equation 17:

A( g ) + bB ( s ) → cC ( g ) + dD ( s )

(17)

the unreacted-core shrinking model gives the following change of the core radius, rp :

−

drp dt

=

2 p 2 p g

r

R k

+

cC A / ρ B ( R p − rc )rc R p De

(18)

1 + ks

where the solids conversion, X s , is given by

1− X s = (

rp Rp

)3

(19)

The three terms in the denominator of Equation 18 represent the external gas film diffusion, the product layer diffusion and the chemical reaction, respectively.

5.2.3 Nucleation and nuclei growth models Many gas-solid reactions with formation of a solid product proceed by the nuclei formation and nucleation process, which is a dynamic process and practically initiates the reaction. There is an induction period for activating the solid phase to form nuclei before nucleation, the reaction rate increases as the number of nuclei increases during the first moment of reaction. The length of the induction period primarily depends on the gas-solid system and reaction temperature. The

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nuclei growth occurs due to the overlap of the nuclei and/or ingestion of a nuclei site. After the induction period, the reaction will occur uniformly over the solid surface, and the reaction front advances uniformly into the inner part of the grain. Thus, the conversion vs. time curves are characterized by a sigmoid behaviour, often described by the Avramie-Erofeev Model (AEM) [327]. For a particular gas-solid reaction, the nuclei growth rate is constant at a given temperature and composition of the gas phase. The relative rate of nucleation, nuclei growth and the concentration of the potential nucleus-forming sites (for generation of metallic nuclei) determine the overall conversion of the reaction. Error! Reference source not found. shows the possible steps during a gas-solid reaction following the nucleation and nuclei growth model. Nucleation effects are often significant in systems such as reduction of metallic oxides. These models have been widely used in reduction of Ni-based catalyst at low temperatures. The nucleation process is accelerated as the temperature increases, for example, the induction period was imperceptible at >340oC [427]. Sedor et al. [270] found that the reaction starts immediately at temperatures above 600oC without an induction period. It was believed that the reduction and oxidation of metal oxide proceed through nucleation and crystal growth, and nucleation and nuclei growth model was applied to describe the kinetics of OC based on the reaction rate controlling Avrami-Erofeev model [28,435], a reaction rate equation can be expressed as following [327,436,437]: dX s = k s' (T )C gn f ( X s ) dt

(20)

The general equation for the function of the solids conversion is f ( X s ) = υ (1 − X s )[− ln(1 − X s )](υ −1) / υ

(21)

The Avrami-Erofeev Model has been applied to the reduction and oxidation of Ni-based oxygen-carriers [327]. Temperature programmed reduction (TPR) or oxidation (TPO) was 75 ACS Paragon Plus Environment

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performed to obtain the kinetic parameters, using a heating rate of 10oC/min. The temperature was 200-500oC for the oxidation reaction, reduction reaction proceeds from 300oC to 600oC, above which the reduction is complete. The best fit was obtained from the Random Nucleation Model (RNM) with υ =1 (Equation 21). When υ =1, an induction period is not present. An equivalent expression to that for the RNM can be obtained from the Power Law Model (PLM) or a Modified Volumetric Model (MVM). These models have been used when the reaction occurs uniformly all through the particle, i.e. no diffusion resistance exists. Son and Kim [261] used the same time-conversion dependence for the Modified Volumetric Model (MVM) to obtain kinetic parameters for the reduction with CH4 of Ni- and Fe-based OCs. In a CLC system, higher temperatures in the range (900-1200oC) should be necessary to get high electrical efficiency although temperatures of about 600-800oC could be sufficient for industrial process [65,73]. For the temperature range involved in CLC system, the nucleation process could be relatively fast and of low relevance regarding the conversion of the bulk solids. When the nucleation occurs rapidly over the entire solid surface, the models which deal with interfacial chemical reactions can be applied (e.g. SCM) [427]. Actually, both the SCM and RNM have been shown to fit the same experimental data reasonably well [270]. The nucleation model only emphasizes on the chemical mechanism and kinetics of the gas–solid reactions, but it does not consider the morphological factors, which may be equally important in determining the kinetics. It has been shown that the reaction rate of the gas–solid reactions can be influenced by the grain size for a particle diameter greater than 10µm [438]. Specifically for the porous particles, the effect of particle size and its state during the reaction is very important in determining reaction rates. To take into account the various aspects discussed above, literature studies investigated both the nucleation and nuclei growth model and unreacted

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shrinking core model, in order to describe the gas–solid reaction kinetics during both the reduction and oxidation cycles of the CLC processes.

5.3 Process simulations and mathematical models for CLC systems Aside from the need of technological improvement of the CLC system, there are some other important issues that have to be addressed for commercial implementation of an integrated CLC plant with CO2 capture. Mass and energy balance, thermodynamic analysis, process synthesis and simulation have been conducted in order to optimize operating parameters and process integration, and minimize cost [49]. Most of these studies consider methane, syngas (CO and H2) or some carbonaceous solids as possible fuels with Ni-, Fe-, Cu-based OCs due to their favorable cost, reactivity and stability.

5.3.1 CLC of syngas Syngas is from gasification of fossil fuels (e.g. coal and petroleum coke) or biomass. After cleaning, syngas of certain gas composition and impurities passes through the FR, where syngas reacts with metal oxides to form CO2 and H2O. This configuration is called Integrated Gasification Chemical Looping Combustion (IG-CLC) [439,440]. There are several process options for this plant configuration: 1) IG-CLC with gas turbine system and heat generation (IGCLC-GT); 2) IG-CLC with steam turbine system (IG-CLC-ST); 3) IG-CLC with combined cycle (IG-CLC-CC). In refineries, refinery gas may be used instead of syngas. A simplified scheme of an integrated CLC-based power plant is shown by Brandvoll and Bolland [99]. In this configuration, the outlet gas stream of the AR drives the gas turbine before it is routed to heat recovery steam generator (HRSG) to produce steam, which is used in the steam turbine to generate extra power, whereas the exhaust of the FR drives the CO2 turbine. After circulating through the CO2 turbine, the gas is further cooled to nearly ambient temperature in order to condense the water, leaving almost pure CO2 (greater than 90%). Finally the 77 ACS Paragon Plus Environment

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concentrated CO2 stream is compressed for transportation and sequestration. Spallina et al. [91] simulated different cycle strategies using a numerical model for the heat management in a dynamically operated packed bed reactor with syngas as the fuel for CLC. Different layouts have been compared in order to discuss the effects on the axial solid temperature and solid conversion profiles, the fuel conversion and the reactor outlet conditions. Several studies analyzed a CLC system of two reactors using syngas from coal gasification for electricity and hydrogen coproduction. The authors already conducted an exergy analysis for the design [441]. Xiang et al. [442] analyzed an IGCC design with electricity and hydrogen co-production using a Fe-based OC in a three-reactor CLC system or using a combination of Ni- and subsequent Fe-based OCs [443]. Both studies used the oxidizer (steam reactor) for the production of hydrogen. To apply fluidization within the steam and air reactor, the feed gas streams are injected at a pressure of 40 bar. Experimental investigations concerning the CLC design mostly use steam or nitrogen to seal the interconnections among the reactors. Sorgenfrei and Tsatsaronis [444] have done design and thermodynamic evaluation of an IGCC process using syngas chemical looping combustion (IGCLC-CC) of three reactors (multistage moving bed for reducer, and FBs for oxidizer and combustor) for generating electricity, as well as for CO2 capture and H2 production. The syngas from coal gasification is cleaned using high-temperature gas desulfurization (HGD) and the Febased OC is selected. The gas turbine downstream combustor is expected to exhibit low NOx emissions due to the high ratio of water in the combustion chamber. BGL slagging gasifier and the Shell entrained flow gasifier were employed to evaluate the net efficiency. The option of using a CO2 turbine after the FR was also investigated. It was found that the best net efficiency of 43% (based on HHV) can be obtained using a BGL gasifier without a CO2 turbine at an AR temperature of 1000oC, including CO2 compression for transport and storage. A sensitivity analysis on the AR outlet temperature was compared based on the power output and the net 78 ACS Paragon Plus Environment

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efficiency. Please note that both the HGD and CLC systems are not under commercial operation so far. Energetic and exergetic analyses of CLC gas turbines fueled with syngas were conducted in previous study. Second-law analysis of syngas-CLC-CC and simulation results of syngas-CLC with different OCs were provided [445]. However, energy savings in the capture of CO2 were not quantified. There is also a potential of integrating CLC with combined cycles (syngas-CLCCC), like the analysis of the trigeneration [446] . Combining CLC and IGCC (IG-CLC-CC) in particular, could achieve highly efficient power generation together with nearly zero greenhouse gas emissions. An accurate thermodynamic modeling is required for the optimization of several design parameters. Simulations to evaluate the energetic efficiency of this CLC based power plant under diverse working conditions have been carried out, and a comparison of a conventional integrated gasification power plant with precombustion capture of carbon dioxide has been made. Two different syngas compositions have been tried to check its influence on the results. The energy saved in carbon capture and storage by IG-CLC-CC process is found to be significant, inducing an improvement of the overall power plant thermal efficiency of around 7% in some cases [447].

5.3.2 CLC of NG NG-CLC is similar to syngas CLC systems except replacing syngas with NG. There are also several process options for this plant configuration: 1) NG-CLC with gas turbine system and heat generation (NG-CLC-GT); 2) NG-CLC with steam turbine system (NG-CLC-ST); 3) NGCLC with combined cycle (NG-CLC-CC) [97,398,439]. Consonni et al [208] analyzed a methane-fueled combined cycle of gas and steam turbines with CLC configurations (NG-CLC-CC). Song et al. [190] studied the viability of nonconventional OC for CLC of methane. Peltola et al.[448] provided a model-based evaluation at 79 ACS Paragon Plus Environment

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pre-commercial stage of a combined cycle and Zhang et al. [449] gived a theoretical exploration of chemical-looping hydrogen (CLH) generation with methane as fuel, which is a CLC variation [88]. Kolbitsch et al. [401] introduced a model to simulate the interconnected fluidized bed CLC system using Ni-based OCs with air and fuel reactors in fast and turbulent fluidization regime, respectively. One of the assumptions was the reacting gas was only in contact with a defined fraction of well-mixed solids in each reactor. Hydrodynamic profile of the reactor was described only by prescribed solids concentration profile along the reactor height. Energy and solids were balanced globally across the whole reactor and plug flow was assumed for the gases. Abad et al. [398] built a model for the FR using a Ni-based OC and CH4 as the fuel, including the conversion of the OC and the gas composition at the reactor exit, the axial profiles of gas concentrations, the fluid dynamic structure of the reactor and combustion efficiency. Although the model was validated in a 120 kWth CLC unit, no systematic process simulation was available. An analysis using Fe-based OCs in a three-reactor configuration focused on a NG fueled CLC integrated in a combined cycle (NG-CLC-CC) [450]. Wolf and Yan [451] analyzed a NG fueled three-reactor CLC system using Ni-based OCs in a steam-injected gas turbine (STIG). Naqvi et al. [97] presents thermodynamic cycle analysis of a NG-CLC power plant for combined cycle and steam cycle, a steady-state model was developed for the gas-solid reactions occurring in the reactor systems with energy consumption of CO2 capture being taken into account. Effects of exhaust recirculation for preventing coking formation and incomplete fuel conversion were also investigated. The results show that an optimum efficiency of 49.7% (LHV) can be achieved under given conditions with a NG-CLC-CC at zero emissions level. The NG-CLC-ST is capable of achieving 40.1% efficiency (LHV) with zero emissions. Fernandez and Alarcon[69] presented a process scheme based on fixed-bed reactors for carrying out the CLC of NG at high pressure with ilmenite as the OC. High pressure and high 80 ACS Paragon Plus Environment

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temperature operations can achieve highly efficient power cycles at at the price of complex heat management, and switching valves able to function at very high temperatures are required. The authours have done a preliminary conceptual design of 500MWth capacity with NG as fuel and ilmenite as OC. The design has shown that a minimum of five reactors, 10 m long, with an inner diameter of 6.7 m, would be required to fulfil the overall process, assuming cycles of 10 min with maximum pressure drops per stage of less than 6%. These results demonstrate the potential of the CLC technology for power generation in combination with CO2 capture. The design and operation strategies of dynamically operated packed-bed reactors of a CLC system were reported [70]. This CLC system was included in an integrated gasification combined cycle (IGCC) for electric power generation. The CLC reactors employed ilmenite as the OC and operated sequentially across the following phases: oxidation, purge, reduction and heat removal. The results indicated that 14–16 units with 5.5 m of internal diameter and 11 m of length are required for continuous operation of a 350–400 MWe coal-fired power plant. Penthor et al. [290] tested the Cu-based OC in a 120 kW chemical looping pilot plant with NG as the fuel, variations of several process parameters like temperature, fuel power, solids inventory and solids circulation rate have been performed. The copper particles showed good performance regarding conversion of CO and H2 (almost full conversion) but only moderate conversion of CH4 (up to 80%) was achieved. Continuous analysis of the OCs revealed an initial decay of active CuO content caused by attrition on the external surface of the particles. The CuO content stabilized after 30 h of operation at around 9 wt% and no further decrease was observed. Kallen et al.[304] carried out CLC experimetnts in a FB reactor with continuous circulation of solids designed for a thermal power of 300 W, using syngas and NG as the fuels. The OCs used were combined oxides of iron, manganese and silicon with varying composition of these three materials. It was shown that full conversion of syngas and above 95% conversion of NG above 900 oC have been achieved. 81 ACS Paragon Plus Environment

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Zeng et al. modeled a countercurrent gas-solid flow pattern using Aspen Plus® software to provide insight into the operating conditions of a CLC system with NG as the fuel. The sensitivity analysis shows that, at 900 oC and 1 atm, a Fe2O3 to CH4 molar ratio of greater than 2.6 is essential for complete conversion of methane in a countercurrent moving-bed reactor, as any values lower than 2.6 will result in unconverted fuel exiting with the reducer flue gas stream. The model was validated experimentally using a bench-scale countercurrent moving bed demonstration unit. However, the experimental gas analysis results show a higher CH4 content and lower CO and H2 concentrations compared with the equilibrium composition from the simulation results, indicating kinetic limitations. Further kinetic simulation is needed to address the discrepancy between the multistage equilibrium model and experimental study [178].

5.3.3 Chemical looping combustion with solid or liquid fuels Solid fuels mainly refer to as coal and biomass [278,452–455], liquid fuels may be various carbonaceous fuels from petroleum and refinery plants (e.g. kerosene, heavy oils, crude oil) [456]. For CLC with solid and liquid fuels, there are also interation between gases and solids in the AR and FR, similar to NG and syngas CLC systems from this point. However, the CLC process with solid and liquid fuels is generally more complex than that using gaseous fuels in terms of issues related to char and pollutants. A model based on semi-empirical correlations was established to simulate the performance of the 1 MWth CLC rig with coal as the fuel and FR in the fast fluidization regime, using ilmenite as OC. The model considered the reactor fluid dynamics, the coal conversion and the reaction of the OC with evolved gases from coal. The efficiency of a carbon separation system (i.e. two cyclones and one carbon stripper) was analyzed for the FR performance [457]. A carbon stripper between air and fuel reactors increased carbon conversion for coal direct chemical looping combustion (CDCL) system [457–459]. The effects of the carbon separation efficiency, the solids inventory and temperature in the FR, ratio 82 ACS Paragon Plus Environment

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of OC to fuel, OC reactivity, coal properties (e.g. size, reactivity) on the carbon capture and combustion efficiency were carried out. For CDCL, the FR temperature and the efficiency of the carbon separation system (i.e. two cyclones and one carbon stripper) are key parameters to improve the performance of the process. Water-gas shift reaction (WGS) was also evaluated [459]. It was shown that the carbon capture efficiency of 98.6% could be reached if the FR temperature of 1100oC, a solid inventory of 1000kg/MWth and a carbon separation efficiency of 98% were met [459]. Aspen Plus® simulations and mathematical modeling of a CLC process consisting of three reactors (i.e. fuel reactor, oxidizer and combustor) indicated that the incorporation of a small amount of copper in the Fe-based OC led to increased hydrogen yield and process efficiency [308]. Li et al. [452] simulated and analyzed biomass direct chemical looping (BDCL) based on Aspen Plus® process simulation. The process also consists of three chemical looping reactors (i.e. the reducer, oxidizer, and combustor) for CO2 capture, H2 production and power generation, with a Fe-based OC circulating among the reactors. The four oxidation states of iron and their distinct thermodynamic properties give rise to a large degree of freedom in the process design and operating parameters. From the process simulation, the suitable reactor design, operating conditions and process configuration for the BDCL process were determined. Integration of CLC system with the existing plant and process equipment is very important to achieve high overall efficiency and low cost for power generation and CO2 capture [439,440]. Fan et al. [460] simulated three chemical looping processes, i.e. Syngas-CLC process, CDCL process, and Calcium Looping process (CLP) utilizing simple reaction schemes to convert carbonaceous fuels into products such as hydrogen, electricity, and synthetic fuels based on Aspen Plus® simulator. As operation with liquid fuels in chemical-looping combustion is still in its infancy, experience from the feeding process is limited [456,461,462]. Mixing the liquid fuels with 83 ACS Paragon Plus Environment

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superheated steam at the moment of injection is a viable option [463]. The steam will evaporate the liquid, resulting in a mixture of steam and gaseous fuel that enters the bed together. For solid and liquid fuels CLC system, experimental study in reaction mechanism and performance of the OCs, process simulation and modeling are still need. Solid and liquid fuels feeding [461,462,464–467], carbon separation and CO2 capture efficiency deserve more attention. More sophisticated and advanced process simulation and process analysis should be conducted considering the following issues: 1) the CLC plant configurations; 2) the possibility of integration with existing power or refinery plants; 3) the operating parameters; 4) energy and exergy efficiencies; 5) economic analysis.

5.3.4 Modelling of alternative CLC concepts As mentioned above, interconnected fluidized beds are the most promising CLC concepts based on previous research, other concepts have also been analyzed to carry out the CLC process although at lower scales. In these concepts, air and methane (or syngas) flow alternatively through the reactor, which can be formed by coated monoliths [146], packed bed [147] or FB [148]. Very few published papers have dealt with modeling of a CLC system as a whole. Noorman et al. [147] developed a one dimensional adiabatic packed bed reactor model where the solids are stationary and are alternately exposed to reducing and oxidizing conditions via periodic switching of the air and fuel gas streams. It was shown that the gas mass flow rate or the oxidation kinetics of the oxygen-carrier in a packed-bed reactor doesn’t affect the maximum temperature much. Furthermore, the CLC process can reach the cyclic steady state after only a small number of oxidation and reduction cycles, and continuous power generation can be reached with only two packed-bed reactors in parallel [147]. Modelling and simulations are required by fixing the time period at which the air and the fuel are fed, and the time whenever the flue gas is directed towards a CO2 storage unit in order to 84 ACS Paragon Plus Environment

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optimize the CLC system [146]. When this configuration is used for solid fuels, char gasification should be taken into account for optimization of carbon capture efficiency [147,148]. Mass and energy balance, temperature profile in the reactor, pressure drop, a high flow rate switching system should be considered for the cyclic process simulation and modeling.

6. Comparison of COE and CO2 cost for different CO2 capture technologies The CLC processes could be designed in a variety of ways. A conceptual design of a CLC power plant using NG as fuel with inherent CO2 capture (655 MWth) was conducted based on the CLC pilot study prototype at UBC (see Table 10). The scheme of the CLC conceptual design is shown in Figure 5, and comparison of cost of electricity (COE) and CO2 cost is shown in Table 10. Depending on the AR exhaust gas temperature, a combined cycle can be employed, similar to NGCC. For FR side, the combined cycle concept may also apply depending on the pressure and temperature of the FR, mainly determined by the OC properties and materials of construction. The CO2 expansion turbine generates electricity while also reduces the flue gas pressure, increasing power consumption for downstream CO2 compression. From Aspen Plus® simulation, electricity generated from the CO2 expansion turbine is much larger than required by CO2 compression from 1.05 to 150 bara (compressor and pump) for sequestration or EOR application. Power consumption for CO2 compression (1-73.9 bara using 5-stage compressors, and 73.9-150 bara using pumps) is 100 kWh/tonne CO2, while CO2 expansion turbine generate 312 kWh/tonne CO2. Heat Recovery and Steam Generation (HRSG) and Steam Turbine (ST) system can also be used to recover the heat from the hot flue gas from the FR. Depending on flue gas temperature and process requirements, there may be only HRSG for steam generation downstream the CO2 expansion turbine in Figure 5. Recall fluid catalytic cracking (FCC) in refinery, after gas-solid 85 ACS Paragon Plus Environment

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separation using multiple two-stage cyclones and swirl tubes, the exhaust gas (~700oC, 2-3 barg) from the catalyst regenerator is expanded through an expansion turbine to provide power for the process, then expanded flue gas is routed through a steam-generating boiler (or CO boiler) where CO is burned as fuel to provide heat needed. The same principle applies for FR flue gas. The levelized cost of electricity (COE) is comprised of three components: capital charge, operation and maintenance costs (O&M, fixed and variable), and fuel costs. Capital cost is generally the largest component of COE for coal power plants, fuel price is significant for a NGCC system because high price of NG. The electrical efficiency, COE and cost of CO2 avoided are compared for different CO2 capture technologies in order to identify the potential application of the CLC process for power generation and CO2 capture. Most of research and development work is based on operation at atmospheric pressure and reactor temperature < 1050oC for the CLC process. Higher energetic efficiency is obtained with higher reactor temperature and higher pressure operation by means of combined cycles for electricity generation and CO2 capture. This is especially important for the use of gaseous fuels because of the competition with NGCC using conventional CO2 capture. There are at least two ways to estimate the economics for a CLC process based on known processes and technologies: 1) CFB combustion using coal; 2) NGCC. The circulating fluidized bed combustion (CFBC) of solid fuels is a well established commercial technology, with boilers of sizes up to 600 MWe designed and 250-320 MWe already in operation. CFB units are best suited to low-value feedstocks such as high-ash coals or coal waste. They are very flexible in feedstocks and can also burn biomass. BFB have also been widely used in drying, gasification and combustion. The CLC process using solids fuels can get industrial experience from these similar technologies. In the EU project ENCAP, a first design of a 455 MWe CLC solid fuel power plant was made. A comparison to a similar FB combustion power plant indicated a very 86 ACS Paragon Plus Environment

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low efficiency penalty as well as a very low capture cost, 10 €/tonne of CO2 [59]. From Figure

5, we can see that the gas turbine, HRSG and steam turbine system (i.e. combine cycle) of the CLC system using NG as fuel is very similar to NGCC. One of the major difference is that the flue gas from a NGCC plant is divided into two gas streams, i.e. N2 stream (N2 and remaining O2) and CO2 stream (CO2 and water vapor) in CLC for power and heat generation. Hence, we can estimate the total cost for the CLC system based on NGCC processes. The advantages of a CLC system using NG as fuel over a NGCC process with advanced post-combustion CO2 capture technologies (e.g. amine processes) are mainly: 1) Inherent CO2 capture without additional CO2 capture process; 2) Post-combustion CO2 capture is energy intensive processes, and need to handle large volumetric flow rate of the flue gas for CO2 capture. Large volumetric flow leads to large equipment and pipe size, increasing capital cost. 3) CO2 concentration in the NGCC flue gas is low and CO2 capture efficiency is usually lower than other power plants; 4) The CLC system is flexible in feedstocks and has the potential to use gaseous, liquid and solids fuels; Retrofitting existing power plants for the CLC process is promising. Compared to NGCC, a CLC process have additional cost penalty as follows: 1) Additional equipment (e.g. air and fuel reactors); 2) Makeup of OCs cause extra cost; 3) Possible large power consumption for air (and fuel) compressors due to possible larger pressure drop through FBs than that in original NGCC plants. Cost estimations for the CLC of NG is based on NGCC with consideration of the extra equipment (e.g. AR and FR) and operating cost (e.g. OC) needed by the CLC process. The capital cost estimation of FB AR and FR is based on previous quotes on FBs using six-tenth 87 ACS Paragon Plus Environment

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factor (consider different capacity), Chemical Engineering Plant Cost Index and other factors (e.g. material and labor factors) to estimate the capital investment related to the added equipment. The capital investment estimation is confirmed by empirical equations for vertical pressure vessels in terms of reactor weight, dimension and materials. The OC cost is estimated based on information provided by Lyngfelt et al. [59], the added cost by the OC to the electricity is ~ 0.2-0.4 cents/kWhe. From Table 10, we can see CLC is a very promising technology for power generation and CO2 capture with potential low cost and high CO2 capture efficiency compared to other CO2 capture technologies. However, Table 10 doesn’t consider CO2 transportation and storage related costs. The electricity consumption for CO2 compression from 1 to 150 bara (compressor and pump) is about ~3% of the net power output from the power plant based on Aspen Plus® simulation. More information about the cost estimation and economic analysis for CLC processes is available in previous study [468]. Please note that the CLC process power efficiency assessments normally assume some ideal conditions such as good quality of OCs without OC deactivation, these idealized assumptions require clarification for proper efficiency calculations and cost estimation. Successful commercialization of integrated CLC power generation processes depends on development of OCs, development of specific process configurations and reactor design.

7. Conclusions and future work CLC holds great promises for power generation with inherent CO2 capture. The electrical efficiency can be up to 52-60% (LHV) based on Aspen Plus® simulation. Research and development of OCs is the key issue for a successful CLC system. Reactivity, thermal stability, mechanical strength, oxygen transport capacity, cost and toxicity are main areas to work on to improve performance and economics of the entire process. Interconnected fluidized beds are 88 ACS Paragon Plus Environment

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very promising technology for the CLC application. FBs provide uniform temperature, excellent gas-solid mixing, good mass and heat transfer, fast reaction rate and large throughput. Circulating solids transfer oxygen and heat from AR to FR to reach mass and heat balance. Low solids circulation rate leads to large solids conversion difference for the same bed inventory and possible low fuel reactor temperature, particularly for endothermic FRs, which reduces the reaction rate in the FR and may cause low conversion to CO2 and H2O. Too large solids circulation rate may pose difficulty in smooth fluid bed operations and may also cause larger particle attrition and entrainment. Solids circulation rate and solids inventory are important parameters for design of the CLC system. Hydrodynamics, mass and energy balance calculations, reaction kinetics and CLC modeling are all important for design, operation and scale-up of the CLC system. Existing CFBCs provide guidance and industrial experience for the CLC process, including reactor performance, process integration, cost estimation and economic evaluation. It is viable to retrofit an existing CFBC plant for a CLC power plant with inherent CO2 separation, avoiding costly post-combustion CO2 capture process (e.g. amine process). For syngas CLC, it is beneficial to integrate coal or biomass gasification unit to the CLC process to improve the process efficiency. NG-CLC system needs to compete with NGCC process. NGCC can reach high electrical efficiency up to ~ 60% (LHV), but the CO2 concentration in the flue gas is low, typically 3-8% by volume, NGCC with CO2 captured is not as efficient and cost effective as the NG-CLC process in terms of Aspen Plus® simulation and process analysis. Cost estimation and economic assessment of the NG-CLC process can rely on existing NGCC system. Coke formation can be minimized if the reactor operates at rather high temperature. However, sintering and agglomeration may occur at high temperature. Decomposition of metal oxides (e.g. CuO) affects the FR performance and more research is still needed in these areas. 89 ACS Paragon Plus Environment

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CLC process minimizes SOx and NOx emissions compared to other power generation processes due to relatively low FR temperature and inherent separation between the high temperature N2containing gas and the FR gas stream. Sulphur and nitrogen in the fuel may bind with OCs, they may also be released with the flue gas. The small volumetric flow rate of the flue gas from the FR eases the removal of these impurities. The main obstacle in commercialization of the CLC technology, as with all CO2 capture technologies, may be the lack of real incentives. Large efforts to scale up more established CO2 capture technologies may divert the interest from the CLC technology which is rather new and less well known. The measures and extra costs needed to reach adequate performance of CLC, CO2 capture efficiency and efficient process integration are not yet fully understood or solved, and there are still technology barriers and some limitations to achieve smooth and reliable operation for a large scale or commercial plant. Although more work is still needed, especially with respect to finding the best design of the CLC system, it is clear that the CLC technology provides a unique potential for avoiding the large costs and energy penalties inherent in gas separation. More experimental research should be done for the CLC hot unit to investigate performance of different types of OCs, reaction rate, conversion of solids and gases, gas yield, solids circulation rates, different operation conditions and process parameters (e.g. temperature and pressure, gas velocity and bed inventories). More detailed process simulation and modeling of CLC, as well as comparison with other power processes in technology and economics, are still required and need further investigation.

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Nomenclature: Ar

Archimedes number, -.

br C1 , C2

Stoichiometric factor for the fuel reaction, mol solid reactant/mol fuel. Constants, -.

Cg

Gas concentration, mol/m3

Cavoided

Cost of CO2 avoided, $/tonne CO2.

COEreference

Cost of electricity for reference plant without CO2 captured, $/kWh.

COEcapture

Cost of electricity for power plant with CO2 captured, $/kWh.

d

Stoichiometric factor in fuel combustion reaction with O2, mol O2 per mol fuel.

dp

Mean particle diameter, m.

D

Diameter of reactor, m.

Ds

Coefficient of gas diffusion in the product solid layer, m2/s

Ff

Molar flow of fuel gas, mol/s.

Fin

Input flow rate of gases and solids in Equations 11 and 12,mol/s;

Fout

Outlet flow rate of gases and solids in Equations 11 and 12,mol/s;

FO2

Molar flow of O2 in air reactor, mol/s.

Gs

Specific solids circulation rate per kg/(m2. s).

h

Enthalpy flow of gases and solids in Equations 11 and 12, kJ/mol;

Ht

Height of the reactor, m.

k re

Chemical reaction rate constant for reduction in Equation 14, 1/s.

ks

Chemical reaction rate constant based on surface area, mol1-n.m3n-2/s

k s'

Chemical reaction rate constant, (mol/m3)-n s-1

K CH 4

Equilibrium constant for CH4 conversion to CO2 and H2O;

K H2

Equilibrium constant for H2 conversion to H2O;

K CO

Equilibrium constant for CO conversion to CO2;

KC

Equilibrium constant for C conversion to CO2;

o

Circulation rate of fully oxidized OC, kg/s.

m oc

ma

Actual mass of OC in its partially oxidized state, Kg.

mox

Mass of fully oxidized OC, kg. 91 ACS Paragon Plus Environment

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mred

Mass of fully reduced OC, kg.

m ,n Me

Constants in Equation 1, -. Metal, -.

MeO

Metal oxides, -.

( M CO2 ) reference

Mass of CO2 generated per kWh for reference plant, tonne CO2/kWh;

( M CO2 ) capture

Mass of CO2 generated per kWh for plant with CO2 captured, tonne

M NiO

Molecular weight of metal oxide, g/mol.

n

Reaction order respect to gas, -.

Qox

Heat rate released from exothermic reaction in air reactor, kW.

Qre

Heat rate absorbed by endothermic reaction in fuel reactor, kW

Qremove, AR

Heat rate removed from air reactor, kW

Qremove, FR

Heat rate removed from fuel reactor, kW

R0 r0

Oxygen ratio of OCs, -. Initial grain radius, m

r1

Grain radius, m

rc

Unreacted core radius, m

− rg

Reaction rate of gas with solid, mol/(m3.s)

Re mf

particle Reynolds number based on minimum fluidization velocity, -;

Rp

Particle radius, m

S0

Specific surface area, m2/m3

T

Temperature, oC.

Tsin

Temperature sintering occurs, oC.

U

Superficial gas velocity, m/s.

U mf

Minimum fluidization velocity, m/s.

Ut

Terminal settling velocity, m/s.

X

Degree of oxidation, -.

Xs

Solid conversion, -.

x NiO

Mass fraction of metal oxide in fully oxidized OCs, -.

x, y

Constants in Equations 1 and 2, -. 92 ACS Paragon Plus Environment

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∆H Ro

Standard heat of reaction at 25oC and 1 bar;

∆H 1173 K

Heat of reaction at 900oC (1173K);

∆X f

Conversion difference of the gas for fuel reaction, -.

∆X O2

Conversion difference of the gas for air reaction, -.

∆X s ,FR

Solid conversion difference in fuel reactor, -.

∆X s

Solid conversion difference between the air reactor and reducer, -.

Greek letters

ρf

Density of fluid, kg/m3.

ρp φ

Particle density, kg/m3.

ω

υ Subscripts 1, 2 f ox red

Sphericity of particles, -. Void fraction of the bed, -. Avrami exponent indicative of the reaction mechanism and crystal growth dimension Air and fuel reactors. Fuel. Oxidizer. Reducer.

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Abbreviations: AR

Air reactor

AEM

Avramie-Erofeev Model

AJI

Air Jet Index

BFB

Bubbling fluidized bed

BFBC

Bubbling fluidized bed combustion

CCR

Capital charge rate

CFB

Circulating fluidized bed

CFBC

Circulating fluidized bed combustion

CFD

Computational fluid dynamic

CDCL

Coal direct chemical looping

CGSM

Changing grain size model

CLC

Chemical looping combustion

CLC-CC

Chemical looping combustion combined cycle

CLC-SC

Chemical looping combustion steam cycle

CLOU

Chemical-looping with oxygen uncoupling

CLP

Calcium Looping process (CLP)

CLR

Chemical looping reforming

COE

Cost of electricity

DCS

Distributed control system

DFB

Dual fluidized bed

EOR

Enhanced oil recovery

FCC

Fluid catalytic cracking

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FB

Fluidized bed

FR

Fuel reactor

GHSV

Gas hourly space velocity

GT

Gas turbine

HGD

High-temperature gas desulfurization

HRSG

Heat Recovery and Steam Generation

IGCC

Integrated gasification combined cycle

IG-CLC

Integrated gasification chemical looping combustion

IG-CLC-GT

IG-CLC with gas turbine system

IG-CLC-ST

IG-CLC with steam turbine system

IG-CLC-CC

IG-CLC with combined cycle

M

Mass of CO2 generated per kWh elecricity

MEA

Monoethanolamine

MVM

Modified Volumetric Model

NG

Natural gas

NGCC

Natural gas combined cycle

NG-CLC

Natural gas chemical looping combustion

NG-CLC-GT

NG-CLC with gas turbine system

NG-CLC-ST

NG-CLC with steam turbine system

NG-CLC-CC

NG-CLC with combined cycle system

NG-NiO-CLC-CC

NG-CLC with combined cycle and NiO as oxygen carrier

NNGM

Nucleation and nuclei growth model

OC

Oxygen carrier

OCM

Oxidative coupling of methane 95 ACS Paragon Plus Environment

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OXY

Oxygen fuel

PC

Pulverized coal

PLC

Programmable logic controller

TRCL

Three Reactor Chemical Looping

TPO

Temperature programmed oxidation

TPR

Temperature programmed reduction

PLM

Power law model

PSA

Pressure swing adsorption

Redox

Reduction and oxidation

RNM

Random Nucleation Model

SCM

Shrinking Core Model

SubC

Subcritical

ST

Steam turbine

TIT

Gas turbine inlet temperature

TSA

Temperature swing adsorption

UBC

University of British Columbia

USCM

Unreacted shrinking core model

U mf

Minimum fluidization velocity

Ut

Teminal velocity

YSZ

Yttria stabilized zirconium

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http://www.sciencedirect.com/science/article/pii/S1750583613000169 (accessed August 15, 2014). [475] D.Y.C. Leung, G. Caramanna, M.M. Maroto-Valer, An overview of current status of carbon dioxide capture and storage technologies, Renew. Sustain. Energy Rev. 39 (2014) 426–443. doi:10.1016/j.rser.2014.07.093. [476] Y. Takamura, J. Aoki, S. Uchida, S. Narita, Application of high-pressure swing adsorption process for improvement of CO2 recovery system from flue gas, Can. J. Chem. Eng. 79 (2001) 812–816. [477] M. Clausse, J. Merel, F. Meunier, Numerical parametric study on CO2 capture by indirect thermal swing adsorption, Int. J. Greenh. Gas Control. 5 (2011) 1206–1213. http://www.sciencedirect.com/science/article/pii/S175058361100096X (accessed August 18, 2014). [478] G. Xu, L. Li, Y. Yang, L. Tian, T. Liu, K. Zhang, A novel CO2 cryogenic liquefaction and separation system, Energy. 42 (2012) 522–529. http://www.sciencedirect.com/science/article/pii/S0360544212001624 (accessed September 3, 2014). [479] K. Atsonios, K.D. Panopoulos, A. Doukelis, A. Koumanakos, E. Kakaras, Cryogenic method for H2 and CH4 recovery from a rich CO2 stream in pre-combustion carbon capture and storage schemes, Energy. 53 (2013) 106–113. http://www.sciencedirect.com/science/article/pii/S0360544213001400 (accessed September 3, 2014). [480] H. Li, Y. Hu, M. Ditaranto, D. Willson, J. Yan, Optimization of Cryogenic CO2 Purification for Oxy-coal Combustion, Energy Procedia. 37 (2013) 1341–1347. doi:10.1016/j.egypro.2013.06.009. [481] C.-F. Song, Y. Kitamura, S.-H. Li, K. Ogasawara, Design of a cryogenic CO2 capture system based on Stirling coolers, Int. J. Greenh. Gas Control. 7 (2012) 107–114. http://www.sciencedirect.com/science/article/pii/S1750583612000151 (accessed August 15, 2014). [482] H. Rabiee, M. Soltanieh, S.A. Mousavi, A. Ghadimi, Improvement in CO2/H2 separation by fabrication of poly(ether-b-amide6)/glycerol triacetate gel membranes, J. Memb. Sci. 469 (2014) 43–58. http://www.sciencedirect.com/science/article/pii/S0376738814004773 (accessed September 3, 2014). [483] X. Wang, H. Chen, L. Zhang, R. Yu, R. Qu, L. Yang, Effects of coexistent gaseous components and fine particles in the flue gas on CO2 separation by flat-sheet polysulfone membranes, J. Memb. Sci. 470 (2014) 237–245. [484] P.K. Kundu, A. Chakma, X. Feng, Effectiveness of membranes and hybrid membrane processes in comparison with absorption using amines for post-combustion CO2 capture, Int. J. Greenh. Gas Control. 28 (2014) 248–256. doi:10.1016/j.ijggc.2014.06.031. [485] J. Franz, P. Maas, V. Scherer, Economic evaluation of pre-combustion CO2-capture in IGCC power plants by porous ceramic membranes, Appl. Energy. 130 (2014) 532–542. http://www.sciencedirect.com/science/article/pii/S0306261914001573 (accessed August 18, 2014). 132 ACS Paragon Plus Environment

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[486] M. Ishida, H. Jin, A novel combustor based on chemical-looping reactions and its reaction kinetics., J. Chem. Eng. JAPAN. 27 (1994) 296–301. doi:10.1252/jcej.27.296. [487] K.-S. Song, Y.-S. Seo, H.K. Yoon, S.J. Cho, Characteristics of the NiO/Hexaaluminate for chemical looping combustion, Korean J. Chem. Eng. 20 (2003) 471–475. [488] Ø. Brandvoll, L. Kolbeinsen, N. Olsen, O. Bolland, Chemical Looping Combustion – Reduction of nickel oxide / nickel aluminate with hydrogen,The Norwegian Research Council program, 2003. [489] T. Cho, P., T. Mattisson, Reactivity of iron oxide with methane in a laboratory fluidized bed-application of chemical-looping combustion, in: Proc. 7th Int. Conf. Circ. Fluid. Beds, Niagara Falls, Ontario, Canada, 2002. [490] T. Mattisson, A. Lyngfelt, P. Cho, Possibility of using iron oxide as an OC for combustion of methane with removal of CO2-application of chemical-looping combustion, in: Proc. 5th Int. Conf. Greenh. Gas Control Technol., Cairns, Australia, 2000. [491] H. Tian, Q. Guo, J. Chang, Investigation into decomposition behavior of CaSO4 in Chemical-Looping Combustion, Energy & Fuels. 22 (2008) 3915–3921. [492] T. Bolhar-Nordenkampf, J., Proll, P. Kolbitsch, H. Hofbauer, Comprehensive modeling tool for chemical looping based processes, Chem. Eng. Technol. 32 (2009) 410–417. [493] S. Cooper, C.J. Coronella, CFD Simulation of particle mixing in a binary fluidized bed, Powder Technol. 151 (2005) 27–36. [494] S. Vaishali, S. Roy, P.L. Mills, Hydrodynamic simulation of gas-solids downflow reactors, Chem. Eng. Sci. 63 (2008) 5107–5119. [495] O. Gryczka, S. Heinrich, N.G. Deen, M.V. Annaland, J.A.M. Kuipers, M. Jacob, et al., Characterization and CFD-modeling of the hydrodynamics of a prismatic spouted bed apparatus, Chem. Eng. Sci. 64 (2009) 3352–3375. [496] J. Ding, D. Gidaspow, A bubbling fluidization model using kinetic theory of granular flow, AIChE J. 32 (1990) 523–538.

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Figure Captions Figure 1. Schematic representation of the chemical-looping combustion process Figure 2. Schematic diagram of the proposed CLC reactor (cold flow) [68] Figure 3. Design procedure of a CLC reactor system[469] Figure 4. Effects of active NiO content on solids circulation flux and fractional conversion rate using NG as fuel [468] Figure 5. Scheme of changing grain size model (CGSM) (adapted from [327]) Figure 6. Scheme of Shrinking Core Model (SCM) (adapted from [327]) Figure 7. Scheme of nucleation and nuclei growth model (adapted from [327]) Figure 8. Schematic of a CLC process for power generation and CO2 capture

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Table 1. Different CO2 capture strategies [2,470–473] 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48

CO2 Capture Descriptions Technology Post combustion 9-14 v% CO2 in flue gas; Amine based solvents, capture for power plants chilled ammonia, sodium hydroxide and potassium carbonate solutions, membrane and cryogenic separation.

Pre-combustion capture

Mainly for separation of CO2 and H2; Physical solvents, Benfield, PSA and cryogenics.

Oxy-fuel combustion for power plants

Recycle of CO2 gas to adjust flame temperature.

Chemical looping combustion Cement Industry

Steel and iron production H2 and NH3 production Natural gas processing Oil refining

CO2 capture from air

Advantages

Disadvantages

Analogous to wet flue gas desulphurization (FGD) and are widely used; Without significant changes in the upstream system, plug and play mode and great flexibility.

Low CO2 partial pressure and need to handle large volume of gas, large equipment size and high capital costs; Regeneration is energy intensive; Contamination of CO2 solvents by sulfur and nitrogen oxides. No long continuous operating times to confirm reliability of CO2 capture technologies for IGCC plants. Large energy and cost penalty because of air separation unit.

Larger CO2 partial pressure; consumption of regeneration.

Low

energy

Low gas flow rate; Retrofitting existing boilers or building a new and more compact boiler; Promising if low cost air separation technology becomes reliable (e.g. oxygen transport membrane). Two step combustion technology. Avoid mixing of air and fuel gas with inherent CO2 separation, and power and heat generation. 14-33v% CO2 concentration in calciner flue gas; High CO2 concentration and suitable for CO2 Amine based solvents, membrane, cryogenics. capture. ~27 v% CO2 concentration; Amine based solvents, membrane, cryogenics. Water gas shift reaction and CO2 capture (e.g. Benfield and PSA) Usually amine based solvents for CO2 capture

A bit higher CO2 concentration and suitable for CO2 capture, has been widely applied High purity CO2 capture and H2 production

CO2 is emitted from fired heaters, power generation, utility, and fluid catalytic cracking regenerator, etc; Chemical solvents or oxy-fuel combustion. Direct CO2 capture from air; Chemical solvents.

CO2 concentration may be high for CO2 capture for some cases (oxy-fuel combustion is feasible).

Meet CH4 pipeline purity

Directly reduce CO2 concentration in air

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Lack long operating time for OCs No low grade heat from cement plant for regeneration; Combine heat and power plant and oxy-fuel combustion are the options See power plant flue gas PSA is not efficient enough to separate H2 and CO2 Captured CO2 is usually not sequestrated Low CO2 concentration may not be economically viable for CO2 capture for some cases. Not economically viable

Energy & Fuels

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Table 2. CO2 capture technologies used [2] 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48

CO2 capture technologies Physical solvents Rectisol

Solvent/sorbent name

Process conditions and descriptions

Low temperature methanol

-10/-70 oC, > 2 MPa

Selexsol

Dimethyl ether of polyethylene glycol

-40oC, 2-3 MPa

Puisol Fluor Solvent Chemical solvents MEA Amine Guard (MEA) MDEA Hindered amine (Flexsorb/KS-1/2/3) Benfield and versions

n-methyl-2-pyrolidone (NMP) Propylene carbonate

-20/+40 oC, > 2 MPa Below ambient temperature, 3.1-6.9 MPa Recovery efficiency over 90%. ~ 40 oC, ambient-intermediate pressures, ~ 40 oC, ambient-intermediate pressures

Monoethanolamine and chemical inhibitors Monoethanolamine and chemical inhibitors Methyl diethanolamine Amine based

Developer/Licensor/Authors Lurgi and Linde, Germany; Lotepro Corporation, USA. Union Carbide, USA; UOP Lurgi, Germany Fluor, EI Paso, USA Dow Chemicals, USA Union Carbide, USA Exxon, USA; MHI, Japan

K2CO3+catalysts Lurgi and Catarcab with arsenic trioxide

70-120oC,2.2-7 MPa

Chilled ammonia

Chilled ammonia

~2-10 oC, ambient-intermediate pressures

Lurgi, Germany, Eickmeyer and Associates, USA, Giammarco Vetrocoke, Italy Alstrom

Sulfinol

Mixture of DIPA or MDEA, water, and tetrahydrothiophene (DIPAM) or diethylamine Mixture of methanol and MEA, DEA, diisopropylamine (DIPAM) diethylamine

>0.5 MPa (mixed physical and chemical solvent)

Shell, Netherlands

5/40oC, > 1 MPa (mixed physical and chemical solvent)

Lurgi, Germany

Zeolitic molecular sieves, activated carbon, silica gel, activated alumina, calcium oxides, hydrotalcites and lithium zirconate Lime, or NiO, CuO, Fe2O3 with support materials

~ ambient temperature, 10-30 bar; based on a physical binding of gas molecules to adsorbent material; high adsorption efficiency achievable (>85%). ~ 600-1000oC, ambient-intermediate pressure

[2,474–477]

Refrigeration enables direct production of liquid CO2; suitable for high concentration (>90%), high pressure gases. Various types of membrane materials are available including polymers, metals and rubber composites

Low temperature operation, refrigeration energy consumption is large; Water has to be removed before the gas stream is cooled. Varying temperature and pressure; Low gas throughputs requiring multistage operation or stream recycling.

[2,478–481]

Amisol Others Adsorption (e.g. PSA and TSA) Chemical looping Cryogenics Membrane

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[28,49]

[2,482–485]

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Table 3. Comparison of CO2 capture cost in a 500MWe plant [101] Items Net plant capacity (MWe) Net plant capacity (kWh/yr) CO2 emission rate (gCO2/kWh) SO2 emission rate (g SO2/kWh) NOx emission rate (g Nox/kWh) CO2 sequestrated (tonne/yr) CO2 sequestrated (g/s) Cost of electricity ($/MWh) CO2 cost ($/tonne CO2 avoided)

Reference plant 462 3.33E+09 941 2.45 0.45 49.2 -

With CO2 capture 326 2.35E+09 133.0 0.0003 0.58 2.58E+06 1.12E+05 97 59.2

Note: Based on coal fired power plant with and without MEA based CO2 capture

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Table 4. Chemical looping combustion reaction schemes (Ni-based OC) Oxidation reaction

Ni oxidation

2Ni + O2(g) = 2NiO

∆H1173K = -469.4 kJ/mol

OC reduction reactions

CH4 oxidation Partial CH4 oxidation Partial CH4 oxidation H2 oxidation CO oxidation Water splitting

CH4(g) + 4NiO = 4Ni + CO2(g) + 2H2O(g) CH4(g) + 2NiO = 2Ni + CO2(g) + 2H2(g) CH4(g) + NiO = Ni + CO(g) + 2H2(g) H2(g) + NiO = Ni + H2O(g) CO(g) + NiO = Ni + CO2(g) Ni + H2O(g) = NiO + H2(g)

∆H1073K = 139 kJ/mol ∆H1073K = 165 kJ/mol ∆H1073K = 212 kJ/mol ∆H1073K = -13 kJ/mol ∆H1073K = -47 kJ/mol ∆H1073K = 13 kJ/mol

Reactions catalyzed by Ni

Steam reforming Water gas shift Dry reforming Methane decomposition Carbon gasification by steam Carbon gasification by CO2 Methanation reaction Methanation reaction

CH4(g) + H2O(g) = CO(g) + 3H2(g) CO(g) + H2O(g) = CO2(g) + H2(g) CH4(g) + CO2(g) = 2CO(g) + 2H2(g) CH4(g) = C + 2H2(g) C + H2O(g) = CO(g) + H2(g) C + CO2(g) = 2CO(g) CO(g) + 3H2(g) = CH4(g) + H2O(g) CO2(g) + 4H2(g) = CH4(g) + 2H2O(g)

∆H1073K= 225 kJ/mol ∆H1073K= -34 kJ/mol ∆H1073K= 259 kJ/mol ∆H1073K= 89 kJ/mol ∆H1073K= 136 kJ/mol ∆H1073K= 170 kJ/mol ∆H1073K= -225 kJ/mol ∆H1073K = -191 kJ/mol

Combustion reaction

CH4 combustion

CH4(g) + 2O2(g) = CO2(g) + 2H2O(g)

∆H1073K = -802 kJ/mol

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1 2 Table 5. Heat of reactions and equilibrium constants for CLC redox system with different oxygen carriers 3 CH4 H2 CO C O2 4 Redox system o o o o o ∆H 1173 K ∆H R ∆H 1173 K K CH 4 ∆H R ∆H 1173 K K H 2 ∆H R ∆H 1173 K K CO ∆H R ∆H 1173 K K C ∆H R 5 6 NiO/Ni 156 136 7.00E+11 -2.1 -14 1.61E+02 -43 -47 1.24E+02 86 74 5.49E+05 -479 -469 7 CuO/Cu -179 -207 2.55E+26 -86 -100 7.02E+05 -127 -133 5.40E+05 -82 -97 1.05E+13 -312 -298 8 9 CuO/Cu2O -239 -278 9.92E+31 -101 -118 1.75E+07 -142 -151 1.35E+07 -112 -133 6.53E+15 -282 -262 10 Cu2O/Cu -120 -137 6.56E+20 -71 -82 2.81E+04 -112 -115 2.16E+04 -52 -62 1.68E+10 -341 -333 11 Fe O /Fe O 149 184 4.75E+22 -3.8 -2.1 8.21E+04 -45 -35 6.31E+04 82 98 1.43E+11 -476 -493 2 3 3 4 12 13 Fe2O3/FeO 351 305 3.25E+10 47 28 7.46E+01 5.5 -5.0 5.73E+01 183 159 1.18E+05 -577 -554 14 Mn O /MnO -48 -70 1.11E+25 -53 -66 3.21E+05 -94 -99 2.46E+05 -16 -29 2.18E+12 -377 -366 2 3 15 -397 -416 1.98E+34 -140 -152 6.59E+07 -182 -185 5.07E+07 -191 -202 9.23E+16 -203 -193 16 Mn2O3/Mn3O4 17 Mn3O4/MnO 126 103 2.62E+20 -10 -22 2.24E+04 -51 -55 1.72E+04 71 58 1.06E+10 -464 -453 18 Co O /Co 107 100 1.40E+15 -14 -23 1.08E+03 -55 -56 8.26E+02 62 56 2.46E+07 -455 -451 3 4 19 -18 11 4.37E+34 -46 -45 8.04E+07 -87 -79 6.18E+07 -1.1 11.7 1.37E+17 -392 -407 20 Co3O4/CoO 21 CoO/Co 149 129 4.46E+08 -3.9 -16 2.55E+01 -45 -49 1.96E+01 82 71 1.39E+04 -476 -466 22 CaSO /CaS 162 145 1.37E+11 -0.7 -12 1.07E+02 -42 -45 8.21E+01 89 79 2.43E+05 -482 -474 4 23 o o 24 Note: 1) Unit of ∆H R and ∆H 1173 K is kJ/mol gas or C.; 2) Equilibrium constants (i.e. K CH 4 , K H 2 K CO , K C ) are obtained at 900 C (i.e. 1173K); 3) All 25 thermodynamic data are from HSC Chemistry 6.1® software. 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 139 45 46 ACS Paragon Plus Environment 47 48

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1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48

MeO (wt%) 100% 80% 60% 40% 20% 10%

CaSO4/Cas 47.01 37.61 28.21 18.80 9.40 4.70

Table 6. Oxygen ratio (%) for different oxygen carriers NiO/Ni CuO/Cu Fe2O3/Fe3O4 Fe2O3/FeO 21.42 20.11 3.34 10.02 17.13 16.09 2.67 8.02 12.85 12.07 2.00 6.01 8.57 8.05 1.34 4.01 4.28 4.02 0.67 2.00 2.14 2.01 0.33 1.00

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Mn3O4/MnO 6.99 5.59 4.20 2.80 1.40 0.70

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Table 7. Summary of investigation on oxygen carriers (metal oxides) for chemical looping combustion Oxygen carrier particles NiO, NiO(60%,80%)+YSZ, Fe2O3(60%)+YSZ o Tsin : 1300 C

d p =1.3, 2.0, 2.8 mm NiO(20,30,40%)+hexaaluminate (NiO+LaAl11O18) (NiO+NiAl2O4) o Tsin : 1000 C NiO (59%)+ bentonite (Al3SiO2) o 3 Tsin : 900 C, ρ p :4038 kg/m , ρ b : 1319 kg/m3, ω : 0.69, d p : 91 µm

NiO(59%)+bentonite o 3 Tsin : 900 C, ρ p :4080/m

ρ b : 1407 kg/m3, ω : 0.59~0.714

d p : 400 µm NiO+NiAl2O4 d p : 300~500 µm , 600~1000 µm , 1200~1700 µm , 2000~3150 µm , Ni-Al-O (Ni/Al=0.5~2.25)= NiO(6%~100%) Ni-Mg-Al-O (Ni/Mg=1), (Ni+Mg)/Al=1) o Tsin : 1000 C NiO(33%)+Al2O3, Mn2O3(29.4%)+Al2O3, Co3O4(34.8%)+Al2O3, CuO(33.8%)+Al2O3 o Tsin : 550 C

Experimental conditions Reactor: TGA Reduction: H2 550 ~ 950 o C Oxidation: Air 1000 o C Pressure: atmospheric pressure

Remarks NiO/YSZ is suitable material, and the reaction temperature is the strongest factor in the reduction.

Reactor: TGA Gases: H2 (5.6%)+Ar (reduction), Air(oxidation)

Showed good reduction and oxidation properties.

[487]

Reactor: TGA Gases: CH4 (5.04%) (reduction), Air(oxidation) Gas flows: 100ml/min (reduction), 100ml/min(oxidation) Temperature: 650~1000o C Pressure: atmospheric pressure Reactor: A fixed bed Gases: 13% CH4 (reduction), 8.6% O2(oxidation) Gas flows: 2.3L/min (reduction), 2.2L/min(oxidation) Temperature: 500~1000 o C Pressure: atmospheric pressure Bed materials: 40g Reactor: fluidized/fixed-bed reactor Gases: H2 Temperature: 850~ 560 o C Pressure: atmospheric pressure Activity run (under constant temperature): Gases: 50 v% CH4 +He (reduction), 50 v% CH4 +H2O (reduction), 20% (v%) O2+He(oxidation) Gas flows: 20 cm3/min (reduction), Temperature: 800 o C Pressure: atmospheric pressure Bed mass: 2g

Carbon deposition, reduction kinetics and regenerative ability were examined; 900 o C is the appropriate for reduction and avoiding carbon deposition

[249]

Carbon deposition, sintering and lump of NiO+bentonite take place at 1000o C; 900 o C is the most appropriate temperature for NiO+bentonite.

[383]

The mass transfer mechanisms, i.e. particleexternal and particle-internal diffusion, control the overall rate of reduction reaction.

[488]

Addition of Mg was found to stabilize Ni2+ in the cubic oxide and spinel phase, increasing the reduction temperature, markedly improves regenerability; CH4/H2O (1:1) could avoid coke formation in Ni-Mg-Al-O; Ni-based systems are poorly selective to H2O and CO2, being CO and H2 if feeding CH4.

[250]

Ni and Cu-based OCs showed high reactivity, with reduction rate up to 100%/min for CuO and 45%/min for Ni, oxidation rate up to 25%/min for them; Mn and Co aren’t suitable; 560~620kg/MW OC would be needed; Solds

[236]

Reactor: a fixed bed Gases: 10% CH4+10% H2O+5%CO2 +75%N2(reduction), 10% O2+N2 (oxidation) Flow rate: 80 NmL/min Temperature: 750~950 o C

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Authors [486]

Energy & Fuels

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48

Oxygen carrier particles ω : 0.25~0.29mL/g NiO (60%)+ NiAl2O4 o Tsin : 1300 C

d p : 97 µm

NiO (26.0%~78.4%)+ bentonite (Al3SiO2) o 3 Tsin : 900 C, ρ p :3562~3731kg/m

ρ b : 1020~1530 kg/m3 ω : 0.59~0.714 Ni(21.6%~74.0%)+ bentonite (Al3SiO2) o 3 Tsin : 1050 C, ρ p :2655~7374kg/m

Experimental conditions Pressure: atmospheric pressure Bed mass: 100 mg 1. Reactor: TGA Gases: H2(reduction), Air(oxidation) Temperature: 900 o C Pressure: atmospheric pressure Bed mass: 10 mg 2. Reactor: fluidized bed Gases: 67% H2+33% Ar (4.7cm/s, STP) (reduction), Air (1.7 cm/s, STP)(oxidation) Temperature: 600, 900, 1200 o C Pressure: atmospheric pressure Reactor: TGA Gases: CH4 (5.04%) (reduction), Air(oxidation). Gas flows: 100 ml/min (reduction), 100 ml/min(oxidation) Temperature: 650~1000 o C Pressure: atmospheric pressure

Remarks circulation rate is 1~8 kg/(MW.s).

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Authors

NiO (60%)+NiAl2O4 has good circulation properties, high reactivity and high mechanical strength. It could be used in the CLC circulation.

[51]

Oxidation reaction rate increases as mass content of Ni and temperature increase; Reduction rate shows maximum point with increase of temperature and mass ratio of NiO; Oxidation is controlled by product layer diffusion, reduction is controlled by chemical reaction rate.

[418]

Reactor: TGA Reduction : 600 o C, H2 100mL/min (STP); 700 o C, CH4+H2O (1:2) Oxidation: 1000 o C, Air Pressure: 1~9atm

NiO+NiAl2O4 will provide an Outstanding performance for CLC. H2O/CH4=2.0 could avoid the carbon deposition

[259]

Reactor: a fixed bed Gases: 33~40 v% CO, 17~20 v% H2, 20~33 H2O vol%, 0~10 vol% CO2, Ar (reduction), Air 100 v% (oxidation) Gas flow: 500~1200 ml/min Temperature: 600~1000 o C Pressure: 1~9 atm

Identify the reaction kinetics of coal gas fueled chemical looping combustion

[103]

ρ b : 847~1515 kg/m3 ω : 0.681~0.795 d p : 80 µm NiO+Al2O3, NiO+TiO2, NiO+MgO; CoO+Al2O3, CoO+TiO2, CoO+MgO; Fe2O3+Al2O3, Fe2O3+TiO2, Fe2O3+MgO; NiO+Al2O3, NiO+NiAl2O4, NiO+YSZ Reactant: binder=60%:40% d p =1800, 2100 µm o

Tsin : 1300 C

NiO(60%)+NiAl2O4, CoONiO(60%)+YSZ o Tsin : 1300 C Size: pellet (dia.=4.0 mm, height=1.5 mm)

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Oxygen carrier particles 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48

Fe2O3 (40%~80%) +MgAl2O3 o Tsin : 950~1300 C, ρ p :1400~3500 kg/m3. d p : 90~250 µm ω : 0.23~0.59

Fe2O3(60%)+Al2O3, Al2O3(32%) + kaolin(8%), MgAl2O4, ZrO2,TiO2 o Tsin : 950~1400 C ρ p :1100~4200 kg/m3 d p : 125~180 µm ω : 0.22~0.76

Fe2O3(60%)+Al2O3, Al2O3(32%) + kaolin(8%), NiO(60%)+NiAl2O4,CuO(60%)+CuAl 2O4, Mn3O4(60%)+MnAl2O4 o Tsin : 1300 C

d p : 125~180 µm 40%~80% Cu, Fe, Mn, Ni oxides with Al2O3, sepiolite, SiO2, TiO2, ZrO as inert. o Tsin : 950~1300 C, ρ p :1400~5000 kg/m3 ω : 0.1~0.77 CuO (40%, 60%, 80%) with Al2O3, sepiolite, SiO2, TiO2, ZrO as inert. o Tsin : 500, 1100 C.

d p : 200~400 µm . Fe2O3 (40%~80%) +Al2O3

Energy & Fuels

Experimental conditions Bed mass: 50g Reactor: FB reactor of quartz Gases: 50% CH4+50% H2O (reduction), 5% O2(oxidation) Velocity: 2~8 U mf (reduction),

Remarks

Authors

60%Fe2O3 on 40% MgAl2O3 is the most suitable (sintered at 1100o C); Bed mass needed in FR was in the order of 150 kg/MWth.

[297]

Fe2O3+ MgAl2O4 sintered at 950 o C, Fe2O3+ZrO2 sintered at 1100 oC and Fe2O3+Al2O3 sintered at 1300 o C are good OCs.

[251]

OCs based on Fe, Ni, Cu showed high reactivity; CuO and Fe2O3 on Al2O3 showed agglomeration; NiO has limited strength; Bed mass needed is 80~330kg/MWth, and recirculation needed is 4~8kg/(s.MWth).

[257]

SiO2 and TiO2 are best for Cu-based OC; Al2O3 and ZrO2 are best for Fe-based OC; ZrO2 is best for Mn-based OC; TiO2 is best for Nibased carriers.

[268]

CuO based carriers which were prepared by wet impregnation using titania and silica as supports have good chemical and mechanical properties for CLC.

[268]

The feasibility of using iron oxide as an OC

[489]

5~11 U mf (oxidation) Temperature: 650~950 o C Pressure: atmospheric pressure Bed mass: 10g, 15g Reactor: FB reactor of quartz Gases: 50% CH4+50% H2O (reduction), 5% O2(oxidation) Velocity: 2~8 U mf (reduction), 5~12 U mf (oxidation) Temperature: ~950 o C Pressure: atmospheric pressure Reactor: FB reactor of quartz Gases: 50% CH4+50% H2O (reduction), 5% O2(oxidation) Velocity: 5~10 U mf (reduction), 10~20 U mf (oxidation) Temperature: 950 o C, 850 o C Pressure: atmospheric pressure Bed mass: 10g Reactor: TGA Gases: 70% CH4+30% H2O (reduction), Air (oxidation) Gas flows: 25 nL/h Temperature: 800~950 o C Bed materials: 20~100mg Reactor: TGA Gases: CH4+H2+H2O (30wt%) CO+H2+H2O (30wt%) (reduction), Air(oxidation) Gas flows: 25 nl/h Temperature: 800 o C Pressure: atmospheric pressure Bed mass: 20~40mg Reactor: FB reactor of quartz

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Energy & Fuels

Oxygen carrier particles 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48

o

Tsin : 1300 C

d p : 125~250 µm NiO(42.6%)+YSZ, NiO(48.7%)+NiAl2O4 CoO-NiO(42.7%)/YSZ o Tsin : 1300 C Size: pellet (dia.=4.0 mm and height=1.5 mm). Fe2O3 (58%-100%) +Al2O3 o Tsin : 1300 C

d p : 120~500 µm

CaSO4 Bulk density: 1500 kg/m3 Specific density: 2900 kg/m3 d p : 0.15~0.2 mm CaSO4 d p : 8.9 µm

CaSO4 Bulk density: 1500 kg/m3 Specific density: 2900 kg/m3 dp : 0.15~0.2 mm

Hematite oxygen carrier BET surface area=3.98 m2/g Porosity=0.03 cm3/g; The average aperture size=22.22 nm. dp : 0.1–0.5 mm.

Experimental conditions Gases: 100% CH4 (reduction), 5% O2(oxidation) Gas flows: 250, 500 NmL/min (reduction), 1000, 1000 NmL/min(oxidation) Temperature: 950 o C Pressure: atmospheric pressure Reactor: a fixed bed, TGA Gases: CH4/H2O=1:2(reduction), 10% O2+N2 (oxidation) Flow rate: 300mL/min (0.2m/s) Temperature: 600~700 o C Pressure: 1~3 atm Bed materials: 50g Reactor: a fixed-bed reactor of quartz Gases: 100% CH4 (reduction), Air(oxidation) Gas flows: 300mL/min (reduction), 900mL/min(oxidation) Temperature: 950 o C Pressure: atmospheric pressure Bed mass: 20g, 60g, 90g Reactor: Fixed bed Gases: 50 mL/min CH4 (reduction) 1000 mL/min Air (oxidation) Temperature:850~ 950 oC Reactor: TGA Gases: 40% CO2+40% N2+20% CO (reduction), 21% O2+79%N2 (oxidation) Temperature: increase the temperature from 25 to 1355o C at different heating rate. Reactor: fluidized bed Gases: 600 mL/min of 50%H2+25%CO+25%CO2 at 7.12-8.15 Umf (reduction) 1200 mL/min of 5%O2+N2 ) at 15.6 Umf (oxidation) Temperature: 950 o C Reactor: TGA Gases: 1.5L L/min of 10%H2+90%N2 (reduction) Temperature: 950 o C

Page 144 of 156

Remarks

Authors

NiO/NiAl2O4 and CoO-NiO/YSZ are good candidates. Carbon deposition could be avoided completely by CoO-NiO/YSZ without addition of water and by NiO/NiAl2O4 at the ratio of H2O to CH4 of 2.0.

[324]

was investigated.

(

dX ) red : 3~23 %/min dt

(

dX ) ox : considerably fast, 20~90%/min dt

[490]

The mass-based reaction rates during the reduction and oxidation also demonstrated the variation of reactivity of CaSO4 oxygen carrier.

[189]

Kinetic parameters of the decomposition reaction were achieved.

[491]

The performance of in the cyclic experiments was tested. The oxygen carrier conversion after the reduction reaction decreased gradually in the cyclic test.

[190]

Relative stable structure was seen in the region rich in SiO2 or Al2O3 contents.

[362]

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Energy & Fuels

Oxygen carrier particles MnO2/Mn3O4 and MgO with optional addition of Ca(OH)2 or TiO2. Freeze granulation dp: 125-180 µm

Experimental conditions Reactor: quartz reactor Gases: 450 mln/min pure CH4 Temperature: 810oC, 850oC, 900oC and 950oC

CaMn0.875Ti0.125O3 Freeze granulation and freeze drying Sintered for 3 h at 1200 ◦C dp:90-212 µm Bulk density=1100 kg/m3 Crushing strength: 1.25 N Fe2O3-CaO Granulation and followed by impregnation dp:425-500µm Perovskite-structured Ca0.8Sr0.2Ti0.8Ni0.2O3, fluoritestructured CeO2, and spinel-structured MgAl2O4 Solid state reaction (SSR), coprecipitation (CP), and citric acid(CA) methods dp: