Chemical Looping Gasification for Producing High Purity, H2-Rich

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Chemical Looping Gasification (CLG) for Producing High Purity, H-rich Syngas in A Co-Current Moving Bed Reducer with Coal and Methane Co-Feeds Tien-Lin Hsieh, Yitao Zhang, Dikai Xu, Chenghao Wang, Marshall Pickarts, Cheng Chung, Liang-Shih Fan, and Andrew Tong Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.7b04204 • Publication Date (Web): 16 Jan 2018 Downloaded from http://pubs.acs.org on January 16, 2018

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Industrial & Engineering Chemistry Research

Chemical Looping Gasification (CLG) for Producing High Purity, H2-rich Syngas in A Co-Current Moving Bed Reducer with Coal and Methane Co-Feeds Tien-Lin Hsieh, Yitao Zhang, Dikai Xu, Chenghao Wang, Marshall Pickarts, Cheng Chung, Liang-Shih Fan, Andrew Tong*

William G. Lowrie Department of Chemical and Biomolecular Engineering, 151 W. Woodruff Avenue, Columbus, OH 43210, USA

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ACS Paragon Plus Environment

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ABSTRACT

A novel design of a coal gasifier using the chemical looping concept is introduced in the present study for high purity, H2-rich syngas generation using coal and methane as co-feeds. In this work, an iron-titanium composite metal oxide (ITCMO), capable of cracking the heavy hydrocarbons produced in coal pyrolysis as well as regulating the product syngas purity, is used as the oxygen carrier. The co-current moving bed avoids backmixing of solid and gas reactants, allowing both phases to interact, reaching thermodynamic equilibrium conditions at the reactor gas outlet. This paper focuses on demonstrating the co-current moving bed reducer with the ITCMO oxygen carrier. A sensitivity analysis is performed to determine the optimal operating conditions for converting PRB coal using ASPEN Plus modelling. The tar-cracking capability is ascertained by the GC-MS analysis. The bench scale moving bed reducer substantiated its capability of achieving near-full conversion of the carbon species. The co-feeding of methane can yield a high purity syngas with H2/CO ratio of 2 or higher, which is suitable for downstream chemical synthesis. The gas and solid compositions obtained at reducer outlets match the predictions from the ASPEN Plus model. The results indicate that the extent of char gasification at the top moving bed is a critical factor for achieving a high coal conversion. The results further indicate that the sulfur in the coal is mostly converted into the gas phase emitted with the syngas product in the reducer, while the remainder is retained in the ash.

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1. Introduction

Chemical looping is the use of chemical intermediates to decompose a desired reaction into multiple subreactions preventing the inert components in each reactant feed stream from coming in contact with the other. The distinction of chemical looping processes from purely catalytic systems is that the chemical intermediate participates in each sub-reaction, though the net result is no change in the chemical state of the chemical intermediate.1,2 Since each sub-reaction occurs in sequence instead of parallel, one can easily achieve product separation by controlling the location of the reactant feed and product outlet points in the reactor system. In the early 2000s, prompted by the awareness of global warming from carbon emission, chemical looping combustion (CLC) and chemical looping reforming (CLR) were used to convert fossil fuels to electricity and hydrogen, respectively.3–8 The specific application of the chemical looping concept of interest in the present manuscript is the use of metal oxides as the chemical intermediate, or oxygen carrier, to transport oxygen from air to the fossil fuel via redox reaction cycles. Such application has gained much popularity as the core concept for developing advanced processes for carbon capture and syngas generation applied to power and chemicals production applications, respectively.2 Due to the high interest, the research into chemical looping continued to increase exponentially, demonstrated by the number of publications on this topic published since 1992 as obtained from the Web of Science. In this article, we consider the use of the chemical looping concept for coal gasification with a composite metal oxide serving as the oxygen carrying chemical intermediate with a series of reduction and oxidation sub-reactions. Figure 1 is a simplified block flow diagram of the chemical looping scheme for coal gasification where a metal oxide (MeOa) is reduced by the feedstock to MeOb in a reducer before being regenerated back to the oxidized form in the combustor with an oxidant. The metal oxide acts as an oxygen carrier which undergoes cyclic reduction and oxidation reactions to eliminate the direct contact between the coal feedstock and N2 diluted air oxidant. As the net reaction is coal gasification with air and steam to produce syngas, the energy penalty of the system is minimized. However, the separation of the gaseous oxidant from the carbonaceous fuel, allows for the direct production of syngas undiluted with N2 present in air. In this work, we first provide a brief review of the

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conventional coal gasifiers and their relative strength and weaknesses. We then investigate the potential of utilizing the chemical looping reaction scheme in place of the traditional coal gasification technology through a unique moving bed reactor design and a composite metal oxide oxygen carrier. [Figure 1] Conventional coal gasifiers, which use molecular oxygen as the oxidant directly with the coal, are classified into three categories based on the flow behavior of the coal: i.e. moving bed, fluidized bed and entrained flow gasifiers. In moving bed gasifiers (MBGs), large coal particles travel downward, via gravity, along the length of the reactor while a counter-current oxidant stream, consisting of molecular oxygen from a direct air feed or an air separation unit, flows upward.9,10 Though the MBGs are advantageous in their ability to process large coal particles with low oxidant demands, high concentrations of heavy hydrocarbon and tars exist in the produced syngas stream due to a large temperature gradient across the reactor bed caused by poor radial and axial coal mixing resulting in poor heat distribution. Therefore, commercialized gasification processes with MBGs, such as the Sasol-Lurgi process developed downstream separation/treatment to recover the hydrocarbon as byproducts.11 A fluidized bed gasifier (FBG) is operated as a well-mixed coal-oxidant reactor where oxidants serve to both partially oxidize the coal particles and to fluidize it, as the superficial velocity in the FBGs are maintained above the minimum fluidization velocity of the coal particles.12,13 The fluidized bed operation promotes uniform heat transfer across the reactor for greater tar reforming than MBGs. However, because FBGs resemble wellmixed CSTR reactors, there is a wide range of residence times of the individual particles. Since the unconverted carbon is distributed across the bed, unreacted carbon is inevitably elutriated with the removal of fully-reacted ash particles. Thus, full conversion of coal is theoretically unachievable in FBGs. One common strategy to improve the carbon conversion is by the staging approach, which is further converting the unreacted carbon removed from the fluidized bed with oxidants in a separate compartment.14,15 In entrained flow gasifiers (EFGs), the oxidants and coal particle travel co-currently as the superficial gas velocity in the EFG reactor is maintained above the terminal velocity of the coal particles throughout.16 EFGs can achieve high conversions of coal to CO and H2 syngas. However, the high particle velocity achieved in EFGs results in a low coal particle residence time and, thus, requires the use of high operating pressures and temperatures as well as small coal feed sizes to 4


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maximize coal gasification kinetics. The high operating pressures and temperatures result in costly reactor designs and long-term operability issues. Measures to extend the coal particle residence time often involves geometrical designs of burner firing patterns such as swirling or tangential firing.17,18 While each gasifier design has an advantage and field of applications, each are challenged with either low CO/H2 syngas yields or costly high temperature, high pressure reactor designs. Further, for chemical production applications, each gasifier requires the use of an ASU to produce a syngas stream of sufficient purity for further downstream catalytic reforming. Chemical Looping Gasification (CLG) processes are becoming an attractive and potentially cost-effective alternative as these processes do not require the use of molecular oxygen to produce a syngas product from coal gasification undiluted with the N2 present in air, thus the need of ASU in conventional gasification systems can be eliminated. Moreover, a tailored CLG system could potentially overcome the limitations observed in conventional gasifiers, such as incomplete hydrocarbon conversion of MBG, unconverted carbon elutriation of FBG, and extreme condition requirement for EFG. The CLG system proposed in this paper aims to address these limitations by (a) utilizing oxygen carrier material capable of catalytically convert hydrocarbons, (b) using a reactor design that prevents solid back mixing which leads to wide residence time distribution, (c) providing sufficient residence time for complete coal conversion. Table 1 summarizes the published performances of CLG processes for syngas production with various combinations of reactor configurations and oxygen carriers. The performance criteria summarized in Table 1 are the molar ratios of H2/CO and CO/CO2. Obtaining quality syngas composition is important for minimizing downstream processing, for example a H2/CO molar ratio of >2 is a desired composition for the production of methanol or gasoline. The CO/CO2 molar ratio in the syngas product stream can be viewed as a direct indicator of the process’ syngas production ratio over the fully oxidized products such as CO2 and H2O.

For the CLG processes demonstrated to date, most processes adopted a fluidized bed reducer reactor design due its excellent properties for uniform reactant and heat distribution. However, fluidized bed design results in 5



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a wide distribution in oxidation states of the oxygen carrier due to the well-mixed nature of this reactor design for solids mixing. For oxygen carriers with multiple oxidation states such as iron-based oxygen carriers, the presence of the higher oxidation states of the metal oxides in the reactor will thermodynamically favor over conversion of the syngas product to CO2/H2O. Further, similar to FBGs, the gas bubble phase that exist in this reducer design can allow for heavy hydrocarbons and tars to bypass the reactor unconverted due to the mass diffusion limitations between the gas bubble phase and the gas-solid emulsion phase. The separation of coal ash from the oxygen carriers in a fluidized bed can pose an additional challenge due to their similarities in terms of particle size. Further, due to the slow solid-solid reaction kinetics between coal and oxygen carrier solids, the conversion of fixed carbon in coal relies on the introduction of an oxygen-containing gasifying agents, such as steam or CO2, to perform the following gasification reactions: + → + + → 2 The oxygen carrier provides the necessary heat from the combustor reactor to perform the above endothermic reactions. The CO and H2 generated can then react and equilibrate with the oxygen carriers and other gaseous species. For fluidized bed reducer operations, an additional amount of gasifying agent (i.e. above stoichiometric requirements) is often required to ensure the superficial gas velocity throughout the reducer is above the minimum fluidization velocity. [Figure 2] In this paper, we present a unique co-current moving bed reducer design developed at Ohio State University (OSU) for CLG processes to achieve high coal conversion and CO and H2 syngas selectivity. In a co-current moving bed, both the oxygen carriers and the fuel flow downward from the top to the bottom of the reducer, respectively as shown in Figure 2. The packed, moving bed operation of this reducer design capitalizes on the advantage of MBGs in providing sufficient residence time for complete coal conversion in a single-pass of the coal feed. In addition, this reducer design capitalizes on the advantage of EBGs where the plug flow operation of the co-current coal and gas flow, with no gas bubble phase formation, ensures high product selectivity to CO and H2. Given sufficient contact and reaction time, the oxygen carrier will reach a uniform oxidation state at the 6


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bottom of the reducer – in thermodynamic equilibrium with the syngas composition at the reducer gas outlet. Thus, co-current, moving bed reducer reactor design can be a promising component for CLG reactor systems and, with the selection of a highly reactive oxygen carrier thermodynamically favored for high H2/CO yields from coal, can be used to produce syngas of high CO and H2 purity from coal. [Figure 3] An oxygen carrier material favorable for reaching syngas-rich equilibrium can be selected using a modified Ellingham Diagram (Figure 3), which compares the reaction Gibbs free energies (∆Gs) within a range of temperatures. Ellingham Diagrams are used extensively in the extractive metallurgy industry to select the best reducing agents to produce metals from ores, but it can also be adopted to analyze metal oxides with favorable redox properties for fuel gasification and inhibited from full combustion of the fuels to CO2/H2O.2 Using thermodynamic software such as ASPEN Plus or HSC Chemistry, the oxidation ∆G curves of different oxygen carrier redox pair can be generated and screened for selecting metal oxide materials within the boundary of C oxidation to CO and H2 oxidation to H2O illustrated in Figure 3 as the region shaded in gray. Curves in this region, when combined with the Gibbs free energy of H2 and CO result in a net negative ∆G indicating the spontaneous production of syngas. Curves in this region are also below the Gibbs free energy for the formation of CO2 and H2O, which indicates the reduction of oxygen carrier material to form CO2 and H2O is not spontaneous and, thus, are thermodynamically unfavorable competing reactions. Therefore, despite that small amount of CO2 and H2O will still exist in the equilibrium product due to the water gas shift reaction, metal oxides in the shade region are favorable for partially oxidizing the hydrocarbon over full combustion, providing high selectivity towards syngas production. In the authors’ previous work, iron-titanium composite metal oxide (ITCMO) was selected as oxygen carriers using the modified Ellingham diagram screening method as shown in Figure 3. Further analysis in ASPEN Plus for syngas production from natural gas/shale gas shows that the ITCMO oxygen carrier can produce syngas of high purity (>90%) and high hydrogen content (H2:CO ~ 2:1) with near full conversion of the carbonaceous fuel in a co-current moving bed reducer operating between 950 and 1050˚C.30 Thus, the simulation results show the co-

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current moving bed reactor design using an ITCMO oxygen carrier can produce high fuel conversion to syngas when natural gas is used as the feedstock. The active components of ITCMO particles used include Fe, FeTiO3, Fe2TiO4, Fe3O4 in the reduced state, Fe2O3, Fe2TiO5 in the fully oxide state, and TiO2 in both states. ITCMO oxygen carrier also contains additional support materials that enhances the mechanical property of the material. More details about materials analysis can be found in the Materials sections of the authors’ previous publications.30,31 The objective of this study is to experimentally analyze the performance of a co-current moving bed reducer using ITCMO oxygen carrier for coal gasification via experimental testing of the reactor design concept. The performance is evaluated in terms of coal conversion, as well as syngas purity and H2/CO ratio of the gas product as functions of gasifying agent input. ASPEN Plus simulations are performed to determine the operating conditions and to predict equilibrium compositions of gas and solid phases at the outlets. Post-experiment solids analysis of the oxygen carrier sampled from the moving bed reactor tests are conducted to confirm oxygen carrier conversion, coal conversion and to investigate potential sulfur deposition on ITCMO. Gas Chromatography–Mass Spectrometry (GC-MS) coupled with a pyrolyzer and a lab-scale fixed bed is used to investigate ITCMO’s capability to decompose heavy hydrocarbons to single carbon species in support of the experimental result of total carbon conversion. 2. Thermodynamic Analysis with ASPEN Plus Simulation

2.1 ASPEN Plus Model Setup

[Table 2] An ASPEN Plus simulation model is constructed to investigate ITCMO’s syngas generation capability as the oxygen carrier in a moving bed reducer. Most of the chemical species involved in the CLG process can be directly simulated in ASPEN Plus with existing property databases. Coal and ash are simulated as Nonconventional components, which is the most common simulation method for complex solid species. The specific model setup is also summarized in Table 2.

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[Figure 4] [Table 3] Figure 4 shows the flow sheet of CLG reducer simulation model. The specific composition of coal input to the model is shown in Table 3, which is defined by the ultimate and proximate analyses of PRB coal procured for bench-scale moving bed experiments. The RYield block, DECOMP, is set to decompose the coal into its constituent elements to enable coal reaction calculation. The ITCMO oxygen carrier is simulated by a conventional solid stream, OCIN, with Fe2O3 and TiO2 as the reactive species with a molar ratio of 1:1. The cocurrent moving bed is simulated using an isothermal RGibbs block at 1050˚C in which the gaseous and solid phases reached thermodynamic equilibrium. The streams which exit the RGIBBS block, REDGOUT and REDSOUT, represent the gaseous and solid product flows, respectively. This model was used to perform the initial thermodynamic analysis in Section 2.2 as well as the reducer performance predictions in Section 4.1. For the latter, two extra feed streams of METHANE and WATER were also used to simulate the CH4 and H2O fed to the reducer in addition to the oxygen carrier and coal. It is also worth mentioning that the energy balance associated with the DECOMP module must be considered in order to properly simulate an adiabatic coal gasification system. In the presented work, the authors are discussing the thermodynamic equilibrium of the CLG reducer under isothermal condition to match the operating condition of the experimental apparatus. Thus, the energy change associated with the DECOMP module does not affect the product composition of an isothermal RGIBBS module. In cases where the heat balance of the CLG reducer is of interest, an energy flow should be drawn from the DECOMP module and directed to the RGIBBS module, such that the accurate heating value of the coal feedstock is taken into account.

2.2 Thermodynamic Analysis

A parametric study of the effect of oxygen carrier to coal mass flow ratio on the gas and solid equilibrium products is performed with the ASPEN model described in Section 2.1. The ratio, , is defined by the flowing equation: 9



=

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where and are the equivalent mass flow of the fully-oxidized iron oxide in ITCMO and the total mass flow of PRB coal, respectively. Using the Sensitivity function in the Model Analysis Tools of ASPEN plus, the equilibrium product compositions can be determined over a range of value of . was set at 3.06 g/min, while the was varied between 0.05 to 15.05 g/min. Figure 5 shows the parametric study results of the distributions of each carbon, iron and hydrogen species in the gas product stream from the reducer. From Figure 5, when is between 0 and 1.45, only metallic iron, Fe, exists as the product for the Fe-based solid phase species. As is increased from 1.45 to 4.17, both Fe and FeTiO3 species exists where concentration of FeTiO3 increases as increases. When is between 4.17 to 5, only FeTiO3 exists as the solid phase product species. The possible carbon species in the product stream are CO, CO2 and solid carbon, C. Solid carbon forms when is below 1.2, which indicates thermodynamically, such low flow ratios either prohibit coal char conversion or promote carbon deposition via the reverse Boudouard reaction and is not a desirable operating condition. When is between the 1.45 and 4.17, the distribution of carbon species is identical throughout due to the thermodynamic effect of co-existing Fe and FeTiO3. Within this window of , the product gas composition is uniquely determined by the temperature pressure, despite the varying ratio between Fe and FeTiO3. When ranges from 0 to 1.45 and 4.17 to 5, the gas species, which include CO, H2, H2O and CO2, generally follow the same trends where the concentration of CO and H2 decreases in proportion as increases, and vice versa for CO2 and H2O.

[Figure 5]

Figure 6 shows the parametric study of the effect of on the syngas purity and H2 to CO ratio in the gas product stream. The syngas purity is defined as

=

+ + + +

10


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where [i] represents the molar concentration of each gaseous specie (i = CO, H2, CO2, H2O) in the outlet stream. As the is increased from 1.2 to 4.2, shown as the shaded region of Figure 6, a syngas purity of greater than 91% is produced. Further increasing the results in a decrease in syngas purity as more lattice oxygen is available and shifts the equilibrium towards CO2 formation. Therefore, the shaded area represents the desired region of for maximum syngas production in the CLG process. [Figure 6] Despite the high purity, however, using coal as the sole feedstock to the CLG process does not provide a sufficient H2/CO ratio for many desired chemicals. According to the simulation result (shown in Figure 6), the H2/CO ratio plateaus at around 0.7, which is significantly less than 2, the ideal ratio for producing chemical products such as liquid transportation fuels or methanol. In order to upgrade the H2 produced from the CLG reaction, a hydrogen-rich reactant must be co-injected with the coal feedstock. One option is to co-inject natural gas, which is predominately methane (CH4). Since the net exothermic reaction of CH4 partial oxidation introduces additional heat to the overall process, co-injecting CH4 can be beneficial for boosting the H2/CO ratio as well as contributing to achieve autothermal operation of CLG. Another option is to introduce steam as a pure hydrogen source to directly increase the H2/CO ratio via water gas shift reaction with the produced CO, steam reforming of the tar volatiles and char, and from the steam-iron reaction with the reduced oxygen carrier. The major drawback of the use of steam upgrade the H2/CO ratio is the addition heat required maintain the process heat balance. As the net reaction in chemical looping processes is the partial oxidation of coal with air and/or steam, the reforming of coal with steam in a chemical looping reaction scheme can be considered a net endothermic reaction. The greater the amount steam used for reforming the coal will result in a reduction in the net heat output from the chemical looping reactor system to a point where the net reaction becomes endothermic. To ensure the chemical looping process can operate autothermally, i.e. with no external heat input to the reactor system, the net reaction must be at least neutral, if not slightly exothermic. Therefore, when coal is used as the primary feedstock, a portion of the CO must be fully oxidized to CO2 to balance the overall heat required in the chemical looping reactor system to produce a 2:1 ratio of H2:CO from the reducer reactor, thereby, reducing the syngas purity from the 91% achieved in the coal only case. The following 11



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experimental work demonstrates the feasibility of co-injecting coal, methane, and steam to produce high purity, high hydrogen content syngas. The gas and solid composition at the reactor outlets were also measured and compared with the ASPEN Plus predictions for discrepancies and further discussion.

3. Experimental Setup

3.1 Moving Bed Apparatus for Syngas Generation

The performance of the ITCMO and co-current moving bed reducer reactor are experimentally tested using a bench-scale unit (shown in Figure 7). The moving bed reactor is constructed from a 2-inch ID stainless steel pipe and encased with external heaters to allow for high temperature, isothermal operation. The solid flow rate is controlled by a screw feeder, whose feeding speed could be adjusted on the Variac transformer that powered the motor. The solid loading system, located near the top of the reactor, consists of a Y-shaped injection port with a lock hopper. The oxygen carrier particles are introduced into the system from this port. The screw feeder, attached to the base of the reactor, transfers the oxygen carriers and residual coal char from the reactor into a collection hopper at a controlled solids flow rate. The bed-level is constantly monitored using a glass window located above the heated section to measure solid flow rate. The temperature profile along the height of the reactor is recorded using eight type-K thermocouples. When the bed temperature reached the designed reaction temperature (>1000°C), the pulverized PRB coal is introduced by pre-mixing with the oxygen carrier at the designed ratio and introducing the mixture through the Y-shaped injection port at the top of the reactor. [Figure 7] Nitrogen and methane are introduced at the top of the reactor through the gas mixing panel with mass flow controllers regulating each flowrate. The nitrogen serves as an inert tracer gas, whose known molar flowrate is used to determine the molar flowrates of the various gas species generated from the gasification reactions. An HPLC pump is used to deliver deionized water directly into the reactor at a top side port labeled in Figure 7. The water consumption rate in the reservoir is monitored to calculate the steam injection rate. 12


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The outlet for the gaseous products is located at the end of the screw feeding system and connected to the gas analysis system. Fine particulates are separated from the gas stream using a disengagement stripper and a water trap. The gas is subsequently cooled using a heat exchanger tube submerged in an ice bath, dried by a desiccant bed and passes through two gas analyzers which measure CO, CO2, CH4 and H2. Finally, the exiting gas is vented to atmosphere. After steady state conditions are reached for each test condition, the external heaters are turned off while a nitrogen blanket flow through the reactor is maintained to prevent further reaction with the oxygen carrier and residual coal during cooldown. When the reactor reaches room temperature, solid materials are sampled from each of the sampling ports along the bed height for oxygen carrier analysis. 3.2 Oxygen Carrier Analysis

A total carbon analyzer (shown in Figure 8a) is used to analyze the oxygen carrier conversion and to determine the presence of unconverted char or carbon deposition. The carbon analyzer (CO2 Coulometer, UIC, Inc.) consists of a sample combustion and gas scrubbing system and a CO2 coulometer. For each sample, approximately 1 gram is weighed and put into the combustion tube to combust with O2 at 900oC. The tail gas is then scrubbed and injected into the CO2 coulometer, which measures the total amount of CO2 in the tail gas, and then translate the integrated value into total mass of carbon in the sample. [Figure 8a, Figure 8b] A total sulfur analyzer (Figure 8b, TS 3000, Thermo Fisher Scientific, Inc.) measures the sulfur residue on the ITCMO oxygen carriers. The sulfur analyzer consists of a sample combustion chamber, a magnetic sample boat, a gas scrubbing system and a SO2-UV fluorescence detector. The cylindrical furnace is set at 1000oC to rapidly oxidize the sample in an oxygen-rich environment. The sample boat loaded with samples is pushed into the heated zone with a programmed depth by magnetic force. As time progresses, the sample boat migrates deeper into the heater, causing the temperature to increase. The sulfur content in the sample is then oxidized to SO2 at various elevated temperatures, depending on different oxidation temperature of the solid sulfur species. By integrating the signal strength over time, the total amount of sulfur in the sample can be obtained. 3.3 Data Evaluation 13



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Since coal is introduced into the reducer by pre-mixing with oxygen carrier particles at a specific mass ratio, the ! is calculated based on the mass flow rate of ITCMO particles (" ). The reducer capacity is defined as the thermal capacity of the total fuel feedstock and is calculated by #$ = ! ∙ & '( $') + *!+, ,./ ∙ 01!+, & '( 2 where *3,./ represents the input molar flow rate of species A. The outlet gas composition is obtained from two gas analyzers during the bench scale moving bed experiment. CO/CO2, H2/CO and methane conversion are calculated by following equations:

/ ') ' =

H / ') ' =

CH2 $'78' = 1 −

! ! + !

*!+, ,; ,./ =1− *!+, ,./ > *!+, ,./

where 3 represents the molar concentration of component A recorded by the gas analyzer. Since ? is used as an inert tracer gas, molar flow rate of other species can be calculated:

*3,; ,./ >

The coal conversion is defined as the extent of the original carbon content converted to the gas phase. Therefore, consider the carbon balance of the reducer to calculate coal conversion, i.e., *!!,./ + *!+, ,./ = *!!,;

The oxygen transfer rate represents the extent of the oxygen carrier reduction towards equilibrium, which can be used as an indicator to validate the conversions of the CH4 and coal obtained previously. The weight measurements obtained can also be used to quantify the amount of unconverted carbon with respect to the original carbon content in the coal fed into the reactor, defined as the carbon residue percentage (J%):

J% =

0! ∙ " 0! ! ∙ % ÷ = " 0B 0B ∙ ! ∙ %

Similar calculation can be performed to obtained the sulfur residue percentage (J%):

J% =

0N ∙ " 0N ! ∙ % ÷ = " 0B 0B ∙ ! ∙ %

where % being the sulfur content of coal from the ultimate analysis, and 0N being the mass of sulfur in each sample measured by the sulfur analyzer. 3.4 Tar Cracking Test of ITCMO Fixed Bed

A Gas Chromatography–Mass Spectrometry (GC-MS, Agilent Technologies, 7890B GC system/5977A-MSD) setup is used to study the tar cracking capability of ITCMO particles. In each test run, a small amount coal is 15



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placed in a pyrolyzer continuously flushed by a carrier gas stream (50ml/min helium). The pyrolyzer then rapidly heats up to 900˚C, generating various types of coal volatiles. The carrier gas carries the volatiles and passes through a quarter-inch-diameter reactor. The reactor contained either no particle or an ITCMO fixed bed, and the volatile was expected to be thermally-cracked or cracked by the particles. The carrier gas further carries the cracked volatile into the GC column, where the various species separate due to different travelling speeds of the molecules. Then, the MS ionizes the species and sorts the ions by their mass-to-charge ratios. Based on the information obtained, the exact formula of the volatile species can be determined. Note that the lighter hydrocarbons and carbonaceous species such as CH4, CO and CO2 cannot be detected by GC-MS. 4. Results and Discussion

4.1 Bench Unit Gas Analysis

The bench unit was first operated at 1025℃ under isothermal conditions, using only PRB coal and ITCMO particles as feed to the reducer, without co-injection of CH4 or steam. The mass flow rate for and ! used were 8.56 and 2.9 g/min, respectively. Therefore, the value of was 2.95, which represents the middle region of Figure 5 where the concentration of the product gas species plateaus. Figure 9a shows the dynamic gas outlet concentration profiles of each species from the CLG bench reducer for Condition A as described in Table 4. Based on the concentration profiles, CO/CO2 and H2/CO gas ratios are calculated and shown in Figure 9b. The corresponding ITCMO composition profiles and product gas species at different stages of the experiment for Condition A is illustrated in Figure 10. Note, the windows with no data recorded were the system service periods, as indicated with gray shading in Figure 9a and 9b. At the beginning of the experiment, the coal devolatilized upon reaching the heated reactor zone. The hydrocarbon volatiles passed through the pre-existing, fully oxidized ITCMO moving bed and were all converted to combustion products, CO2 and H2O, because the gases generated equilibrated with Fe2O3/Fe2TiO5 at the gas outlet. The effective at this instant was much larger than the designed value of 2.95 because the equivalent included the fully oxidized ITCMO in the reactor inventory. As H2O condensed and was separated from the gas stream, only CO2 was observed at the gas analyzer with the rest being the balanced nitrogen. Around the 20-min mark, the CO2 16


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concentration rapidly increased as more coal reached the heated zone and converted while the nitrogen flow rate remained constant. The CO2 concentration stopped increasing as the upper detection limit of 52% was reached. The high CO2 concentration observed during heatup is due to high concentrations of fully oxidized oxygen carrier during to initial reactor charging as part of process startup. Significant increase in CO concentration could be observed around 40 mins after the CO2 concentration began to rise, which indicated that the product gases equilibrated with more reduced ITCMO phases such as Fe or FeTiO3 as the reducer reactor bed of oxygen carrier becomes more reduced. Note that at this instant, less than 2 hours had passed since the very first batch of coal reached the top of the reactor. The steady state gas concentration profile was reached roughly 2.5 hours after the initial CO2 concentration rise. Slight fluctuations could be observed in the gas concentration reading, possibly due to incomplete mixing of coal and oxygen carrier particles in the feed. The CO/CO2 ratios were close to the predicted value from the ASPEN thermodynamic analysis and proved that the gas residence time in the designed condition was sufficient for CLG using ITCMO. Note, the gas residence time could be potentially shortened even further in a commercial unit because it will not require additional nitrogen mixed in the product stream as a tracer gas. [Figure 10] [Table 4] Table 4 summarizes the steady-state experimental results of the syngas generation from the bench unit under different conditions. From Condition A to C, the ratio between the hydrogen and carbon content in the feed increases, either by increasing CH4 or steam flow rate while maintaining the same process capacity. In all the conditions, the methane concentration in the gas product was less than 0.5%, which translated to high methane conversion (>90%) comparable to the previously reported work.30 In Condition A, coal conversion of 85%, comparable to the values reported in the literatures in Table 1, was achieved with no additional gasifying agent input. Thus, the results suggest a self-gasifying mechanism exists in which the CO2/H2O is produced in the reducer reactor and then used as the gasifying agent for the residual fixed carbon. This mechanism is proposed in Figure 11. In an integrated CLG system with a co-current moving bed reducer, the oxygen carrier also serves as a heat carrier to transfer the heat produced from the combustor reactor to perform the endothermic 17



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reactions in the reducer, including coal pyrolysis, producing char with fixed carbon and gaseous volatile. The volatiles react with the high oxidation state oxygen carriers and generate CO2/H2O, while the char continues to travel downward with the moved bed of oxygen carrier. The co-current design forces the CO2/H2O product gas to travel downward and gasify the residual coal char with the heat provided by the oxygen carrier. By coinjecting CH4 and steam, the estimated coal conversion was further increased to >90%. The enhancement was attributed to the increased partial pressures of the gasifying agents at the top of the reactor from methane combustion products (i.e. CO2 and H2O) and additional steam input. [Figure 11] 4.2 Solid Conversion and Carbon/Sulfur Residue Analysis

The solid conversions along the bed height are shown in Figure 12. In Condition A, without co-injection, the solid conversion exceeded 33% along the bed, which indicates the oxidation state of ITCMO is between FeTiO3 and Fe. Large deviations of solid conversion were also observed, which were attributed to incomplete premixing of coal and oxygen carrier at the solid feed. Theoretically, sections with locally high coal concentration will have high solid conversion, which explains the periodical surges of solid conversion at 32%, 61% and 87% bed height. The volume solid mixture fed in each batch is roughly equivalent to 30% of the moving bed in height, and hence the equal distances between of the high coal concentration layers. [Figure 12] The two solid conversion profiles from the conditions with CH4 and steam co-injection were relatively smoother than from the coal-only condition. Both profiles plateaued around 33%, which suggested the oxidation states of the ITCMO particles stabilized at FeTiO3. Since the extra H2O input in Condition C resulted in higher H2 production rate than Condition B, it was reasonable to assume that the oxygen content of the extra H2O stimulated the hydrogen-producing reactions such as water gas shift, methane reforming and coal gasification. This hypothesis is supported by the enhanced methane conversion and coal conversion from Condition B to Condition C. Furthermore, the plateaus of these profiles suggested that the solid phase reached equilibrium at around 30% of the bed height. Therefore, the size of the reactor can potentially be reduced by 18


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half while maintaining the same purity of syngas produced, assuming the gas phase has also reached equilibrium at the 30% bed height. With a closer inspection of Condition B, a noticeable increase in solid conversion was observed at the last 20% of the bed height from 32% to 41%. This value exceeded the ASPEN simulation prediction of 39% by 2%, which might indicate that a more reducing environment was present in this section. One potential reason might be a higher concentration of unconverted char at the bottom of the reducer. According to the carbon balance for Condition B, roughly 8% unconverted carbon would exit the reducer continuously at steady state. Since the particle size of pulverized PRB coal was much smaller than the oxygen carrier particles, the unconverted char powder likely had much lower terminal velocity than the oxygen carrier and was carried by the gas flow in the reactor, travelling faster than the moving bed towards the bottom. The unconverted char then gradually accumulated at the reactor bottom, causing the fuel to oxygen carrier ratio to increase locally and hence a more reducing environment. For Condition C, such phenomenon was less prominent (

Chemical Looping Gasification for Producing High Purity, H2-Rich

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