Chemical Looping Gasification for Hydrogen Production a Comparison of Two Unique Processes Simulated Using ASPEN Plus 2014 International Journal of Hy

7/25/2019 Chemical Looping Gasification for Hydrogen Production a Comparison of Two Unique Processes Simulated Using A

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Chemical looping gasication for hydrogenproduction: A comparison of two unique processessimulated using ASPEN Plus

Stephen G. Gopaul, Animesh Dutta *, Ryan ClemmerMechanical Engineering Program, School of Engineering, University of Guelph, Guelph, Ontario N1G 2W1, Canada

a r t i c l e i n f o

Article history:Received 25 October 2013Received in revised form17 January 2014Accepted 26 January 2014Available online 18 February 2014

Keywords:ASPEN PlusBiomass gasicationChemical looping gasication

HydrogenSteam gasicationTar reforming

a b s t r a c t

The research compares the simulations of two chemical looping gasication (CLG) typesusing the ASPEN Plus simulation software for the production of H 2. The simulated biomasstype was poultry litter (PL). The rst CLG type used in situ CO 2 capture utilizing a CaOsorbent, coupled with steam utilization for tar reforming, allowing for the production of aCO2-rich stream for sequestration. Near-total sorbent recovery and recycle was achievedvia the CO 2 desorption process. The second type utilized iron-based oxygen carriers inreduction e oxidation cycles to achieve 99.8% Fe 3O4 carrier recovery and higher syngasyields. Temperature and pressure sensitivity analyses were conducted on the main re-actors to determine optimal operating conditions. The optimal temperatures ranged from500 to 1250 C depending on the simulation and reactor type. Atmospheric pressure provedoptimal in all cases except for the reducer and oxidizer in the iron-based CLG type, which

operated at high pressure. This CLG simulation generated the most syngas in absoluteterms (2.54 versus 0.79 kmol/kmol PL), while the CO 2 capture simulation generated muchmore H 2-rich syngas (92.45 mol-% compared to 62.94 mol-% H 2).Copyright 2014, Hydrogen Energy Publications, LLC. Published by Elsevier Ltd. All rights

reserved.

Introduction

Hydrogen (H 2) has the potential to shift the global reliance on

fossil fuel energy sources to cleaner, more efcient forms of energy as it presents a viable alternative. The issues of globalcarbon dioxide (CO 2) and other greenhouse gas (GHG) emis-sions and overall atmospheric concentration can also beaddressed with the utilization of H 2 technologies for thegeneration of energy. Furthermore, H 2 presents an advantageover other conventional alternative energiessuch as wind andsolar due to its energy storage and transport capabilities.

Thus, H 2 utilization has the potential to become widespread inthe near future.

Current methods of H 2 production involve the reforming of fossil fuels; processes that ultimately contribute to the socie-tal and environmental issues mentioned above. Correspond-ingly, contemporary H 2 energy systems are neither GHG-neutral nor sustainable since these non-renewable fossilfuels are responsible for equivalent CO 2 emissions. Thus, it isessential that H 2 be produced from a renewable, carbon-neutral energy source. The thermochemical gasication-based conversion of biomass in the presence of steam pre-sents a viable renewable H 2 source and is a strong contenderfor the replacement of fossil fuel-based H 2 production.

* Corresponding author .E-mail addresses: [email protected] (S.G. Gopaul), [email protected] (A. Dutta), [email protected] (R. Clemmer).

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Chemical looping gasication (CLG) using biomass fuel isan example of such H 2 energy technologies as it capitalizes onrenewable, environmentally friendly, and abundant sourcesof energy. Biomass is an organic fuel source containing car-bon, hydrogen, oxygen, nitrogen, and sulfur, and can come inthe form of agricultural wastes, municipal solid wastes, ani-mal wastes, saw dust, etc. [1]. The gasication of biomass in-

volves thermochemical conversion to H 2, as well as varioushydrocarbons, for subsequent use in H 2 conversion technol-ogies with the aim of energy production. Furthermore,biomass gasication is a complex, endothermic process con-sisting of many chemical reactions. These reactions dependon the gasication agent being used which is commonlysteam. The reactions are summarized in Table 1 [2].

At the present time, minimal research has been conductedregarding CLG systems development. Moreover, the majorresearch focus involves the chemical looping combustion(CLC) process, with emphasis on oxygen carrier developmentfor use in CLC. Though widely recognized by many re-searchers, work conducted on the CLG process developed byFan et al. at Ohio State University [3] requires further researchand development to become a fully-implemented, renewableH2 production technology. Furthermore, other advancedchemical looping systems exemplar of CLG processes for thepurpose of H 2 production include the HyPr-RING (HydrogenProduction by Reaction-Integrated Novel Gasication), fuel-exible advanced combustion e gasication, ALSTOM hybridgasication e combustion, Advanced Gas-ication e Combustion (AGC), and Zero Emission Coal Alliance(ZECA) processes [3e 5].

The mitigation and subsequent elimination of sorbent/catalyst performance degradation over time is an ongoing challenge in most chemical looping systems. Successfuloperation of chemical looping systems is often jeopardized byunpredictable sorbent behaviour. Thus, research has beenconducted to identify the mechanisms contributing to sorbentlosses and develop performance improvement methods toensure sufcient system operation and longevity. Subse-quently, sintering of the chemical looping sorbent particles asa result of high-temperature operation and cyclic heating/cooling cycles, in conjunction with char and tar deposition,signicantly reduces sorbent capture and regeneration capa-bilities. The detrimental effects of these phenomena are mostnotably felt in CLG systems. Another challenge presented bychemical looping systems involves the continuous ow of solid materials between a multitude of interconnected re-actors operating at high temperatures and pressure [6].

The CLG of biomass is advantageous in that it is a clean,renewable source of H 2; however some disadvantages remainprominent. For example, product syngas streams resulting from the CLG of biomass contain many impurities. Conse-quently, much research has been conducted regarding syngasimpurity removal, whether it be the end-of-pipe or in situapproaches. Therefore, the research compares two different

CLG processes under similar feed and operating conditions forthe purpose of H 2 production using a computer simulation.The novelty of the conducted research lies in the developmentof the two CLG processes using the ASPEN Plus simulationsoftware and subsequent comparison of the results using thesame biomass feedstock in both cases.

The paper also includes temperature and pressure sensi-tivity analyses conducted on all relevant reactors and the re-sults compared between the CLG 1 and CLG 2 processes.Relevant parameters included optimal operating temperatureand pressure for each reactor and product syngas molar yieldand composition. Comparison of the H 2 production and puritycapabilities of the two processes is emphasized upon.

Simulated processes

Two CLG mechanisms were simulated and analysed using theASPEN Plus V7.3 software. The simulation results werecompiledforpoultrylitter(PL),a nonconventionalbiomass type.

Chemical looping gasication type 1The rst CLG simulation (CLG 1) incorporated in situ productCO2 capture in the absorbing reactor with the use of a CaOsorbent. A representative block diagram can be seen in Fig. 1.CO2 absorption occurs according to the following chemicalreaction [7]:

CaOs CO2g / CaCO3s

This reaction is exothermic with a heat of reaction of 178.3 kJ/mol [7]. In addition, near-total CaO sorbent recovery

and recycle was inherent to the simulation setup. CO 2desorption is the reverse of CO 2 absorption and is therefore anendothermic reaction. The energy required for the reaction tooccur is provided by the higher operating temperature of thedesorbing reactor. The resulting theoretical product stream ispure CO2 which can be sent for sequestration. Furthermore,the overall reaction prior to sorbent regeneration is as follows:

CnHmO p 2n pH2O n CaO/ n CaCO3 m=2 2n pH2

This overall reaction is endothermic with a heat of reactionof 107.5 kJ/mol [6]. The constants n, m , and p represent therespective carbon, hydrogen, and oxygen contents in thebiomass being gasied.

Another key aspect of the CLG 1 simulation is the use of steam (H 2O(g)) to address the issue of tar and char formation.Tar and char were modelled as pure carbon and the reforming reactions correspond to water e gas (i) and water e gas (ii) fromTable 1 as follows:

Cs H2Og / COg H2g

Cs 2H2Og / CO2g 2H2g

Table 1 e Chemical reactions in the steam gasication of biomass. Adapted from Ref. [2].

Reaction Chemical equation D H 923 (kJ mol 1)Water e gas shift CO H2O / CO2 H2 35.6 (exothermic)Methane reforming CH 4 H2O / CO 3H2 225 (endothermic)Water e gas (i) C H2O / CO H2 136 (endothermic)Water e gas (ii) C 2H2O / CO2 2H2 100 (endothermic)Oxidation (i) C O2 / CO2 394 (exothermic)Oxidation (ii) C 0.5 O2 / CO 112 (exothermic)Boudouard C CO2 / 2CO 171 (endothermic)

Methanation C 2H2 / CH4 89.0 (exothermic)

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These reactions were assumed to reform all tar com-pounds exiting the desorbing reactor. Furthermore, the use of steam is preferable to conventional air since the thermo-oxidative reforming of pure carbon in air results in a nitro-gen (N2) stream which must be subsequently separated fromthe CO2 product stream. This process is energy-intensive andtherefore costly [8], however the condensation of steam froman exit gas stream can also be a costly process. A more prac-tical process schematic illustrating an overview of the CLG 1process can be seen in Fig. 2.

Chemical looping gasication type 2Subsequently, syngas chemical looping is another H 2-pro-duction process involving the gasication of biomass. Thisprocess separates the gasication and looping stages andproduces H 2 via reduction, oxidation, and combustion cyclesinvolving iron (Fe), haematite (Fe 2O3), and magnetite (Fe 3O4)[6]. This process was also simulated using ASPEN Plus V7.3using poultry litter as the chosen biomass type (CLG 2). Theblock diagram for the CLG 2 process can be seen in Fig. 3.

Similarly to CLG 1, CLG 2 is comprised of many chemicalreactions involving a multitude of reactors. These reactionsand corresponding reactors are outlined in Table 2 [3]. Thereactions persist under theoretical conditions. In actuality,tar formation in the reducer is unavoidable. Tar reforming in the CLG 2 simulation is carried out in the oxidizer, wherelarge amounts of steam are introduced. Here, tar iscompletely reformed by the steam to CO, CO 2, and H 2 via theaforementioned water e gas reactions presented in Table 1 .Additionally, pure O 2 is utilized in the combustor asopposed to air for the same reasons explained for the CLG 1simulation [8].

Feedstock used

Composition of biomass types

The characteristics of three biomass types were compared toidentify a suitable feedstock for utilization in the two simu-lations. The types are: poultry litter, wood pellets, and oakpellets. The biomass types vary in their chemical compositionand are thus representative of a spectrum of biomass types,with poultry litter representing a nonconventional type. Table3 summarizes ultimate analyses of poultry litter while Table 4summarizes its proximate analysis.

Chemical equations for gasication of biomass types

Each biomass type was assumed to gasify according to thefollowing chemical formula [9]:

CnHm O p n pH2O/ n CO m=2 n pH2

The ultimate analysis of each biomass type assuming anabsence of both sulfur and nitrogen was used to calculate then, m, and p values in the equation above. Moreover, thechemical composition of each biomass type varied in terms of hydrogen and oxygen content, and was calculated relative tocarbon molar content. The corresponding chemical formulasfor each of the biomass typeswerefound to be CH 0.01286 O0.1831 ,CH0.00973 O0.6331 , and CH0.01061 O0.8667 for poultry litter, willowpellets, and oak pellets, respectively. Poultry litter can be seento contain the greatest hydrogen and lowest oxygen contentper mole of biomass. Consequently, the chemical equation forthe steam gasication of poultry litter is as follows:

Fig. 1 e CLG 1 simulation block diagram.



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Fig. 2 e CLG 1 process schematic [6].

Fig. 3 e CLG 2 simulation block diagram [3].



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CH0:01286O0:1831s 0:8169 H2Og / COg 0:8234H2g

It is evident that the steam gasication equations for wil-low and oak pellets would be of the same form as poultry litterbut rather with different molar coefcients for steam con-sumption and lesser values for H 2 generation. Poultry litter

generated the greatest theoretical H 2 yield as roughly 0.82 molwere generated permole of biomass andwasthereforechosenas the biomass type to be used for both the CLG 1 and CLG 2simulations.

Simulation input parameters and description

The following sections outline the input data to the ASPENPlus simulation engine as well as the chosen calculationmethods. Moreover, detailed descriptions of the utilizedowsheets are provided.

Setup and calculation methods

The owsheet type was chosen as Solids with metric units,allowing for the analysis and results presentation for solid-state input and output streams. The setup of the owsheetinvolved assigning the MIXCINC stream class to the simula-tion. This allowed uid and aqueous streams (MIXED), con-ventional solid streams (CISOLID), and non-conventional solidstreams (NC) to be input and analysed during simulation runsand calculations.

Subsequently, the process type was chosen as COMMON.This allotted a generic industry type to the simulation, asopposed to chemical, petrochemical, pharmaceutical, etc. The

IDEAL base calculation method wasselected forsimplicityand

thus phase equilibrium calculations were conducted using Raoults Law, Henrys Law, ideal gas law, etc.

Component denition and input

Solid biomass was modelled using a user-dened, non-con-ventional solid based on ultimate, proximate, and sulfur an-alyses. Thus, the input for the poultry litter biomass type wasbased on these parameters. Sulfur analyses e including py-ritic, sulfate, and organic e were set to zero. Furthermore, theenthalpy and density of biomass were approximated using

coal properties. The methods used by ASPEN Plus for thesecalculations are HCOALGEN and DCOALIGT, respectively.Fluid streams were modelled using conventional compo-

nents which have thermophysical data stored in ASPEN Plusdatabanks. Therefore, no data input were required for thesecomponents. The components include: hydrogen (H 2), water(H2O), carbon monoxide (CO), carbon dioxide (CO 2), methane(CH4), and oxygen (O 2).

Additionally, solid components were modelled using con-ventional solids which also have necessary thermophysicaldata stored in the databanks. Tar formation was approxi-mated as solid carbon (i.e. graphite) in the simulation. Thecomponents include: tar (C), calcium oxide (CaO), calcium

carbonate (CaCO 3), iron (Fe), haematite (Fe 2O3), and magnetite(Fe3O4). Calcium-based components were exclusive to the CLG1 simulation while iron-based components were exclusive tothe CLG 2 simulation, with uid components being involved inboth.

CLG 1 owsheet description

The CLG 1 simulationowsheet can be seen in Fig. 4. Theinputand operating conditions for all feed streams and block unitsare summarized in Tables 5 and 6, respectively.

Biomass and water at ambient conditions were fed to thegasier after being heated to the reactor temperature. Thisblock gasied the biomass based on user-dened output forH 2and CO. The output from the gasier was then fed to thereforming reactor where further gasication reactionsoccurred, resulting in H 2, CO, CO2, CH4, tar, and steam for-mation. Subsequently, the output from the reformer was fedto the absorbing reactor in conjunction with a CaO feedstream used for CO 2 absorption and capture. The solid andgaseous products from the absorber were then separated witha gas e solid separator with an efciency of 99.9%. The productgases were heated to the WGS reactor temperature andfurther gasied to convert the majority of the remaining CO toH2. Tar products contained in the WGS reactor exit streamwere removed at a removalefciency of 99.9%, and sent to thedesorbing reactor for steam reforming. The gaseous products

Table 2 e Chemical reactions in syngas chemical looping[3].

Reactor Chemical equation DescriptionGasier Biomass H2O / H2 CO Steam gasication

of biomass.Combustor 4Fe 3O4 O2 / 6Fe2O3 Forms Fe 2O3 for reducer.Reducer 3CO Fe2O3 / 3CO2 Fe

3H2 Fe2O3 / 3H2O 2FeForms Fe for oxidizerand CO and H 2O to beseparated.

Oxidizer 3Fe 4H2O / Fe3O4 H2 Forms product H 2 andFe3O4 for recycle tocombustor.

Table 3 e Ultimate analysis of poultry litter in both thepresence and absence of sulfur and nitrogen.

Element Mass composition (wt.-%)

Sulfur andnitrogen present

Sulfur andnitrogen absent

Carbon 43.30 46.49Hydrogen 6.62 7.11Oxygen 5.95 6.39Nitrogen 5.72 eSulfur 1.15 e

Ash 37.26 40.01

Table 4 e Proximate analysis of poultry litter.

Parameter Mass composition (wt.-%)Moisture Content 20.10Fixed Carbon 3.33Volatile Matter 54.29Ash 22.28



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of this stream were subsequently condensed to remove mostof the remaining steam and small amounts of tar from thesyngas, consequently increasing H 2 purity in the product gas.It is important to note that the gasication process wasmodelled with the combined use of the gasifying, reforming,absorbing, and WGS reactor blocks rather than in a single

reactor step. This was a result of limitations inherent tostandard ASPEN Plus block unit input options and capabilities.Furthermore, the solids stream exiting the initial gas e solid

separator containing CaCO 3, tar, and unused CaO was fed tothe desorbing reactor, along with tar products from the WGSreactor exit. Here, CaCO 3 was thermally dissociated into CaOand CO2, whereas all of the present tar was reformed to CO,CO2, and H 2 with the use of steam. An additional gas e solidseparator was utilized to separate the resulting solids (CaOand remaining tar) and desorber gases. The unused and re-generated CaO stream was cooled to ambient temperatureand sent for re-use in the absorber. Further, the resulting CO 2-rich stream was also cooled to ambient temperature and sent

for sequestration.

CLG 2 owsheet description

The CLG2 simulation owsheet can be seen in Fig. 5. The inputand operating conditions for all feed streams and block unitsare summarized in Tables 7 and 8, respectively.

Biomass and water at ambient conditions were fed to thegasier after being heated to the reactor temperature. Thisblock gasied the biomass based on user-dened output forH 2and CO. The output from the gasier was then fed to thereducer. Conversion of Fe 3O4 to Fe2O3 via combustion using pure O 2 was carried out to facilitate Fe-generating reductionreactions, of which the product Fe was further oxidized using

H2O for H2 production. As such, the Fe 2O3 resulting fromcombustion was combined with the gasication products inthe reducer. Here, Fe 2O3 was reducedto Fe and thegasicationproducts underwent the various gasication reactions,resulting in H 2, CO, CO2, CH4, tar, and steam. The solid-stateand gaseous products were separated in the cyclone at a

solids removal efciency of 99.9%. The gases were then fed tothe reforming reactor while the Fe and tar were fed to theoxidizing reactor for Fe 3O4 regeneration. Furthermore, CH 4completely reformed to CO and H 2 in the oxidizer. Theoxidizer exit gases and solids were then separated in a sec-ondary cyclone, with over 99.8% regenerated Fe 3O4 to berecycled to the combustor and gases fed to the reformer.

Consequently, the reducer and oxidizer exit gases werefurther reformed to H 2 and CO2 in the reformer to increase theH2 yield of the system. Small amounts of CH 4 were regener-ated in the reformer. The reformer exit gases were next fed toa condenser to remove most of the remaining steam and re-sidual solids from the resulting syngas. This increased H 2

purity in the syngas. CO 2 capture was not inherent to CLG 2 aswas the case with CLG 1.

Results and discussion

Important assumptions were initially applied to the simula-tion to ensure that it ran smoothly and produced results.These included:

1. Chosen biomass types contained no nitrogen or sulfur.2. Biomass and steam reacted completely in the rst gasi-

cation reactor (i.e. the gasier) and the only products wereH2 and CO.

Fig. 4 e CLG 1 simulation owsheet.



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3. Tar by-products formed in the various reactors were purecarbon.

4. Tar reforming using steam in the CLG 1 simulation was100% efcient.

5. CO2 desorption from CaCO 3 in the CLG 1 simulation was100% efcient.

6. Gas-solid separators were 99.9% efcient.

The following sections detail the results obtained fromboth simulations using poultry litter. Temperature andpressure sensitivity analyses were conducted for each of themain reactors to determine the optimal operating condi-tions for those reactors. The outputs to the gasier werespecied as part of the simulation in both cases, and thus nosensitivity analysis could be carried out on those blockunits.

Furthermore, the reactor type utilized for the absorber,WGSreactor, desorber, combustor, reducer, oxidizer, andbothreformers was RGibbs. This reactor is an ASPEN Plus block unitwhich calculates its output using the Gibbs free energy mini-mization method. The calculations are based on the chemicalequilibrium reactions of the components being input to thereactor under the specied operating conditions.

Determining the optimal operating conditions

CLG 1 resultsThe reformer in the CLG 1 simulation aimed to produce moreH2 following the gasier via CO conversion. The aforemen-tioned gasication equations occurred simultaneously,resulting in by-product formation in the output stream. Theresults of the reformer temperature sensitivity analysis can beseen in Fig. 6, which shows that H 2 and CO yield at thereformer exit increased with increasing operating tempera-ture. Further, by-product formation, aside from CO, tended todecrease with temperature, and sharply so at the higher endof the temperature scale. These phenomena are the result of the gasication equations proceeding in the reformer. Theforward reactions favour H 2 and CO production and areendothermic. Thus, increasing temperature favoured theformation of these species. As a result, the optimal operating temperature was subject to debate based on the criterion of higher H 2 yield versus greater H 2 purity. The temperature of 750 C was chosen as optimal since H 2 yield was relativelyhigh at this point while CO yield had not yet reached itshighest value. Moreover, by-product yield was relatively lowat this temperature. The trends observed for the reformer

Table 5 e Feed stream input conditions for CLG 1 simulation.

Feed stream Input conditionsTemperature ( C) Pressure (atm) Flowrate (kmol/h) Component

Biomass 25 1 1 a Biomass (Non-conventional)H2O-feed 25 1 1 H2O (Conventional)CaO-feed 25 1 6b CaO (Conventional Solid)

Steam 240 1 85b

H2O (Conventional)a Input as mass owrate (kg/h) using biomass molecular weight.b Fed in excess of required stoichiometric amount.

Table 6 e Block unit operating conditions for CLG 1 simulation.

Block information Operating conditions

Name Type Temperature ( C) Pressure (atm) OtherGasier RYield 750 1 Output based on set values for H 2 and CO (units of kg/kg total feed)Reformer RGibbs 750a 1d eAbsorber RGibbs 500b 1d eWGS-RCTR RGibbs 750c 1d e

Desorber RGibbs 650 1 eHeater Heater 750 1 eHeat-gas Heater 750e 1 eCool-CaO Heater 25 1 eCool-gas Heater 25 1 eG-S-SEP1 Sep e e Separated gaseous components (H 2, CO, CO2, CH4, and H 2O) from 99.9%

of solid components (C, CaO, and CaCO 3).G-S-SEP2 Sep e e Separated DESORBER gases from 99.9% of CaO.G-S-SEP3 Sep e e Separated syngas products from 99.9% of tar (C).Condense Flash2 20 5 e

a Represents optimal operating temperature. Reformer temperature was varied from 500 to 900 C.b Represents optimal operating temperature. Absorber temperature was varied from 300 to 800 C.c Represents optimal operating temperature. WGS reactor temperature was varied from 400 to 1000 C.d Represents optimal operating pressure. Pressure was varied from 1 to 20 atm.e

HEAT-GAS heater temperature was set to the WGS reactor temperature being simulated.

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temperature sensitivity analysis closely reected those foundbyMahisi et al. [8]. However, ethanol(C 2H5OH) was usedas thesource biomass in that case.

Next, the results of the pressure sensitivity analysis on thereformer can be seen in Fig. 7, which illustrates that theopposite phenomenon from temperature sensitivity occurred.Both H2 and CO yield began to decrease rapidly withincreasing pressure, especially within the rst ve atmo-spheres of pressure increase. Correspondingly, other by-product formation increased rapidly within the rst ve at-

mospheres. Due to the relatively large reduction in H 2 yieldand rapid rise of by-product formation at higher pressures,atmospheric pressure was determined to be optimal for thereforming reactor. The trends observed for the pressuresensitivity analysis also closely reected those observed byMahisi et al. [8].

Furthermore, the main goal of the absorber was to capturegenerated CO 2 using CaO as the sorbent. As previouslymentioned, this is an exothermic equilibrium reaction whichtypically occurs at temperatures from 600 to 650 C. Absorberexit yield, CO 2 capture efciency, and CO 2 product yield (i.e.CO2 sent for sequestration) were parameters of consideration.It was noted that H 2 yield tended to decrease at the absorber

exit relative to H 2 in the feed to the reactor. This phenomenon

was deemed a necessary sacrice to ensure total CO 2 capture,and was compensated for further downstream with the use of the WGS reactor.

A temperature sensitivity analysis was also conducted onthe absorbingreactor. Theanalysis showedthat H 2 productionincreased with increasing absorber temperature. Conversely,CO2 production decreased until reaching a local minimum at750 C. Accordingly, CO2 capture efciency slowly decreasedwith temperature before being rendered totally ineffective at750 C. Approximately 500 C was concluded as the optimal

operating temperature for the absorber, based on H 2 and CO2production andCO 2 capture efciency. HigherCO 2 production,relatively high H 2 yield, and a CO 2 capture efciency of over99% were allpresentat this temperature, which occurred priorto the rapid decline in CO 2 capture. Moreover, the pressuresensitivity analysis conducted on the absorber yielded similarresults to that of the reformer, and thus atmospheric pressurewas deemed optimal for the absorber. Both CO 2 productionand capture efciency slightly increased with pressure; how-ever minimal gains, on the order of 0.01%, were observed andwere insufciently benecial to merit operating the absorberat higher pressures.

In addition, CO 2 desorption was totally effective in the

simulation. CaO sorbent recovery was not, since 0.01% of the

Fig. 5 e CLG 2 simulation owsheet.

Table 7 e Feed stream input conditions for CLG 2 simulation.

Feed stream Input conditionsTemperature ( C) Pressure (atm) Flowrate (kmol/h) Component

Biomass 25 1 1 a Biomass (Non-conventional)H2O-in 25 1 1 H2O (Conventional)Fe3O4-in 25 1 0.0370b Fe3O4 (Conventional Solid)O2 25 1 0.10b O2 (Conventional)Steam 240 32 20 H2O (Conventional)a Input as mass owrate (kg/h) using biomass molecular weight.b

Required stoichiometric amount for down-stream reducer reactions to occur.

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sorbent feed was lost due to inefciencies in the gas e solidseparators throughout the process. In an experimentalsetting, however, greater amounts of sorbent would berendered ineffective as a result of calcium sulfate (CaSO 4)formation from any sulfur components present in the

biomass feedstock [8], which was not considered in the con-ducted research. Thus, greater amounts of sorbent regenera-tion would be required in experimental and real-life scenariosthan what is implied by the CLG 1 simulation.

Subsequently, the purpose of the WGS reactor was toprovide a nal block unit for H 2 production as well as to regainH2 that was lost as a result of CO 2 absorption in the absorbing reactor. The lowered temperature required for CO 2 absorptionto proceed was unfavourable to H 2 production and retentionin the absorbing reactor. Both H 2 yield and content signi-cantly increased as a result of the gasication reactions thatoccurred within the WGS reactor. Furthermore, H 2 content inthe syngas product stream was further upgraded with

condenser utilization in the following step. This block uniteffectively liqueed the majority of the steam present in theWGS reactor exit stream, thereby altering the syngas compo-sition in favour of H 2.

The results of the temperature sensitivity analysis con-ducted on the WGS reactor can be seen in Fig. 8. Both H2 andCO yield increased signicantly after 500 C with increasing

temperature, and then began to plateau at higher tempera-tures. Although the CO yield increase with temperature waspronounced, the H 2 to CO yield ratio remained very high dueto low concentrations of CO. Other by-product formationgenerally declined with increasing temperature and was even

negligible in the case of CO 2. Accordingly, the optimal oper-ating temperature for the WGS reactor was found to be 750 C,which provided relatively high H 2 yield while maintaining alower total by-product yield. This temperature is considerablyhigher than the 300 C value observed by Mahisi et al. [8].However, H 2 yield increased following WGS reactor utilizationin both CLG 1 and the simulation conducted by Mahisi et al.The reactor in that simulation accounted for an 8% increase inH2 yield, while the CLG 1 simulation exhibited an approximate170% increase. The substantially larger increase in CLG 1 canbe attributed to the requirement for the regeneration of reduced amounts of H 2 following the CO 2 absorption step.Additionally, the pressure sensitivity analysis conducted on

the WGS reactor demonstrated that atmospheric pressurewas optimal. H 2 production decreased and by-product for-mation increased with pressure, similarly to trends observedfor the previous reactors.

Overall, the optimal reactor temperatures for the CLG 1simulation were 750 C, 500 C, and 750 C for the reformer,absorber, and WGS reactors, respectively. Therefore, a value

Table 8 e Block unit operating conditions for CLG 2 simulation.

Block information Operating conditionsName Type Temperature ( C) Pressure (atm) OtherGasier RYield 750 1 Output based on set values for H 2 and CO (units of kg/kg total feed)Combust RGibbs 1250 1 eReducer RGibbs 870 30 e

Oxidizer RGibbs 720 30 eReformer RGibbs 500 1 eHeater Heater 750 1 eCyclone Sep e e Separated 99.9% of Fe and tar (C) from reducer exit gases (H 2, CO, CO2,

H2O, and CH4).Cyclone2 Sep e e Separated 99.9% of Fe 3O4 from oxidizer exit gases (H 2, CO, CO2, and H 2O)Condense Flash2 20 1 e

Fig. 6 e CLG 1 reformer temperature sensitivity analysis.



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of 750 C was chosen for the gasier as a conservative esti-mate. Moreover, this value is in agreement with some gasiertemperatures utilized in coal-type or wood-type biomassgasication studies outlined in the literature [10e 20], and is80 C warmer than the value of 670 C reported by Acharya [6].

CLG 2 resultsCombustor operating conditions were chosen as 1250 C and1 atm based on data presented by Fan [3]. These conditionsproved sufcient to fully convert Fe 3O4 to Fe2O3 via thermo-oxidation given a stoichiometric excess of pure O 2.

The CLG 2 simulation reducer aimed to generate Fe byreacting Fe 2O3 with the H 2 and CO gasication products. By-product formation was evident in the simulation. Thesephenomena are better illustrated in the temperature sensi-tivity analysis of Fig. 9. The analysis was conducted at 30 atmrather than atmospheric pressure as this was the recom-mended pressure proposed by Fan [3]. Also, Fe formation wasdetermined to be undesirably low at atmospheric pressure.Fig. 9 displays that minimal increases in H 2 and CO yield wereinitially observed before slowly decreasing at the critical point

of roughly 875 C. Conversely, at this critical point, by-productformation (CO 2, H2O, CH4, and tar) tended to increase withtemperature. Fe formation was virtually unaffected by tem-perature variation for the tested range. Therefore, 870 C waschosen as the optimal reducer temperature. This value waschosen due to its slightly lower cost implications whencompared to the critical point of 875 C which was too proxi-mate to thedecreasing portion of the H 2 yieldcurve.Moreover,the optimal operating pressure of 30 atm was conrmed via apressure sensitivity analysis on the reducer, ranging from 30to 37 atm, conducted at the optimal temperature of 870 C. H2and CO production decreased with increasing pressure whileby-product formation increased, similar to the trends of theCLG 1 pressure sensitivity analyses.

In addition, the general purpose of the CLG 2 oxidizer wasboth to produce H 2 in greater quantities than the gasier andto regenerate spent Fe 3O4. The temperature sensitivity anal-ysis for the oxidizer was again conducted at 30 atm ratherthan atmospheric pressure as this was the recommendedpressure proposed by Fan [3]. The analysis demonstrated thattemperature variation only had a signicant effect on CO and

Fig. 7 e CLG 1 reformer pressure sensitivity analysis.

Fig. 8 e CLG 1 WGS reactor temperature sensitivity analysis.



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CH4 yield. CO increased relatively quickly with temperatureand quantities of CH 4 were negligible below about 680 C.Thus, 720 C was chosen as the optimal oxidizer temperatureas only trace amounts CH 4 were present, and negligiblechanges in H 2 yield were observed. Oxidizer pressure wasincrementally increased to 40 atm to determine pressure ef-fects. However, minimal changes in component yield wereobserved in any case, and 30 atm was conrmed as theoptimal operating pressure. Finally 20 kmol/h was chosen asthe design feed steam owrate to the oxidizer after owratevariation analysis was conducted from 5 to 25 kmol/h todetermine the effects on component yields. The analysis wascarried out under steam conditions of 240 C and 32 atm. By-product CO yield tended to decrease with increasing ow-rate, with other component species virtually unaffected, andso a higher owrate was chosen to minimize CO yield.

Subsequently, the CLG 2 reformer was meant to furtherincrease H 2 yield and content via reforming of the remaining by-products. The temperature sensitivity analysis was con-ducted and the optimal operating temperature chosen based

on H 2 yield andby-product levels presentin the syngasstreamexiting the condenser unit at the process termination. Theobserved results from the analysis can be seen in Fig. 10.Itcanbe seen that H 2 and CO2 yield in the syngas begin to decreaseafter reformer temperatures reach roughly 500 C. CO yieldincreased rapidly over the simulated range, albeit at lowerconcentrations throughout. CH 4 yield rapidly decreased be-tween 400 and 500 C. H2 yield tended to peak at approxi-mately 1.60 kmol/kmol PL. In addition, syngas compositionwas closely examined under varying temperature conditions.The respective component curves had roughly the sameshape as the syngas yield curves which are illustrated inFig. 10. H2 and CO2 comprised the majority of the syngas withCO, H2O, and CH4 by-products accounting for smaller per-centages. Again, H 2 composition peaked in the 450 e 500 Crange and began to decrease with further temperatureincrease.

Based on the aforementioned trends, 500 C was chosen asthe optimal operating temperature for the CLG 2 reforming reactor. This is due to both higher H 2 yield and content in the

Fig. 9 e CLG 2 reducer temperature sensitivity analysis.

Fig. 10 e CLG 2 reformer syngas yield temperature sensitivity analysis.



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resulting syngas stream, as well as lower by-product yieldsandcompositions in proximity to this temperature. Increasing operating pressure above atmospheric conditions tended todecrease desirable product yield and correspondinglyincreased by-product yield and concentration. Thus, adetailed pressure sensitivity analysis was not further pursuedfor the CLG 2 reformer.

Comparison of simulation results

The following section details a comparison between the re-sults of the two biomass gasication simulations. Compari-

sons are made based on syngas yield and composition.

Syngas yield comparisonA comparison of the absolute syngas yields for the CLG 1 andCLG 2 simulations can be seen in Fig. 11. Itis evident thatCLG2generated more syngas than CLG 1, at values of roughly 2.54and 0.79 kmol/kmol PL, respectively. However, it is also

important to note that CO 2 removal was not a focal point of CLG 2, and thus almost 0.87 kmol CO 2 /kmol PL adds to theabsolute CLG 2 syngas value. Furthermore, H 2 can be seen tobe the main constituent in both cases, with the CLG 2 syngasproducing more H 2 in absolute terms.

CLG 1 generated 0.73 kmol H 2 /kmol PL while CLG 2 gener-ated 1.60 kmol H 2 /kmol PL. These values are less than thosereported in the literature for either case. Processes similar toCLG 1 reported H2 yields ranging from 1.6 to roughly5.7 kmol H2 /kmol biomass [2,8], while Fan outlined a processsimilar to CLG 2 capable of producing approximately11.73 kmol H2 /kmol coal [3]. The latter process, however, used

coal as the solid fuel and greater amounts of iron-based oxy-gen carriers than CLG 2.Furthermore, the assumption that all biomass and steam

were completely converted to H 2 and CO in the gasier waschallenged, and its effect on syngas yield in both cases wasdetermined. This was done in terms of conversion efciency,labelled as gasier efciency. The analyses showed that the

Fig. 11 e Comparison of simulation syngas yields.

Fig. 12 e Comparison of simulation syngas compositions.



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yields of all syngas components linearly decreased with adecrease in gasier efciency. It is also noted that the reduc-tion in H 2 yield from 100 to 50% gasier efciency is relativelydrastic; 0.73 to 0.36 kmol/kmol PL.

Syngas composition comparisonA comparison of the syngas compositions in either simulation

can be seen in Fig. 12, which also includes individualcomponent compositions. It can be seen that H 2 is the maincomponent in either syngas stream, at 92.45 mol-% and62.94 mol-% for CLG 1 and CLG 2, respectively. The formervalue is greater than the 80.94% H 2 value reported by Acharyaunder similar conditions. Further, the CO 2 concentration of 0.01% in CLG 1 syngas is signicantly lower than that reportedin the literature 5.71% [6]. Other studies similar to CLG 1 re-ported lower obtained H 2 concentrations, ranging fromroughly 70 e 90 mol-% [2,7,8].

However, as previously mentioned, CO 2 was not removedfrom the syngas stream in CLG 2, and thus its relatively largeconstituency of 34.11 mol-% is accounted for by that fact. HadCO2 removal been inherent to the design of CLG 2, andassuming removal efciency upwards of 99%, the H 2 compo-sition ofthe CLG2 syngas would be over 95 mol-%. Further, the62.94 mol-% value from CLG 2 is similar to the 62.1 mol-%found in the literature [3].

Future research

Potential for future research stems in-part from the simula-tion owsheets and block unit setups themselves. Deter-mining a method to model biomass gasication in a singlegasifying reactor, as opposed to four separate reactors in thecase ofthe CLG1 process, is exemplar of this. This would allowfor biomass gasication to occur spontaneously in anequilibrium-based reactor, i.e. RGibbs, rather than in a reactorwith user-dened outputs, as is the case for both simulations.Further, thiswould allow foran investigationof the sensitivityof the gasifying reactor to steam-to-biomass, calcium-to-biomass, and equivalence ratios.

In addition, the nitrogen and sulfur components of eachbiomass type could be included in calculations to furtherenhance the accuracy of the simulation. Consequently, thiswould require modication of the owsheet and input pa-rameters to deal with any sulfur dioxide (SO 2), hydrogen sul-de (H2S), or nitrogen (N 2) streams that may be present fromgasication of biomass containing these elements [21].

Moreover,more detailed design of the gas e solidseparators(e.g. cyclones) utilized in the simulation would further in-crease its accuracy. For example, cyclones, baghouses, orelectrostatic precipitators (ESPs) could be designed forgase solid separation. These process units would be morerepresentative of an industrial-scale scenario, as opposed tothe generic separator blocks used in this simulation.Furthermore, the energy requirements of such block unitscould be accurately represented as well.

Finally, a cost analysis for each block unit could be con-ducted to obtain an overall syngas production cost for eachsimulated process type. The cost analysis could be based onenergy requirements and feed stream raw materials supply

costs. Subsequently, comparing the cost estimates wouldprovide an idea of the feasibility of the proposed processesfrom a practical standpoint.

Conclusions

In conclusion:

1. Poultry litter was the chosen biomass type to be simulatedboth because it is a nonconventional biomass type and dueto its greater H 2 yield potential when compared with wil-low pellets and oak pellets.

2. The optimal operating condition estimates determined forthe main reactors in both simulations were in line withthose presented in the literature.a. CLG 1 simulation:

i. Reformer: 750 C, 1 atm.ii. Absorber: 500 C, 1 atm.

iii. WGS reactor: 750 C, 1 atm.b. CLG 2 simulation:

i. Combustor: 1250 C, 1 atm.ii. Reducer: 870 C, 30 atm.

iii. Oxidizer: 720 C, 30 atm.iv. Reformer: 500 C, 1 atm.

3. CLG 1 and CLG 2 syngas yields were 0.79 and 2.54 kmol/kmol PL, respectively. CLG 2 generated the most H 2 in theproduct syngas stream, 1.60 kmol/kmol PL, based on ab-solute yield, with CLG 1 producing only 0.73 kmol/kmol PL.H2 production was signicantly less than that outlined inthe literature for both simulations.

4. CLG 1 produced purer syngas with an H 2 concentration of 92.45 mol-%, while CLG 2 had 62.94 mol-% H2. The lowerCLG 2 concentration was due to the presence of CO 2 in thatsyngas stream, as its removal was not a focus of thatsimulation. CLG 1 exhibited more H 2 rich syngas than otherstudies while CLG 2 produced results similar to those foundin the literature.

5. Future research could focus on increasing the accuracy andscalability of the simulations through assumption mitiga-tion or removal. Examples include modelling biomassgasication in a single gasifying reactor, detailed design of gas e solid separators, and inclusion of nitrogen and sulfurelements in biomass ultimate analyses.

Acknowledgements

The authors would like to acknowledge funding from NSERCvia the Discovery Grant Program. The authors would like tothank Miss Leanna Harnarain for proofreading the manu-script and for improving its appearance.

Appendix A. Supplementary data

Supplementary data related to this article can be found athttp://dx.doi.org/10.1016/j.ijhydene.2014.01.178 .

http://dx.doi.org/10.1016/j.ijhydene.2014.01.178http://dx.doi.org/10.1016/j.ijhydene.2014.01.178http://dx.doi.org/10.1016/j.ijhydene.2014.01.178http://dx.doi.org/10.1016/j.ijhydene.2014.01.178http://dx.doi.org/10.1016/j.ijhydene.2014.01.178http://dx.doi.org/10.1016/j.ijhydene.2014.01.178


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