Development of a process model for coal chemical looping combustion and validation against 100 kWth tests

Applied Energy 157 (2015) 433–448

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Development of a process model for coal chemical looping combustion and validation against 100 kWth tests Peter Ohlemüller ⇑, Falah Alobaid, Adrian Gunnarsson, Jochen Ströhle, Bernd Epple Technische Universität Darmstadt, Institut Energiesysteme und Energietechnik, Otto-Berndt-Str. 2, 64287 Darmstadt, Germany

h i g h l i g h t s  Coal CLC process simulation model was developed, considering main unit operations.  Experimental data obtained from 100 kWth CLC tests is used for model validation.  Good agreement was found between the modeling results and the measurements.  Sensitivity analyses for different parameters of the validated model were performed.

a r t i c l e

i n f o

Article history: Received 20 November 2014 Received in revised form 22 May 2015 Accepted 23 May 2015 Available online 23 June 2015 Keywords: Chemical looping combustion Process simulation Aspen plus 100 kWth CLC pilot plant Validation study Sensitivity analysis

a b s t r a c t Chemical looping combustion is a very efficient CO2 capture technology utilizing two interconnected circulating fluidized beds. Despite promising results in basic research on chemical looping combustion at laboratory scale, the technical process has to be further developed, especially toward a potential industrial application. In order to accelerate this development, it is essential to simulate chemical looping combustion with process simulation software. In this study a process simulation model is validated using experimental data obtained from a 100 kWth pilot plant at Chalmers University of Technology (Sweden). The targeted configuration of the process is created and all substances and boundary conditions are defined. The solids distribution and also the chemical conversion of gases and solids are simulated by a user defined model. This model is based on mathematical equations for fluidized beds and kinetic data of the chemical reactions is taken into consideration. The one-dimensional solids distribution is validated with experimental measurements using empirical correlations according to Kunii and Levenspiel. Conversions in the dense and the lean zone of the fluidized bed are determined using kinetic data from literature. The elaborated model is validated against the concentrations of gases at the fuel reactor exit. The calculated results are in very good agreement with experimental data. Sensitivity analyses are performed to optimize the operational conditions of the pilot plant. Ó 2015 Elsevier Ltd. All rights reserved.

1. Introduction 1.1. Concept of chemical-looping combustion Chemical-looping combustion (CLC) is one of the most promising carbon capture and storage technologies [1]. Conventional CO2 capture processes like oxyfuel combustion or post-combustion CO2 capture by MEA scrubbing are always associated with substantial energy losses [2]. In the chemical looping process, the furnace of a conventional coal fired power plant is replaced by two circulating fluidized bed reactors (see Fig. 1). In this process a particulated oxygen carrier (MexOy) circulates between these two interconnected ⇑ Corresponding author. Tel.: +49 (0) 6151/16 76966; fax: +49 (0) 6151/16 5685. E-mail address: [email protected] (P. Ohlemüller). http://dx.doi.org/10.1016/j.apenergy.2015.05.088 0306-2619/Ó 2015 Elsevier Ltd. All rights reserved.

reactors. The oxygen carrier becomes oxidized by the oxygen in the air which is used for the fluidization of the first reactor (AIR REACTOR).

2Mex Oy1 þ O2 ! 2Mex Oy

ð1Þ

The oxidized particles are subsequently transferred into the second reactor (FUEL REACTOR). In this reactor a solid fuel such as hard coal is gasified by either steam (cf. Eq. (2)) or carbon dioxide (cf. Eq. (3)) and the oxygen carrier is reduced by the products from devolatilization and gasification products (CO, H2, CH4) leading to the formation of carbon dioxide and water (cf. Eqs. (4)–(6)).

C þ H2 O ! CO þ H2

ð2Þ

C þ CO2 ! 2CO

ð3Þ

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P. Ohlemüller et al. / Applied Energy 157 (2015) 433–448

Nomenclature a A A0 Ar bi

decay constant cross section area of the fuel reactor (m2) area of the gas-distributor per nozzle (m2 per nozzle) Archimedes number stoichiometric coefficient for metal oxide in combustion of gas i ci concentration of gas i (mol m3) dp particle diameter (m) db diameter of bubble (m) fb empirical function GS⁄ saturated mass flux of solids (kg m2 s1) H height (m) Hd, Hl, Hr height of the dense zone, the lean zone and the reactor (m) k0,OC preexponential factor, OC reduction (mol1n m3n2 s1) k0,i preexponential factor, gasification reaction (s1) K0,i kinetic constants in the gasification rate (bar1) ni reaction order of component i molar flow of component i (kmol h1) n_ i OC oxygen carrier pi partial pressure of the component i (bar) rN,i reaction rate of ilmenite or char (N) with the component i (s1) rg grain radius (m) ROC oxygen transport capacity of the oxygen carrier u0 superficial gas velocity (m s1) ub1 velocity of a single bubble (m s1)

CO þ Mex Oy ! CO2 þ Mex Oy1

ð4Þ

H2 þ Mex Oy ! H2 O þ Mex Oy1

ð5Þ

CH4 þ 4Mex Oy ! CO2 þ 2H2 O þ 4Mex Oy1

ð6Þ

After oxidation of unburned flue gases downstream the fuel reactor, power generation, cleaning of the exhaust gases and the condensation of steam, the remaining pure CO2 can be compressed and stored. The oxidation would take place in a so-called post-oxidation chamber (POC) where pure oxygen would be provided for this process. 1.2. Experimental work Chemical looping combustion with solid fuels is researched from laboratory scale [3–12] up to MWth scale [13,14]. All CLC

umf ut uvis XC XOC yi z db

es es ⁄ esd ese gOO gCC lg qm qg or qs s U0

w XOD XT

gas velocity at minimum fluidization (m s1) terminal velocity of a falling particle (m s1) visible gas flow in the bubbles (m s1) conversion of char conversion of the oxygen carrier dry concentration of the component i height of vertical position in the reactor (z) volumetric fraction of bubbles in the bottom bed volume fraction of solids maximum volume fraction of solids that can be pneumatically transported volume fraction of solids in the lower dense region of a fluidized bed volume fraction of solids at the reactor exit oxide oxygen fraction (%) carbon capture efficiency (%) viscosity of gas (kg m1 s1) molar density (mol m3) density of gases or solids (kg m3) time for complete solid conversion of the oxygen carrier (s) ratio of moles of oxygen needed to convert the fuel completely per moles of carbon in the fuel ratio of the visible bubble flow to the total flow through the bubbles oxygen demand (%) total oxygen demand (%)

pilots with a thermal power up to 100 kWth are electrically heated. There are many results from different lab scale CLC units. At IFP in France, a 10 kWth unit was constructed and successfully operated with methane as fuel gas [15]. It consists of two air reactors and one fuel reactor which are operated as bubbling beds to achieve optimum reaction conversion. During operation of a 25 kWth CLC unit at Hamburg with a two-stage fuel reactor high carbon capture rates >96 % were achieved [16]. A 50 kWth CLC was designed and constructed to be operated in both iG-CLC and CLOU modes [17]. High CO2 capture efficiencies are predicted depending on the carbon stripper efficiency and OC material. There are further promising results from the electrically heated 100 kWth unit at Chalmers University of technology (Sweden) [5] and the 1 MWth pilot plant at Technische Universität Darmstadt (Germany) with partial chemical looping conditions [13]. During partial chemical looping conditions a certain amount of air was fed into the fuel reactor together with steam.

Oxygen depleted air

CO2 and H2O MexOy

AIR REACTOR

FUEL REACTOR Fuel

MexOy-1 Steam

Air

Fig. 1. General scheme of the CLC process.

P. Ohlemüller et al. / Applied Energy 157 (2015) 433–448

There are various oxygen carriers like active oxides of iron, copper, nickel or manganese [18,19]. Ilmenite, a natural titanium-iron oxide (FeTiO3), is one of the most studied oxygen carriers [4,5,12,13]. The high availability, low environmental and health impairments and the low costs are the advantages of this material. 1.3. Chemical-looping process simulation Beside experiments, chemical-looping process simulation models are of great significance of the further development of the process and a possible application in the future. Validated chemical-looping process simulation models can be applied to optimize the operational conditions of a pilot plant or can be used as a basis for scale-up studies. Many publications about CLC process simulation for solid fuels exist in the literature [20–32] but only a few of these have been compared to experimental results from continuously operated units. Abad et al. [21,22] developed a 1.5 D fuel reactor model of a 100 kWth CLC unit. The numerical results were compared toward experimental data reported from Markström et al. [5], showing a good agreement. A carbon separation unit was introduced by Gayán et al. and several technological options were investigated [30]. The integration of both fuel and air reactor are only found for gaseous fuels and validated with results obtained in a 120 kWth CLC unit for gaseous fuels at Vienna [31,32]. However, a process simulation model for coal chemical looping including fuel and air reactor, carbon stripper, cyclones and post-oxidation chamber is not available in literature. In this work, a coal CLC process simulation model is developed and validated against experimental results obtained from the 100 kWth pilot plant at Chalmers University of Technology [5]. 34 different operational points with e.g. different mass flows of

435

coal, different pressure drops of the fuel reactor or different compositions of fluidization agents were used for this validation. The predictions of the model are in very good agreement with experimental data, even for operational points with very different boundary conditions. To optimize the operating conditions of the 100 kWth pilot plant, some sensitivity analyses were performed. For a future work or the evaluation of the CLC process this model could be used as a basis for upscale studies. Therefore, it could be extended by e.g. fans, a water-steam-cycle and CO2 compression. 2. Model development 2.1. Description of the process flow sheet The developed CLC process model is based on the 100 kWth unit at Chalmers University of Technology [5]. 2-D and 3-D sketches of this unit are shown in Fig. 2, and the corresponding flow sheet of the process simulation model is presented in Fig. 3. The 100 kWth unit consists of two reactors, namely the air reactor (AR) and the fuel reactor (FR), a carbon separation unit (CS/carbon stripper), and a circulation riser (CR) that enables the transport of material from fuel reactor to carbon stripper/air reactor. Furthermore there are four loop seals (LS 1-4) and three cyclones (CY1-3). The flow sheet rebuilds this test facility:  The coal (COAL) that is fed into the system is decomposed into its products from decomposition (DECOMP).  The products from devolatilization including char (DEVOLAD), the oxygen carrier (OCO-TOFR) from the air reactor, steam (STEAM), char particles from the carbon stripper (CS-FR) and nitrogen for fluidization of the loop seal LS2 and the screw

Fig. 2. A 2-D sketch (left) and a 3-D sketch (right), drawn to scale, of the 100 kWth unit at Chalmers [5].

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P. Ohlemüller et al. / Applied Energy 157 (2015) 433–448 DECOMP POC COAL FRFLUE

DEVOLAD

OCLOSS GASES

ARFLUE

CS-FR

AR

CYCLONE

CS AR-OUT

OCR-TOAR

FR

SOL-TOCS FR-OUT

SPLIT

MAKEUP AIR

ASH SCREWFLU OCO-TOFR STEAM

Fig. 3. Flow sheet of the process simulation model.









 



conveyor (SCREWFLU, ‘‘screw fluidization’’) enter the fuel reactor. Loop seal LS2 was fluidized with nitrogen to avoid the gas production from char gasification inside the downcomer from the fuel reactor cyclone [5]. Inside the fuel reactor, char is gasified and the products from devolatilization and gasification are converted by the oxygen carrier. To take the characteristics of a fluidized bed and also kinetics into consideration, all conversion rates are calculated by a FORTRAN code. The products of the fuel reactor are separated in gases (GASES) and solids (SOL-TOCS). A part of the char leaves the process to the post-oxidation chamber (POC). The carbon stripper (CS) works as a split: Ash is removed from the system (ASH), char is either led back to the fuel reactor (CS-FR) or burned in the air reactor (OCR-TOAR) and the oxygen carrier is also transferred to either the air reactor or the fuel reactor to take the internal circulation into consideration. To control the circulating mass flow, 1% of the circulating oxygen carrier is removed from the process (CS, OCLOSS) and the same mass is fed to the air reactor (AR, MAKEUP). This is unavoidable in Aspen Plus to achieve the target circulation flow. Inside the air reactor, the oxygen carrier is fully oxidized with air and the char from the carbon separation unit is burned. The products of the air reactor (AR-OUT) are separated at the cyclone (CYCLONE): Flue gases leave the system (ARFLUE) and solids (oxidized oxygen carrier) are transferred to the fuel reactor. Inside the post oxidation chamber carbon (CPOC) and unconverted gases (GASES) are converted with oxygen.

2.2. Air reactor In the air reactor two reactions take place, namely the oxidation of ilmenite (cf. Eq. (7)) and the combustion of carbon entering the reactor (cf. Eq. (8)):

4FeTiO3 ðsÞ þ O2 ðgÞ ! 2Fe2 TiO5 ðsÞ þ 2TiO2 ðsÞ

ð7Þ

CðsÞ þ O2 ðgÞ ! CO2 ðgÞ

ð8Þ

The conversions of both reactions are set to 100% because the reactions are very fast [33]. There is an average air–fuel ratio of 2, and the average residence time of solids in the AR of the 100 kWth unit is very high (6 min, calculated from [5]). Hence, a stochiometric reactor is used for the air reactor in the process simulation model. 2.3. Fuel reactor When feeding coal, steam and oxygen carrier in the fuel reactor, three processes take place: Pyrolysis of coal (cf. Eq. (9)), gasification of char (fixed carbon, see Eqs. (10)–(13)), and oxidation of reducing gases by the oxygen carrier (cf. Eqs. (14)–(16)), which corresponds to the reduction of the oxygen carrier. All hydrocarbons are assumed to be methane. Pyrolysis of the coal:

CoalðsÞ ! H2 OðgÞ þ Volatile matterðgÞ þ CðsÞ þ ashðsÞ

ð9Þ

Gasification of char:

CðsÞ þ CO2 ðgÞ ! 2COðgÞ

ð10Þ

CðsÞ þ H2 OðgÞ ! H2 ðgÞ þ COðgÞ

ð11Þ

COðgÞ þ H2 OðgÞ ! CO2 ðgÞ þ H2 ðgÞ

ð12Þ

Eqs. (11) and (12) can be combined to Eq. (13) since the water– gas-shift is not modeled in detail [34]. Thus, the conversion from Eq. (11) is applied to calculate both, the gasification of char particles with steam [22] as well as the water–gas-shift reaction.

CðsÞ þ 1:25H2 OðgÞ ! 0:75COðgÞ þ 0:25CO2 ðgÞ þ 1:25H2 ðgÞ

ð13Þ

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P. Ohlemüller et al. / Applied Energy 157 (2015) 433–448

Oxidation of reducing gases/reduction of the oxygen carrier:

CH4 ðgÞ þ 4Fe2 TiO5 ðsÞ þ 4TiO2 ðsÞ ! 2H2 OðgÞ þ CO2 ðgÞ þ 8FeTiO3 ðsÞ

ð14Þ

H2 ðgÞ þ Fe2 TiO5 ðsÞ þ TiO2 ðsÞ ! H2 OðgÞ þ 2FeTiO3 ðsÞ

as well as the actual mass (mOC) of the oxygen carrier. The kinetic constants for the oxidation reactions are summarized in Table 2 [35], and the properties of ilmenite are given in Table 3 [5,35].

r ilmenite;i ¼

ð15Þ

si ¼ COðgÞ þ Fe2 TiO5 ðsÞ þ TiO2 ðsÞ ! CO2 ðgÞ þ 2FeTiO3 ðsÞ

ð16Þ

All these reactions take place in the fuel reactor. The kinetics of the chemical reactions (gasification, oxidation) have a major influence on the flue gas composition of the fuel reactor. Kinetic data from the literature were integrated in the fuel reactor code. The reaction rates of the gasification of char are calculated by Eqs. (17) and (18) based on the partial pressure (pi) of steam or CO2, the kinetic constants (ki, Ki) and the conversion of carbon (XC). The homogenous reaction model with control by chemical reaction is used [22].

r C;H2O ¼

dX C;H2O kH2O pH2O ¼  ð1  X C Þ dt 1 þ K H2O pH2O þ K H2 pH2

ð17Þ

r C;CO2 ¼

dX C;CO2 kCO2 pCO2 ¼  ð1  X C Þ dt 1 þ K CO2 pCO2 þ K CO pCO

ð18Þ

The kinetic constants are calculated by the Arrhenius equation (cf. Eq. (19)). The carbon conversion is needed for the determination of the reaction rates and is obtained from Eq. (20). The kinetic constants for coal gasification are summarized in Table 1. EA;i

DH i

ki ¼ k0;i  e RT XC ¼

K i ¼ K 0;i  e RT

ð19Þ

mC;in  mC;out mC;out ¼1 mC;in mC;in

ð20Þ

The reaction rates of ilmenite with reductive gases (cf. Eq. (21)) are calculated using the transport capacity of the oxygen carrier (ROC), time for complete solid conversion (si), and the conversion of the oxygen carrier (XOC). The time for complete solid conversion (cf. Eq. (22)) is calculated by the molar density (qm), the grain radius (rg), the stoichiometric coefficient for metal oxide conversion (bi), the kinetic constants (ki), the concentration of the reductive gases (ci) and the reaction order of each component (ni). The conversion of ilmenite is calculated by Eq. (23). The mass of the complete oxidized (mOC,ox) and reduced ilmenite (mOC,red) is needed Table 1 Kinetic constants of gasification reactions [22]. H2O k0;H2O EA;H2O K 0;H2O DHH2O K 0;H2 DHH2

CO2 52.6 95.1 2.81  106 135.1 8.1  109 218.5

k0;CO2 ECO2 K 0;CO2 DHCO2 K 0;CO DHCO

Unit 3

s1 bar1 kJ mol1 bar1 kJ mol1 bar1 kJ mol1

4.53  10 160.1 3.28  107 158.5 1.84  106 157.6

dX OC;i 3 ¼ ROC ð1  X OC Þ2=3 dt si

qm r g n bi ki c i i

X OC ¼

ð22Þ

mOC;ox  mOC RO mOC;red

ð23Þ

2.3.1. Hydrodynamic model For the execution of the simulations, a fluidized bed reactor model was used, where the reactor is divided in two reaction zones: A dense zone at the bottom of the reactor and a lean zone above the dense. All input streams enter the fuel reactor at the bottom. An average conversion is calculated for both the oxygen carrier and the char for both zones. The conversion of reductive gases is first calculated for the dense zone and then for the lean zone considering the conversions in the bottom part. The heights of the dense and the lean zone are calculated according to the model of Kunii and Levenspiel [36]. The bottom zone is characterized by a very high volume fraction of solids (esd) while the lean zone is located above the dense zone with a decaying solid concentration with the reactor height (es (H), H > Hd) (cf. Eqs. (24) and (25)). The volume fraction of solids in the lean zone (es) depends on the maximum volume fraction of solids that can be pneumatically transported (es ), the volume fraction of solids in the dense zone (esd), the height of the dense zone (Hd) and a decay constant (a). The volume fraction of solids at the fuel reactor outlet (ese) can be calculated by Eq. (25) using the full height of the fuel reactor (Hr, see Eq. (26)).

es ðHÞ ¼ esd ; H 6 Hd

ð24Þ

es ðHÞ ¼ es þ ðesd  es ÞaðHHd Þ ; H > Hd

ð25Þ

es ðH ¼ Hr Þ ¼ ese

ð26Þ

The height of the lean zone is calculated by a numerical approximation using Eq. (27). This equation includes the mass inventory (minv), the decay constant (a), the cross-sectional area of the fuel reactor (AFR), the density of solids (qs), different specific volume fractions of solids in the fluidized bed (esd, e⁄s , ese) and the height of the lean zone (Hl) and the reactor (Hr). Eq. (27) is derived by Eq. (28). The maximum volume fraction of solids that can be pneumatically transported (es , cf. Eq. (29)) is determined by the saturated mass flux of solids (Gs , cf. Eq. (30)), the superficial gas velocity (u0), the terminal velocity of a falling particle (ut) and the gas and the solids density (qi). The terminal velocity (ut) is determined according to Kunii and Levenspiel [36] by the dimen sionless particle diameter (dp ), and the dimensionless gas velocity  (ut , see Eq. (31)).

CONST ¼ eaHl  aHl ¼

Table 2 Kinetic constants of reductive gases [35].

ð21Þ

minv a AFR qs

 esd þ es  Hr esd a

ð27Þ

ðesd  es Þ

Table 3 Properties of ilmenite [5,35].

Kinetic constant

CH4

CO

H2

Unit

Property

Symbol

Value

Unit

b k0;OC EOC n

5.78 9.8 135 1.0

1.45 0.01 81 0.8

1.45 0.062 65 1.0

– mol1n m3n2 s1 kJ/mol –

Average particle diameter Effective particle density Oxygen transfer capacity (activated) Grain radius

dOC

171 3600 3.3 1.25

lm

qOC ROC rg

kg/m3 wt.% lm

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P. Ohlemüller et al. / Applied Energy 157 (2015) 433–448

minv esd  ese ¼ þ Hr  esd  Hl ðesd  es Þ AFR qs a

es ¼

Gs ðu0  ut Þqs

ð29Þ ut

Gs ¼ 23:7qg u0 eð5:4uo Þ " ut ¼ ut

ð28Þ

lðqs  qg Þ  g q2g

ð30Þ #1=3 ð31Þ

The lean zone of the fluidized bed is assumed to be perfectly mixed and the dense zone is divided in bubbles (only gas) and an emulsion phase [20]. The bubbles are considered to be free from solids and no solid/gas reaction takes place inside the bubbles. It is therefore important to calculate the volume fraction of bubbles in the bottom bed, since no mixture of gases between bubbles and emulsion phase is assumed. This has a major influence of the conversion of unburned gases in the dense zone. The volumetric fraction of bubbles (db) is calculated by Eq. (32), where uvis is the visible bubble flow and ub1 is the single bubble velocity calculated according to Eqs. (33) and (34) [20].

db ¼

uv is uv is þ ub1

ð32Þ

uv is ¼ wðug  umf ð1  db ÞÞ

ð33Þ

qffiffiffiffiffiffiffiffi ub1 ¼ 0:71 gdb

ð34Þ

The symbol w is the ratio of visible flow and the total flow through the bubbles (cf. Eq. (35)) and db (cf. Eq. (36)) is the bubble diameter as a function of the bed height z (z = Hd/2), the area of the gas-distributor per nozzle (A0) and a coefficient fb (cf. Eq. (37)) which depends on the particle diameter (dp), superficial (u0) and minimum fluidization velocity (umf, see Eq. (38)). The Archimedes number (Ar) is defined by Eq. (39). The fraction of bubbles which is the most important information is calculated by iteration of the named equations.

pffiffiffiffiffi 0:4 w ¼ f b ðz þ 4 A0 Þ

ð35Þ

pffiffiffiffiffi 0:8 db ¼ 0:54ðug  umf Þ0:4 ðz þ 4 A0 Þ g 0:2

ð36Þ

fb ¼

0:26 þ 0:70  e3300dp

ð37Þ

1

ð0:15 þ ug  umf Þ3 qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi  C 21 þ C 2 Ar  C 1  lg

umf ¼

qg dp

;

C 1 ¼ 27:2; C 2 ¼ 0:0408

ð38Þ

to be lost as fly ash at the fuel reactor cyclone. In the model, it is removed from the process in the carbon separation unit since this is easier to execute (cf. Fig. 3). 2.3.2. Products from devolatilization The products from devolatilization are recalculated according to Matthesius et al. [37]. Propane and the tar are considered to be transformed to methane and reformed to carbon monoxide and hydrogen by steam and carbon dioxide (see Table 4). 2.3.3. Split factors The split factors of the fuel reactor cyclone (SPLIT, see Fig. 3) and the carbon separation unit (CS) include the loss of char at the cyclone and the internal circulation char particles in the fuel reactor. To obtain good results it is absolutely necessary to select good values for these factors. In the model the efficiency of the fuel reactor cyclone regarding char particles is 96% and the carbon separation efficiency was chosen to be 99.8% including the internal recirculation of the fuel reactor. The carbon separation efficiency was not experimentally determined [5]. Hence, a value was chosen to achieve good results and to be in agreement with experimental results found in literature [21]. 2.4. Experimental data Experimental data were collected and reported by Markström et al. [5] for a 100 kWth pilot plant at Chalmers Technical University, Sweden, using ilmenite oxygen carrier and a Columbian bituminous coal (Cerrejón coal, analysis see Ref. [5]). The configuration was shown in Fig. 2. Experimental data are used for validation of process models. Operational parameters for set-up of the simulations are gathered from the report including temperature, pressure drops and mass flows at the two reactors. To achieve a good comparison between the model and the experimental data, it is desirable to perform the comparison at stable operative conditions. Therefore some stable experimental points are selected for the comparison and validation of the model. The experiments consist of five operational periods from VI to X, and a total of 34 points are selected from the operational periods VI to X where the measured values are stable for at least 5 min. The necessary parameters are collected from Tables 1 and 4–7 as well as from Figs. 6, 9, 10 and 13 or as stated in the text [5]. The nitrogen flow to the fuel reactor is given as an range of 4.5–6.8 kg/h. It was chosen to be 5.2 kg/h which is in the given range independent on the coal feed. At operational point X-1, the steam flow to the carbon stripper was replaced by nitrogen. The nitrogen mass flow was calculated assuming a constant molar flow of the gases carbon monoxide, hydrogen, methane and carbon dioxide of operational points X-1 and X-2 (see Eq. (41)) resulting in a nitrogen mass flow of 30.3 kg/h.

3

Ar ¼

dp  qg  ðqs  qg Þ  g

l2

ð39Þ

The mass inventory has a great influence on the hydrodynamic model (see Eqs. (27) and (28)). It is obtained from the pressure drop and the cross-sectional area of the fuel reactor [36]:

minv ¼

0:85  Dp  A g

ð40Þ

The inventory of the fuel reactor consists of both oxygen carrier (ilmenite) and char including fixed carbon and ash. The ratio of char and oxygen carrier particles in the reactor is not known and calculated from the inlet streams of each component. Char particles are assumed to consist of ash and fixed carbon. Ash is assumed

Table 4 Products from devolatilization of coal (in wt.%, El Cerrejón coal) [37]. Product

Mass fraction (wt.%)

Carbon Ash CH4 CO CO2 H2 H2O N2 SO2

48.20 8.60 7.76 25.44 1.77 4.00 1.74 1.30 1.20

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n_ N2 ;OP X-1 ðy þ yH2 þ yCH4 þ yCO2 ÞOP X-2 ¼  CO ðyCO þ yH2 þ yCH4 þ yCO2 ÞOP X-1 n_ N2 ;OP X-2 

3. Results and validation

1  ðyCO þ yH2 þ yCH4 þ yCO2 ÞOP X-1 1  ðyCO þ yH2 þ yCH4 þ yCO2 ÞOP X-2

ð41Þ

As an approximation, the steam flow to the fuel reactor was calculated from the steam flows to the loop seals 1 and 3, to the carbon stripper, and to the circulation riser. The circulating mass flow of oxygen carrier is set to be the mass flow calculated from the pressure drop in the air reactor riser [5], even if it could contain errors. The temperature is set as an averaged value if not exactly known, and the pressure drop is given for each operational point. The oxidation of oxygen carrier in the air reactor is assumed to be 100% because of the long residence time and the fast reaction rate (see Section 2.2). All operational conditions are summarized in Table 6. The oxygen demand, XOD, is the fraction of oxygen lacking for a complete conversion of gases formed by coal pyrolysis and gasification (cf. Eq. (42)). U0 is the ratio of moles of oxygen needed to convert the fuel completely per moles of carbon in the fuel [5] (U0 = 1.179 for El Cerrejón coal). Oxide oxygen fraction, gOO, describes the ratio of oxygen needed for oxidizing the oxygen carrier in the air reactor divided by the total amount of oxygen consumed (see Eq. (43)) [3]. The gases H2S and SO2 are neglected for these calculations.

XOD ¼

gOO ¼

ð0:5  yCO þ 0:5  yH2 þ 2  yCH4 ÞFR;out

U0  ðyCO2 þ yCO þ yCH4 ÞFR;out

;

U0 ¼ 1:179

ð0:21  yO2  yCO2 ÞAR;out

ð42Þ

ð43Þ

ð0:21  yO2  0:21yCO2 ÞAR;out

Two further evaluation parameters are the total oxygen demand, XT, and the carbon capture efficiency, gCC. These have already been calculated by Abad et al. [21]. However, the loss of fixed carbon should be considered in the calculation of the oxygen demand since oxygen would be needed to oxidize these particles.

In this section the predictions of the CLC process simulation model are validated against experimental data from a 100 kWth unit [5] and furthermore compared to a 1.5 D fuel reactor model [21]. The input parameters from Table 6 are used to predict the concentrations of the components CO2, H2, CO and CH4 as well as the oxygen demand, XOD, and oxide oxygen fraction, gOO. All results are calculated using exactly the same model by only changing the input parameters. 3.1. Hydrodynamic model To determine the solids distribution inside the fuel reactor the volume fraction of solids in the dense zone (esd = 0.4) and the product of the decay constant and the superficial gas velocity (au0 = 5) were determined (see Fig. 4). The one dimensional pressure profile calculated from the model equals the measured pressure drops at different heights. The height of the dense region was calculated separately for all operational conditions. For this specific case it was 0.7 m. 3.2. Flue gas concentrations The dry concentrations of the gases carbon dioxide, hydrogen, carbon monoxide and methane at the fuel reactor outlet as well as the oxygen demand and the oxide oxygen fraction are graphically shown from Figs. 5–10 (CLC process model: red diamonds). The experimental results and the results from a 1.5 D model are also represented in order to evaluate the models predictions (experimental results: black squares; 1.5 D model: blue triangles). It should firstly be noted that there is a really good agreement between the CLC process model developed in this work and the experimental results. The courses of the different components depending on the operational points with different conditions are described below.

5

U0  n_ C;fuel feed

¼ 1:179

gCC ¼

ðn_ CO þ n_ CH4 þ n_ CO2 þ n_ Cfix ÞFR;out n_ C;fuel feed

;

U0 ð44Þ

4

ð45Þ

3

The gas concentrations of CO2, CO, H2 and CH4 in the outlet stream from the fuel reactor as well as the oxygen demand and the oxide oxygen fraction were used for the comparison and validation of the model. These values are taken from Figs. 4, 5, 7, 8 and 11–13 in Ref. [5]. Since the oxygen demand is quite sensitive to gas concentrations of gases in the outlet stream, it is recalculated rather than read out from a figure. Furthermore, sulfur has been neglected for these calculations. These results are summarized in Table 7. The predictions of the CLC process model and the experimental findings are also compared to the results from the fuel reactor model from Abad et al. using a 1.5 CFB model for the fuel reactor. These values can be found in Figs. 6 and 7 in Ref. [21]. The oxygen demand is recalculated from the flue gas composition. Sulfur has been neglected for these calculations. Data are summarized in Table 8. All values of the oxygen demand (XOD) are recalculated according to Eq. (42).

H [m]

XT ¼

ð0:5  n_ H2 þ 0:5  n_ CO þ 2  n_ CH4 þ n_ C fix ÞFR;out

2

1

0 0

5

10

15

20

25

Δp [kPa] Experimental results

CLC process model

Fig. 4. Pressure profile from the fuel reactor at operational point VII-5 and the corresponding predictions of the CLC process model elaborated in this work.

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100

y(CO2) [%]

80

60

40

Experimental results

20

1.5 D fuel reactor model CLC process model X-16 X-17 X-18 X-19 X-20

X-14

X-5 X-6 X-7 X-8 X-9

X-1 X-2 X-3

IX-1 IX-2 IX-3 IX-4 IX-5 IX-6 IX-7 IX-8 IX-9 IX-10

VII-3 VII-4 VII-5 VII-6 VII-7 VII-8 VII-9

VI-1 VI-2 VI-3

0

Operational point Fig. 5. Dry concentrations of CO2 at the fuel reactor outlet obtained for different operational points and modeled in this work.

10

y(H2) [%]

8

6

4

Experimental results

2

1.5 D fuel reactor model CLC process model X-16 X-17 X-18 X-19 X-20

X-14

X-5 X-6 X-7 X-8 X-9

X-1 X-2 X-3

IX-1 IX-2 IX-3 IX-4 IX-5 IX-6 IX-7 IX-8 IX-9 IX-10

VII-3 VII-4 VII-5 VII-6 VII-7 VII-8 VII-9

VI-1 VI-2 VI-3

0

Operational point Fig. 6. Dry concentrations of H2 at the fuel reactor outlet obtained for different operational points and modeled in this work.

14 12

y(CO) [%]

10 8 6 4

Experimental results 1.5 D fuel reactor model

2

CLC process model

X-16 X-17 X-18 X-19 X-20

X-14

X-5 X-6 X-7 X-8 X-9

X-1 X-2 X-3

IX-1 IX-2 IX-3 IX-4 IX-5 IX-6 IX-7 IX-8 IX-9 IX-10

VII-3 VII-4 VII-5 VII-6 VII-7 VII-8 VII-9

VI-1 VI-2 VI-3

0

Operational point Fig. 7. Dry concentrations of CO at the fuel reactor outlet obtained for different operational points and modeled in this work.

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5

y(CH4) [%]

4

3

2

Experimental results

1

1.5 D fuel reactor model CLC process model X-16 X-17 X-18 X-19 X-20

X-14

X-5 X-6 X-7 X-8 X-9

X-1 X-2 X-3

IX-1 IX-2 IX-3 IX-4 IX-5 IX-6 IX-7 IX-8 IX-9 IX-10

VII-3 VII-4 VII-5 VII-6 VII-7 VII-8 VII-9

VI-1 VI-2 VI-3

0

Operational point Fig. 8. Dry concentrations of CH4 at the fuel reactor outlet obtained for different operational points and modeled in this work.

30

25

ΩOD [%]

20

15

10 Experimental results

5

1.5 D fuel reactor model CLC process model X-16 X-17 X-18 X-19 X-20

X-14

X-5 X-6 X-7 X-8 X-9

X-1 X-2 X-3

IX-1 IX-2 IX-3 IX-4 IX-5 IX-6 IX-7 IX-8 IX-9 IX-10

VII-3 VII-4 VII-5 VII-6 VII-7 VII-8 VII-9

VI-1 VI-2 VI-3

0

Operational point Fig. 9. Oxygen demand (XOD) in the fuel reactor obtained for different operational points and modeled in this work.

100

ηOO [%]

95

90

85

Experimental results 1.5 D fuel reactor model CLC process model X-16 X-17 X-18 X-19 X-20

X-14

X-5 X-6 X-7 X-8 X-9

X-1 X-2 X-3

IX-1 IX-2 IX-3 IX-4 IX-5 IX-6 IX-7 IX-8 IX-9 IX-10

VII-3 VII-4 VII-5 VII-6 VII-7 VII-8 VII-9

VI-1 VI-2 VI-3

80

Operational point Fig. 10. Oxide oxygen fraction (gOO) obtained for different operational points and modeled in this work.

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Carbon dioxide is the main product of the CLC process. This is the reason why the dry concentrations are typically higher than the concentrations of unconverted gases. In Fig. 5 almost all values are between 40% and 60%. There are two operational points with either a significantly higher (OP IX-5) or lower (OP X-1) concentration of CO2. This is due to the boundary conditions of the operational points: At OP IX-5 the screw was purged with carbon dioxide and at OP X-1 the fluidization of the carbon stripper was changed from steam to nitrogen. In general the models predictions fit the curve shape of the experimental data. The deviations are between 0.0 and 3.5 mol.% which equals relative deviations below 10 %. One of the unconverted gases which leave the CLC process is hydrogen. The dry concentrations of the 34 different operational points are graphically presented in Fig. 6. The deviations of the calculated concentrations from experimental results are up to 1.9 mol.% (OP VI-3) but for the most of the conditions below 1 mol.%. The decreasing concentration of hydrogen from operational point IX-1 to IX-4 can be explained by the increasing pressure drop (see Table 6) which enhances the conversion of H2, CO and CH4 to CO2 and H2O by the oxygen carrier. Hydrogen is systematically overpredicted for operational points VI1-3 and X-16–X-20. For operational period VI this could be explained by an over-estimation of the mass inventory of char particles leading to a high amount of the gasification product hydrogen. Just before operational points X-16–X-20, new oxygen carrier was refilled into the pilot. This could be less reactive than activated oxygen carrier leading to lower conversion of e.g. hydrogen. Another reason could be the amount on nitrogen which was set to a fixed value for all simulations (5.2 kg/h) but is reported to be in the range between 4.5 and 6.8 kg [5]. A higher feed of nitrogen would dilute other gases like steam or hydrogen and would influence the simulation results. The concentrations of carbon monoxide calculated from the CLC process model are presented in Fig. 7. The models predictions follow the experimental results from the 100 kWth CLC unit. There is a deviating tendency for only 2 operational points, namely OP VII-7 and OP X-20. The underestimation of the CO-concentration at OP VII-7 was also found using a 1.5 D model [21]. The rather low (1600 kg/h, OP X-20) or high (3300 kg/h, OP VII-7) circulating mass flow of oxygen carrier could lead to slight deviations of the models predictions from experimental results. Furthermore there is an offset of the CO-concentrations after the refill of fresh ilmenite (cf. OPX-5, X-16 – X-20). The predicted concentrations of carbon monoxide are higher than the measured

values. This could be explained by fresh oxygen carrier material which had not yet been activated, leading to a lower conversion of unburned gases. In Fig. 8 the models predictions of the concentrations of methane are compared to experimental findings. The models predictions fit the experimental concentrations accurately. The deviations of all operational points are lower than 0.6 mol.% (OP X-8). The rather low concentration does not indicate a high conversion of volatile matter. It is caused by the low amount of methane which is produced during pyrolysis (cf. Table 4). For all concentrations of unburned gases there is an opposite trend of the experimental results and the models predictions between operational point OP X-19 and OP X-20. Between these points the circulating mass off oxygen carrier was decreased during likely constant conditions. This leads to higher conversions of carbon monoxide, hydrogen and methane. From the model a higher conversion of these gases is predicted. Deviations could be explained by a short period of operation and some erroneous estimation of the solids circulation rate. 3.3. Oxygen demand and oxide oxygen fraction The oxygen demand is one of the key parameters of the chemical looping process since it describes the oxygen needed for the total oxidation of unburned gases. The models predictions of the oxygen demand as well as the experimental results and the predictions of the 1.5 D fuel reactor model are presented in Fig. 9. Due to the fact that the calculated dry concentrations from the CLC process model are generally in good agreement with the experimental findings, the models predictions of the oxygen demand correspond to the recalculated values of the 100 kWth tests. The calculated oxygen demand of OP IX-5 is lower in comparison to the other operational conditions. This can be explained by replacing the purge nitrogen by carbon dioxide. A further comparison variable of the chemical looping process is the oxide oxygen fraction. The experimental results and the CLC process models predictions are compared in Fig. 10. For the CLC tests an average oxide oxygen fraction of 97.9% can be calculated. From the CLC process model the same value is obtained. When the carbon stripper is fluidized with nitrogen instead of steam (OPX-1), the experimentally found oxide oxygen fraction is significantly lower. This was explained by the additional gasification inside the carbon stripper. Since an additional gasification in the carbon stripper is not considered in the model the oxide oxygen

Average relative deviation [%]

14 12

11.4

10 8 7.0

7.5

6 4.7

4 2.7

2 0.5

0 Carbon dioxide

Hyrogen

Carbon monoxide

Methane

Oxygen demand Oxide oxygen fraction

Fig. 11. Comparison of the fuel reactor model: average relative deviations of the models predictions from experimental findings.

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fraction calculated by the CLC process model is slightly overestimated (97.4 instead of 95.5%). There is a low variation of the predicted oxide oxygen fraction (97.3–98.5%). In reality, this efficiency seems to be slightly more sensitive on the operational conditions (96.7–99.0% with the exception of OPX-1). However, the same value of separation was used for all operational points since there is no extensive study concerning the separation efficiency of the carbon stripper and no experimental data are available.

operational points with different operational conditions (see also Table 6). 4. Sensitivity analysis After the validation of the CLC process model, it is possible to investigate the influence of different parameters on the 100 kWth pilot. The total oxygen demand (XT) and the carbon capture efficiency (gCC) are selected for evaluation purpose. The operational point IX-4 was chosen as a basic point because the calculations are in very good agreement with the experimental findings for this operational point. In the following section, sensitivity analyses regarding the influence of different boundary conditions are carried out. Furthermore, the efficiency of the fuel reactor cyclone is varied to investigate the influence on the loss of char particles at the fuel reactor cyclone and the total oxygen demand. It is possible to investigate the influence of the thermal power on the process performance since the model was validated with different fuel feeds. The results of this sensitivity analysis are shown in Fig. 12. For constant boundary conditions the total oxygen demand increases with increasing thermal power while the

3.4. Evaluation of the CLC process model All relevant dry concentrations of the fuel reactor outlet as well as the oxygen demand and the oxide oxygen fraction of the CLC process model were compared so far. To further evaluate the validity of the model and to get a more precise overview, the average relative deviations of the calculated values from the experimental findings are calculated and graphically shown in Fig. 11. The relative deviations of all considered concentrations and characteristic quantities are below 8% with the exception of hydrogen. It should be pointed out that these deviations are calculated for all

Thermal power [kW] 0

20

40

60

80

100

120

100

40

99

30

98

20

97

10

96

0

0

4

8

12

16

ηCC [%]

ΩT [%]

50

95

Coal feed [kg/h]

50

100

40

99

30

98

20

97

10

96

0

150

170

190

210

230

250

ηCC [%]

ΩT [%]

Fig. 12. Influence of the coal feed (or thermal power) on the total oxygen demand (XT) and the carbon capture efficiency (gCC).

95

Δp [mbar] Fig. 13. Influence of the pressure drop in the fuel reactor on the total oxygen demand (XT) and the carbon capture efficiency (gCC).

P. Ohlemüller et al. / Applied Energy 157 (2015) 433–448

carbon capture efficiency decreases slightly. This can be explained by the loss of char particles at the fuel reactor cyclone which depends directly on the fuel feed. Furthermore, the oxygen carrier to fuel ration becomes smaller with increasing power resulting in lower conversion of unburned gases. The pressure drop, the temperature, the circulating mass of ilmenite, the steam flow to the fuel reactor and the efficiency of the fuel reactor cyclone are selected to optimize the operational conditions of the 100 kWth unit. These parameters are varied in the range of the given values (see Table 6). The total oxygen demand and the carbon capture efficiency were chosen for comparison because these parameters are very important quantities of the CLC process.

ΩT [%]

 By increasing the pressure drop (see Fig. 13) of the fuel reactor from 170 to 230 mbar, the total oxygen demand decreases from 46.3 to 42.4%. Synchronous with the decrease of the oxygen demand the carbon capture efficiency increases. This can be explained by higher conversion of unconverted gases due to a higher inventory of material (oxygen carrier). This equals a higher oxygen carrier to fuel ratio leading to a better contact between oxygen carrier particles and unburned gases.

 The variation of the temperature between 957 and 977 °C (see Fig. 14) leads to total oxygen demands between 40.8% (977 °C) and 46.5% (957 °C) while the carbon separation efficiency increases with temperature. Higher temperatures enhance the reaction rates for gasification and conversion of unburned gases.  By increasing the mass flow of ilmenite (see Fig. 15) from 1800 to 3300 kg/h the oxygen demand increases from 41.5% to 47.7% and the carbon capture efficiency decreases from 98.5 to 98.0%. In the model, higher mass flows lead to higher losses of material at the fuel reactor cyclone due to higher total circulation. This results in a higher total oxygen demand. Furthermore, higher mass flows lead to a shorter residence time inside the air reactor. However, the average conversion of the oxygen carrier in the fuel reactor is lower caused by a higher circulating mass flow. This leads to a shorter time needed to fully oxidize the oxygen carrier in the air reactor.  Higher steam flow lead to slightly lower total oxygen demands (range: 44.2–43.4%) and to a slightly higher carbon capture efficiency (98.3–98.4%, see Fig. 16). This can be explained by a higher concentration of steam leading to higher gasification rates and to higher conversion of the fuel.

50

100

40

99

30

98

20

97

10

96

0

950

955

960

965

970

975

980

985

ηCC [%]

444

95

T [°C]

50

100

40

99

30

98

20

97

10

96

0 1500

2000

2500

3000

ηCC [%]

ΩT [%]

Fig. 14. Influence of the fuel reactor temperature on the total oxygen demand (XT) and the carbon capture efficiency (gCC).

95 3500

Circulating mass of OC [kg/h] Fig. 15. Influence of circulating mass flow of oxygen carrier on the total oxygen demand (XT) and the carbon capture efficiency (gCC).

445

50

100

40

99

30

98

20

97

10

96

0

25

27

29

31

33

35

37

39

ηCC [%]

ΩT [%]

P. Ohlemüller et al. / Applied Energy 157 (2015) 433–448

95

Steam flow [kg/h] Fig. 16. Influence of the steam flow to the fuel reactor on the total oxygen demand (XT) and the carbon capture efficiency (gCC).

100

60 95

50 40

90 30 20

ηCC [%]

ΩT [%], loss of fixed caron [wt.-%]

70

85

10 0

91

93

95

97

99

80

Efficiency of the fuel reactor cyclone [%] Fig. 17. Influence of the fuel reactor cyclone on the loss of fixed carbon, the total oxygen demand (XT) and the carbon capture efficiency (gCC).

Table 5 Modeling results of flue gas concentrations, the oxygen demand, the total oxygen demand, the oxide oxygen fraction and the carbon capture efficiency for an optimum operational point. Value

CO2

Unit

mol.%

Optimum operational point

57.30

H2

CO

CH4

XOD

gOO

XT

gCC

98.31

19.69

98.18

% 8.55

11.53

 The fuel reactor cyclone has the greatest effect on the total oxygen demand and the carbon capture (see Fig. 17). The total oxygen demand can be reduced from 53.6% (92% cyclone efficiency) to only 19.23% (100% cyclone efficiency). At the same time, the carbon capture efficiency decreases from 98.8% to 96.5%. These trends can be explained by the loss of fixed carbon at the fuel reactor cyclone. A higher efficiency of the cyclone leads to a lower loss of char particles and a lower total oxygen demand (cf. Eq. (44)). The lower carbon capture can be explained by higher amount of char entering the air reactor due to accumulation in the fuel reactor system.

2.40

17.70

Based on these investigations it is possible to predict an optimum operational point for the 100 kWth test rig. The highest pressure drop (230 mbar), the highest temperature (977 °C), the lowest circulating mass flow (1800 kg/h) as well as the highest steam flow (37.5 kg/h) from the sensitivity analysis are used. Furthermore, the separation efficiency of the fuel reactor cyclone is set to 100%. The results of this optimum operational point are summarized in Table 5. For optimized operational conditions, the total oxygen demand can be reduced below 20%. More than 98.2% of the carbon could be captured under these conditions. It can be reasonably assumed that the total oxygen demand could be significantly

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P. Ohlemüller et al. / Applied Energy 157 (2015) 433–448

Table 6 Operational conditions of the 100 kWth CLC unit [5]. Operational point

N2 (kg/h)

CO2 (kg/h)

Steam (kg/h)

_ OC (kg/h) m

_ Coal (kg/h) m

T (°C)

Dp (bar)

Air (kg/h)

VI-1 VI-2 VI-3 VII-3 VII-4 VII-5 VII-6 VII-7 VII-8 VII-9 IX-1 IX-2 IX-3 IX-4 IX-5 IX-6 IX-7 IX-8 IX-9 IX-10 X-1 X-2 X-3 X-5 X-6 X-7 X-8 X-9 X-14 X-16 X-17 X-18 X-19 X-20

5.2 5.2 5.2 5.2 5.2 5.2 5.2 5.2 5.2 5.2 5.2 5.2 5.2 5.2 0 5.2 5.2 5.2 5.2 5.2 30.3 5.2 5.2 5.2 5.2 5.2 5.2 5.2 5.2 5.2 5.2 5.2 5.2 5.2

0 0 0 0 0 0 0 0 0 0 0 0 0 0 5.2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0

31.5 37.5 37.5 37.5 37.5 27.5 27.5 32.5 27.5 27.5 27.5 27.5 27.5 27.5 27.5 27.5 28.5 29.5 30.5 32.5 13.5 27.5 27.5 27.5 27.5 32.5 32.5 37.5 27.5 27.5 27.5 37.5 27.5 26.5

1000 1000 1000 2160 2700 2520 2520 3300 2220 2100 1500 1800 2400 2400 2400 2400 2520 2520 2800 2900 1750 1700 2100 1450 2400 3000 2350 2350 1850 1800 1700 2100 2000 1600

6.3 6.3 6.3 12.6 12.6 12.6 12.6 12.6 12.6 12.6 12.6 12.6 12.6 12.6 12.6 12.6 12.6 12.6 12.6 12.6 12.6 12.6 12.6 12.6 12.6 12.6 12.6 12.6 12.6 12.6 12.6 12.6 12.6 12.6

942.5 942.5 942.5 956.0 957.0 959.0 967.0 974.0 970.0 962.0 960.0 960.0 960.0 965.0 965.0 965.0 965.0 965.0 965.0 965.0 955.0 955.0 955.0 955.0 955.0 955.0 955.0 955.0 955.0 955.0 955.0 955.0 955.0 955.0

0.13 0.13 0.13 0.23 0.23 0.21 0.2 0.17 0.19 0.19 0.17 0.19 0.2 0.2 0.2 0.2 0.18 0.18 0.17 0.15 0.18 0.18 0.19 0.18 0.20 0.17 0.15 0.15 0.18 0.20 0.19 0.21 0.19 0.20

143.0 143.0 157.3 178.8 178.8 178.8 178.8 178.8 178.8 178.8 143.0 160.0 178.8 178.8 178.8 178.8 178.8 178.8 178.8 178.8 161.0 161.0 178.8 143.0 178.8 178.8 161.0 161.0 178.8 161.0 161.0 178.8 178.8 178.8

Table 7 Experimental results of flue gas concentrations [5]. The oxygen demand was calculated according to Eq. (42). Operational point

CO2 (mol.%)

H2 (mol.%)

CO (mol.%)

CH4 (mol.%)

XOD (%)

gOO (%)

VI-1 VI-2 VI-3 VII-3 VII-4 VII-5 VII-6 VII-7 VII-8 VII-9 IX-1 IX-2 IX-3 IX-4 IX-5 IX-6 IX-7 IX-8 IX-9 IX-10 X-1 X-2 X-3 X-5 X-6 X-7 X-8 X-9 X-14 X-16 X-17 X-18 X-19 X-20

41.5 42.5 43.0 52.0 53.5 55.0 56.0 53.0 54.0 53.0 49.5 51.5 54.0 55.0 82.0 55.0 55.0 54.0 53.0 52.0 19.4 49.0 51.0 49.8 53.2 53.0 49.5 48.4 51.4 52.7 49.3 49.7 50.4 52.3

6.5 6.5 6.0 6.5 5.0 5.0 5.5 6.5 6.0 5.5 8.5 7.5 6.5 6.0 6.0 5.5 6.0 6.0 6.2 6.5 2.4 7.0 6.7 6.9 5.8 6.3 7.4 7.5 6.6 5.9 6.2 5.8 6.0 5.5

8.5 8.5 8.5 9.0 8.0 7.5 7.5 8.5 8.0 8.5 11.0 10.0 8.5 8.0 8.5 7.5 7.5 7.5 7.6 8.0 3.9 8.5 8.5 8.6 7.7 8.1 9.1 9.2 8.5 7.7 8.6 8.3 8.0 7.4

3.3 3.5 3.5 3.5 3.5 3.0 3.0 3.0 3.0 3.0 3.5 3.0 3.0 3.0 3.0 3.0 3.0 3.0 2.9 3.0 1.5 3.3 3.4 3.4 3.1 3.1 3.2 3.5 3.7 3.2 3.7 3.8 3.5 3.2

22.3 22.6 22.0 19.4 17.6 15.9 15.9 17.8 17.0 17.1 22.2 19.4 17.5 16.7 12.0 16.2 16.5 16.8 17.0 17.8 21.2 20.0 19.3 20.0 17.1 17.7 20.2 21.2 19.9 17.7 20.5 20.0 19.1 17.2

98.5 98.0 97.0 97.8 98.0 98.0 98.5 97.0 99.0 98.0 99.0 98.5 98.0 98.0 97.0 98.0 98.0 97.5 97.7 97.3 95.5 98.5 98.6 98.3 97.8 97.1 96.9 96.7 98.3 98.2 98.3 98.7 98.7 98.8

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Table 8 Modeling results of flue gas concentrations from a 1.5 D fuel reactor model [21]. The oxygen demand was calculated according to Eq. (42). Operational point

CO2 (mol.%)

H2 (mol.%)

CO (mol.%)

CH4 (mol.%)

XOD (%)

gOO (%)

VII-3 VII-4 VII-5 VII-6 VII-7 VII-8 VII-9

53 55 55 57 56 55 54

5.8 5.4 5.8 6.0 6.0 6.3 6.8

8.5 8.0 7.4 7.5 7.0 8.2 9.8

4.0 3.8 3.7 3.2 2.8 3.3 3.6

19.6 18.2 18.0 16.5 15.6 17.7 19.5

97 98 98 97 97 98 98

improved in a large-scale unit: The residence time of gases would be longer leading to higher conversions of unburned gases. Thus, a total oxygen demand significantly below 20% can be expected. 5. Conclusions A process simulation model was developed and validated with experimental results from a 100 kWth CLC unit at Chalmers University of Technology. A Colombian coal (El Cerrejón) and ilmenite (oxygen carrier) were used for the investigation of 34 different operational points. The predictions of the model are in very good agreement with the experimental results. Simulations were performed using different mass flows of coal, different pressure drops of the fuel reactor, different temperatures of the fuel reactor and various other changes. The average relative deviation of the concentrations of CO2, H2, CO and CH4 as well as the oxygen demand and the oxide oxygen fraction are below 12% for all values. The validated process model was used to investigate the influence of the pressure drop, the temperature, the circulating mass of ilmenite, the steam flow to the fuel reactor, the thermal power and the efficiency of the fuel reactor cyclone on the total oxygen demand and the carbon capture efficiency of the 100 kWth test rig. An optimum operational point with the relevant operational conditions was elaborated and the flue gas concentrations as well as the key parameters of the chemical looping process were predicted for these conditions. It might be reasonably assumed that the total oxygen demand of the 100 kWth CLC unit can be reduced to 19.7%. Some parts like the air reactor, the carbon stripper or the post-oxidation chamber are not modeled in detail and could to be included in a more precise process model. However, the elaborated CLC process model could be used as a basis for further investigations and scale-up studies. Acknowledgement This work was supported by the European Commission within the RFCS project ACCLAIM (RFCP-CT-2012-00006). References [1] Figueroa JD, Foudt T, Plasynski SP, McIlvried H, Srivastava RD. Advantages in CO2 capture technology – the U.S. Department of Energy´s Carbon Sequestration Program. International Journal of Greenhouse Gas Control 2008;2(1):9–20. [2] Buhre BJP, Elliott LK, Sheng CD, Gupta RP, Wall TF. Oxy-fuel combustion technology for coal-fired power generation. Progress in Energy and Combustion Science 2005;31:283–307. [3] Berguerand N, Lyngfelt A. Design and operation of a 10 kWth chemical-looping combustor for solid fuels – testing with South African coal. Fuel 2008;87(12):2713–26. [4] Cuadrat A, Abad A, García-Labiano F, Gayán P, de Diego LF, Adanez J. The use of ilmenite as oxygen-carrier in a 500 Wth chemical looping coal combustion unit. Int J Greenhouse Control 2011;5(6):1630–42. [5] Markström P, Linderholm C, Lyngfelt A. Chemical-looping combustion of solid fuels – design and operation of a 100 kW unit with bituminous coal. International Journal of Greenhouse Gas Control 2013;15:150–62. [6] Leion H, Mattisson T, Lyngfelt A. The use of petroleum coke as fuel in chemicallooping combustion. Fuel 2007;86:1947–58.

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