Carbonate looping experiments in a 1 MWth pilot plant and model validation

Fuel 127 (2014) 13–22

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Carbonate looping experiments in a 1 MWth pilot plant and model validation Jochen Ströhle ⇑, Markus Junk, Johannes Kremer, Alexander Galloy, Bernd Epple Energy Systems and Technology, Technische Universität Darmstadt, Otto-Berndt-Str. 2, 64287 Darmstadt, Germany

h i g h l i g h t s  1 MWth carbonate looping pilot plant has been operated for >400 h with CO2 capture.  Total CO2 capture rates above 90% were achieved.  Process simulations with 1D carbonator model show good agreement with experiments.

a r t i c l e

i n f o

Article history: Received 3 July 2013 Received in revised form 12 November 2013 Accepted 15 December 2013 Available online 10 January 2014 Keywords: CO2 capture Calcium looping Dual fluidized bed Pilot plant Process model

a b s t r a c t Carbonate looping is an efficient post-combustion CO2 capture technology using limestone based sorbents. A carbonate looping pilot plant consisting of two interconnected circulating fluidized bed (CFB) reactors with a thermal capacity of 1 MWth has been designed and erected at Technische Universität Darmstadt. The pilot plant has been operated for >1500 h in fluidized bed mode, thereof >400 h with CO2 capture. The heat for the endothermic regeneration of CaO in the calciner was provided by combustion of either propane or pulverised coal with O2 enriched air. High CO2 absorption efficiencies of up to 85% in the carbonator were achieved for long periods. Taking the CO2 produced by oxyfuel-combustion in the calciner into account, the pilot plant was operated with total CO2 capture rates above 90%. A process model for the carbonate looping pilot plant has been developed with ASPEN PLUS™. A 1D CFB model has been implemented in the process model to determine the effect of hydrodynamics within a fast fluidized bed on the CO2 absorption rate in the carbonator. Operating conditions of a selected test campaign where used as boundary conditions. The results of process simulations show good agreement of calculated CO2 absorption rate with experimental data. Hence, this process model can be considered as a reliable tool for scale-up of the process. Ó 2013 Elsevier Ltd. All rights reserved.

1. Introduction CO2 capture and its subsequent storage or reuse has emerged as one of the possible solutions to reduce CO2 emissions from fossil fuels. Intensive research has been carried out to meet the technological and economic challenges of capturing CO2 from fossil-fuelbased electricity production. First generation CO2 capture processes (MEA scrubbing, Oxyfuel combustion, IGCC) have the disadvantage of substantial net efficiency losses of 8–14% points (incl. compression) and high CO2 avoidance costs [1–4]. Hence, second generation CO2 capture processes with improved efficiency are currently being investigated. The carbonate looping process (CL) is a promising post-combustion CO2 capture technology using limestone as sorbent, suitable for retrofitting existing plants [5–7]. CO2 in a flue gas is absorbed by CaO in the carbonator and transferred as CaCO3 to the calciner, ⇑ Corresponding author. Tel.: +49 6151 16 4791. E-mail address: [email protected] (J. Ströhle). 0016-2361/$ - see front matter Ó 2013 Elsevier Ltd. All rights reserved. http://dx.doi.org/10.1016/j.fuel.2013.12.043

where the CO2 is released by a temperature increase, forming a gas stream of highly concentrated CO2. The regenerated CaO is transferred back to the carbonator. The most straightforward heat source for the endothermic calcination reaction is a supplementary firing with coal and oxygen. The use of technical oxygen at an appropriate oxygen excess results in small amounts of O2, N2, and Ar in the product stream, which requires a purification step. A system of two interconnected circulating fluidized beds (CFB) has been proposed for CL, since good mixing between solids and gas, high solids circulation rates, and high heat transfer rates leading to uniform temperature are required [8]. CL is associated with rather low efficiency penalties of 5–7% points (incl. CO2 compression and purification) and additional power [9–11]. The low price and high availability of limestone in various locations on earth is a further advantage of CL. These features lead to relatively low CO2 avoidance costs below 20 €/tCO2 [12]. The efficiency of the CL process can be further improved by using an indirectly heated calciner [9,13].

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Nomenclature a a0 C CO2 ;d C CO2 ;ex C CO2 ;in C CO2 ;eq fm, fw F0 FR Hd Hl Ht Kbe Kff fa Kr Kg Kri R0 ks Lmf S0 u0 umf uT

decay constant for solid fraction in the lean region (m1) decay constant of clusters in the freeboard (m1) CO2 concentration at the top of the dense phase (mol/ m3) CO2 concentration at the reactor exit (mol/m3) CO2 concentration at the reactor inlet (mol/m3) CO2 equilibrium concentration (mol/m3) limestone specific constants (–) make-up flow (kmol/s) sorbent flow between the reactors (kmol/s) height of the dense region (m) height of the lean region (m) total height of the reactor (m) overall gas interchange coefficient between bubble and emulsion phases (s1) overall rate for fast fluidization (s1) fraction of active CaO at the reactor exit (–) overall carbonation rate of particle in the emulsion phase (s1) mass transfer coefficient of CO2 toward the particles in the emulsion phase (m/s) carbonation reaction rate (s1) carbonation reaction rate (s1) rate constant for the carbonation reaction at the surface of CaO (m4/s mol) height of bed at minimum fluidization (m) initial surface area of CaO per unit mass of solid CaO (m2/g) superficial gas velocity (m/s) minimum fluidization velocity (m/s) terminal velocity (m/s)

The feasibility of the CL process has been confirmed through operation of various test rigs at different scales. Experiments in a continuously operated 30 kWth test rig at CSIC (Spain) showed CO2 capture rates up to 85% that decreased with time due to deactivation [14]. Steady-state CO2 capture efficiencies greater than 90% were achieved in a continuous 10 kWth test rig at Stuttgart University (Germany) for different combinations of operational parameters [15]. Results of a 10 kWth CL test rig at CANMET (Canada) operating in discontinuous mode showed that the moisture content has a very positive effect on carbonation [16]. This effect was confirmed by tests in a continuously operated 200 kWth test rig at Stuttgart University [17]. In a 1 MWth pilot plant erected at Technische Universität Darmstadt, batch tests with alternating carbonation and calcination proved that the CO2 concentration at the carbonator exit could be reduced to equilibrium conditions [18]. Using two interconnected CFB reactors in the same plant, continuous CL tests were performed in 1 MWth scale for the first time worldwide [19]. Firing the calciner with hard coal and O2 enriched air, up to 88% of CO2 from the flue gas was captured in the carbonator corresponding to an overall CO2 capture efficiency of up to 95% [20,21]. Recently, a 1.7 MWth test plant has been erected and operated at La Pereda (Spain). Up-scaling of CL technology to industrial scale confronts challenges, since there is no previous experience, and large scale experimental investigation is expensive. The properties of fluidized beds play an important role in the performance of the CL process due to the high reaction rates required, and the high enthalpy of the carbonation reaction. Some empirical 1D carbonator models have recently been developed [22–24]. Detailed flow analyses can be obtained through Computational Fluid Dynamics (CFD) modelling. Few published works have implemented the carbonation and

XCarb X CO2 ;total X Xave E Rt T PCO2 PCO2eq nloops mInv _M m _S m _ CO2 ;fuel m

CO2 absorption efficiency of the carbonator (–) total CO2 capture efficiency of the process (–) fractional conversion of active CaO (–) average fraction of active CaO (–) activation energy (kJ/mol) gas constant (J/molK) temperature (K) partial pressure of CO2 (Pa) equilibrium partial pressure of CO2 (Pa) age of the total inventory in loops (–) solids inventory (kg) make-up mass flow (kg/h) circulating solids mass flow (kg/h) CO2 production due to fuel combustion in the calciner (kg/h) _ CO2 ;Calc;out CO2 mass flow leaving the calciner (kg/h) m _ CO2 ;M m CO2 production due to make-up flow (kg/h) _ CO2 ;in CO2 introduced into the carbonator from the upstream m plant (kg/h) Greek letters ef void fraction in a fluidized bed (–) emf void fraction in a bed at minimum fluidizing conditions (–) esd volume fraction of solids in the dense region (–) ese volume fraction of solids at the reactor exit (–) es saturated carrying capacity of a gas (–) gbed contact efficiency (–) ccore fraction of solids in the core region (–) cwall fraction of solids in the wall region (–)

calcination reactions of the dual fluidized bed system in a CFD model [25]. The present paper summarizes the main results from operation of the 1 MWth CL pilot plant at Technische Universität Darmstadt achieved so far. Furthermore, process simulations of a selected 1 MWth test campaign using a 1D CFB model for the carbonator are presented, and the results are compared with experimental data for model validation. 2. Theory The carbonate looping (CL) process is based on the reversible carbonation–calcination reaction of limestone as described by Eq. (1).

CaOðsÞ þ CO2ðgÞ ¢ CaCO3ðsÞ þ 178:2 kJ=mol

ð1Þ

The carbonation (forward) reaction is exothermic, whereas the calcination (backward) reaction is an endothermic reaction. The chemical equilibrium of Eq. (1) is shifted towards the right hand side at high temperatures and high CO2 partial pressure. The equilibrium partial pressure of CO2 as function of temperature, C CO2 ;eq , can be approximated by Eq. (2) [26].

C CO2 ;eq ¼ 1:462

  1011 19130 exp T T

ð2Þ

The principle of carbonate looping follows the scheme of Fig. 1. In the so-called carbonator, CO2 contained in a flue gas from a fossil-fired plant is absorbed by CaO forming CaCO3. An ideal carbonation temperature regarding kinetics and the chemical equilibrium for the carbonation reaction is approximately 650 °C [27]. Lower

J. Ströhle et al. / Fuel 127 (2014) 13–22

temperatures lead to a more advantageous chemical equilibrium but a slower reaction rate and vice versa. At a temperature of 650 °C and a pressure of 1 bar, an equilibrium concentration of 0.158 mole CO2/m3 gas or 1.2 vol%. CO2 in gas, respectively, can be achieved. Hence, at least 1.2 vol% of CO2 will remain in the flue gas of the carbonator. Assuming a typical flue gas composition from a coal-fired power plant with 15 vol% of CO2, the maximum capture efficiency for the carbonator, XCarb, is 92%, calculated by Eq. (3).

X Carb ¼ 1 

C CO2 ;ex C CO2 ;in

ð3Þ

Decarbonized flue gas leaves the system, while CaCO3 is transferred to the calciner. The endothermic calcination reaction occurs at a temperature above 900 °C at a CO2 partial pressure close to 1 bar. The required heat for the calcination reaction and for heating the sorbent from 650 °C to 900 °C is supplied by fuel combustion. In the calciner, fuel is burnt in an oxyfuel atmosphere avoiding dilution of the flue gas with nitrogen, since it is desired to receive a flue gas which solely consists of CO2 and H2O. The flue gas leaving the calciner will be post-treated before an almost pure CO2 product stream can be either used for multiple applications or can be stored. The CaO that has been formed will be transferred to the carbonator to close the loop. The CO2 stream leaving the oxyfuel-fired calciner consists of three fractions: (a) the CO2 absorbed from the flue gas of the upstream plant and released in the calciner, (b) the CO2 produced by the combustion of additional fuel in the calciner, and (c) the CO2 released by fresh limestone entering the calciner. The entire CO2 stream leaving the calciner will be processed and can therefore be considered as captured CO2. Hence, the total CO2 capture efficiency of the CL process, X CO2 ;total , is defined as the CO2 mass flow _ CO2 ;calc;out , in relation to the CO2 entering leaving the calciner, m _ CO2 ;in , the fuel the system via the flue gas of the upstream plant, m _ CO2 ;fuel , and the make-up introduced introduced into the calciner, m _ CO2 ;M , according to Eq. (4). into the calciner, m

X CO2 ;total ¼

_ CO2 ;calc;out m _ CO2 ;in þ m _ CO2 ;fuel þ m _ CO2 ;M m

ð4Þ

A certain amount of sorbent has to be replaced constantly by a make-up flow of fresh limestone since some material will be lost as fly ash due to attrition [28], and the sorbent chemically deactivates over time [29]. The decrease in CO2 uptake with repeated carbonation/calcination cycles can be attributed to sintering, which is primarily caused by the low melting point of CaCO3 and accompanied by large changes in the micro-granular network, gradually shifting the sorbent’s porosity from micro- and mesopores to macropores [30]. Dolomite shows a less pronounced decay in cyclic CO2 capture capacity [31,32], but has lower mechanical strength compared to limestone resulting in higher attrition rates [33]. Several strategies have recently been proposed to reduce the decay in CO2

Fig. 1. Carbonate looping process principle.

15

uptake of limestone, e.g. regeneration by hydration [34]. The synthesis of calcium based materials stabilized by high-melting point supports has gained momentum recently [35–37]. These recent develop ments have only been obtained in bench scale, e.g. by thermogravimetric analysis. However, thorough investigation in dual fluidized bed CL test rigs is required before moving to pilot and demonstration scale. Hence, the first approach for scaling-up the process to larger plants will be the use of a natural sorbent, i.e. limestone. SO2 entering with the flue gas and via coal entering the calcination reactor also affects the carbonation reaction due to the formation of gypsum according to Eq. (5):

CaOðsÞ þ SO2ðgÞ þ 0:5O2ðgÞ ! CaSO4ðsÞ

ð5Þ

The reaction with SO2 consumes CaO since the produced gypsum cannot be regenerated below 1000 °C. The formation of a gypsum layer on CaO grains reduces the CO2 capture capability of lime at a high number of cycles, although the SO2 concentration is by orders of magnitude lower than the CO2 concentration in the flue gas [32]. Regeneration of CaSO4 to CaO above 1000 °C is not advantageous due to increased sorbent deactivation. Particle reactivation at lower temperatures is possible, but associated with process inefficiencies. On the other hand, carbonation/calcination cycles lead to an increase of macropores, enhancing the diffusivity with respect to sulfation, so that SO2 capture increases. 3. Experimental A CL pilot plant with a nominal power of 1 MWth consisting of two interconnected CFB reactors has been erected at Technische Universität Darmstadt [38]. The reactor system is fully refractory lined to minimize thermal losses allowing autothermal operation. Fig. 2 shows the setup of the pilot plant. A coal-fired combustor has also been installed for production of a real flue gas containing CO2. However, this combustion chamber was not in operation during the test campaign presented in this paper. Instead, a mixture consisting of air from a fan and CO2 from a tank was used as synthetic flue gas. The synthetic flue gas can be electrically pre-heated up to 350 °C and enters the carbonator through a nozzle grid. The carbonator is equipped with adjustable, internal cooling tubes in order to extract the heat of the exothermic carbonation reaction and to adjust the temperature level of the reactor. Make-up limestone can be introduced into the carbonator by a gravimetric dosing system. The advantage of introducing make-up into the carbonater is that the make-up is heated to 650 °C in the carbonator before it enters the calciner, so that less coal and O2 are required in the calciner. A decarbonized flue gas leaves the system via a heat exchanger, a fabric filter for dust precipitation, and an induced draft (ID) fan. Solids from the carbonator are transferred to the calciner by a screw conveyor attached to the bottom of the loop seal. The solids transfer between the reactors can be adjusted by the rotation speed of the screw conveyor. The calciner is fluidized by a mixture of air from a fan and oxygen from a tank, which can be electrically preheated up to 450 °C. The addition of oxygen allows enhancing the oxygen content in the reactor while maintaining reasonable fluidization velocities. Two different fuels can be burnt in the calciner in order to provide the heat for sorbent regeneration. The reactor can be fired with propane, either by a burner or a bed lance. The gas burner is used for heating up the plant. The bed lance allows additional introduction of propane at the bottom of the reactor. Furthermore, the reactor can be fired with pulverized coal. A maximum flow of 150 kg/h of coal, corresponding to a thermal power of approximately 1 MWth depending on the lower heating value of coal, can be introduced to the calciner by a gravimetric dosing system. Analogue to

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Fig. 2. Experimental setup of the test facility.

Table 1 Dimensions and instrumentation of the carbonator and calciner.

Inner diameter of riser (m) Outer diameter of riser (m) Height of riser (m) Gas analysis Temperature measurement (reactor height) (m) Pressure measurement (reactor height) (m)

Carbonator

Calciner

0.59 1.3 8.66 CO, CO2, O2, SO2, NO 0.25; 1.14; 5.1; 7.23 0.14; 0.61; 2.13; 4.76; 8.08

0.4 1.0 11.35 CO, CO2, O2, SO2, NO, H2, CH4, H2O 0.28; 1.78; 7.04; 9.73 0.15; 0.85; 2.15; 4.59; 5.8; 10.11

the carbonator, the calciner flue gas is released to the environment via a heat exchanger, a fabric filter, and an ID fan. A double loop seal configuration has been installed to control transfer of solids from the calciner to the carbonator. However, the internal circulation of the calciner was closed for operation simplification, so that all solid material entrained to the cyclone of the calciner was directly transferred to the carbonator. The pilot plant is equipped with pressure transducers and thermocouples along the height of the reactors, in the wind boxes, and in the connecting peripherals. The flow and gas composition of the flue gas from the carbonator and calciner are constantly measured. Details on dimensions and instrumentation of the carbonator and calciner are provided in Table 1. 4. Modelling Heat and mass balances of the 1 MWth pilot plant are calculated with a steady state carbonate looping model developed in Aspen Plus™. The CO2 absorption efficiency of the carbonator is calculated by a 1D CFB Fortran code that is executed before the carbonator unit operation in Aspen. The model is based on the fast fluidization relationships by Kunii and Levenspiel [39] and has been previously published [23,40]. For the calculation of the CO2 absorption efficiency, gas and solids streams are imported from the global Aspen model to the Fortran code. As a result the calculated CO2 concentration and the solids entrainment at the exit of the carbonator are exported from the Fortran code back to the

Aspen model. The main parts of the carbonator model are the determination of particle distributions in the riser of the CFB, the calculation of the CO2 absorption rate including kinetics and chemical equilibrium considerations, and the determination of the overall CO2 conversion of the carbonator. The basic equations of the carbonator model are summarized below. For a more detailed description, the reader is referred to [23]. 4.1. Solids distribution The pressure drop in the carbonator during the experiments is used for the calculation of the solids inventory. With the corresponding bed height and solids fraction at minimal fluidization condition, Lmf and emf, and the saturation solids fraction, es , as input variables, the height of the lean region, Hl, and the solids fraction at the exit, ese, are calculated by an iteration loop according to Eqs. (6) and (7) [39], with the solids fraction in the dense region, esd, set constant at 0.24.

Lmf ð1  emf Þ ¼

esd  es a

þ Ht esd  Hl ðesd  es Þ

ese ¼ es þ ðesd  es ÞeaHl

ð6Þ ð7Þ

4.2. CO2 absorption rate Investigations of Sun et al. [41] proved that the driving force of the CO2 absorption reaction, i.e. Eq. (1), changes from first-order to zero-order in dependency of the CO2 partial pressure. A driving force of 10 kPa is taken as a reasonable approximation of the turning point or transition region according to these experimental investigations. The reaction rate for Eq. (1) in [41] is described with Eqs. (8)–(10).

R0 ¼

 dX  ¼ 3  r 0 ¼ 56  ks ðPCO2  PCO2 ;eq Þn  S0 dt t¼0

ks ¼ 1:67  103 exp



E Rt  T

> 10 kPa and n ¼ 0

ð8Þ



   mol for PCO2  PCO2 ;eq m2 s ð9Þ

J. Ströhle et al. / Fuel 127 (2014) 13–22

ks ¼ 1:67  104 exp



E Rt  T





  mol for P CO2  PCO2 ;eq 2 m s

< 10 kPa and n ¼ 1

ð10Þ 2

3

As a specific initial surface area S0 = 40 m /cm and a particle porosity of 0.5 were specified, yielding typical areas of 12 m2/g of active CaO [22,27]. In [41] the initial surface area of the used limestone was measured to 29 m2/g. The equations are actually valid for the surface area of the particles at t = 0. In the present study, the transition of the reaction rate order and the change of the rate constant by a factor of 101 times the driving force was used for the whole kinetically controlled early stage of carbonation. The CO2 equilibrium is calculated according to Eq. (2). 4.3. CO2 conversion The CO2 conversion in the carbonator is calculated subsequently for the dense and lean regions by means of a model for catalytic reactions assuming the reaction rate to be first order [39]. The investigations of [27] and [41] showed slight differences in the values for the reaction rate, whereas the reaction rate determined by Sun et al. [41] was approximately 1.3 times higher than the values from Bhatia and Perlmutter [27]. For this work and driving forces of ðP CO2  PCO2 ;eq Þ < 10 kPa (first-order reaction) the values for the rate in the KL-model were increased. For driving forces of ðPCO2  P CO2 ;eq Þ > 10 kPa the reaction rate constant was adjusted for the KL-model and considered to be zero-order. The CO2 concentration at the top of the dense region, C CO2 ;d , is derived from the general rate constant, Kr, as a function of the height of the dense region, Hd, and the superficial velocity, u0, as stated in Eqs. (11)– (14).

ln

C CO2 ;in dHd ¼ K ff C CO2 ;d u0

K ff ¼ ycore fa K r þ

Kr ¼

ð11Þ

1 K 0be

1 þy

ð12Þ

1

wall fa K r

1 dp 6K g

þ K1ri

X ave S0 qCaO K ri ¼ kS ð1  XÞ2=3 M CaO

ln

gbed ¼ 4ycore þ fa K r K 0be

3

þ y1

wall

5

d ð1  ef Þ

The average value of the maximum sorbent conversion is calculated as a function of circulating solids mass flow, FR, and the make-up flow, F0, according Eq. (17) [9].

X ave ¼

fm ð1  fw ÞF 0 þ fw F 0 þ F R ð1  fm Þ

" # n n  X Y _ S;i _ M;i Dt m m ¼ Dt 1 mInv;i mInv;i i¼0 i¼0

ð14Þ

ð16Þ

Due to the transition of reaction order and the change of reaction rate along the reactor height, it is deemed necessary to determine the CO2 concentration in several sections of the reactor. Therefore, the carbonator is divided in sections of 1 cm height, and the CO2 concentration is calculated in each section. Since the CO2-concentrations at the carbonator inlet during the test

ð17Þ

However, this equation is only valid for steady-state conditions. A change in make-up flow in a real plant affects the process longterm due to the long residence time of the solids in the system. A varied make-up flow of 5%, 10%, and 30% of the total solids inventory per hour takes approximately 44 h, 21 h, and 6 h, respectively, to exchange more that 90% of the inventory, so that the reactivity converges very slowly towards a new design value. Using Eq. (17) directly in the simulations has a strong and sudden effect on the calculated reactivity due to the disparity between transient plant behavior and steady state simulation model. To account for the long-term behavior of the make-up flow, a routine was developed that calculates an equivalent make-up flow better agreeing with the transient behavior of the plant. The equation calculates the age of the total inventory in loops, nloops , depen_ M , the dent on the temporal progressions of the make-up flow, m _ S . The product inventory, mInv , and the solids circulating flow, m operator in Eq. (18) calculates the fraction of the inventory from the period Dt at time step -i that is still in the plant at i = 0. This fraction is multiplied by the amount of loops. The amount of loops is calculated by the quotient of inventory and loops at time step -i multiplied by the time period Dt. The results of each period are added forming the age of plant inventory in loops at i = 0. In the next step of the routine, a constant make-up flow leading to the same age of the inventory at time step i = 0 is iteratively calculated by Eq. (19).

nloops ¼

ð15Þ 1

4.4. Calculation of make-up flow

nloops

   C CO2 ;d ð1  ef Þbed K r 1  gbed  ðaþa0 ÞHf ¼ ð1  eaHf Þ  1  e C CO2 ;ex u0 a 1 þ a0 =a 2

campaigns were between 10 and 12 vol.% the reaction order was usually first-order and therefore the effect of the change in the reaction order was minor.

ð13Þ

The deactivation of CaO is considered by the average value of the maximum sorbent conversion Xave calculated by Eq. (17) and is integrated into the carbonation rate constant of the KL model in Eq. (14). The CO2 concentration at the carbonator exit, C CO2 ;ex , is calculated according to Eq. (15) with the contact efficiency for fluidized beds, gbed , in Eq. (16).

17

" # n n  X Y _ S;0 _ M;constant Dt m m Dt 1 mInv;0 mInv;0 i¼0 i¼0

ð18Þ

ð19Þ

5. Results and discussion The 1 MWth pilot plant at Technische Universität Darmstadt has been operated for more than 1500 h in fluidized bed mode, thereof around 400 h with CO2 capture. This paper focusses on three experimental campaigns performed between July 2011 and February 2012. The calciner was fired with propane in the first two campaigns and with coal in the third campaign. After the first campaign, some improvements of the plant were implemented, i.e. the thermal power of the calciner was increased by installation of a new propane lance, and the separation efficiency of the calciner cyclone was improved by a modification of the loop seals. The first campaign is used for model validation in this paper. 5.1. First test campaign – propane fired calciner In July 2011, continuous CL tests were performed in 1 MWth scale for the first time worldwide. During a period of 72 h, CO2 from a synthetic flue gas (10–12 vol% CO2 in air; volume flow of 1300 Nm3/h corresponding to the flue gas flow from a 1 MWth

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coal-fired furnace) was continuously captured in the carbonator in autothermal operation and released in the calciner. The calciner was fired with propane and O2 enriched air. The use of O2 enriched air instead of an O2/CO2 mixture was due to operational reasons. Of course, this will not be the case in an industrial plant, where O2 will be mixed with recirculated flue gas. The maximum thermal power of the burners and lances summed up to approximately 700 kWth. The initial bed material consisted of limestone with a mean diameter of 430 lm provided by the company Rheinkalk. Prior to the experiments, several tons of limestone had been calcined and tempered for around 1 h at 1000 °C to increase the mechanical strength of the particles. The tempered CaO was used as make-up during the first period of the test campaign. During the second period, fresh limestone was added as make-up. The make-up flow was within 70 and 150 kg/h of CaCO3 equivalent. That means that the mass flow of CaO is calculated with a considered density equal to CaCO3 to provide a better comparability of both experimental phases. However, this high make-up flow was caused by a rather high material loss from the calciner due to a poor performance of the calciner cyclone. The circulation between carbonator and calciner was varied between 1500 and 3000 kg/h. The fluidization air for the calciner was enriched with oxygen to provide a complete combustion of the propane while keeping the riser velocity below an acceptable level of 8 m/s. The carbonator was not heated due to the exothermic carbonation reaction. Results of this test campaign during the entire 72 h period are shown in Fig. 3. The most obvious result is that the CO2 capture efficiencies with tempered CaO as makeup (Phase 1) are lower than with CaCO3 as make-up (Phase 2). It is likely that the tempering of the sorbent prior to the experiment had a negative effect on its reactivity. This is indicated by the fact that the capture efficiency even from the beginning of the experiment is lower than in Phase 2. From the middle of Phase 1, the make-up did not consist of fresh CaO exclusively but contained an increasing fraction of used CaO that was removed from the filter of the calciner. The decrease in capture efficiency over Phase 1 can therefore be explained by enrichment of older sorbent in the bed of the carbonator regarding the decrease of reactivity with number of calcination/carbonation cycles. This behavior changes at the beginning of Phase 2. The enrichment of solids with CaO generated from fresh (not tempered) limestone leads to an increase of the capture efficiency to an even higher level than in the beginning of Phase 1. Evaluation of the collected data revealed that of the varied conditions, the composition of the sorbent had the strongest influence on the capture rate during this experiment. This shows that the properties of the sorbent together with a correct treatment have a strong impact on the capture efficiency. On the contrary, a variation of the makeup flow did not show a significant influence on the capture

Fig. 3. CO2 concentrations at carbonator inlet, C CO2 ;in;Carb , and outlet, C CO2 ;en;Carb , equilibrium CO2 concentration, C CO2 ;eq (all in% vol.), CO2 absorption efficiency of carbonator, Xcarb, and total CO2 capture efficiency, X CO2 ;total , during the first test campaign.

efficiency during this experiment, most likely because the level of make-up flow was relatively high due to the poor performance of the calciner cyclone. The CO2 concentration at the carbonator inlet also affects the capture efficiency. The minimum CO2 concentration at the carbonator outlet is determined by the CO2 equilibrium that primarily depends on the temperature. Hence, a higher CO2 concentration at the carbonator inlet results in a higher maximum capture rate as long as there is a sufficient amount of reactive CaO in the carbonator. The experiments were performed with a lower CO2 concentration at the carbonator inlet than it can be expected from a power plant generated flue gas. The low CO2 inlet concentration was a consequence of the limited firing power in the calciner. Thus, the achieved capture efficiency usually is below the reachable level generally stated for the CL process. Nevertheless, a total efficiency of more than 90% was achieved at the end of the test campaign. 5.2. Second test campaign – propane fired calciner In January 2012, further continuous CL tests were performed in the 1 MWth plant with propane fired calciner. The maximum thermal power of the burners and lances was around 500 kWth. CO2 from a synthetic flue gas was continuously captured during a total period of 131 h. The initial bed material consisted of limestone with a mean diameter of 180 lm provided by the company Rheinkalk. The make-up flow was significantly lower than in the first test campaign due to an improved performance of the calciner cyclone. As explained in Section 3, the internal circulation of the calciner was closed, so that all solid material entrained in calciner was directly transferred to the carbonator. The solids entrainment is directly related to the riser velocity. Hence, the fluidization air had to be enriched with up to 50% oxygen to keep the riser velocity below an acceptable level of 4 m/s while ensuring complete combustion of propane. Nevertheless, this high velocity in the calciner in combination with the small particle size resulted in a rather low bed inventory during this test campaign, leading to a relatively short solids residence time in the calciner. Cooling tubes had to be inserted into the carbonator to keep the temperature level below 660 °C. During a selected 12 h operation period, 160 kg/h of CO2 entered the carbonator leading to a CO2 concentration of 12 vol% at the carbonator inlet. The circulating mass flow between carbonator and calciner was about 2000 kg/h corresponding to an average CaO/CO2 molar ratio of 11.6 with respect to the CaO entering the carbonator from the calciner. Fig. 4 shows the profiles for CO2 flow, temperature, pressure and CO2 capture in the carbonator as well as total CO2 capture for the selected period. During the first 3 h, the CO2 capture efficiency was inversely correlated with temperature, which indicates that CO2 absorption was limited by chemical equilibrium. As the carbonator temperature was decreased from 660 °C to 640 °C after 1 h, the CO2 absorption efficiency in the carbonator increased from 80% to about 85%. Taking the oxyfuel combustion in the calciner into account, a total CO2 capture efficiency of up to 92% was achieved. After 6 h of operation, the CO2 capture efficiency starts to decrease, which cannot be attributed to chemical equilibrium restrictions since the temperature also decreases. This effect was caused by a decrease of thermal power in the calciner due to icing of the propane tank at low ambient temperature. The combustion of propane in the calciner could not provide sufficient heat for complete sorbent regeneration, i.e. the amount of CaO decreased while the amount of CaCO3 in the system increased. The low residence time of solids in the calciner may also have contributed to poor CaO regeneration. The reduction of available lime for CO2 capture in the carbonator led to a lower CO2 capture efficiency.

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Fig. 4. CO2 flow, temperature, pressure and CO2 capture efficiency in the carbonator, as well as total CO2 capture efficiency during the second test campaign with propane fired calciner.

5.3. Third test campaign – coal fired calciner In February 2012, first continuous CL tests were performed in the 1 MWth plant with coal fired calciner. The maximum thermal power of coal was around 750 kWth. CO2 from a synthetic flue gas was continuously captured during a period of 30 h. The same limestone as in the previous test with a mean diameter of 180 lm was used. The make-up flow was close to zero throughout this period. The fluidization air was enriched with up to 45–50% oxygen to limit the riser velocity while ensuring complete combustion of coal. As in the previous campaign, this high velocity in the calciner in combination with the small particle size resulted in a rather low bed inventory during this test campaign. Cooling tubes had to be inserted into the carbonator to keep the temperature level below 670 °C. During the period of CO2 capture, 160 kg/h of CO2 were added to the primary air flow of the carbonator to produce a synthetic flue gas leading to a CO2 concentration of 12 vol% at the carbonator inlet. The circulating solids mass flow between carbonator and calciner was about 2800 kg/h, which corresponds to an average CaO/CO2 molar ratio of 17.2 assuming that the solids flow entering the carbonator completely consisted of CaO. However, the actual CaO mole flow is smaller since the solids also contain gypsum and ash originating from the coal. Unfortunately, chemical analysis of solid samples is not available for this period. Fig. 5 displays the measured profiles for CO2 flow, temperature, pressure and CO2 capture in the carbonator as well as total CO2 capture for a selected 22 h period of operation. No CO2 is introduced in the beginning of this period. After 2 h, CO2 is added to the primary air flow of the carbonator leading to a temperature increase up to 670 °C as the exothermic CO2 absorption starts. After 4 h, heat is removed from the carbonator by inserting the cooling tubes in order to reach the desired temperature level of 650 °C. A CO2 absorption efficiency in the carbonator of 85% corresponding to a total CO2 capture efficiency of 91% is reached at this point. In the period between 2 and 7 h, the CO2 concentration at the carbonator exit was almost equal to the equilibrium CO2 concentration at the actual operating temperature. Hence, the CO2 absorption efficiency was limited by chemical equilibrium, which resulted in rather stable operating conditions. After that, the temperature slightly dropped to around 640 °C, and a fluctuating

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Fig. 5. CO2 flow, temperature, pressure and CO2 capture efficiency in the carbonator, as well as total CO2 capture efficiency during the third test campaign with coal fired calciner.

behavior of temperature and CO2 absorption rate was observed. The CO2 concentration at the carbonator exit was well below the equilibrium CO2 concentration, so that the CO2 absorption efficiency was most likely limited by chemical kinetics, leading to a lower CO2 absorption rate at more unstable operating conditions. After 16 h, the temperature in the carbonator was decreased to 610 °C by further introducing the cooling tubes. As a consequence, the CO2 capture efficiency dropped to about 60% due to a reduced chemical reaction rate.

5.4. Model validation The first test campaign in the 1 MWth pilot plant with propane fired calciner (see Section 5.1) is used to validate the process model described in Section 4. Measurements from the 1 MWth test rig (i.e. temperatures, pressure drops, flows, etc.) are automatically imported from the data acquisition system to the process model providing the boundary conditions for the calculations. Steady-state simulations were performed for every 15 min of the test campaign. Fig. 6 indicates that the calculated solids distribution is in good agreement with time averaged experimental results for an inventory of 200 kg. The calculated solids fraction is slightly higher at the bottom of the lean phase and slightly lower at the exit of the reactor compared to the experimental results. The calculated entrainment of solids in the carbonator is twice as high as the maximum flow rate of the screw conveyor between carbonator and calciner. Fig. 6 also shows profiles of calculated reaction rate and CO2 concentration along the reactor height. There is an inflection point observable for the reaction rate in the dense region that is due to the transition of the reaction order. Additionally, it can be noted that the gradient of the reaction rate and CO2 concentration changes with the transition from the dense to the lean region. The reaction rate in a CFB is expected to be highest in the splash zone, i.e. the region between the bottom dense zone and the freeboard. However, a simplified model is used here, which only considers a dense and a bottom zone. Although the detailed conversion profile might not be captured precisely, the model can serve as an approximation to calculate the overall conversion in the CFB and has been verified by a large number of previous experiments. The simplicity of the model allows an analytic

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Fig. 8. Calculated CO2 absorption efficiency using the real make-up flow compared to experimental data during the first test campaign.

Fig. 6. Calculated solids fraction compared to the averaged experimental results, as well as calculated CO2 absorption rate and CO2 concentration plotted along the height of the reactor for an inventory of 200 kg.

solution of the conversion, so that it can easily be integrated into a process model. Fig. 7 presents the calculated make-up flow using the model for equivalent make-up flow compared to the results with real makeup flow during the test campaign. The peak maxima and minima could obviously be eliminated, and the characteristic is significantly smoothened. Results of simulations using the real and equivalent make-up flow, respectively, are presented in the following sub-sections.

5.4.1. Simulation with real make-up flow For the first set of process simulations, the real make-up flow was used instead of the calculated equivalent make-up flow. Fig. 8 depicts the calculated CO2 absorption efficiency in the carbonator compared to experimental data. The calculated CO2 absorption efficiencies were significantly higher than in the experiments during the first one and a half days, when tempered CaO was used as fresh make-up stream. This discrepancy can be explained by the reduced reactivity of the tempered CaO in the experiments, whereas the model generally uses the reactivity of fresh, highly reactive limestone for the make-up. The simulations agree better with the experimental results during the second one and a half days when fresh CaCO3 was introduced into the pilot plant, which confirms that the disagreement in the first period was due the low reactivity of tempered CaO. The qualitative characteristic in the second period proves good conformity. However, the overall

Fig. 7. Equivalent make-up flow calculated by Eqs. (16) and (17) compared to the real make-up flow to the pilot plant during the first test campaign.

simulation does not converge when the make-up flow becomes zero, i.e. at the end of the test campaign. The total test period is divided into six sections to discuss the results in more detail. In Section 1, the simulation shows qualitatively reasonable results during the period of constant make-up flow. However, the decrease of CO2 absorption due to sharp drop in make-up flow at the end of Section 1 is overpredicted by the simulations. In Section 2, the make-up flow is first increased from 73 kg/h to 90 kg/h and then reduced to 67 kg/h. The calculated CO2 absorption efficiency directly follows this trend. A variation of the make-up stream has a sudden effect on the capture efficiency since the model calculates steady state. The effect cannot be observed during the experiments in the plant, since this change has no short-term impact. The increase of the measured CO2 absorption rate at the end of Section 2 is due to a decrease of the primary air temperature of around 33 K. There is an opposite trend observable for the simulations, since the decrease of make-up outranges the decrease of carbonator temperature. Section 3 shows the reaction of the plant and the simulations on a decrease of the CO2 partial pressure at the inlet of the carbonator from 11 to 9 vol.%. The characteristics show the same trend, whereas the calculated value increases more rapidly. In Section 4, the make-up flow (as fresh limestone) is increased from 60 kg/h to 120 kg/h and the inventory from 180 kg to 200 kg. The corresponding increase in CO2 absorption rate is well captured by the simulations. At the end of Section 4, the make-up is decreased by 30 kg/h leading again to a sudden decrease of the simulated capture efficiency. Afterwards the make-up is increased by 40 kg/h to an overall make-up flow of 140 kg/h, increasing the calculated efficiency to the level of the plant results. In Section 5 the capture efficiencies of the simulation and experiment are mostly identical except of slight deviations depending in variations in make-up and inventory. During this period the plant was operated at very stable conditions. The minimum at the beginning of Section 5 might be due to a sudden decrease of the inventory of around 50 kg in the plant that was adjusted afterwards by an increase of the make-up flow about 40 kg/h. This short-term effects have more impact on the simulation than on the measurements. At the end of Section 5 the make-up flow was increased by 30 kg/h, explaining the increase of the simulated capture efficiency. In Section 6, the make-up flow instantaneously drops to zero, which directly affects the calculated CO2 absorption rate, but not the measured values.

5.4.2. Simulation with equivalent make-up flow Simulations were also carried out with the equivalent make-up as depicted in Fig. 9. No significant improvements of the results could be obtained in Section 1, since the equivalent make-up only

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Fig. 9. Calculated CO2 absorption efficiency using the equivalent make-up flow compared to experimental data during the first test campaign.

decreases slightly during this period, and the influence on the simulation is minor. The slight decrease of CO2 absorption efficiency at the end of Section 1 is caused by a temperature increase in the carbonator. The different gradients in CO2 absorption efficiency indicate that the influence of the temperature in the model is not as high as measured. In Section 2, the simulations with the equivalent make-up stream qualitatively follow the trend of the plant results. A short-term increase of the make-up flow does not affect the simulation anymore. Now, the temperature decrease in the carbonator (increased capture efficiency at the end of Section 2) can be also simulated, but with less intensity than in the experiments. The generally weaker influence of the temperature in the simulations is due to the fact that the temperature profile of the carbonator has not been implemented. The bed temperature is assumed as nominal overall temperature for the total reactor. However, cooling lances (dissipation of exothermic reaction heat) and the startup burner air cooling induce inhomogeneous temperature profiles during the experiments. In Section 3, the simulation results run in parallel to the experimental results. The equivalent make-up converges 60 kg/h during this period. The spread between simulation and experimental results is caused by the low reactivity of tempered make-up used in the experiments compared to fresh CaCO3 in the simulations. The effect of the equivalent make-up can be also identified in Section 4. A short-term peak of the make-up flow now leads only to a smooth increase of the carbonator absorption efficiency. Section 5 indicates the influence of the continuous increase of equivalent make-up. The CO2 absorption efficiency of the simulation now converges more slowly towards the experimental results compared to simulations with the real make-up flow. Section 6 shows reasonable conformity, and it is obvious that a make-up flow of zero has no influence on the convergence of the simulations any more. 6. Conclusion The carbonate looping process has been successfully tested in a 1 MWth pilot plant at Technische Universität Darmstadt. The heat for the endothermic regeneration of the sorbent in the calciner was either provided by combustion of propane or by the combustion of pulverized coal, while CO2 was continuously captured in the carbonator. High CO2 absorption efficiencies of up to 88% in the carbonator were achieved with propane fired calciner. Taking the oxyfuel-fired calciner into account, the pilot plant was operated with total CO2 capture rates up to 92%. CO2 absorption efficiencies of up to 85% in the carbonator corresponding to total CO2 capture rates up to 91% were reached in coal-fired mode. The conditions in the carbonator are stable when the CO2 absorption is limited by the

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chemical equilibrium, whereas a limitation by chemical kinetics can lead to fluctuations of CO2 absorption rate and temperature. The results of the tests campaigns proved the feasibility of the process. However, further testing is required to bring CL technology to larger scale. In the next step, the 1 MWth pilot plant will be operated under conditions closer to industrial application, i.e. the flue gas from a coal fired combustion chamber will be introduced into the carbonator. The performance of the plant can be expected to improve since the moisture content has a positive CO2 absorption efficiency in the carbonator. A flue gas recycle will be installed for the calciner to operate this fluidized bed reactor under real oxyfuel conditions. Furthermore, the influence of the circulating solid flow between the reactors and the influence of the makeup flow of fresh limestone on the CO2 capture efficiency will be examined in more detail. A process model for the carbonate looping pilot plant has been developed using a 1D CFB model to determine the effect of hydrodynamics within a fast fluidized bed on the CO2 absorption rate in the carbonator. The results of process simulations show good agreement of calculated CO2 absorption rate with experimental data. The application of a calculation procedure for an equivalent make-up flow instead of the real make-up flow accounts for the long-term effect of make-up addition, leading to improved qualitative agreement with the experiments. Hence, this process model can be considered as a promising tool for scale-up of the process. However, further validation and optimization is required to prove the reliability of the model. An improvement of the process model by the inclusion and evaluation of an adequate kinetic model for the calciner is subject to further work.

Acknowledgements This work was funded by the German Federal Ministry of Economics and Technology within the ‘‘LISA’’ project (FKZ 0327771C) under the framework of the COORETEC programme. Further funding and assistance was provided by Alstom, E.ON, Fisia Babcock, Grosskraftwerk Mannheim AG, Hitachi Power Europe, Linde, Rheinkalk (Lhoist) and RWE.

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