Investigation into a gas–solid–solid three-phase fluidized-bed carbonator to capture CO2 from combustion flue gas

Chemical Engineering Science 66 (2011) 375–383

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Investigation into a gas–solid–solid three-phase fluidized-bed carbonator to capture CO2 from combustion flue gas Changqing Cao, Kai Zhang, Chenchen He, Yanan Zhao, Qingjie Guo n College of Chemical Engineering, Qingdao University of Science and Technology, Qingdao 266042, People’s Republic of China

a r t i c l e in f o

abstract

Article history: Received 8 November 2009 Received in revised form 13 April 2010 Accepted 26 October 2010 Available online 7 November 2010

A two-dimensional (2D) transient model was developed to simulate the local hydrodynamics of a gas (flue gas)–solid (CaO)–solid (CaCO3) three-phase fluidized-bed carbonator using the computational fluid dynamic method, where the chemical reaction model was adopted to determine the molar fraction of CO2 at the exit of carbonator and the partial pressure of CO2 in the carbonator. This investigation was intended to improve an understanding of the chemical reaction effects of CaO with CO2 on the CO2 capture efficiency of combustion flue gases. For this purpose, we had utilized Fluent 6.2 to predict the CO2 capture efficiency for different operation conditions. The adopted model concerning the reaction rate of CaO with CO2 is joined into the CFD software. Model simulation results, such as the local time-averaged CO2 molar fraction and conversion of CaO, were validated by experimental measurements under varied operating conditions, e.g., the fraction of active CaO, chemical reaction temperature, particle size, and cycle number at different locations in a gas–solid–solid three-phase fluidized bed carbonator. Furthermore, the local transient hydrodynamic characteristics, such as gas molar fraction and partial pressure were predicted reasonably by the chemical reaction model adopted for the dynamic behaviors of the gas–solid–solid three-phase fluidized bed carbonator. On the basis of this analysis, capture CO2 strategies to reduce CO2 molar fraction in exit of carbonator reactor can be developed in the future. It is concluded that a fluidized bed of CaO can be a suitable reactor to achieve very effective CO2 capture from combustion flue gases. & 2010 Elsevier Ltd. All rights reserved.

Keywords: Chemical reactors Fluidization CO2 capture Dynamic simulation Combustion flue gas Conversion

1. Introduction Increasing atmospheric concentration of CO2, and concerns over its effect on climate, are powerful driving forces for the development of new advanced energy cycles, incorporating CO2 capture and storage. It is generally accepted that the cost associated with the separation of CO2 from flue gases introduces the largest economic penalty to this mitigation option (Herzog, 2001). At present, CO2 is separated from low pressure flue gases or from high pressure natural gas by using a liquid or solution as an absorbent (Nunge and Gill, 1963). The process typically works in the temperature interval between 313 K (absorption) and 393 K (desorption). Main drawbacks of the solvent based absorption processes are high energy requirement and possible environmental issues due to the loss of alkanolamine as a consequence of its high volatility. Separation processes based on solid sorbents might be an alternative, and during the past decade a number of silica, zeolite, carbon, and polymer based sorbents have been developed (Chang et al., 2003; Harlick and Tezel, 2004; Harlick and Sayari, 2006; Huang et al., 2003; Khatri et al., 2005; Kim et al., 2005; Satyapal

n

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et al., 2001; Siriwardance et al., 2001; Xu et al., 2002; Zheng et al., 2005). However, a main drawback of many of these is their relatively limited surface areas, their limitation in the structural design, and the limited possibility for surface modification. For current combustion systems, the only proven commercially available technology to separate CO2 is based on amine-based absorption systems. However, this technology introduces severe efficiency penalties and added costs (Rao and Rubin, 2002), and this justifies a range of emerging approaches that claim to be more energy efficient and cost effective than low temperature absorption-based systems. From the different approaches available for capturing CO2, we focus in this paper on the separation of CO2 from a combustion flue gas using regenerable solid sorbents based on the carbonation/calcinations loop of CaO/CaCO3. The basic scheme of the process proposed is depicted in Fig. 1. Proposed by Shimizu et al. (1999), it has been the subject of detailed analysis elsewhere (Wang et al., 2004). The system consists of a carbonation reactor, where a flue gas from a power plant meets a flux of CaO ready to react with the CO2 present in the gas to form CaCO3. A second reactor is used for the regeneration of the sorbent (calcinations of CaCO3). The carbonation reaction can take place in a reducing atmosphere to enhance H2 formation (Lin et al., 2002; Lopez Ortiz and Harrison, 2001; Wang et al., 2004) or in a combustion flue gas (Abanades and Alvarez, 2003; Wang et al., 2004).

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Fig. 1. CaO carbonation/calcinations cycle to capture CO2.

One of the critical units in these CO2 capture systems is the carbonator reactor itself. High-reaction rates between the CO2 in the flue gas and the sorbent particles are necessary in order to design compact absorbers. Fresh calcined lime is known to be able to carbonate very readily at appropriate temperatures. But the average sorbent particle in the system must experience many carbonation/calcinations cycles (Abanades and Alvarez, 2003) and sorbent capture capacity will decrease rapidly in these conditions. Previous studies that have investigated the reversibility of the carbonation/calcinations reaction have shown that carbonation is far from reversible in practice (Abanades and Alvarez, 2003; Bathia and Perlmutter, 1983; Grasa and Abanades, 2006; Silaban and Harrison, 1995). After a fast, chemically controlled, initial reaction stage, a second slower reaction stage controlled by diffusion in the product layer, CaCO3, takes place (Bathia and Perlmutter, 1983). It was also observed that the transition between the fast and slow regimes takes place quite suddenly at a given level of conversion and this level of conversion decreases, when the number of carbonation/calcinations cycles increased. The evolution of the capture capacity of natural sorbents, over a number of cycles, has been studied in the previous works (Abanades and Alvarez, 2003; Grasa and Abanades, 2006) with varying the process variables. It was found that capture capacity decreases dramatically during the first 20 cycles and then tends to stabilize to a residual conversion of around (7.5–8%), which remains constant up to at least 500 cycles (Grasa and Abanades, 2006). Calcinations temperatures over 950 1C, and/or extended calcinations times accelerate sorbent degradation, and a residual capture capacity with a lower number of cycles is reached. Detailed observation by mercury porosimetry and SEM of the textural changes of the limestone throughout cycling (Alvarez and Abanades, 2005a, b) concluded that the main mechanism of sorbent deactivation is the progressive sintering or grain growth of the originally rich texture of the material resulting from the first calcinations. According to this mechanism, the CaCO3 formed during carbonation will fill up all the available porosity made up of small pores plus a small fraction of the large voids, where the product layer grows until a critical thickness is reached marking the onset of the slow carbonation period. Second-order effects (pore mouth blockage, isolated voids in the calcined material, particle shrinkage) can also be detected in some special sorbents and conditions (Alvarez and Abanades, 2005a, b). However, what marks the reduction in the capture capacity is the grain sintering mechanism as the number of cycles increases combined with the modest product layer thickness allowed on the surface surrounding the large voids. To prevent this decay in activity, some methods are being proposed to manufacture improved CaCO3 sorbents (Aihara et al., 2001; Gupta and Fan, 2002). However, these methods might overshadow one of the key advantages of methods following a lime carbonation–calcinations route, since natural limestones are very cheap sorbents that allow for the large makeup flows contemplated in Fig. 1. In all previous studies of the capture capacity, the quantification of the carbonation rate during the initial fast reaction period under typical CO2 absorber conditions (combustion flue gases at an atmospheric pressure) has not been studied in detail. A basic

reaction model (homogeneous model) proposed by Grasa et al. (2008) has been proved to be sufficient for interpreting the reactivity data obtained under different conditions: partial pressure of CO2, particle size and other relevant operation variables for the carbonation/calcinations loop. The intrinsic rate parameter was found to be 3.2 and 8.9  10  10 m4/mol s. Kinetic data for the carbonation reaction have usually been adopted (Shimizu et al., 1999; Wang et al., 2004) from studies conducted with CaO particles that have undergone only one calcinations cycle. However, it is reasonable to expect that the drastic texture changes behind the decay in the sorbent capture capacity at the end of the fast carbonation period will also produce drastic changes in the carbonation rates. Here, we propose the application of computational fluid dynamics (CFD) to quantitatively analyze CO2 capture behavior based on the application of fundamental physics along with turbulence and chemical reaction models. A commercial software package (Fluent 6.2) is applied for this purpose. A parametric analysis has been carried out for several fractions of active CaO (fa) and cycle number of CaO (N) and reaction temperature, and particle size. Material balances and momentum balance are presented to compute the CO2 capture efficiency. Fluidized beds are a natural choice for carbonator reactor to capture CO2, because of the high-reaction rates required, and the high enthalpy of the carbonation reaction. Fluidized beds have already been used in practice to capture CO2 with CaO, operating at high pressure in the acceptor gasification process (Curran et al., 1967). Despite the increasing number of published works that deal with different aspects of such systems (sorbent performance and reactivation studies, batch experiments and modeling, process simulation work, etc.), there is lack of information available on the actual performance of a fluidized bed of CaO working as CO2 absorber at the typical low CO2 partial pressures of combustion flue gases. Therefore, the first objective of this work was to obtain experimental information on reactivity for highly cycled lime particles that would represent the average sorbent particle in the proposed capture system. These data are the critical parameters for the design and performance of the carbonation reactor. The second objective was to fill this knowledge gap by proposing a basic reaction model based on simple assumptions about the fluiddynamics of the reactors involved and by integrating existing knowledge about sorbent capture capacity and reactivity and chemical reaction constant that seem to be valid for a wide range of operating conditions. Ultimately, it was our aim to show that a fluidized bed of CaO is a very effective device to capture CO2 emitted from small pilot plants, hence, supporting the development of CO2 capture concepts as outlined in Fig. 1. These computer simulations can be further used for predictive purposes to test different scenarios that may be too risky or expensive to try in an existing operational industrial power and combustion plants. Recently, various environmental regulatory authorities have been concerned with the assessment of CO2 control strategies to minimize air quality impacts. Such strategies could include changes in operational practices as well as active feedback control strategies based on actual plant measurements. Our simulation models could be useful in developing these control strategies.

2. Model description CFD relies on solving conservation or transport equations for mass, momentum, energy, and participating species. If the flow is turbulent, model equations for specific turbulent quantities have to be solved in addition. Since even with today’s supercomputers resolving turbulent length scales directly results in tremendous efforts. Reynolds averaged equations are typically applied to include the physics of turbulence. Hence the basic model equations for a fluid in the

C. Cao et al. / Chemical Engineering Science 66 (2011) 375–383

turbulent flow are the Reynolds Averaged Navier–Stokes (RANS) equations. For steady state, these equations are given below.

individual conversion NX ¼1

Xave ¼ 2.1. Continuity equation   r rv ¼ 0

ð1Þ

e

ð5Þ

2.3. Species transport equations Finally, for reacting systems, the species transport equations must be solved. In general form, this equation is given by     @ ðrYi Þ þ r rvYi ¼ r J i þ Ri ð6Þ @t Fluent applies the finite volume method to discretize and solve the governing flow equations described above. A fitting exercise was carried out to obtain kinetic constants for the carbonation reaction, using the entire experimental results with conditions ranging Tcarb 823–973 K; Tcal 1123–1223 K; particle diameter 0.25–1.0 mm; PCO2 0.0–0.08 Mpa; sorbent type limestone (Zibo) and cycle numbers up to 500. The carbonation rate is described by the first-order Eq. (7), similar to the one used by earlier works referred to in Bathia and Perlmutter (1983) and consistent with a grain model for the carbonation reaction in the interior of the particle ð7Þ

In this equation, Save has been estimated from the Xave . Xave ðrCaO =PM,CaO Þ emax ðrCaCO3 =PM,CaCO3 Þ

ð11Þ

For a perfect mixed model, fa is defined as  fa ¼ ð1et =t Þ

ð12Þ

where t is the average particle residence time in the carbonator, which is defined as NCa WCaO ¼ : FR 56FR

ð13Þ

This rate expression of Eq. (7) is similar to the one used in previous works (Abanades et al., 2004; Grasa and Abanades, 2006), but without the term ð1XÞ2=3 , which is characteristic of grain models. This modification is a minor one in quantitative terms, considering that conversions are typically low except in the first few cycles and that there will be few particles of this type in the continuous process of Fig. 1. Bathia and Perlmutter (1983) also used a similar first-order expression and found that the intrinsic carbonation rate constant was around 5.95  10  10 m4/mol s. The rate constant for highly cycled particles has recently been determined by Grasa and Abanades (2006) and it is consistently within a range 3.2  10  10–8.9  10  10 m4/mol s. Unless stated otherwise, in this work, we adopt a conservative value of 4.5  10  10 m4/mol s. With the rate of reaction of the active particles in the carbonator defined by Eq. (7), it is possible to formulate the carbon mass balance in the gas phase in a differential element of the carbonator reactor (Alonso et al., 2009). Assuming that the gas passes in plug flow through a bed of perfectly mixed solids, the balance for a differential element is FCO2 ,0

Save ¼

fa lnð1=ð1fa ÞÞ

ð4Þ

Robustness, economy, and reasonable accuracy for a wide range of turbulent flows explain the popularity of this model in industrial flow and heat transfer simulations.

Rave ¼ kS Save ðCCO2 Ceq Þ

ð10Þ

ðF0 þFR ÞN

XCaO ¼ Xave

t¼

The Boussinesq approach is used in the k–e model, which is the turbulence model applied in this work. The advantage of this approach is the relatively low computational cost associated with the computation of the kinetic viscosity. For the k–e model, the eddy viscosity is obtained from Prandtl–Kolmogorov relation Cm k2

F0 FRN1

ð2Þ

where k is the kinetic energy of turbulence, defined by

vt ¼

ð9Þ

where (see reference Abanades and Alvarez, 2003) rN ¼

The Reynolds stresses, rvuvu, are extra terms that stem from decomposing solution turbulent variables into the mean and fluctuating components; these terms must be modelled in order to close Eq. (2). A common approach employs the Boussinesq hypothesis to relate the Reynolds stresses to the mean velocity gradients   @v i @v j 2 vuvu ¼ vt þ ð3Þ  ðvt ðr vÞ þ kÞdij 3 @xj @xi

1 k ¼ vuvu 2

rN XN

N¼1

2.2. Momentum conservation equation

r ðrvvÞ ¼ rp þ r ðmðrv þðrvÞT ÞrvuvuÞ

377

¼ Afa

The average maximum carbonation conversion that can be achieved by the particles in the carbonator is the average of the

rCaO PMCaO

kS Save rM,g

  ðf0 fe Þ þ ðf0 fe f0 ÞEcarb : ð1f0 Ecarb Þ

ð14Þ

Using a dimensionless variable for height, the integrated form of this equation is Z¼

"  # FCO2 ,0 PMCaO f0 f0 ðf0 1Þ ðf0 fe Þ þ ðf0 fe f0 ÞEcarb Ecarb þ ln WCaO kS Save rM g fa ðf0 fe f0 Þ ðf0 fe Þ ðf0 fe f0 Þ2

ð15Þ where f0 is an inlet molar fraction of CO2. The equilibrium molar fraction of CO2 is calculated from the expression (Barker, 1973) fe ¼

ð8Þ

dEcarb r ¼ Afa CaO Rave dz PMCaO

pe 10ð7:0798308=TÞ : Ptotal

ð16Þ

So that the carbonation efficiency in the carbonator reactor can be defined as a simple function of fa Ecarb ¼

FR fa : Xave FCO2 ,0 lnð1=ð1fa ÞÞ

ð17Þ

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2.4. Simulation methods and boundary conditions The computations were carried out using Fluent 6.2 after making the mesh in GAMBIT 2.0. The simulation was performed using a 2D system axisymmetric model. The whole carbonator domain was modelled with CFD code to save the computational work. Steady-state-averaged solutions were obtained. Pressure operation was taken as 0.01 MPa. The flow has very low Mach number and therefore could be considered incompressible. Turbulence was modelled by using the classical k–e turbulence model. The flow domain mesh was packed towards the axis of symmetry. To describe the reaction of CaO with CO2, appropriate chemical reaction mechanisms were included. For this purpose, the detailed reaction kinetics was successfully implemented in Fluent. The simulation procedure followed as follows (David and Thomas, 2006). An initial solution was obtained by using a simple combustion model in Fluent (mixture fraction model). The mixture fraction model assumes that the species and enthalpy transport equations collapse into a single transport equation for the mixture fraction under some specific assumptions. The mixture fraction can be directly related to species mass fraction, mixture density, and mixture temperature. Interaction of turbulence chemistry is accounted for with a probability density function (PDF). This model also performs equilibrium calculations if an infinitely fast chemical reaction is assumed. From that initial solution, the Eddy Dissipation Concept (EDC) model was used to refine the results. The EDC model takes advantage of the assuming part of the fluid to be thoroughly mixed within a particular cell as well as to be the main driver for the chemical reaction. These well-mixed portions of a sub-volume, the so-called ‘‘fine scale’’, resemble a constant pressure reactor cell. Using the resulting simplified equations for species conservation, the corresponding source terms are defined by Arrhenius-type reaction mechanism. To speed up the calculation, the ISAT algorithm can be applied. Notice that the EDC model can incorporate detailed chemical mechanisms into turbulent reacting flows, although it requires more computational effort than the mixture fraction and laminar flamelet models. Boundary conditions for the reaction of CaO with CO2 simulations are briefly summarized below. For the inlet streams, velocity inlet was used in the Fluent. The species mass fractions were set based on the inlet gas compositions. Moreover boundary conditions at the outside must be defined to represent an unconfined boundary. This was typically achieved by using the pressure outlet condition in the Fluent. The segregated, steady state solver was used for all computations. After the preliminary solution was obtained with the non-premixed/mixture fraction approach, the species model was set-up as an EDC model. The chemical reaction was updated at each flow iteration, and the energy equation was enabled. The standard k–e viscous model was used with default values for the model constants. The ISAT method was used to save the computational work. The absolute and relative error tolerances were set as 1  10  6 and 1  10  9, respectively. The ISAT error tolerance was set as 1  10  4. The maximum storage capacity was set as 200 MB. The reference time step was set as 0.001. A second order upwind scheme was used for all equations. The SIMPLE algorithm with the multi-grid solver was used to couple pressure and velocities. Under-relaxation factors were initially set as default values, although they were adjusted at the final iterations in order to get the solution converged. The flow was always initialized from flue gases inlet conditions. Equation residuals were checked after iteration for convergence. The convergence criterion based on residuals was set as 1  10  6 for each equation.

designed for long multicycle carbonation/calcinations tests and to derive reactivity data during carbonation. The mini-circulating fluidized bed used in this work is presented in Fig. 2. This has been described in detail elsewhere (Abanades et al., 2004). It consists of a dense bed region, riser section, and cyclone. The so-called dense bed region is 1 m high, and has an internal diameter of 100 mm. This section is also surrounded by four electric heaters (18 kW total), which can provide supplemental heat during operation. Heaters can maintain the dense bed region at a maximum temperature of 1223 K. Air entering the dense bed region passes through a plenum or windbox, and is forced up through a distributor plate. Situated approximately 1 m above the distributor plate are the solid feed inlet port, return-leg port, and the natural gas burner. The natural gas burner is used to provide heat to the riser on startup. The riser has an internal diameter of 100 mm, is 5 m long, and is insulated with 75 mm of refractory. The circulating fluidized bed is equipped with a data acquisition system, which records the system temperature, pressure drop, and gas composition. Temperatures in the dense bed region are measured at four different points by K-type thermocouples: 120, 240, 360, and 480 mm from the distributor plate. Thermocouples are also situated along the riser, cyclone, and return leg. The pressure drop in the riser is measured by a series of pressure taps. Gas sampling is performed at the top of the riser; detectors can record the levels of O2, CO2, CO, SO2, and NOX. Solid samples can be collected at the base of the return leg or via a valve immediately above the distributor plate in the dense bed region. The carbonation/calcinations tests conducted in this work were run in the batch mode for the solids. Different variables affecting

3. Experimental set The cyclic carbonation and calcinations reactions were experimentally studied in a mini-circulating fluidized bed (CFB) specially

Fig. 2. Scheme diagram of experimental equipment set-up for CO2 capture, using the carbonation/calcinations loop of CaO.

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379

Fig. 3. SEM micrographs of CaO particle after one calination (a) and after 30 carbonation/calcinations cycles (b) in the fluidized bed.

the carbonation process were studied in this work (particle diameter, partial pressure of CO2, volume fraction of CO2, carbonation temperature) in a series of calcinations/carbonation tests range 1–300 of cycles. They were initiated by loading 5 kg of limestone to form a bed around 0.5 m height at minimum fluidization conditions. Initial experiments were carried out to determine the total gas flow needed to eliminate external diffusion effects around the sample basket (this was finally set to 1.0 m/s superficial gas velocity at 923 K). The temperature and sample weight were continuously recorded on the computer. Experiments were performed with an empty sample and inert material in order to determine the disturbances in the weight readings when solid sample was moved from the calcinations to the carbonation reactor. Limestone named ‘‘Zi Bo’’ from Shandong province of China was used, and the particle size ranges were between 0.25 and 1.0 mm. The limestone beds were subjected to cyclic calcinations in air at temperatures around 1193 K, and carbonation at 922 K, with a synthetic gas mixture of air and 15% by volume of CO2. The calcinations part of every cycle was carried out by switching on the external bed heaters and allowing for a small flow of gas through the bed. There was 6 and 11 calcinations–carbonation cycles completed for each limestone, similar to those presented in Fig. 3 for the latest cycles with Zi Bo. During the calcinations of Zi Bo, a few short periods of fluidization were applied to mix the solids in the bed. The implications of this change in calcinations conditions will be discussed below. During the carbonation part of the experiments, the bed was fully fluidized at 1.0 m/s with a synthetic gas mixture of air and CO2. The concentration at the exit of the system showed that capture of CO2 was very effective in these conditions, with values of CO2 concentration apparently lower than equilibrium at the reference bed temperature under certain conditions. This was attributed to the residual recarbonation in the upper, cooler parts of the riser. Samples were collected after completion of the carbonation part of the cycle, which was clearly marked by a rapid increase in the CO2 concentration profile (break through) at the exit of the bed, as can be seen in the example of Fig. 5. Carbonated samples were collected after every calcinations– carbonation experiments, and subjected to a range of techniques to aid in the interpretation of results. A scanning electron microscope (SEM, Zeiss DSM 942) was used for an examination of the internal structure of the carbonated sorbent particles and their calcines. In order to obtain fresh fracture surfaces for observation of the interior of the particles, the particles were lightly crushed by pressing between two pieces of glass, and the sample thus obtained was dispersed on a sticky graphite tab placed on an aluminum stub. The samples were gold-coated (  20 nm thickness) in order to improve the electronic signal, and also, to protect the surfaces from hydration and/or carbonation. The SEM images were formed from

the backscattered electron signal, which normally yielded better quality pictures than the secondary electrons.

4. Results and discussion Fig. 3 provides an SEM image of the carbonated samples of Zi Bo after different calcinations/carbonations cycles. The particle has been broken in order see the particle surface texture. The SEM image shows that the particle surface is heavily sintered in comparison to the particle interior. This reduced porosity effectively created a shell around the particle, which may explain the relatively short kinetically controlled regime observed during different calcinations/carbonations cycles. Comparing Fig. 3(a) and (b), it can be clearly seen that the particle surface has far more large pores with increase in calcinations/carbonations cycles. 4.1. Effect of cycle numbers on CaO conversion Fig. 4 shows the typical conversion curves vs. reaction time obtained in the experiments. The figure contains curves corresponding to different cycles and, as can be seen, the main features of the carbonation reaction are presented in all the cycles (initial fast reaction period followed by an abrupt change to a slower reaction stage). It is clear too that although the maximum capture capacity decreases with increasing the number of cycles, the path followed to reach the maximum conversion at different cycles is very similar. In these conditions, reaction times in the order of a couple of minutes are sufficient to reach the end of the fast reaction period, but clear differences in the slope of these curves are also evident. Fig. 5 presents an example of the CO2 concentrations measured at the exit of the fluidized-bed carbonator in three different carbonation cycles and model predictions under average carbonation conditions. 4.2. Effect of particle size on CaO conversion To investigate the possible effect of particle size on the carbonation reaction rate, four narrow particle size fractions (0.25–0.4, 0.4–0.6, 0.6–0.8, and 0.8–1.0 mm) of a limestone named ‘‘Zi Bo’’ were tested in the carbonator reactor of the fluidized bed. Experimental results are plotted in Fig. 6. From previous studies on the carbonation reaction of CaO (Silaban and Harrison, 1995; Dennis and Hayhurst, 1987), it was expected that the particle size would have a strong effect on the overall particle carbonation rates. Particle with a pore structure similar to Fig. 3(b) were expected to show an increasing resistance to CO2 diffusion towards the free CaO surface in the interior of the particle as the particle sizes increased. Fig. 6(a) confirms this tendency, as it shows an example of the

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results obtained under the typical reaction conditions expected in the CO2 absorber (928 K and CO2 volume fraction below 15%) and typical particle size ranges for fluidized bed systems. For the first calcinations cycle, it can be concluded that diffusion effects inside the particle must be responsible for the slower carbonation rates of the larger particles. However, the quantitative differences in the slopes of curves are modest, and only small differences in time were needed to achieve 55% conversion of particles (especially for the 0.25–0.4, 0.4–0.6, and 0.6–0.8 mm intervals). Furthermore, the maximum conversion attained was very similar for the particle

range studied. This indicates that the main rate resistance to CO2 reacting with CaO was of a chemical nature. Fig. 6(b) shows the conversion curve for the cycle number 30. At this stage, the sorbent capture capacity (conversion at the end of the fast reaction period) is lower, as expected, and the differences in slope for the fast reaction period between the different particle sizes have disappeared. This indicates that there are no diffusion effects in the interior of the particles and the reaction rate must be controlled by the reaction mechanism taking place uniformly on the free surface CaO. 4.3. Effect of reaction temperature on CaO conversion The carbonation reaction was studied in a range of temperature from 823 to 973 K, close to the operation conditions in the proposed capture system. Fig. 7(a) and (b) shows curves obtained for cycles 35 and 140, respectively. As can be seen, the slopes corresponding to the linear stage of the carbonation curve are very similar for the range of temperatures studied. This indicates the poor dependency of the kinetic parameter on temperature in agreement with the results obtained in literatures. 4.4. Effect of the fraction of an active CaO on the axial CO2 concentration profiles

Fig. 4. Conversion curves vs. time for different cycle numbers.

Fig. 5. Experimental CO2 concentrations measured at the exit of the fluidized-bed carbonator in three different carbonation cycles.

Fig. 8 presents an example of the CO2 concentration profile in the fluidized-bed carbonator calculated with Eq. (15), and the auxiliary equations above, for conditions resembling those used during the experiments. As can be seen in this figure, the CO2 concentration profiles are insensitive to values of fa higher than 0.1. This corresponds to the early stages in the carbonation cycle during the experiments, where the emulsion phase is acting as a very effective sink for CO2 and the overall carbonation process is controlled by the transfer of CO2 from the bubble phase to the emulsion phase. With lower values of fa, the concentration at the exit of the bed is appreciable, and this corresponds to the beginning of the breakthrough curves. The integration starts by calculating the concentration at the exit of the bed at the beginning of the experiment (fa ¼Xb,N for t ¼0). This exercise was undertaken to produce the simulated CO2 concentrations at the exit of the carbonator that were included in Fig. 5 as continuous dotted lines. As can be seen in this figure, there is a reasonable agreement with the experimental results, when considering the number of simplifications and assumptions adopted to build the fluidized-bed carbonator model. It is important to emphasize that the sensitivity of these curves is low with respect to the assumptions and parameters adopted by the carbonation rates at the particle level, at the conditions tested; it is sufficiently large to guarantee a rapid change (both experimental

Fig. 6. Conversion curves vs. time for different particle size.

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381

Fig. 7. Conversion curves vs. time for different Tcarb.

Fig. 8. Effect of the fraction of active CaO present in the bed on the axial CO2 concentration profiles as a simulation under carbonation conditions used during the experiments.

and theoretical) in the CO2 concentration at the exit of the bed in the proximity of the breakthrough conversion. The comparison between predicted and experimental CO2 concentration data is poor for the lower sampling ports (0.25 m above the distributor). While the model predicts significant concentration of CO2 in the gas phase even at the beginning of the experiment (maximum fa), all the experimental tests showed that, at the beginning of the carbonation period of each cycle, the bed was very effectively absorbing CO2 even at this low sampling port. This discrepancy is due to the correlation adopted for the bubbleto-emulsion transport of CO2 that seems to be too conservative for the actual bed conditions, since it does not allow more pronounced CO2 concentration profiles, even when this transport is the only resistance to progress of the carbonation reaction in the bed. It is, however, beyond the scope of this work to refine this correlation for the limited number of fluidized-bed experiments conducted so far. It can also be noted that the breakthrough curves in Fig. 5 show good agreement on the expected breakthrough times, but some different from the model predictions in the shape of the CO2 profile. As previously mentioned, recarbonation in cooler parts of the freeboard and some air leak into the freeboard were detected in some carbonation tests. This explains why the intrinsic scattering of data seems higher in the experimental concentration of CO2 measured at the exit of the riser than in that CO2 concentration measured inside and just above the fluidized bed. 4.5. Model prediction of local transient hydrodynamics As the measuring technology of the local transient dynamic behaviors of the whole carbonator is rather limited, model predictions of the local transient hydrodynamics as well as CO2 concentrations at

the exit of the carbonator are necessary for understanding and analysis of the factors, which may affect the multiphase reaction processes, such as gas velocity, reaction temperature, CO2 partial pressure, and mass transfer characteristics. Fig. 9 is the model prediction of the local transient hydrodynamic characteristics of the carbonator under different operation conditions and fixed initial solid loading of 5 kg and time of 20 s at the vertical section. The nine charts in Fig. 9 show the model prediction local transient CO2 molar fraction and CO2 partial pressure drop from down to up, and at different CaO cycle numbers and the reaction temperature. These charts visually exhibit the instantaneous dynamic behaviors of the local transient hydrodynamics of the carbonator. Gas bubbles are carried downward by the CaO particles circumfluence near the wall of the carbonator. A gas bubble plume is formed and moves wiggly in the carbonator resulted from the interphase momentum transfer forces. Large vortices alongside the bubble plume are engendered, which further results in the nonuniform distribution of the CaO particles within the carbonator. It can be concluded from the figure that it is the pseudo-periodical wiggle of the gas bubble plume that originates the dynamic changes in the CaO particle velocities as well as the gas holdup distribution. 4.6. Prospects The carbonation rate parameters obtained in this work need to be put in the context of ongoing developments in order to design a full carbonation/calcinations pilot system following the scheme in Fig. 1. The reaction rates measured for particles cycled a high number of times, the fact that they follow a homogenous reaction model during their carbonation, and the range of intrinsic reaction constants obtained suggest that fluidized bed systems can be effective absorbers for capturing the CO2 present in the flue gases from power plants, even when the particles in the bed have experienced many carbonation/calcinations cycles. Fluidized bed reactors can accommodate solid residence times of several minutes, axial hold ups of several hundreds of kg/m2 and solid circulation rates in the order of 10 kg/m2 s of CaO (equivalent molar flow of approximately 180 mol CaO/m2 s). This larger flow of CaO should be able to match the poor conversion of CaO to CaCO3 and still capture a relevant fraction of the CO2 in the flue gas flowing through a CFB absorber. Therefore, a range of reasonable operating conditions for the absorber, not far from the standard in existing large scale CFB systems, should allow an effective CO2 capture from the flue gas from a power plant in a system similar to the one depicted in Fig. 1. Finally, it is necessary to highlight that, from a practical point of view, the most interesting part of the experiments and the model simulations are those with low values of fa. This is because, in continuous carbonation/calcinations systems to separate CO2, it will be a design objective to maximize utilization of the

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from the calcinations is much more pronounced. Under these conditions, the general bed characteristics (superficial gas velocities, bed temperature, bed heights, bubble behavior, and so on) strongly affect performance of the fluidized bed carbonator in terms of CO2 capture efficiency in the gas phase. Despite these remarks, and in view of the results obtained so far, it has been shown that some attractive operating conditions exist, where a fluidized bed of free CaO is an effective absorber of CO2 from a combustion flue gas at high temperature. Further work on the CFD modeling of the real multiphase reaction processes coupling CaO velocity, mass transfer, and intrinsic kinetics in the gas–solid–solid fluidized bed carbonator is under way.

5. Conclusions A 2D transient CFD model for simulating the local hydrodynamics of a gas–solid–solid three-phase fluidized bed reactor, with the reaction model of CaO with CO2 adopted to determine conversion of CaO to CaCO3 and CO2 molar fraction at the exit of the carbonator, was validated to be reliable by time-averaged data of the experimental measurements. Local transient hydrodynamics of the gas–solid–solid three-phase fluidized bed reactor, such as CO2 molar fraction, conversion of CaO to CaCO3, CO2 partial pressure were also predicted reasonably by the adopted model. The study on the dynamic behaviors of the CO2 molar fraction may provide a useful method for understanding the mass transfer characteristics of the gas–solid–solid three-phase reaction in the fluidized bed carbonator. The carbonation reaction rates seem to be suited to the range of residence times typical of circulating fluidized beds. This method is under improvement for simulating real combustion behavior, where interphase mass transfer and intrinsic kinetics should be coupled together. Moreover this developed method can be further applied in many fields, such as to optimize the design and construction of the gas–solid–solid three-phase fluidized bed reactor and the operation of some main processes (like power plants, cement plants, steel mills, refineries, etc.), as well as to facilitate the scaleup strategies. Detailed knowledge of the local fluid-dynamics is very important in design and scaleup of the multiphase reactors. However, non-invasive optical techniques, such as LDA and particle imaging velocimetry, are restricted to characterizing velocity fields with low molar fraction of CaO. Thus nonoptical techniques, like computer-automated radioactive particle tracking, and computer-assisted tomography, may provide useful data for opaque systems at higher values of CaO molar fraction, which will be used in future study on the three-phase combustion flue gases process.

Nomenclature A CCO2 Ceq emax Fig. 9. Prediction of local transient hydrodynamic characteristics of the carbonator: (a) dp ¼ 0.85 mm, UG ¼1.0 m/s, T¼ 925 K, t¼ 20 s, WCaO ¼ 5 kg (FCO2); (b) dp ¼ 0.85 mm, UG ¼1.0 m/s, N¼ 20, t¼ 20 s, WCaO ¼ 5 kg (FCO2) and (c) dp ¼0.85 mm, UG ¼ 1.0 m/s, N¼25, t¼20 s, WCaO ¼ 5 kg (PCO2).

Ecar f0 fa fe

sorbent and minimize losses of an active CaO. Therefore, the value of fa in a continuous operation must be kept low. For low values of fa, the sensitivity of the model to the reactivity of the sorbent arising

F0 FCO2,0

carbonator section, m2 CO2 bulk concentration, mol/m3 the CO2 equilibrium concentration at reaction conditions, mol/m3 maximum thickness of the layer of CaCO3 on the pore wall, 50 nm CO2 capture efficiency in the carbonator inlet molar fraction of CO2 volumetric fraction of CaO that reacts in the carbonator in the fast reaction regime molar fraction of CO2 at the point of equilibrium in the reaction conditions molar flow rate of fresh limestone, mol/s CO2 molar flow rate at the inlet of the carbonator, mol/s

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FCO2 FR Ji k ks N p pe PM,CaCO3 PM,CaO Ptotal Rave Ri rN Save tn t Tcal Tcarb UG WCaO Xave Xb, N XCaO XN z

CO2 molar fraction molar flow rate of CaO coming from the calciner, mol/s diffusion flux of species i, kg/m2 s turbulence kinetic energy, m2/s2 rate constant for the carbonation reaction at the surface, m4/mol s CaO cycle number local pressure, pa Equilibrium pressure, pa molecular weight of CaCO3, g/mol molecular weight of CaO, g/mol total pressure, pa first-order average reaction rate of the active material, 1/s net rate of production of species i by the chemical reaction, kg/m3 s particle fraction in the N cycle maximum average reaction surface, m2/m3 characteristic time at which the reaction rate becomes zero, s reaction time, s calcinations temperature, K carbonation temperature, K superficial gas velocity, m/s solid inventory in the carbonator, kg maximum average conversion of particle carbonation conversion at the end of fast reaction period in the Nth calcinations/carbonation cycle conversion of CaO to CaCO3 at the exit of the carbonator maximum conversion of particles in the Nth calcinations/ carbonation cycle dimensionless height from the distributor

Greek letters

e r rCaO rM,g m dij v v0 vt

turbulence energy dissipation, m2/s3 fluid density, kg/m3 CaO particle density, g/m3 molar density of the gas, mol/m3 dynamic molecular viscosity of the fluid, Pa s ¨ Kronecker delta, dimensionless ensemble-averaged velocity vector, m/s turbulent fluctuation of the velocity vector, m/s eddy kinetic viscosity, m2/s

Abbreviations CFB CFD EDC ISAT LDA PDF RANS SEM

circulating fluidized bed computation fluid-dynamics eddy dissipation concept in situ adaptive tabulation laser doppler anemometry probability density function Reynolds averaged Navier–Stokes Scanning electron microscopy

Acknowledgements The financial support of this research by the Taishan Mountain Scholar Constructive Engineering Foundation of China (Js200510036) and by the National Natural Science Foundation of China (20876079).

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