Progress of Combustion in an Oxy-fuel Circulating Fluidized-Bed

Article pubs.acs.org/EF

Progress of Combustion in an Oxy-fuel Circulating Fluidized-Bed Furnace: Measurements and Modeling in a 4 MWth Boiler Sadegh Seddighi K.,* David Pallarès, Fredrik Normann, and Filip Johnsson Division for Energy Technology, Department of Energy and Environment, Chalmers University of Technology, SE-412 96 Göteborg, Sweden ABSTRACT: Oxy-fuel combustion, which is a promising technology for the abatement of carbon dioxide emissions, can be applied in circulating fluidized-bed (CFB) power plants. In this study, the effects of operational conditions on the progress of oxy-fuel CFB combustion were investigated by means of a mathematical model for CFB oxy-fuel combustion together with experimental data from a 4 MWth oxy-fuel CFB, currently representing the largest oxy-fuel CFB combustion experiments in the literature. Modeled in-furnace profiles for carbon monoxide (CO) and oxygen (O2) were compared to the corresponding measurements, yielding a general good agreement for both air- and oxy-fuel-fired conditions. The developed model was also used to investigate the effects of varying the inlet O2 concentration over a wider range than that applied in the experiments. The experimental results show that, for an equivalent inlet O2 concentration, the peak CO concentration is higher under oxy-fuel-fired conditions than under air-fired conditions. The model result shows that a higher inlet O2 concentration generates combustion of greater intensity up through the furnace with a lower level of CO at the furnace exit.

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INTRODUCTION Oxy-fuel combustion is being investigated as a technology for carbon capture in carbon capture and storage (CCS) schemes. Studies on oxy-fuel pulverized coal (PC) have, for example, been carried out by Horn et al.1 and Abraham et al.,2 while studies on oxy-fuel circulating fluidized beds (CFBs) have been conducted by Nsakala et al.3 and Seddighi et al.4 In oxy-fuel combustion, the fuel is burnt in pure oxygen (O2) instead of air and the flue gases are recycled to control the furnace temperature, resulting in a flue gas that, in principle, consists exclusively of carbon dioxide (CO2) and water. The water is readily condensed, and CO2 is transported for compression and storage. O2 is produced in an air separation unit (ASU), which typically employs cryogenic air separation,5 and this is the main contributor to the energy penalty (cost for capture) in oxy-fuel combustion. The CFB technology is characterized by high fuel flexibility, possibility for in-bed sulfur capture, and a low combustion temperature, yielding low thermal NOx emissions. In addition, the high thermal capacity of the circulating bed solids in the CFB enables heat extraction outside of the furnace, i.e., via an external heat exchanger in the return leg. This allows for operation with a high inlet concentration of O2 and a restricted furnace temperature, which is achieved by heat extraction in the return leg.4 However, to assess oxy-fuel-fired CFB boilers, modeling tools that elucidate the progress of combustion in the CFB loop under various oxy-fuel conditions are needed, to provide information for the design and scaling-up of the process. In this work, CO is used to compare modeling and experimental results and as an indicator to track the progress of combustion. The progress of combustion is evaluated using peak CO concentration, horizontal variation in peak CO concentration, and CO concentration at the furnace outlet. CO is a good indicator of the progress of combustion because CO oxidation is an important step in carbon conversion. To date, little information has appeared in the © 2013 American Chemical Society

literature in relation to the progress of combustion in largescale, oxy-fuel-fired CFB units. Seddighi K. et al.,6 using modeling and experiments in a 100 kW oxy-fuel-fired CFB, concluded that for the same O2 concentration, the CO oxidation rate is lower in CO2-dominated environments; i.e., that for an equivalent inlet concentration of O2, the combustion proceeds at a slower rate in an oxy-fuel atmosphere than in an air-fired atmosphere. However, there is, thus far, no experimental data from oxy-fuel-fired CFB units of industrial scale published in the literature. Thus, to gain a better understanding of oxy-fuel-fired CFB and to investigate the parameters that affect the progress of combustion in an oxy-fuel CFB, modeling and experimental exercises in large-scale units are necessary. The present work combines modeling and experiments that are conducted under both air- and oxy-fuel-fired conditions. The modeling draws on an existing three-dimensional (3D) model for large-scale fluidized-bed combustion, which was originally developed and validated for air-fired conditions.7 The aims of the present work are (1) to develop the model for oxyfuel CFB by including a combustion model with expression for axial gas mixing and global combustion reactions that are generalized for both air- and oxy-fuel-fired conditions (the resulting model is validated with unique in-furnace measurements of CO and O2 in a 4 MW oxy-fuel-fired CFB unit) and (2) by applying the derived model, to investigate the effects of setting inlet O2 across a wider concentration range than applied in the experiments on the progress of combustion.

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EXPERIMENTAL SECTION

The experiments were performed in the Metso Power 4 MWth oxyfuel-fired CFB boiler (Figure 1). The furnace has a height of 13 m and Received: June 25, 2013 Revised: September 11, 2013 Published: September 17, 2013 6222

dx.doi.org/10.1021/ef4011884 | Energy Fuels 2013, 27, 6222−6230

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Table 3. Operational Conditions in the Cases Studied case

AF-C1

OF21-C1

OF25-C1

AF-C2

OF25-C2

fuel firing mode [O2]1y (%) [O2]exit (%) stoichiometric ratio share of primary gas Pthermal (MW) Ufluidizationa (m/s) Tbedb (K)

coal 1 air 21 3.1 1.13 0.67 3.93 3.98 1178

coal 1 oxy-fuel 21 2.4 1.02 0.76 3.98 4.27 1064

coal 1 oxy-fuel 25.5 2.4 1.02 0.70 4.71 4.24 1173

coal 2 air 21 3.3 1.19 0.59 4.79 4.44 1169

coal 2 oxy-fuel 25.4 3.8 1.15 0.68 4.73 3.96 1117

a

Fluidization velocity, Ufluidization, is based on the total primary gases injected, including both O2 and RFG. bThe dense bed temperature is measured by a thermocouple measurement at 0.26 m above the primary gas distributor. Figure 1. Oxy-CFB (4 MW) test rig. parameters by extending the range of operational parameters beyond what was applied in the experiments. Modeling. The modeling combines a combustion model by Seddighi K. et al.6 that was developed and validated at a lab-scale 100 kW unit operated under both oxy-fuel and air conditions and a previously developed 3D semi-empirical model for the gas−solid structure in air-fired fluidized-bed combustion7 validated against experimental data from different large-scale units. In the present work, the combustion model6 is included in the 3D CFB model, so that the global reaction mechanism accounts for CO2 atmospheres together with inclusion of the axial in-furnace mixing process in a simplified and, thus, computationally cost-efficient way. In addition, handling of the flue gas recirculation (FGR) is included, with optional partial condensation of the flue gas, as well as air in-leakage to the FGR system. The 3D model used in this work is categorized among semiempirical models, which are based on the experimental findings of the gas−solid structure in large-scale CFB furnaces.12 Use of a semiempirical approach for calculation of the gas−solid structure is used because this yield reasonable calculation times compared to modeling from first principles using computation fluid dynamic (CFD) methods.12 In general, semi-empirical models can only be assumed to be valid for conditions similar to the experiments in which the empirical parameters are found.6 CO is used as an indicator of the progress of combustion in the CFB furnace. The formation of CO is attributed to both heterogeneous and homogeneous reactions, and it is influenced by the gas mixing in the furnace. In fluidized-bed combustion, a favorable operational temperature is 1123 K, to maximize in-bed sulfur capture and prevent agglomeration. This temperature is considerably lower than the operational temperatures applied in PC units, which are around 1800 K. Hence, at the temperature range of the CFB units, gasification reactions (which are important in PC combustion) are generally slower than combustion reactions and are negligible in oxyfuel-fired CFB. The combustion chemistry, including heterogeneous and homogeneous reactions, in relation to CFB conditions is discussed in detail by Seddighi K. et al.6 The major source of CO formation in coal-fired CFB combustors is char oxidation, which can be described as6

a cross-section of 1 × 1 m2. The unit is equipped with a single cyclone. O2 is provided to the system from an O2 tank and mixed with the recirculated flue gases (RFGs) before being injected into the unit. Two types of coal, with low (coal 1) and medium (coal 2) sulfur content (see Tables 1 and 2 for fuel analyses), were used in the tests. The coal

Table 1. Proximate Analyses of the Bituminous Coals Used moisture (wt %) ash (wt %, dry) volatile matter (wt %, dry) fixed carbon (wt %, dry) lower heating value (MJ/kg, ar)

coal 1

coal 2

9.0 11.8 35.9 52.3 26.0

8.8 7.2 34.3 58.5 28.1

Table 2. Ultimate Analysis of the Bituminous Coals Used (wt %, Dry) C H N S O

coal 1

coal 2

70.3 5.0 2.3 0.47 10.1

67.8 4.6 1.3 2.3 7.9

particle size ranges from 0.05 to 10 mm, with a mean particle size of 2.5 mm. Although the coals were not chosen specifically for the present study (i.e., the fate of sulfur is not part of the present study), they were used to extend the number of operational points. Fuel preparation involved crushing and sieving. Table 3 lists the operational conditions for the two air-fired and three oxy-fired cases presented in this work. The concentrations of gases were measured with gas suction probes connected to a Fourier transform infrared (FTIR) analyzer for CO2 and CO and a conventional gas analyzer for O2 at four heights in the furnace, as shown in Figure 2, in which the secondary gas injection levels are also indicated. Suction probes have shown to be simple but fairly accurate measurement devices.8,9 However, the large fluxes of particles in the CFB furnace do not allow for long sampling times and, thus, reduce the accuracy.10 In addition, the pulsating nature of upcoming flow, in particular, at lower parts of the furnace, affects the accuracy of the measurements.11 As shown in Table 3, it was not possible to change one parameter while keeping constant all of the remaining parameters during the experiments. For example, under oxy-fuel conditions, a change in the inlet concentration of O2 implies changes in other parameters, such as RFG properties, fuel feed, and temperature. As indicated above, the aim of the modeling exercise was to assess the influence of individual

αC + O2 → 2(α − 1)CO + (2 − α)CO2

(1)

The heterogeneous reaction expressed by eq 1 involves a complex set of reactions,13 and the products of char oxidation may comprise any combination of CO and CO2. Two major methods have been reported in the literature for calculating the shares of CO and CO2 in eq 1: (1) nth-order Arrhenius models14 and (2) detailed reaction sets (e.g., the Langmuir−Hinshelwood mechanism15). The nth-order Arrhenius models have shown good agreement with experimental results for a wide range of fuels, temperatures, and fuel sizes.14,16−18 In the nthorder Arrhenius approach, the CO2/CO ratio is correlated to one global reaction rate, activation energy, temperature, and partial pressure of O2, as described by14,16−18 6223


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Figure 2. Schematic view of the unit with the injection and probe planes. The fuel feed is at Y = −0.55 m.

CO2 ⎛ −E ⎞ ⎟ = ApOn exp⎜ ⎝ RT ⎠ 2 CO

presented by Thunman et al.,20 yielding the compositions of combustible volatile gases listed in Table 4.

(2)

where the values of A = 0.02, n = 0.21, and E/R = 3070 K are taken from the work by Tognotti et al.14 The modeling of the char conversion rate19 correlates the rate of change in particle mass to the conversion factor (Ω), effective rate constant of Reff (which includes both kinetics and diffusion), particle area (Apart), and surrounding O2 concentration (CO2,∞)20

dmc = −ΩR eff A part CO2, ∞ dt

Table 4. Modeled Composition of the Volatiles in the Two Bituminous Coals Used (Which Are Assumed To Have the Same Volatile Composition)

(3)

where the effective rate constant (Reff) is the combined effect of the kinetic rate (Rkin) and mass-transfer rate (Km) as given by

R eff =

1 1 R kin

+

1 km

(4)

component

weight percent

CO CO2 H2O H2 C1.16H4 C6H6.2O0.2 SO2 NH3

21.9 11.4 3.1 7.9 4.6 40.1 2.8 8.2

6

The major consumption path of CO in CFB is CO oxidation, which can be expressed as

CO + 0.5O2 → CO2

Because the compositions of the two coals used in the present work are similar, the composition of volatiles determined for coal 1 is applied to both coals. As given above, the main difference between coals 1 and 2 is in relation to the sulfur content, which is assumed not to have any influence on the results of the combustion modeling. The composition of the RFG was calculated prior to the run of the 3D model through a zero-dimensional combustion model, which follows an iterative procedure and includes all of the gases in all parts of the CFB loop (riser, seal, bottom ash coolers, and leakage) and fuel injections and which assumes full fuel conversion. Accordingly, the 3D model needs to be run only once, which means that computational costs can be kept at reasonable levels. In oxy-fuel-fired combustion, both wet and dry RFG can be applied. Varol et al.23 have indicated that a high level of moisture may cause operational problems with the fluidized-bed furnace, such as bed agglomeration, fouling, slagging, and corrosion of heat-transfer surfaces when using fuels with a high alkali content. Khan et al.24 have pointed out that a high moisture content may lead to ignition problems. In addition, drying of the RFG in the condenser leads to energy loss in oxy-fuel combustion, and this should be minimized. Overall, the use of dry RFG is the preferred option for the primary fluidization flow. The model in the present work can handle partial drying of RFG, and it calculates the temperature and composition of the dry and wet RFG. The water content of the RFGs leaving the condenser (dry RFG) is calculated according to the equilibrium given by the equation25

(5)

The CO oxidation rate depends upon the temperature (T), activation energy (E), and concentrations of CO, O2, and H2O.21 The CO oxidation is also influenced by CO2, which obviously is present at higher concentrations under oxy-fuel conditions than under air-fired conditions. Previous work by the authors6 compared the measured and modeled CO concentration profiles in a narrow 100 kW CFB unit operated under both oxy-fuel- and air-fired conditions, and the following expression for CO oxidation was derived:

dCCO ⎛ −E ⎞ α β γ θ ⎟c c c c = K 0,CO exp⎜ ⎝ RT ⎠ CO O2 H2O CO2 dt

(6) 6

In the present work, a global reaction mechanism is used, in which the CO oxidation rate is calculated from eq 6, while the oxidation of other combustible gases is assumed to be mixing-controlled. The following values or K0 and E/R in eq 6 determined by Seddighi K. et al.6 are used: K0 = 5.56 × 107 (m3)0.43 S−1 mol−0.43, E/R = 15145 K, α = 1, β = 0.25, γ = 0.5, and θ = −0.32. The reaction zone is modeled as a plug flow, which is a commonly applied approach in CFB combustion modeling, such as that described by Petersen et al.22 The composition of volatiles is calculated from the fuel analyses given in Tables 1 and 2, according to the volatile composition model 6224


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B C+T

μmix =

(7)

where A = 8.07, B = 1730.06, C = 233.426, and T is the temperature in degrees Celsius. Gas injections to the unit consist of mixtures of gas streams: O2 from the ASU, dry RFG from the condenser, and wet RFG from the exhaust gases. Each of these gas streams has a different composition and temperature, which means that the resulting temperature and temperature of each gas injection are obtained through an energy balance that assumes adiabatic conditions

∑ M mixture,ihiT = Tmixture = ∑ MihiT = Ti

∑ Xi(Mi)1/2

(12)

where μi is the viscosity, Xi is the molar fraction, and Mi is the molecular weight of component i. The temperatures are derived from the experimental values. All of the cases show essentially constant temperature profiles over the furnace height, with temperature deviations of no more than 30 K from the average for the furnace. The stoichiometric ratio of the furnace is based on the incoming fuel and the total incoming O2 from the ASU and air leakage into the system. The air leakage to the system is calculated from the mass balance over the boiler. Although the exact location of the air leakage to the system is not known, it is likely that leakage occurs in the FGR system. Therefore, in both the combustion and fluid dynamic modeling, air leakage is distributed between all of the injection points involving FGR. Initial Model Adjustment. In the present work, the experimental CO concentration data for case AF-C1 are used for fitting the empirical coefficients in eqs 9−11. The experimental data for the remaining four cases are used as validation data, to verify the consistency of the model results using the derived coefficients; i.e., no fitting is used during the modeling of these four cases. The result of the initial fitting yielded a horizontal gas dispersion coefficient of 0.008 m2/s, which was then applied in eq 10, and a horizontal fuel dispersion coefficient of 0.08 m2/s, which was used in eq 11. The initial fitting yielded the following values for the coefficients in eq 9: a = 1.02, b = 0.01, c = 0.12, d = 1.9, and m = 60. The fitted penetrations of the secondary gas injections were 0.40 and 0.15 m in the normal and tangential directions, respectively. The measured vertical profiles of the CO concentration for case AF-C1 and the corresponding fitted profiles are shown in Figure 3.

(8)

where M is the mass and h is the enthalpy of the gases. Efforts directed at modeling gas concentration profiles in CFB combustors have revealed that gas combustion in such units is stringently limited by mixing. CFB modeling typically applies coefficients in the expressions for the reaction rates to account indirectly for the limited mixing, thereby enabling fitting of the modeled gas concentration to the experimental values.22,26,27 Axial gas mixing is strongly coupled to the main solids flow pattern, which yields different fluid-dynamical regions in CFB units,6 with limited gas mixing in the dense bed,6,28 strong mixing close to the site of secondary gas injection and splash zone,6,29 and limited mixing in the upper parts of the freeboard.6,30,31 Gas mixing in the dense bed is governed by the existence of an O2-rich, high-velocity bubble phase, which engages in limited gas exchange with the surrounding emulsion phase, which is due to the fact that the bubble phase is characterized by high-velocity bubbles that enable significant “by-passing” of the gas through the dense bed, thereby decreasing the gas mixing. The emulsion phase contains fuel particles and, consequently, combustible gases,6,28,32 depleting the emulsion phase of oxygen. In the zone located immediately above the dense bed, secondary gas injection and back mixing in the splash zone enhance the axial gas mixing. Finally, in the upper freeboard, axial gas mixing is governed by a dispersion-like process. The axial gas mixing applied in the present work is based on previous modeling and experimental studies on CFB mixing,6 in which the fraction of incoming O2 mixed with the reactants, dS, along a differential height element, dz, is expressed as

dS(z)/dz = (a − b(z − c)2 )n + d

∑ Xiμi (Mi)1/2

(9) 6

where a, b, c, d, and n are empirical parameters that define the extension and shape of the three above-mentioned fluid-dynamical zones, i.e., the dense region, the splash zone, and the upper freeboard. The lateral gas mixing in the riser is described using the lateral gas dispersion coefficient Dg in the transport equation for the gas flux of the different gas species, as given by Pallarès et al.33

∂Gi = Dg ∇2 Gi ∂t

(10)

where Gi is the gas flux of gas species i, t is the time, and Dg is the gas horizontal dispersion coefficient. The horizontal mixing of the fuel and the rate of fuel conversion determine the heterogeneity level of the fuel concentration over the bed.7 The parameter of solid mixing is taken from the model described by Pallarès7 ∂Cs = Dfuel ∇2 Cs + Sfuel ∂t

Figure 3. Measured vertical profiles of the CO concentration and the fitted profile for case AF-C1.

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RESULTS AND DISCUSSION Figures 4−7 compare the modeled and measured O2 and CO concentrations for the four cases used for the validation of the modeling, i.e., OF21-C1, OF25-C1, AF-C2, and OF25-C2. There is reasonable accordance between the experimental and modeled values for both coals (C1 and C2). Both the modeled and experimental data show that the highest CO concentration is in the lower parts of the freeboard. The high levels of CO in the lower part of the furnace reflect the predictions from the modeling; i.e., in this region, the char concentration is high and

(11)

where Cs is the concentration of solids, t is the time, Dfuel is the diffusion coefficient for fuel particles, and Sfuel is the source term. The fuel conversion model used in the present work is explained in detail by Pallarès et al.34 The fuel conversion kinetics is modeled through pseudo-steady-state solving of the heat diffusion into the fuel particle, using the expressions proposed by Palchonok et al.35 The viscosity of the gas at different locations in the modeled domain is calculated on the basis of the temperature and gas composition as36 6225


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Figure 4. Modeled and measured concentrations of O2 and CO at different heights in the furnace for case OF21-C1.


Figure 6. Modeled and measured concentrations of O2 and CO at different heights in the furnace for case AF-C2.

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accumulation of CO observed in all of the figures. Above the secondary gas injection height, O2 provided by the secondary gas increases the rate of combustion; thus, the total flow of CO decreases gradually with the increase in the furnace height. The experimental results show that, for the same inlet O2 concentration, the CO concentration in the lower part of the riser is higher for the OF21-C1 case than for the air-fired condition (Figures 3 and 4). The same conclusion is reached using eq 6, which indicates that a high concentration of CO2 slows CO oxidation. In addition, it should be noted that, in the present experimental setup, the gas velocities for the oxy-fuel combustion cases with coal 1 were 6−7% higher than those for the air-fired case (see Table 3), shortening the residence time for CO oxidation and, thereby, contributing to an increased concentration of CO. The experimental results presented in Figures 4 and 5 show that the oxy-fuel case with 25%vol inlet O2 concentration has a lower peak concentration of CO than the case with 21%vol inlet O2 concentration. The explanation for this phenomenon is that, keeping a similar volumetric flow of the gas injected in the riser (as is the case shown in Figures 4 and 5), a higher O2 inlet concentration involves a higher fuel feeding rate and, consequently, increased production of CO, which is diluted in approximately the same total gas volumetric flow. In addition, a higher inlet O2 concentration yields an increase in the bed temperature, which enhances CO production by char oxidation and increases the rate of CO oxidation. The relationship between the increases in magnitude of these competing effects from an increase in the temperature, i.e., the increase in CO production from char combustion and the increase in the CO oxidation rate from eq 6, determines the shape of the CO concentration curves. The outcomes for the two oxy-fuel cases in Figures 4 and 5 are that the higher temperature in case OF25-C1, as compared to that in case OF21-C1, leads to a lower level of CO. Measured concentrations of CO for cases AF-C1 and OF25C1 shown in Figures 3 and 5 are similar because of the similarity of dense bed temperatures in these two cases, while inlet O2 concentrations do not differ much between these two cases. However, the model overestimates the CO concentration for the OF25-C1 case shown in Figure 5 because the approximated stoichiometric ratio is too low. In other words,

heterogeneous char oxidation produces CO faster than it is consumed, resulting in CO accumulation and an increase in the CO concentration with the height in the furnace. As the char concentration decreases with the height level in the furnace, CO production declines; meanwhile, CO is oxidized, yielding a net decrease in the CO concentration. Therefore, the modeling results show a peak in the CO concentration in the bottom region (Figures 4−7). The measured and modeled concentrations of CO and O2 in Figures 4−7 are plotted for three horizontal locations: (1) close to the fuel feed wall (Y = −0.35 m), (2) at the centerline (Y = 0), and (3) close to the wall opposite the fuel feeding (Y = 0.35 m). Large horizontal variations in the experimental and modeled CO concentrations are observed in the bottom region, mainly because of the limited lateral fuel mixing. As a consequence and as evidenced by the modeled CO concentrations in Figures 4−7, the location and magnitude of the CO peak vary with the horizontal location. The experimental results presented in Figures 4−7 reveal that O2 from the primary gas injection is depleted rapidly in relation to the height in the furnace and is negligible before the site of the secondary gas injection. This is reasonable, because the major proportion of char combustion in the CFB occurs in the dense bed. However, it is noteworthy that the lowest concentration of gas was detected at a height of 2 m, which means that neither the gas profiles nor the progress of combustion in the dense bed and splash zone can be resolved below this level of the furnace. Performing measurements in or near the dense bed of a CFB furnace is challenging, owing to the high concentration of solids, and the results are difficult to interpret because of the dynamic nature of the gas solids flow.11,33 Thus, modeling results can improve the understanding of the fate of combustion and gas concentration fields in the lower part of the furnace. In Figures 4−7, it is clear that the model predicts a rapid decrease in O2 in relation to the height at the fuel feed wall, with all of the O2 being consumed below the site of secondary gas injection. This rapid depletion of O2 near the fuel feed wall is due to the relatively high concentration of fuel, as compared to other locations in the furnace cross-section, which is a direct consequence of the limited lateral fuel mixing. This absence of O2 close to the fuel feed wall also contributes to the 6227


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if enough excess O2 is maintained, the level and peak of the modeled CO concentration in the OF25-C1 case is expected to be similar to those of the AF-C1 case and lower compared to those of the OF21-C1 case because of the considerable lower dense bed temperature in the latter case. There is a considerable difference between measured and modeled CO concentrations at Y = 0 at the height of 1.8 m in Figures 5−7. However, one should note that the gas concentration measurements close to the dense bed and splash zone are largely affected by the bubble eruptions at the bed surface.37 Such bubble eruptions produce strong fluctuations in the upcoming gases38 and, therefore, make the measurements at the locations close to the dense bed and splash zone less reliable, and model comparisons to such data are more difficult.38 As shown in Figures 3−7, for the upper locations in the furnace, the cases that involve higher temperatures, i.e., with higher inlet O2 concentrations, show a lower CO concentration because of the higher rate of CO oxidation. According to the modeling results, a high temperature results in a high CO oxidation rate, as seen in Figure 5 at Y = 0.35 m, in which the CO concentration declines rapidly after the peak, as compared to a case involving a lower temperature (Figure 4). However, the modeling results show that the low excess of O2 in Figure 5 leads to rapid depletion in O2 at Y = 0 and −0.35 m, thus preventing CO oxidation and yielding higher levels of CO than would otherwise be expected for a high-temperature run. Modeling: Extending the Range of Operational Parameters. The model is used to investigate the effects of the inlet O2 concentration over a wide range (up to 100%), with the assumption that heat extraction is sufficient to maintain the furnace temperature at 1117 K. This assumption facilitates a study focused on the effects of inlet O 2 concentrations on combustion, while ensuring constant fluid dynamics and temperature profiles. A previous study by the authors4 investigated the heat balance during oxy-fuel operation of a 287 MW CFB combustor using hard coal as the fuel with and without external heat extraction. It was concluded that heat must also be extracted from the externally recirculated solids flow when the inlet oxygen concentration is >30%. Furthermore, the original external solids flux provided sufficient external heat extraction for an oxygen inlet concentration of up to approximately 40% (maintaining mixing and fluid dynamics). Thus, in the modeling exercise performed below, the modeling results can be seen quantitatively valid for inlet O 2 concentrations for which the fluid dynamics and external solids flux can be maintained. Three major parameters are proposed as indicators of the progress of combustion: (1) the peak CO concentration, (2) the difference in the peak CO concentration over the furnace cross-section, and (3) the CO concentration at the furnace outlet. The peak CO concentration indicates the intensity of combustion. Thus, considering the increase in fuel flow with the increase in the inlet O2 concentration, as investigated in the present work, it is clear that, if the outgoing CO emission level is similar, a higher CO peak value indicates more intense combustion for levels in the furnace higher than the CO concentration peak. The cross-sectional variation in the peak CO concentration indicates the degree of maldistribution of gases during the combustion, and it is important to minimize this effect. The horizontal variation in the CO concentration, ΔCCO, is defined as

ΔCCO = abs(CCO, Y =−0.35 m − CCO, Y = 0.35 m)

(13)

The outlet CO concentration is obviously a measure of the extent to which complete burnout is achieved within the furnace. Figure 8 shows how the CO concentration close to the fuel feed wall (Y = −0.35 m) varies with the inlet O2 concentration.

Figure 8. Modeled vertical profile of CO for different inlet O2 concentrations under conditions similar to those applied in case OF25-C2 at Y = −0.35 m.

The horizontal location of Y = −0.35 m is chosen for plotting the concentration of CO because the measured CO concentrations close to the fuel feed wall in Figures 4−7 are at Y = −0.35 m. The major implication of the increased inlet O2 concentration, in addition to the increase in the peak CO concentration, is that the CO concentration decreases faster with the height above the CO peak height, as shown in Figure 8. It is evident from Figure 8 that the peak CO concentration increases steadily with the increase in the inlet O 2 concentration. The increase in the peak CO concentration occurs because the increase in the inlet O2 concentration leads to an increase in the thermal power and, consequently, in fuel feed (if the unit size and fluidization velocity are kept constant). This increased fuel feed results in the production of greater quantities of combustible gases, which are diluted in approximately the same volumetric gas flow rate, leading to a higher concentration of CO. Figure 9 shows the modeled horizontal difference in peak CO concentrations for various inlet O2 concentrations. It is clear that ΔCCO increases with an increase in the inlet O2 concentration. As mentioned above, an increase in the inlet O2 concentration under a similar fluidization velocity corresponds to a higher fuel feed rate, which, combined with limited fuel mixing, leads to an increase in the difference in the fuel concentration between the vicinity of the fuel feed and the opposite wall. As indicated above, the fuel dispersion coefficient is assumed to be kept constant at 0.08 m2/s, as obtained from the above-mentioned initial model adjustment. Going from OF25 to OF50, as an example, means that the fuel feed is doubled. Thus, assuming a constant fuel dispersion coefficient 6228


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CONCLUSION

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AUTHOR INFORMATION

Article

In the present study, a 3D mathematical model for CFB oxyfuel combustion is developed and validated against unique experimental data from a large CFB oxy-fuel combustor and applied to investigate the effects of operational conditions on the progress of combustion, using CO as an indicator. The modeled concentration profiles of CO and O2 are in reasonable agreement with the experimental results. The modeling and experimental results show that, for similar inlet concentrations of O2, the oxy-fuel-fired condition results in a higher peak CO concentration, as compared to air-fired conditions. Moreover, the modeling shows that both the CO peak concentration and lateral gas maldistribution increase with an increasing O2 inlet concentration, owing to a higher fuel input in the same volumetric flow of gas. Thus, the inlet concentration of O2 in CFB oxy-fuel-fired operation is likely to be limited by the lateral maldistribution of gas species in the furnace. In addition, the modeling results show that an increase in the partial pressure of O2, resulting from an increase in the inlet O2 concentration, leads to a decrease in the level of specific CO (g/MJ) at the furnace outlet.

Figure 9. Modeled horizontal differences in peak CO concentrations for different inlet O2 concentrations under conditions similar to those applied in case OF25-C2.

for the unit, doubling the fuel feed leads to a doubling of the fuel concentration difference between the vicinity of the fuel feed and the opposite wall. This leads, in turn, to more pronounced maldistribution of the gas species. This is of special importance for the heat extraction efficiency of oxy-fuel combustion, because high horizontal differences in the concentrations of combustible gases lead to significant variations in the heat load close to the fuel feed wall, as compared to the opposite wall. Figure 10 plots the modeled specific emissions (g/MJ) of CO at the furnace outlet for different inlet O2 concentrations. It

Corresponding Author

*E-mail: [email protected]. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS Financial support from the Swedish Energy Agency and financial support and experimental data provided by Metso Power OY are acknowledged. Ville Ylä-Outinen is acknowledged for processing the experimental data. The authors thank Dr. Klas Andersson for valuable discussions on the modeling and Marko Palonen for discussions regarding the experimental procedure.

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REFERENCES

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Figure 10. Modeled specific emissions of CO for different inlet O2 concentrations under conditions similar to those applied in case OF25C2.

is evident that an increase in the inlet O2 concentration leads to a reduction in the specific CO emissions from the furnace. This decrease in specific outlet CO confirms that an increase in the O2 inlet concentration yields an increase in the intensity of combustion (cf. CO concentration peak in Figure 8). However, at O2 inlet concentrations exceeding 70%, there is almost no further reduction in specific emission of CO, indicating that the mixing and furnace height start to limit the burnout of the fuel. 6229


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