Fluidized-Bed Combustion of Single Coal Char Particles - American

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Fluidized-Bed Combustion of Single Coal Char Particles: An Analysis of the Burning Rate and of the Primary CO/CO2 Ratio Fabrizio Scala* Istituto di Ricerche sulla Combustione - CNR P.le Tecchio, 80 80125 - Napoli, Italy ABSTRACT: The fluidized-bed combustion of large coal char particles was studied with a focus on the burning rate and the primary CO/CO2 ratio at the particle surface. To this end, single particle laboratory-scale experiments were carried out using a low-attrition and high-reactivity coal char. A previously proposed indirect experimental technique was applied, which was modified to cover a wider range of experimental conditions. The experiments were conducted at different bed temperatures (800-900 °C), fluidization velocities (0.3 and 0.53 m/s), inlet oxygen concentrations (0.5%-8.0%), and inert bed particle size ranges (100-212, 500-600, and 900-1000 μm). The actual sphericity and temperature of the particle were also measured during selected experiments and used in the calculations. Results showed that, under the experimental conditions investigated, carbon was completely oxidized to CO2 within the particle boundary layer. The experiments confirmed that the char particles burned under boundary layer diffusion control in the temperature range of 800-900 °C. It was demonstrated that single particle burning rate experiments with a high-reactivity char can be used to estimate the particle Sherwood number in fluidized beds, but only if char attrition can be assumed to be negligible.

’ INTRODUCTION Combustion of coal char particles is generally controlled by a combination of intrinsic kinetics and intraparticle and external diffusion of oxygen. However, it is typically assumed that large char particles burn in fluidized-bed (FB) combustors under the external (or boundary layer) diffusion control regime. Under these conditions, the particle burning rate is controlled by the rate of oxygen transfer through the particle boundary layer, and by the size and shape of the particles (i.e., by the external surface). This assumption has been often used to justify the use of char combustion experiments to estimate the particle mass-transfer coefficient in fluidized beds. As thoroughly reviewed by Scala,1 these char combustion experiments have obtained only limited success, often leading to uncertain and controversial results. Apart from possible inaccuracies in measurements and in parameter evaluations, the most severe limitations in the use of this technique are as follows:1 (i) not accounting for the possible influence of intrinsic kinetics and intraparticle diffusion on the combustion rate (especially for low-reactivity fuels); (ii) the accuracy of the char particle temperature; (iii) the assumption of the CO/CO2 primary ratio and of the CO oxidation location; (iv) neglect of how the influence of attrition and fragmentation, the change of carbon particle size with time, and nonspherical carbon particles may affect the results. Among the factors listed above, one that has attracted much attention is the stoichiometry of the carbon oxidation. In fact, carbon can be oxidized either to carbon monoxide or to carbon dioxide. The proportion to which carbon converts to either of the two products at the particle surface (the so-called “primary CO/ CO2 ratio”) has been the subject of an extensive number of research papers, and it is still open to debate.2 A further complication derives from the possibility of CO to be oxidized to CO2 in the gas phase at or near the carbon surface. From a practical point of view, it is important to quantify the CO/CO2 ratio at the exit of the particle boundary layer, since CO oxidation that occurs r 2011 American Chemical Society

within this layer can be considered equivalent to the primary CO2 product. In fact, both the oxygen transfer rate and the heating rate of the particle depend on the relative quantities of CO and CO2 that are produced close to the particle surface. Measurement of the product CO/CO2 ratio is very difficult since carbon monoxide can be further oxidized in the reactor far from the burning particle, even if the sand bed in a fluidized combustor is reported to provide a CO oxidation quenching effect.3 The result of the above complexities is that very few works (as reviewed by Scala4) attempted to measure the product CO/CO2 ratio for large char particles burning in a fluidized bed. Most of the techniques used in these works were based on many assumptions and approximations and produced results with a considerable degree of uncertainty. Recently, a new indirect technique was proposed that is relatively simple and accurate.4 It is based on the measurement of the burning rate (following the CO and CO2 concentrations at the reactor outlet) of a single char particle under low oxygen concentration conditions. Under these conditions, two advantages are obtained. First, the boundary-layer mass-transfer coefficient can be calculated without the need to account for high-mass-transfer-rate and/or possible nonequimolarcounterdiffusion corrections. Second, heat effects are very limited, and the char particle temperature can be assumed to be approximately equal to that of the bed. Comparison of the measured particle burning rate with an independently estimated particle Sherwood number, with the aid of a simple fluidizedbed combustor model, allowed the author to find the product CO/ CO2 ratio, without having to estimate the homogeneous CO oxidation rate in the reactor. Following a recent work,1 the relevant particle Sherwood number was calculated by means of a correlation based on experimental data gathered in the same experimental apparatus as that used in the present work. In that Received: September 2, 2010 Revised: January 12, 2011 Published: February 03, 2011 1051

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Table 1. Properties of Snibston Coal parameter

value Proximate Analysis, % (As-Received)

moisture

14.6

ash

4.0

volatile matter

35.2

fixed carbon

46.2

Ultimate Analysis, % (Dry Basis and Ash-Free Basis)

Figure 1. Experimental apparatus. Legend: (1) gas preheating section; (2) electrical furnaces; (3) ceramic insulator; (4) gas distributor; (5) thermocouple; (6) fluidization column; (7) gas suction probe; (8) stack; (9) cellulose filter; (10) membrane pump; (11) gas analyzers; (12) personal computer; (13) manometer; and (14) digital mass flowmeters.

work, the mass-transfer coefficient around a few freely moving Pt catalyst spheres in the bed was measured by following the CO oxidation reaction at 450 °C at different fluidization velocities, catalyst sphere sizes, and inert bed particle sizes. Experimental data were excellently fitted by the following correlation:
 Remf 1=2 1=3 Sc ð1Þ Sh ¼ 2:0εmf þ 0:70 εmf where Sh is the particle Sherwood number, εmf the bed voidage under minimum fluidization conditions, Remf the particle Reynolds number calculated at the minimum fluidization velocity, and Sc the Schmidt number. The objective of this work was to thoroughly characterize the burning rate and the primary CO/CO2 ratio during the fluidizedbed combustion of large coal char particles and validate the suitability of single particle combustion experiments to estimate the mass-transfer coefficient in fluidized beds. An extensive experimental campaign was performed using the technique proposed by Scala,4 which was improved and modified to explore a wider range of experimental conditions. In particular, inlet oxygen concentrations above 2% were investigated. Also, this work examined the influence of the fluidization velocity, inert bed particle size, and bed temperature. The actual sphericity and temperature of the particle were measured during the experiments and taken into account in the data analysis.

’ EXPERIMENTAL SECTION Apparatus. A circular stainless steel (Sandvik 253 MA) atmospheric bubbling fluidized-bed reactor with an inner diameter (ID) of 40 mm and a height of 1 m was used for the experiments (see Figure 1). The fluidization gas distributor was a 2-mm-thick perforated plate with 55 holes 0.5 mm in diameter disposed in a triangular pitch. The grid region had no influence on the particle mass-transfer coefficient, as checked previously.1 A 0.6-m-high stainless steel column, containing several

carbon

81.3

hydrogen

5.3

nitrogen oxygen

1.6 10.8

sulfur

1.0

char density

940 kg/m3

steel nets for gas preheating and mixing, was placed under the distributor. The fluidization column and the preheating section were heated by two semicylindrical 2.2-kW electric furnaces. The temperature of the bed, measured by means of a chromel-alumel thermocouple placed 40 mm above the distributor, was kept constant by a proportional-integraldifferential (PID) controller. Temperature variations during the runs were always within (1 °C of the setpoint. A stainless steel circular basket could be inserted from the top to retrieve char particles from the bed. A basket mesh of 2 mm was used, so that bed material could easily pass through the net openings. Gases were fed into the column via two high-precision digital mass flowmeters/controllers (accuracy of (1% full scale). Each flowmeter/ controller was frequently calibrated with a bubble flowmeter. Gases were supplied from two cylinders, containing nitrogen and oxygen, respectively. The top section of the fluidization column was left open to the atmosphere and the exit gas was sucked by a hood. A stainless steel probe was inserted from the top of the column, to convey a known fraction (0.06 m3/h) of the exit gas directly into the gas analyzers. A high-efficiency cellulose filter was inserted in the line, to avoid dust entrainment into the analyzers. The probe, 2 mm ID, was positioned 0.6 m above the distributor, approximately at the axis of the column. The absence of any gas leakage and/or suction to/from the surrounding environment in the sampling probe and line was carefully checked. A paramagnetic analyzer and two NDIR analyzers (accuracy of (1% full scale) were used for online measurement of O2, CO, and CO2 concentrations, respectively, in the exhaust gases. The measuring ranges of the CO and CO2 analyzers used in the experiments were 1000 and 2000 ppmv, respectively, to minimize the measurement error. The analyzers were frequently calibrated with specific calibration cylinders containing a gas mixture of N2 and either CO (1000 ppmv) or CO2 (2000 ppmv). Data from the analyzers were logged and further processed on a personal computer (PC). Materials. Snibston bituminous coal was selected as the test fuel. Char from this coal has been reported to have the lowest propensity to attrition and fragmentation among the coals studied in previous investigations.5 Table 1 reports the fuel properties. The fuel particles were first devolatilized by dropping them in the fluidized bed with nitrogen at 850 °C. After ∼5 min, char particles were retrieved from the bed and machined into approximately spherical particles with an average size of ∼7 mm (see Figure 2A). The particles were then preprocessed in air for 8 h in the fluidized bed at ambient temperature with a fluidization velocity of 0.3 m/s to smoothen the particle surface and remove asperities. Following earlier work,6 the sphericity of the particles was estimated as the ratio between the external surface of a sphere having the same volume as the particle and the actual external surface of the particle. The 1052

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Figure 2. (A) Typical coal char particles used for the experiments. (B and C) Char particle with a thermocouple inside it. sphericity was consistently larger than 0.85. Char density was estimated by working several char particles to the shape of regular parallelepipeds and then measuring their weight and volume. The bed material consisted of 180 g of quartz sand, corresponding to an unexpanded bed height of 0.1 m. Sand was double-sieved in the three following particle size ranges: 100-212, 500-600, and 900-1000 μm. The minimum fluidization velocity was 0.01, 0.13, and 0.36 m/s, respectively. The particle density of the quartz sand was 2560 kg/m3, and the bed voidage at minimum fluidization was 0.44. Procedures. All tests performed in this experimental campaign had an inlet oxygen concentration lower than 10% (v/v) to simulate relevant local O2 concentrations near a burning particle in a full-scale fluidizedbed combustor. This is based on experimental evidence reported by Leckner7 that shows very limited oxygen penetration in the bottom dense zone of an air-fired CFB combustor, where most of the char particles burn. It has been demonstrated that the particle temperature is approximately equal to that of the bed if oxygen concentrations are lower than 2% v/v.4 In this work, some tests were performed with a higher O2 concentration. The particle temperature was measured in some of these tests by inserting a thermocouple inside the particle during combustion. A long (1.5 m) and thin (250 μm outer diameter (OD)) thermocouple was fitted and cemented inside a hole drilled through the particle (see Figures 2B and 2C). The thermocouple tip was placed approximately at the center of the particle. The thermocouple was sufficiently thin and flexible, such that it did not appreciably restrict the free motion of the particle inside the bed.8 This was confirmed by repeating the same experiment with similar char particles with and without the thermocouple; the carbon burning rate, as a function of the particle equivalent diameter, was practically the same for both runs. The CO and CO2 traces measured by the analyzers were corrected for the influence of the sampling line residence time distribution (RTD). The RTD of the sampling line was measured by injecting step changes of either CO or CO2 in the sampling line and recording the analyzer measurements. It was found that, for both CO and CO2, the RTD of the sampling line could be assumed to be the sum of that for a plug-flow reactor and for a continuously stirred reactor, with characteristic times of 18 and 4 s, respectively. Possible segregation of the char particles, with respect to the bed sand, was also considered. Visual inspection of the bed surface from the top of the freeboard during experiments revealed that the char particles appeared on the surface only for limited time periods, indicating the absence of significant segregation to the upper bed section. This result is consistent with Nienow and Chiba,9 who reported that segregation is not expected for the particles and operating conditions used in the present work. Given an average bulk density of the sand bed of ∼1430 kg/m3, the density ratio between the char particles and the bed was 0.66 (see Table 1). Measurement of the Carbon Burning Rate. Single particle combustion experiments were conducted in the fluidized bed operated at 800, 850, or 900 °C at atmospheric pressure with an inlet gas mixture of N2 and O2. The inlet oxygen concentration was fixed to 0.5%, 1.0%, 2.0%, 4.5%, or 8.0%. The fluidization velocity was either 0.3 or 0.53 m/s, corresponding to bubbling/slugging conditions.1

Each test consisted of the injection of one char particle into the fluidized bed, followed by continuous measurement of the CO and CO2 concentrations at the outlet. In selected experiments, at definite time intervals, the reaction was quenched by switching the gas feed to N2. The particle was slowly retrieved from the bed by means of the steel basket, and, after cooling, it was weighed and photographed, and its size was accurately measured with a micrometer in the three principal directions. The particle then was reinjected in the bed and the gas feed was switched again to the oxygen-containing mixture. This procedure was repeated several times until complete burnoff. For each run, mass balance closure on carbon was within (3%. The oxygen conversion degree was always below 5%. These conditions approach differential conditions, with respect to O2 concentration, and thus minimize possible errors that are due to bed fluid-dynamics assumptions (see the next section). Data Analysis. Scala4 reported a method to analyze char combustion experimental data using a simple reactor model based on the two-phase fluidization theory.10 Following experimental evidence,1 the assumption was made that the char particle resides and carbon oxidation occurs only in the dense phase of the fluidized bed. The bubble-emulsion phase mass transfer index, X, which represents the number of times a bubble exchanges its volume during its rise through the bed, was estimated10,11 to be X > 6 (except for the 100212 μm sand bed, where X ≈ 3). These X values indicate almost perfect mixing in the reactor, and, therefore, the fluidized bed was modeled as a perfectly mixed reactor. Increased accuracy of the reactor model is not worth the added complexity, in the light of the differential nature of the experiments. In fact, negligible differences were found in results obtained with the present model, compared to those obtained with a morecomplex two-phase model (such as that reported by Scala4). Carbon consumption by oxygen in the particle can be expressed by the following general stoichiometry:
 n ð2Þ ð1 þ nÞC þ 1 þ O2 f nCO þ CO2 2 where n is the CO/CO2 molar product ratio. If we now introduce the quantity λ = (1 þ n)/[1 þ (n/2)], representing the molar ratio between the consumed carbon and oxygen, the carbon and the oxygen consumption rates in the particle are related by RC = λRO2. The molar production rate of CO and CO2 combined is RCOX = -RC. Using the sum COX in the calculations, instead of the single species CO and CO2, is advantageous, because COX is not dependent on the homogeneous gas-phase CO oxidation reaction that transforms CO to CO2. Under external diffusion control (assuming a vanishing O2 concentration at the surface of the char particle) and pseudo-stationary combustion, the total COX production rate in the fluidized bed for a carbon particle can be expressed as RCOX ¼ λAp kg Cbed O2

ð3Þ

where AP represents the total external surface of the char particle (AP = πdP2/ψ), dp is its actual equivalent diameter (the diameter of a sphere having the same volume as the particle), and ψ is the particle sphericity. 1053

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kg is the boundary-layer mass-transfer coefficient and CO2bed is the oxygen concentration in the bed. It is assumed here that that the particle shape does not influence the particle mass-transfer coefficient appreciably, with respect to that relative to a sphere with the same equivalent diameter. This assumption is reasonable, because all particles used in the experiments had a sphericity larger than 0.85, and the boundary layer formed around the particle should not significantly differ from that formed around a perfect sphere. By solving the material balance on COX in the bed under the above assumptions, and using eq 3, the following expression is derived: Cbed COX ¼

λAp kg Cbed O2 Ar U

ð4Þ

bed where CCO is the COX concentration in the bed, U the total fluiX dization gas velocity, and Ar the reactor cross section. Assuming perfect mixing, the outlet COX concentration is simply

bed Cout COX ¼ CCOX

ð5Þ

since COX is not influenced by CO homogeneous gas-phase oxidation in the freeboard, as noted earlier. If we assume that homogeneous oxidation of CO is negligible within the dense phase (as suggested in the literature3), then the COX concentration in the bed can be linked to that of oxygen by in bed Cbed COX ¼ λðCO2 - CO2 Þ

ð6Þ

It must be underlined that the above assumption is not as severe as it might appear, since the present experiments are carried out under differential conditions, with respect to oxygen, so that the bed oxygen concentration would be hardly influenced by any homogeneous CO oxidation. Rearrangement and combination of eqs 4, 5, and 6, finally gives ! Cout COX Ar U þ1 ð7Þ λ ¼ in CO2 Ap kg By measuring the sum of CO2 and CO concentrations at the reactor outlet, the use of eq 7 gives the value of λ as a function of the char particle size. The mass-transfer coefficient in eq 7 can be expressed in terms of the particle Sherwood number as kg = ShDO2/dp, where DO2 is the oxygen molecular diffusion coefficient in the boundary layer. The actual mass of carbon in the particle at a generic time t* during the run is given by Z t  Cout ð8Þ wðt Þ ¼ w0 COX QMC dt 0

where w0 is the initial carbon mass, Q the gas flow rate, and MC the carbon molecular weight. The actual particle equivalent diameter at time t* is !1=3 6wðt Þ  ð9Þ dp ðt Þ ¼ πFC where FC is the particle carbon density (assumed to be constant during the run). The O2 molecular diffusivity in the gaseous mixture used in the experiments was evaluated to be given as DO2 = 2.11  10-4 m2/s at 850 °C and was corrected for the influence of temperature for tests at 800 and 900 °C. The particle Sherwood number, as a function of the char particle equivalent diameter, was estimated using the correlation proposed by Scala for spherical particles (eq 1).1 To use eq 7, the actual particle sphericity and temperature must be either known or estimated. Conversely, if the value of λ is known or assumed, eq 7 can be used to estimate the particle mass-transfer coefficient (or Sherwood number). If

Figure 3. Typical CO and CO2 measured outlet profiles during two single char particle combustion tests: (A) no fragmentation and (B) a fragmentation event occurred. For both tests, the conditions are as follows: T = 850 °C; U = 0.3 m/s; ds = 500-600 μm; [O2] = 2.0%. we insert the definition of the mass-transfer coefficient and the particle external surface in eq 7 and rearrange, we obtain ! Cout ψAr U COX Sh ¼ ð10Þ out πd λCin C p DO2 COX O2 Note that, for the experiments conducted at higher oxygen concentration, the high-mass-transfer-rate and/or nonequimolar-counterdiffusion corrections to the mass-transfer coefficient might be non-negligible. In a recent paper, the mass-transfer coefficient around a burning carbon particle in an atmosphere of O2, N2, CO2, CO, and H2O was considered and the complete set of Stefan-Maxwell equations was analytically solved under the assumption of negligible homogeneous reaction in the boundary layer.12 The results were given in terms of a correction factor to the equimolar-counterdiffusion mass-transfer coefficient. This correction factor was evaluated for all experiments reported in this work. The corrected mass-transfer coefficient differed from that relative to equimolar-counterdiffusion conditions by 2%, the error due to the assumption that the char particle has the same temperature as that of the bed is larger than 10 °C. Based on these results, in the following analysis, the particle temperature was assumed to be the same as that of the bed only for the runs conducted at inlet oxygen concentrations of 0.5% and 1%. For the runs at inlet oxygen concentrations of 2%, 4.5%, and 8%, the gas properties (density and viscosity) and the oxygen diffusion coefficient were calculated at the film temperature (i.e., at the arithmetic mean between the bed and the particle temperature). The particle temperature values measured in the above experiments were used in the calculations. For particle conversion degrees larger than that at the thermocouple detachment point, the particle temperature was assumed to be equal to the maximum measured value. The char particle temperature was assumed to be uniform along its radius during combustion, neglecting possible temperature gradients within the particle. This assumption is based on evaluation of the Prater number,15 which expresses the relative magnitude of the rate of heat generation by reaction and the rate of heat transport inside the particle. For the particle sizes considered in the present experiments, the Prater number always resulted in a value much lower than 0.1, justifying the above assumption. Estimation of the CO/CO2 Ratio from Single Particle Combustion Experiments. The sum of the measured CO and CO2 outlet concentrations was used to calculate the value of λ by eq 7, as a function of time. The actual particle equivalent diameter at each time was calculated by integrating the CO and CO2 curves, according to eqs 8 and 9. The Sherwood number (Sh), as a function of the equivalent diameter, was calculated using eq 1. Particle actual temperature and sphericity were taken into account as detailed before. Figures 7 and 8 report the results of the analysis in the form of λ as a function of the particle equivalent diameter. Only data for dp g 2 mm are reported, to ensure that the particle burning rate is under external diffusion control. For some experiments, the reported data stop at particle diameters of >2. In these experiments, as described earlier, a fragmentation event occurred (see Figure 3B), and measurements following this event were not used for the burning rate analysis. Two horizontal dashed lines are also drawn in the figures, which indicate the limit values that λ can assume. A value of λ = 1 corresponds to complete carbon 1056

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Figure 7. Experimental values of λ, as a function of the char particle diameter, for different inlet oxygen concentrations. Conditions: T = 850 °C; U = 0.3 m/s; and ds = 500-600 μm.

Figure 8. Experimental values of λ, as a function of the char particle diameter, for different operating conditions. ([O2] = 2.0%.)

conversion to CO2 within the boundary layer, and a value of λ = 2 corresponds to complete carbon conversion to CO. Figure 7 reports the measured values of λ, as a function of the char particle diameter, for different inlet oxygen concentrations. Figure 8 shows the effect of bed temperature, inert particle size, and fluidization velocity. Analysis of Figures 7 and 8 highlights the following results: 1 For all the operating conditions investigated, the value of λ closely approaches the lower dashed line (λ = 1). This indicates that CO2 is the prevailing species exiting the boundary layer of the burning char particle. A change of bed temperature between 800 °C and 900 °C does not influence the measured CO/CO2 ratio. 2 In some of the experiments, λ > 1 was measured for the largest char particle sizes, as was also noted by Scala.4 These events are not correlated with a particular char particle size. They appear to be slightly more significant at the lowest oxygen concentration ([O2] = 0.5%) and at the smallest inert particle size (100-212 μm). These observations suggest that the apparent values λ > 1 are not linked to an

increase of primary CO production, but rather to the unavoidable influence of particle attrition on char consumption. In fact, the attrition rate is known to assume the largest value at the beginning of the experiments, as a consequence of a rounding-off mechanism, decaying thereafter with time.5 In addition, the attrition rate is likely to give the most significant relative contribution to char consumption in those experiments where the combustion rate is lower, i.e., at low oxygen concentration and small inert particle size. Attrition generates char fines that are partly burned during their residence time in the bed before elutriation. Therefore, the net effect of attrition is an increase of the apparent burning rate of the particle, which reflects an increase of the measured λ (see eq 7). Estimation of the Particle Sherwood Number from Single Particle Combustion Experiments. If the value of λ is known or assumed, eq 10 can be used to estimate the particle Sherwood number from the single particle combustion experiments. On the basis of the previous results, a value λ = 1 was assumed in the present calculations. The sum of the measured CO and CO2 outlet 1057

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Figure 9. Experimental Sherwood number (assuming λ = 1), as a function of the char particle diameter for different inlet oxygen concentrations. T = 850 °C; U = 0.3 m/s; ds = 500-600 μm. The black curve is the theoretical Sherwood number, calculated with eq 1.

Figure 10. Experimental Sherwood number (assuming λ = 1), as a function of the char particle diameter for different operating conditions. [O2] = 2.0%. The black curve is the theoretical Sherwood number, calculated with eq 1.

concentrations was used to calculate the value of Sh for each experiment. The actual particle equivalent diameter, as a function of time, was calculated according to eqs 8 and 9. Particle actual temperature and sphericity were taken into account in the calculation. Figures 9 and 10 report the results of the analysis in the form of Sh, as a function of the particle equivalent diameter, for the different operating conditions investigated. As previously noted, only data for dp g 2 mm are reported to ensure that the particle burning rate is under external diffusion control. The Sherwood number predicted by eq 1 for the same equivalent diameters is also reported in the figures for comparison. The following results can be noted from this comparison: (1) The measured Sherwood number compares excellently with the prediction by eq 1 for all the operating conditions. It is worth noting that no adjustable parameter is present in the calculation, since the values of all parameters were calculated from independent tests.

(2) The value λ = 1 is confirmed to be a very good estimate for coal char combustion in a fluidized bed under all operating conditions tested. (3) The assumption of boundary-layer diffusion control for char combustion in a fluidized bed appears to be fully justified (at least for a high-reactivity fuel such as Snibston coal). Note that if the intrinsic kinetics and/or intraparticle diffusion contributions to the overall burning rate were significant, the measured apparent Sh values would have shifted below the theoretical Sh line, since the boundary layer diffusion control represents the maximum possible particle burning rate at each particle size. (4) The excellent comparison of the present experimental data with values predicted by eq 1, which was based on data obtained in a different reaction system and at a different temperature,1 further confirms the accuracy of this mass-transfer correlation. 1058

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(5) Single particle burning rate experiments can be used to estimate the particle Sherwood number in a fluidized bed only if attrition can be assumed to be negligible. Significant carbon attrition during the experiments would lead to a fictitious increase of the measured apparent particle Sherwood number.6

’ AUTHOR INFORMATION Corresponding Author

*Tel.: þ39 081 7682969. Fax: þ39 081 5936936. E-mail: [email protected].

’ ACKNOWLEDGMENT The author is indebted to Ms. F. Raganati for her support in performing the experimental tests. The support of Mr. S. Russo for making the photograph of the char particles is gratefully acknowledged. ’ REFERENCES (1) Scala, F. Chem. Eng. Sci. 2007, 62, 4159–4176. (2) Agarwal, P. K.; La Nauze, R. D. Chem. Eng. Res. Des. 1989, 67, 457–480. (3) Hayhurst, A. N.; Tucker, R. F. Combust. Flame 1990, 79, 175– 189. (4) Scala, F. Proc. Combust. Inst. 2009, 32, 2021–2027. (5) Chirone, R.; Massimilla, L.; Salatino, P. Prog. Energy Combust. Sci. 1991, 17, 297–326. (6) Scala, F.; Chirone, R.; Salatino, P. Energy Fuels 2006, 20, 91–102. (7) Leckner, B. Prog. Energy Combust. Sci. 1998, 24, 31–61. (8) Hayhurst, A. N.; Parmar, M. S. Combust. Flame 2002, 130, 361– 375. (9) Nienow, A. W.; Chiba, T. Fluidization of dissimilar materials. In Fluidization, Second Ed.; Davidson, J. F., Clift, R., Harrison, D., Eds.; Academic Press: London, 1985; pp 357-382. (10) Davidson, J. F.; Harrison, D. Fluidised Particles; Cambridge University Press: Cambridge, U.K., 1963. (11) Sit, S. P.; Grace, J. R. Chem. Eng. Sci. 1981, 36, 327–335. (12) Scala, F. Combust. Flame 2010, 157, 137–142. (13) Lin, J. L.; Keener, H. M.; Essenhigh, R. H. Combust. Flame 1995, 100, 271–282. (14) Clift, R.; Grace, J. R.; Weber, M. E. Bubbles, Drops, and Particles; Academic Press: New York, 1978. (15) Scala, F.; Salatino, P. Chem. Eng. Sci. 2002, 57, 1175–1196.

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