Studies of the Flash Carbonization Process. 1. Propagation of the

Ind. Eng. Chem. Res. 2006, 45, 585-599

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Studies of the Flash Carbonization Process. 1. Propagation of the Flaming Pyrolysis Reaction and Performance of a Catalytic Afterburner Teppei Nunoura, Samuel R. Wade, Jared P. Bourke,† and Michael Jerry Antal, Jr.* Hawaii Natural Energy Institute, School of Ocean and Earth Science and Technology, UniVersity of Hawaii at Manoa, Honolulu, Hawaii

The flash carbonization process quickly and efficiently produces biocarbon (i.e., charcoal) from biomass. This process involves the ignition of a flash fire at elevated pressure in a packed bed of biomass. The fire moves upward through the bed, against the downward flow of air, triggering the transformation of biomass into gas at elevated pressure and charcoal with fixed-carbon yields that can reach the thermochemical equilibrium “limit” within 20 min of reaction time. Research described in this paper had two foci. The first concerned the propagation of the flaming pyrolysis reaction. The spatial distribution of the products’ properties confirmed that the carbonization reaction propagated upward from the bottom of the bed to its top, and revealed that the charcoal products were further carbonized and combusted from the top downward after the entire bed was converted to charcoal. In addition, we found that higher pressures increased the upward speed of the flame through the bed. The second focus concerned the use of an afterburner that operated at elevated pressures, and used a commercially available catalytic monolith to reduce emissions. When the afterburner was supplied with sufficient secondary air, virtually all carbon monoxide emissions were eliminated. 1. Introductory Section Biomass charcoal finds applications in an extraordinary number of high-value markets. In addition to its use as a barbeque fuel, charcoal is an important metallurgical reductant1 and is used to manufacture activated carbons. Charcoal is also widely used as a soil amendment2 and is sold as a digestive aid in health stores. Very low-priced charcoal can replace coal as a boiler fuel. A recent comprehensive review3 of the art, science, and technology of charcoal production is available. Ongoing research at the University of Hawaii (UH) has led to the development of a new flash carbonization (FC) process that quickly and efficiently produces biocarbon (i.e., charcoal) from biomass.4 This process involves the ignition of a flash fire at elevated pressure in a packed bed of biomass. Because of the elevated pressure, the fire quickly spreads through the bed, triggering the transformation of biomass to biocarbon. Fixed-carbon yields can attain the thermochemical equilibrium limit after reaction times of 20 min. Biomass feedstocks have included woods (e.g., leucaena and oak), agricultural byproducts (e.g., macadamia nutshells and corncobs), and wet green wastes (wood sawdust and Christmas tree chips). Research described in this paper had two principal foci that were not addressed in our earlier work.4 The first concerned the propagation of the flaming pyrolysis reaction. Results confirmed the upward movement of the flaming pyrolysis reaction, followed by downward combustion and simultaneous carbonization of the hot charcoal. Insights were gained into the movement and disposition of mineral matter (ash) during the FC process, as well as the dependence of the flame speed on the reactor’s pressure. The second focus concerned the performance of an afterburner that operated at elevated pressures and used a commercially available catalytic monolith to reduce emissions. In addition to destroying carbon monoxide, the afterburner was able to combust undetected organic compounds that were present in the gaseous effluent. †

Present affiliation: The University of Waikato, New Zealand. * To whom correspondence should be addressed. E-mail address: [email protected].

Figure 1. Schematic diagram of the laboratory-scale flash carbonization reactor and the catalytic afterburner. Legend is as follows: A, air accumulator; C, compressor; CI, canister with insulation; CM, catalyst monolith; DDV, downdraft valve; FI, funnel with insulation; GSP, gas sampling port; H, electric heater; IV, isolation valve; MMV, micrometer valve; O, oxygen analyzer; PG, pressure gauge; PRV, pressure relief valve; PT, pressure transducer; PV, pressure vessel; R, regulator; RM, rotometer; SRD, safety rupture disk; T, air tank; TC, thermocouple; UDV, updraft valve; WT, water trap.

2. Apparatus and Experimental Procedures A schematic diagram of the apparatus used in this study is shown in Figure 1, and the reaction conditions are listed in Table 1. In this section, we detail the apparatus and describe the procedures for both the experiment and the analysis. 2.1. Flash Carbonization Reactor. The carbonization procedures applied in this work were similar to those reported previously.4 However, we made two modifications: we used a newly built canister that had a larger aspect ratio than the

10.1021/ie050854y CCC: $33.50 © 2006 American Chemical Society Published on Web 12/13/2005

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Table 1. Reaction Conditions of the Flash Carbonization Experiments

a Date of experiment. Expressed as yy/mm/dd. b Moisture content (wet basis) determined as per explained in the Experimental Section. c Time during which the biomass ignition heater was turned on. d Time from the activation of ignition heater to the cessation of primary air delivery. e Average throughout the run. f Contained in a metal beaker and put on the top of the corncob bed. g Run was conducted without a catalytic afterburner. h Put on the top of the biomass bed. i Not analyzed.

Table 2. Elemental Analyses of Feedstocks

a

Ash content (dry basis) determined by ASTM E 1755-95. b Dry-basis values measured by Huffman Laboratories, Inc., USA. c Not analyzed.

previous one, and we used a single electric heater with a spiral shape to ignite the biomass feedstock. These new features will be described below. Corncobs obtained from Waimanalo, Oahu and macadamia nutshells (macshells) from the Big Island of Hawaii were used as the feed materials. Note that corncobs as well as corn stover are the most abundant agricultural residues in the United States. We also remark that macshells are one of the typical, plentiful agricultural wastes in the State of Hawaii and they are usually burned in boilers or sent to landfill. The shape of the corncobs was roughly cylindrical and their dimensions were ∼3 cm in diameter and 10-20 cm in length. Sometimes, the corncobs were cut into half-lengths or even shorter to facilitate their loading into the canister. The shape of the macshells was approximately hemispherical, and their diameters were in the range of 2-3 cm. Sometimes, a small amount of other types of feedstocks (such as pistachio nutshells and orange peel) was also added to the top of the canister for a special interest. The properties of the feedstocks are summarized in Table 2. A measured amount of corncobs (∼1 kg) or macshells (∼3 kg) was loaded into a newly built, cylindrical stainless-steel canister with an inner diameter of 7.6 cm and a length of 95 cm. In addition, a metal screen and a weight were placed on the top of the bed, to prevent the bridging of feedstocks inside the canister. At the same time as the loading, some additional samples of the feedstock were taken for analysis of their moisture content. The canister was subsequently placed in a pressure vessel, and the entire system was pressurized by air to the reaction pressure (0.791-2.86 MPa). In the pressure vessel, there was an electric heater (ARi Industries, model BXX-19B-50-11T), whose heating element was located between the grated bottom of the canister and the

bottom flange of the pressure vessel, as shown in Figure 1. To heat the circular bottom of the canister, the heating element was bent to form a spiral shape with a diameter of ∼9 cm. After the pressure vessel reached the reaction pressure, the spiralshaped heater, which we call the ignition heater, was turned on to ignite the biomass at the bottom of the bed. The ignition heater consumes 1.0 kW when operated at our usual setting. The ignition heater was kept on for 7-18 min and then was turned off. Two or three minutes after the ignition heater was turned on, primary air was fed into the canister within the pressure vessel. This primary air was delivered from an air accumulator (9.33 L of inner volume) that was pressurized to ∼30 MPa. A micrometer valve controlled the delivery of primary air, and the flow rate was monitored by both rotometer (Matheson, FM1050) readings and the rate of pressure decrease in the accumulator. Especially, the temporal change of pressure in the accumulator was continuously recorded by means of a data logging system (National Instruments), and these data were interpreted later to obtain exact mass flow rates of air. Primary air entered the canister from its top, passed downward through the bed of biomass, and subsequently flowed to the catalytic afterburner (see below). This downward airflow in the bed of biomass feedstock caused the upward movement of the flame front, which triggered the conversion of biomass to charcoal. After sufficient air was fed into the canister, the air delivery was halted. The reactor was then depressurized and cooled, and finally, the products of carbonization were taken out for the analysis. 2.2. Catalytic Afterburner. The gas stream leaving the carbonization reactor was fed to the catalytic afterburner at the same operating pressure as the FC reactor. A cylindrical,

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honeycomb-type catalyst monolith (Condar Company, model CC-001) was used in this study. The diameter and height of the monolith are 14.3 and 3.81 cm, respectively. The monolith has square channels with a density of 3.88 cells/cm2. The catalyst is primarily a mixture of platinum and palladium suspended in a binding agent and is supported on a monolith structure manufactured from mullite extrudate. The monolith was set in a galvanized steel cylinder and funnel surrounded by insulation, and this assembly was housed in a pressure vessel, as shown in Figure 1. To preheat the catalyst, an electric heater (ARi Industries, model BXX-19B-50-11T), which was also bent into a spiral with a diameter of 15 cm, was placed 5.1 cm below the catalyst. The effluent gas from the carbonization reactor entered the afterburner from its bottom and passed through the monolith upward. After leaving the afterburner, the gas stream was depressurized at a downdraft valve. In some cases, secondary air was introduced into the afterburner prior to the catalyst. This secondary air was supplied from a cylinder tank of compressed air and its flow was controlled by means of a micrometer valve. The flow rate of the secondary air was monitored by another Matheson model FM-1050 rotometer. Gas samples before and after the afterburner were periodically collected for analysis by gas chromatography. For this sampling, the gas stream at elevated pressure was first directed into a preevacuated sampling loop, and a portion of the gas within this loop was subsequently collected in a pre-evacuated, gas-tight glass tube that was equipped with a septum (a “vacutainer”). In addition, an oxygen analyzer (Bacharach, Inc., model OXOR II) monitored the oxygen concentrations in the afterburner’s effluent gas. Three K-type thermocouples were inserted into the afterburner to measure temperatures at 2.5 cm below, 2.5 cm above, and 18 cm above the catalyst monolith. These temperatures were monitored by means of the SCXI data acquisition systems manufactured by National Instruments. 2.3. Analyses. After each run, the product charcoal samples were taken out of the canister and exposed to room air overnight. The samples then were first divided into “top”, “middle”, and “bottom” sections, considering the samples’ locations in the canister. The charcoal in each section was weighed and a small amount of a representative sample from each section was subjected to proximate analysis (ASTM D 1762-84). From the results of proximate analysis for the three samples, properties of the overall samples were obtained by calculating the massweighted averages. The moisture content of the feedstock was determined by drying grab samples of the material in an oven at 105 °C until no further decrease in weight was observed. The ash content of the feedstock was analyzed by following the procedures described in ASTM E 1755-95. As shown in Figure 1, the apparatus had two water traps: one between the carbonization vessel and the afterburner and the other below the final stack. After each run, liquid condensate accumulated in these traps was taken out and its volume was measured. We usually collected 100-300 mL of liquid in total. The liquid was composed primarily of water, but its color was black, indicating that some organics were dissolved in it. The gas samples taken at the inlet and outlet of the catalytic afterburner were analyzed via gas chromatography (GC), using two gas chromatographs that were equipped with thermal conductivity detectors (Hewlett-Packard (HP) model 5890 Series II and model HP 6890). A HP model 5890 Series II gas chromatograph was used for the analysis of hydrogen, methane, carbon monoxide, carbon dioxide, and other hydrocarbons. An

Alltech Carbosphere column packed with 80-100 mesh carbon molecular sieve was used as the stationary phase, and the carrier gas was a mixture of 8% hydrogen in helium balance. A HP 6890 gas chromatograph was used for the quantification of oxygen and nitrogen. An Alltech Molsieve column packed with 80-100 mesh molecular sieve 13X was installed in the gas chromatograph, and helium gas was used as the carrier gas. The calibration curve of each gaseous component was obtained using air or standard gas mixtures manufactured by Matheson. We verified the accuracy of the calibration curves by analyzing the compositions of alternate Matheson standard gas mixtures and comparing the results with the declared compositions of the standards. If the recovery for each component, which was defined as the mole fraction determined by GC divided by the mole fraction as specified by Matheson, was within the range of 100% ( 5%, the gas analysis was accepted. Some charcoal samples were analyzed to assess their porosity. The samples were first ground and sieved into a particle size of 425-850 µm and then analyzed by a Quantachrome Autosorb-1 gas sorption analyzer. After a pretreatment under vacuum at 210 °C for 4 h, the adsorption and desorption isotherms of the sample were obtained using nitrogen as an adsorbate. From the isotherm, we calculated the BrunauerEmmett-Teller (BET) specific surface area and the total pore volume of the charcoal samples. To validate our measurement, we analyzed the surface area of commercially available Barnebey and Sutcliffe (B&S) coconut shell activated carbon. Our result (1119 m2/g) was similar to the value reported by B&S for bulk samples (1106 m2/g). 3. Results and Discussion 3.1. Results of Carbonization. After the ignition heater was turned on, usually ∼30-60 min were needed to carbonize all the biomass loaded in the canister and obtain a full bed of charcoal. More time was required than observed in the previous results (20-30 min),4 mainly because the aspect ratio of the carbonization canister used in this study (12.5:1) was larger than that of the previously used canister (3:1). 3.1.1. Yields of Products. Yields and properties of the product charcoal are summarized in Table 3. The charcoal yield (ychar) is defined as

ychar )

mchar mbio

(1)

where mchar and mbio represent the dry mass of charcoal and feedstock, respectively. This index gives a quick (but superficial) understanding of the efficiency of the carbonization run. Unfortunately, the chemical composition of the product charcoal is not taken into account in this index, and, thus, we also used the fixed-carbon yield (yfC):

yfC ) ychar

(100 -%fC %feed ash)

(2)

where %fC and %feed ash denote the percentage of fixed-carbon contained in the charcoal and the percentage of ash in the feedstock, respectively.1 This yield shows the efficiency of the conversion from the ash-free feedstock to fixed carbon contained in the charcoal product. The fixed-carbon yield from corncob feedstock reached values as high as 29.3% and that from macshell reached 32.0%. Based on the chemical composition of the biomass feedstock, it is possible to predict the theoretical limit of the fixed-carbon

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Table 3. Results of the Flash Carbonization (FC) Experiments

a HHV of charcoal calculated using eq 4. Dry basis. b Charcoal products were not separated into three sections. c Mass-weighted average of three sections (for run 4, mass-weighted average of middle and bottom sections). d Not analyzed, because of incomplete carbonization. e Contained in a metal beaker and put on the top of the corncob bed. f Separated into three sections prior to the run using two metal screens. g Charcoal from orange peel was not analyzed. h Contribution of orange peel was not included.

yield. Details of this calculation are shown elsewhere.1 Briefly, the theoretical limit is calculated by thermochemical equilibrium software as the mass fraction of solid carbon present at

equilibrium when the elements C, H, and O with known molar fractions are allowed to react at a given temperature and pressure. Using this calculation, the theoretical limits of yfC for

Ind. Eng. Chem. Res., Vol. 45, No. 2, 2006 589 Table 4. Comparison of the Proximate Analysis Results of Charcoal Samples

a HHV of charcoal measured by Huffman Laboratories, Inc., USA. Dry basis.

corncob and macshell, whose elemental compositions are listed in Table 2, are calculated to be 28% and 37%, respectively, at 400 °C and 1 MPa. Note that the theoretical limit predicted by this calculation is not significantly affected by pressure within the range of 0.001-5 MPa.4 Finally, we define the energy conversion efficiency (ηchar) as

ηchar ) ychar

(

)

HHVchar HHVbio

(3)

where HHVbio is the higher heating value (HHV) of biomass feedstock (shown in Table 2) and HHVchar is the HHV of charcoal product. Because the determination of HHVchar can be expensive when many samples are tested, we used a correlation to predict the value of HHVchar. Table 4 lists proximate analysis and HHVs of six charcoal samples produced by FC experiments. HHV was measured by Huffman Laboratories, Inc. Using our proximate analysis data and the Huffman measurement of HHV, we obtained the following correlation:

HHVchar ) 35.10 - 0.1784%VM - 0.4292%char ash (4) where %VM and %char ash represent the volatile matter content and ash content of charcoal, respectively. Note that %VM and %char ash are directly measured in the proximate analysis, whereas %fC is obtained as a difference:

%fC ) 100 - %VM - %char ash

(5)

This is why we preferred to use these two parameters in the correlation. The form of this equation is similar to that of Cordero et al.5 and Parikh et al.,6 and this equation offered a good fit (R*2 ) 0.974) to our data. Table 3 lists the values of HHVchar and ηchar for all runs. As an example, the fixed-carbon yield from corncob in run 3, in which the highest yield was obtained, reached the theoretical limit even after considering possible experimental error involved in the proximate analysis.1 This result is slightly superior to our previously4 reported result (28.0%) that was obtained under similar reaction conditions. Concerning run 3, as much as 66% of the carbon in corncobs was converted to fixed carbon in the charcoal, whereas 68% of the moist mass of feedstock was gasified. Using the correlation determined as eq 4, the charcoal product in run 3 retained 63% of the energy in the feedstock. Similarly, the fixed-carbon yield from macshell in run 12 was ∼86% of the theoretical limit. Again, the fixed-carbon yields

obtained from macshell in this study were comparable to the previous result (30.9%).4 In this run, 61% of the carbon in macshells was converted to fixed carbon in charcoal, and 67% of the moist mass of macshells was gasified. The energy conversion efficiency into charcoal was calculated as 56%. 3.1.2. Propagation of the Reaction. Unlike a batch process in a closed system, this FC process requires a forced flow of air into the system, and it is essential to decide when to halt the air delivery. If the air delivery were not sufficient, we would obtain some uncarbonized or partially carbonized pieces of feedstock; but, if the air delivery were too much, part of the bed would be overcarbonized and we would obtain a lower yield of charcoal. The spatial distribution of the products’ properties shown in Table 3 offers a comprehensive explanation concerning how the carbonization reaction propagates in the canister. Runs 3-6 were performed under similar reaction conditions as shown in Table 1. Concerning run 3, where the highest fixed-carbon yield was obtained, there was a slight variance from the bottom to the top in the proximate composition of the charcoal products. Charcoal in the middle section had the highest volatile matter content and the lowest fixed-carbon content, which was a typical result, whereas charcoal at the top showed the lowest volatile matter and the highest fixed-carbon content, suggesting that the feedstock in this section was exposed for a relatively long time to higher temperatures. Data concerning the temperature distribution in the carbonization reactor was displayed in our earlier paper.4 Run 4 was the experiment in which we intentionally delivered less air than needed for the carbonization of the entire bed. The amount of delivered air can be found in the column of delivered ABR (where ABR denotes the air/biomass ratio) in Table 5 that is defined as the mass of primary air delivered into the reactor divided by the dry mass of biomass feedstock. In this run, corncobs at the top of the bed were partially carbonized; that is, the outer surface was black, because of carbonization, but the inside remained raw and brown. The charcoal products taken from the middle section were not as well-carbonized as those obtained in run 3: a considerable amount of volatile matter remained. Because carbonization was not complete, the mass fraction of ash in the charcoal of the middle section was low, compared to the bottom sample and the run 3 samples. Incomplete carbonization can also be confirmed by the decrease of the height of the corncob bed. Generally, the height of the biomass bed decreases during the carbonization process, because biomass normally shrinks when it is converted to charcoal. In the successful run 3, the decrease of the bed height was 18 cm. On the other hand, the decrease in run 4 was 17 cmssmaller than that in run 3, despite the larger amount of feedstock. This result also indicates that the corncobs in run 4 were not completely carbonized. The vertical distribution of the products in this run proves that carbonization begins at the bottom of the bed and propagates upward against the downward airflow. Note that the charcoal yield in this run is comparable to that achieved in run 3 but the fixed-carbon yield is noticeably lower. This is an example of why we discount values of the charcoal yield in favor of the fixed-carbon yield. High charcoal yields can simply reflect incomplete carbonization. By comparison, run 5 was the test where a larger amount of air was delivered than was needed for the carbonization of the bed. As expected, excessive air delivery reduced the values of both ychar and yfC, because of the combustion of the product charcoal. The decrease of the bed height was 25 cm. Concerning the proximate compositions: charcoal at the top had a remark-

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Table 5. Air/Biomass Ratio of the Flash Carbonization Experiments

c

a Air/biomass ratio based on the delivered amount of primary air. b Air/biomass ratio based on the amount of initial air in the canister and delivered air. Approximately 0.1 kg of orange peel was put on the top. ABR was calculated using the mass of macshell only.

ably higher fixed-carbon content than the other sections. This distribution suggests that, after the pyrolysis flame reached the top of the bed, the carbonization reaction proceeded further, mainly by the combustion of charcoal. This combustion causes the remaining charcoal to reach very high temperatures, driving off any residual volatile matter and increasing the ash content of the charcoal. This perspective is in agreement with the experimental observation that charcoal at the top had a tendency to have the highest fixed-carbon content in successful runs and that a noticeable amount of white ash was deposited on the surface of the charcoal in the top section. Usually, charcoal in the top section is the hardest, having a distinctive metal sound when shaken in a container. The data of run 6 show the result with an even greater excess amount of air delivered. Notice the high ABR value shown in Table 5. In this run, the yields of charcoal and fixed carbon were significantly lower, because of overcarbonization, although charcoal with a high fixed-carbon content was obtained throughout the bed. The decrease of the bed height was 58 cm. Compared with the results of run 5, the distribution of the proximate composition shows that the combustion of charcoal propagated further down through the packed bed. This is why the charcoal in the top and middle sections had a relatively high ash content, compared to the bottom sample. The discussion above illustrates the fact that an optimal fixedcarbon yield is achieved only when an optimal quantity of air has been delivered. Consequently, air delivery must be halted at the right time. To determine when to stop the air delivery, a measurement of temperature in the carbonization canister is useful, as was done in our previous work.4 However, because of the geometry of the new canister, we did not insert thermocouples into it. Instead, we paid attention to the profile of the temperatures measured in the catalytic afterburner. As will be explained in the following section, the temperatures above the catalyst monolith had a tendency to decrease when the flaming pyrolysis reactions had consumed most of the volatile matter in the bed. Therefore, we used this temperature profile, and also the profile of the oxygen concentration measured at the stack, to determine the time to halt the delivery of primary air. As a byproduct of our study of the reaction propagation process, we observed a relationship between the location of charcoal in the canister and its porous properties. Table 6 lists the BET specific surface area and total pore volume of some representative samples of corncob charcoal. Concerning the three charcoal samples from run 9, charcoal located in the two upper levels showed larger surface areas and pore volumes than charcoal located in the bottom of the canister. We suppose that this trend is related to the extent of carbonization of charcoal products. Pores and cracks in the charcoal develop as volatile matter is eliminated at high temperatures, resulting in increased

Table 6. Porous Properties of Charcoal Samples

a Obtained from the adsorption isotherm at the relative pressure of 0.010.1. b Obtained from the adsorption data at a relative pressure of >0.995.

porosity. This explains why the “top” sample in run 9, which had the lowest volatile matter content (%VM), exhibited the highest surface area and pore volume, whereas the “bottom” sample with the highest volatile matter content showed the lowest surface area. Two additional charcoal samples (the top of run 11 and the bottom of run 2) corroborate this explanation. The top sample from run 11 (%VM ) 15.2%) had a lower porosity than the top from run 9 (%VM ) 3.2%), and the bottom sample from run 2 (%VM ) 4.1%) exhibited a higher porosity than the bottom from run 9 (%VM ) 15.1%). Evidently, the porosity and surface area of a charcoal are strongly dependent on its volatile matter content. Nevertheless, the porosity of the charcoal product was related not only to the extent of carbonization but also to the charcoal’s location itself. For example, the bottom charcoal sample from run 9 and the top sample from run 11 had almost the same volatile matter content. However, the top sample from run 11 had a much higher specific surface area and pore volume. Similarly, the bottom sample from run 2 had a low volatile matter contentscomparable to that of the top sample from run 9sbut it had a lower surface area. In summary, charcoal located in an upper position had a tendency to have higher porosity, even when the volatile matter content was comparable to that of charcoal located lower in the bed. One possible reason for this variance may be the effect of tarry vapors on the charcoal. Tarry vapors produced in the canister are forced downward through the charcoal bed, condensing on the bottom charcoal and blocking pores. We hope that closer observation of charcoal samples via microscopic techniques will offer more insights into these variances. Another interesting point to note concerns the recovery of ash. Here, we define the ash recovery (yash) by the following equation:

yash ) ychar

%char ash (%feed ash)

(6)

This measure represents how much ash was retained in charcoal, compared with the initial amount of ash contained in the feedstock. As for the runs listed in Table 1, the value of yash was in the range of 21.2%-76.7%, indicating that some part

Ind. Eng. Chem. Res., Vol. 45, No. 2, 2006 591 Table 7. Recovery of Ash in Run 9

of ash in the feedstock moved out of the canister. We estimate the relative error in ash determination to be 700 °C.8 Furthermore, chlorine aids potassium release by acting as a facilitator, transporting potassium to active biomass surfaces. Sulfur release as sulfur dioxide is observed at 500 °C upon combustion and pyrolysis.8 Note here that the FC reactor is known to reach maximum temperatures in the range of 500-770 °C.4 Furthermore, Knudsen et al.9 stated that these volatilized minerals can be captured by charcoal, which corroborates our interpretation of the distribution of yash discussed previously. 3.1.3. Effect of Pressure. As explained in the previous section, the proximate composition of charcoal, and the yield of fixed carbon, are strongly dependent on when the air delivery was halted, which is somewhat arbitrary; thus, discussion of

the effect of pressure on the products’ properties is difficult. As can be seen from Table 3, pressures of >1.14 MPa did not show any significant advantage, in terms of fixed-carbon yield, but this might be due to fewer trials at the higher-pressure conditions. Nevertheless, we emphasize that the fixed-carbon yields obtained at pressures of 1.14 MPa and above are significantly higher than those realized at ambient pressure.1 Although operating pressures in excess of 1 MPa showed little effect on fixed-carbon yields, they did influence the air utilization efficiency. The total ABR is here defined as mass of air present around biomass feedstock during the carbonization process per unit dry mass of the feedstock. This includes both the delivered amount of primary air and the amount of air initially present in the carbonization canister at the operating pressure. The calculation of primary air amount was based on the initial and final pressure of the air accumulator:

mpri.air )

(

Vac Pac,i Pac,f RairTac Zair,i Zair,f

)

(7)

where Vac, Tac, Pac, Rair, and Zair represent the inner volume of the accumulator, temperature of air in the accumulator, pressure of air in the accumulator, gas constant for air, and compressibility factor of air10 under the corresponding conditions, respectively. For the calculation of the initial air amount present in the carbonization canister, the void volume (i.e., volume that was not occupied by feedstock in the canister) was first determined using the density of the feedstock pieces, and then the amount of air was calculated as

minit.air )

PreacVvoid RairTreacZair

(8)

where Vvoid, Preac, and Treac represent void volume, pressure, and temperature of air in the canister under the initial conditions, respectively. Finally, the delivered and total ABRs were calculated using the following equations and these values for the representative experiments are listed in Table 5.

delivered ABR ) total ABR )

mpri.air mbio

mpri.air + minit.air mbio

(9)

(10)

Note that the stoichiometric ABR values for the complete combustion of corncob and macshell are 5.14 kg/kg and 6.12 kg/kg, respectively.4 The ABR value is practically determined by the time when air delivery was halted, and, thus, it became smaller in the undercarbonized runs and larger in the overcarbonized runs. For the purpose of studying the effect of pressure on the ABR values, we will focus on the “successful” runs (not undercarbonized or overcarbonized runs). Recall that we halted air delivery after the temperature above the monolith reached a peak and began to fall. As the operating pressure increased, the value of delivered ABR had a tendency to decrease, as shown in Table 5. This means that carbonization at higher pressure requires less primary air delivery and, thus, the run time can be shortened at higher pressure, provided that the delivery rate of primary air is identical. The tendency observed here is partly due to the fact that there was more air initially available in the canister at higher pressure. Nevertheless, we can still see a decrease in total ABR values as the operating pressure increases. As for

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run 13, the ABR values in Table 5 were calculated by dividing the amount of air needed for carbonization of both macshell and orange peel by the mass of macshell feedstock, only for the sake of comparison with the data of run 12. This means that the ABR values of run 13 for the carbonization of only macshell were even smaller than the values shown in Table 5. Therefore, we can clearly see that ABR for macshell decreases as the operating pressure increases. In the range of pressures studied here, higher pressures offered more rapid propagation and completion of the carbonization reaction. This is, first, because the flame was able to spread quickly upward in the bed, as a result of the elevated pressure where oxygen becomes highly available for the flame. Second, high pressure favored the exothermic formation of charcoal from biomass under the conditions of this process, as discussed previously.4 Also, the formation of primary carbon from biomass may be catalytically accelerated by the existence of water vapor or chemisorbed moisture at elevated pressures.11-13 Moreover, the formation of secondary carbon from tarry vapors is favored at higher pressures, because of the higher concentration and the longer intraparticle residence time of the tarry vapor. Consequently, less air was required for carbonization of the entire bed at higher pressures. 3.1.4. Effect of Ignition Time. In the macshell runs 14 and 15, we obtained a partially carbonized bed of feedstock, because of the relatively short ignition time. The products were black at the outer surface but brown inside, which is similar to the sample obtained in the top section of the undercarbonized run 4. These partially carbonized samples were obtained at the middle and bottom sections, whereas the top section was fully carbonized, as shown in Table 3. Based on the mechanism of reaction propagation discussed previously, this distribution of the products indicates that the carbonization flame passed through the bed from bottom to top but the intensity of the flame was insufficient to carbonize each macshell particle completely, and, after the flame reached the top, particles were further carbonized at the top of the bed. It is likely that the carbonization of a biomass particle first occurs at its surface and then the reaction gradually spreads inside, as the appearance of the partially carbonized samples suggests. In the FC process, the carbonization reaction propagates in two ways: toward the interior of each biomass particle and upward through the bed. The forced downward airflow draws the carbonization flame upward and, thus, the flame continuously moves toward the top of the bed, leaving the carbonized particles behind. Therefore, if the ignition time is short, the flame front would be smaller and less intense, and the time during which each biomass particle is exposed to the flame would be shorter; consequently, the time for carbonization would also be shorter for each particle. Especially, if the run is performed at higher pressures where the flame propagates faster, and if the feedstock is a dense material, which makes it difficult for heat to propagate into the center of the particle, the likelihood of incomplete carbonization increases. Although it is dependent on reaction conditions, we believe that the optimal ignition time for most feedstocks is ∼10 min. 3.2. Performance of the Afterburner. The effluent gas from the FC reactor usually contained 0%-10% carbon monoxide, 2%-20% carbon dioxide, 0%-3% methane, 0%-2% hydrogen, 0%-14% oxygen, and 60%-80% nitrogen (in mole fraction). Although the biomass used in this study contained both O and N atoms, almost all of the O2 and N2 in the effluent gas was derived from air delivered to the canister.4 Because some of the oxygen was consumed in the canister, its concentration in

the effluent was less than that in ambient air. On the other hand, nitrogen is regarded to be inert for the carbonization reaction, and the observed decrease of its mole fraction in the effluent was considered to be a result of the evolution of gaseous products from biomass pyrolysis and carbonization reactions. These effluent gases from the carbonization canister were directed into the catalytic afterburner. Although it did not evidence activity at lower temperatures, the catalyst worked well at temperatures above ∼260 °C. Therefore, an electric heater, which was located just below the monolith, was used to maintain the temperature of the catalyst at ∼300-400 °C. Thermocouples inserted in the afterburner showed that temperatures at 2.5 and 18 cm above the catalyst were often much higher than those below the catalyst after the catalyst had ignited the gas. Temperatures above the catalyst usually were in the range of 300-700 °C. In our reaction system, the molar flow rate of gas toward the monolith was not constant throughout the run. This was due to (1) fluctuation in the flow rate of the effluent from FC reactor and (2) temporal variation in the secondary airflow rate. During every run, the flow rate of primary air was kept almost constant, to maintain a steady supply of air into the FC reactor. Nevertheless, the flow rate of effluent coming out of the FC reactor was not constant and was larger than that of primary air, because of the generation of pyrolytic gases. Considering the nitrogen content in the effluent, the increase was, at most, 30%. On the other hand, the flow of secondary air was varied manually, based on the final oxygen concentration in the stack (explained later in detail). The ratio of the secondary air molar flow rate to that of the primary air (γ) was in the range of 0 e γ < 2. Because the total flow rate at the monolith was not constant, we used a space velocity that was based on the primary airflow rate (GHSV1) to describe our flow conditions. GHSV1 is calculated as the volumetric flow rate of primary air, corrected to normal conditions (NTP) and divided by the gross volume of the catalyst monolith (612 cm3). The actual GHSV at the monolith is obtained as

GHSV ) GHSV1(1 + γ)

( ) fN2,air fN2,in

(11)

where fN2,air and fN2,in denote the mole fractions of nitrogen in ambient air and in the gas stream at the afterburner inlet (analyzed by GC), respectively. Note that fN2,in is obtained after the mixing point of the FC reactor effluent and the secondary air, as shown in Figure 1. Under the conditions used in this study, the GHSV was determined to lie in the range of 100%330% of GHSV1. The GHSV1 used in this study was 23003900 h-1. Even after considering the variances in the actual value of GHSV, our flow condition was comparable to that of Corella and co-workers, who studied the oxidation of organic compounds by catalyst monoliths in an incineration plant14 and a biomass gasification process.15 In the following subsections, we describe the performance of the afterburner without and with secondary air delivery. 3.2.1. Results with No Secondary Air. In the absence of secondary air, the effluent gas from the preheated catalytic afterburner contained 0%-10% carbon monoxide, 9%-23% carbon dioxide, 0%-3% methane, 0%-2% hydrogen, 0%8% oxygen, and 70%-85% nitrogen in mole fraction. Generally, the concentrations of carbon monoxide and oxygen decreased by passing through the afterburner, because of the oxidation of carbon monoxide on the catalyst; thus, the concentration of carbon dioxide increased. In addition, the concentrations of


Figure 2. Temporal changes of temperature in the afterburner and exit oxygen concentration: (a) run 7, corncob at 1.14 MPa without secondary air; (b) run 3, corncob at 1.14 MPa with secondary air; (c) run 8, corncob at 1.14 MPa with secondary air; and (d) run 15, macshell at 2.17 MPa with secondary air. Symbol legend: (O) temperature 2.5 cm below the monolith, (×) temperature 2.5 cm above the monolith, (4) temperature 18 cm above the monolith, and (+) oxygen concentration.

methane and hydrogen were observed to decrease, because they were also oxidized in the afterburner. In some cases, however, considerable concentrations of unconverted carbon monoxide were released from the afterburner, because of insufficient oxygen. Figure 2a shows the temporal changes of temperatures in the afterburner during run 7 (see Table 1 for reaction conditions). In addition, the oxygen concentration in the effluent of the afterburner, as measured by an oxygen analyzer, is displayed. The effluent gas composition from the afterburner is given in Table 8a. Gas compositions at the inlet of the afterburner were not analyzed in this run. The electric heater below the monolith was turned on at the same time that the ignition heater in the carbonization canister was turned on, i.e., at 0 min of time-onstream. Primary air was delivered after 2 min of time-on-stream, with GHSV1 ) 2600 h-1. Molar flow rates at the afterburner inlet and outlet can be obtained from this GHSV1 value. For example, using the monolith volume and the definition of GHSV1, the primary airflow rate is 1.09 mol/min; and the molar flow rate at the afterburner outlet at 5 min of time-on-stream is 1.16 mol/min (also using the nitrogen outlet concentration shown in Table 8a). Until ∼16 min of time-on-stream, the temperature of the catalyst was not high enough to activate the catalyst, and, thus, the compositions of the afterburner effluent are believed to be the same as those of the effluent from the carbonization canister. At that time, the temperature below the monolith was higher than temperatures above it, just because of the difference in distance between each thermocouple and the electric heater. At 16 min of time-on-stream, the temperature of the catalyst became sufficiently high to activate the catalyst, and, soon thereafter, temperatures above the catalyst monolith exceeded those below the monolith, indicating that the catalyst monolith had ignited the combustibles. Note that the temperature at 18 cm above the catalyst was even higher than that at 2.5 cm above it, indicating that exothermal combustion extended into the region well above the monolith. The difference between the

temperatures 2.5 cm above and below the monolith reached as much as 300 °C and then began to decrease at ∼30 min of time on stream. This decrease indicates that the amount of combustible pyrolytic gases coming out of the carbonization reactor was decreasing because the entire bed of biomass feedstock was converted to charcoal and the carbonization reaction was reaching its end. Thus, we shut off the primary air delivery at 41 min of time on stream when the two temperatures above the monolith (2.5 and 18 cm above) became almost the same. Concerning the gas compositions, the oxygen concentration in the effluent decreased right after the catalyst lit and, accordingly, the carbon dioxide concentrations increased. However, considerable concentrations of carbon monoxide were still detected in the effluent, suggesting that an additional supply of oxidant was needed to achieve a complete elimination of carbon monoxide. 3.2.2. Results with Secondary Air. To enhance the combustion of carbon monoxide, secondary air was introduced into the catalytic afterburner. Secondary air was mixed with the effluent of the carbonization canister before entering the afterburner, and the delivery of secondary air was controlled so that the effluent oxygen level was maintained at a desired value. Figure 2b shows temperatures in the afterburner and oxygen concentrations in the afterburner effluent during run 3. The gas compositions at both the inlet and the exit of the afterburner are displayed in Table 8b. The reaction conditions of this run were almost the same as those of run 7 discussed previously; however, in this experiment, secondary air was supplied so that the effluent oxygen concentration was maintained at ∼4%. In this run, we activated the afterburner heater ∼30 min before the ignition heater, to preheat the catalyst monolith so that it would be active from the beginning of the carbonization run. Primary air delivery was initiated at 2 min of time on stream (GHSV1 ) 2400 h-1), and, soon thereafter, temperatures above the monolith exhibited a steep rise. At the same time, the exit oxygen concentration began to decrease rapidly and we started

594


Table 8. Gas Compositions at the Inlet and Exit of the Afterburner (a) Run 7, Corncob at 1.14 MPa without Secondary Air

(b) Run 3, Corncob at 1.14 MPa with Secondary Air

(c) Run 8, Corncob at 1.14 MPa with Secondary Air

Ind. Eng. Chem. Res., Vol. 45, No. 2, 2006 595 Table 8. (Continued) (d) Run 15, Macshell at 2.17 MPa with Secondary Air

a

Time after the ignition heater was turned on. b Not detected. c Not analyzed. d See the Gas Balance section. e Minimum values are listed. See eq 26.

secondary air delivery when the oxygen level decreased to 1%. The heater in the afterburner was controlled so that temperature of the catalyst gradually decreased to ∼300 °C. Temperatures above the catalyst monolith soon exceeded the temperature below the monolith, and they finally became more than 300 °C higher than the temperature below the catalyst. Because secondary air delivery was controlled by manually operating a micrometer valve, the effluent oxygen concentrations fluctuated at ∼4%. However, the concentrations of carbon monoxide in the effluent were, in almost all cases, lower than the detection limit of the GC analysis performed in this study (∼300 ppm). At the same time, methane was also almost completely combusted in the afterburner. Similarly, Figure 2c and Table 8c show the results of run 8 with a delivery of secondary air. Again, the reaction conditions of this run were comparable to the other two runs, as shown in Table 1. The GHSV1 in this run was 2300 h-1. In this experiment, we intended to keep the effluent oxygen concentration at 2%, although the result shows that the oxygen concentration profile was less steady than that previously displayed in Figure 2b. The concentrations of carbon monoxide and methane in the effluent of the carbonization canister were reduced in the afterburner as a result of the secondary air supply. However, carbon monoxide was detected in a few samples taken at the exit of the afterburner and, thus, the efficiency of carbon monoxide elimination was not as good as the previous run where the effluent oxygen concentration was kept at ∼4%. Figure 2d and Table 8d show the results of run 15 where macshell was used as the feedstock. The temperature 2.5 cm above the monolith is not shown in the figure because a thermocouple malfunctioned. Again, the afterburner was preheated and thus the catalyst exhibited activity right after the initiation of primary air delivery (GHSV1 ) 3500 h-1). At 4 min of time on stream, we started to feed secondary air so that the exit oxygen concentration was maintained at ∼4%. When the temperature above the monolith exceeded that below the monolith, we turned off the afterburner heater to observe the behavior of the monolithic reactor at lower temperatures. Soon, the temperature below the monolith fell to ∼100 °C, whereas that above the monolith remained ∼500 °C, showing that the catalyst was still working to oxidize carbon monoxide exother-

mally. However, at 17 min of time-on-stream, the temperature above the monolith began to fall and, at the same time, the exit oxygen concentration increased, indicating that fewer combustibles were entering the afterburner. As shown in Table 8d, the inlet concentrations of carbon monoxide, methane, and hydrogen decreased, in terms of time-on-stream, from 15 to 25 min. Also notice that the mole fraction of nitrogen in the inlet gas increased, which means that the amount of pyrolytic gases leaving the canister decreased. When the exit oxygen concentration greatly exceeded 4%, the delivery of secondary air was halted. Temperatures both above and below the monolith exhibited minima at 30 min of time-on-stream, and the carbon monoxide concentration in the afterburner reached 1.17% despite the high concentration of oxygen at the exit. This was because the catalyst was not active at such low temperatures. However, the temperature above the monolith began to increase again as the amount of combustibles entering the afterburner increased. At the same time, the exit concentration of oxygen began to fall and, subsequently, the delivery of secondary air was resumed. The temperature above the monolith continued to increase, finally attaining a temperature of 810 °C. At this stage, the amount of combustibles, as well as oxygen, was sufficient to maintain combustion without help from the afterburner heater. The large fluctuation of temperature below the monolith at 3945 min of time on stream was due to an abrupt collapse of the macshell bed in the carbonization canister that evoked a vigorous pyrolytic reaction involving a large amount of heat release. Because the carbonization of biomass causes it to shrink, sometimes the biomass feedstock bridges in the canister and then falls onto the hot charcoal below. However, such a large collapse as observed in this run was quite rare: it occurred only once during more than thirty runs. Because of the abrupt collapse, the condition in the afterburner became temporarily unsteady and we were unable to keep the exit oxygen concentration at 4 %. This is why carbon monoxide was not completely oxidized at that time where free oxygen was virtually unavailable in the afterburner. However, other than at 30 and 40 min of time on stream, carbon monoxide was completely eliminated throughout the run. The previously described observations prove that concentrations of carbon monoxide emitted from the afterburner can be

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Table 9. Example of the Gas Balance Calculation, for Gas Compositions at 36 min Time on Stream in Run 3 (a) Without Considering the Unknown Compounds

(b) Including the Contribution of the Unknown Compounds

a

Composition obtained from the GC analysis. b Composition calculated from the inlet composition and the oxidation of each compound.

reduced to below the detection limit of the GC system used in this study via the delivery of secondary air, especially when the final oxygen level is kept at ∼4%. In addition, it is possible to operate the afterburner without an inlet heater, provided that sufficient combustibles and oxygen are continuously supplied into the system. 3.2.3. Gas Balance. Without any normalization, the mole fractions of the six gas components shown in Table 8 sum to 100 ( 5%. Furthermore, no C2 or C3 hydrocarbons were detected in any of the samples. Nevertheless, closer investigation of the experimental data implies that there were some undetected compounds that were oxidized in the afterburner. For example, the inlet sample taken at 36 min of time on stream during run 3 contained 2.25 mol % carbon monoxide, 13.3 mol % oxygen, and 7.23 mol % carbon dioxide as well as some hydrogen and methane, as shown in Table 8b (see Table 9a also). Because no carbon monoxide was detected in the exit sample, it is evident that this inlet carbon monoxide was completely oxidized in the afterburner. The complete oxidation of 2.25 mol carbon monoxide requires 1.13 mol of oxygen and it produces 2.25 mol of carbon dioxide. Similar calculations can be made for the oxidations of methane and hydrogen, using their concentrations at the inlet and the exit. As a result, oxygen and carbon dioxide concentrations in the effluent were estimated to be 11.2 mol % and 10.3 mol %, respectively. However, the observed oxygen concentration was much lower and that of carbon dioxide was much higher than these estimated values. This disagreement was found in more than half of the samples; consequently we concluded that there were undetected compounds that were oxidized in the afterburner. Because such compounds were not detected in our GC analysis, we guessed that these unknown compounds were present in the gas stream entering the afterburner but were not collected in the vacutainer tube, probably because they condensed in the sampling loop of our apparatus. Actually, some black condensate has been observed in the line around the afterburner. Also, ∼100 mL of black-colored liquid, which was primarily comprised of water, has been collected in the water trap under the stack after some runs, indicating that some organic substances did pass through the afterburner. We assumed that the overall chemical composition of these unknown compounds can be expressed as COm(H2O)n. Note that

m ) 0 represents a carbohydrate, and m ) -n represents a hydrocarbon in this formula. In the afterburner, we consider the following four independent combustion reactions:

1 CO + O2 f CO2 2

[1]

CH4 + 2O2 f CO2 + 2H2O

[2]

1 H2 + O2 f H2O 2

[3]

COm(H2O)n +

2-m O2 f CO2 + nH2O 2

[4]

The aim is to estimate the value of m in the formula for the unknown compounds, as well as the extents of the four reactions. Unfortunately, the value of n cannot be evaluated, because we do not quantitatively recover all the water effluent from the reactor. The values of unknown parameters were determined as follows. First, we define fj,in and fj,out as the mole fractions of species j at the inlet and exit of the afterburner, respectively, determined by the GC analysis. Also, we define the following vectors:

fin ) (fH2,in, fO2,in, fN2,in, fCO,in, fCH4,in, fCO2,in)T

(12)

fout ) (fH2,out, fO2,out, fN2,out, fCO,out, fCH4,out, fCO2,out)T (13) The unknown compounds are not included in these vectors, because the GC analysis did not detect them. Water is also not included for the reasons stated above. Here, we consider a gas mixture whose composition is represented by fin. The molar amount of species j present in an arbitrary volume is given by

nj,in ) λfj,in

(14)

(where λ is an arbitrary constant), or in vector form, as

nin ) λfin

(15)

When these gases are mixed with a certain amount of COm-


(H2O)n and they experience reactions 1-4, the molar amount of each constituent (nj) is obtained from the following equation:

nj ) nj,in +

∑k νjkξk

(16)

where ξk is the molar extent of reaction k and νjk represents the stoichiometric coefficient of species j in reaction k. Using a matrix of stoichiometric coefficients, the equation above can be written as

(

where

n ) nin + Nξ

0 -1/2 N ) (νjk) ) 0 -1 0 1

-1 - 1/ 2 0 0 0 0

0 -2 0 0 -1 1

0 (m - 2)/2 0 0 0 1

)

ξ ) (ξ1, ξ2, ξ3, ξ4)T

(17)

(18)

(19)

For the resulting gas composition n to be identical with that experimentally determined at the afterburner exit, the following equation must be satisfied:

fout )

1

∑j

n nj

(20)

From the equations given previously, the values of the unknown parameters are obtained as

ξ/2 )

(21)

fN2,in ξ2 ) fCH4,in - fCH4,out λ fN2,out

(22)

fN2,in ξ3 ) fH2,in - fH2,out λ fN2,out

(23)

ξ/3 ) ξ/4 )

( ) ( ) ( )

fN2,in ξ1 ) fCO,in - fCO,out λ fN2,out

ξ/1 )

ξ4 fN2,in ) (f + fCH4,out + fCO2,out) λ fN2,out CO,out (fCO,in + fCH4,in + fCO2,in) (24)

m ) [fN2,in( - fH2,out + 2fO2,out + fCO,out - 2fCH4,out + 2fCO2,out) - fN2,out( - fH2,in + 2fO2,in + fCO,in - 2fCH4,in + 2fCO2,in)]/[fN2,in(fCO,out + fCH4,out + fCO2,out) fN2,out(fCO,in + fCH4,in + fCO2,in)] (25) where ξ/k is the dimensionless parameter defined as ξk/λ. Table 9b shows an example of the results obtained for the sample taken at 36 min of time on stream in run 3. The values of ξ/k and m are obtained as ξ/1 ) 0.023, ξ/2 ) 0.0053, ξ/3 ) 0.0024, ξ/4 ) 0.053, and m ) 0.11 for this case. In the aforementioned calculations, the number of moles of COm(H2O)n that are completely oxidized in the system is equal to ξ4. Therefore, the amount of unknown in the gas mixture at the afterburner inlet should be equal to or larger than ξ4. Thus,

the fraction of undetected carbon contained in the effluent from the carbonization canister (f ′uC,in) is expressed as

f ′uC,in g

ξ4 ) nCO,in + nCH4,in + nCO2,in + ξ4 1-

fN2,out(fCO,in + fCH4,in + fCO2,in) fN2,in(fCO,out + fCH4,out + fCO2,out)

(26)

Note that f denotes the mole fraction determined by the GC analysis of gas samples, whereas f ′ represents fractions in the actual gas streams around the afterburner. In the example shown in Table 9b, the value of f ′uC,in was calculated to be 0.35 or larger, which means the unknown compounds carried 35% or a larger amount of carbon, whereas the rest of carbon was contained in carbon monoxide, methane, and carbon dioxide. Because we cannot determine the amount of unknown compounds at the exit, we cannot obtain their initial amount, either. However, in the case where carbon monoxide and methane were completely eliminated and a substantial concentration of oxygen remained at the exit of the afterburner, we may assume that (i) the amount of unknown was also zero at the exit and, thus, (ii) the initial amount of COm(H2O)n was equal to ξ4 at the inlet. Tables 8b-d list the values of m and f uC,in ′ obtained by this procedure. The calculation was performed for the data sets in which the effluent oxygen concentration was much lower and that of carbon dioxide was much higher than the estimated values, as discussed previously. The value of f ′uC,in usually ranged from g0.1 to g0.3. The highest value of f ′uC,in g 0.49 was obtained at 40 min of time on stream in run 15, where an abrupt collapse of macshell bed caused an unsteady condition in the afterburner. If we express the molecular formula of the unknown compound as CpH2npO(m+n)p, based on its chemical composition, its molar amount in the gas stream at the afterburner inlet is ξ4/p or larger. Again, we may assume the molar amount is equal to ξ4/p in the case of the complete oxidation of carbon monoxide and methane. The mole fraction of the unknown compound in the entire gas stream (excluding water vapor) (f ′u,in) then is expressed as

f ′u,in g

ξ4/p ξ/4 ) λ + ξ4/p p + ξ/

(27)

4

The mole fraction of unknown compounds decreases as the value of p increases. Based on this equation, the mole fraction of unknown compounds in the case of complete oxidation of carbon monoxide and methane was calculated to range between 0 and 0.1, from the data shown in Table 8. The determination of the value of p requires a detailed investigation on the liquidphase products, which was not the scope of this study. As shown in Table 8, the value of m varied from sample to sample, suggesting that the unknown present in the effluent was not a single compound and that its composition may have changed with time. However, in many cases, the value of m was close to 0, which means the composition of unknown can be approximately expressed as CH2nOn. This composition is similar to that of the carbohydrates levoglucosan (C6H10O5) and hydroxyacetaldehyde (C2H4O2), which are representative products of the cellulose pyrolysis.3,16-29 Also, it can represent the composition of acetic acid (C2H4O2), which is a major product of biomass pyrolysis.29-31 Note that levoglucosan, hydroxyacetaldehyde, and acetic acid are the major components present in the liquid product obtained from the pyrolysis of wood.29

598


Also, the composition CH2nOn can be regarded as C + nH2O, or just as C in the case of n ≈ 0. It is well-known that the effluent gas from charcoal kilns contains a substantial amount of particulate matter.32-34 Therefore, it is possible that the effluent from the carbonization reactor in this study contained some particulate matter and that solid carbon adsorbed on the particulate matter was oxidized in the afterburner. Although data concerning the emission of particulate matter are available for conventional charcoal kilns, no information has been obtained so far for particulate matter emissions from the FC process. We are going to investigate the emission factor of particulate matter from the FC process and assess its removal efficiency in the catalytic afterburner soon. 4. Conclusions (1) Using the flash carbonization (FC) process, charcoal with fixed-carbon yields as high as 29.3% was obtained from corncob at elevated pressures of 0.791 to 2.86 MPa. This value attained the theoretical limit predicted by thermochemical equilibrium calculations. As much as 66% of the carbon in the corncobs was converted to charcoal fixed carbon, whereas 68% of the moist mass of the feedstock was gasified. The charcoal retained 63% of the energy in the feedstock. (2) Similarly, charcoal with fixed-carbon yields as high as 32.0% was obtained from macadamia nutshells (macshells) at pressures of 1.14-2.17 MPa, which was ∼86% of the theoretical limit. Sixty-one percent of the carbon in the macshells was converted to charcoal fixed carbon, and 67% of the moist mass of the macshells was gasified. The energy conversion efficiency was 56%. (3) The spatial distribution of the products’ properties indicated that the flaming pyrolysis reaction propagated upward, beginning from the bottom of the bed to its top, and that the charcoal products were further combusted and carbonized from the top downward after all the biomass was converted to charcoal. (4) Corncob charcoal with low volatile matter content had a tendency to have a larger specific surface area and total pore volume than corncob charcoal with high volatile matter content. Also, corncob charcoal that was located at a higher position in the carbonization canister had a tendency to have higher porosity. (5) The recovery of ash was always less than unity. During the FC runs, ash contained in the biomass feedstock migrated downward, probably in the form of fly ash and vapor, and some part of the ash was expelled from the canister. (6) Operation at higher pressure required less primary air delivery, and, hence less operation time, to achieve complete carbonization. This, in part, was because the reactor contained more air in its initial state. However, the total air/biomass ratio still showed a tendency to decrease as the pressure increased, suggesting that higher pressures (in the range of this study) are advantageous for the FC process. (7) Insufficient ignition time sometimes caused incomplete carbonization at higher pressures and with denser feedstocks (e.g., macshell). The distribution of the undercarbonized products indicated that the flaming pyrolysis reaction propagated upward faster at higher pressure and that the interior of the feedstock remained uncarbonized, because the duration of time each particle was exposed to the flame was not long enough. (8) The GC analyses showed that the effluent gas from the carbonization canister for both feedstocks consisted of 0%2% hydrogen, 0%-14% oxygen, 60%-80% nitrogen, 0%-

10% carbon monoxide, 0%-3% methane, and 2%-20% carbon dioxide. No C2 or C3 hydrocarbons were detected. (9) The monolithic catalyst ignited the effluent gas at temperatures of >260 °C. Following ignition, almost no carbon monoxide was detected when enough secondary air was supplied to keep the effluent oxygen concentration above ∼4%. Some of the methane and hydrogen was also combusted in the afterburner. (10) The exothermic combustion of the exhaust gas raised the temperature above the monolith. The difference between the temperatures above and below the monolith was frequently >300 °C. Even after the heater below the monolith was turned off, the afterburner was able to sustain steady combustion of the effluent gas. (11) The changes in gas compositions between the inlet and exit of the afterburner suggested the existence of some compounds that were not detected by GC analyses of the gas samples. Elemental balances and stoichiometric considerations indicated that the chemical composition of this unknown compound is similar to that of a carbohydrate (e.g., levoglucosan or hydroxyacetaldehyde). Carbonaceous particulate matter is also a possible explanation for this unknown compound. Although the identification of this unknown is not easy in our system, we plan to initiate an evaluation of particulate matter emission from the FC reactor soon. Acknowledgment We thank the following sponsors of our work for their support: Pacific Carbon and Graphite, LLC (Edward Griffiths); Carbon Diversion Corp. (Michael Lurvey); the UH Accelerated Research Commercialization (ARC) Program (Keith Mattson); the Coral Industries Endowment of the University of Hawaii; and the University of Hawaii. We also thank Lloyd Paredes (Hawaii Natural Energy Institute) for his assistance with the experimental work, Richard Cox (UH Office of Technology Transfer and Economic Development) for his continued support, and three reviewers for their constructive comments. Nomenclature ABR ) air/biomass ratio f ) vector of mole fractions fj ) mole fraction of species j determined by the GC analysis fN2,air ) mole fraction of nitrogen in air f u′ ) mole fraction of unknown compounds in the gas stream f ′uC,in ) fraction of carbon contained in unknown compounds at the afterburner inlet GHSV ) gas hourly space velocity (h-1) GHSV1 ) gas hourly space velocity based on primary air flow rate (h-1) HHVbio ) higher heating value of biomass feedstock (MJ/kg) HHVchar ) higher heating value of charcoal product (MJ/kg) HHVchar,c ) higher heating value of charcoal product calculated by eq 4 (MJ/kg) HHVchar,m ) higher heating value of charcoal product measured experimentally (MJ/kg) m ) coefficient used in the assumed composition of unknown compounds mbio ) dry mass of biomass feedstock (kg) mchar ) dry mass of charcoal (kg) minit.air ) mass of air initially present in the reactor (kg) mpri.air ) mass of delivered primary air (kg) N ) matrix of stoichiometric coefficients n ) coefficient used in the assumed composition of unknown compounds


n ) vector of molar amounts nj ) molar amount of species j (mol) P ) pressure (Pa) p ) coefficient used in the assumed formula of unknown compounds R*2 ) coefficient of multiple determination adjusted for degree of freedom Rair ) gas constant for air (Pa m3 kg-1 K-1) T ) temperature (K) V ) volume (m3) yash ) recovery of ash ychar ) charcoal yield yfC ) fixed-carbon yield Zair ) compressibility factor of air %char ash ) ash content in charcoal (%) %fC ) fixed-carbon content in charcoal (%) %feed ash ) ash content in biomass feedstock (%) %VM ) volatile matter content in charcoal (%) Greek Letters γ ) ratio of molar flow rate of secondary air to that of primary air ηchar ) energy conversion efficiency λ ) constant multiplier (mol) νjk ) stoichiometric coefficient of species j in the reaction k ξ ) vector of molar extents of reactions ξk ) molar extent of the reaction k (mol) ξ/k ) parameter defined as ξk/λ Subscripts ac ) air accumulator f ) final state i ) initial state in ) inlet of the afterburner out ) exit of the afterburner reac ) carbonization reactor void ) void space in the carbonization reactor Literature Cited (1) Antal, M. J.; Allen, S. G.; Dai, X.; Shimizu, B.; Tam, M. S.; Grønli, M. Attainment of the Theoretical Yield of Carbon from Biomass. Ind. Eng. Chem. Res. 2000, 39, 4024-4031. (2) Glaser, B.; Lehmann, J.; Zech, W. Ameliorating physical and chemical properties of highly weathered soils in the tropics with charcoals a review. Biol. Fertil. Soils 2002, 35, 219-230. (3) Antal, M. J.; Grønli, M. The Art, Science, and Technology of Charcoal Production. Ind. Eng. Chem. Res. 2003, 42, 1619-1640. (4) Antal, M. J.; Mochidzuki, K.; Paredes, L. S. Flash Carbonization of Biomass. Ind. Eng. Chem. Res. 2003, 42, 3690-3699. (5) Cordero, T.; Marquez, F.; Rodriguez-Mirasol, J.; Rodriguez, J. J. Predicting heating values of lignocellulosics and carbonaceous materials from proximate analysis. Fuel 2001, 80, 1567-1571. (6) Parikh, J.; Channiwala, S. A.; Ghosal, G. K. A correlation for calculating HHV from proximate analysis of soild fuels. Fuel 2005, 84, 487-494. (7) Bjo¨rkman, E.; Stro¨mberg, B. Release of Chlorine from Biomass at Pyrolysis and Gasification Conditions. Energy Fuels 1997, 11, 1026-1032. (8) Knudsen, J. N.; Jensen, P. A.; Dam-Johansen, K. Transformation and Release to the Gas Phase of Cl, K, and S during Combustion of Annual Biomass. Energy Fuels 2004, 18, 1385-1399. (9) Knudsen, J. N.; Jensen, P. A.; Lin, W.; Dam-Johansen, K. Secondary Capture of Chlorine and Sulfur during Thermal Conversion of Biomass. Energy Fuels 2005, 19, 606-617. (10) Perry, R. H.; Green, D. W. Perry’s Chemical Engineers’ Handbook; Seventh Edition; McGraw-Hill: New York, 1997. (11) Mok, W. S.-L.; Antal, M. J.; Szabo, P.; Varhegyi, G.; Zelei, B. Formation of Charcoal from Biomass in a Sealed Reactor. Ind. Eng. Chem. Res. 1992, 31, 1162-1166.

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ReceiVed for reView July 21, 2005 ReVised manuscript receiVed October 12, 2005 Accepted October 17, 2005 IE050854Y

Studies of the Flash Carbonization Process. 1. Propagation of the

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