Ind. Eng. Chem. Res. 2008, 47, 5957–5965
5957
Air-Steam Gasification of Biomass in a Fluidized Bed under Simulated Autothermal and Adiabatic Conditions Manuel Campoy,* Alberto Go´mez-Barea, Angel L. Villanueva, and Pedro Ollero Bioenergy Group, Chemical and EnVironmental Engineering Department, Escuela Superior de Ingenieros, UniVersity of SeVille, Camino de los Descubrimientos s/n. 41092 SeVille, Spain
Biomass gasification tests with air-steam mixtures were conducted in a bubbling fluidized-bed (FB) gasification pilot plant. The effect of steam addition on the performance of gasification was studied for different air-tobiomass ratios and biomass throughputs. The reactor was operated near adiabatic conditions, with the aim of reproducing the behavior of full-scale FB units. The temperature of the air-steam mixture was fixed at 400 °C for all tests, a value that can be achieved by energy recovery from the off-gas without tar-condensing problems. The stoichiometric ratio (ratio of actual to stoichiometric air flow rates) and steam-to-biomass flow rate ratio were varied from 0.19 to 0.35 and from 0 to 0.45, respectively. It is shown that the proper selection of operating conditions makes it possible to increase the gasification efficiency from 40% to 60%, while maintaining the low heating value of the gas at around 5 MJ/Nm3. In addition, an appreciable increase of char and tar conversion is achieved. The air-steam gasification concept explored in this work is an interesting option for the improvement of direct air-blown biomass gasification because it can improve the process efficiency considerably without significantly increasing capital costs. 1. Introduction Biomass gasification is considered one of the more promising technologies for the thermochemical conversion of biomass and waste. Atmospheric gasification with air in a bubbling fluidizedbed (BFB) reactor is the simplest way to produce a fuel gas at sufficient scale, keeping the operating costs relatively low. This gasification concept is a feasible technology that has been extensively investigated for industrial applications oriented toward direct combustion in thermal applications; co-combustion in existing pulverized-coal (PC) boilers; and if the gas is sufficiently cleaned, power production using engines and turbines.1 In air gasification, the composition of the gas is affected by high nitrogen dilution, yielding a gas with a low heating value, i.e., 4-6 MJ/Nm3. To produce a gas with a medium heating value, 9-13 MJ/Nm3, the use of steam has proven to be effective, although a considerable amount of heat has to be supplied to the gasifier. Steam gasification, however, is a more complex process to scale up, because the problem of heat supply is critical in the design of large-scale reactors. Two gasification concepts have been conceived to bring steam gasification to practical operation. The first is based on the use of oxygen (or oxygen-enriched air) and steam.2 This method, however, needs a large investment for oxygen supply equipment, making commercial-scale application less attractive. The second approach attempts to avoid oxygen by separating the combustion and gasification/pyrolysis processes into two parallel zones, one fed with air, the other with steam. Heat transfer from the combustion zone to the gasification zone is achieved by circulation of the bed inventory through the two zones. Some demonstration units operate today with this concept,1 but the capital costs are too high and the acceptance of this technology is, at the moment, uncertain. There is a vast literature on biomass gasification with air.3 Extensive work has also been published on biomass gasification with pure steam4–6 and steam-oxygen or steam-oxygen* To whom correspondence should be addressed. E-mail: mcampoy@ esi.us.es. Tel.: +34-954481391. Fax: +34-954461775.
enriched air mixtures.2,7,8 In contrast, relatively little is found related to biomass air-steam gasification.9–11 The cited works were aimed at producing a medium-heating-value gas or hydrogen-rich gas. Although there is some disagreement on the role of steam, its effect under different conditions is rather wellknown: the heating value and the H2 content in the product gas are generally higher than with air. Also, reforming and cracking of tar and C2-C3 hydrocarbons have been verified to be enhanced upon use of a sufficiently high temperature or a catalyst.12 A generalized characteristic of research in this area is the use of small fluidized-bed units with heat supplied to the reactor by an electric heater and/or with the preheating of the entering air-steam mixture. Therefore, no work investigating air-steam gasification has varied the operating conditions during autothermal operation of the reactor. Some studies have involved autothermal gasification with air,13–15 but the small size of the systems used made it impossible to operate the tests adiabatically. Ocampo et al.15 showed that carbon conversion and gas heating value increased nearly 20% as the steam-to-coal flow rate ratio was changed from 0.58 to 0.71 kg/kg in a autothermally operated coal fluidized bed (FB). However, around 20% of the total energy input was estimated as energy losses due to the small size of the reactor and the insulation quality. Mathieu and Dubuisson16 and Schuster et al.17 investigated theoretically the influence of various operating conditions based on equilibrium predictions. However, the equilibrium assumption is questionable for biomass gasification in a FB under practical operating conditions. Although some authors have pointed out the importance of carrying out tests at the pilot scale to mimic large-scale operation,18,19 most studies at the laboratory or pilot scale have been conducted allothermically: the temperature, airto-biomass ratio, and steam-to-biomass ratio have been varied independently. In general, the addition of heat to the reactor by electric heaters is neither technical nor economical to implement if the process is to be scaled up to commercial scale. Therefore, most of the existing results on air-steam gasification could be useful, in the best case, for optimizing the operating conditions for indirect gasifiers and for oxygen-steam blown gasification.
10.1021/ie800220t CCC: $40.75 2008 American Chemical Society Published on Web 07/02/2008
5958 Ind. Eng. Chem. Res., Vol. 47, No. 16, 2008
Figure 1. Layout of the FBG pilot plant.
However, for autothermal gasification at the industrial scale, the a priori best conditions achieved at the laboratory scale might not be attained in the scaled-up unit because the variables are interdependent. Moreover, there is no information in the literature on the improvement resulting from the addition of low-quality (low-pressure, near-saturation conditions) steam to an air-blown autothermal fluidized-bed gasifier (FBG). When the operation is undertaken autothermically, the addition of low-quality steam leads to a decrease in the reactor temperature. This option could potentially be interesting, however, because the presence of small amounts of steam can significantly alter the gas composition by enhancing the reforming reactions of hydrocarbons, as well as the tar and heterogeneous gasification reactions.15 Obviously, the thermal level in the gasifier needs to be kept reasonably high to promote the aforementioned reactions. This work is focused on the experimental investigation of the impact of steam addition in an airblown FBG simulating adiabatic operation, thereby mimicking the performance of a large-scale FBG, i.e., a gasifier with low wall heat losses. In addition, all of the tests were conducted at a fixed temperature of the incoming air-steam mixture, a value that could be achieved by recovering the energy from the produced gas such that tar-condensing problems were avoided; i.e., the outlet gas temperature from the system was above the tar dew point of the heavier hydrocarbons. The main objective of the tests was to study the effects of the stoichiometric ratio (SR), defined as the ratio between the actual air fed to the gasifier and the air necessary for stoichiometric combustion of the biomass, and the steam-to-biomass ratio (SBR), on the quality and composition of the produced gas.
2. Experiment Materials. The fuel used for the tests was wood pellets with the empirical formula CH1.4O0.64 [dry and ash-free (daf)], similar to that of others woods.10,13 The moisture and ash content were 6.3% and 0.5% (mass basis), respectively, and the lower heating value of the fuel (as received) was 17.1 MJ/kg. The pellets were cylindrically shaped with a mean diameter of 6 mm and a height of 5-10 mm. The apparent density of the pellets was 1300 kg/ m3, whereas the bulk density was 600 kg/m3. The bed material used in the FBG was ofite, a silicate subvolcanic rock with formula (Ca, Mg, Fe, Ti, Al)2(SiAl)2O. The ofite had a density of 2650 kg/m3 and an average size of 290 µm. The theoretical minimum fluidization velocity for the average particle size was 0.12 m/s, calculated using Wen and Yu’s correlation.20 Ultimate and elemental analyses as well as particle size distribution of the ofite are reported in Gomez-Barea et al.21 Facility. The reactor was a bubbling FB with two zones, namely, bed and freeboard, having diameters of 150 and 250 mm and lengths of 1.40 and 2.15 m, respectively. Figure 1 shows a layout of the facility in the current state, and Table 1 presents the most important parameters of the pilot plant under the tests conditions of this work. The plant has been described in detail in previous publications.14,21 Only the most outstanding modifications conducted to date are described here. The feeding system was pressurized by connection with primary air to avoid backflow of hot gas from the reactor. An additional knife valve was installed and a depressurization system was included to support the downflow of difficult fuels. An electrical resistance of 1 kW was installed in the air line to
Ind. Eng. Chem. Res., Vol. 47, No. 16, 2008 5959 Table 1. Technical and Operating Data for the Facility and Main Test Operating Data parameter
value
bed inside diameter bed height freeboard inside diameter freeboard height bed material fuel fuel flow rate gasification agent operating temperature operating pressure regime of fluidization
0.15 m 1.40 m 0.25 m 2.15 m ofite wood pellets 12-21 kg/h air, air + steam 730-815 °C atmospheric bubbling
preheat the air to about 200 °C. The steam was generated near saturation conditions (105-120 °C) by a 30-kW electrical boiler having a maximum throughput of 45 kg/h. The hot air and saturated steam were brought into contact and preheated in a 7-kW electrical heater before entering the plenum. This temperature was set to 400 °C, in an effort to simulate favorable energetic integration as discussed below. A new distributor was used, having 32 holes (2.5 mm in diameter) in a square arrangement with a 19-mm triangular pitch. A pipe for ash removal was connected to the distributor by a central 38-mmdiameter hole. The ash-removal pipe had two ball valves with a small bin between them, enabling removal of ash from the bed in a controlled way. The already-existing overflow system for ash removal was not used during the tests presented here. A 27-kW electrical furnace was installed as a substitute for the insulation blanket used previously. The oven had three heating zones whose set points were set independently: one in the bed with a power of 21 kW and two in the freeboard with a power of 3 kW each.The power supplied was measured using signal electronic duplicators, connected to PID controllers, transmitting the signal to a programmable logic controller (PLC). The two gas cyclones were insulated to avoid tar and steam condensations, problems detected during earlier operation of the plant. The thickness of the insulation guaranteed temperatures higher than 500 °C at the cyclones and higher than 300 °C at the tar sampling point situated downstream of the cyclones. A tar sampling system was designed following the recommendations of the European Tar Protocol.22 The sample point was situated upstream of the wet scrubbing system, whereas the gas analyzer port was downstream of the scrubber. Test Procedure. At the beginning of each test, a batch of bed material (8 kg) was added to the reactor. The bed was heated with hot air and the electrical furnace. After 1 h, the bed temperature reached 750 °C, and the freeboard temperature was higher than 400 °C. It was necessary, however, to wait for an additional 2-3 h before starting the feeding of the biomass in order to prevent tar deposition downstream of the gasifier. For this purpose, the feeding of the biomass was postponed until the gas temperature upstream of the scrubber reached 300 °C. Once these requirements were fulfilled, a small amount of biomass was fed into the reactor. Under this excess of oxygen, the biomass was oxidized completely, and the reactor was rapidly heated to the desired process temperature. The transition from combustion to gasification was made by increasing the biomass flow rate to decrease the air-to-biomass ratio. Once this ratio was established and the plant was verified to be under steady-state conditions, the steam was fed into the system (in the case of air-steam tests). In every test, an initial transitory period of 3-4 h was followed by a steady-state period of 5-7 h. From the beginning of the test, the computer-based data acquisition system monitored and recorded the temperatures,
the pressures, the gas composition (H2, CO, CO2, CH4, O2), the power supplied to the heating equipments, the flow rates of the fluidizing agent and the produced gas, and the instantaneous weight of the biomass in the feeding system. The tar, particle, and water vapor contents of the product gas were also measured twice during some tests. The operation was completed by taking a sample from the bed inventory and combusting the remaining char in the bed. After each test, the two cyclone bins and the extraction ash bin were sampled and analyzed. Operating Conditions. The experimental program comprised tests with air and mixtures of air and steam, some of which were conducted twice, showing good reproducability. The air tests were carried out as a reference to test the effects of steam addition. As the fluidized bed was operated close adiabatically, the necessary heat for the gasification reactions was provided by combustion of part of the biomass. In this mode of operation, given the type of biomass (wood pellets), the bed material (ofite), and the FB design, there are three variables that can be independently varied: the flow rates of biomass, air, and steam. Two ratios, for a given biomass flow rate, can be defined for the analysis of the system: (1) the stoichiometric ratio (SR), defined as the mass ratio between the amount of air fed and the stoichiometric amount of air required for combustion, and (2) the steam-to-biomass ratio (SBR), defined as the flow rate of steam fed to the reactor divided by the flow rate of biomass [dry and ash free (daf)]. The biomass throughput itself is a third variable to analyze; however, for a limited range of this parameter, the system can be analyzed by means of the two above ratios, as the biomass flow rate is expected to have a minor influence on the results. The SR and SBR were varied from 0.19 to 0.35 and from 0 to 0.45, respectively, by adjusting both the steam flow rate (from 0 to 5.1 kg/h) and the biomass flow rate (from 12.2 to 20.5 kg/h) while keeping the air flow rate roughly constant (17 Nm3/ h). An experimental matrix was constructed with three values of SR (0.19-0.23, 0.27, and 0.33-0.35) and three values of SBR for each SR value. Table 2 provides a summary of the tests and the main results. The matrix was tentatively designed with the aim of assessing the operation at low, medium, and high SR and SBR levels, while keeping the temperature high enough (727-812 °C) for proper gasification under the two following constraints: (1) adiabatic operation and (2) a temperature of 400 °C for the air-steam mixture fed to the gasifier. This is the result of a simulated system with a heat recovery exchanger that is feasible in practical operation without tar condensation. The achievement of constraints 1 and 2 is discussed below. The gas residence times in the bed and freeboard are important operating data, and therefore, they should be specified clearly. Significant differences have been found with regard to the reporting of residence times in earlier works. Ideally, the gas residence time in the bed should account for the gas from biomass devolatilization as well as that from the gasification of char. Similarly, the gas residence time in the freeboard must be clearly defined to avoid the uncertainty introduced by the significant drop in temperature usually found in the freeboard of laboratory-scale rigs, typically 200-300 °C. In this work, we calculated the residence time in the bed based on the total amount of gasification agent fed to the reactor and the mean bed temperature (the four thermocouples situated along the bed registered essentially the same temperature). For the gas residence time in the freeboard, the calculation was made by considering the measured outlet dry gas and the mean temperature of the freeboard (water vapor was not considered because
5960 Ind. Eng. Chem. Res., Vol. 47, No. 16, 2008 Table 2. Test Results test
L1-A
L2-AS
L3-AS
M1-A
M2-AS
M3-AS
H1-A
H2-AS
H3-AS
SBR SR air flow rate (Nm3/h) biomass flow rate (kg/h) steam flow rate (kg/h) mean bed temperature (°C) furnace set point (°C) mean freeboard temperature (°C) gas flow rate (Nm3/h, dry) CO2 (%v/v) (dry) CO (%v/v) (dry) H2 (%v/v) (dry) CH4 (%v/v) (dry) tar content (g/Nm3) GY tar yield (g/kg daf) char yield (g/kg daf) CH4 yield (g/kg daf) LHV gasification efficiency (%) carbon conversion (%) char conversion (%) tbedb (s) tfreeboardc (s)
0 0.19 17.0 20.5 0 780 770 687 22.3 14.2 18.2 13.2 6.0 25.8 0.6 30 56 50 5.9 37 87 65 1.4 4.8
0.18 0.23 17.0 17.5 3.0 752 740 675 23.2 16.9 13.8 14.6 5.2 NAa 0.7 NAa 55 53 5.2 40 89 70 1.1 4.7
0.28 0.19 15.5 19.1 5.0 727 720 660 26.1 18.6 11.5 16.2 5.9 NAa 0.8 NAa 56 61 5.3 42 89 70 1.1 4.3
0 0.27 17.0 15.0 0 805 800 718 21.4 14.9 17.6 12.6 5.2 23.8 0.8 37 64 57 5.4 45 88 65 1.3 4.9
0.23 0.27 17.0 15.0 3.2 786 780 708 25.9 16.2 15.0 14.0 4.7 NAa 0.9 NAa 51 62 5.1 51 90 73 1.1 4.1
0.43 0.27 17.0 15.0 6.0 755 750 709 26.5 18.6 11.9 16.2 5.3 NAa 1.0 NAa 46 71 5.1 53 91 75 1.0 4.0
0 0.35 17.0 11.5 0 812 800 716 23.3 15.1 15.8 8.7 5.1 17.6 1.0 38 53 79 4.8 56 89 71 1.3 4.5
0.22 0.33 17.0 12.2 2.5 804 790 721 27.3 15.9 15.4 11.9 4.8 16.7 1.2 40 48 82 4.9 65 90 74 1.1 3.8
0.45 0.33 17.0 12.2 5.1 789 780 709 28.4 17.0 13.8 13.3 4.6 15.4 1.2 38 39 82 4.8 66 92 79 1.0 3.7
a Not available. b tbed based on the total amount of gasification agent fed to the reactor and the average bed temperature. c tfreeboard based on the gas outlet and the average freeboard temperature.
of the uncertainty of this value). The freeboard temperature was computed as the mean of the three measured temperatures, having a maximum drop of 100 °C at steady-state operation. The reason for this drop was probably that the heating element situated in the freeboard zone was not sufficient to avoid some heat loses, because of the considerable size of the freeboard. Mean temperatures in the bed and freeboard for all tests are reported in Table 2. The gas velocity calculated in this way ranged from 1.03 to 1.45 m/s for the bed and from 0.44 to 0.58 m/s for the freeboard. Experimental pressure fluctuations observed during operation evidenced a bubbling regime in the bed. This was theoretically confirmed by using fluidization maps20 entering with a minimum fluidization velocity of 0.12 m/s and gas superficial velocities ranging from 1 to 1.45 m/s. To study the risk of entering the spouted-bed regime, we calculated the maximum bubble size expected in the bed to be 0.04 m according to Mori and Wen’s correlation20 and checked the maximum length-to-diameter ratio for the spouted-bed regime by Baeyens and Geldart’s equation.20 These experimental and theoretical calculations allow us to confirm that our FB was operated under bubbling conditions. The corresponding gas residence times for the bed and freeboard were calculated taking into account the lengths given in Table 1 and ranged between 1.0 and 1.4 s and between 3.7 and 4.9 s, respectively. We assume that the ranges are sufficient small that we can neglect the effect of the gas residence time on the results presented below. 3. Results and Discussion Achievement of Adiabatic Tests. One significant aspect of the tests presented in this work is that they were conducted nearly adiabatically: the heating system was controlled to provide just the necessary amount of heat to compensate for the heat losses. This operation was achieved by keeping the temperature of the furnace slightly lower than the temperature inside the reactor. Figure 2a shows the experimental setup used. The signal from the thermocouple situated in the bed was continuously transmitted to the PLC, whereas the temperature of the reactor external wall (close to the heating element of the furnace) was sent to the PID, where the value was compared to
Figure 2. (a) Arrangement used for the measurement and control of temperature. (b) Response of bed temperature (solid line) to an increase in the flow rate of biomass. The dotted line is the set point of the furnace established during the test.
a set point. Figure 2b shows the evolution of the bed (solid line) and set-point (dashed line) temperatures during a typical test. In the figure, three steady-state temperatures can be distinguished: 806, 788, and 777 °C. The external wall temperature of the reactor was not transmitted to the PLC, and so, it is not included in the figure. The details presented in Figure 2b help to explain the control approach applied: The solid line drops at time t1 because of the increase in the biomass flow rate. Once the operator observes this drop, which is monitored online, the set point of the PID is decreased manually from 780 to 770 °C at time t2. The bed temperature evolves with time, reaching a new steady-state
Ind. Eng. Chem. Res., Vol. 47, No. 16, 2008 5961
Figure 3. Simulation of the concept of preheating for the determination of the temperature of the air-steam mixture fed to the gasifier.
Figure 4. Bed temperature as a function of SBR for different SR levels.
temperature (around 778 °C). The setting of the set point is made by the operator based on the experience acquired in past trials and by prior theoretical calculations applying mass and energy balances. In this example, the set point was adjusted to 770 °C, and the eventual temperature attained in the bed was 777 °C. This difference was considered good enough, fulfilling the criterion of a maximum absolute difference temperature of 10 °C. Another feature of the tests is that the air-steam mixture was preheated in order to simulate the heat that could be recovered from a heat exchange with the exiting product gas. The preheating was undertaken by two electrical resistances as shown in Figure 1. The cooling of the product gas was constrained in such a way that tar condensation was avoided; i.e., the outlet temperature of the product gas was higher than 330 °C. Figure 3 shows a flowsheet simulating a tentative exchange in which the product gas temperature drops by 416 °C (from 750 to 334 °C). Calculations yielded a maximum temperature of the air-steam mixture of 400 °C. The simulation arrangement presented in Figure 3 probably does not represent actual practice, where the gas cooler necessary for conducting the heat exchange might be difficult to design and further considerations must be taken into account.23 However, the potential heat from this energy recovery system was taken as the basis for setting the temperature of the air-steam mixture fed to the reactor at 400 °C for all of the trials conducted in this work. In this way, the tests were considered autothermal, even though the fluidization agent fed to the reactor was at 400 °C. Influence of SR and SBR. The variables analyzed include the gas composition, tar content, gas yield, heating value, carbon and char conversions, and gasification efficiency. The effects of the SBR and SR on the gasification temperature are shown in Figure 4, indicating that the temperature in the bed decreases as the SBR increases for the three SR levels analyzed. In fact, steam addition decreases the temperature of the bed under autothermal operation, with the decrease being
smaller at high SR, when the biomass throughput is the lowest (around 12 kg/h, see Table 2). Quite different behavior was found in previous studies.8–10 These differences are explained by the operating procedures followed in the different works: In the studies of Turn et al.8 and Sadaka et al.,9 to vary SBR, the steam rate was kept constant while the biomass feed rate was changed. In the work of Lv et al.,10 the biomass and air flow rates were kept constant, but the heat supply was varied from one test to another. In this work, the air flow rate was maintained constant, while the biomass and steam flow rates were changed (see Table 2). In addition, the tests were conducted adiabatically, and the temperature of the air-steam mixture was 400 °C in all tests. The effect of the SBR on the gas composition is presented in Figure 5. Figure 5a indicates that, at all SR levels, the concentration of carbon dioxide increases as the SBR is increased. The opposite trend is observed for carbon monoxide (Figure 5b). Analysis of the slopes suggests that the effect of the SBR becomes less significant as the SR increases. Figure 5c shows the effect of the SBR on the hydrogen content of the product gas. An increase in the SBR from 0 (air gasification) to 0.3-0.4 increases the H2 content significantly. The results reported by Lv et al.10 showed similar trends, but the effect of the SBR on the gas composition was much smaller. The trends found here are also not in agreement with the results of Sadaka et al.,9 who showed that higher SBR values resulted in lower CO2 concentrations and that the CO content was affected by the superficial velocity of the gas. Figure 5d shows the molar fraction of methane as a function of SBR for different SR levels. As can be seen, the CH4 content is fairly constant, ranging between 4.5% and 6%, in agreement with previous works. From an analysis of the ordinate axis (SBR ) 0, i.e., air tests) in the graphs of Figure 5, one can deduce that CO2 is fairly constant for the three SR levels, whereas CO and H2 decrease significantly as the SR increases. This decrement is especially relevant for H2 at the highest SR value. Finally, the addition of steam increases the ratio of H2 to CO significantly, as is clearly seen in Figure 6, where this ratio is displayed as a function of SBR for the three SR levels investigated. In brief, the results from Figures 5 and 6 suggest an enhancement of the water-gas shift reaction (CO + H2O T CO2 + H2) with the steam concentration in the bed. This seems to be true despite the decrease in bed temperature caused by the steam injection. A proper way to assess this effect would have been the calculation of the steam conversion in the bed. Unfortunately, the moisture content of the gas was not known because of clogging problems in the probe during measurements; in addition, the trials made by closing mass and heat balances were not reliable enough to predict proper values of steam conversion. Figure 7a shows the tar concentration of the product gas as a function of temperature for the five tests in which tar was analyzed (see Table 2): the three air tests and the two air-steam tests conducted at the highest SR level. The trend was as expected for the air tests: the higher the temperature in the bed, the lower the tar content in the produced gas. A significant reduction of tar, from 25.8 to 17.6 g/Nm3, was achieved by increasing the temperature of the tests from 780 to 812 °C. The tests using steam, i.e., SBR * 0 (represented by squares in Figure 7a) showed, in addition, a slight reduction of tar content as the SBR increased. This is explained by the fact that the temperature of the test carried out at the highest SBR (0.45) was lower than that conducted at an SBR of 0.22, specifically, 790 vs 804 °C. Therefore, both the steam content and the temperature affect the tar content of the product gas. Clearly,
5962 Ind. Eng. Chem. Res., Vol. 47, No. 16, 2008
Figure 5. Gas composition as a function of SBR for different SR levels: (a) carbon dioxide, (b) carbon monoxide, (c) hydrogen, and (d) methane.
Figure 6. H2/CO ratio (v/v) as a function of SBR for different SR levels.
for the two air-steam tests presented in Figure 7a, the opposite effects on tar conversion caused by temperature and steam concentration seem to be roughly counterbalanced, making the observed tar content similar from one test to another (16.7 vs 15.4 g/Nm3). Another fact that is worth noting in relation to Figure 7a is that the three tests using air only were conducted for very different biomass throughputs: 20.5, 15, and 11.5 kg/h. In principle, if the throughput of biomass is varied over too wide of a range, it might affect the results. To explore this possibility, the tar content was calculated per unit of biomass throughput, i.e., as tar yield expressed in grams of tar per kilogram of waterand-ash free biomass). The results are displayed in Figure 7b as a function of temperature. As can be seen, except for the test conducted at the highest biomass throughput, the yields are similar in all tests, being roughly 38 g/kg (daf). Unfortunately, there are not enough tar measurements to further quantify this behavior, but we can summarize the trends as follows: For a limited range of biomass throughput, bed temperature and steam injection seem to affect the tar concentration of the gas considerably, whereas they affect the tar yield only weakly. In contrast, if the biomass throughput is increased significantly, the tar yield seems to be affected considerably even for tests with equal SR and SBR values. Figure 7c shows the effect of the SBR on the yield of methane. The methane yield varies considerably, from 50 to
80 g/kg of biomass (daf), over the range of operating conditions analyzed. This is a surprising result because the presence of steam is usually accompanied by higher reforming of methane. In this case, this effect would have minor importance because of the relatively low thermal level and the absence of catalyst. In contrast, the opposite behavior was observed, which was quite unexpected. We conclude that, under the conditions tested, the use of steam favors the production of methane during devolatilization, but its effect on steam reforming reactions is limited. In any case, Figure 7 confirms the methane yields found by other investigators for similar operating conditions and types of biomass.24 The gas yield is defined as the ratio of the volumetric flow rates of the dry and N2-free product gas and the biomass (daf). In Figure 8, it is shown that the gas yield increases with increasing SBR for all SR levels tested. This result apparently agrees with some of the previous works analyzed.2,6 However, differences were found with regard the gas yield behavior at different steam and oxygen concentrations: Lv et al.10 reported a decrease of gas yield with increasing SBR. Walawander et al.4 identified different regimes for the range of temperatures 600-780 °C. Franco et al.6 reported that, for three different biomasses, the gas yield reached maximum values for an SBR of 0.6-0.7, with no changes occurring for SBR values higher than 0.7. Lv. et al.10 and Rapagna et al.25 analyzed the effects of particle size and type of biomass and found that gas production is greatly affected by these variables, especially at lower temperature, i.e., 750 °C, where the gas composition and gas yield are strongly affected by biomass devolatilization. Our observations indicate a continuous increase in the gas yield with SR and SBR, showing the benefits of the two variables for producing higher gas yields. Figure 9 shows the lower heating value (LHV) of the product gas as function of SBR for the three SR levels tested. It can be seen that the SBR lowers the LHV of the product gas for low and medium SR levels, whereas the variation seems to be smaller at higher SR. In the latter case, the LHV attains a maximum within the SBR range of 0.2-0.3. The observations are as expected: in the air tests, the LHV decreases as the SR is increased because of the higher amount of gas combusted.
Ind. Eng. Chem. Res., Vol. 47, No. 16, 2008 5963
Figure 7. Effects of temperature on tar and methane content in the produced gas: (a) Tar concentration. The SBRs applied to the air-steam tests (9) were SBR ) 0.45 and SBR ) 0.22 for the test at 789 and 804 °C, respectively. (b) Tar yield for different operating conditions [mbiomass(kg/h)-SR-SBR]: b, 20.5-0.19-0; 2, 12.2-0.33-0.45; 9, 12.2-0.33-0.22; 0, 15-0.27-0; O, 11.5-0.35-0. (c) CH4 yield for different SBR levels.
Figure 8. Gas yield as a function of SBR for different SR levels.
Figure 9. LHV of the produced gas as a function of SBR for different SR levels.
This effect is partially counterbalanced by steam addition in the air-steam tests, as more tar (and char to some extent, as we show below) is reformed, and so, higher amounts of light fuel gases are generated. It is worth noting that in the calculation of the LHV, the energetic content of H2, CO, and CH4 was taken into account. Other fuel gases such as C2 and other light hydrocarbons (which were not measured), as well as tar, were not considered in the computed LHV, so the actual LHV of the gas is higher than
the value reported here. In addition, the LHV is presented in Figure 9 for the dry gas, including the diluting effect of nitrogen. The reason is that the gas produced in this process (an autothermal gasifier using air as a source of oxygen) is thought to be used for burning in thermal applications or, if the gas is properly cleaned, for power production. Consistently, the LHV of the gas must include the gross gas, including nitrogen. This is not the usual practice for air-steam gasification in the literature, where the LHV is expressed on a nitrogen-free basis because, in this case, the gas is mainly aimed for mediumheating-value (MHV) gas production or for high hydrogen gas production. This explains the differences in the range of LHVs obtained in this work, from 4.7 to 6 MJ/Nm3, compared to other air-steam gasification publications, up to 12 MJ/Nm3. The carbon conversion expresses the transformation of the carbon contained in the fuel to gaseous species. Figure 10a shows that the carbon conversion ranges from 87% to 92% and increases with SBR for the three SR levels. These values are in agreement with those of other works in which similar fuels, operating conditions, and reactor types were used.10,11,13 These trends are in agreement with the results of Lv et al.,11 but disagree with those of Sadaka et al.,9 who reported that the carbon conversion decreased with increasing SBR. The differences can again be explained by the different operating conditions of the tests. The char conversion, defined as the conversion of the carbon in the char generated after the biomass devolatilization, is a useful indication of the extent of solid conversion attained in the gasifier. Figure 10b presents the char conversion attained in the tests calculated from the analysis of the cyclone ashes. The char conversion was computed by assuming that 20% w/w of the biomass is converted into char after devolatilization, having a carbon content of 96.5% (w/w). These data were obtained experimentally by Go´mez-Barea et al.,26 for the devolatilization of single pellets in the presence of nitrogen at various temperatures in a batch-operated laboratoryscale FB. As shown in Figure 10b, the char conversion ranges from 60% to 75%, giving clear quantification of the energetic
5964 Ind. Eng. Chem. Res., Vol. 47, No. 16, 2008
Figure 10. (a) Carbon conversion and (b) char conversion as functions of temperature for different SBR levels.
Figure 11. Cold gasification efficiency as a function of SBR for different SR levels.
inefficiency caused by the unconverted char in the bed. Despite the relatively low temperature, the char conversion achieved is rather high. In contrast, the effect of steam on char gasification is not as high as expected: only an increase of 15% of the char input was achieved by increasing the SBR from 0 to 0.45. The gasification efficiency is defined as the ratio of the heat of combustion of the produced gas, calculated through the LHV defined above, to the heat of combustion of the feedstock. The cold gasification efficiency assumes that the temperature of the product gases is 25 °C, so the sensible heat of the gas is not included. Figure 11 displays the cold gasification efficiency as a function of SBR for the three levels of SR investigated. As seen, the addition of steam increases the efficiency from 55% to 65% for the highest SR value tested. This is because gasification efficiency accounts for the combined role of the LHV and the gas yield, which have been shown to increase. These trends are in agreement with the results of Sadaka et al.9 and Lv et al.10 The low values found for the efficiency can be explained as follows: (1) Light and heavier hydrocarbons are not taken into account in the calculation of the the LHV of the gas. It has been estimated that, based on the gas composition reported by earlier investigators, the gasification efficiency could increase an additional 10%. (2) Despite the special attention paid on conducting the test adiabatically, some wall heat losses could
still have occurred. From data reconciliation of these (and other tests), the wall heat loss was estimated to be around 5% of the input energy in the input fuel. (3) The aforementioned inefficiency caused by the unconverted char, which can be quantified as an additional 4-8% of the input energy. In any case, the trends obtained in this experimental study are rather clear, allowing establishing the effect of steam on the gasification efficiency. Overall Discussion. Overall, the many effects involved in the process analyzed make it difficult to identify clear mechanisms responsible for the results. In general, an increase in gas yield is favored by high temperature and high steam concentration in the bed during allothermal gasification. Temperature and steam concentration favor (1) higher production of gas in the devolatilization step, (2) enhancement of gas production by char gasification with steam, and (3) an increase in gas yield due to steam reforming and cracking of hydrocarbons and tars. The addition of low-quality steam in autothermal gasification lowers the bed temperature but increases the steam concentration in the bed; therefore, a compromise of the two opposing effects can occur. The results of this work suggest that char gasification with steam (effect 2 above) is not enhanced as much by the presence of steam as could be expected in steam gasification. Clearly, the temperature is not as high as would be needed for rapid char gasification for the residence time of char in this work, whereas it is sufficient to enhance devolatilization and secondary gas-gas reactions such as the reforming of tars, CO (water-gas shift reaction), and light hydrocarbons (mainly C2 and C3). Variations in the biomass flow rate in some tests resulted in different char residence times, but the observed influence on char conversion was limited. This confirms that mechanisms 1 and 3 are the main processes responsible for the benefits of using steam observed in this work. The increase of efficiency by the presence of steam is explained, therefore, by the higher yield in devolatilization and the subsequent cracking and reforming of hydrocarbons. Finally, comparison with earlier works was difficult because of the differences in the way that the tests were conducted. It was especially difficult to evaluate the heat supplied by the oven in published works to validate our results by extrapolation, because this variable is rarely reported. Further work must be done to refine our results, especially in terms of measuring tar and light hydrocarbons. The role of the biomass throughput also has to be refined, particularly its effects on the tar concentration in the gas. Improvement of the heat supplied to the system is underway through automatic control of the system by different thermocouples located along the reactor. This should reduce the temperature drop in the freeboard and might improve the simulation of full-scale plants. Modeling of the system has also been done, and the results will be presented in the future. 4. Conclusions Tests in a bubbling FBG were conducted in an effort to explore autothermal FB air-steam gasification technology with steam-air mixtures. The reactor was operated adiabatically, thereby making the laboratory-scale results more useful for scaling up to large-scale gasification units. The results obtained differed considerably from those of earlier investigations in laboratory and pilot plants, usually conducted allothermally, i.e., with heat addition. It is shown that the optimization of the flow rates of biomass, steam, and air increases the gasification efficiency significantly. In this work, the tests were conducted by keeping the air flow rate constant while varying the biomass
Ind. Eng. Chem. Res., Vol. 47, No. 16, 2008 5965
and steam flow rates. As a consequence, as steam addition was increased, the biomass throughput was decreased in order to maintain a sufficient temperature level in the reactor. The autothermal air-steam gasification studied in this work seems to be an interesting compromise between direct air-blown atmospheric FBG operation and indirect gasification. This option reduces the capital costs and, as shown, can improve the process efficiency considerably. Nomenclature daf ) dry and ash-free FB ) fluidized bed FBG ) fluidized-bed gasification GY ) gas yield (Nm3 of dry, N2-free gas/kg daf) LHV ) lower heating value (MJ/Nm3 of dry gas) PC ) pulverized coal PLC ) programmable logic controller SR ) stoichiometric ratio SBR ) steam-to-biomass ratio tbed ) gas residence time in the bed zone tfreeboard ) gas residence time in the freeboard zone
Acknowledgment This work was partly financed by the European Commission, the Commission of Science and Technology of Spain, and the Junta de Andalusia. The authors acknowledge Eduardo Sa´nchezPen˜a and Rafael Moriana for their assistance in the sampling and analysis of tars. Literature Cited (1) Maniatis, K. Progress in Biomass Gasification: An Overview. In Progress in Thermochemical Biomass ConVersion; Bridgwater, A. V., Ed.; Blackwell Science: London, 2001. (2) Gil, J.; Corella, J.; Aznar, M. P.; Caballero, M. A. Biomass Gasification in Atmospheric and Bubbling Fluidized Bed: Effect of the Type of Gasifying Agent on the Product Distribution. Biomass Bioenergy 1999, 17, 389. (3) Narvaez, I.; Orio, A.; Corella, J.; Aznar, M. P. Biomass Gasification with Air in an Atmospheric Bubbling Fluidized Bed. Effect of Six Operational Variables on the Quality of the Produced Raw Gas. Ind. Eng. Chem. Res. 1996, 34, 2110. (4) Walawender, W. P.; Hoveland, D. A.; Fan, L. T. Steam Gasification of Pure Cellulose. 1. Uniform Temperature Profile. Ind. Eng. Chem. Process. Des. DeV. 1985, 24, 813. (5) Herguido, J.; Corella, J.; Gonzalez-Sanz, J. Steam Gasification of Lignocellulosic Residues in a Fluidized Bed at a Small Pilot Scale. Effect of the Type of Feedstock. Ind. Eng. Chem. Res. 1992, 31, 1274. (6) Franco, C.; Pinto, F.; Gulyurtlu, I.; Cabrita, I. The Study of Reactions Influencing the Biomass Steam Gasification Process. Fuel 2003, 82, 835. (7) Wang, Y.; Kinoshita, C. M. Experimental Analysis of Biomass Gasification with Steam and Oxygen. Sol. Energy 1992, 49, 153. (8) Turn, S.; Kinoshita, C.; Zhang, Z.; Ishimura, D.; Zhou, J. An Experimental Investigation of Hydrogen Production from Biomass Gasification. Int. J. Hydrogen Energy 1998, 8, 641.
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ReceiVed for reView February 7, 2008 ReVised manuscript receiVed April 29, 2008 Accepted May 2, 2008 IE800220T
