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Pressurized gasification of lignite in a pilot scale bubbling fluidized bed reactor with air, oxygen, steam and CO2 agents
Applied Thermal Engineering 130 (2018) 203–210
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Applied Thermal Engineering journal homepage: www.elsevier.com/locate/apthermeng
Research Paper
Pressurized gasification of lignite in a pilot scale bubbling fluidized bed reactor with air, oxygen, steam and CO2 agents Serhat Gül ⇑, Fehmi Akgün, Emir Aydar, Namık Ünlü TUBITAK Marmara Research Center Energy Institute, 41470 Gebze, Kocaeli, Turkey
h i g h l i g h t s Air, oxygen, steam and CO2 gases are used as gasification agent. Experiment was carried out with 500 kWth capacity pilot scale pressurized gasifier. 3
3
Syngas LHV was varied between 8.2 MJ/Nm and 8.6 MJ/Nm . Cold gas efficiency is approximately 58.3%. Carbon conversion is varied between 72.0% and 84.6%.
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Article history: Received 7 June 2017 Revised 3 October 2017 Accepted 4 November 2017 Available online 6 November 2017 Keywords: Bubbling fluidized bed Pressurized Lignite Gasification Pilot scale
a b s t r a c t In this study, 80 h continuously operated gasification experiment was carried out with pilot scale pressurized bubbling fluidized bed gasifier which was operated with the coal feed capacity of 500–520 kWth (80–83 kg/h coal feed rate). The mixture of air, steam, oxygen and CO2 was used as gasification agent with different ratios. The operating pressure of gasifier in this experiment was between 2.4 and 2.7 barg. The effects of ER, steam/carbon ratios and CO2/carbon ratios on syngas composition, carbon conversion ratios and cold gas efficiencies were investigated. According to the results, using CO2 or H2O as gasification agent, may have both advantages and disadvantages for the operation. If CO2 was increased in gasification agent, shift reaction was effected negatively and H2 products were decreased. If steam was increased, cold gas efficiency was slightly decreased. Ó 2017 Elsevier Ltd. All rights reserved.
1. Introduction Bubbling fluidized beds (BFB) are widely used for thermal processes, such as drying, combustion, gasification and mineral processing. In industrial applications, gasification process is used for producing syngas from coal or biomass for various applications such as generate heat and power, some chemicals, synthetic natural gas (SNG), hydrogen or liquid fuels production. The properties and composition of the synthetic gas produced by gasification process depend on many factors of design and operational parameter such as the reactor type, the properties and flow rates of feedstock materials and gasifying agents, pressure and temperature in the reactor. In this study, pressurized pilot scale bubbling fluidized bed gasification was investigated for different gasification agents with various ratios.
⇑ Corresponding author. E-mail address: [email protected] (S. Gül). https://doi.org/10.1016/j.applthermaleng.2017.11.021 1359-4311/Ó 2017 Elsevier Ltd. All rights reserved.
According to the previous studies published in the literature, air was generally used for gasification agent as the reference base case. The effect of oxygen and steam addition to the gasification agent was investigated as well. Kumar et al. [1] studied the gasification of coal with 80–100 kWth pilot facility using air and steam as gasification agent. Steam to coal mass flowrate and equivalence ratios (ER) are varied between 0.15–0.25 and 0.25–0.35, respectively. According to the results of this study, the calorific value of syngas increase with the decrease of ER value and increase of steam to coal ratio. Similarly, cold gas efficiency increases with the decrease of ER and increase of steam to coal ratio. Gil et al. [2] reported the result of biomass gasification with the atmospheric bubbling fluidized bed using different gasification agents (air, steam and steam-oxygen mixtures). Their results show that, with the mixture of H2O and O2, H2 and CO concentrations are between 25–30% and 43–47%, respectively. When the pure steam is used as gasifying agent, H2 and CO concentrations are become 53–54% and 21–22%, respectively. Campoy et al. [3] represents the effect of oxygen concentration in the gasification agent using oxygen enriched air-steam
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of 5195 kCal/kg. The proximate and elemental analyses of selected lignite are shown in Table 1. The particle size of coal fed to the gasifier was in between 0.5 mm and 1 mm.
gasification in a bubbling fluidized bed. The ER value and the steam to biomass ratios are varied from 0.24–0.38 to 0–0.63, respectively. They reported that the increase of the oxygen concentration from 21% to 40% increase the cold gas efficiency from 54% to 68%. Karimipour et al. [4] investigated the effect of steam to O2 ratio on the quality of syngas produced with fluidized bed gasification of lignite. Their study was performed with the steam to O2 ratio between 0.5 and 1.0. They concluded that the carbon conversion increases with the increase of steam/O2 ratio. H2/CO ratio strongly related with the steam/O2 ratio. Kern at al. [5] studied the gasification of lignite with the 90 kWth dual fluidized bed gasifier. Steam to carbon mass ratio is varied between 1.3 and 2.1. According to their results, decreasing the steam/carbon ratios increase the cold gas efficiency. When this result is compared with the previous studies, it could be concluded that there is an optimum value of steam/coal ratio, apart from this at which higher or lower ratios result in the decreases in the cold gas efficiency. Arena et al. [6] studied the air gasification process of municipal waste in a pilot scale bubbling gasifier with the capacity of 70 kg/h and 400 kWth. They studied with the ER value and the bed temperature of 0.25– 0.33 and 850–930 °C, respectively. The lower heating value (LHV) of syngas and the cold gas efficiency of gasifier are about 5 MJ/ Nm3 and 60%, respectively. Carbon conversion achieved during the experiments is varied between 80 and 90%. There are many other studies related with the investigation of the effect of the temperature of gasification reactor [7–9], equivalence ratio [10–13] and steam to carbon ratio [14–16]. Most of the these research in literature carried out on the bench-scale and/or laboratory scale systems, except in some of these studies [8,9,15,17,18] used pilot scale gasification system with thermal capacities relatively higher than the others (80–100 kWth). In this study, the test results of a pressurized bubbling fluidized bed gasification system with 500–520 kWth fuel feeding capacity was analyzed for 80 h continuously operation period by using air, oxygen, steam and CO2 mixture as gasification agent with different ratios. Normally, gasifier was designed for generating syngas for coal to liquid process including gas cleaning, gas conditioning and FT processes. Just before the FT reactor, there was a high pressure compressor in order to pressurize the syngas (up to 30 barg). Due to the pressure losses up to the FT compressor, gasifier has to be work at approximately 2.5 barg. Because of that, 2.5 barg pressure level was selected as operating pressure in this study and gasifier was tested independently from the other processes. The capacity of the system is relatively higher than that of most of the research published in literature. So, the test results of pilot system may provide valuable insights into the process and operation of larger scale gasification systems.
2.2. Experimental setup The schematic of the pilot scale pressurized bubbling fluidized bed gasifier system is shown in Fig. 1. The system consists of BFB reactor, cyclone separator, syngas cooler, quencher, syngas burner, fuel feeding system, ash removal system and gasification agent (air, steam, O2, CO2) preheating systems. The BFB reactor is 3500 mm in high. Inner diameters of dense bed and freeboard regions of the reactor are 300 mm and 450 mm respectively. There is a cyclone separator with 140 mm inner diameter. Just after the cyclone separator, there is a syngas cooler for cooling down the syngas temperature from approximately 800 °C to 300 °C. The final stage of the system is quencher which is used for cooling syngas to approximately 30–40 °C with spraying the cooling water. An air compressor was used for the supply of pressurized air. The flowrate of the air was measured and controlled with a mass flow controller (MFC). Oxygen and CO2 gasses were supplied by two different gas storage tanks. A steam generator was used for the steam supply at the saturated conditions (at the pressure of 6 barg and the temperature of 165 °C). In the BFB reactor, the mixture of fluidization gas passed through the windbox and the distributor plate to the reactor. The distributor plate was made out of stainless steel and has 100 standpipes with 4 nozzles which are 1 mm in diameter. There were ten temperature measurement points located on the system in order to measure the temperatures of ambient air, 1st heater exit, 2nd heater exit, lower dense bed, upper dense bed, lower freeboard, upper freeboard, cyclone exit, heat exchanger exit and quencher exit, respectively. Pressure measurement device were also located on the same points with the temperature sensors. The reactor pressure was controlled by the pressure control valve located at the syngas exit just before the syngas burner. The system was monitored and controlled with the distributed control system (DCS). The gas analyzer was used for measuring the syngas composition. SIEMENS Process Chromatography Maxum Edition II was used to analyze syngas composition. It measures H2, O2, CO, CO2, N2, CH4 and C6+ with a specific Thermal Conductivity Detector (TCD). 2.3. Experimental method 2.3.1. Startup of the gasification reactor In the gasification experiment, 80 kg bed material (silica sand) was loaded to the reactor which had the initial static bed height of approximately 530 mm. Air flowrate was adjusted in order to get 0.8 m/s superficial gas velocity at the dense bed region of gasifier at the atmospheric conditions (20 °C, 0 barg). At this condition, the bed material was in bubbling regime. In order to heat up the bed material to the ignition temperature of coal, the electric heater was used for preheating the fluidization air. There were 2 groups of electric heaters; each consists of three serial connected 20 kW electric heaters. With these electric heaters, the fluidized bed temperature was increased to 350 °C which was the practical limit of
2. Experimental 2.1. Materials Silica sand was used as bed material which had size distribution between 0.3 mm and 0.7 mm and weighted arithmetic mean diameter of 644 mm according to the sieve analysis. Bulk and particle densities were 1450 kg/m3 and 2545 kg/m3, respectively. Turkish lignite (Soma lignite) was used as feedstock which had the LHV
Table 1 Proximate and elemental analysis of selected Turkish coal (mass basis). Proximate analysis
Elemental analysis
Moisture (%)
Volatile matter (%)
Fixed carbon (%)
Ash (%)
6.6
38.15
42.6
12.65
C (%) 60.56
H (%) 3.89
O (%) 14.32
S (%) 0.92
N (%) 1.06
H2O (%) 6.60
Ash (%) 12.65
LHV (kcal/kg) 5195
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Fig. 1. Schematic of gasification system.
ignition of coal particles. After reaching the temperature of 350 °C in the bed, coal was started to feed (approximately 10 kg/h) and combustion reactions started and the temperature increased very rapidly to 870 °C. Just after the reaching steady state combustion condition, the fuel feeding rate was approximately increased 3 times (30 kg/h) in order to get the gasification conditions at the constant air flowrate. The temperature slightly increased during this process and then decreased to its initial value. After reaching steady state gasification conditions, fine tuning of temperature was done with adjusting the fuel flowrate. 2.3.2. Startup of syngas cooler Before heating of bed material, thermal oil cycle at syngas cooler (shell&tube) was started in order to increase the temperature of thermal oil which was passing through the tube side of syngas cooler. At the steady state conditions, thermal oil temperature must be approximately 200 °C due to prevent the heavy hydrocarbon condensation on the surface of tube of heat exchanger. The temperature of thermal oil was held constant at desired value with adjusting the level of boiling water tank which was used for cooling the thermal oil coil immersed to the boiling water at atmospheric conditions.
constant automatically at desired level. The quencher was used not only decreasing the temperature of the syngas, but also decreasing the particle and heavy hydrocarbon contents in the syngas in order to protect pressure control valve located at the end of the process. 2.3.4. Pressurizing gasifier Gasifier pressure was held at desired value with pressure control valve which has PID controller. The desired value of gasifier pressure was set and PID controller of the control valve adjusted the pressure value and maintained the gasifier pressure at set point. As can be seen at Fig. 1, pressure control valve was located just after the syngas quencher and before the syngas burner. Gasifier pressure was increased gradually with 0.1 bar increments with pressure control valve. During the pressurization process, actual volumetric flow of fluidization agent was decreased. In order to held constant superficial gas velocity at the dense bed region, the gasification agent must be increased. The mass flowrate of coal also increased to operate the gasifier at the same temperature. At each 0.1 bar increments, it was waited to get stable operational conditions (especially in temperature levels) in gasifier. 3. Results and discussions
2.3.3. Startup of quencher After the reaching the steady state atmospheric gasification conditions, quencher pumps were started to feed 1000 kg/h water to quencher with two nozzles (totally 2000 kg/h) at different heights of quencher. The water level at the bottom of the quencher was controlled with the control valve and the water level was held
3.1. Operating conditions of experiment 3.1.1. Temperatures and pressures variations during the operation In this experiment, the continuous operation period of pressurized BFB gasifier system was 80 h. The time variation of
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temperature in the dense bed is shown in Fig. 2. As shown in this figure, the temperature of bed was fluctuating between 870 °C and 880 °C. The possible reason of these fluctuations could be the changes of coal properties during the operation. There were two different temperature measurement points for dense bed. One was located at the bottom of the dense bed which is shown in Fig. 2, and the other one was at the top of the dense bed. During the operation of gasifier, it was observed that the average temperature difference between the bottom and the top of the bed was 0.5 °C and the maximum difference was 1.6 °C which means the fluidization and the homogeneity of the bed was very good. The temperature profile of the system from the beginning of the process to the end is shown in Fig. 3. During the experiment, the average of the ambient air temperature was approximately 22 °C. Gasification agent passed through the 1st and 2nd electric heater and temperature was increased to the 192 °C and 291 °C, respectively. Preheated agent entered to the gasifier and the bottom of the dense bed temperature was 873 °C. As mentioned previously, top of the bed was approximately same (873.5 °C) with the bottom temperature. The next two temperature measurement points were lower freeboard and upper freeboard temperatures which were 872 °C and 832 °C, respectively.
Syngas temperature decreased from 832 °C to 738 °C while passing through the cyclone separator due to the heat losses. After the cyclone separator, syngas reached to the syngas cooler which was shell&tube heat exchanger and syngas passed through the shell side. Temperature was decreased from 738 °C to 434 °C and flew through the quencher. At the quencher, 2 1000 kg/h of water was sprayed to the quencher column with two nozzles (1000 kg/h for each). Temperature of syngas was decreased from 434 °C to 37 °C while passing through the quencher. The majority of the moisture content of syngas was taken at quencher. During the operation, the pressure of gasifier was in between 2.4 barg and 2.7 barg with little fluctuations, as shown in Fig. 4. The pressure was tried to be held constant at desired value. There were some fluctuations due to the fluctuations of syngas flow rate or some resistance/blockage through the process. The averages of pressure values through the process are shown in Fig. 5. 1st pressure measurement point was at the exit of 1st electric heater and its value was 2.92 barg. The pressure difference between the 2nd heater exit and lower dense bed was approximately 40 mbar which was the pressure loss at the distributor plate. The pressure loss due to the dense bed was approximately
Fig. 2. Temperature of the dense bed.
Fig. 4. Gasifier reactor pressures during the operation.
Fig. 3. Temperature profile through the process.
Fig. 5. Pressure profile through the process.
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80 mbar which was the difference between the pressure of lower dense bed and lower freeboard. At freeboard region, there was almost no pressure loss due to the very low solid entrainment and low solid concentration. Cyclone pressure loss was approximately 6 mbar which was the difference between upper freeboard and cyclone exit pressures. The pressure losses of syngas cooler and quencher were 5 mbar and 30 mbar, respectively. The pressure just before the pressure control valve was 2.74 barg and syngas was discharged to the atmosphere by pressure control valve. 3.1.2. Feeding rate changes of coal and gasification agents during the operation The flowrates of coal, steam, air, oxygen and CO2 during the experiment are shown in Fig. 6. As shown in this figure, the air flowrate was held constant at 10 Nm3/h. Air was fed to the gasifier in order to make nitrogen balance between inputs and syngas in order to determine the flowrate of syngas. Oxygen flowrate was changed in order to control the temperature and ER value, between 21 Nm3/h and 28 Nm3/h. CO2 was fed to the gasifier in order to determine the effect on gasifier performance. CO2 addition was tested for 4 different flowrates (25, 22, 35, 30 Nm3/h). In addition, H2/CO ratio could be also controlled at acceptable level with CO2 addition. Significant change in CO2 flowrate was occurred at 53th, 62nd and 71st hours. Coal feed rate during the experiment was between 80 kg/h and 83 kg/h, which correspond to 500 kWth and 520 kWth, respectively. Steam was mainly fed to the gasifier in order to get required actual volumetric flowrate for proper superficial velocity (0.8 m/s) through the dense bed. During the experiment, the effect of steam to carbon ratio was also investigated with changing the flowrate of steam between 100 kg/h and 140 kg/h. Considerable change in steam flowrate was occurred at 27th, 33th and 65th hour. 3.1.3. The variation of ER and steam to carbon ratio during the operation The equivalence ratio (ER) during the experiment was changed between 0.25 and 0.33, as shown in Fig. 7. In fact, the ER and the temperature of the reactor are coupled operation parameters. In this study, temperature of the reactor was selected as control parameter and ER value was considered as the function of the reactor temperature. Another important parameter for the operation of the gasifier is steam to carbon ratio which is shown in Fig. 8. Steam/carbon ratio
Fig. 7. ER value during the experiment.
Fig. 8. Steam to carbon ratio.
(S/C) was changed between 2.0 and 2.8 during the experiment. The system requires more fluidization agent for proper fluidization due to the pressurized operation of the gasifier. In this experiment, majority of the fluidization agent came from superheated steam. Because of that, S/C ratio for this experiment was over than 2. 3.2. Syngas composition
Fig. 6. Gasification agent flowrates.
Syngas composition was measured periodically during the experiment with 5 min periods with online gas analyzer. As shown in Fig. 9, CO2, H2, CO, N2 and CH4 were measured. Syngas CO2 and H2 contents were varied between 33.3–37.8% and 31.5–34.2%, respectively, which depended on the operation conditions. Considerable changes in percentage of the CO2 content were numbered in Fig. 9. At point 1, CO2 concentration increased while H2 was approximately constant. At that point, both ER and steam/carbon ratios(S/C) were increased together. CO2 content was proportional to the both ER and S/C ratios. Because with the increase of ER value which means also increase of oxygen content in gasification agent, combustion reaction increased and CO2 formation increased. Similarly, when the S/C ratio increased, water gas shift reaction was promoted and CO2 content increased. H2 content was also proportional to the S/C but inversely proportional to the ER value. In this situation, due to the increase of both ER and S/C, CO2 increased and H2 was not changed. At point
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Fig. 9. Syngas composition during the experiment. Fig. 10. Syngas flowrate during the experiment.
2, CO2 content increased while H2 content decreased due to the increase of ER value without the increase of S/C ratio. At point 3, 4 and 5, CO2 flowrate at gasification inlet was significantly decreased, increased and decreased, respectively. There were two reasons for this situation. First, with the increase or decrease of CO2 flowrate at the inlet of the gasifier, unreacted CO2 content in the syngas increased and decreased, respectively. Second, the shift reaction was affected by the CO2 content at gasification agent. If the CO2 content at gasification agent was increased, the shift reaction equilibrium conditions were changed and products of reaction (H2 and CO2) were decreased. The CO content of syngas was varied between 18.6 and 22.0%. Considerable changes in percentage of the CO contents were also numbered in Fig. 9. At point 6 and 7, CO content increased and decreased, respectively. While the change of CO percentage, S/C ratio was decreased and increased for point 6 and 7, respectively. S/C ratio promoted the shift reaction and CO was decreasing. The N2 and CH4 contents of syngas were varied between 6.7– 7.8% and 4.0–4.6%, respectively. N2 content was approximately constant during the experiment. CH4 content very slowly increased from 4.0% to the 4.6% at the end of the experiment. Syngas C6 content was also measured and varied between 0.28% and 0.46%. At first half of the experiment; C6 content was lower than the second half of the experiment which was related to the S/C ratio. S/C ratio was also lower at first half and higher at second half. Syngas flowrate was calculated with the N2 balance between input N2 (coming from air and coal) and syngas N2 content. Nitrogen content in coal fed to the gasifier can be mainly released as N2, NH3 and HCN. Small amount of fuel N is not released and stay in ash [19]. During the experiment, NH3 and HCN content were not measured. Due to this, it was assumed that fuel N was released as N2 and corresponded to approximately 0.7 Nm3/h N2 content in the syngas. Flowrate of syngas during the experiment is shown in Fig. 10. As shown in this figure, the syngas flowrate varied between 108.2 Nm3/h and 129.0 Nm3/h. The calculation of syngas flowrate was very sensitive to the change in N2 content of syngas. Because of that, the small fluctuations in measurement of N2 content of syngas causes very big fluctuations at syngas flowrate calculations. The lower heating value (LHV) of syngas was calculated according to the syngas composition, as shown in Fig. 11. The LHV of syngas varied between 8.2 MJ/Nm3 and 8.6 MJ/Nm3. These values were higher than that of the air gasification of coal (3–5 MJ/ Nm3) and lower than that of the oxy-steam gasification (9–12 MJ/Nm3). In this experiment, there were some differences in LHVs comparing with that of the pure oxy-steam gasification, because
Fig. 11. Syngas LHV (MJ/Nm3) during the experiment.
some amount of air and CO2 was also added to the gasification agent. Therefore, nitrogen from the air and excess CO2 content in the syngas resulted in a decrease in the LHV of the syngas. 3.3. Cold gas efficiency and carbon conversion The cold gas efficiency of gasification process during this experiment is shown in Fig. 12. As shown in this figure, cold gas efficiency was approximately 58.3% (average of the data). At first half of the experiment, efficiency was slightly higher than that of the second half due to the lower steam feed rate at first half. Steam addition resulted in an increase the mass of hot syngas and more combustion reaction take place at reactor in order to hold temperature constant. The increase in combustion reaction resulted in decrease of cold gas efficiency. Operation pressure also affected the efficiency of gasifier. When the operation pressure is higher than the atmospheric pressure, which was approximately 2.6 barg in this experiment, much more steam was fed to the gasifier than the required for chemical reactions in order to get required fluidization velocity at dense bed. In fact there is a limit of coal conversion load (throughput) of bed which depends on design capacity, coal reactivity, coal particle size and temperature. For this pilot scale gasifier, the practical limit
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S. Gül et al. / Applied Thermal Engineering 130 (2018) 203–210 Table 2 Average values of carbon balance between coal and syngas (mass basıs). C fuel (kg/h)
CCO2 (kg/h)
CCO (kg/h)
CCH4 (kg/h)
CC6 (kg/h)
CSyngas (kg/h)
49.86
22.54
12.82
2.76
1.52
39.64
conversion, the ratio of carbon content of syngas to the fresh carbon was calculated. Syngas carbon content was calculated with the syngas composition which includes CO2, CO, CH4 and C6 molecules (Eq. (2)). Average values of carbon balance between coal and syngas is shown in Table 2. 4. Conclusions
Fig. 12. Cold gas efficiency of gasification process.
of bed coal conversion load observed during the experiments was approximately 11 MWth/m3. Over this limit of conversion load, the carbon concentration in fluidized bed which was 8.2% for this experiment is increased and the homogeneity of bed was broken and agglomeration problem occurs. Due to this reason, when pressurizing the gasifier, the coal feed rate and the gasification agent flowrate should not be increased proportionally. The coal feed rate approximately held constant and the steam flowrate was increased while pressurizing process. Another important performance criterion for gasification process is carbon conversion ratio. Carbon conversion calculated with the carbon balance between coal and syngas is shown in Fig. 13. As shown in figure, carbon conversion varied between 72.0% and 84.6% (average value 80.3%). Majority of carbon was lost with fly ash which the major part was captured by the cyclone. Carbon balance between streams was determined with the equations below (mass basis). Carbon flowrate fed to the gasifier with coal must be equal to the total amount of carbon in the streams of syngas, bottom ash and fly ash (Eq. (1)).
Cfuel ¼ Csyngas þ Cbottom ash þ Cfly ash
ð1Þ
Csyngas ¼ CCO2 þ CCO þ CCH4 þ CC6þ
ð2Þ
During the experiment, the amount of carbon in the bottom ash and fly ash were not determined. In order to determine the carbon
The present work attempted to analyze the effect of some operational parameters on pilot scale pressurized bubbling fluidized bed gasification system. The properties and composition of the synthetic gas produced by gasification process depend on many factors of design and operational parameter such as the reactor type, the properties and flow rates of feedstock materials and gasifying agents, pressure and temperature in the reactor. The following conclusions can be drawn from the experimental results. – With the increase or decrease of CO2 flowrate at the inlet of the gasifier, unreacted CO2 content in the syngas increased and decreased, respectively. Shift reaction was also affected by the CO2 content at gasification agent. If the CO2 content at gasification agent was increased, shift reaction equilibrium conditions were changed and products of reaction (H2 and CO2) decreased. – Syngas CO content was varied between 18.6 and 22.0%. While the increase and decrease of S/C ratio, CO percentage of syngas decreased and increased, respectively. S/C ratio promoted the shift reaction and CO concentration decreased. – Syngas LHV was varied between 8.2 MJ/Nm3 and 8.6 MJ/Nm3. Different from the oxy-steam gasification, some air and CO2 was added to the gasification agent. Because of nitrogen content from the air and excess CO2 content in the syngas decreased the syngas LHV. – Cold gas efficiency average value was approximately 58.3%. At first half of the experiment, efficiency was slightly higher than the second half due to the lower steam feed rate at first half. Steam addition increased the mass of hot syngas and more combustion reaction took place at reactor in order to held temperature constant. Increase in combustion reaction resulted in decrease of cold gas efficiency. – Carbon conversion calculated with the carbon balance between coal and syngas was varied between 72.0% and 84.6%. Majority of carbon was lost with flyash.
Acknowledgments This study has been realized within the project entitled ‘‘Liquid Fuel Production from Biomass and Coal Blends”. TUBITAK is greatly acknowledged for the support of this project under the ‘‘Support Programme for Research Projects of Public Institutions” (the contract number 108G043). References
Fig. 13. Carbon conversion of gasification process.
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Applied Thermal Engineering journal homepage: www.elsevier.com/locate/apthermeng
Research Paper
Pressurized gasification of lignite in a pilot scale bubbling fluidized bed reactor with air, oxygen, steam and CO2 agents Serhat Gül ⇑, Fehmi Akgün, Emir Aydar, Namık Ünlü TUBITAK Marmara Research Center Energy Institute, 41470 Gebze, Kocaeli, Turkey
h i g h l i g h t s Air, oxygen, steam and CO2 gases are used as gasification agent. Experiment was carried out with 500 kWth capacity pilot scale pressurized gasifier. 3
3
Syngas LHV was varied between 8.2 MJ/Nm and 8.6 MJ/Nm . Cold gas efficiency is approximately 58.3%. Carbon conversion is varied between 72.0% and 84.6%.
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Article history: Received 7 June 2017 Revised 3 October 2017 Accepted 4 November 2017 Available online 6 November 2017 Keywords: Bubbling fluidized bed Pressurized Lignite Gasification Pilot scale
a b s t r a c t In this study, 80 h continuously operated gasification experiment was carried out with pilot scale pressurized bubbling fluidized bed gasifier which was operated with the coal feed capacity of 500–520 kWth (80–83 kg/h coal feed rate). The mixture of air, steam, oxygen and CO2 was used as gasification agent with different ratios. The operating pressure of gasifier in this experiment was between 2.4 and 2.7 barg. The effects of ER, steam/carbon ratios and CO2/carbon ratios on syngas composition, carbon conversion ratios and cold gas efficiencies were investigated. According to the results, using CO2 or H2O as gasification agent, may have both advantages and disadvantages for the operation. If CO2 was increased in gasification agent, shift reaction was effected negatively and H2 products were decreased. If steam was increased, cold gas efficiency was slightly decreased. Ó 2017 Elsevier Ltd. All rights reserved.
1. Introduction Bubbling fluidized beds (BFB) are widely used for thermal processes, such as drying, combustion, gasification and mineral processing. In industrial applications, gasification process is used for producing syngas from coal or biomass for various applications such as generate heat and power, some chemicals, synthetic natural gas (SNG), hydrogen or liquid fuels production. The properties and composition of the synthetic gas produced by gasification process depend on many factors of design and operational parameter such as the reactor type, the properties and flow rates of feedstock materials and gasifying agents, pressure and temperature in the reactor. In this study, pressurized pilot scale bubbling fluidized bed gasification was investigated for different gasification agents with various ratios.
⇑ Corresponding author. E-mail address: [email protected] (S. Gül). https://doi.org/10.1016/j.applthermaleng.2017.11.021 1359-4311/Ó 2017 Elsevier Ltd. All rights reserved.
According to the previous studies published in the literature, air was generally used for gasification agent as the reference base case. The effect of oxygen and steam addition to the gasification agent was investigated as well. Kumar et al. [1] studied the gasification of coal with 80–100 kWth pilot facility using air and steam as gasification agent. Steam to coal mass flowrate and equivalence ratios (ER) are varied between 0.15–0.25 and 0.25–0.35, respectively. According to the results of this study, the calorific value of syngas increase with the decrease of ER value and increase of steam to coal ratio. Similarly, cold gas efficiency increases with the decrease of ER and increase of steam to coal ratio. Gil et al. [2] reported the result of biomass gasification with the atmospheric bubbling fluidized bed using different gasification agents (air, steam and steam-oxygen mixtures). Their results show that, with the mixture of H2O and O2, H2 and CO concentrations are between 25–30% and 43–47%, respectively. When the pure steam is used as gasifying agent, H2 and CO concentrations are become 53–54% and 21–22%, respectively. Campoy et al. [3] represents the effect of oxygen concentration in the gasification agent using oxygen enriched air-steam
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of 5195 kCal/kg. The proximate and elemental analyses of selected lignite are shown in Table 1. The particle size of coal fed to the gasifier was in between 0.5 mm and 1 mm.
gasification in a bubbling fluidized bed. The ER value and the steam to biomass ratios are varied from 0.24–0.38 to 0–0.63, respectively. They reported that the increase of the oxygen concentration from 21% to 40% increase the cold gas efficiency from 54% to 68%. Karimipour et al. [4] investigated the effect of steam to O2 ratio on the quality of syngas produced with fluidized bed gasification of lignite. Their study was performed with the steam to O2 ratio between 0.5 and 1.0. They concluded that the carbon conversion increases with the increase of steam/O2 ratio. H2/CO ratio strongly related with the steam/O2 ratio. Kern at al. [5] studied the gasification of lignite with the 90 kWth dual fluidized bed gasifier. Steam to carbon mass ratio is varied between 1.3 and 2.1. According to their results, decreasing the steam/carbon ratios increase the cold gas efficiency. When this result is compared with the previous studies, it could be concluded that there is an optimum value of steam/coal ratio, apart from this at which higher or lower ratios result in the decreases in the cold gas efficiency. Arena et al. [6] studied the air gasification process of municipal waste in a pilot scale bubbling gasifier with the capacity of 70 kg/h and 400 kWth. They studied with the ER value and the bed temperature of 0.25– 0.33 and 850–930 °C, respectively. The lower heating value (LHV) of syngas and the cold gas efficiency of gasifier are about 5 MJ/ Nm3 and 60%, respectively. Carbon conversion achieved during the experiments is varied between 80 and 90%. There are many other studies related with the investigation of the effect of the temperature of gasification reactor [7–9], equivalence ratio [10–13] and steam to carbon ratio [14–16]. Most of the these research in literature carried out on the bench-scale and/or laboratory scale systems, except in some of these studies [8,9,15,17,18] used pilot scale gasification system with thermal capacities relatively higher than the others (80–100 kWth). In this study, the test results of a pressurized bubbling fluidized bed gasification system with 500–520 kWth fuel feeding capacity was analyzed for 80 h continuously operation period by using air, oxygen, steam and CO2 mixture as gasification agent with different ratios. Normally, gasifier was designed for generating syngas for coal to liquid process including gas cleaning, gas conditioning and FT processes. Just before the FT reactor, there was a high pressure compressor in order to pressurize the syngas (up to 30 barg). Due to the pressure losses up to the FT compressor, gasifier has to be work at approximately 2.5 barg. Because of that, 2.5 barg pressure level was selected as operating pressure in this study and gasifier was tested independently from the other processes. The capacity of the system is relatively higher than that of most of the research published in literature. So, the test results of pilot system may provide valuable insights into the process and operation of larger scale gasification systems.
2.2. Experimental setup The schematic of the pilot scale pressurized bubbling fluidized bed gasifier system is shown in Fig. 1. The system consists of BFB reactor, cyclone separator, syngas cooler, quencher, syngas burner, fuel feeding system, ash removal system and gasification agent (air, steam, O2, CO2) preheating systems. The BFB reactor is 3500 mm in high. Inner diameters of dense bed and freeboard regions of the reactor are 300 mm and 450 mm respectively. There is a cyclone separator with 140 mm inner diameter. Just after the cyclone separator, there is a syngas cooler for cooling down the syngas temperature from approximately 800 °C to 300 °C. The final stage of the system is quencher which is used for cooling syngas to approximately 30–40 °C with spraying the cooling water. An air compressor was used for the supply of pressurized air. The flowrate of the air was measured and controlled with a mass flow controller (MFC). Oxygen and CO2 gasses were supplied by two different gas storage tanks. A steam generator was used for the steam supply at the saturated conditions (at the pressure of 6 barg and the temperature of 165 °C). In the BFB reactor, the mixture of fluidization gas passed through the windbox and the distributor plate to the reactor. The distributor plate was made out of stainless steel and has 100 standpipes with 4 nozzles which are 1 mm in diameter. There were ten temperature measurement points located on the system in order to measure the temperatures of ambient air, 1st heater exit, 2nd heater exit, lower dense bed, upper dense bed, lower freeboard, upper freeboard, cyclone exit, heat exchanger exit and quencher exit, respectively. Pressure measurement device were also located on the same points with the temperature sensors. The reactor pressure was controlled by the pressure control valve located at the syngas exit just before the syngas burner. The system was monitored and controlled with the distributed control system (DCS). The gas analyzer was used for measuring the syngas composition. SIEMENS Process Chromatography Maxum Edition II was used to analyze syngas composition. It measures H2, O2, CO, CO2, N2, CH4 and C6+ with a specific Thermal Conductivity Detector (TCD). 2.3. Experimental method 2.3.1. Startup of the gasification reactor In the gasification experiment, 80 kg bed material (silica sand) was loaded to the reactor which had the initial static bed height of approximately 530 mm. Air flowrate was adjusted in order to get 0.8 m/s superficial gas velocity at the dense bed region of gasifier at the atmospheric conditions (20 °C, 0 barg). At this condition, the bed material was in bubbling regime. In order to heat up the bed material to the ignition temperature of coal, the electric heater was used for preheating the fluidization air. There were 2 groups of electric heaters; each consists of three serial connected 20 kW electric heaters. With these electric heaters, the fluidized bed temperature was increased to 350 °C which was the practical limit of
2. Experimental 2.1. Materials Silica sand was used as bed material which had size distribution between 0.3 mm and 0.7 mm and weighted arithmetic mean diameter of 644 mm according to the sieve analysis. Bulk and particle densities were 1450 kg/m3 and 2545 kg/m3, respectively. Turkish lignite (Soma lignite) was used as feedstock which had the LHV
Table 1 Proximate and elemental analysis of selected Turkish coal (mass basis). Proximate analysis
Elemental analysis
Moisture (%)
Volatile matter (%)
Fixed carbon (%)
Ash (%)
6.6
38.15
42.6
12.65
C (%) 60.56
H (%) 3.89
O (%) 14.32
S (%) 0.92
N (%) 1.06
H2O (%) 6.60
Ash (%) 12.65
LHV (kcal/kg) 5195
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Fig. 1. Schematic of gasification system.
ignition of coal particles. After reaching the temperature of 350 °C in the bed, coal was started to feed (approximately 10 kg/h) and combustion reactions started and the temperature increased very rapidly to 870 °C. Just after the reaching steady state combustion condition, the fuel feeding rate was approximately increased 3 times (30 kg/h) in order to get the gasification conditions at the constant air flowrate. The temperature slightly increased during this process and then decreased to its initial value. After reaching steady state gasification conditions, fine tuning of temperature was done with adjusting the fuel flowrate. 2.3.2. Startup of syngas cooler Before heating of bed material, thermal oil cycle at syngas cooler (shell&tube) was started in order to increase the temperature of thermal oil which was passing through the tube side of syngas cooler. At the steady state conditions, thermal oil temperature must be approximately 200 °C due to prevent the heavy hydrocarbon condensation on the surface of tube of heat exchanger. The temperature of thermal oil was held constant at desired value with adjusting the level of boiling water tank which was used for cooling the thermal oil coil immersed to the boiling water at atmospheric conditions.
constant automatically at desired level. The quencher was used not only decreasing the temperature of the syngas, but also decreasing the particle and heavy hydrocarbon contents in the syngas in order to protect pressure control valve located at the end of the process. 2.3.4. Pressurizing gasifier Gasifier pressure was held at desired value with pressure control valve which has PID controller. The desired value of gasifier pressure was set and PID controller of the control valve adjusted the pressure value and maintained the gasifier pressure at set point. As can be seen at Fig. 1, pressure control valve was located just after the syngas quencher and before the syngas burner. Gasifier pressure was increased gradually with 0.1 bar increments with pressure control valve. During the pressurization process, actual volumetric flow of fluidization agent was decreased. In order to held constant superficial gas velocity at the dense bed region, the gasification agent must be increased. The mass flowrate of coal also increased to operate the gasifier at the same temperature. At each 0.1 bar increments, it was waited to get stable operational conditions (especially in temperature levels) in gasifier. 3. Results and discussions
2.3.3. Startup of quencher After the reaching the steady state atmospheric gasification conditions, quencher pumps were started to feed 1000 kg/h water to quencher with two nozzles (totally 2000 kg/h) at different heights of quencher. The water level at the bottom of the quencher was controlled with the control valve and the water level was held
3.1. Operating conditions of experiment 3.1.1. Temperatures and pressures variations during the operation In this experiment, the continuous operation period of pressurized BFB gasifier system was 80 h. The time variation of
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temperature in the dense bed is shown in Fig. 2. As shown in this figure, the temperature of bed was fluctuating between 870 °C and 880 °C. The possible reason of these fluctuations could be the changes of coal properties during the operation. There were two different temperature measurement points for dense bed. One was located at the bottom of the dense bed which is shown in Fig. 2, and the other one was at the top of the dense bed. During the operation of gasifier, it was observed that the average temperature difference between the bottom and the top of the bed was 0.5 °C and the maximum difference was 1.6 °C which means the fluidization and the homogeneity of the bed was very good. The temperature profile of the system from the beginning of the process to the end is shown in Fig. 3. During the experiment, the average of the ambient air temperature was approximately 22 °C. Gasification agent passed through the 1st and 2nd electric heater and temperature was increased to the 192 °C and 291 °C, respectively. Preheated agent entered to the gasifier and the bottom of the dense bed temperature was 873 °C. As mentioned previously, top of the bed was approximately same (873.5 °C) with the bottom temperature. The next two temperature measurement points were lower freeboard and upper freeboard temperatures which were 872 °C and 832 °C, respectively.
Syngas temperature decreased from 832 °C to 738 °C while passing through the cyclone separator due to the heat losses. After the cyclone separator, syngas reached to the syngas cooler which was shell&tube heat exchanger and syngas passed through the shell side. Temperature was decreased from 738 °C to 434 °C and flew through the quencher. At the quencher, 2 1000 kg/h of water was sprayed to the quencher column with two nozzles (1000 kg/h for each). Temperature of syngas was decreased from 434 °C to 37 °C while passing through the quencher. The majority of the moisture content of syngas was taken at quencher. During the operation, the pressure of gasifier was in between 2.4 barg and 2.7 barg with little fluctuations, as shown in Fig. 4. The pressure was tried to be held constant at desired value. There were some fluctuations due to the fluctuations of syngas flow rate or some resistance/blockage through the process. The averages of pressure values through the process are shown in Fig. 5. 1st pressure measurement point was at the exit of 1st electric heater and its value was 2.92 barg. The pressure difference between the 2nd heater exit and lower dense bed was approximately 40 mbar which was the pressure loss at the distributor plate. The pressure loss due to the dense bed was approximately
Fig. 2. Temperature of the dense bed.
Fig. 4. Gasifier reactor pressures during the operation.
Fig. 3. Temperature profile through the process.
Fig. 5. Pressure profile through the process.
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80 mbar which was the difference between the pressure of lower dense bed and lower freeboard. At freeboard region, there was almost no pressure loss due to the very low solid entrainment and low solid concentration. Cyclone pressure loss was approximately 6 mbar which was the difference between upper freeboard and cyclone exit pressures. The pressure losses of syngas cooler and quencher were 5 mbar and 30 mbar, respectively. The pressure just before the pressure control valve was 2.74 barg and syngas was discharged to the atmosphere by pressure control valve. 3.1.2. Feeding rate changes of coal and gasification agents during the operation The flowrates of coal, steam, air, oxygen and CO2 during the experiment are shown in Fig. 6. As shown in this figure, the air flowrate was held constant at 10 Nm3/h. Air was fed to the gasifier in order to make nitrogen balance between inputs and syngas in order to determine the flowrate of syngas. Oxygen flowrate was changed in order to control the temperature and ER value, between 21 Nm3/h and 28 Nm3/h. CO2 was fed to the gasifier in order to determine the effect on gasifier performance. CO2 addition was tested for 4 different flowrates (25, 22, 35, 30 Nm3/h). In addition, H2/CO ratio could be also controlled at acceptable level with CO2 addition. Significant change in CO2 flowrate was occurred at 53th, 62nd and 71st hours. Coal feed rate during the experiment was between 80 kg/h and 83 kg/h, which correspond to 500 kWth and 520 kWth, respectively. Steam was mainly fed to the gasifier in order to get required actual volumetric flowrate for proper superficial velocity (0.8 m/s) through the dense bed. During the experiment, the effect of steam to carbon ratio was also investigated with changing the flowrate of steam between 100 kg/h and 140 kg/h. Considerable change in steam flowrate was occurred at 27th, 33th and 65th hour. 3.1.3. The variation of ER and steam to carbon ratio during the operation The equivalence ratio (ER) during the experiment was changed between 0.25 and 0.33, as shown in Fig. 7. In fact, the ER and the temperature of the reactor are coupled operation parameters. In this study, temperature of the reactor was selected as control parameter and ER value was considered as the function of the reactor temperature. Another important parameter for the operation of the gasifier is steam to carbon ratio which is shown in Fig. 8. Steam/carbon ratio
Fig. 7. ER value during the experiment.
Fig. 8. Steam to carbon ratio.
(S/C) was changed between 2.0 and 2.8 during the experiment. The system requires more fluidization agent for proper fluidization due to the pressurized operation of the gasifier. In this experiment, majority of the fluidization agent came from superheated steam. Because of that, S/C ratio for this experiment was over than 2. 3.2. Syngas composition
Fig. 6. Gasification agent flowrates.
Syngas composition was measured periodically during the experiment with 5 min periods with online gas analyzer. As shown in Fig. 9, CO2, H2, CO, N2 and CH4 were measured. Syngas CO2 and H2 contents were varied between 33.3–37.8% and 31.5–34.2%, respectively, which depended on the operation conditions. Considerable changes in percentage of the CO2 content were numbered in Fig. 9. At point 1, CO2 concentration increased while H2 was approximately constant. At that point, both ER and steam/carbon ratios(S/C) were increased together. CO2 content was proportional to the both ER and S/C ratios. Because with the increase of ER value which means also increase of oxygen content in gasification agent, combustion reaction increased and CO2 formation increased. Similarly, when the S/C ratio increased, water gas shift reaction was promoted and CO2 content increased. H2 content was also proportional to the S/C but inversely proportional to the ER value. In this situation, due to the increase of both ER and S/C, CO2 increased and H2 was not changed. At point
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Fig. 9. Syngas composition during the experiment. Fig. 10. Syngas flowrate during the experiment.
2, CO2 content increased while H2 content decreased due to the increase of ER value without the increase of S/C ratio. At point 3, 4 and 5, CO2 flowrate at gasification inlet was significantly decreased, increased and decreased, respectively. There were two reasons for this situation. First, with the increase or decrease of CO2 flowrate at the inlet of the gasifier, unreacted CO2 content in the syngas increased and decreased, respectively. Second, the shift reaction was affected by the CO2 content at gasification agent. If the CO2 content at gasification agent was increased, the shift reaction equilibrium conditions were changed and products of reaction (H2 and CO2) were decreased. The CO content of syngas was varied between 18.6 and 22.0%. Considerable changes in percentage of the CO contents were also numbered in Fig. 9. At point 6 and 7, CO content increased and decreased, respectively. While the change of CO percentage, S/C ratio was decreased and increased for point 6 and 7, respectively. S/C ratio promoted the shift reaction and CO was decreasing. The N2 and CH4 contents of syngas were varied between 6.7– 7.8% and 4.0–4.6%, respectively. N2 content was approximately constant during the experiment. CH4 content very slowly increased from 4.0% to the 4.6% at the end of the experiment. Syngas C6 content was also measured and varied between 0.28% and 0.46%. At first half of the experiment; C6 content was lower than the second half of the experiment which was related to the S/C ratio. S/C ratio was also lower at first half and higher at second half. Syngas flowrate was calculated with the N2 balance between input N2 (coming from air and coal) and syngas N2 content. Nitrogen content in coal fed to the gasifier can be mainly released as N2, NH3 and HCN. Small amount of fuel N is not released and stay in ash [19]. During the experiment, NH3 and HCN content were not measured. Due to this, it was assumed that fuel N was released as N2 and corresponded to approximately 0.7 Nm3/h N2 content in the syngas. Flowrate of syngas during the experiment is shown in Fig. 10. As shown in this figure, the syngas flowrate varied between 108.2 Nm3/h and 129.0 Nm3/h. The calculation of syngas flowrate was very sensitive to the change in N2 content of syngas. Because of that, the small fluctuations in measurement of N2 content of syngas causes very big fluctuations at syngas flowrate calculations. The lower heating value (LHV) of syngas was calculated according to the syngas composition, as shown in Fig. 11. The LHV of syngas varied between 8.2 MJ/Nm3 and 8.6 MJ/Nm3. These values were higher than that of the air gasification of coal (3–5 MJ/ Nm3) and lower than that of the oxy-steam gasification (9–12 MJ/Nm3). In this experiment, there were some differences in LHVs comparing with that of the pure oxy-steam gasification, because
Fig. 11. Syngas LHV (MJ/Nm3) during the experiment.
some amount of air and CO2 was also added to the gasification agent. Therefore, nitrogen from the air and excess CO2 content in the syngas resulted in a decrease in the LHV of the syngas. 3.3. Cold gas efficiency and carbon conversion The cold gas efficiency of gasification process during this experiment is shown in Fig. 12. As shown in this figure, cold gas efficiency was approximately 58.3% (average of the data). At first half of the experiment, efficiency was slightly higher than that of the second half due to the lower steam feed rate at first half. Steam addition resulted in an increase the mass of hot syngas and more combustion reaction take place at reactor in order to hold temperature constant. The increase in combustion reaction resulted in decrease of cold gas efficiency. Operation pressure also affected the efficiency of gasifier. When the operation pressure is higher than the atmospheric pressure, which was approximately 2.6 barg in this experiment, much more steam was fed to the gasifier than the required for chemical reactions in order to get required fluidization velocity at dense bed. In fact there is a limit of coal conversion load (throughput) of bed which depends on design capacity, coal reactivity, coal particle size and temperature. For this pilot scale gasifier, the practical limit
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S. Gül et al. / Applied Thermal Engineering 130 (2018) 203–210 Table 2 Average values of carbon balance between coal and syngas (mass basıs). C fuel (kg/h)
CCO2 (kg/h)
CCO (kg/h)
CCH4 (kg/h)
CC6 (kg/h)
CSyngas (kg/h)
49.86
22.54
12.82
2.76
1.52
39.64
conversion, the ratio of carbon content of syngas to the fresh carbon was calculated. Syngas carbon content was calculated with the syngas composition which includes CO2, CO, CH4 and C6 molecules (Eq. (2)). Average values of carbon balance between coal and syngas is shown in Table 2. 4. Conclusions
Fig. 12. Cold gas efficiency of gasification process.
of bed coal conversion load observed during the experiments was approximately 11 MWth/m3. Over this limit of conversion load, the carbon concentration in fluidized bed which was 8.2% for this experiment is increased and the homogeneity of bed was broken and agglomeration problem occurs. Due to this reason, when pressurizing the gasifier, the coal feed rate and the gasification agent flowrate should not be increased proportionally. The coal feed rate approximately held constant and the steam flowrate was increased while pressurizing process. Another important performance criterion for gasification process is carbon conversion ratio. Carbon conversion calculated with the carbon balance between coal and syngas is shown in Fig. 13. As shown in figure, carbon conversion varied between 72.0% and 84.6% (average value 80.3%). Majority of carbon was lost with fly ash which the major part was captured by the cyclone. Carbon balance between streams was determined with the equations below (mass basis). Carbon flowrate fed to the gasifier with coal must be equal to the total amount of carbon in the streams of syngas, bottom ash and fly ash (Eq. (1)).
Cfuel ¼ Csyngas þ Cbottom ash þ Cfly ash
ð1Þ
Csyngas ¼ CCO2 þ CCO þ CCH4 þ CC6þ
ð2Þ
During the experiment, the amount of carbon in the bottom ash and fly ash were not determined. In order to determine the carbon
The present work attempted to analyze the effect of some operational parameters on pilot scale pressurized bubbling fluidized bed gasification system. The properties and composition of the synthetic gas produced by gasification process depend on many factors of design and operational parameter such as the reactor type, the properties and flow rates of feedstock materials and gasifying agents, pressure and temperature in the reactor. The following conclusions can be drawn from the experimental results. – With the increase or decrease of CO2 flowrate at the inlet of the gasifier, unreacted CO2 content in the syngas increased and decreased, respectively. Shift reaction was also affected by the CO2 content at gasification agent. If the CO2 content at gasification agent was increased, shift reaction equilibrium conditions were changed and products of reaction (H2 and CO2) decreased. – Syngas CO content was varied between 18.6 and 22.0%. While the increase and decrease of S/C ratio, CO percentage of syngas decreased and increased, respectively. S/C ratio promoted the shift reaction and CO concentration decreased. – Syngas LHV was varied between 8.2 MJ/Nm3 and 8.6 MJ/Nm3. Different from the oxy-steam gasification, some air and CO2 was added to the gasification agent. Because of nitrogen content from the air and excess CO2 content in the syngas decreased the syngas LHV. – Cold gas efficiency average value was approximately 58.3%. At first half of the experiment, efficiency was slightly higher than the second half due to the lower steam feed rate at first half. Steam addition increased the mass of hot syngas and more combustion reaction took place at reactor in order to held temperature constant. Increase in combustion reaction resulted in decrease of cold gas efficiency. – Carbon conversion calculated with the carbon balance between coal and syngas was varied between 72.0% and 84.6%. Majority of carbon was lost with flyash.
Acknowledgments This study has been realized within the project entitled ‘‘Liquid Fuel Production from Biomass and Coal Blends”. TUBITAK is greatly acknowledged for the support of this project under the ‘‘Support Programme for Research Projects of Public Institutions” (the contract number 108G043). References
Fig. 13. Carbon conversion of gasification process.
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