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Gasification of high-ash Indian coal in bubbling fluidized bed using air and steam – An experimental study
Applied Thermal Engineering 116 (2017) 372–381
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Research Paper
Gasification of high-ash Indian coal in bubbling fluidized bed using air and steam – An experimental study K. Vijay Kumar a, M. Bharath a, Vasudevan Raghavan a,⇑, B.V.S.S.S. Prasad a, S.R. Chakravarthy b, T. Sundararajan a a b
Department of Mechanical Engineering, Indian Institute of Technology Madras, Chennai 600036, India Department of Aerospace Engineering, Indian Institute of Technology Madras, Chennai 600036, India
h i g h l i g h t s Bubbling bed gasification of high-ash Indian coals is systematically studied. Design, fabrication and installations details are reported. Effects of air to coal and steam to coal ratios studied in detail. Carbon conversion, synthetic gas composition and cold gas efficiency are presented for all cases.
a r t i c l e
i n f o
Article history: Received 18 November 2016 Revised 21 January 2017 Accepted 26 January 2017 Available online 30 January 2017 Keywords: Coal gasification High-ash Indian coals Bubbling fluidized bed Steam-coal ratio Carbon conversion Cold gas efficiency
a b s t r a c t In this experimental study, gasification characteristics of Indian coals having high ash content (>35%) in a bubbling fluidized bed reactor, using air and steam as the gasification agents under atmospheric pressure, are presented. The thermal power output of the gasification device has been varied in the range of 80– 100 kW. The ratio of the actual mass flow rate of air supplied to the stoichiometric air mass flow rate (required for complete combustion) has been varied in the range of 0.25–0.35. The ratio of mass flow rate of steam to that of coal has been varied from 0.15 to 0.25. Effects of air and steam flow rates on the calorific value of product gas, carbon conversion and cold gas efficiency have been studied in detail. For highash Indian coals, supply of steam produces positive effect as all the metrics such as calorific value of product gas, carbon conversion and cold gas efficiency increases with increase in the steam supply rate. Ó 2017 Elsevier Ltd. All rights reserved.
1. Introduction India has a lot of coal reserves – ranks third following China and USA. About 61% of the total energy requirement of India is met by coal-based power plants. Indian coals have good amount of fixed carbon and volatiles (around 30% each) and very little sulphur (less than 1%). However, Indian coals have large ash content, due to which, burning or gasification of Indian coals poses lots of problems in terms of ash disposal, combustion or gasification efficiency and carbon conversion. In particular, pulverized burning of Indian coals, which is carried out in several power plants, leads to significant amount of particulate emission and a costly after treatment and techniques to capture fine ash particles are required. Gasification is an important process in clean coal technology such as Inte-
⇑ Corresponding author. E-mail address: [email protected] (V. Raghavan). http://dx.doi.org/10.1016/j.applthermaleng.2017.01.102 1359-4311/Ó 2017 Elsevier Ltd. All rights reserved.
grated Gasification Combined Cycle (IGCC). In general, gasification process leads to efficient recovery of combustible gas mixtures (called synthetic gas) from coal (or any other solid fuels such as biomass), paving way to better cleaning and ash handling methods. First pioneering research work with respect to gasification of Indian coals has been reported by Iyengar and Haque [1]. They systematically investigated the gasification process of high-ash Indian coal employing lab-scale experimental setups. They used different types of gasification devices such as moving bed, bubbling fluidized bed and entrained bed. They performed the experiments using four types of Indian based coals, in which the ash content varied from around 12–31% on weight basis. They observed that the calorific value of the synthetic gas diminishes as the ash content in the coal increased. They also concluded that the adverse effect of high ash content in coal is minimally detected in moving bed gasification system followed bubbling bed and entrained flow gasification devices. Mathematical model of gasification process in
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moving bed using Indian high-ash coals was reported by Singh et al. [2]. In this study, the effect of addition of steam was brought out clearly. Datta et al. [3] studied the agglomeration characteristics of high-ash Indian coals in fluidized beds. Cause-effect relationship with respect to agglomerate formation with high ash Indian coals was reported. Recently, Jayaraman et al. [4] reported the production of char from high-ash Turkish and Indian coals using CO2 – steam agents in high heating rate Thermo Gravimetric Analyzer (TGA). Engelbrecht et al. [5] reported gasification of South African coals, which also have high ash content. They used a bubbling fluidized bed device for gasification. Watkinson et al. [6] carried out experiments with three types of Canadian coals in a fluidized and fixed bed reactors, with steam and air as gasification agents. They obtained similar gas heating values, gas yields and thermal efficiencies for both reactors. Lee et al. [7] studied the gasification on Australian coal in a fluidized bed with air and steam mixtures at atmospheric pressure. The cold gas efficiency, gas heating values and gas yields increased with an increase in the steam to coal ratio, however, the carbon conversion was almost unchanged. Gutierrez and Watkinson [8] analyzed the gasification of Canadian coal in a fluidized bed reactor with air and steam mixtures at atmospheric pressure. Ocampo et al. [9] studied the gasification characteristics of Colombian coal in a fluidized bed reactor at different air to coal and steam to coal ratios at atmospheric pressure. In these studies parameters such as cold gas efficiency, gas heating value and carbon conversion were reported. Xiao et al. [10] carried out the gasification of Chinese coal in laboratory and pilot scale reactors. In that study, the major differences in cold gas efficiency, gas heating value and carbon conversion were observed between laboratory and pilot scale reactors due to the heat loss and differences in the hydrodynamics conditions. Relatively low temperature operation is possible in a bubbling fluidized bed when compared to that of a moving bed. Based on these studies, it is clear that either a moving bed (also called fixed bed) gasification device or a slow bubbling fluidized bed gasification device can serve the purpose of attaining good gasification using Indian coals. Ash removal is little bit complicated in fixed beds, which handle large chunks of coal particles (6– 50 mm). Due to high temperature at the bottom of the reactor, ash may melt and provision for molten ash removal is required. Similarly, the fluidized bed gasification technology is also highly challenging for the Indian coals because of the issues such as improper fluidization, sudden raise in temperature due to ash agglomeration, carbon loss by bed drain, entrainment of fine carbon particles in the fly ash and treatment of cohesive coal in fly ash. These have been discussed in literature [6–14], however, not specific to high-ash coals. In summary, it is required to fix the regime of fluidization carefully in order to avoid the adverse effects of the ash content in the coal on both fluidization and gasification aspects. Therefore, in this study, a relatively slow bubbling bed gasification reactor has been designed, fabricated and experimented with Indian high-ash coals. The calorific value of product gas, carbon conversion and cold gas efficiencies have been systematically evaluated for wide range of air to coal and steam to coal ratios. The design procedure, dimensions of the reactor, experimental procedure and results have been discussed subsequently.
bigger initial particle sizes have been employed. The results from proximate and ultimate analyzes of Indian coal used in this study is reported in Table 1. The higher calorific valve of the coal is calculated by using Dulong and Petit equation, given as,
O þ 9418 S kJ=kg; HHV of coal ¼ 33823 C þ 144249 H 8 where C, H, O and S, are the mass percentages of carbon, hydrogen, oxygen and sulphur, respectively, and are obtained from ultimate analysis (Table 1). HHV is evaluated as 17.76 MJ/kg. Assuming a thermal efficiency of around 75%, the coal feed rate has been calculated for a given power rating. Air flow rate has been fixed as a percentage of stoichiometric air required for completely burning the coal. Equivalence ratio (ER), which is the ratio of the actual mass flow rate of air supplied to the stoichiometric air mass flow rate (required for complete combustion) has been varied in the range of 0.25–0.35. The ratio of mass flow rate of steam to that of coal has been varied from 0.15 to 0.25. In the cases where steam has been supplied the equivalence ratio has been kept constant as 0.25. Table 2 reports the parameters varied in the present experimental study. For all the cases, a regime corresponding to that of bubbling fluidized bed has been considered. The hydrodynamic calculations have been carried out to establish the dimensions, particle sizes and operating conditions of the gasification device [15,16]. The minimum fluidization velocity (Umf), superficial velocity (Uo), transport disengaging height (TDH) and total height have been calculated for different cases for a range of average particle sizes. The minimum fluidization velocity (Umf) is calculated using the procedure given in Kunii and Levenspiel [16] and transport disengaging height (TDH) is calculated using the procedure given in Basu [15]. Superficial gas velocities are calculated theoretically at the experimental operating conditions (1 atm, 900 °C). Table 3 reports these quantities for different cases. The diameter of the gasification device has been fixed as 0.15 m (150 mm). Based on the hydrodynamic calculations, for bubbling fluidization regime and for the range of power rating, the height, particle sizes and superficial velocities have been fixed. Superficial velocity (Uo) is 1–1.89 times larger than minimum fluidization velocity (Umf). For example, it has been observed that for 100 kWth case and for ER of 0.35, the value of TDH is around 2.07 m, considering an average particle size of 2.5 mm, which is the maximum among all the cases. Based on this, the total height of the reactor has been fixed as 2.5 m. The bed height is varied based on the air and steam flow rates.
3. Experimental setup The reactor has been fabricated in three portions, namely, the top portion having a height of 0.5 m, the middle portion with a height of 1 m and the bottom portion of 1 m height. The exit of top portion is welded with a flange, which facilitates the connection of pressure regulator, pipes for gas scrubber and synthetic gas burner. The gasification unit has been fabricated using two cylindrical refractory linings each having a thickness of 100 mm.
Table 1 Analysis of Indian coal (air dried).
2. Design procedure The thermal power rating of bubbling fluidized bed gasification reactor has been fixed initially at 100 kW based on the choice of regime of fluidization and particle size. However, after the first test, noting a low carbon conversion due to carbon loss from flue gases, the thermal power rating has been reduced to 80 kW and
Proximate analysis
wt (%)
Ultimate analysis
wt (%)
Ash content Moisture Volatiles Fixed carbon
36.4 5.4 28.7 29.5
Carbon Hydrogen Nitrogen Oxygen Sulphur
43.7 3.8 0.9 14.5 0.7
Calorific value (MJ/kg)
17.76
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Table 2 Experimental tests. Test No.
Power output (kWth)
Coal feed rate (kg/h)
Air flow rate (kg/h)
ER
Steam flow rate (kg/h)
Steam/coal ratio (S/C)
1 2 3 4 5 6
100 80 80 80 80 80
20.2 16.2 16.2 16.2 16.2 16.2
40.6 32.5 27.8 23.2 23.2 23.2
0.35 0.35 0.30 0.25 0.25 0.25
– – – 2.4 3.2 4.0
– – – 0.15 0.20 0.25
Table 3 Hydrodynamic calculations. Umf (m/s)
d*p
U*
TDH (m)
U0 (m/s)
Dense bed (m)
Total height (m)
Test run 1: 100 kWth, ER = 0.35 3.0 28 2.5 18.3
1.42 1.12
38.35 31.96
1.09 1.09
1.59 2.07
2.12 2.12
0.3 0.3
1.89 2.37
Test run 2: 80 kWth, ER = 0.35 3.0 28 2.5 18.3
1.42 1.12
38.35 31.96
0.86 0.86
0.96 1.40
1.70 1.70
0.3 0.3
1.26 1.70
Test run 3: 80 kWth, ER = 0.3 3.0 28 2.5 18.3
1.42 1.12
38.35 31.96
0.74 0.74
0.53 1.04
1.45 1.45
0.3 0.3
0.83 1.34
Test run 4: 80 kWth, ER = 0.25, S/C = 0.15 3.0 28 1.42 2.5 18.3 1.12
38.35 31.96
0.72 0.72
0.33 0.99
1.41 1.41
0.3 0.3
0.63 1.29
Test run 5: 80 kWth, ER = 0.25, S/C = 0.20 3.0 28 1.42 2.5 18.3 1.12
38.35 31.96
0.76 0.76
0.61 1.08
1.48 1.48
0.3 0.3
0.91 1.38
Test run 6: 80 kWth, ER = 0.25, S/C = 0.25 3.0 28 1.42 2.5 18.3 1.12
38.35 31.96
0.79 0.79
0.75 1.19
1.55 1.55
0.3 0.3
1.05 1.49
dp (mm)
Re
The refractory linings have different thermal conductivities as listed in Table 4. A mild steel shielding of 10 mm thickness has been provided on the outer surface of the second refractory to provide strength to the column. The entire experimental setup is shown in Fig. 1. Distributor plate has a conical section with an angle of 45°. Top internal diameter of a cone is 0.15 m and the bottom internal diameter is 0.04 m. It is made of stainless steel 310 grade and has a thickness of 30 mm. It has several perpendicular holes of 2 mm diameter, separated by a pitch distance of 4 mm. Distributor plate is placed in between plenum (wind box) and bottom portion of the gasifier. Coal particles from a hopper is fed to the reactor at a height of 0.5 m above the distributor plate. A screw conveyor is used to feed the coal. Pitch size, diameter and length of screw conveyor to feed coal are 35 mm, 50 mm and 650 mm, respectively. The feed rate is controlled by a variable speed DC motor. A water jacket is provided outside the screw conveyor in order to cool it sufficiently. Seven thermocouples have been provided at different heights (Table 5) in order to measure the axial temperature distribution inside the reactor.
Based on the test case, either pre-heated air alone or pre-heated air and steam are injected into the plenum (wind box). The gasification agent enters the reactor through the holes in the conical distributor plate and bubble through the bed material and coal particles. Initially the bed material (sand particles) is filled up in the pipe below the distributor plate, leading to the ash collector. After providing sufficient time for conversion of the coal particles, bed ash is discharged into the ash collector by opening the valve connected to the pipe below the plenum. Fine fly-ash particles in the flue gas are removed in the filter unit as well as in the water scrubber. Filter box of size 350 mm inner diameter and 1 m long is made up of stainless steel. Three layers of meshes, with holes of sizes 2 mm, 1 mm and 0.5 mm, respectively, are placed inside the filter unit at different heights. The synthetic gas passes through these meshes and the particulate material is collected at the bottom of the filter unit. The synthetic gas is then allowed to pass through a scrubber, which is filled with water to cool the flue gas as well as to separate the left-over fine fly ash contained in the gas. At regular time intervals, small amount of gas from the exit pipe attached to the water scrubber is taken into a gas analyzer and its
Table 4 Properties of refractory linings. Properties of refractory
Material composition
%
Thermal conductivity (W/m K)
Maximum service temperature (°C)
Hot face refractory (self flow casting having low iron)
Al2O3 CaO Fe2O3
60.8 3.45 0.77
1.32
1500
Cold face refractory (high purity insulating castable)
Al2O3 SiO2 CaO Fe2O3
46.1 22.6 7.72 1.33
0.31
1300
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Fig. 1. Schematic of assembled experimental set up along with instrumentation.
Air pre-heater, shown in Fig. 2(a), is a specially designed heat exchanger that utilizes a premixed LPG-air burner. It is made up
Table 5 Location of the thermocouples. S. No.
Location of thermocouples measured from distributor plate (m)
1 2 3 4 5 6 7
0.2 0.5 0.8 1.2 1.5 1.8 2.4
composition is determined. The synthetic gas is burned in a syngas burner, and the flue gases are left to the atmosphere. The volumetric flow rate of the flue gas is measured by passing it through a precalibrated rotameter. Gas analyzer measures the volumetric percentages of components present in the synthetic gas such as CO, CO2, O2, CH4 and H2. The gas analyzer has been specially calibrated to measure the product of gasification, where high amounts of CO and H2 are present. The volume fractions can be measured with an uncertainty of 0.1%. The range of volumetric percentage of each component of synthetic gas and the type of sensors used in the gas analyzer are listed in Table 6.
(a) Air pre-heater
Table 6 Range of composition and sensor type in gas analyzer. S. No.
Gas
Range (% V/V)
Sensor type
1 2 3 4 5
CO CO2 O2 CH4 H2
0–50 0–50 0–30 0–20 0–40
Non-dispersion infrared Non-dispersion infrared Electrochemical Non-dispersion infrared Thermal conductivity
(b) Steam generator Fig. 2. Schematics of (a) air pre-heater and (b) steam generator.
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of stainless steel. Air flows through a long stainless steel tube arranged in a spiral form as shown in Fig. 2(a). The inner diameter of the air tube is 19 mm and its length is 11 m. This arrangement is kept in a shell, on top of which, the Liquefied Petroleum Gas (LPG) burner is placed. The products of combustion of LPG-air flows from the top towards the bottom, passing through the coiled air tube. The capacity of the burner is 50 kW. The LPG is supplied continuously to LPG-air burner at a rate of 0.068 kg/h. A refractory block is placed below the burner to guide the hot gases in proper direction. Air heater is designed to heat the required flow rate of air up to 500 °C. Steam generator, as shown in Fig. 2(b), consists of a boiler unit at the bottom of a shell and a super-heater unit having bank of tubes placed above the shell. An inverted LPG-air burner, similar to that used in the air pre-heater, is placed in between the boiler and super heater units. Stainless steel 304 is used for fabrication. Super heater unit consists of 40 tubes, each of length 200 mm and inner diameter 19 mm. Water enters into the boiler around the walls in the bottom portion and evaporates due to the heat transfer from the flame impinging surfaces in the bottom portion. The saturated steam passes through the super heater unit for further heating by the raising hot flue gases. The maximum flow rate in the steam generator is 10 kg/h with a temperature up to 400 °C at atmospheric pressure.
4. Experimental procedure Coal is pulverized and the particles are sieved using standard sieves. The maximum and minimum sizes of that sieve are reported as particle size range. In order to achieve ignition and to preheat the bed, the gasification unit has been initially filled with 5 kg of bed material (sand particles of size 1–1.5 mm) and 3 kg of fine coal particles (1–2 mm). Air at a rate of 20.8 kg/h has been supplied through the plenum to achieve proper fluidization of solid particles. LPG has been injected into the gasifier at a height of around 0.2 m above the distributor plate. The mixture of LPG and air has been ignited using a pilot flame. LPG flame rapidly heats up the refractory and the bed materials. The coal particles have been subsequently ignited. When the bed temperature, as measured by the thermocouple T1 reaches around 600 °C, LPG injection has been stopped. Combustion of coal continues the raise the bed temperature. When the temperature reaches around 1050 °C, coal particles of size 2–3 mm have been fed using the screw conveyor. The bed temperature momentarily decreases during the initial addition of coal particles and subsequently increases to reach a steady state value. The rate of coal feeding has been slowly and gradually increased to the required value. Steady operating regime has been established within around 2 h. At regular time intervals, ash and bed material particles have been removed. The mass of the removed material has been measured using a load cell. It should be noted that the feeding of coal and gasification agent have been continuous processes, however, the ash removal has been a manual batch process. Care has been taken to remove the bed material and ash without quenching the bed as a result of over-extraction. In the test cases where steam has been also used along with air as a gasification agent, once stable conditions have been achieved, steam has been introduced into the reactor. The product gas has been sampled into the gas analyzer after the steady operation has reached and composition of synthetic gases have been recorded at regular intervals. The reactor has been operated for around 5 h such that steady operation has been achieved for at least 2.5–3 h in each test case. All the test cases have been repeated at least three times to ensure similar operating conditions and output. It has been observed that
the data has been repeatable within 1.5% around the average values. The bed ash and fly ash have been kept in furnace at around 950 °C for around 5 h and the mass loss has been recorded to estimate the amount of unburnt carbon present in these samples. Total and fixed carbon conversions have been calculated using the procedure reported by Engelbrecht et al. [5]. The formulae for calculating the fixed carbon conversion and total carbon conversion are given by
Fixed carbon conversion ¼
fða bÞ ½ðd eÞ þ ðf gÞg ; ða bÞ
Total carbon conversion ¼
fða cÞ ½ðd eÞ þ ðf gÞg ; ða cÞ
where a is coal feed rate in kg/h, b is the percentage of fixed carbon content in coal, c is the percentage of carbon content in coal, d is the flow rate of bed ash in kg/h, e is the percentage of carbon content in bed ash, f is the flow rate of filter ash in kg/h and g is the percentage of carbon content in filter ash. Cold gas efficiency is calculated using the formulae [15] given as 3
Cold gas efficiency ¼
Syn gas flow rate in m =s Gas HHV in MJ=m3 Coal feed rate in kg=s Coal HHV in MJ=kg 100;
Gas higher heating value (HHV) depends on product volume percentages (Vi) of CO, H2 and CH4 and their calorific values (CVi). It is calculated as
Gas HHV ¼
VCO CV CO þ VH2 CV H2 þ VCH4 CV CH4 100
5. Results and discussion The reaction temperature is one of the significant parameters that affects the performance of a gasification unit. The processes in gasification should be coupled in such a way that the exothermic reactions, which utilize oxygen and release heat, should transfer the heat to the endothermic reactions, which predominantly utilize gases such as CO2 and H2O. As mentioned earlier, the temperature in a fluidized bed reactor is almost uniform and is relatively lesser than that of a fixed or moving bed reactor. The maximum temperature of the gasification unit is usually maintained below the ash-melting temperature for easier operation and this become significant when the ash content in coal is high. The main chemical reactions in the coal gasification are summarized below [15,16]. 1. Oxidation reactions: C + 0.5O2 M CO2, C + O2 M CO2 and H2 + 0.5O2 M H2O (exothermic) 2. Boudouard reaction: C + CO2 M 2CO (endothermic) 3. Water gas reaction: C + H2O M CO + H2 (endothermic) 4. Methane formation reaction: C + 2H2 M CH4 (exothermic) 5. Water gas shift reaction: CO + H2O M CO2 + H2 (exothermic) 6. Steam methane reforming reaction: CH4 + H2O M CO + 3H2 (endothermic) 5.1. Test run-1–100 kWth In test run-1, coal feed rate has been calculated for a power output of 100 kWth and the air flow rate has been calculated considering an equivalence ratio of 0.35. No steam has been added in this case. The variation of temperature with time as measured by thermocouples located at different heights of the reactor is shown in Fig. 3 for the transient operation period and the same for steady operation period is shown in Fig. 4. Time = 0 s in Fig. 3 indicates
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Fig. 3. Transient temperature variation at different heights for test run 1.
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Fig. 5. Variation of synthetic gas composition with time for test run 1.
78 and 89% and for the coal with 31.7% ash, the total carbon conversion varied from 60 to 87%, depending upon the air and steam flow rates. The cold gas efficiency reported by Engelbrecht et al. [5] has been around 65.5% (maximum). 5.2. Test runs at 80 kWth power with air alone as gasification agent
Fig. 4. Variation of temperatures with time at different heights for test run 1 under steady operation regime.
the starting of the ignition process. Ignition and steady operating conditions have been established as discussed in the previous section. The average bed temperature, as measured by T1, under steady operating conditions (2.5–5 h) is around 922 °C. This value is comparable to that reported by Engelbrecht et al. [5], where two types of South African coals with ash content 40.7% and 31.7% have been utilized. The temperatures measured by T1, T2, T3 and T4 are close to each other with only around 70 °C drop until a height of 1.2 m. The uniform temperature distribution in this height range ensures that the operation is in bubbling bed regime and also indicates a good solid–gas mixing. However, above around 1.5 m height, the temperature decreases significantly to around 450 °C, indicating heat losses in freeboard. For the smaller scale reactors such as this, the heat losses are usually higher than the actual scaled up reactors. The amounts of unburned char present in fly ash and ash removed from the bed have determined by placing the collected samples in a furnace at 950 °C for 5 h. It is observed that the unburned char present in the fly ash is around 25.6% and that in the ash extracted from the bed is around 2%. The deteriorating effect of high-ash content (36.4%) is clearly observed for this case, as indicated by trapped char in fine ash particles (100–500 lm). Synthetic gas composition has been measured under steady operating conditions as shown in Fig. 5, where the average values obtained from three trial have been plotted. On dry basis, the volumetric percentages of CO, CO2, H2 and CH4, are 15.8, 14.5, 7.5 and 0.1 respectively. Total carbon conversion is around 74.3% and the cold gas efficiency is around 61.2%. Engelbrecht et al. [5] used steam in addition to air. Their gas composition showed higher percentages of hydrogen. They have reported that for coal with around 40.7% ash, the total carbon conversion has been between around
Based on the results obtained from 100 kWth run, in order to minimize the unburned char carried away by the fly ash, the feed size of the coal particles has been increased to 2.5–3 mm and the superficial velocity is reduced from 2.12 m/s to 1.7 m/s. For these conditions, the coal feed rate has also been decreased 16.2 kg/h, which corresponds to a power rating of 80 kWth, corresponding to an equivalence ratio of 0.35 (test run 2). This case has moved to a slower bubbling regime when compared to the previous case. The temporal variations of temperatures measured along the height of the gasification reactor for this case have been similar to that of the previous case. However, the average bed temperature, as measured by T1, under steady operating conditions (2.5–5 h) has been around 999 °C, which is higher than the previous case. The temperature in the bed has increased because of an increase in the average size of coal particles as well as due to the decrease in the velocity of air. There has been a phenomenal decrease in the amount of unburned char present in the fly ash (from 25.6 to 18.4%) and that in the ash removed from the bed (from 2 to 1%). For this case, the average volumetric percentages of CO2, CO and H2, have been around 18.6%, 15.3% and 6.7%, respectively. The total carbon conversion has increased from 74.3% to 82.4% and the cold gas efficiency has remained almost the same (around 59.4%, when compared to 61% in test run 1). In order to study the effect of ER, a test with a lower ER value has been conducted. The air flow rate is reduced such that ER is equal to 0.3 (Table 1 – test run 3). Coal feed rate in maintained to provide a power output of 80 kWth and the coal particles are fed with the size range of 2.5–3 mm as in the previous case. The superficial velocity for this case is around 1.45 m/s. Fig. 6 shows the temporal variations of temperatures at different heights under steady operation regime. It is observed that in the slow bubbling regime, the temperature gradient along the reactor height has increase when compared to the earlier cases. However, as a result of slow bubbling process, the average bed temperature, as measured by T1, under steady operating condition (2.5–5 h) has increased to around 1040 °C. In this slow bubbling regime, the unburned char in the fly-ash has further reduced from 18.4% to 16.8%. The average composition of synthetic gas under steady operating condition is shown in Fig. 7. It is observed that the volumetric percentages of CO2, CO and H2 are 17, 19.1 and 6.9, respectively. The concentration of CH4 is almost zero. The total carbon
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0.15–0.25 respectively. The regime for the test cases 4, 5 and 6 are very much similar to that of test cases 2 and 3. As mentioned earlier, the reactor is started only using air and after establishing steady operating condition, steam injection is started gradually. Fig. 8 presents the temporal variation of temperature at several locations along the reactor height under steady operating conditions for the case of steam to coal ratio of 0.15. It can be noted that the maximum temperature has slightly decreased (from 1040 °C to around 968 °C) as a result of steam addition. In this slow bubbling regime, there is a notable temperature gradient along the height of the gasifier. The unburned char present in the fly ash is around 18.3%, almost same as in the previous case. Fig. 9 presents the variation Fig. 6. Variation of temperatures at different heights with time for test run 3.
Fig. 8. Variation of temperatures at different heights with time for the test run 4. Fig. 7. Variation of average composition of synthetic gas with time for test run 3.
conversion has increased further to 84.1% for this test and the cold gas efficiency has increased to around 65.7%, comparable to that reported in Engelbrecht et al. [5]. The test run 3 clearly shows that there is an optimal operating regime for a given type of coal in a given reactor. Careful variation of coal and air flow rates and proper choice of size of the feed coal particles contribute significantly to the gasification performance. The deteriorating effects of high-ash content in the coal can be overcome by providing sufficient bubbling (fixing the regime properly) as well as providing enough residence time. As compared to a moving/fixed bed, the operating temperature is quite low in bubbling beds (1040 °C in slow bubbling regime, when compared to around 1350 °C in a moving bed), which is quite advantageous [1]. Fig. 9. Variation of the composition of the synthetic gas with time for test run 4.
5.3. Test runs at 80 kWth power with air and steam as gasification agents Addition of steam results in water-gas reactions. The gas-phase water-gas shift reaction is slightly exothermic and the solid phase water-gas reaction is endothermic. Water-gas shift reaction provides increased amounts of hydrogen in the synthetic gas. Therefore, like coal and air feed rates, the steam feed rate also affects the gasification process significantly. Keeping this in mind, three trials of tests 4, 5 and 6 have been conducted, where the steam flow rate has been increased gradually. Since steam is added, the ER value is decreased to 0.25 for these cases. The power rating and thus, the coal feed rate, has been maintained the same as in test runs 2 and 3. Also, the feed coal particles size has been in the range of 2.5–3 mm, as in the previous case. Steam flow rate is varied from 2.4 kg/h to 4 kg/h, for a coal flow rate of 16.2 kg/h. This corresponds to steam to coal ratio of
Fig. 10. Variation of temperatures at different heights with time for the test run 5.
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Fig. 11. Variation of the composition of the synthetic gas with time for test run 5.
Fig. 13. Variation of the composition of the synthetic gas with time for test run 6.
gas mixture. The total carbon conversion and cold gas efficiency for this test are 83.3% and 63.6%, respectively. In test run 5, keeping the same coal and air feed rates, and the feed coal particle size, as in the previous case, the steam flow rate has been increased to attain a steam to coal feed ratio of 0.2. Fig. 10 presents the temporal variation of temperature at several locations along the reactor height under steady operating conditions. A slight reduction in the bed temperature indicated by T1 is observed (958 °C as compared to the previous case of 968 °C). While the dense portion of the reactor show an almost steady temperature distribution, the freeboard portion shows slight transients as a result of changes undergone by the gas-phase reactions. The fly ash carbon content is almost the same at 18.1%, as in the previous case. Fig. 11 presents the volumetric fraction of CO, CO2 and H2 in the synthetic gas under steady operating conditions. There is a further increase in the volumetric fraction of hydrogen from 8.4% in the previous case to around 10.3% in the present case. Similarly, the volume fraction of CO has also increased from 18.5% to 20.6%. The volumetric fraction of methane has also increased slightly to 1%. With these, the calorific value of synthetic gas has further increased. While the total carbon conversion (83.4%) is almost the same as in the previous case, the cold gas efficiency has increased significantly from 63.6% to 70.6%. This case gives a scope of further increasing the steam flow rate as the maximum bed temperature is still higher than 950 °C.
Fig. 12. Variation of temperatures at different heights with time for the test run 6.
of the synthetic gas composition with time. The volumetric percentages of CO2, CO and H2, are 15.9, 18.5 and 8.4, respectively. An apparent increase in the volume fraction of hydrogen from 6.9% to 8.4% due to steam addition is observed. Also, methane formation reaction rate has increased due to increased availability of hydrogen. This is apparent from the volumetric percentage of CH4, which has increased to around 0.92%. The gas calorific value has increased due to the increase in the content of hydrogen in the
Table 7 Summary of fluidized bed gasification tests using high-ash Indian coal. Test number
1
2
3
4
5
6
Capacity (kWth) Coal feed rate (kg/h) Air flow rate (kg/h) Steam flow rate (kg/h) ER Air to coal ratio Steam to coal ratio Average temperature of air/steam at inlet (°C) Average lower bed temperature (°C) Coal particle size (mm)
100 20.2 40.6 – 0.35 2 – 300 922 2–3
80 16.2 32.5 – 0.3 1.73 – 340 999 2.5–3
80 16.2 27.8 – 0.3 1.48 – 340 1040 2.5–3
80 16.2 23.2 2.4 0.25 1.23 0.15 350 968 2.5–3
80 16.2 23.2 3.2 0.25 1.23 0.2 365 958 2.5–3
80 16.2 23.2 4.0 0.25 1.23 0.25 360 951 2.5–3
Dry gas composition (Volume %) CO H2 CH4 CO2 N2 + Other Gas calorific value (MJ/Nm3) Unburned carbon in fly ash (%) Fixed carbon conversion (%) Total carbon conversion (%) Cold gas efficiency (%)
15.8 7.5 0.1 14.5 62.13 2.84 25.6 61.95 74.26 61.2
15.3 6.7 0 18.6 59.4 2.99 18.4 73.98 82.39 59.4
19.2 6.9 0 17.0 56.9 3.14 16.8 76.51 84.1 65.7
18.5 8.4 0.9 15.9 56.3 3.58 18.3 75.21 83.23 63.6
20.6 10.3 1.0 14.2 53.9 4.1 18.1 75.52 83.43 70.6
19.4 15.4 1.6 12.4 51.2 4.81 17.5 76.43 84.05 77.7
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Table 8 Comparison of coal gasification results with literature. Work by
Type of coal
Ash content%
Air to coal ratio
Steam to coal ratio
HHV of product gas MJ/m3
Present study
Indian
36.4
Iyengar and Haque [1] Engelbrecht et al. [5]
Indian New Vaal Grootegeluk Sukunka Forestburg Roselyn Australian
40.2 40.7 31.7 13.31 8.71 5.45 14.9
Canadian Colombian Chinese (Xuzhou)
8 1.5 26.9
1.46–2 1.23 1.63–1.7 0.74–0.81 0.96–1.07 2.53 3.45 2.90 1.6 1.6 –3.2 4.46 –6.97 2.12–2.73 2.86 2.55
No steam 0.15–0.25 0.29–0.37 0.43–0.61 0.72–0.87 0.30 0.33 0.39 0.63–1.26 0.63 0.34 –1.04 0.58–0.71 0.38 0.42
2.84–3.14 3.58–4.81 4.38–4.77 5.69–6.12 4.51–5.48 3.01 3.94 3.34 2.5–2.8 2.75–1.6 1.37–2.94 2.7–3.3 3.6 4.4
Watkinson et al. [6]
Lee et al. [7] Gutierrez and Watkinson [8] Ocampo et al. [9] Xiao et al. [10]
In the final case (test run 6), three trials of experiments keeping the steam to coal ratio as 0.25 have been carried out. All other input parameters have been the same as in the test 5. The variation of temperature with time under steady operation condition at different heights for test 6 is shown in Fig. 12. Due to increased quantity of steam inflow, the average bed temperature (T1) under steady operating condition has slightly decreased to around 951 °C. It should be noted that the temperature read by T7 has increased when compared to the previous case. This is due to the slight exothermic nature of the water-gas shift reaction occurring in the freeboard due to enhanced steam supply. Fig. 13 shows the variation of synthetic gas composition with time. Significant rise in the volume fraction of hydrogen from 10.3% to 15.4% has occurred. A slight decrease in the volume fraction of CO to 19.4% is observed. Methane’s volume fraction has increase to around 1.7%. The calorific value of the synthetic gas has increased due to increase in the hydrogen content. The total carbon conversion has been around 84%. However, there is a notable increase in the cold gas efficiency from 70.6% to 77.7%. It is interesting to note that the total carbon conversion percentage of 84% is almost same as the average value reported by Engelbrecht et al. [5]. On the other hand, the cold gas efficiency of 77.7% is much higher than 65.5%, which is maximum value reported by Engelbrecht et al. [5]. The summary of the results from all the tests are presented in Table 7. Table 8 reports the comparison of gasification characteristics of Indian coal with that of other coals reported in literature. The gas heating value is found to be higher when air and steam are used as gasification agents and this is even higher than the heating values of syngas obtained during the gasification of low-ash coals [6–9]. 6. Conclusions A lab-scale bubbling fluidized bed gasification reactor has been designed, fabricated and experimented using Indian coal having high ash content. The thermal power output of the gasification reactor has been varied in the range of 80–100 kW. The equivalence ratio, which is the ratio of actual mass flow rate of air supplied to the stoichiometric air mass flow rate required for complete combustion, has been varied in the range of 0.25–0.35. The ratio of mass flow rate of steam to that of coal has been varied from 0.15 to 0.25. Effects of air and steam flow rates on the calorific value of product gas, carbon conversion and cold gas efficiency have been studied in detail. The salient conclusions are summarized as follows: (1) The calorific value of synthetic gas decreases with increase in air to coal feed rate ratio. However, it increases with an increase in the steam to coal feed ratio, keeping the equiva-
lence ratio a constant. In this study, when the steam to coal ratio is increased from 0.15 to 0.25, the calorific value of the synthetic gas increased from 3.6 MJ/m3 to 4.8 MJ/m3. This is consistent with the reports available in literature [1,5,7]. (2) As the air to coal ratio is decreased, more residence time is available for ash bound coal particle to react. Therefore, the total carbon conversion increases with a decrease in the air to coal ratio. As the air to coal ratio is decreased from 2 to around 1.5, the total carbon conversion increases by around 10%. The steam to coal ratio does not affect the total carbon conversion. As the steam to coal ratio is increased from 0.15 to 0.25, the total carbon conversion is seen to remain almost around a value of 83.4% as there is only a slight variation in the value of unburnt carbon in the fly ash and the bed ash. This is consistent with the report available in literature [7]. (3) The cold gas efficiency increases with a decrease in air to coal ratio. For 80 kWth cases, when the air to coal ratio is decreased from around 1.75 to around 1.5, the cold gas efficiency is seen to increase from around 59% to 66%. Cold gas efficiency increases with an increase in the steam to coal ratio. As the steam to coal ratio is increased from 0.15 to 0.25, the cold gas efficiency increases from around 63.6% to 77.7% due to an increase in the calorific value of synthetic gas. This study shows that there is an optimal operating regime for a given type of coal in a given reactor. Careful variation of coal, air and steam flow rates and proper choice of size of the feed coal particles contribute significantly to the gasification performance. The deteriorating effects of high-ash content in the coal can be overcome by providing sufficient bubbling (by fixing the regime properly) as well as providing enough residence time. As compared to a moving/fixed bed, the operating temperature is quite low in bubbling beds (1040 °C in slow bubbling regime, when compared to around 1350 °C in a moving bed), which is quite advantageous. Acknowledgement Authors wish to acknowledge the funding for this work from National Center for Combustion Research and Development (NCCRD) at IIT Madras and from OPTIMASH project. References [1] R.K. Iyengar, R. Haque, Gasification of high-ash Indian coals for power generation, Fuel Process. Technol. 27 (1991) 247–262. [2] Neeraj Singh, V. Raghavan, T. Sundararajan, Mathematical modeling of gasification of high-ash Indian coals in moving bed gasification system, Int. J. Energy Res. 38 (2013) 737–754.
K. Vijay Kumar et al. / Applied Thermal Engineering 116 (2017) 372–381 [3] S. Datta, P. Sarkar, P.D. Chavan, S. Saha, G. Sahu, A.K. Sinha, V.K. Saxena, Agglomeration behaviour of high ash Indian coals in fluidized bed gasification pilot plant, Appl. Therm. Eng. 86 (2015) 222–228. [4] K. Jayaraman, I. Gökalp, S. Jeyakumar, Estimation of synergetic effects of CO2 in high ash coal-char steam gasification, Appl. Therm. Eng. 110 (2017) 991– 998. [5] A.D. Engelbrecht, B.C. North, B.O. Oboirien, R.C. Everson, H.W.P.J. Neomagus, Fluidized bed gasification of high ash South African coals: an experimental and modeling study, Proc. Ind. Fluid. South Africa 145–160 (2011). [6] A.P. Watkinson, G. Cheng, C.B. Prakash, Comparison of coal gasification in fluidized and spouted beds, Can. J. Chem. Eng. 61 (1983) 468–477. [7] Woon-Jae Lee, Sang-Done Kim, Byung-Ho Song, Steam gasification of an australian bituminous coal in a fluidized bed, Korean J. Chem. Eng. 19 (2002) 1091–1096. [8] Luis A. Gutierrez, A. Paul Watkinson, Fluidized bed gasification of some Western Canadian coals, Fuel 61 (1982) 133. [9] A. Ocampo, E. Arenasb, F. Chejnea, J. Espinela, C. Londonoa, J. Aguirrea, J.D. Pereza, An experimental study on gasification of Colombian coal in a fluidized bed, Fuel 82 (2003) 161–164. [10] Rui Xiao, Laihong Shen, Mingyao Zhang, Baosheng Jin, Yuanquan Xiong, Yufeng Duan, Zhaoping Zhong, Hongcang Zhou, Xiaoping Chen, Yaji Huang, Partial
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gasification of coal in a fluidized bed reactor: comparison of a laboratory and pilot scale reactors, Korean J. Chem. Eng. 24 (2007) 175–180. Ravi Inder Singh, Anders Brink, Mikko Hupa, CFD modeling to study fluidized bed combustion and gasification, Appl. Therm. Eng. 52 (2013) 585–614. Pinto Filomena, Franco Carlos, Andre´ Rui Neto, C. Tavares, M. Dias, I. Gulyurtlu, I. Cabrita, Effect of experimental conditions on co-gasification of coal, biomass and plastics wastes with air/steam mixtures in a fluidized bed system, Fuel 82 (2003) 1967–1976. Silvera Scaccia, Stefona Stendardo, Giuseppina Vanga, Leandro Pagliari, Stefano Cassani, Mirko Nobili, Giuseppe Messina, Andrea Assettati, Giuliano Guidarelli, Salvatore Attanasi, Caterino Stringola, Andrea Grasso, Ivano Cassani, Antonio Calabr’o, Pier Ugo Foscolo, Steam-O2 coal gasification in the Italian ZECOMIX bubbling fluidized bed gasifier unit: spent bed material characterization, Open J. Chem. Eng. Sci. 1 (2014) 88–98. Hakan Karatas, Hayati Olgun, Fehmi Akgun, Coal and coal and calcined dolomite gasification experiments in a bubbling fluidized bed gasifier under air atmosphere, Fuel Process. Technol. 106 (2013) 666–672. Prabir Basu, Combustion and Gasification in Fluidized Beds, Taylor & Francis Group, 2016. D. Kunii, O. Levenspiel, Fluidization Engineering, Butterworth-Heinemann, 1991.
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Research Paper
Gasification of high-ash Indian coal in bubbling fluidized bed using air and steam – An experimental study K. Vijay Kumar a, M. Bharath a, Vasudevan Raghavan a,⇑, B.V.S.S.S. Prasad a, S.R. Chakravarthy b, T. Sundararajan a a b
Department of Mechanical Engineering, Indian Institute of Technology Madras, Chennai 600036, India Department of Aerospace Engineering, Indian Institute of Technology Madras, Chennai 600036, India
h i g h l i g h t s Bubbling bed gasification of high-ash Indian coals is systematically studied. Design, fabrication and installations details are reported. Effects of air to coal and steam to coal ratios studied in detail. Carbon conversion, synthetic gas composition and cold gas efficiency are presented for all cases.
a r t i c l e
i n f o
Article history: Received 18 November 2016 Revised 21 January 2017 Accepted 26 January 2017 Available online 30 January 2017 Keywords: Coal gasification High-ash Indian coals Bubbling fluidized bed Steam-coal ratio Carbon conversion Cold gas efficiency
a b s t r a c t In this experimental study, gasification characteristics of Indian coals having high ash content (>35%) in a bubbling fluidized bed reactor, using air and steam as the gasification agents under atmospheric pressure, are presented. The thermal power output of the gasification device has been varied in the range of 80– 100 kW. The ratio of the actual mass flow rate of air supplied to the stoichiometric air mass flow rate (required for complete combustion) has been varied in the range of 0.25–0.35. The ratio of mass flow rate of steam to that of coal has been varied from 0.15 to 0.25. Effects of air and steam flow rates on the calorific value of product gas, carbon conversion and cold gas efficiency have been studied in detail. For highash Indian coals, supply of steam produces positive effect as all the metrics such as calorific value of product gas, carbon conversion and cold gas efficiency increases with increase in the steam supply rate. Ó 2017 Elsevier Ltd. All rights reserved.
1. Introduction India has a lot of coal reserves – ranks third following China and USA. About 61% of the total energy requirement of India is met by coal-based power plants. Indian coals have good amount of fixed carbon and volatiles (around 30% each) and very little sulphur (less than 1%). However, Indian coals have large ash content, due to which, burning or gasification of Indian coals poses lots of problems in terms of ash disposal, combustion or gasification efficiency and carbon conversion. In particular, pulverized burning of Indian coals, which is carried out in several power plants, leads to significant amount of particulate emission and a costly after treatment and techniques to capture fine ash particles are required. Gasification is an important process in clean coal technology such as Inte-
⇑ Corresponding author. E-mail address: [email protected] (V. Raghavan). http://dx.doi.org/10.1016/j.applthermaleng.2017.01.102 1359-4311/Ó 2017 Elsevier Ltd. All rights reserved.
grated Gasification Combined Cycle (IGCC). In general, gasification process leads to efficient recovery of combustible gas mixtures (called synthetic gas) from coal (or any other solid fuels such as biomass), paving way to better cleaning and ash handling methods. First pioneering research work with respect to gasification of Indian coals has been reported by Iyengar and Haque [1]. They systematically investigated the gasification process of high-ash Indian coal employing lab-scale experimental setups. They used different types of gasification devices such as moving bed, bubbling fluidized bed and entrained bed. They performed the experiments using four types of Indian based coals, in which the ash content varied from around 12–31% on weight basis. They observed that the calorific value of the synthetic gas diminishes as the ash content in the coal increased. They also concluded that the adverse effect of high ash content in coal is minimally detected in moving bed gasification system followed bubbling bed and entrained flow gasification devices. Mathematical model of gasification process in
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moving bed using Indian high-ash coals was reported by Singh et al. [2]. In this study, the effect of addition of steam was brought out clearly. Datta et al. [3] studied the agglomeration characteristics of high-ash Indian coals in fluidized beds. Cause-effect relationship with respect to agglomerate formation with high ash Indian coals was reported. Recently, Jayaraman et al. [4] reported the production of char from high-ash Turkish and Indian coals using CO2 – steam agents in high heating rate Thermo Gravimetric Analyzer (TGA). Engelbrecht et al. [5] reported gasification of South African coals, which also have high ash content. They used a bubbling fluidized bed device for gasification. Watkinson et al. [6] carried out experiments with three types of Canadian coals in a fluidized and fixed bed reactors, with steam and air as gasification agents. They obtained similar gas heating values, gas yields and thermal efficiencies for both reactors. Lee et al. [7] studied the gasification on Australian coal in a fluidized bed with air and steam mixtures at atmospheric pressure. The cold gas efficiency, gas heating values and gas yields increased with an increase in the steam to coal ratio, however, the carbon conversion was almost unchanged. Gutierrez and Watkinson [8] analyzed the gasification of Canadian coal in a fluidized bed reactor with air and steam mixtures at atmospheric pressure. Ocampo et al. [9] studied the gasification characteristics of Colombian coal in a fluidized bed reactor at different air to coal and steam to coal ratios at atmospheric pressure. In these studies parameters such as cold gas efficiency, gas heating value and carbon conversion were reported. Xiao et al. [10] carried out the gasification of Chinese coal in laboratory and pilot scale reactors. In that study, the major differences in cold gas efficiency, gas heating value and carbon conversion were observed between laboratory and pilot scale reactors due to the heat loss and differences in the hydrodynamics conditions. Relatively low temperature operation is possible in a bubbling fluidized bed when compared to that of a moving bed. Based on these studies, it is clear that either a moving bed (also called fixed bed) gasification device or a slow bubbling fluidized bed gasification device can serve the purpose of attaining good gasification using Indian coals. Ash removal is little bit complicated in fixed beds, which handle large chunks of coal particles (6– 50 mm). Due to high temperature at the bottom of the reactor, ash may melt and provision for molten ash removal is required. Similarly, the fluidized bed gasification technology is also highly challenging for the Indian coals because of the issues such as improper fluidization, sudden raise in temperature due to ash agglomeration, carbon loss by bed drain, entrainment of fine carbon particles in the fly ash and treatment of cohesive coal in fly ash. These have been discussed in literature [6–14], however, not specific to high-ash coals. In summary, it is required to fix the regime of fluidization carefully in order to avoid the adverse effects of the ash content in the coal on both fluidization and gasification aspects. Therefore, in this study, a relatively slow bubbling bed gasification reactor has been designed, fabricated and experimented with Indian high-ash coals. The calorific value of product gas, carbon conversion and cold gas efficiencies have been systematically evaluated for wide range of air to coal and steam to coal ratios. The design procedure, dimensions of the reactor, experimental procedure and results have been discussed subsequently.
bigger initial particle sizes have been employed. The results from proximate and ultimate analyzes of Indian coal used in this study is reported in Table 1. The higher calorific valve of the coal is calculated by using Dulong and Petit equation, given as,
O þ 9418 S kJ=kg; HHV of coal ¼ 33823 C þ 144249 H 8 where C, H, O and S, are the mass percentages of carbon, hydrogen, oxygen and sulphur, respectively, and are obtained from ultimate analysis (Table 1). HHV is evaluated as 17.76 MJ/kg. Assuming a thermal efficiency of around 75%, the coal feed rate has been calculated for a given power rating. Air flow rate has been fixed as a percentage of stoichiometric air required for completely burning the coal. Equivalence ratio (ER), which is the ratio of the actual mass flow rate of air supplied to the stoichiometric air mass flow rate (required for complete combustion) has been varied in the range of 0.25–0.35. The ratio of mass flow rate of steam to that of coal has been varied from 0.15 to 0.25. In the cases where steam has been supplied the equivalence ratio has been kept constant as 0.25. Table 2 reports the parameters varied in the present experimental study. For all the cases, a regime corresponding to that of bubbling fluidized bed has been considered. The hydrodynamic calculations have been carried out to establish the dimensions, particle sizes and operating conditions of the gasification device [15,16]. The minimum fluidization velocity (Umf), superficial velocity (Uo), transport disengaging height (TDH) and total height have been calculated for different cases for a range of average particle sizes. The minimum fluidization velocity (Umf) is calculated using the procedure given in Kunii and Levenspiel [16] and transport disengaging height (TDH) is calculated using the procedure given in Basu [15]. Superficial gas velocities are calculated theoretically at the experimental operating conditions (1 atm, 900 °C). Table 3 reports these quantities for different cases. The diameter of the gasification device has been fixed as 0.15 m (150 mm). Based on the hydrodynamic calculations, for bubbling fluidization regime and for the range of power rating, the height, particle sizes and superficial velocities have been fixed. Superficial velocity (Uo) is 1–1.89 times larger than minimum fluidization velocity (Umf). For example, it has been observed that for 100 kWth case and for ER of 0.35, the value of TDH is around 2.07 m, considering an average particle size of 2.5 mm, which is the maximum among all the cases. Based on this, the total height of the reactor has been fixed as 2.5 m. The bed height is varied based on the air and steam flow rates.
3. Experimental setup The reactor has been fabricated in three portions, namely, the top portion having a height of 0.5 m, the middle portion with a height of 1 m and the bottom portion of 1 m height. The exit of top portion is welded with a flange, which facilitates the connection of pressure regulator, pipes for gas scrubber and synthetic gas burner. The gasification unit has been fabricated using two cylindrical refractory linings each having a thickness of 100 mm.
Table 1 Analysis of Indian coal (air dried).
2. Design procedure The thermal power rating of bubbling fluidized bed gasification reactor has been fixed initially at 100 kW based on the choice of regime of fluidization and particle size. However, after the first test, noting a low carbon conversion due to carbon loss from flue gases, the thermal power rating has been reduced to 80 kW and
Proximate analysis
wt (%)
Ultimate analysis
wt (%)
Ash content Moisture Volatiles Fixed carbon
36.4 5.4 28.7 29.5
Carbon Hydrogen Nitrogen Oxygen Sulphur
43.7 3.8 0.9 14.5 0.7
Calorific value (MJ/kg)
17.76
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Table 2 Experimental tests. Test No.
Power output (kWth)
Coal feed rate (kg/h)
Air flow rate (kg/h)
ER
Steam flow rate (kg/h)
Steam/coal ratio (S/C)
1 2 3 4 5 6
100 80 80 80 80 80
20.2 16.2 16.2 16.2 16.2 16.2
40.6 32.5 27.8 23.2 23.2 23.2
0.35 0.35 0.30 0.25 0.25 0.25
– – – 2.4 3.2 4.0
– – – 0.15 0.20 0.25
Table 3 Hydrodynamic calculations. Umf (m/s)
d*p
U*
TDH (m)
U0 (m/s)
Dense bed (m)
Total height (m)
Test run 1: 100 kWth, ER = 0.35 3.0 28 2.5 18.3
1.42 1.12
38.35 31.96
1.09 1.09
1.59 2.07
2.12 2.12
0.3 0.3
1.89 2.37
Test run 2: 80 kWth, ER = 0.35 3.0 28 2.5 18.3
1.42 1.12
38.35 31.96
0.86 0.86
0.96 1.40
1.70 1.70
0.3 0.3
1.26 1.70
Test run 3: 80 kWth, ER = 0.3 3.0 28 2.5 18.3
1.42 1.12
38.35 31.96
0.74 0.74
0.53 1.04
1.45 1.45
0.3 0.3
0.83 1.34
Test run 4: 80 kWth, ER = 0.25, S/C = 0.15 3.0 28 1.42 2.5 18.3 1.12
38.35 31.96
0.72 0.72
0.33 0.99
1.41 1.41
0.3 0.3
0.63 1.29
Test run 5: 80 kWth, ER = 0.25, S/C = 0.20 3.0 28 1.42 2.5 18.3 1.12
38.35 31.96
0.76 0.76
0.61 1.08
1.48 1.48
0.3 0.3
0.91 1.38
Test run 6: 80 kWth, ER = 0.25, S/C = 0.25 3.0 28 1.42 2.5 18.3 1.12
38.35 31.96
0.79 0.79
0.75 1.19
1.55 1.55
0.3 0.3
1.05 1.49
dp (mm)
Re
The refractory linings have different thermal conductivities as listed in Table 4. A mild steel shielding of 10 mm thickness has been provided on the outer surface of the second refractory to provide strength to the column. The entire experimental setup is shown in Fig. 1. Distributor plate has a conical section with an angle of 45°. Top internal diameter of a cone is 0.15 m and the bottom internal diameter is 0.04 m. It is made of stainless steel 310 grade and has a thickness of 30 mm. It has several perpendicular holes of 2 mm diameter, separated by a pitch distance of 4 mm. Distributor plate is placed in between plenum (wind box) and bottom portion of the gasifier. Coal particles from a hopper is fed to the reactor at a height of 0.5 m above the distributor plate. A screw conveyor is used to feed the coal. Pitch size, diameter and length of screw conveyor to feed coal are 35 mm, 50 mm and 650 mm, respectively. The feed rate is controlled by a variable speed DC motor. A water jacket is provided outside the screw conveyor in order to cool it sufficiently. Seven thermocouples have been provided at different heights (Table 5) in order to measure the axial temperature distribution inside the reactor.
Based on the test case, either pre-heated air alone or pre-heated air and steam are injected into the plenum (wind box). The gasification agent enters the reactor through the holes in the conical distributor plate and bubble through the bed material and coal particles. Initially the bed material (sand particles) is filled up in the pipe below the distributor plate, leading to the ash collector. After providing sufficient time for conversion of the coal particles, bed ash is discharged into the ash collector by opening the valve connected to the pipe below the plenum. Fine fly-ash particles in the flue gas are removed in the filter unit as well as in the water scrubber. Filter box of size 350 mm inner diameter and 1 m long is made up of stainless steel. Three layers of meshes, with holes of sizes 2 mm, 1 mm and 0.5 mm, respectively, are placed inside the filter unit at different heights. The synthetic gas passes through these meshes and the particulate material is collected at the bottom of the filter unit. The synthetic gas is then allowed to pass through a scrubber, which is filled with water to cool the flue gas as well as to separate the left-over fine fly ash contained in the gas. At regular time intervals, small amount of gas from the exit pipe attached to the water scrubber is taken into a gas analyzer and its
Table 4 Properties of refractory linings. Properties of refractory
Material composition
%
Thermal conductivity (W/m K)
Maximum service temperature (°C)
Hot face refractory (self flow casting having low iron)
Al2O3 CaO Fe2O3
60.8 3.45 0.77
1.32
1500
Cold face refractory (high purity insulating castable)
Al2O3 SiO2 CaO Fe2O3
46.1 22.6 7.72 1.33
0.31
1300
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Fig. 1. Schematic of assembled experimental set up along with instrumentation.
Air pre-heater, shown in Fig. 2(a), is a specially designed heat exchanger that utilizes a premixed LPG-air burner. It is made up
Table 5 Location of the thermocouples. S. No.
Location of thermocouples measured from distributor plate (m)
1 2 3 4 5 6 7
0.2 0.5 0.8 1.2 1.5 1.8 2.4
composition is determined. The synthetic gas is burned in a syngas burner, and the flue gases are left to the atmosphere. The volumetric flow rate of the flue gas is measured by passing it through a precalibrated rotameter. Gas analyzer measures the volumetric percentages of components present in the synthetic gas such as CO, CO2, O2, CH4 and H2. The gas analyzer has been specially calibrated to measure the product of gasification, where high amounts of CO and H2 are present. The volume fractions can be measured with an uncertainty of 0.1%. The range of volumetric percentage of each component of synthetic gas and the type of sensors used in the gas analyzer are listed in Table 6.
(a) Air pre-heater
Table 6 Range of composition and sensor type in gas analyzer. S. No.
Gas
Range (% V/V)
Sensor type
1 2 3 4 5
CO CO2 O2 CH4 H2
0–50 0–50 0–30 0–20 0–40
Non-dispersion infrared Non-dispersion infrared Electrochemical Non-dispersion infrared Thermal conductivity
(b) Steam generator Fig. 2. Schematics of (a) air pre-heater and (b) steam generator.
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of stainless steel. Air flows through a long stainless steel tube arranged in a spiral form as shown in Fig. 2(a). The inner diameter of the air tube is 19 mm and its length is 11 m. This arrangement is kept in a shell, on top of which, the Liquefied Petroleum Gas (LPG) burner is placed. The products of combustion of LPG-air flows from the top towards the bottom, passing through the coiled air tube. The capacity of the burner is 50 kW. The LPG is supplied continuously to LPG-air burner at a rate of 0.068 kg/h. A refractory block is placed below the burner to guide the hot gases in proper direction. Air heater is designed to heat the required flow rate of air up to 500 °C. Steam generator, as shown in Fig. 2(b), consists of a boiler unit at the bottom of a shell and a super-heater unit having bank of tubes placed above the shell. An inverted LPG-air burner, similar to that used in the air pre-heater, is placed in between the boiler and super heater units. Stainless steel 304 is used for fabrication. Super heater unit consists of 40 tubes, each of length 200 mm and inner diameter 19 mm. Water enters into the boiler around the walls in the bottom portion and evaporates due to the heat transfer from the flame impinging surfaces in the bottom portion. The saturated steam passes through the super heater unit for further heating by the raising hot flue gases. The maximum flow rate in the steam generator is 10 kg/h with a temperature up to 400 °C at atmospheric pressure.
4. Experimental procedure Coal is pulverized and the particles are sieved using standard sieves. The maximum and minimum sizes of that sieve are reported as particle size range. In order to achieve ignition and to preheat the bed, the gasification unit has been initially filled with 5 kg of bed material (sand particles of size 1–1.5 mm) and 3 kg of fine coal particles (1–2 mm). Air at a rate of 20.8 kg/h has been supplied through the plenum to achieve proper fluidization of solid particles. LPG has been injected into the gasifier at a height of around 0.2 m above the distributor plate. The mixture of LPG and air has been ignited using a pilot flame. LPG flame rapidly heats up the refractory and the bed materials. The coal particles have been subsequently ignited. When the bed temperature, as measured by the thermocouple T1 reaches around 600 °C, LPG injection has been stopped. Combustion of coal continues the raise the bed temperature. When the temperature reaches around 1050 °C, coal particles of size 2–3 mm have been fed using the screw conveyor. The bed temperature momentarily decreases during the initial addition of coal particles and subsequently increases to reach a steady state value. The rate of coal feeding has been slowly and gradually increased to the required value. Steady operating regime has been established within around 2 h. At regular time intervals, ash and bed material particles have been removed. The mass of the removed material has been measured using a load cell. It should be noted that the feeding of coal and gasification agent have been continuous processes, however, the ash removal has been a manual batch process. Care has been taken to remove the bed material and ash without quenching the bed as a result of over-extraction. In the test cases where steam has been also used along with air as a gasification agent, once stable conditions have been achieved, steam has been introduced into the reactor. The product gas has been sampled into the gas analyzer after the steady operation has reached and composition of synthetic gases have been recorded at regular intervals. The reactor has been operated for around 5 h such that steady operation has been achieved for at least 2.5–3 h in each test case. All the test cases have been repeated at least three times to ensure similar operating conditions and output. It has been observed that
the data has been repeatable within 1.5% around the average values. The bed ash and fly ash have been kept in furnace at around 950 °C for around 5 h and the mass loss has been recorded to estimate the amount of unburnt carbon present in these samples. Total and fixed carbon conversions have been calculated using the procedure reported by Engelbrecht et al. [5]. The formulae for calculating the fixed carbon conversion and total carbon conversion are given by
Fixed carbon conversion ¼
fða bÞ ½ðd eÞ þ ðf gÞg ; ða bÞ
Total carbon conversion ¼
fða cÞ ½ðd eÞ þ ðf gÞg ; ða cÞ
where a is coal feed rate in kg/h, b is the percentage of fixed carbon content in coal, c is the percentage of carbon content in coal, d is the flow rate of bed ash in kg/h, e is the percentage of carbon content in bed ash, f is the flow rate of filter ash in kg/h and g is the percentage of carbon content in filter ash. Cold gas efficiency is calculated using the formulae [15] given as 3
Cold gas efficiency ¼
Syn gas flow rate in m =s Gas HHV in MJ=m3 Coal feed rate in kg=s Coal HHV in MJ=kg 100;
Gas higher heating value (HHV) depends on product volume percentages (Vi) of CO, H2 and CH4 and their calorific values (CVi). It is calculated as
Gas HHV ¼
VCO CV CO þ VH2 CV H2 þ VCH4 CV CH4 100
5. Results and discussion The reaction temperature is one of the significant parameters that affects the performance of a gasification unit. The processes in gasification should be coupled in such a way that the exothermic reactions, which utilize oxygen and release heat, should transfer the heat to the endothermic reactions, which predominantly utilize gases such as CO2 and H2O. As mentioned earlier, the temperature in a fluidized bed reactor is almost uniform and is relatively lesser than that of a fixed or moving bed reactor. The maximum temperature of the gasification unit is usually maintained below the ash-melting temperature for easier operation and this become significant when the ash content in coal is high. The main chemical reactions in the coal gasification are summarized below [15,16]. 1. Oxidation reactions: C + 0.5O2 M CO2, C + O2 M CO2 and H2 + 0.5O2 M H2O (exothermic) 2. Boudouard reaction: C + CO2 M 2CO (endothermic) 3. Water gas reaction: C + H2O M CO + H2 (endothermic) 4. Methane formation reaction: C + 2H2 M CH4 (exothermic) 5. Water gas shift reaction: CO + H2O M CO2 + H2 (exothermic) 6. Steam methane reforming reaction: CH4 + H2O M CO + 3H2 (endothermic) 5.1. Test run-1–100 kWth In test run-1, coal feed rate has been calculated for a power output of 100 kWth and the air flow rate has been calculated considering an equivalence ratio of 0.35. No steam has been added in this case. The variation of temperature with time as measured by thermocouples located at different heights of the reactor is shown in Fig. 3 for the transient operation period and the same for steady operation period is shown in Fig. 4. Time = 0 s in Fig. 3 indicates
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Fig. 3. Transient temperature variation at different heights for test run 1.
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Fig. 5. Variation of synthetic gas composition with time for test run 1.
78 and 89% and for the coal with 31.7% ash, the total carbon conversion varied from 60 to 87%, depending upon the air and steam flow rates. The cold gas efficiency reported by Engelbrecht et al. [5] has been around 65.5% (maximum). 5.2. Test runs at 80 kWth power with air alone as gasification agent
Fig. 4. Variation of temperatures with time at different heights for test run 1 under steady operation regime.
the starting of the ignition process. Ignition and steady operating conditions have been established as discussed in the previous section. The average bed temperature, as measured by T1, under steady operating conditions (2.5–5 h) is around 922 °C. This value is comparable to that reported by Engelbrecht et al. [5], where two types of South African coals with ash content 40.7% and 31.7% have been utilized. The temperatures measured by T1, T2, T3 and T4 are close to each other with only around 70 °C drop until a height of 1.2 m. The uniform temperature distribution in this height range ensures that the operation is in bubbling bed regime and also indicates a good solid–gas mixing. However, above around 1.5 m height, the temperature decreases significantly to around 450 °C, indicating heat losses in freeboard. For the smaller scale reactors such as this, the heat losses are usually higher than the actual scaled up reactors. The amounts of unburned char present in fly ash and ash removed from the bed have determined by placing the collected samples in a furnace at 950 °C for 5 h. It is observed that the unburned char present in the fly ash is around 25.6% and that in the ash extracted from the bed is around 2%. The deteriorating effect of high-ash content (36.4%) is clearly observed for this case, as indicated by trapped char in fine ash particles (100–500 lm). Synthetic gas composition has been measured under steady operating conditions as shown in Fig. 5, where the average values obtained from three trial have been plotted. On dry basis, the volumetric percentages of CO, CO2, H2 and CH4, are 15.8, 14.5, 7.5 and 0.1 respectively. Total carbon conversion is around 74.3% and the cold gas efficiency is around 61.2%. Engelbrecht et al. [5] used steam in addition to air. Their gas composition showed higher percentages of hydrogen. They have reported that for coal with around 40.7% ash, the total carbon conversion has been between around
Based on the results obtained from 100 kWth run, in order to minimize the unburned char carried away by the fly ash, the feed size of the coal particles has been increased to 2.5–3 mm and the superficial velocity is reduced from 2.12 m/s to 1.7 m/s. For these conditions, the coal feed rate has also been decreased 16.2 kg/h, which corresponds to a power rating of 80 kWth, corresponding to an equivalence ratio of 0.35 (test run 2). This case has moved to a slower bubbling regime when compared to the previous case. The temporal variations of temperatures measured along the height of the gasification reactor for this case have been similar to that of the previous case. However, the average bed temperature, as measured by T1, under steady operating conditions (2.5–5 h) has been around 999 °C, which is higher than the previous case. The temperature in the bed has increased because of an increase in the average size of coal particles as well as due to the decrease in the velocity of air. There has been a phenomenal decrease in the amount of unburned char present in the fly ash (from 25.6 to 18.4%) and that in the ash removed from the bed (from 2 to 1%). For this case, the average volumetric percentages of CO2, CO and H2, have been around 18.6%, 15.3% and 6.7%, respectively. The total carbon conversion has increased from 74.3% to 82.4% and the cold gas efficiency has remained almost the same (around 59.4%, when compared to 61% in test run 1). In order to study the effect of ER, a test with a lower ER value has been conducted. The air flow rate is reduced such that ER is equal to 0.3 (Table 1 – test run 3). Coal feed rate in maintained to provide a power output of 80 kWth and the coal particles are fed with the size range of 2.5–3 mm as in the previous case. The superficial velocity for this case is around 1.45 m/s. Fig. 6 shows the temporal variations of temperatures at different heights under steady operation regime. It is observed that in the slow bubbling regime, the temperature gradient along the reactor height has increase when compared to the earlier cases. However, as a result of slow bubbling process, the average bed temperature, as measured by T1, under steady operating condition (2.5–5 h) has increased to around 1040 °C. In this slow bubbling regime, the unburned char in the fly-ash has further reduced from 18.4% to 16.8%. The average composition of synthetic gas under steady operating condition is shown in Fig. 7. It is observed that the volumetric percentages of CO2, CO and H2 are 17, 19.1 and 6.9, respectively. The concentration of CH4 is almost zero. The total carbon
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0.15–0.25 respectively. The regime for the test cases 4, 5 and 6 are very much similar to that of test cases 2 and 3. As mentioned earlier, the reactor is started only using air and after establishing steady operating condition, steam injection is started gradually. Fig. 8 presents the temporal variation of temperature at several locations along the reactor height under steady operating conditions for the case of steam to coal ratio of 0.15. It can be noted that the maximum temperature has slightly decreased (from 1040 °C to around 968 °C) as a result of steam addition. In this slow bubbling regime, there is a notable temperature gradient along the height of the gasifier. The unburned char present in the fly ash is around 18.3%, almost same as in the previous case. Fig. 9 presents the variation Fig. 6. Variation of temperatures at different heights with time for test run 3.
Fig. 8. Variation of temperatures at different heights with time for the test run 4. Fig. 7. Variation of average composition of synthetic gas with time for test run 3.
conversion has increased further to 84.1% for this test and the cold gas efficiency has increased to around 65.7%, comparable to that reported in Engelbrecht et al. [5]. The test run 3 clearly shows that there is an optimal operating regime for a given type of coal in a given reactor. Careful variation of coal and air flow rates and proper choice of size of the feed coal particles contribute significantly to the gasification performance. The deteriorating effects of high-ash content in the coal can be overcome by providing sufficient bubbling (fixing the regime properly) as well as providing enough residence time. As compared to a moving/fixed bed, the operating temperature is quite low in bubbling beds (1040 °C in slow bubbling regime, when compared to around 1350 °C in a moving bed), which is quite advantageous [1]. Fig. 9. Variation of the composition of the synthetic gas with time for test run 4.
5.3. Test runs at 80 kWth power with air and steam as gasification agents Addition of steam results in water-gas reactions. The gas-phase water-gas shift reaction is slightly exothermic and the solid phase water-gas reaction is endothermic. Water-gas shift reaction provides increased amounts of hydrogen in the synthetic gas. Therefore, like coal and air feed rates, the steam feed rate also affects the gasification process significantly. Keeping this in mind, three trials of tests 4, 5 and 6 have been conducted, where the steam flow rate has been increased gradually. Since steam is added, the ER value is decreased to 0.25 for these cases. The power rating and thus, the coal feed rate, has been maintained the same as in test runs 2 and 3. Also, the feed coal particles size has been in the range of 2.5–3 mm, as in the previous case. Steam flow rate is varied from 2.4 kg/h to 4 kg/h, for a coal flow rate of 16.2 kg/h. This corresponds to steam to coal ratio of
Fig. 10. Variation of temperatures at different heights with time for the test run 5.
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Fig. 11. Variation of the composition of the synthetic gas with time for test run 5.
Fig. 13. Variation of the composition of the synthetic gas with time for test run 6.
gas mixture. The total carbon conversion and cold gas efficiency for this test are 83.3% and 63.6%, respectively. In test run 5, keeping the same coal and air feed rates, and the feed coal particle size, as in the previous case, the steam flow rate has been increased to attain a steam to coal feed ratio of 0.2. Fig. 10 presents the temporal variation of temperature at several locations along the reactor height under steady operating conditions. A slight reduction in the bed temperature indicated by T1 is observed (958 °C as compared to the previous case of 968 °C). While the dense portion of the reactor show an almost steady temperature distribution, the freeboard portion shows slight transients as a result of changes undergone by the gas-phase reactions. The fly ash carbon content is almost the same at 18.1%, as in the previous case. Fig. 11 presents the volumetric fraction of CO, CO2 and H2 in the synthetic gas under steady operating conditions. There is a further increase in the volumetric fraction of hydrogen from 8.4% in the previous case to around 10.3% in the present case. Similarly, the volume fraction of CO has also increased from 18.5% to 20.6%. The volumetric fraction of methane has also increased slightly to 1%. With these, the calorific value of synthetic gas has further increased. While the total carbon conversion (83.4%) is almost the same as in the previous case, the cold gas efficiency has increased significantly from 63.6% to 70.6%. This case gives a scope of further increasing the steam flow rate as the maximum bed temperature is still higher than 950 °C.
Fig. 12. Variation of temperatures at different heights with time for the test run 6.
of the synthetic gas composition with time. The volumetric percentages of CO2, CO and H2, are 15.9, 18.5 and 8.4, respectively. An apparent increase in the volume fraction of hydrogen from 6.9% to 8.4% due to steam addition is observed. Also, methane formation reaction rate has increased due to increased availability of hydrogen. This is apparent from the volumetric percentage of CH4, which has increased to around 0.92%. The gas calorific value has increased due to the increase in the content of hydrogen in the
Table 7 Summary of fluidized bed gasification tests using high-ash Indian coal. Test number
1
2
3
4
5
6
Capacity (kWth) Coal feed rate (kg/h) Air flow rate (kg/h) Steam flow rate (kg/h) ER Air to coal ratio Steam to coal ratio Average temperature of air/steam at inlet (°C) Average lower bed temperature (°C) Coal particle size (mm)
100 20.2 40.6 – 0.35 2 – 300 922 2–3
80 16.2 32.5 – 0.3 1.73 – 340 999 2.5–3
80 16.2 27.8 – 0.3 1.48 – 340 1040 2.5–3
80 16.2 23.2 2.4 0.25 1.23 0.15 350 968 2.5–3
80 16.2 23.2 3.2 0.25 1.23 0.2 365 958 2.5–3
80 16.2 23.2 4.0 0.25 1.23 0.25 360 951 2.5–3
Dry gas composition (Volume %) CO H2 CH4 CO2 N2 + Other Gas calorific value (MJ/Nm3) Unburned carbon in fly ash (%) Fixed carbon conversion (%) Total carbon conversion (%) Cold gas efficiency (%)
15.8 7.5 0.1 14.5 62.13 2.84 25.6 61.95 74.26 61.2
15.3 6.7 0 18.6 59.4 2.99 18.4 73.98 82.39 59.4
19.2 6.9 0 17.0 56.9 3.14 16.8 76.51 84.1 65.7
18.5 8.4 0.9 15.9 56.3 3.58 18.3 75.21 83.23 63.6
20.6 10.3 1.0 14.2 53.9 4.1 18.1 75.52 83.43 70.6
19.4 15.4 1.6 12.4 51.2 4.81 17.5 76.43 84.05 77.7
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Table 8 Comparison of coal gasification results with literature. Work by
Type of coal
Ash content%
Air to coal ratio
Steam to coal ratio
HHV of product gas MJ/m3
Present study
Indian
36.4
Iyengar and Haque [1] Engelbrecht et al. [5]
Indian New Vaal Grootegeluk Sukunka Forestburg Roselyn Australian
40.2 40.7 31.7 13.31 8.71 5.45 14.9
Canadian Colombian Chinese (Xuzhou)
8 1.5 26.9
1.46–2 1.23 1.63–1.7 0.74–0.81 0.96–1.07 2.53 3.45 2.90 1.6 1.6 –3.2 4.46 –6.97 2.12–2.73 2.86 2.55
No steam 0.15–0.25 0.29–0.37 0.43–0.61 0.72–0.87 0.30 0.33 0.39 0.63–1.26 0.63 0.34 –1.04 0.58–0.71 0.38 0.42
2.84–3.14 3.58–4.81 4.38–4.77 5.69–6.12 4.51–5.48 3.01 3.94 3.34 2.5–2.8 2.75–1.6 1.37–2.94 2.7–3.3 3.6 4.4
Watkinson et al. [6]
Lee et al. [7] Gutierrez and Watkinson [8] Ocampo et al. [9] Xiao et al. [10]
In the final case (test run 6), three trials of experiments keeping the steam to coal ratio as 0.25 have been carried out. All other input parameters have been the same as in the test 5. The variation of temperature with time under steady operation condition at different heights for test 6 is shown in Fig. 12. Due to increased quantity of steam inflow, the average bed temperature (T1) under steady operating condition has slightly decreased to around 951 °C. It should be noted that the temperature read by T7 has increased when compared to the previous case. This is due to the slight exothermic nature of the water-gas shift reaction occurring in the freeboard due to enhanced steam supply. Fig. 13 shows the variation of synthetic gas composition with time. Significant rise in the volume fraction of hydrogen from 10.3% to 15.4% has occurred. A slight decrease in the volume fraction of CO to 19.4% is observed. Methane’s volume fraction has increase to around 1.7%. The calorific value of the synthetic gas has increased due to increase in the hydrogen content. The total carbon conversion has been around 84%. However, there is a notable increase in the cold gas efficiency from 70.6% to 77.7%. It is interesting to note that the total carbon conversion percentage of 84% is almost same as the average value reported by Engelbrecht et al. [5]. On the other hand, the cold gas efficiency of 77.7% is much higher than 65.5%, which is maximum value reported by Engelbrecht et al. [5]. The summary of the results from all the tests are presented in Table 7. Table 8 reports the comparison of gasification characteristics of Indian coal with that of other coals reported in literature. The gas heating value is found to be higher when air and steam are used as gasification agents and this is even higher than the heating values of syngas obtained during the gasification of low-ash coals [6–9]. 6. Conclusions A lab-scale bubbling fluidized bed gasification reactor has been designed, fabricated and experimented using Indian coal having high ash content. The thermal power output of the gasification reactor has been varied in the range of 80–100 kW. The equivalence ratio, which is the ratio of actual mass flow rate of air supplied to the stoichiometric air mass flow rate required for complete combustion, has been varied in the range of 0.25–0.35. The ratio of mass flow rate of steam to that of coal has been varied from 0.15 to 0.25. Effects of air and steam flow rates on the calorific value of product gas, carbon conversion and cold gas efficiency have been studied in detail. The salient conclusions are summarized as follows: (1) The calorific value of synthetic gas decreases with increase in air to coal feed rate ratio. However, it increases with an increase in the steam to coal feed ratio, keeping the equiva-
lence ratio a constant. In this study, when the steam to coal ratio is increased from 0.15 to 0.25, the calorific value of the synthetic gas increased from 3.6 MJ/m3 to 4.8 MJ/m3. This is consistent with the reports available in literature [1,5,7]. (2) As the air to coal ratio is decreased, more residence time is available for ash bound coal particle to react. Therefore, the total carbon conversion increases with a decrease in the air to coal ratio. As the air to coal ratio is decreased from 2 to around 1.5, the total carbon conversion increases by around 10%. The steam to coal ratio does not affect the total carbon conversion. As the steam to coal ratio is increased from 0.15 to 0.25, the total carbon conversion is seen to remain almost around a value of 83.4% as there is only a slight variation in the value of unburnt carbon in the fly ash and the bed ash. This is consistent with the report available in literature [7]. (3) The cold gas efficiency increases with a decrease in air to coal ratio. For 80 kWth cases, when the air to coal ratio is decreased from around 1.75 to around 1.5, the cold gas efficiency is seen to increase from around 59% to 66%. Cold gas efficiency increases with an increase in the steam to coal ratio. As the steam to coal ratio is increased from 0.15 to 0.25, the cold gas efficiency increases from around 63.6% to 77.7% due to an increase in the calorific value of synthetic gas. This study shows that there is an optimal operating regime for a given type of coal in a given reactor. Careful variation of coal, air and steam flow rates and proper choice of size of the feed coal particles contribute significantly to the gasification performance. The deteriorating effects of high-ash content in the coal can be overcome by providing sufficient bubbling (by fixing the regime properly) as well as providing enough residence time. As compared to a moving/fixed bed, the operating temperature is quite low in bubbling beds (1040 °C in slow bubbling regime, when compared to around 1350 °C in a moving bed), which is quite advantageous. Acknowledgement Authors wish to acknowledge the funding for this work from National Center for Combustion Research and Development (NCCRD) at IIT Madras and from OPTIMASH project. References [1] R.K. Iyengar, R. Haque, Gasification of high-ash Indian coals for power generation, Fuel Process. Technol. 27 (1991) 247–262. [2] Neeraj Singh, V. Raghavan, T. Sundararajan, Mathematical modeling of gasification of high-ash Indian coals in moving bed gasification system, Int. J. Energy Res. 38 (2013) 737–754.
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