GT

Article pubs.acs.org/EF

New Strategy of Fuel-Slurry Integrated Gasifier/Gas Turbine (FSIG/ GT) Alternative for Power Generation Applied to Biomass Marcio L. de Souza-Santos* and Wilson de A. Beninca Faculty of Mechanical Engineering Department of Energy, University of Campinas, São Paulo CEP 13081-970, Brazil ABSTRACT: The new strategy of the fuel-slurry integrated gasifier/gas turbine (FSIG/GT) concept for thermoelectric power generation is applied to the case of sugar cane bagasse (SCB). The FSIG/GT process allows fuel feeding to a power unit based on gasification using commercially available slurry pumps, thus avoiding the usual sequential lock-hoppers, also known as cascade feeding systems. It also dispenses with the need of pure oxygensometimes combined with hydrocarbonsto promote ignition of particles in the injected slurry. The fuel slurry is prepared to high dry-solid content and pumped into a dryer, from which the solid particles are fed into the gasifier. Because both units operate under similar pressures, simple rotary valves and Archimedes’ screws might carry the secondary feeding. The gas is cleaned to bring the particle content and size as well alkaline concentration within the acceptable limits for injections into standard gas turbines. The work shows that no steam is needed as gasification agent, thus simplifying the process and decreasing the capital costs. Moreover, the strategy leads to higher efficiency when compared with power generation based on high-pressure steam turbines, BIG/GT process consuming SCB, and combined cycles using pressurized-chamber boilers.

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INTRODUCTION The use of biomass as a renewable source of power generation with near zero overall greenhouse gases emissions1 has increased, particularly in the case of sugar cane bagasse (SCB).2−4 Usually, traditional Rankine cycles are applied, but more efficient biomass integrated gasification/gas turbines have been used.5−13 On the other hand, this last process introduced technical hurdles; among them, the gas cleaning to remove particulates and alkaline species to meet acceptable levels for injection into turbines was one the major hurdles.14−17 That obstacle has been overcome.18,19 Another important technical barrier remains, which is the difficulty of feeding solid particulate fuels into pressurized vessels. It is well known that it is not possible to feed particulate solid fuels from atmospheric conditions into pressurized vessels in a single stage without meeting unsurmountable problems. For instance, attempts to do so using feeding screws would compact the particles into high-density blocks that would not disintegrate into smaller particles again inside the reactor. The process can even surpass the maximum torque of the feeding screws leading to mechanical failures. Cascade or sequential feeding systems, composed of two or more levels of pressurized lock hoppers are the most common alternative to avoid such problems.20 The particulate fuel is fed at the top hopper, from which it is conducted to a second one below through a rotary valve. The pressure in the second hopper is higher than the first above; however, the difference of pressures between the two hoppers is within the capacity of the rotary valve to keep the variance between the two environments. Furthermore, partial devolatilization of the fuel may start due to temperature increases when the pressure is raised in the hopper. If so, tar would be released, causing the particles to stick together, thus preventing them to proceed or drop into the rotary valve. Usually, an inert gas, such as nitrogen, is employed to keep the pressurized atmosphere inside the hoppers. This prevents the onset of pyrolysis and even © 2014 American Chemical Society

combustion of the particulate fuel in the hoppers. Then, the solid fuel goes through another rotary valve to a third hopper below, which is kept to an even higher pressure than the one above it. The pressure at the final destination, the capacity of rotary valves to keep environments under different pressures without leaks, and the maximum gradual compression applied to inert gas injected into each hopper without provoking fuel devolatilization determine the number of stages or hoppers. The whole procedure consumes expensive inert gases, thus introducing costs to the power generation unit, not to mention losses on the overall efficiency of the unit due to power diverted to inert gas compressions and cooling. Additionally, it relies on complex sequential operations, which are prone to failures mainly due to interruptions of continuous flow of fuel downward to the next hopper and respective rotary feeding valve. Static electricity build up among particles and entanglements of neighboring particle extremities might cause such problems, particularly in the case of fibrous materials such as sugar cane bagasse (SCB). Pumping fuel slurries into pressurized vessels has been applied for a long time.21 It greatly simplifies the feeding process and, very likely, decreases the capital, operational, and maintenance costs when compared with methods based on a cascade systems of hoppers. On the other hand, the application of slurry feeding to power generation has been confined to the use of boilers,21−25 because the vaporization of the fuel original moisture, added to the water to prepare the slurry, demands burning a substantial fraction of the fuel. That would render very low efficiencies to the gasification and overall power generation process. Other processes need to apply pure oxygensometimes combined Received: February 4, 2014 Revised: March 10, 2014 Published: March 12, 2014 2697

dx.doi.org/10.1021/ef500317a | Energy Fuels 2014, 28, 2697−2707

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with hydrocarbonsto promote ignition of particles in the injected slurry.26 The previous study of FSIG/GT process27 considered the following two alternatives: (a) configuration A, shown in Figure1, where part of the exiting stream from the drying unit (stream 20) is used to provide steam to enhance the gasification. (b) configuration B, shown in Figure 2, where an intermediary extraction (stream 12) from the main steam turbine cycle provides steam for gasification.

Figure 3. Configuration C of the proposed FSIG/GT process. C = compressor, CB = combustor, CD = condenser, CL = cleaning system, CY = cyclone, D = dryer, DF = dried fuel, FE = screw feeding, FS = fuel-slurry pumping, G = gasifier, GT = gas turbine, SG = steam generator, ST = steam turbine, P = water pump, V= valve or splitter.

content slurry (stream 26), which is pumped by equipment 17 into the bubbling fluidized bed dryer (D) that operates at around 2 MPa. Gas stream 28 is used for the drying process. That stream is part of the gas turbine (equipment 3) exhaust after driving the Rankine heat-recovering cycle (equipment 4 to 8). Stream 15 is compressed, leaving stream 28 at suitable temperatures for the slurry drying process. Because the dryer (D) and gasifier (G) operate at similar pressures, the dried fuel can be fed into the gasifier using simple rotary valves combined with Archimedes’ screws. Cyclones and dust collectors drop the particle maximum particle diameter content in stream 16 to values acceptable for injections into gas turbines. Then, stream 16 exchanges heat at equipment 11 in order to reach temperatures below the dew points of alkaline species; such a procedure decreases their concentrations to values acceptable for injection of stream 4 into the gas turbine (equipment 3). The energy recovered from that gas cooling drives another Rankine Cycle composed by equipment 11 to 15. Configurations A and B just add the injections of steam into the gasifier.

Figure 1. Configuration A of the proposed FSIG/GT process. C = compressor, CB = combustor, CD = condenser, CL = cleaning system, CY = cyclone, D = dryer, DF = dried fuel, FE = screw feeding, FS = fuel-slurry pumping, G = gasifier, GT = gas turbine, SG = steam generator, ST = steam turbine, P = water pump, V= valve or splitter.

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METHODOLOGY The conceptual development of the present process demanded extensive simulations and optimizations of the gasification and drying processes as well of the power generation architecture. Because all involved processes are coupled, many simulation trials were required, including revisiting one stage after improvements on others. The Comprehensive Simulator of Fluidized and Moving Bed equipment [CeSFaMB is also known as CSFMB©] (http:// www.csfmb.com) has been validated28−38 and applied27−46 to various types of equipment, including gasifiers, consuming a wide range of fuels. Therefore, it was used here for optimizing the gasification and drying units. Details of the mathematical model behind the latest version can be found elsewhere.38 The Exergetic (or second Law) Efficiency could be applied as objective function for dryers and gasifiers. Nonetheless, cold efficiency seems more appropriate as objective function for cases of gasifiers, because the produced gas requires cooling in

Figure 2. Configuration B of the proposed FSIG/GT process. C = compressor, CB = combustor, CD = condenser, CL = cleaning system, CY = cyclone, D = dryer, DF = dried fuel, FE = screw feeding, FS = fuel-slurry pumping, G = gasifier, GT = gas turbine, SG = steam generator, ST = steam turbine, P = water pump, V = valve or splitter.

This new round of studies includes configuration C, illustrated in Figure 3, which dispenses with the use of steam in the gasification process and here applied to a typical biomass, namely, the sugar cane bagasse (SCB). As seen, water is added to the wet SCB (leaving the mill) to form a high dry-solid 2698


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CeSFaMB can be applied to optimize the dryer and gasifier designs and operational conditions to fit cases within a wide range of particle size distribution. (4) Cylindrical (typical of fibrous materials) has be chosen as the basic form of SCB after gridding. (5) Apparent particle density of SCB has been assumed as 720 kg/m3 and real particle densities as 1400 kg/ m3. (6) Consumption rate of moist SCB was set at 36 kg/s (wet basis) or 18 kg/s (dry basis). Such should be the approximate rate of generated by large sugar−alcohol mills. Nevertheless, the process can be scaled up or down to fit consumption rates differing from that assumed here, and such changes should not contradict the main conclusions arrived in the present study. (7) Before injection into the dryer, water is added to the wet SCB in order to form slurry with final drysolid percentage of 40%, which was the value adopted in a previous studies.22−24 Commercially available equipment should be able to pump such sort of slurry.49,50 This has been confirmed by a large piston pump manufacturer (http:// www.schwingbioset.com/), which even provided some assurance that slurries with up to 50% dry-solid content can be pumped into high-pressure vessels. Therefore, the value assumed here might be conservative. (8) The average internal pressure of the gasifier is set at 2 MPa, and the dryer operates at slighter higher pressure to ensure that the dried solid would be able to be fed into the gasifier using commercially available rotary valves combined with Archimedes’ screws. (9) Alkaline species, usually present in combustion and gasification gases, might bring serious problems of combined erosion and corrosion to gas turbine blades.14−17 Those components can be removed by cooling the gas stream to values below their dew points, which fall above 800 K.17 The present work applies that lower limit to ensure proper cleaning of flue gas. Having in mind the very low concentration of alkaline species in the gas stream, it is safe to assume that the energy involved in their condensation is negligible when compared with the energy involved in the heat exchanging to bring the gas leaving the gasifier to temperatures around 800 K. (10) Axial air compressor isentropic efficiencies are equal to 87%. This is likely a conservative value.51 (11) Axial gas turbine isentropic efficiencies assumed as 87%. (12) Steam isentropic efficiencies equal 80%.52 (13) Pump isentropic efficiencies assumed as 90%.53 (14) Minimum temperature difference between parallel streams entering or leaving heat exchangers is taken as 10 K. (15) Maximum injection temperature into turbines is set at 1700 K. Future works might review those assumptions in order to improve the accuracy of simulations. Moreover, equipment dimensions may be further optimized to achieve lower capital costs and higher efficiencies.

order to condense alkaline compounds before its injection into gas turbines. The Industrial Process and Equipment Simulator (IPES©) software has been applied to many previous works,22−25,27,47 particularly to develop and optimize power generation processes. That software has also been used in many R&D projects (www.desouzasantos.info). Balances of mass and energyaccording to the first and second Law of Thermodynamicsare performed around each equipment or control volume. Those provide a matrix with temperatures, pressures, and compositions of streams involved in the whole process. Once solved, the temperature, pressure, composition, and other physical-chemical properties of each stream is printed. In addition, the overall process parameterssuch as efficiencies are computed.

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ASSUMPTIONS The main assumptions used in the present work are listed in what follows: (1) The Bubbling Fluidized Bed technique has been chosen for the fuel gasification and drying; however similar results might be achieved using Circulating Beds, Entrained Flow reactors, or other equivalent processes. Nonetheless, it is important to mention that the Bubbling Fluidization technique allows high flexibility regarding variations in fuel particle sizes, density, and properties,38,48 therefore applicable to various cases of biomass. Furthermore, small rocks and sand are usually carried with the harvested biomass and might enter the dryer and even the gasifier. Proper air or gas distributor designs of bubbling beds allow removing heavier particles that might fall on the bed base, without interrupting the dryer or gasifier operations.38 (2) SCB properties were taken from previous works22−24,30,34,38 and reproduced in Table 1. (3) The SCB particle size distribution was set to provide average diameter around 1.6 mm. That value was reached after preliminary simulations in order to allow good operational conditions for the dryer and gasifier, while not requiring great expenditures in gridding. Nonetheless, Table 1. Main Characteristics of the Fuel (Sugar-Cane Bagasse) Consumed by the Process property high heating value (dry basis) proximate analysis (wet basis) moisture volatile fixed carbon ash ultimate analysis (dry basis) C H N O S Ash particle size distribution sieve opening (mm) 1.680 0.841 0.354 0.250 0.177

value 19.14 MJ/kg 50.00% 40.78% 7.57% 1.65%

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RESULTS AND DISCUSSION Gasifier. The optimization of gasifier operations were conducted using CeSFaMB and having the cold efficiency as an objective function. This choice seems more appropriate than the Hot Efficiency or Exergetic Efficiency, because the produced gas requires cooling in order to condense alkaline compounds before its injection into gas turbines. Once the above assumptions regarding that equipment were set, the simulations proceed by varying the bed and freeboard internal diameter, bed height, and the flow of injected air through the distributor. In addition, the rate of steam injection was added to verify the range of efficiencies when configuration B was applied.

49.66% 5.71% 0.21% 41.08% 0.03% 3.31% retained mass % 82.00 3.91 9.86 3.23 1.00 2699


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Table 2. Gasifier Main Characteristics and Operational Conditions of Gasifiers Operating under Both Configurations main input conditions of parameters

configuration B

bed internal diameter (m) 5.0 bed height (m) 4.0 freeboard internal diameter (m) 5.0 freeboard height (m) 6.0 100 insulation thickness around the bed and freeboard (mm) number of flutes in the distributor 5.0 × 104 number of orifices per flute 10 diameter of orifices (mm) 3.0 fuel feeding position (above the 1.0 distributor) (m) mass flow of feeding fuel (dry) (kg/s) 17.99 mass flow of injected air (kg/s) 11.0 temperature of injected air (K) 766.0 mass flow of injected steam (kg/s) 1.0 temperature of injected steam (K) 728.0 average pressure inside the equipment 2.0 (MPa) main output conditions or parameters

Figure 4. Optimization of gasifier cold efficiency.

After a preliminary search, the bed diameter around 5 m led to the best efficiencies within the present range of fuel characteristics and feeding rate. Then, the investigation tested the influence of air and steam flow ratios on the gasifier efficiencies. Figure 4 illustrates the search. Few points at and most below 10 kg/s of airflow rate led to difficult or unfeasible operations due to very low equivalence ratio or conditions below minimum fluidization limits. More refined optimization grids regarding the bed diameter as well as rates of air and steam flows might be applied in coming works. Nonetheless, it is believed that the conclusions arrived at the present instance would not fall too far from the possible achievable in the future. Figure 4 shows that the maximum cold efficiency was achieved for air flow around 11 kg/s with no steam injection. This is due to the relatively high hydrogen content of the present biomass, which after oxidation provides all the water for important gasification reactions. (The complete set of reactions taken into account by CeSFaMB can be found elsewhere.34,38) Any surplus injected steam dilutes the fuel gas and lowers the temperatures in the bed and freeboard, which in turn decreases the rates of gasification reactions. In addition, as no steam is injected into the gasifier, higher bed temperatures can be achieved, thus leading to higher gasification reaction rates. The best point found here was used to optimize the process under configuration C. Despite that, the work included the best condition with steam injection, which was achieved for a steam flow rate at 1 kg/s combined with airflow of 11 kg/s. Such was used for the configuration B process optimization. The most important input and output parameters for gasifiers operating under both configurations are summarized in Table 2, whereas Table 3 presents the respective produced gas compositions. As expected, the injection of steam (configuration B) led to higher H2 concentration than the achieved at configuration C. However, greater CO concentration was found for the gasification under this last option. Among many factors, the high concentration is mainly due to the higher average temperatures arrived during gasification without steam injections. The resulting heating values (Table 2) and cold efficiency favored the gasification at configuration C. Because the gasification operations do not deviate too much from each other, the figures ahead depict the details of configuration C. The temperature profiles of various phases throughout the bed are shown in Figure 5. (At several positions, the temperatures of many phases coincide.)

mass flow of gas leaving the equipment (kg/s) mass flow of solids discharged from equipment (kg/s) mass flow of elutriated solids (kg/s) fluidization voidage (bed middle) fluidization superficial velocity (bed middle) (m/s) carbon conversion (%) average temperature at the middle of the bed (K) pressure loss at the distributor (kPa) pressure loss in the bed (kPa) TDH-transport disengaging height (m) rate of energy input by fuel to the equipment (MW) total rate of energy input to the equipment (MW) combustion enthalpy of hot gasa (MJ/kg) combustion enthalpy of cold gasb (MJ/kg) rate of energy output by hot gas (MW) rate of energy output by cold gas (MW) hot efficiency (%) cold efficiency (%) exergy flow brought with the dry fuel (MW) exergy flow brought with the injected gas (MW) exergy flow brought with the injected steam (MW) total entering exergy flowc (MW) exergy flow leaving with the gas (MW) total exiting exergyd (MW) ratio between total leaving and entering exergy flows (%) ratio between the exergy leaving with the produced gas and the total entering exergy (%)

configuration C 5.0 4.0 5.0 6.0 100 5.0 × 104 10 3.0 1.0 17.99 11.0 766.0 0.0 2.0

values

values

28.60 1.206 0.184 0.729 0.220 81.43 1066.8 0.01 21.61 4.094 324.98 331.23 10.16 9.43 290.58 256.18 87.73 77.34 491.2 5.33 1.25 497.8 283.4 287.9 57.83

27.62 1.179 0.191 0.730 0.218 81.50 1070.7 0.01 21.81 4.078 324.98 330.38 10.53 9.67 290.85 257.99 88.03 78.09 491.2 5.33 0.00 496.6 283.1 287.3 57.86

56.93

57.00

“Hot gas” refers to the temperature, pressure, and composition as found at the exiting point from the gasifier. b“Cold gas” refers to the gas properties if at 298 K, 101.325 kPa, dry and tar free. cSum of exergies brought by gases, liquids, or solids injected or fed into the gasifier. dSum of exergies carried by gases, liquids, or solids leaving the gasifier. a

Coordinate Z represents the position above the surface of the gas distributor. To clearly illustrate the values near the distributor area (z = 0), the same graph is shown in Figure 6 using logarithm abscissa scale. The difference of temperatures of fuel particles and injected airflowing through the emulsion and bubble phasesat the oxygen-rich region near the 2700


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Table 3. Composition of the Gas Exiting the Gasifiers Operating under Both Configurations molar percentage chemical species

configuration B

configuration C

H2 H2O H2S NH3 NO NO2 N2 N2O O2 SO2 CO CO2 HCN CH4 C2H4 C2H6 C3H6 C3H8 C6H6 tar

25.9167 6.0319 0.0099 0.2185 0.0000 0.0000 22.6838 0.0000 0.0000 0.0017 32.7863 8.1698 0.0112 3.7980 0.1686 0.1311 0.0062 0.0060 0.0606 0.0000

25.1873 4.0749 0.0102 0.2207 0.0000 0.0000 23.5687 0.0000 0.0000 0.0018 36.5104 6.2384 0.0190 3.7816 0.1752 0.1362 0.0065 0.0062 0.0629 0.0000

Figure 7. Concentration of tar (among other gases) throughout the gasifier operating under configuration C.

Figure 8. Few reaction rates in the bed section of the gasifier operating under configuration C. Notation: V = fuel volatiles.

Figure 5. Temperature profiles in the bed region of the gasifier operating under configuration C. Notation: EMULS.GAS = gas in the emulsion phase, BUBBLE = gas in the bubbles, CARBONAC = carbonaceous fuel particles, AVERAGE = average among all phases.

Figure 9. Temperature profiles in the freeboard region of the gasifier operating under configuration C. Notation: GAS = gas in the freeboard, CARBONAC = carbonaceous fuel particles, INERT = inert or ash particles, AVERAGE = average among all phases.

Figure 6. Temperature profiles in the bed region of the gasifier operating under configuration C (abscissa logarithm scale). Notation: EMULS.GAS = gas in the emulsion phase, BUBBLE = gas in the bubbles, CARBONAC = carbonaceous fuel particles, AVERAGE = average among all phases.

Figure 10. Concentration profiles of CO, CO2, and O2 through the gasifier operating under configuration C.

coking of tar released during fuel pyrolysis. Figure 7 illustrates the tar concentration profile throughout the gasifier, and Figure 8 shows the rates of devolatilization as well tar cracking and coking. (The legend shows just short representations of reactions. Detailed stoichiometry and rates computation can

distributor surface are due the fast combustion of the carbonaceous solid. Among many influences, the surge of temperatures around the fuel feeding position is mainly due to the cracking and 2701


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Figure 11. Concentration profiles of CO, CO2, and O2 throughout the gasifier operating under configuration C (abscissa log scale).

Figure 15. Main heterogeneous reaction rates in the emulsion phase of the gasifier operating under configuration C (log−log scale).

Figure 12. Concentration profiles of H2O, H2, and CH4 throughout the entire gasifier operating under configuration C.

Figure 16. Main homogeneous reactions in the emulsion phase of the gasifier operating under configuration C.

Figure 13. Concentration profiles of H2O, H2, and CH4 in the emulsion phase of the gasifier operating under configuration C (log− log scale).

Figure 17. Main homogeneous reactions in the emulsion phase of gasifier operating under configuration C (log−log).

leaving the equipment, thus bringing serious operational and maintenance problems to the gas cleaning system. Figure 9 shows the temperature profiles in the freeboard region. Ash segregated from original carbonaceous fuel is represented as inert. Figures 10 and 11 illustrate the concentration profiles of important gases; the latter at logarithm scale. Throughout the bed region z from 0 to 4 m, the values averaged between the emulsion and bubble phases. As seen, the oxygen concentration decreases rapidly, just above the distributor region. Despite that, the temperatures of those particles are well below the ash-softening point; otherwise, they could agglomerate forming larger blocks that would fall to the bed bottom. If such process were too severe, it would lead to the complete collapse of the bed. Because the particles are in direct contact with the emulsion, its gas phase experiences a faster increase in temperature when compared to particle-free bubbles. Nonetheless, the accumulation of fuel gases (such as hydrogen and carbon monoxide) in the relatively cold bubbles would provide conditions for a

Figure 14. Main heterogeneous reaction rates in the emulsion phase of the gasifier operating under configuration C.

be found elsewhere.34,38) It is also worthwhile to comment on the importance of feeding the fuel at positions near the distributor. Such would allow enough residence time for tar destruction, which otherwise might be present in the gas stream 2702


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Table 4. Main Characteristics and Operational Conditions of Dryers Operating under Both Configurations main input conditions of parameters

configuration B

bed internal diameter bed height freeboard internal diameter freeboard height insulation thickness around the bed and freeboard number of flutes in the distributor number of orifices per flute diameter of orifices slurry feeding position (above the distributor) mass flow of feeding fuel (50% wet) mass flow of water added to form the fuel slurry percentage of dry fuel in the slurry mass flow of injected gas temperature of injected gas average pressure inside the equipment main output conditions or parameters

configuration C

5.0 m 3.0 m 9.0 m 7.0 m 100 mm

5.0 m 3.0 m 9.0 m 7.0 m 100 mm

5 × 104 10 3.0 mm 0.5 m

5 × 104 10 3.0 mm 0.5 m

36.0 kg/s 9.0 kg/s

36.0 kg/s 9.0 kg/s

40.00% 52.0 kg/s 839 K 2.01 MPa values

40.00% 50.0 kg/s 864 K 2.01 MPa values

mass flow of gas leaving the equipment mass flow of solids discharged from the bed concentration of water in the leaving solid fluidization voidage (bed middle) fluidization superficial velocity (bed middle) mixing index in the bed tar flow at the top of the freeboard pressure loss at the distributor pressure loss in the bed exergy flow brought with the slurry exergy flow brought with the injected gas total entering exergy flow exergy flow leaving with the gas total exiting exergy ratio between leaving and entering exergy flows ratio between exergy leaving with gas and total entering

79.01 kg/s 17.99 kg/s 0.0% 0.824 0.283 m/s 1.000 0.000 kg/s 0.06 kPa 2.28 kPa 505.8 MW 32.30 MW 538.1 MW 46.04 MW 335.2 MW 62.29% 8.56%

Figure 19. Temperature profiles in the dryer freeboard region of the dryer operating under configuration C. Notation: GAS = gas flowing through the freeboard, CARBONAC = carbonaceous fuel particles, AVERAGE = average among all phases.

77.01 kg/s 17.99 kg/s 0.0% 0.822 0.276 m/s 1.000 0.000 kg/s 0.06 kPa 2.30 kPa 505.8 MW 31.79 MW 537.6 MW 45.09 MW 334.3 MW 62.17% 8.39%

Figure 20. Concentration profiles of H2O, H2, and CH4 throughout the dryer operating under configuration C.

above the distributor (Figure 11), allowing the accumulation of fuel gasesproduced by gasification reactionsat regions above that location. The average concentration profiles of various other main gases throughout the entire gasifier are presented in Figures 12 and 13. As seen in Figure 12, hydrogen is mostly produced by pyrolysis near the fuel feeding position. The profiles of the main heterogeneous reaction rates in the bed emulsion phase are shown in Figure 14. The same picture is presented in double logarithmic scale in Figure 15 to explicitly illustrate the fast decline of fuel oxidation reaction at regions near the distributor surface (z = 0). Additionally, it reveals that although simultaneous combustion and gasification reactions occur within the oxidation region, no appreciable hydrogen is produced due to its fast oxidation (Figure 13). This emphasizes the importance of feeding the fuel above the oxidation region as well of the pyrolysis in the whole gasification process. The feeding should also take place well below the bed top to provide enough residence time for the development of relatively slow gasification reactions. The main homogeneous reactions are shown in Figure 16. It allows observing the variation of the shift reaction (CO + H2O = H2 + CO2) near the fuel feeding position. That process introduces more hydrogen and carbon monoxide, thus decreasing and even reverting the shift reaction near that location. Figure 17 repeats the picture using logarithm scale for the abscissa as well and illustrates the oxidation of fuel gases (formed by competing gasification reactions) near the air distributor. Dryer. Table 4 summarizes the most important characteristics of the dryer geometry and operational conditions. It should be noticed that configuration B requires a larger flow rate of hot gas to accomplish the drying than configuration

Figure 18. Temperature profiles in the dryer bed region of the dryer operating under configuration C. Notation: EMULS.GAS = gas in the emulsion phase, BUBBLE = gas in the bubbles, CARBONAC = carbonaceous fuel particles, AVERAGE = average among all phases.

sudden combustion and consequent temperature peak of that phase (Figure 6). The two peaks of CO occurred: a small one around 5 mm above the distributor surface (clearly shown in Figure 11) and the other around the fuel feeding position (Figure 10). This last surge on carbon monoxide is mainly due to the fuel devolatilization. Oxygen is completely consumed near 10 mm 2703


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Table 5. Description of Conditions at Each Stream of the Proposed Process (Configuration B) stream 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34

fluid nature air air gasa gas gas gas waterb water water steam steam steam steam steam water water water gas gas water water gasc steam steam water water water water water air air gas slurry slurry

temperature (K) 298.0 763.4 800.0 1700.0 1085.4 350.0 298.0 298.0 351.1 944.8 732.4 732.1 732.1 430.1 330.0 253.9 254.1 349.9 349.9 298.0 298.0 1066.0 950.0 373.6 340.0 340.3 298.0 298.0 331.2 298.0 766.4 839.9 298.1 298.1

mass flow (kg/s)

pressure (kPa) 0.10133 0.20000 0.20000 0.19900 0.14000 0.12000 0.10133 0.13000 0.11000 0.10000 0.24000 0.23500 0.23500 0.14000 0.12000 0.11500 0.10010 0.11500 0.11500 0.10133 0.12000 0.20100 0.10000 0.70000 0.60000 0.10010 0.10133 0.13000 0.11000 0.10133 0.22000 0.22000 0.10133 0.22000

× × × × × × × × × × × × × × × × × × × × × × × × × × × × × × × × × ×

3

10 104 104 104 103 103 103 103 103 105 104 104 104 103 103 103 105 103 103 103 103 104 105 102 102 105 103 103 103 103 104 104 103 104

enthalpy (kJ/kg)d −0.21044 0.48833 −0.32332 −0.35179 −0.84298 −0.16841 −0.15906 −0.15906 −0.15683 −0.12175 −0.12600 −0.12600 −0.12600 −0.13186 −0.15772 −0.16090 −0.16080 −0.16841 −0.16841 −0.15906 −0.15906 −0.27893 −0.12162 −0.13293 −0.15730 −0.15718 −0.15906 −0.15906 −0.15766 −0.22843 0.49169 −0.11397 −0.15906 −0.15904

184.50 184.50 28.60 213.10 213.10 213.10 500.00 500.00 500.00 45.00 45.00 1.00 44.00 44.00 44.00 45.00 45.00 161.10 52.00 1.00 1.00 28.55 3.50 3.50 3.50 3.50 60.00 60.00 60.00 11.00 11.00 52.00 36.00 36.00

× × × × × × × × × × × × × × × × × × × × × × × × × × × ×

10 104 102 103 104 105 105 105 105 105 105 105 105 105 105 105 104 104 105 105 104 105 105 105 105 105 105 105

× × × ×

103 104 105 105

3

entropy (kJ/kg/K) 0.67402 0.68598 0.86796 0.79672 0.80538 0.68239 0.21844 0.21844 0.28732 0.10635 0.10762 0.10771 0.10771 0.11023 0.26123 0.15155 0.15196 0.68359 0.68359 0.21844 0.21844 0.91558 0.10648 0.11075 0.27379 0.27412 0.21844 0.21844 0.26278 0.67164 0.68366 0.69548 0.21844 0.21852

× × × × × × × × × × × × × × × × × × × × × × × × × × × × × × × × × ×

101 101 101 101 101 101 101 101 101 102 102 102 102 102 101 101 101 101 101 101 101 101 102 102 101 101 101 101 101 101 101 101 101 101

a

After cleaning to set alkaline concentration within acceptable levels. bWater = liquid water. cAfter cleaning to set particle size and content within acceptable levels. dEnthalpy values include the formation and sensible terms.

C. This is not only because the higher temperature of Steam 28 of configuration C than stream 32 at configuration B but also because the latter alternative applies steam injection into the gasifier, which leads to higher moisture content in stream 32 (Figure 2) than in stream 28 (Figure 3). Despite the required compression of Steams 19 (Figure 2) and stream 15 (Figure 3), the dryer exiting gas carries less than 9% of the total exergy injected into the dryers. The temperature profiles in the bed and freeboard regions are illustrated in Figures 18 and 19. Figure 20 shows the sudden increase of water vapor concentration near the slurry feeding position. Process. The optimizations of both processes were accomplished through the IPES software. Table 5 presents the main parameters of each streams of configuration B, and Table 6 shows the characteristics of configuration C. Table 7 shows the overall energy balance for both alternatives. It becomes clear that configuration C is more advantageous than B. The main reasons are the following: (1) Despite no steam injection, the gasification under configuration C led to higher cold efficiencies. (2) Unlike configuration B, configuration C does not divert part of produced steam from one of the involved Rankine cycles to be used in the gasification

unit. (3) Configuration B requires more power than configuration C to compress the larger mass flow of gas to the drying unit. Therefore, configuration C is not just more efficient but simpler than configuration B. The achieved efficiency is well above the range of 20%, which is presently obtained from sugar-alcohol mills using Rankine cycles. (Information provided by the R&D team of large boiler manufacturer.54). Additionally, the value attained here surpasses the 33% efficiency estimated for BIG/GT processes12 and are within the range of maxima achieved for processes that usually require pure oxygen as gasification agent.13 It also matches the efficiencies reached at previous proposed alternatives based on complex and costly boilers with highly pressurized combustion chambers.22−24

■

CONCLUSIONS

The FSIG/GT process consuming sugar cane bagasse has been studied. As shown in a previous work,27 that strategy for power generation is advantageous when compared with other alternatives using pressurized gasification5−12 because it allows feeding particulate fuel as slurry, thus simplifying that operation when compared with traditional cascade systems. Moreover, it 2704


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Table 6. Description of Conditions at Each Stream of the Proposed Process (Configuration C) stream 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28

fluid nature air air gasa gas gas gas waterb water water steam steam water water gas gas gasc steam steam water water water water water air air slurry slurry gas

temperature (K) 298.0 763.4 800.0 1700.0 1046.7 350.0 298.0 298.0 350.6 999.8 386.6 330.0 330.3 350.0 350.0 1070.0 854.0 407.2 317.0 317.3 298.0 298.0 308.6 298.0 766.4 298.0 298.1 864.9

mass flow (kg/s)

pressure (kPa) 0.10133 0.20000 0.20000 0.19900 0.12000 0.10500 0.10133 0.13000 0.11000 0.10000 0.60000 0.50000 0.10100 0.10395 0.10395 0.20000 0.10000 0.21000 0.20000 0.10010 0.10133 0.13000 0.11000 0.11000 0.22000 0.10133 0.22000 0.22000

× × × × × × × × × × × × × × × × × × × × × × × × × × × ×

3

10 104 104 104 103 103 103 103 103 105 102 102 105 103 103 104 105 103 103 105 103 103 103 103 104 103 104 104

enthalpy (kJ/kg)d −0.21044 0.48833 −0.28487 0.37982 −0.80965 −0.15993 −0.15906 −0.15906 −0.15685 −0.12043 −0.13268 −0.15772 −0.15760 −0.15993 −0.15993 −0.24063 −0.12393 −0.13233 −0.15826 −0.15815 −0.15906 −0.15906 −0.15861 −0.22843 0.49169 −0.15906 −0.15904 −0.10280

188.50 188.50 27.62 216.12 216.12 216.12 500.00 500.00 500.00 45.00 45.00 45.00 45.00 166.12 50.00 27.63 3.50 3.50 3.50 3.50 200.00 200.00 200.00 11.00 11.00 36.00 36.00 50.00

× × × × × × × × × × × × × × × × × × × × × ×

10 104 102 103 104 105 105 105 105 105 105 105 104 104 104 105 105 105 105 105 105 105

× × × ×

103 105 105 104

3

entropy (kJ/kg/K) 0.67402 0.68598 0.86460 0.79381 0.80303 0.68450 0.21844 0.21844 0.28672 0.10770 0.11212 0.26123 0.26157 0.68478 0.68478 0.91217 0.10394 0.10727 0.24436 0.24470 0.21844 0.21844 0.23312 0.67164 0.68366 0.21844 0.21852 0.69700

× × × × × × × × × × × × × × × × × × × × × × × × × × × ×

101 101 101 101 101 101 101 101 101 102 102 101 101 101 101 101 102 102 101 101 101 101 101 101 101 101 101 101

a

After cleaning to set alkaline concentration within acceptable levels. bWater = liquid water. cAfter cleaning to set particle size and content within acceptable levels. dEnthalpy values include the formation and sensible terms.

The present work has been applied to the case of units consuming sugar cane bagasse, but can be employed to other biomass as well as coal and residues. Furthermore, future investigations should explore the influence of drying and gasification operational pressures on the overall process efficiency.

Table 7. Overall Efficiency main parameter

configuration B

configuration C

mechanical power inputa mechanical power outputb net mechanical power output energy rate input by fuelc efficiency based on 1st Lawd

124.46 MW 221.03 MW 96.57 MW 306.06 MW 31.55%

126.71 MW 241.27 MW 114.56 MW 306.06 MW 37.43%

■

a

Due to compressors and pumps. bFrom steam and gas turbines. Based on LHV. dDefined as (useful mechanical power output)/(rate of energy input by fuel).

AUTHOR INFORMATION

Corresponding Author

c

*E-mail: [email protected]. Fax: +55-19-3513278. Tel.: +55-1997107134. Notes

The authors declare no competing financial interest.

■

dispenses the need of pure oxygen or hydrocarbons required to ignite the fuel particles when injected as slurry.26 Among the alternatives of FSIG/GT studied here, the best choice is the one dispensing the injection of steam into the gasifier, or configuration C. It is not just more efficient but also simpler than configuration B, which should provide savings in capital as well operational and maintenance costs. The achieved efficiency for configuration C is well above the range of 20% that is presently obtained from sugar-alcohol mills using Rankine cycles. Additionally, the value attained here surpasses the 33% efficiency estimated for BIG/GT processes12 and are within the range of maxima achieved for processes that usually demand pure oxygen as gasification agent.13 It also matches the efficiencies reached at previous proposed alternatives based on complex and costly boilers with highly pressurized combustion chambers.22−24

ACKNOWLEDGMENTS The authors are grateful for the grant provided by CAPES (Brazilian Federal Agency for the Support and Evaluation of Graduate Education).

■

REFERENCES

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