Gas Turbine

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New Developments on Fuel-Slurry Integrated Gasifier/Gas Turbine (FSIG/GT) Alternative for Power Generation Applied to Biomass; Configuration Requiring No Steam for Gasification Marcio L. de Souza-Santos* University of Campinas, Faculty of Mechanical Engineering Department of Energy, Campinas, São Paulo, Brazil

Andres F. B. Bernal University of Campinas, Faculty of Mechanical Engineering Department of Energy, Campinas, São Paulo, Brazil

Andres F. Rodriguez-Torres University of Campinas, Faculty of Mechanical Engineering Department of Energy, Campinas, São Paulo, Brazil ABSTRACT: The work presents additional improvements on the strategy of the Fuel-Slurry Integrated Gasifier/Gas Turbine (FSIG/GT) concept for electric power generation applied to the case of Sugar Cane Bagasse (SCB). The process allows feeding solid fuel particles into pressurized reactors without the need for complex systems of sequential lock-hoppers. For that, the particulate fuel is mixed with water to produce a high dry solid content slurry that is pumped into a pressurized dryer. Since both the bubbling fluidized bed dryer and gasifier operate at similar pressures, the transference of dried particles to the gasifier can be accomplished with simple Archimedes screws. As the particles are dried before gasification, the process also circumvents the difficulties of fuel slurry ignition, which usually require pure oxygen combined with hydrocarbons. The gas stream leaving the gasifier is cleaned and injected into a gas turbine combustor or combustors. Heat recovering is accomplished by Rankine cycles. The present phase follows preliminary studies where the rates of air and steam injections were variables on the optimization of the dryer and gasifier. Now, the studies include the dryer and gasifier diameters as additional variables for optimizations. The work considers exergetic and cold efficiencies as objective functions. The investigation demonstrates the possibility of achieving higher overall power generation efficiency, which surpasses that achieved by other strategies such as high-pressure Rankine, Biomass Integrated Gasification/Gas Turbines, and combined cycles using pressurized-chamber boilers. The present work concentrates on a possible configuration for the FSIG/GT process that does not require steam for fuel gasification; thus just compressed air is injected into that equipment. Future works will explore other alternatives.

1. INTRODUCTION Biomass as a renewable source of power generation with near zero overall greenhouse gas emissions1−5 has increased, particularly in the case of Sugar Cane Bagasse (SCB).6−11 For instance, today near 10 GWwhich represents around 8% of the Brazilian electric power capacitycomes from thermoelectric units based on sugar cane bagasse.11 That represents almost 5 times the amount generated in 2006.10,11 So far, all that generation applies high-pressure Rankine cycles, but much more efficient Biomass Integrated Gasification/Gas Turbines (BIG/GT) have been devised.12−20 On the other hand, that last process introduces technical barriers, among them the gas cleaning required to decrease particulates and alkaline contents in order to meet acceptable levels for injection into turbines.21−24 Such a hurdle has already been surpassed.25,26 Nonetheless, another significant technical obstacle, represented by the problem of feeding solid particulate fuels into pressurized vessels, remains. Feeding systems based on two or more levels of pressurized lock hoppers are usually applied to such cases.27 Yet, that alternative involves cumbersome and noncontinuous operations, which are susceptible to stoppages. © XXXX American Chemical Society

In the case of biomasses, the most common problem is the interlacing of fibers of neighboring particles leading to domes inside the hopper that prevent the particles from dropping into rotating valves connecting two hoppers. Moreover, those arrangements require cooling inert gases, which have to be introduced into the hoppers to avoid fuel pyrolysis or even combustion. Those difficulties and measures to avoid them increase the power unit invested capital as well operational and maintenance costs. Those problems can be avoided when the fuel is mixed with water to form a slurry, which is just pumped into the pressurized vessel. Such an alternative has been used for some time28 and has been employed in prior studies evaluating proposals of processes employing complete fuel combustion.29−32 However, injecting the fuel slurry directly into a gasifier leads to critical problems, among them, the need to evaporate the slurry water demands burning a significant fraction of the carbonaceous fuel, thus significantly decreasing Received: April 10, 2015 Revised: May 18, 2015

A

DOI: 10.1021/acs.energyfuels.5b00775 Energy Fuels XXXX, XXX, XXX−XXX

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conditions of the dyer and the gasifier. That software has been validated37−47 and applied34−55 to various categories of equipment consuming a wide range of fuels. A summary of the basic assumptions and strategy behind the simulator as well lists of its required inputs and provided outputs can be found elsewhere.34,47 The simulation of remaining equipment involved in the process and their combined operations was possible using IPES or Industrial Process and Equipment Simulator. Such a tool has been employed in many previous works.29−32,34−36,56 IPES routines include the mass and energy balances around each piece of equipment or control volume. When applied to the whole process, a matrix with temperatures, pressures, and compositions of streams is generated. After solving the system of equations, the various properties of each stream are printed as well overall process parameters, among them the efficiencies according to the first and second laws of thermodynamics. Figure 1. Scheme of the studied FSIG/GT process. C = compressor, CB = combustor, CD = condenser, CL = cleaning system, CY = cyclone, D = dryer, DF = dried fuel, FE = screw feeding, FS = fuelslurry pumping, G = gasifier, GT = gas turbine, SG = steam generator, ST = steam turbine, P = water pump, V = valve or splitter.

3. PREMISES The main assumptions used in the present work are listed below: Table 1. Main Characteristics of the Fuel (Sugar-Cane Bagasse) Consumed by the Process

the portion that could be gasified to produce fuel gases. Frequently, pure oxygen or hydrocarbons are employed33 in such situations, but that entails additional costs related to the power generation processes. Those difficulties can be avoided by drying the slurry prior to gasification. The FSIG/GT applies that strategy, which may be accomplished by various alternatives.34−36 The present work concentrates on one that does not require steam for the gasification stage. Its scheme is shown in Figure 1. According to that, stream 16, leaving the pressurized gasifier (G), is cleaned to decrease the particle content to levels acceptable by gas turbine injections. Its temperature is also decreased by exchanging heat with water in equipment 11 in order to reach the dew points of alkaline species. Such heat exchanging drives a Rankine cycle composed by equipment 11 to 15. The gas cooling brings its alkaline content also within the range required by gas turbines. Stream 3 is then injected into the gas turbine combustor (equipment 2). Another Rankine cycle (equipment 4 to 8) is used to recover energy from stream 5 leaving the gas turbine (equipment 3). Stream 15, which is a fraction of stream 6, has its temperature raised by compression, and the resulting stream 28 is employed in the fuel slurry drying stage at the bubbling fluidized bed dryer (D). The slurry is prepared by adding water to the wet fuel particles and pumped by equipment 17 into the drier. From that, the dried particles are transferred to the bubbling fluidized bed gasifier (G). Since the dryer and gasifer operate at similar pressures, that transference can be accomplished using simple rotary valves and Archimedes screws.

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 average particle sphericity

value 19.14 MJ/kg 50.00% 40.78% 7.57% 1.65% 49.66% 5.71% 0.21% 41.08% 0.03% 3.31% retained mass % 82.00 3.91 9.86 3.23 1.00 0.3

(1) The alternative of bubbling fluidized bed was chosen for the gasification and drying processes. Nonetheless, other techniques might be employed, among them circulating fluidized bed and entrainment flow. However, it should be noticed that bubbling fluidization permits great flexibility in terms of fuel composition, density, particle size distribution, and other properties.47,57 Additionally, the bagasse leaving the mill usually contains small rocks that have been carried during the sugar cane harvesting. Such materials might enter the dryer and even proceed to the gasifier. This represents no major problem because a properly designed gas distributor would be able to remove heavier particles falling on its surface47 without the need of interrupting the bubbling bed drying or gasifying operations.

2. METHODOLOGY Like all evaluations of industrial processes, the one described here required extensive simulations. Due to the close coupling of various operations, simulations and optimizations of dimensions and operational conditions of the dryer and gasifier should be combined with simulations and optimizations of the whole process. The Comprehensive Simulator of Fluidized and Moving Bed equipment (CeSFaMB; http://www.csfmb.com) was applied during the tasks of optimization of dimensions and operational B


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Energy & Fuels (2) The properties of SCB properties have been taken from the literature as well as employed in previous works.29−31,39,43,47 The main characteristics are listed in Table 1. The particle size distribution (PSD) is not too far from that typically found in SCB leaving sugar-alcohol mills. Therefore, that would preclude expenditures on grindings. Nevertheless, future works might include the PSD as an optimization variable. (3) The shape of cylindrical fibers has been selected for the feeding SCB. (4) The apparent and real densities of SCB particles have been taken as 720 kg/m3 and 1400 kg/m3, respectively. (5) The SCB consumption rate was set at 36 kg/s (wet basis) or 18 kg/s (dry basis), which is within the range yielded by sizable sugar-alcohol industrial units. However, the results depicted here can be adapted in cases of larger or smaller scales. Additionally, the conclusions arrived at this instance should not be significantly affected for different scales. (6) Liquid water is added to the SCB to produce the slurry with 40% of dried solid. Such a condition has been applied in prior investigations.29−31,35 The literature58,59 allows some assurance that such a slurry could be pumped using commercially available equipment. Moreover, a large piston pump manufacturer [http://www.schwingbioset.com/] has expressed some confidence that even slurries with 50% dry biomass could be injected into reactors operating at high pressures. Thus, the present assumption regarding the dry solid content in the slurry might even be conservative. (7) A total of 2 MPa has been used as average pressure in the gasifier. The dryer should operate somewhat above that value to allow the transference of dried particles from it to the gasifier. Such transference can be accomplished with commercial rotary valves and Archimedes’ screws. The assumed pressure level has been used in previous works29−32,34−36 as well. However, future studies would include the dyer and gasifier operational pressures as variables. (8) Common gases leaving SCB combustion or gasification reactors contain alkaline species. The levels of such contaminants should be dropped to allow injections into gas turbines.21−24 That can be achieved through cooling in order to drop the gas stream temperature to values below the dew points of those alkaline species, which occurs at temperatures above 800 K.24 To be on the conservative side, that value has been set for the gas temperature after the cleaning process. (9) The assumed isentropic efficiencies are listed in the literature as follows: (a) 87% for axial air compressors;60 the same value has been set for the gas turbine (b) 80% for steam turbines61 (c) 90% for pumps62 (10) A minimum temperature difference of 10 K has been imposed between exchanging heat streams. (11) A temperature of 1700 K was set as the maximum gas turbine injection temperature. Those assumptions might be reviewed in upcoming works, but the main conclusions arrived at the present instance should not be drastically modified if modest changes on the above were taken.

Figure 2. Exergetic efficiencies for gasification processes.

Figure 3. Cold efficiencies for gasification processes.

ηexe =

exergy of exiting gas total exergy of streams (gaseous or solids) entering the gasifier

(1)

The exergy of each stream is represented by its mass flow (kg/ s) times the respective specific exergy (kJ/kg). (b) The cold gas efficiency, or just cold efficiency, defined by ηcold =

enthalpy of exiting gas at 298 K total enthalpy of streams (gaseous or solids) entering the gasifier

(2)

The enthalpy of each stream is represented by its mass flow (kg/s) times the respective specific enthalpy (kJ/kg). The enthalpies include the low heating values of streams (when appropriate) plus their sensible terms. Those objective functions have been applied in preceding works34−36,63 as well. This work improves on the developed before,35 for it includes the bed diameter as variable of gasifier and dryer optimization studies. It should be noticed that viable bubbling bed operations are only possible within ranges of operational variables because the following situations might arise:47 • There are conditions that would result in operations outside the range of bubbling fluidization such as superficial

4. RESULTS AND DISCUSSION 4.1. Gasifier. The CeSFaMB software was employed for the gasifier optimization. For that, two parameters have been adopted as objective functions: (a) The exergetic efficiency given by C


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Energy & Fuels Table 2. Gasifier Main Characteristics and Operational Conditions main input conditions of parameters bed internal diameter (m) bed height (m) freeboard internal diameter (m) freeboard height (m) insulation thickness around the bed and freeboard (mm) number of flutes in the distributor number of orifices per flute diameter of orifices (mm) fuel feeding position (above the distributor) (m) mass flow of feeding fuel (dry) (kg/s) mass flow of injected air (kg/s) temperature of injected air (K) mass flow of injected steam (kg/s) average pressure inside the equipment (MPa) main output conditions or parameters mass flow of gas leaving the equipment (kg/s) mass flow of solids discharged from the bed (kg/s) mass flow of elutriated solids (kg/s) fluidization voidage (bed middle) minimum fluidization velocity (bed middle) (m/s) fluidization superficial velocity (bed middle) (m/s) carbon conversion (%) average temperature at the middle of the bed (K) average temperature at the top of freeboard (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) rate of energy output by hot gasa (MW) rate of energy output by cold gasb (MW) combustion enthalpy of the cold gas (MJ/kg) hot efficiency (%) cold efficiency (%) exergy flow brought with the dry fuel (MW) exergy flow brought with the injected gas (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 (%)

HCE operation

HEE operation

5.5 4.0 5.5 6.0 100

5.0 4.0 5.5 6.0 100

5.0 × 104 10 3.0 1.0 17.99 10.0 766.0 0.0 2.0 values

Table 3. Composition of the Gas Exiting the Gasifier molar %

5.0 × 104 10 3.0 1.0 17.99 10.0 766.0 0.0 2.0 values

26.56 1.695 0.095 0.672 0.092 0.150 81.32 936.0 936.0 0.01 22.84 3.88 324.98 329.89 286.21 260.23 10.03 86.76 78.88 491.2 4.84 496.1 280.42 284.70 57.39

26.53 1.636 0.178 0.724 0.101 0.215 81.04 1098.3 1097.9 0.01 22.03 4.02 324.98 329.89 292.86 259.75 10.13 88.78 78.74 491.2 4.84 496.1 284.77 289.60 58.39

56.53

57.41

chemical species

HCE operation

HEE operation

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

26.7339 2.7154 0.0105 0.2338 0.0000 0.0000 22.0688 0.0000 0.0000 0.0018 36.7761 6.8710 0.0142 4.1758 0.1804 0.1403 0.0067 0.0064 0.0648 0.0000

25.6173 3.9878 0.0106 0.2231 0.0000 0.0000 22.0747 0.0000 0.0000 0.0017 37.9341 5.6701 0.0233 4.0586 0.1805 0.1403 0.0067 0.0064 0.0648 0.0000

withdrawals of solid particles using overflow pipes. However, at the given rates of fuel feeding and conversion, if the rate of particle elutriation is too high, the bed height might drop below the position where the top of the overflow pipe is situated. Therefore, the bed height cannot be maintained constant. The simulations have shown that the ranges of bed internal diameter and rate of air injection into the gasifier studied here are within the conditions of bubbling bed operational feasibility. From this point, the results achieved using the gas from gasifier operating under the highest exergetic efficiency would be called HEE, while it would be called HCE when the highest cold efficient gasification is applied. The range of injected air mass flow covered values between 10 and 16 kg/s and internal bed diameters between 3.5 and 6 m. Those conditions were found to be within the field of viable operations as well where HEE and HCE operations occurred. Figures 2 and 3 display the results of optimum searches having the gasifier exergetic and cold efficiencies as objective functions. As seen, HEE and HCE occur at smaller air injection rates. At such situations, the superficial velocities approach the minimum fluidization values, and smaller bubbles are found in the bed. Such allows for intense mass transfers between bubbles and emulsion, which leads to quick consumption of the oxygen entering the gasifier. This leaves larger portions of the bed at reducing conditions and provides higher concentrations of combustible species in the produced gas.47 HEE operation was found for a 5.0 m internal diameter gasifier to which air is injected at 10 kg/s, while HCE was found for the 5.5 m internal diameter gasifier injected with the same airflow rate. Of course, that search could be refined, but such should not significantly change the results and conclusions arrived at the present work. The gasifier main geometric characteristics and respective operational conditions are listed in Table 2. Table 3 shows the produced gas composition obtained at those operations. As seen, despite using only air as a gasifying agent, the operations should provide very good gas quality. Such results can be understood by the relatively high

“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

velocities below the minimum fluidization one at any point of the bed or velocities above the second turbulent limit. • The temperature is beyond the ash-melting limit, which would lead to particle agglomerations followed by bed collapsing. • There is an impossibility of achieving steady-state operation. For instance, in many bubbling fluidizations, the bed height is usually kept constant by continuous or batch D


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Figure 4. Temperature profiles in the bed region of the gasifier. HCE operation. Notation: EMULS. GAS = gas in the emulsion phase, BUBBLE = gas in the bubbles, CARBONAC = carbonaceous fuel particles, AVERAGE = average among all phases.

Figure 5. Temperature profiles in the freeboard region of the gasifier. HCE operation. Notation: GAS = gas in the freeboard, CARBONAC = carbonaceous fuel particles, INERT = inert or ash particles, AVERAGE = average among all phases.

above regarding operations near the minimum fluidization velocity leading to small bubbles, which bestow intense mass transfers between bubbles and emulsion with consequently quick consumption of the oxygen injected in the gasifier. Additionally, the relatively high temperature of the injected air (stream 25) also contributes to the fast combustion reactions. Figure 9 shows the concentration profiles of H2O, H2, and CH4. It is possible to observe a sharp increase on fuel gas concentrations near the fuel feeding position. That is mainly due to fuel pyrolysis. More details on the role played by that process and other homogeneous and heterogeneous chemical reactions can be found elsewhere.47 4.2. Dryer. The dryer bed internal diameter and the gas flow injected into it were used as variables during the drying process optimization. That optimization aimed to minimize the required gas flow to complete drying of the feeding slurry. This would provide power savings on the compression of

hydrogen content of the bagasse, which after combustion provides enough water to important gasification reactions. The HCE case is used to illustrate a few details of the gasification process. Figure 4 presents the temperature profiles of various phases in the gasifier bed section, while the profiles at the freeboard region are portrayed in Figure 5. The temperature increase around the fuel feeding position (centered at 1 m above the distributor) is mainly due to the slightly exothermic process of pyrolysis, such as tar cracking and coking. The tar evolution and complete destruction by cracking and coking can be observed in Figure 6. The concentration profiles of CO, CO2, and O2 throughout the equipment are shown in Figure 7. The same profiles are repeated in Figure 8 using logarithmic scale abscissa to allow a better view of those profiles near the bottom of the bed and the almost total consumption of all oxygen very near the air distributor surface (z = 0). This confirms the observation made E


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Figure 6. Concentration of tar (among other gases) throughout the gasifier. HCE operation.

Figure 7. Concentration profiles of CO, CO2, and O2 through the gasifier. HCE operation.

therefore outside the range of bubbling bed as assumed here. Turbulent regimes are characteristic of circulating fluidized beds and might be the aim of future investigations. As either HEE or HCE options provided similar dryer operations, Table 4 only describes the main operational parameters using the HEE operation. From that, it is possible to observe relatively low exergy carried by the leaving gas stream when compared with the amount entering the equipment. This favors the presently proposed power generation process. Figures 10 and 11 refer to the case when the HEE gasifier operation was applied to the power generation process. Figure 10 shows the temperature profiles in the dryer fluidized bed, while Figure 11 illustrates the sudden increase of moisture in the gas flowing upwardly in the dryer. Almost all drying takes place between the slurry feeding position (0.5 m) and 3.0 m above the distributor. 4.3. Process. After the drying and gasifier optimization, new rounds of optimizations were carried using the IPES software.

stream 15 (Figure 1). The properties and temperature of stream 28 were found after the simulation of the whole process using the properties of the gas leaving the gasification step, both under HEE and HCE options. Details of the simulation results achieved for the whole process are described in the next section. It should be noticed that, below a critical minimum flow of hot gas injection, only partial removal of moisture is possible. However, relatively small increases of that flow provide the conditions for almost complete drying of the feeding slurry. Of course, more refined grids for the mass flow of injected gas would allow observing further intermediary drying operations. Nonetheless, such refinements should not drastically change the conclusions arrived at here. After the dryer optimizations, both aiming HEE and HCE objectives led to similar results with a dryer bed internal diameter of 3.8 m and a minimum flow rate of 46 kg/s of gas (stream 28) to complete the drying. Smaller bed internal diameters entailed a highly turbulent fluidization regime, F


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Figure 8. Concentration profiles of CO, CO2, and O2 through the gasifier. HCE operation (logarithm scale abscissa).

Figure 9. Concentration profiles of H2O, H2, and CH4 throughout the entire gasifier. HCE operation.

Variations of water flow rates running through both Rankine cycles were tested with the objective of maximizing the overall power generation efficiency. Iterative rounds of dryer and process optimization were also conducted. As seen before, the HEE gasification provided gas with a greater temperature than HCE operation. The reverse occurred regarding the gas heating value at 298 K. To decide which option would provide the best overall power efficiency, tests were performed by simulating the whole process using each gasification optimum. Those demonstrated that the HEE operation led to the highest overall efficiency. This is because the gas exiting the gasifier requires cooling to condense alkaline compounds. Higher gas temperature provides larger heat recovering rates that drive the combined Rankine cycle (equipment 11 to 14 in Figure 3). That advantage overcame the slightly greater power delivered by the gas turbine (equipment 3) in the case when HCE gasification was applied. Using the gas achieved at HEE gasification as well the best drying operation, the whole process was simulated again, and

the main properties of the streams are exhibited in Table 5. The same was done using the gas from the HCE operation. Table 6 summarizes the parameters used to compute the overall process efficiencies in both situations. The present optimizations allowed increasing the efficiencies achieved in prior works.35,63 It is important to notice that the first law efficiency attained here is well above the range of 20%, which is presently obtained at sugar-alcohol mills using Rankine cycles [information provided by the R&D team of a large boiler manufacturer64]. Additionally, it surpasses the 33% level estimated for BIG/GT processes19 and is within the range of maxima achieved for processes that usually require pure oxygen as a gasification agent.33 The current results also are comparable to the efficiencies reached at previously proposed alternatives based on complex and costly boilers with highly pressurized combustion chambers.29−31 Finally, it should be noticed that the high temperature heat exchanging operations of equipment 11 are possible due to materials specially designed for such conditions.65−67 G


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cane bagasse has been developed. The process applies a combination of pressurized drying and gasification of the feeding biomass to provide fuel gas to gas turbines. Heat recovering is achieved through cycles based on steam turbines. The feeding of biomass as slurry simplifies and very likely brings savings to the feeding of particulate fibrous fuels into pressurized vessels when compared to usual feeding methods that apply sequences of pressurized lock-hoppers. At this stage, the gasifier and dryer dimensions were included as variables as well as compared to the operations using gas from the best exergetic and the best cold efficient gasification. The results show a significant improvement on the power unit overall first law efficiency, now reaching values above 38%. This almost doubles the current 20% attained at modern sugar cane mills as well as surpasses the 33% estimated for BIG/GT processes using the same fuel. Despite applying just air as a gasification agent, FSIG/GT matches the efficiencies from elaborate and costly alternatives such as those employing pure oxygen, which usually operate at very high gasification temperatures. Apart from studies of other FSIG/GT alternatives,34−36 forthcoming investigations would test the effects of the gasifier and dryer pressure levels, feeding fuel particle sizes, as well as the limits of dry solid content of pumpable slurries on the overall power efficiency. In addition, improvements on the Rankine cycles might also contribute to further improvements on the overall power efficiency. Other gasification techniques, such as circulating fluidized bed and entrained flow reactor, which could replace the bubbling alternative, might be tested as well.

Table 4. Main Characteristics and Operational Conditions of the Dryer under HEE Operation main input conditions of parameters

value

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

3.8 m 3.0 m 5.0 m 7.0 m 100 mm 4 × 104 10 3.0 mm 0.5 m 36.0 kg/s 9.0 kg/s 40.00% 46.0 kg/s 865 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) 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

73.00 kg/s 17.96 kg/s 0.0% 0.876 0.463 m/s 0.000 kg/s 0.08 kPa 1.90 kPa 505.8 MW 29.21 MW 535.0 MW 43.75 MW 333.3 MW 62.29% 8.18%

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AUTHOR INFORMATION

Corresponding Author

*Tel.: +55-19-97107134. Fax: +55-19-3513278. E-mail: dss@ csfmb.com.

5. CONCLUSIONS A new round of optimizations of the FSIG/GT (Fuel-Slurry Integrated Gasifier/Gas Turbine) process consuming sugar

Notes

The authors declare no competing financial interest.

Figure 10. Temperature profiles in the dryer bed region of the dryer. HHE operation. Notation: EMULS. GAS = gas in the emulsion phase, BUBBLE = gas in the bubbles, CARBONAC = carbonaceous fuel particles, AVERAGE = average among all phases. H


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Figure 11. Concentration profiles of H2O, H2, and CH4 throughout the dryer. HHE operation.

Table 5. Description of Conditions at Each Stream under HEE Operation stream

fluid nature

temperature (K)

pressure (kPa)

mass flow (kg/s)

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

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

298.0 763.4 800.0 1700.0 1046.1 350.0 298.0 298.0 331.1 1010.1 391.7 330.0 330.3 350.0 350.0 1097.0 954.2 364.1 317.0 317.3 298.0 298.0 339.1 298.0 766.4 298.0 298.1 865.4

0.10133E3 0.20000E4 0.19900E4 0.19800E4 0.12000E3 0.10500E3 0.10133E3 0.13000E3 0.11000E3 0.10000E5 0.60000E2 0.50000E2 0.10100E5 0.10395E3 0.10395E3 0.20000E4 0.10000E5 0.60000E2 0.50000E2 0.10010E5 0.11000E3 0.13000E3 0.11000E3 0.11000E3 0.22000E4 0.11000E3 0.22000E4 0.22000E4

191.50 191.50 26.53 218.03 218.03 218.03 800.00 800.00 800.00 45.00 45.00 45.00 45.00 172.03 46.00 26.53 3.50 3.50 3.50 3.50 50.00 50.00 50.00 10.00 10.00 36.00 36.00 46.00

Table 6. Process Overall Efficiencies main parameters a

mechanical power input mechanical power outputb net mechanical power output rate of energy input by fuel slurryc efficiency based on 1st lawd rate of exergy input by fuel slurry efficiency based on 2nd lawe

HCE operation

HEE operation

125.64 MW 242.42 MW 116.78 MW 306.06 MW 38.16% 505.80 MW 23.09%

125.42 MW 244.41 MW 118.99 MW 306.06 MW 38.88% 505.80 MW 23.52%

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 slurry). eDefined as (useful mechanical power output)/(rate of exergy input by fuel slurry).

c

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REFERENCES

(1) Roth, I. F.; Ambs, L. L. Incorporating Externalities into a Full Cost Approach to Electric Power Generation Life-cycle Costing. Energy 2004, 29, 2125−2144 DOI: 10.1016/j.energy.2004.03.016. (2) Basu, P. Biomass Gasification, Pyrolysis and Torrefaction. Practical Design and Theory, 2nd ed.; Elsevier: San Diego, 2013. (3) Krzywański, J.; Rajczyk, R.; Nowak, W. Model research of gas emissions from a lignite and biomass co-combustion in a large scale CFB boiler. Chem. Proc. Eng. 2014, 35 (2), 217−231. (4) Krzywanski, J.; Rajczyk, R.; Bednarek, M.; Wesolowska, M.; Nowak, W. Gas emissions from large scale circulating fluidized bed boilers burning lignite and biomass. Fuel Process. Technol. 2013, 116, 27−34. (5) Werther J. Potentials of biomass co-combustion in coal-fired boilers. 20th International Conference on Fluidized Bed Combustion, 2009, Xi’an China.China, May 18−21, 2009; pp 27−42. (6) Maues J. A. Optimization of Power Generation from Bagasse and Sugarcane Waste in a Brazilian Sugar and Alcohol Mill. Proceedings of ISES World Congress 2007 (Vol. I−V), 2009; pp 2444−2448. link. springer.com/chapter/10.1007%2F978-3-540-75997-3_495#page-1 (accessed on 03/03/2014). (7) Macedo, I. C.; Seabra, J. E. A.; Silva, JEAR. Greenhouse gases emissions in the production and use of ethanol from sugarcane in Brazil: The 2005/2006 averages and a prediction for 2020. Biomass Bioenergy 2008, 32 (7), 582−595 DOI: 10.1016/j.biombioe.2007.12.006.

a

After cleaning to set alkaline concentration within acceptable levels. Water = liquid water. cAfter cleaning to set particle size and content within acceptable levels.

b

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


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