Modeling and Analysis of an Indirect Coal Biomass to Liquids Plant

Article pubs.acs.org/EF

Modeling and Analysis of an Indirect Coal Biomass to Liquids Plant Integrated with a Combined Cycle Plant and CO2 Capture and Storage Yuan Jiang and Debangsu Bhattacharyya*

Downloaded by NEW YORK MEDICAL COLLEGE on September 9, 2015 | http://pubs.acs.org Publication Date (Web): August 11, 2015 | doi: 10.1021/acs.energyfuels.5b00360

Department of Chemical Engineering, West Virginia University, Morgantown, West Virginia 26505, United States ABSTRACT: A model of an indirect coal biomass to liquids (CBTL) plant integrated with CO2 capture and storage (CCS) and a combined cycle plant has been developed in Aspen Plus 7.3.2. In the CBTL plant, syngas produced in the gasifier is sent to a water-gas shift (WGS) reactor to obtain the desired H2/CO ratio in the syngas for Fischer−Tropsch (FT) synthesis. The product upgrading section is configured to satisfy the desired product specifications. CO2 generated in the syngas and syncrude production units are captured by different CO2 capture technologies. Fuel gas and steam generated in the plant are used as either utilities or sent to the combined cycle plant for power generation. The CBTL plant includes a novel integrated hydrotreating unit, which has the potential to reduce both operating and capital costs of the product upgrading section, in comparison to conventional separate hydrotreating units for gasoline and diesel. Impacts of a number of key input and operational variables, such as the steam/carbon ratio in the autothermal reformer (ATR) inlet, H2/CO ratio in the FT inlet, extent of carbon capture in the CCS unit, biomass/coal ratio, and biomass type in the feedstock, are evaluated in this paper. It is observed that the thermal efficiency can be increased by decreasing the biomass/coal ratio, increasing the H2/CO ratio in the FT inlet stream, and/or decreasing the extent of CCS. The thermal and carbon efficiencies of the base case are about 46.6 and 36.4%, respectively, which are slightly higher than the data reported in the literature for a similar product yield and extent of CO2 capture. with CCS,15−17 the impact of key global design parameters, such as the H2/CO ratio in the FT inlet stream, is not well evaluated. In a previous study of Jiang and Bhattacharyya,1 it has been shown that the penalty of the CCS section in a once-through indirect liquefaction plant can be significantly reduced by increasing the H2/CO ratio in the FT inlet stream. However, because the product upgrading section and utilization of the light hydrocarbon product and unconverted syngas were not considered in the previous study, the effect of the H2/CO ratio on the fuel yield and plant efficiency was not addressed. Hence, further studies are conducted in this paper for a better understanding of the impact of the H2/CO ratio. In the indirect liquefaction plant, as the H2/CO ratio is increased, more light hydrocarbons are produced in the FT reactor. A portion of these light hydrocarbons can be used to produce gasoline through the C3−C5 alkylation unit and C4 isomerization unit. However, consideration of these technologies may not be desired for small-scale FT plants, because these technologies add to the complexity and are expected to have low-temperature distillation systems with large penalty.18,19 On the other hand, the light hydrocarbons can be either used as process fuel in the furnace, sent to the combined cycle for power production, or recycled back to the FT reactor through an autothermal reformer (ATR) to produce syngas, which, in turn, increases the fuel yield.20 In the ATR, a combination of exothermic partial oxidation and endothermic steam reforming reactions usually take place on Ni/Al2O3 catalysts at thermally neutral conditions.21 Many studies of ATR available in the

1. INTRODUCTION Because of insecurity of the crude oil supply, there is strong interest in developing technologies for production of alternative fuels, such as bioethanol, biodiesel, and various synfuels. The Fischer−Tropsch (FT)-based indirect liquefaction process is one of the most mature technologies for producing synfuels.1 In this process, the syngas produced from coal, biomass, and/or natural gas is converted to syncrude, which can be upgraded into on-spec transportation fuel.2 While the FT-based coal-toliquids (CTL) and gas-to-liquids (GTL) technologies are already commercialized,3,4 no FT plant has been built in the United States mainly because of the high greenhouse gas (GHG) emission and the uncertainty of the economic feasibility.5,6 To reduce the carbon footprint of the FT process, one possible option is to consider CO2 capture and storage (CCS). Further reduction in the carbon footprint is possible by adding a small amount of biomass into the system, resulting in the so-called coal-biomass to liquids (CBTL) process because of the carbon neutrality of biomass.7−9 CCS can significantly reduce the carbon footprint at the cost of higher capital and operating costs, which can significantly affect the economic feasibility of the CBTL technology. In the existing literature, a number of studies on modeling and optimization of physical and chemical absorption processes for preand post-combustion CCS has been conducted for power plant applications.10−13 Other application areas and CCS technologies, such as membrane, pressure swing adsorption (PSA), and vacuum swing adsorption, have also been investigated.12,14 CCS in the CBTL process differs from the pre- or post-combustion CCS in power plants as a result of the partial pressure of CO2, presence of light hydrocarbons, and sulfur level. Even though some outstanding studies have been conducted for the FT plant © 2015 American Chemical Society

Received: February 14, 2015 Revised: July 19, 2015 Published: July 28, 2015 5434

DOI: 10.1021/acs.energyfuels.5b00360 Energy Fuels 2015, 29, 5434−5451

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Energy & Fuels

amounts of fuel gas and reduce the overall gasoline yield. Previous studies indicate that, with tolerable fuel gas production, the isomerization technologies can only increase the octane number to about 80−90.15,27,29,30 Therefore, as the key design parameters, such as the H2/CO ratio in the FT plant, are changed, the product upgrading section needs to be appropriately designed to satisfy all product specifications. H2 required in the product upgrading section is considerable because the technologies, such as hydrotreating and hydrocracking, consume a large amount of H2 and operate under a H2-rich environment. However, typical routes for H2 production are associated with large GHG emissions.31 In the CBTL plant, H2 can be recovered from the unreacted syngas and purged gas from the upgrading section, while the remaining gases can be sent to the combined cycle plant. The H2 production and recovery units are expected to have a strong impact on the thermodynamic efficiency of the CBTL plant. Hence, a good estimation of H2 consumption is required for the efficiency analysis of the indirect CBTL plant. In the CBTL plant, the process fuel is supplied by the purged light gas and the unreacted syngas. The excess light gas and unreacted syngas can be sent to the combined cycle plant for electricity generation. An optimized heat recovery and steam generation (HRSG) unit can considerably increase the efficiency of the power plant. Many researchers have focused on the optimal design of the triple-pressure HRSG unit with reheat as part of an integrated gasification combined cycle (IGCC) plant.10,11,32−34 In typical IGCC plants, the high pressure (HP) section pressure is higher than 140 bar and the exhaust steam from the HP section is reheated for power plant application. For the once-through CBTL process, the HRSG unit can be designed similar to the IGCC plant as described in the literature.10,34,35 Unlike the IGCC plants, the fuel to the gas turbine in the CBTL plant is mainly the off-gas produced from the refinery and FT synthesis, which is usually not a large quantity. Typical differences in the design of the HRSG unit between the IGCC and CBTL plants have been reported in the literature.35,36 If a high amount of FT gas is recycled, very little high-temperature heat would be available for superheating the intermediate pressure (IP) steam produced in the FT reactor as a result of the large FT exotherm.34,35 However, a high recycle would help in improving the product yield. In this study, the HRSG unit is designed especially for cases when a large amount of FT gas is recycled for a higher fuel yield, resulting in deficiency of a high-temperature heat source. The authors of this paper have presented a study on a oncethrough CBTL plant with CCS technology, where the syncrude upgrading section was not considered.1 In this paper, our focus is on the modeling of the product upgrading section and overall plant performance analysis with new plant configuration. In particular, the contributions of this work can be summarized as follows: (1) A steady-state plant-wide model of an indirect CBTL plant with CCS and a combined cycle plant has been developed in Aspen Plus version 7.3.2 for production of onspec transportation fuel. (2) A simplified configuration of the product upgrading section has been developed for small-scale (10 000 bbl/day) application, which produces on-spec fuel with reasonable yield. (3) A novel integrated hydrotreating approach is considered for FT product upgrading. The model is developed on the basis of the atom balance, and the plant performance data are available in the literature. (4) The impacts of the steam/carbon ratio in the ATR inlet, H2/CO ratio in the FT inlet, extent of CCS, biomass/coal ratio, and biomass type

literature focus on the hydrogen production from natural gas or light hydrocarbons, which usually has a high steam/carbon ratio in the feed to obtain a high H2/CO ratio in the product. For the CBTL application, with moderate H2/CO ratio requirement in syngas, a low steam/carbon ratio can be used in the ATR unit to reduce the utility cost.22 In some studies conducted for the FT application, the ATR unit is modeled as an equilibrium reactor.4,15,17 Because of the key role that the ATR plays, a kinetic model is more appropriate for this unit, especially when its feed composition varies widely. The liquid product from the FT reactor is sent to the product upgrading section. In the conventional product upgrading section, syncrude is first separated into naphtha, diesel, and wax and then sent to two different hydrotreating units and a hydrocracking unit. Instead, integrated hydrotreating of the syncrude can increase the thermodynamic efficiency and reduce the footprint of the upgrading section, where the entire syncrude is first hydrotreated and then separated into different products for further upgrading. Although there is hardly any work in the existing literature on the use of an integrated hydrotreater for upgrading the FT syncrude, it should be noted that the integrated hydrotreating has been considered for upgrading of hydrocracked residuum petroleum crude oil,23 whole crude oil,24 and syncrude from coal direct liquefaction.25 It is, therefore, reasonable to consider that integrated hydrotreating can also be applied to upgrading of the FT syncrude because the type of components, such as paraffin, olefin, and oxygenate, carbon number, and boiling point range of FT syncrude and the main desired reactions, such as hydrodeoxygenation, hydrodemetallization, and hydrogenation of alkenes, are similar to those in the applications cited before.23−25 In the literature, some rigorous models have been developed for optimization and scaling up of the integrated hydrotreater based on the hydrodynamics, kinetics, and heat and mass balance.24 Other studies provide simple correlation for estimating the performance of the conventional separated hydrotreating unit.26 This work provides the perspective of a simplified yield model of the integrated hydrotreating unit in Excel. In the FT plant, the hydrotreated diesel can automatically satisfy most of the property specifications for commercial diesel. However, the straight-run FT naphtha mainly contains n-paraffin, resulting in a very low octane number, and needs to be further upgraded. The FT naphtha upgrading technology has been well-described in the Bechtel reports18,19 and has been considered in most of the recent studies on the FT plant.15,27 In these designs, the isomerization unit increases the research octane number (RON) of the light naphtha to about 82−85, while the catalytic reforming unit increases the RON of the heavy naphtha to about 95−100.19 Typical selection of technologies in commercial plants can also be found in the literature.2,28 However, as the gasoline and diesel specifications continue to change, especially with respect to their environmental impacts, suitable technologies should be selected. For example, the designs considered in the Bechtel reports18,19 can lead to violation of the aromatic content in the gasoline pool27 mainly as a result of a large quantity of high aromaticcontaining gasoline from the catalytic reforming unit. One of the alternative approaches is to apply the heavy naphtha isomerization technology that can increase the octane number of the straight-run heavy naphtha without producing aromatics. However, because the heavy naphtha is not only active for the isomerization reactions but also for the cracking reactions, the heavy naphtha isomerization technology will produce high 5435


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Figure 1. BFD of the indirect CBTL plant.

Figure 2. BFD of the novel product upgrading section. through the ATR, while the remaining portion is sent to a PSA unit to satisfy the H2 requirement for the product upgrading section. The liquid products (stream 6) are sent to the product upgrading section to produce on-spec gasoline and diesel (stream 10). Part of the off-gas (stream 11) from the entire process is used as utilities, while the remaining portion is used in a combined cycle plant. The combined cycle plant uses a gas turbine integrated with the HRSG unit that operates under three different pressure levels. In the product upgrading section, shown in Figure 2, a novel integrated hydrotreater is proposed. This integrated hydrotreater is expected to have a higher process thermodynamic efficiency and more compact design than the traditional approach. The light gases (LG) and H2-rich stream from the product upgrading section are sent to the PSA unit to produce pure H2 for hydroprocessing. A hydrocracking unit is used to produce naphtha and diesel from wax. A combination of the isomerization unit and catalytic reforming unit is considered for satisfying the current specifications for gasoline.2,19 2.1. Steady-State Modeling and Simulation Approach. In this section, the steady-state modeling and simulation of the indirect CBTL plant with a low biomass/coal ratio are discussed. The modeling approaches for the syngas production section, Selexol unit, Claus unit, FT synthesis unit, post-FT CO2 capture unit, and CO2 compression unit are summarized in Table 1 with detailed description available in the previous study of Jiang and Bhattacharyya.1 This paper mainly focuses on the ATR unit, the combined cycle plant, and the product upgrading section, which are required for producing on-spec

in the feedstock on process efficiency, CO2 emission, and product selectivity are evaluated.

2. PROCESS DESIGN AND MODELING The block flow diagram (BFD) of the indirect CBTL plant is shown in Figure 1. In the indirect CBTL plant, the raw syngas (stream 3) is produced in a gasifier co-fed with coal and biomass slurry (stream 2) and oxygen (stream 1) produced in an air separation unit (ASU).37 After scrubbing, the syngas is split between a single-stage sour watergas shift (WGS) reactor and a carbonyl sulfide (COS) hydrolysis unit, so that a desired H2/CO ratio in the clean syngas (stream 5) is achieved at the inlet of the FT reactor.1 The dual-stage Selexol technology is selected for the acid gas removal (AGR) unit to remove H2S and CO2 selectively, from where H2S is sent to the Claus unit and CO2 is sent to the CO2 compression section.10,38 Here, the Selexol unit is selected over the Rectisol unit mainly because of its relatively low capital and operating costs for the expected range of operating CO2 partial pressures.39,40 In the low-temperature Fischer−Tropsch (LTFT) reactor, the clean syngas (stream 5) from the AGR unit and the recycled gas (stream 8) from the ATR unit are converted to syncrude. The vapor-phase product (stream 7) is sent to the post-FT CO2 removal unit, where MDEA/PZ is used as the solvent.1 Removed CO2 is sent to the CO2 compression unit for pressurization and sequestration. A significant portion of the clean light gas (stream 9) from the post-FT CO2 removal unit is sent back to the FT reactor 5436


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Energy & Fuels Table 1. Operating Condition and Simulation Approach of Syncrude Production blocks

simulating approach

reference

fluidized-bed gasifier at 850 °C and 2380 kPa

WGS

sour WGS reactor with an inlet temperature of 250 °C

Selexol unit

H2S absorbers at 2 °C and about 2048 kPa and CO2 absorber at 2 °C and 2013 kPa slurry bed reactor at 257 °C and 2000 kPa yield model with Anderson−Schulz−Flory theory absorber at 1965 kPa and 38 °C and stripper at 172 kPa with RadFrac model with rate-based model for both absorber reboiler at 116 °C and stripper

Fischer−Tropsch post-FT CO2 removal


operating condition

gasification

Table 2. Proximate and Ultimate Analyses of Coal and Biomass bagasse41

Proximate Analysis (Dry Basis) moisture 11.12 10.6 fixed carbon 49.72 14.8 volatile material 39.37 82.1 ash 10.91 3.1 Ultimate Analysis (Dry Basis) ash 10.91 3.10 carbon 71.72 47.90 hydrogen 5.06 6.20 nitrogen 1.74 0.60 sulfur 2.82 0.01 oxygen 7.75 42.19 HHV (dry, kJ/kg) 30505 19575

1, 10, and 41 1, 10, 42, and 43 1 and 10 1 and 44 1

and oxygen/carbon ratios of the two streams are maintained to be the same as in the feed. Availability of information on reforming kinetics of C2+ hydrocarbons is scarce in the literature. However, several studies indicate that reforming of C2+ hydrocarbons are faster than methane reforming and results in methane formation.45−47 Hence, it is assumed that chemical equilibrium is reached for C2+ hydrocarbon, and therefore, these reactions are modeled using the RGibbs block. The product of the RGibbs block is mixed with the C1 stream and sent to a PFR, where the methane reforming reaction is considered. The key reactions of methane autothermal reforming on Ni/Al2O3 catalysts are shown in Table 3, with the kinetic parameters obtained from the

transportation fuels. Models of individual sections are developed on the basis of the experimental or operational data, whenever available, in the literature. If yield models are developed for a unit/section in Excel, then User2 Blocks in Aspen Plus are used to integrate the yield models, with other models developed in Aspen Plus. A stage-by-stage calculation of the steam turbine is performed in MATLAB to obtain a reasonably accurate estimate of the power output from the steam turbine under different operating conditions. The proximate and ultimate analyses of Illinois No. 6 coal, bagasse, and hardwood used in this study can be found in Table 2.1,10,41

coal10

equilibrium model for coal gasification and yield model for biomass gasification adiabatic PFR model with kinetics available in the literature RadFrac model for absorbers with equilibrium stages

Table 3. Reactions Considered in the ATR Kinetic Model21

wood chips41

number

9.58 16.55 82.51 0.94 0.94 48.51 6.17 0.12 0.04 44.22 19376

name

reaction

reaction heat

1

oxidation

CH4 + 2O2 → CO2 + 2H 2O

exothermic

2

CH4 + H 2O ↔ CO + 3H 2

endothermic

3

steam reforming dry reforming

CH4 + CO2 ↔ 2CO + 2H 2

endothermic

4

WGS

CO + H 2O ↔ CO2 + H 2

slight exothermic

literature.21 A high steam/carbon ratio is usually used to increase the H2 yield. If a moderate H2/CO ratio is required in the syngas, a low steam/carbon ratio can be used in the ATR unit to reduce the utility cost.22 The steam/carbon ratio from 0.5 to 3.0 is studied in this study. The oxygen flow rate is manipulated to achieve a reactor outlet temperature of 982 °C. The Soave−Redlich−Kwong (SRK) equation of state (EOS) is used to calculate the thermodynamic properties. Figure 4 shows that the simulation results agree well with the data available in the literature for the ATR unit as part of a CTL plant for different feed compositions and operating conditions.16,19,48 It should be noted that, in the CTL plant, the recycle gas to the ATR unit contains not only light hydrocarbons but also some unconverted syngas that strongly impacts the product distribution because of the WGS reaction. The data considered for model validation cover the range of feed compositions and operating conditions listed in Table 4. Appendix A provides detailed stream information for various cases that have been considered for model validation. 2.3. Hydrotreating and Hydrocarbon Recovery. In the product upgrading section, an integrated hydrotreating approach is proposed, as shown in Figure 5, for increasing the thermodynamic

2.2. ATR. The ATR unit uses a combination of exothermic partial oxidation and endothermic steam reforming reactions while operating under thermally neutral conditions to achieve optimum efficiency with less complicated facilities and less or no external energy in comparison to the steam reforming units. The process can practically approach adiabatic conditions if appropriately designed. Figure 3 gives the configuration of the ATR unit, where LG from the post-FT CO2 captured unit is first preheated by the hot ATR product, before sending to the ATR. For modeling purposes, the ATR reactor is simulated as a combination of a RGibbs reactor and a plug flow reactor (PFR). The ATR reactor feed is separated in a dummy component separator, where C1 and C2+ hydrocarbons are separated, and the steam/carbon

Figure 3. Configuration of the ATR unit. 5437


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by the hot treated syncrude and then heated by a furnace to reach the required temperature before being sent to the reactor. After being cooled, the treated syncrude is sent to a high-pressure flash (HPF) drum, followed by a low-pressure flash (LPF) drum to recover H2 and LG. Then, it is sent to the main distillation column through a series of heat exchangers. In this study, the correlations given by Bechtel18,19 are applied for the material and energy balance estimation of the conventional hydrotreating units for naphtha and diesel, while a simple yield model is developed in Excel for the integrated hydrotreater unit for obtaining reasonable estimates of H2 and utility consumption. To simplify the calculation of H2 requirement in the novel integrated hydrotreating unit, a number of assumptions has been made. The operating condition is considered to be similar to that in the conversional diesel hydrotreater (58 bar and 297 °C), which is much more severe than the operating conditions in the naphtha hydrotreaters. Hence, it is assumed that the naphtha cut becomes completely hydrotreated, and the amount of diesel cut that becomes hydrotreated depends upon the catalyst type and experimental bromine number of


Figure 4. ATR model validation.16,19,48 efficiency and making the plant footprint smaller, in comparison to the conventional separated hydrotreating approach shown in Figure 6. In the conventional separated hydrotreating approach, the crude oil is first separated into different streams in flash drums and distillation columns. Then, naphtha and diesel are sent to two different hydrotreating units, while wax is sent to a hydrocracking unit. In contrast, in the integrated hydrotreating unit, the raw syncrude is first preheated

Table 4. Range of Feed Composition and Operating Conditions for ATR Model Validation16,19,48 minimum maximum

steam/carbon (mol/mol)

oxygen/carbon (mol/mol)

syngas/hydrocarbons (mol/mol)

outlet temperature (°C)

3.76 1.23

0.509 0.157

3.125 13.15

971 982

Figure 5. Configuration of the novel integrated hydrotreating approach.

Figure 6. Configuration of the conventional separated hydrotreating approach. 5438


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to ensure that the final product pools satisfy the desired gasoline and diesel specs. The cut points of light naphtha are specified for satisfying the gasoline specs. The cut points of heavy naphtha and diesel are specified to satisfy the specs of the gasoline and diesel pools, respectively. The PetroFrac model is used to design and simulate the main distillation column, where BK10 EOS is used as the thermodynamic model because the distillation system contains species of wide boiling point range.52,53 A stabilizer is simulated via the RadFrac model using SRK EOS as the thermodynamic model because the system mainly contains lighter hydrocarbons. The column and heat exchanger specifications are provided in Appendix C. 2.4. Other Product Upgrading Units. The wax stream from the main distillation column is sent to the wax hydrocracking unit to produce shorter chain hydrocarbons that are then separated into light naphtha, heavy naphtha, and diesel. A simple yield model is developed by multivariable regression using the experimental data reported by UOP for their single-stage HC Unibon process.54 The HC Unibon technology is a fixed-bed catalytic process that uses a high-activity bifunctional catalyst and has been developed to maximize diesel production for full conversion application.54 The H2 reacted per barrel of wax (FH2) depends upon the gasoline/diesel ratio if the conversion is the same. Equation 1 gives an estimation of FH2 of the wax hydrocracking unit correlated to the weight percentage of the C7+ product (yC7+), where FH2 is in standard cubic feet per barrel (SCFB) of wax.54 Information on utility consumption is available in the literature54 and assumed to be proportional to the feed flow rate. It is noted that the wax hydrocracking model does not provide the isomer distribution of the naphtha cut required for modeling the naphtha upgrading units. Hence, a typical composition of naphtha cut from the literature is used in this study.55,56 The yield model developed on the basis of the data by UOP is consistent with the experimental data reported by Sasol shown in Table 7.57,58

hydrotreated diesel. Typically, the bromine number of the hydrotreated FT diesel is lesser than 6.0 g of Br/100 g when catalyzed by NiMo/Al2O3.49 Hence, in the yield model developed, we have considered 5 wt % of unsaturated diesel that corresponds to 6.0 g of Br/100 g. Because Fe-catalyzed FT syncrude contains only a small amount of oxygenates and no sulfur and nitrogen, the main reaction considered is hydrogenation of alkenes and hydrodeoxygenation. With the detailed component distribution in the reactor inlet,1 the H2 consumption can be estimated by atom balance with the following assumptions: (1) Reacted olefins are converted to the corresponding saturated paraffin compound. (2) Wax remains mainly unreacted in this integrated hydrotreater because wax hydrotreating needs much severe reaction conditions. (3) Yields of LG produced by the side hydrocracking reaction are assumed to be similar to the conventional hydrotreating units. (4) All oxygenates are hydrotreated and converted to water and corresponding paraffin compound. Most of the heat required for preheating the hydrotreater feed can be recovered by exchanging heat between the feed stream and hydrotreater outlet stream, while the remaining heat is supplied by the feed furnace. Because of the wide variation in the thermodynamic properties of isomers of C5−C8, a statistical model of the isomer distribution of paraffin in the LTFT product developed by Weller and Friedel50 is considered for more accurate energy calculation. The detailed isomer distribution is reported in Appendix B. Because of the limited information available on hydrotreating of the FT liquids, the yield model is validated by comparing the calculated product distribution and hydrogen consumed to those reported by Bechtel18 with the same feed composition. The composition of oxygenates in the feed was not specified in the Bechtel report. Hence, to generate the final product distribution, we have assumed that oxygenates in naphtha and diesel are represented by C4.78H11.14O1.1 and C9.08H18.94O1.1, respectively.44,51 Table 5 lists the results and shows

Table 5. Validation of the Model of the Integrated Hydrotreater18,19 wt % H2 consumption major products LG naphtha diesel

Bechtel

model

error (%)

1.10

1.07

−2.8

2.97 39.3 57.8

2.96 39.1 57.9

0.34 0.33 0.29

Table 7. Comparison of the Model Results to the Available Data57,58 for the FT Wax Hydrocracking Unit

product light naphtha heavy naphtha diesel wax

52−94 104−174 190−316 327−FBP

separated approach product

ASTM D86 cut point (°C)

naphtha diesel wax

50−174 190−316 327−FBP

Leckel57,58

error (%)

C1−C4 C5−C9 C10−C22

7.55 33.8 58.6

7.6 34 58

−0.65 −0.46 1.05

7+

(1)

For naphtha upgrading, the UOP Penex process is considered for light naphtha isomerization as a result of its low cost. A simplified yield model has been reported by Bechtel for this process.19 The selectivity of the isomer is about 98.3 wt %, and the makeup hydrogen rate is about 0.14 wt % of the light naphtha feed rate. Utility consumption is assumed to be proportional to the feed flow rate.18,19 The UOP CCR Platforming technology is selected to increase the octane number of FT heavy naphtha by converting it into aromatics. According to the experimental data provided by UOP, this technology for catalytic reforming is able to increase the research octane number (RON) of FT heavy naphtha to about 100.19,59 The Aspen Tech Reformer model under the Aspen One package is used for estimating the process yield and product properties. First, the target RON, flow rate, and composition of the feed are specified in the Aspen Tech Reformer model. Then, the simulation is run, and the results are compared to the data provided in the report by Bechtel,19 as shown in Table 8. It shows that the results obtained from the Aspen Tech Reformer are satisfactory. 2.5. H2 Network and PSA Unit. In the product upgrading section, H2 produced in the catalytic reforming unit and the purged gases from the hydroprocessing units, shown in Figure 7, are sent to the H2 recovery unit, a polybed PSA process, to produce a portion of pure H2 for hydroprocessing. The remaining H2 requirement can be satisfied by sending a portion of the FT vapor (stream 7 in Figure 1) to the PSA unit to recover H2 from the unconverted syngas. In this study, a

Table 6. Product Specification of the Hydrocarbon Recovery System ASTM D86 cut point (°C)

model

FH2 = 2215 − 15.427yC

that the errors in yields of major products are within 5%. It should be noted that the syncrude composition reported by Bechtel18,19 is similar to the base case of this study. It is assumed that the hydrocarbon distribution does not change significantly in the range of operating conditions considered in the sensitivity studies conducted in this work. In both hydrotreating approaches, the raw or hydrotreated syncrude is cooled to about 40 °C and sent to the HPF (38 bar) to recover the H2-rich gas. The bottom stream of HPF is sent to the LPF drum (8 bar) from where the LG are recovered and sent to the fuel gas header. Then, a complex distillation column is used to separate the syncrude into products with different boiling point ranges, as shown in Table 6. A stabilizer is used to separate light gases from the naphtha stream. The ASTM D86 cut points of the hydrocarbons are specified

integrated approach

wt %

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where the terms represent volumetric average values of properties as follows: R = RON, M = MON, J = RON − MON, RJ = R × J, MJ = M × J, O = olefins volume percent, and A = aromatics volume percent. It has been reported that the FT diesel has a high cetane number and can satisfy the specification of diesel without any further upgrading except hydrotreating. The diesel pool in our design is a blend of diesel cuts from the hydrocarbon recovery section and hydrocracking section. The properties of the diesel mixture are estimated by volumetric average, and the properties of individual diesel blend are available in the Aspen Plus report and the literature.19,54 The cetane index (CI) is the substitute measure of the cetane number and can be estimated by the ASTM D976 method shown in eq 5

Table 8. Comparison between Aspen Model and Bechtel Data19 for Catalytic Reforming


H2 (wt %) C1−C5 (wt %) reformate (wt %) specific gravity RON benzene (wt %) aromatic (wt %)

Aspen

Bechtel19

error (%)

4.14 8.68 87.00 0.80 95 0.66 66.14

3.44 10.67 85.89 0.77 95 0.70 65.90

1.39 3.49 0 −5.71 0.36

CI = 454.74 − 1641.416D + 774.74D2 − 0.554B

component separator block is used for simulating the PSA unit with the H2 purity and recovery efficiency of the PSA unit assumed to be 99.9 and 50.7%, similar to the Bechtel design that uses a standard 10 bed system.19 In should be noted that the PSA unit is an unsteadystate process, where a number of adsorber vessels are cycled in a desired sequence, changing their pressure typically between 2620 and 690 kPa for adsorption and desorption, respectively.19 In this study, it is assumed that the number of beds in the PSA unit and the sequence have been appropriately designed so that H2 is available continuously at the desired rate. A model of the H2 network is developed to estimate the flow rate of the makeup H2 stream and the amount of FT vapor that can be recycled back to the FT reactor. The high H2 partial pressure in the hydroprocessing reactors is usually maintained by recycling unreacted H2. The product from the hydroprocessing reactor is cooled and sent to a H2 recovery flash drum. The majority of the vapor stream is sent back to the reactor, and the rest is purged and sent to the PSA H2 recovery unit to avoid light gas accumulation in the reactor. The model of the product yield and H2 reacted in the hydrotreating and hydrocracking units are developed in Excel, as reported in sections 2.3 and 2.4. The purge rate is manipulated to maintain the H2 partial pressure required by the corresponding hydroprocessing unit, while the flow rate of the makeup H2 is manipulated to achieve the required H2/oil ratio in the reactor. BK10 EOS is used to estimate the vapor−liquid equilibrium in the flash drum. 2.6. Blending Rules for Fuel Property Estimation. The final gasoline product is a blend of the isomers produced from the isomerization unit and the reformates produced from the catalytic reforming unit. The nonlinear blending rules used to estimate the Reid vapor pressure (RVP), RON, and motor octane number (MON) of the gasoline blend are shown in eqs 2−4, provided by the Ethyl Corporation,60 which is one of the widely used rules for petroleum products. Other properties of the blends are estimated by the linear blending model or Aspen Plus petroleum characterization. 1/1.25 (RVP)mix = [∑ vi(RVP)1.25 i ]

(2)

R = R̅ + 0.03324[RJ − R̅ × J ̅ ] + 0.00085[( O2) − (O̅ )2 ]

(3)

+ 97.803(log B)2

where D is the density at 15 °C in g/mL and B is mid-boiling temperature (D86) in °C. 2.7. Combined Cycle Power Plant. The fuel gas from the PSA unit and the hydrocarbon upgrading section provides fuel required in the FT synthesis and the entire hydrocarbon upgrading units. The remaining portion is sent to the gas turbine (GT) for electricity production, as shown in Figure 8. The appropriate GT frame for this CBTL plant is selected to be GEE MS7001EA, which has a designed power rating of 85 MW and a simple cycle efficiency of 32.7% (for natural gas firing). This frame can be used for H2-rich (H2 > 50%) gas. Chiesa et al. have evaluated the possibility of burning H2-rich gas in a large heavy-duty gas turbine designed for natural gas.61 If H2-rich gas is fed into GT, steam or nitrogen dilution is required to control NOx emissions.61 In this study, nitrogen dilution is selected for NOx emission control, taking advantage of the existing ASU. It is assumed that the GT has been appropriately modified, so that it can be operated by firing H2-rich gas, where the pressure ratio and first rotor inlet temperature are similar to the case for firing natural gas, while the turbine outlet temperature is about 10−15 °C lower.61 The modeling approach reported by Bhattacharyya et al.10 is used to estimate the GT performance. The operating conditions for MS7001EA are obtained from the literature for firing natural gas.62 For firing H2-rich (about 60% H2 concentration) gas in the GT for CBTL application, the N2/fuel ratio is manipulated to reduce the stoichiometric flame temperature to 2027 °C to control the NOx emission.61 The combustion air is compressed to 12.7 atm in an axial flow compressor. The GT combustor temperature is maintained at 1150 °C with a specified heat loss of 1.5% of the fuel gas LHV by manipulating the combustion air flow. The GT firing temperature is maintained at 1125 °C by manipulating the air flow rate to the combustor outlet gas before the first expansion stage. The exhaust temperature is maintained at 528 °C by manipulating the isentropic efficiency.10,61,62 A model of the triple-pressure HRSG with reheat is developed for the indirect CBTL process, with the configuration shown in Figure 9 and Table 9. The steam for power generation is mainly produced by recovering heat from the gas turbine exhaust flue gas, syngas cooler, heat recovery exchangers, and FT reactor cooling system. Part of the steam produced is sent to other units for operation. The pressure levels and steam turbine inlet conditions are specified on the basis of the studies conducted recently for FT application,18,35,36 while a 6% pressure drop

M = M̅ + 0.04285[MJ − M̅ × J ̅ ] + 0.00066[( O2) − (O̅ )2 ] ⎡ (A2 ) − (A)2 ⎤2 ⎥ − 0.00632⎢ ⎢⎣ ⎥⎦ 100

(5)

(4)

Figure 7. General configuration of the hydroprocessing unit. 5440


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Figure 8. Configuration of the combined cycle power plant (fuel side).

Figure 9. Configuration of the combined cycle power plant (steam side).

the base case can be found in Tables 11 and 12. The costs of raw materials, products, and utilities are listed in Table 13.65 Table 12 indicates that syngas production and CCS are the two major utility consumers in the indirect CBTL plant, which is consistent with the literature.1,3,4 It should be noted that the process fuel required in the CBTL plant is supplied by the fuel gas header, while the steams and electricity are supplied by the combined cycle plant. A number of studies have been conducted for analyzing the effects of key design parameters that are listed in Table 10. First, the effect of the steam/carbon ratio on the ATR unit is

Table 9. Configuration and Operating Conditions of the HRSG Section and Steam Header steams HP steam to ST IP steam to ST LP steam to ST HP steam to header IP steam to header LP steam to header

pressure (kPa)

temperature (°C)

7419

373.9

SHR, HRSG

2172

346.1

365

141.7

SHR, Claus, FT ST IP section (through reheater) SHR, HRSG ST LP section

from

4137

ST

931

ST

365

SHR, HRSG

to ST HP section

upgrading unit, ATR SWS, Selexol unit MDEA/PZ unit, upgrading unit

Table 10. Key Design Parameters (Base Case)

is considered for the reheat section.63 The minimum temperature of flue gas to the stack is set at 120 °C.10 For the performance of a threepressure-level steam turbine with multiple steam addition and extraction points, a simple stage-by-stage calculation is performed in MATLAB based on the algorithm presented by Lozza,64 shown in Appendix D.

3. RESULTS AND DISCUSSION Table 10 lists the base case values of the key design parameters investigated in this study. The material and utility summaries of 5441

design parameter

value

biomass type plant capacity (bbl/day) biomass/coal (wt/wt, dry) hydrotreating approach steam/carbon ratio in the ATR inlet (mol/mol) H2/CO ratio in FT inlet stream (mol/mol) CO2 captured in the Selexol unit (%) CO2 captured in the MDEA/PZ unit (%) CO2 stream to compression section (%)

bagasse 10000 8:92 integrated 0.63 2 90 98 100


Article

Energy & Fuels Table 11. Summary of Material Balance (Base Case)


stream temperature (°C) pressure (kPa) flow rate (kg/h) coal biomass H2O CO2 O2 N2 CH4 CO COS H2 H2S C2−C4 C5−C10 C11−C20 wax oxygenates gasoline diesel

1

2

3

4

5

6

7

8

9

32 2380

16 2380

850 2380

21 2289

6 1999

49 1965

49 1965

261 1999

38 1965

10

11

12

25 101

38 138

81 15270

57005 124409

355 204043

409 20451

109 1398

267 54057

5281 2202

195 397

4584 1865 183130 208 12570 4512 177

4583 1863 132431 109 16203 4547 177

1676 130823

30 39

3743 11897

721 22168

2751 8732

1130 3187

309 4745

16154

7

5509

5594

4043

1474

64

910 13933 12036 25632 3279

6512 3251 7

14 1

3479 1531 4

6853 1725

141

774

553

553

173747 15021

141184 2139

30 229545

18391 30579

Table 12. Summary of Major Utility Consumption (Base Case)a sections syngas production syncrude production CO2 capture and storage product upgrading fuel gas header others gas turbineb HRSG a

power (MW)

74 bar steam (kg/h)

59.48 0.88 42.24 1.16

(119895)

12.94 (53.95) (72.44)

42 bar steam (kg/h)

22 bar steam (kg/h) (25667) (186727)

6784

9.3 bar steam (kg/h)

3.7 bar steam (kg/h)

fuel (GJ/h)

58340

(119680)

8.31 19.97

79538

50841 1190

213

87.31 (699.51)

(4266) 583.92 119895

(6997)

216660

(137878)

67649

Parentheses mean utility generation. bWith air and nitrogen compressors.

Table 13. Cost of Raw Materials, Products, and Utilities65 cost coal ($/dry ton) biomass ($/dry ton) gasoline ($/gallon) diesel ($/gallon)

46 80 3.024 2.902

Table 14. Effect of the Steam/Carbon Ratio on the Performance of the ATR Unit

cost LP steam ($/GJ) MP steam ($/GJ) electricity ($/GJ) cooling water ($/GJ)

steam/carbon (mol/mol)

13.28 14.19 16.8 0.354

0.5 Performance 759 486 1.6

H2 produced (kmol/h) CO produced (kmol/h) H2 produced/CO produced (mol/mol) H2/CO in the ATR outlet (mol/mol) 3.4 Utilities O2 consumed (kg/h) 7335 steam consumed (kg/h) 5460 CO2 captured by the Selexol unit (%) 79.3

discussed in section 3.1. Then, the advantage of the novel integrated hydrotreating approach is discussed in section 3.2, where it is compared to the conventional separated hydrotreating approach. Sections 3.3−3.5 show evaluation of the impact of the H2/CO ratio in the FT inlet stream, biomass/coal ratio, and extent of CCS on the thermal efficiency and fuel yield of the novel indirect CBTL plant. Section 3.6 discusses the effect of different biomass types on the plant performance. The properties of the upgraded FT fuels are discussed in section 3.7. Finally, in section 3.8, the results of the base case are compared to other related studies available in the literature. 3.1. Effect of the Steam/Carbon Ratio at the ATR Inlet. The effect of the steam/carbon ratio in the ATR unit is evaluated by fixing the H2/CO ratio in the FT inlet to 2, the same as the base case condition. As seen in Table 14, the results indicate that the H2/CO ratio in the ATR outlet and the utility consumptions of the ATR unit increase with the increase in the

1.0

2.0

3.0

791 454 1.7

822 394 2.1

857 359 2.4

3.6

4.1

4.8

7947 10729 78.8

9075 21368 77.5

10598 33440 75.1

steam/carbon ratio. Because the H2 demand should be satisfied, a higher H2/CO ratio in the ATR outlet would require a lower extent of reactions in the WGS reactor, and therefore, the percent of CO2 captured by physical solvent in the Selexol unit decreases with the increasing steam/carbon ratio. As a results, the penalty of CCS increases as the steam/carbon ratio increases. Furthermore, the FT reactor is usually operated with an inlet H2/CO ratio less than 2.1. Therefore, a low steam/carbon ratio is recommended at the ATR inlet for FT application.22 To prevent coking, the steam/carbon ratio is set to be 0.63 for the base case.22 5442


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Energy & Fuels Table 15. Major Utility Consumptions of the Two Hydrotreating Approaches


integrated hydrotreating

separated hydrotreating

unit

description

GJ/h

unit

description

GJ/h

F1 F2 R1 STM COM total

hydrotreating preheater furnace of the main column reboiler of the stabilizer stripping steam hydrogen compressor

4.26 23.70 3.67 2.51 3.87 38.01

F3 F5 + R3 + R4 F4 + R2 STM COM total

furnace of main column naphtha hydrotreating diesel hydrotreating stripping steam hydrogen compressor

24.75 19.83 3.44 2.18 3.87 54.07

3.3. Effect of the H2/CO Ratio in the FT Inlet Stream. In the CBTL plant, the H2/CO ratio in the syngas can be adjusted in the WGS reactor before sending to the Selexol unit, as shown in Figure 1. Studies indicate that the H2/CO ratio in the FT inlet stream not only affects the penalty of CCS but also the fuel product yield and distribution.1,22 Hence, in this study, a sensitivy study is conducted by changing the H2/CO ratio from 1 to 2.25 and keeping the raw material flow rate and other design parameters the same as the base case. A previous study from our group indicates that, with an increase in the H2/CO ratio in the FT inlet stream, the penalty of CCS keeps reducing in a once-through indirect CBTL plant without product upgrading.1 A similar trend, shown in Figure 10, can be found in the CBTL plant producing on-spec

3.2. Advantage of the Integrated Hydrotreating Unit. By comparison of the configuration of the integrated hydrotreating approach (Figure 5) to the conventional separated hydrotreating approach (Figure 6), it clearly shows that the integrated hydrotreating approach can reduce the plant footprint and make the plant more compact. In the integrated hydrotreating approach, the entire hydrotreated syncrude is sent to the main distillation column to separate the product to light naphtha, heavy naphtha, diesel, and wax, which is similar to the main distillation column in the separated hydrotreating approach design. The only difference in the main distillation columns is that the heavy naphtha side stripper is not considered in the separated approach, because the entire naphtha cut is sent to the naphtha hydrotreating unit together and then separated in another distillation column. One advantage of the integrated hydrotreating approach is to eliminate some distillation columns from the conventional approach, which are required to remove light gases from the products and separate light naphtha from heavy naphtha, thus consuming a considerable amount of plant fuel because of the large reboiler duty (R2, R3, and R4 in Figure 6). The disadvantage of the integrated hydrotreating approach is that the wax, which does not necessarily need to be hydrotreated, is also sent to the hydrotreating unit, resulting in the increase in the preheat furnace duty and the hydrotreater reactor size. However, the temperature increase in the furnace is very low, just about 20 °C, and the wax mostly remains in the liquid phase. Therefore, the increase in the heat duty and the volumetric flow rate to the reactor is not very large. For the separated hydrotreating approach, the utility consumptions in and capital investment for naphtha and diesel hydrotreating units are given by Bechtel,18,19 and then the capital investment is escalated with the chemical engineering plant cost index (CEPCI).65 For the remaining units, the utility consumptions and capital investment are estimated using Aspen Plus and Aspen process economic analyzer (APEA), respectively. Detailed specifications of the APEA model for capital investment estimation can be found in Appendix E. Tables 15 and 16 show the comparison of heat consumption and capital investment between the two hydrotreating approaches. It is observed that the integrated hydrotreating approach can reduce the heat consumption and the capital investment by about 30%.

Figure 10. Effect of the H2/CO ratio on the penalty of CCS.

gasoline and diesel. For the Selexol unit, the solvent circulation rate reduces with the increasing H2/CO ratio because of the higher partial pressure of CO2, which can provide more driving force for the physical absorption process. For the MDEA/PZ unit, the solvent circulation rate decreases because the CO2 selectivity in the FT reactor decreases with the increasing H2/CO ratio.1 In addition, CO2 can be recovered from the Selexol unit at different pressure levels, usually higher than the pressure of CO2 released in the chemical absorption unit, which indicates that the penalty of the CO2 compression section can be reduced as a larger potion of CO2 is captured in the Selexol unit. In Figure 10, the total utility cost (CuCCS) of CCS is calculated by eq 6 u CCCS =

separated hydrotreating

section

MM$

section

MM$

integrated hydrotreating loop hydrocarbon recovery

8.17 3.43

hydrocarbon recovery naphtha hydrotreating diesel hydrotreating total

2.56 4.70 9.45 16.71

total

11.60

(6)

where Fu is the utility consumption of the uth type of utility in GJ/h and Cu is the unit cost of the uth type of utility listed in Table 13 in $/GJ. Because the H2/CO ratio in the FT inlet has a strong impact on the hydrocarbon selectivity in the FT reactor,1,22,66 the product distribution and fuel yield of the indirect CBTL plant highly depend upon the H2/CO ratio in the FT inlet. Figure 11 indicates that the gasoline/diesel ratio keeps increasing with the increasing H2/CO ratio, because the FT reaction produces lighter hydrocarbon with a higher H2/CO ratio in the inlet.1,22,66 Figures 12−14 also show that the fuel yield, overall plant

Table 16. Capital Investment of the Two Hydrotreating Approaches integrated hydrotreating

∑ CuFu

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in the H2/CO ratio. However, it is expected that, with a very high H2/CO ratio, the fuel yield will decrease as more amount of carbon in the feedstock becomes converted to CO2 and removed by the Selexol unit before being sent to the FT unit for fuel production. In this study, a H2/CO ratio larger than 2.25 is not considered because of the absence of the experimental data of the FT reactor operated at a very high H2/CO ratio. It should be noted that, in Figure 13, the thermal efficiency is defined as the energy output (fuels and electricity) to input (coal and biomass) ratio on a higher heating value (HHV) basis, while the carbon efficiency is defined as the percent of carbon in the feedstock converted into fuels. The profit function (PF) in Figure 14 is defined as eq 7

Figure 11. Effect of the H2/CO ratio on the product distribution.

PF =

∑ CpFp − ∑ Cf Ff


prod

−

feed

∑

CuFu (7)

utility

where Ci is the unit cost of the ith item listed in Table 13 and Fi is the material or energy flow rate of the ith item. 3.4. Effect of the Biomass/Coal Ratio. As mentioned before, the carbon footprint of the indirect CBTL plant can be decreased by increasing the biomass content in the feedstock. In this study, a sensitivity analysis is conducted for biomass/ coal weight ratios of 8:92, 15:85, and 20:80 (dry) to estimate the effect of feedstock composition on the plant performance, especially the product yield and plant efficiency. A relatively low biomass content is considered in this study mainly considering sustainability of the plant.67 For the alternative cases, the total amount of dry feed and other design parameters are fixed to be the same as the base case. The simulation results are presented in Table 17. It shows that as the biomass content keeps

Figure 12. Effect of the H2/CO ratio on the fuel yield.

Table 17. Effect of the Biomass/Coal Ratio in the Indirect CBTL Plant biomass/coal

Figure 13. Effect of the H2/CO ratio on the plant efficiency.

feedstock coal (dry) biomass (dry) product gasoline diesel total FT liquid net electricity thermal efficiency FT liquid net electricity total

Figure 14. Effect of the H2/CO ratio on the plant profit and CO2 emission.

dry weight

8:92

15:85

20:80

ton/h ton/h

153.4 13.3

141.8 24.9

133.5 33.2

bbl/h bbl/h bbl/h MW HHV % % %

4050 5950 10000 9.71

3848 5656 9504 7.25

3721 5465 9186 5.06

45.9 0.7 46.6

44.7 0.5 45.2

44.0 0.4 44.4

increasing, the overall fuel production and the plant thermal efficiency decrease, mainly because of the relatively high oxygen content in the biomass. Our previous study has shown that an increase in the biomass/coal ratio results in an increase in the H2/CO ratio in the raw syngas (stream 3 in Figure1).1 As a consequence, the extent of the WGS reaction and the heat recovery decreases if the H2/CO ratio keeps increasing in the raw syngas, while the H2/CO ratio at the WGS outlet (stream 4 in Figure1) remains constant.1 3.5. Effect of the Extent of CCS. For all case studies, the acid gases, CO2 and H2S, are removed from the syngas and FT vapor product by the same extent as the base case. CO2 removal is required to improve the kinetics and economics of the downstream synthesis and upgrading process.4 H2S removal is required to avoid catalyst poisoning.4 In this study, the extent

efficiency, and plant profit increase with the increasing H2/CO ratio but with decreasing slope. That is because, with a higher H2/CO ratio, the H2 conversion decreases in the FT reactor. As a result, the recycled light gases from the post-FT CO2 capture unit has a higher H2 percentage, and a smaller portion is needed to be sent to the H2 plant to produce H2 required for the product upgrading section. A larger portion can be sent back to the FT unit through the ATR to produce more syncrude. In the meanwhile, fewer amounts of light gases are purged from the H2 unit, which is then sent to the combined cycle plant for power production, where no CO2 capture facilities are considered for the flue gas. Hence, with the same extent of CO2 removal in the Selexol unit and the MDEA/PZ unit, the electricity production and overall CO2 emission in plant also decrease with the increase 5444


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3.7. Properties of the Gasoline and Diesel Product. As discussed in sections 2.3 and 2.4, with the simplified refinery design shown in Figure 2, the required specifications of gasoline and diesel can be achieved by adjusting the D86 95 vol % cut point of the light and heavy naphtha stream of the main distillation column. In the base case, the D86 95 vol % cut point of the light and heavy naphtha stream is set to be 94 and 174 °C, respectively. Table 20 shows the values of the final gasoline

Table 18. Effect of the Extent of CCS (i.e., Amount of CO2 That Is Compressed) case


CO2 stream to compression section (%) net electric power (MW) CO2 emission from plant (g of CO2/MJ) thermal efficiency (%, HHV)

high

intermediate

low

no CCS

100

75

50

0

9.71 12.14

14.14 25.95

19.03 39.76

28.65 67.37

46.6

47.0

47.3

48.0

Table 20. Estimated Properties of the Gasoline Pool and Specifications of U.S. Gasoline68−71

of CCS is manipulated by changing the fraction of the CO2 streams sent to the CO2 compression section and not the extent of CO2 captured. The results are shown in Table 18. In the base case, all CO2 streams removed from the system are sent to the CO2 compression section. If CCS technology is not considered (i.e., no compression), all CO2 streams are vented to the atmosphere. With an increase in the extent of CCS, CO2 streams that are sent to the CO2 compression section depend upon their pressure levels. The underlying philosophy is that if one would like to vent a portion of CO2, then this portion will be bled off from the lowest pressure CO2 stream available. For example, in the low CCS case, CO2 streams released from the HP and MP flash drums in the Selexol unit and a portion of the CO2 stream released from the LP flash drum are sent to the compression section, while CO2 streams released from the MDEA/PZ unit and the remaining portion of the CO2 stream from the LP flash drum in the Selexol unit are vented. It is noticed that, in the case without CCS, the CO2 emission from the plant is about 25 g of CO2/MJ less than the data reported in the work of Edwards et al.,6 which is reasonable because the WGS technology is not considered in that work.6 Figure 14 indicates that the CO2 emission from the plant can be reduced by about 22 g of CO2/MJ when the H2/CO ratio in the FT inlet is increased from 1 to 2. Table 18 also indicates that the CO2 emission from the plant can be reduced from 67.37 to 12.14 g of CO2/MJ at the cost of 1.4% decrease in the overall thermal efficiency. The penalty as a result of capital investment for CCS is not discussed in this paper but will be considered in a separate study focusing on techno-economic analysis. 3.6. Impact of Biomass Type. The impact of biomass type on the performance of the CBTL process is shown in Table 19.

U.S.A. specification68−71 fuel property D86 50 vol % (°C) D86 90 vol % (°C) RVP (kPa) aromatics (vol %) benzene (vol %) sulfur (ppm, wt) road octane number ([R + M]/2)

feedstock coal (dry) biomass (dry) product gasoline diesel net electricity analysis C captured by FTL thermal efficiency

wood chips

bagasse

ton/h ton/h

153.8 13.5

153.4 13.3

bbl/day bbl/day MW

4050 5950 2.50

4050 5950 9.71

% % (HHV)

36.3 45.9

36.4 46.6

minimum maximum

source ASTM D4814 ASTM D4814 ASTM D4814 CA RFGa 40 CFR 80 40 CFR 80

a

Flat limit of a small refinery from California reformulated gasoline (RFG), phase 3.

Table 21. Estimated Properties of the Diesel Pool and Specifications of No. 2 Diesel72 Sasol73 fuel property density at 15 °C (kg/m3) flash point (°C) aromatic (vol %) sulfur (ppm, wt) cetane number cetane index

U.S.A. specification72

product LTFT minimum maximum

source

Restrictions on Boiling Range 769 772 876

ASTM D975

60 60 52 Restrictions on Composition 0 0.7 35 0 70 >70 40 >70 40

ASTM D975 ASTM D975 ASTM D975 ASTM D975

blend properties and the selected U.S.A. standard of gasoline.68−71 Table 21 shows that the conceptual design developed in this study can produce on-specification diesel,72 and the estimated properties from our model are consistent with the industrial data.73 3.8. Model Validation and Comparison of the Novel Indirect CBTL Plant. Table 22 shows a comparison of the material and energy balances of the indirect CBTL plant with CCS (base case) to the data available in the literature for the indirect CTL plant.15,16,18 As shown in Table 22, the overall thermal efficiency and carbon efficiency of the base case analyzed in this project are similar to those of the previous studies. The efficiency obtained in this study is slightly higher than the data reported by other studies with a similar extent of CO2 capture mainly as a result of the difference in feedstock, CO2 capture technology, extent of CO2 capture, product upgrading technologies, and their operating conditions, as discussed in the previous sections.

Table 19. Alternative Biomass as Feedstock biomass type

product

Restrictions on Boiling Range 92.8 76.7 121 139.4 190 47.9 54 Restrictions on Composition 34.1 35 0.4 1 0 20 87.2 87

The results indicate that the thermal efficiency of wood chips is lower than bagasse as a result of the higher oxygen content and lower hydrogen/carbon ratio in wood chips, as shown in Table 2. The carbon efficiency remains similar because all of the other key design parameters remain the same and the biomass/coal ratio is small in the feedstock.

4. CONCLUSION In this study, the process model of a CBTL plant with a combined cycle plant and CCS has been developed in Aspen 5445


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Energy & Fuels Table 22. Material and Energy Balance of the Indirect CBTL Plant NETL16,a


Bechtel18 report year feedstock coal (dry) biomass (dry) product propane butanesc gasoline diesel total FT liquid net electricity analysis CO2 removal technology CCS C captured by FTL C captured by CCS thermal efficiencyd

Liu et al.15

Liu et al.15

base caseb

1993

2007

2011

2011

2014

ton/h ton/h

702.1 0

908.5 0

892.0 0

94.9 126.8

153.4 13.3

ton/h ton/h bbl/day bbl/day bbl/day MW

6.45 (11.98) 23943 24686 48629 −54.32

0 0 22173 27819 49992 124.25

0 0 N/A N/A 50000 295

0 0 N/A N/A 9845 53

0 0 4050 5950 10000 9.71

% % % (HHV)

Rectisol, MDEA no N/A 0 51.8

Selexol, MDEA yes 35.5 56.6 42.4

Rectisol yes 34.1 51.6 46.0

Rectisol yes 33.7 53.7 47.5

Selexol, MDEA/PZ yes 36.4 56.9 46.6

a

An additional refinery is required for producing on-specification gasoline. Efficiency is expected to be higher. bData generated in this study. cIn the refinery design by Bechtel, purchased n-butane is required for the upgrading section, such as C4 isomerization and alkylation unit.18,19 dThe HHVs of FT-derived gasoline and diesel are assumed to be 45 471 and 47 655 kJ/kg.

Plus 7.3.2. This model can reasonably estimate the plant performance with different design parameters and be used for techno-economic analysis. The modeling approach for the autothermal reforming, product upgrading section, and combined cycle plant is presented in this paper, while the modeling approach for the syngas and syncrude production units and pros and cons of several different CCS technologies have been discussed in our previous publication.1 Sensitivity studies have been conducted to analyze the impact of key design parameters on the performance of the novel indirect CBTL plant. The results indicate that a low steam/ carbon ratio in the ATR inlet is preferred for FT application because a high H2/CO ratio in the ATR outlet would result in higher penalty for CCS. The integrated hydrotreating approach can reduce the utility and capital investment of the product upgrading section. The fuel yield is found to increase with the decrease in the biomass/coal ratio and the increase in the H2/CO ratio in the FT inlet stream. The thermal efficiency is found to increase with the decrease in the biomass/coal ratio, the increase in the H2/CO ratio in the FT inlet stream, and the decrease in the extent of CCS. The thermal and carbon efficiencies of the base case (with 56.9% of carbon in the feed stored in captured CO2 and 6.7% of carbon vented to the atmosphere) are about 46.6 and 36.4%, respectively, which are higher than the data reported in the literature for a similar product yield and extent of CO2 capture. It should be noted that, to optimize the key design parameters, the thermal and carbon efficiencies should not be the only criteria. A techno-economic study that captures the impact of the key design parameters on the capital and operating costs needs to be considered.

■

Table A1. Results from the ATR Model in Comparison to the Bechtel Data19 feed flow rate (kmol/h) H2O CO2 H2 CO CH4 O2 N2 C2−C4

■

recycle

steam

product oxygen

3068 38 4507 4316 465 1128 169

415 2

model

reported

2378 944 6088 4017 205

2365 950 6114 4018 200

1128 0.44

1128

APPENDIX B

Isomer Distribution of the LTFT Product

In this study, the isomer distribution of paraffin in the LTFT product reported by Weller and Friedel,50 shown in Table B1, is considered for the estimation of utility consumption in the integrated hydrotreating unit.

■

APPENDIX C

Specification of the Hydrocarbon Recovery System

The specifications of the hydrocarbon recovery system are listed in Tables C1 and C2, which are obtained on the basis of the traditional crude oil distillation technology74−77 and the multicomponent distillation column used in the FT process design by Bechtel19 with limited information. In the hydrocarbon recovery system, the syncrude passes through a preheating train with several heat exchangers using the pumparound streams and the product streams that need to be cooled before entering the main distillation column. A feed furnace is used for the crude oil distillation tower instead of a reboiler, evaporating only a small portion of the wax. The feed furnace is specified by applying a fractional overflash of 3.2% LV. Stripping steam is used for decreasing the partial pressure of the hydrocarbons to prevent decomposition, which occurs at a high temperature (about 371 °C). A commonly used value for

APPENDIX A

ATR Model Validation

The model of the ATR unit is validated by comparing to the data reported in the literature, as shown in Tables A1−A3.16,19,48 5446


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Energy & Fuels Table A2. Results from the ATR Model in Comparison to the Commercial-Scale CTL Plant of NETL16 feed flow rate (kmol/h)

recycle

H2O CO2 H2 CO CH4 O2 N2 C2−C4

steam

product oxygen

3367 38 7289 576 2344

model

reported

1814 269 12855 2527 548

1904 232 12679 2521 591

4673

4673

430 8

4673 185

Table C1. Column Specification of the Hydrocarbon Recovery Section integrated approach number of trays main column heavy naphtha side stripper diesel side stripper stabilizerb locations feed to main column stripping steam to main column heavy naphtha product draw and return diesel product draw and return pump-around 1 draw and return pump-around 2 draw and return feed to stabilizer


Table A3. Results from the ATR Model in Comparison to the Small-Scale CTL Plant of NETL48 feed flow rate (kmol/h)

recycle

H2O CO2 H2 CO CH4 O2 N2 C2−C4

steam

product oxygen

482 8 1251 132 75 644 30

model

reported

452 48 1454 207 23

456 42 1455 216 20

642

644

62 1

isomer

molar fraction

isomer

molar fraction

1 0.95 0.05 1 0.896 0.057 0.047 1

n-heptane 2-methylhexane 3-methylhexane 1-octene n-octane 2-methylheptane 3-methylheptane 4-methylheptane

0.877 0.046 0.077 1 0.845 0.039 0.072 0.044

23a NA 5 20

26a 30a 15, 14

19a 23a NA

24, 23 15, 12 24, 21 10

17,a 16 NA 17, 14 10

Table C2. Specification of the Column Operating Condition

main column condenser temperature (°C)a overhead pressure (kPa)a pressure drop per tray (kPa)a feed furnace fractional overflash (% LV) bottom product to feed ratio (kg/kg) stripping steam to bottom product ratio (kg/bbl) side strippers stripping steam to heavy naphtha ratio (kg/bbl) stripping steam to diesel ratio (kg/bbl) pump-around and preheating train pump-around 1 return temperature (°C) pump-around 2 return temperature (°C) heavy naphtha heat exchanger hot stream temperature drop (°C) diesel heat exchanger hot stream temperature drop (°C) wax heat exchanger hot stream temperature drop (°C)

74−77

stripping steam to product ratio is about 2.27−4.54 kg/bbl. Pump-arounds are used as main means to obtain intermediate heat recovery. Liquid is withdrawn from the tray on or above the lower product draw tray, cooled, and returned to a tray, 2−3 trays above but below the upper product draw.74−77 As a result, the size and heat duty of the feed furnace and the overhead condenser could be reduced significantly. Meanwhile, the top reflux and the column diameter could also be reduced. In this study, the outlet temperatures of the two pump-around exchangers are selected to increase the heat recovery as much as possible within operating constraints.

■

30a 5 5 20

a Numbers are obtained from the literature. bStabilizers are designed using the shortcut model in Aspen Plus.

Table B1. Isomer Distribution of Hydrocarbons in the LTFT Product50 1-pentene n-pentane isopentane 1-hexene n-hexane 2-methylpentane 3-methylpentane 1-heptene

separated approach

a

integrated approach

separated approach

37.8 600 1.38 3.2 0.48 4.54

37.8 600 1.38 3.2 0.48 4.54

2.27

NA

2.27

2.27

82.2

NA

282.2

83.3

66.7

NA

85.6

51.7

193.3

194.4

Numbers are obtained from the literature.

isentropic enthalpy drop (Δhis), and outlet steam condition can be solved by eqs D1 and D2 and IAPWS IF97 correlations. The stage power output (Wi) is calculated by eq D3, where isentropic efficiency (ηst) is a known function of Ns, and the average moisture content across the stage is given by Lozza.64 The net power output of the steam turbine is shown in eq D4. If no information is available, the exhaust velocity of the last stage (vex) is assumed to be 250 m/s79

APPENDIX D

Stage-by-Stage Model of Steam Turbine

A simple stage-by-stage calculation is performed in MATLAB for estimating the performance of a three-pressure-level steam turbine with multiple steam addition and extraction points based on the algorithm presented by Lozza.64 In the model, the steam properties are evaluated by the IAPWS IF97 correlations and coded in MATLAB.78 Given the flow rate, pressure, and temperature of the stage inlet, specific speed (Ns), stage

Ns = (RPM/60) Vex /(Δh is)0.75

(D1)

Δh is = k isu 2 /2

(D2)

where Vex is the volumetric flow rate at stage outlet under isentropic conditions in m3/s, Δhis is the stage isentropic enthalpy drop in J/kg, u is the mean diameter peripheral 5447


Article

Energy & Fuels Table E1. Specifications of Project Components of the Integrated Hydrotreating Unit in APEA

a

description

number required

number of spares

model in APEA

sizing

MOC

×103 $ (2013)

syncrude pump to hydrotreator hydrotreator feed furnace feed/product heat exchanger hydrotreating reactora product cooler high-pressure flash H2 recycle compressor catalyst

1 1 1 1 1 1 1 1

1 0 0 0 0 0 0 0

CP CENTRIF FU BOX HE FLOAT HEAD VT MULTI WALL HE FLOAT HEAD VT CYLINDER GC CENTRIF C

Icarus Icarus EDR correlation EDR Icarus Icarus NA

CS casing A213F A285C, A214 SS347 A285C, A214 A516 SS316 NA

334 662 489 3370 129 188 1745 1251

MOC

×103 $ (2013)

A516 A285C, A214 A285C, A214 A285C, A214 A285C, A214 A285C, A214 A285C, A214 A516 CS casing A516, A285C A516, A285C A516, A285C A213C CS casing A285C, A214 A516 A285C, A214 CS casing A516, A285C

144 88 102 78 84 121 120 105 103 693 167 162 765 87 75 73 92 97 275

The hydrotreater is sized by assuming the same space velocity and L/D ratio as reported in the literature.24


Table E2. Specifications of Project Components of the Hydrocarbon Recovery Unit in APEA description

number required

number of spares

model in APEA

low-pressure flash pump-around 1 pump-around 2 heavy naphtha heat exchanger diesel heat exchanger wax heat exchanger main column, condenser main column, drum main column, reflux pump main column, tower side stripper, diesel side stripper, heavy naphtha main column, feed furnace pump to the stabilizer stabilizer, condenser stabilizer, drum stabilizer, reboiler stabilizer, reflux pump stabilizer, tower

1 1 1 1 1 1 1 1 1 1 1 1 1 1 1 1 1 1 1

0 0 0 0 0 0 0 0 1 0 0 0 0 1 0 0 0 1 0

VT CYLINDER HE FLOAT HEAD HE FLOAT HEAD HE FLOAT HEAD HE FLOAT HEAD HE FLOAT HEAD HE FIXED T S HT HORIZ DRUM CP CENTRIF TW TRAYED TW TRAYED TW TRAYED FU BOX CP CENTRIF HE FIXED T S HT HORIZ DRUM RB U TUBE CP CENTRIF TW TRAYED

Pnet = ηg ηmηl (∑ Wi − mlast stageΔhek ) i

(D3)

■

(D4)

where, ηg, ηm, and ηl are the generator loss, mechanical loss, and sealing loss, which are a function of the steam turbine power rating64 and Δhek = νex2/2 is the energy loss as a result of axial exhaust velocity.

■

Icarus EDR EDR EDR EDR EDR EDR Icarus Icarus Aspen Aspen Aspen Icarus Icarus EDR Icarus EDR Icarus Aspen

Plus Plus Plus

Plus

stream, and industry application experience. In the hydrotreating unit, the reactor and H2 compressor are constructed by stainless steel and the hydrotreater feed furnace is constructed by Cr−Mo low-alloy steel, while the other components are constructed by carbon steel. Tables E1 and E2 lists the specification and capital investment of each component in the integrated hydrotreating loop and the hydrocarbon recovery unit.

velocity of steam turbine in m/s, which is given by a function of the stage number,64 and kis is the stage head coefficient and correlated with Ns Wi = Δhi = miΔh is, iηst, i

sizing

AUTHOR INFORMATION

Corresponding Author

*Telephone: 1-304-293-9335. E-mail: debangsu.bhattacharyya@ mail.wvu.edu. Notes

The authors declare no competing financial interest.

APPENDIX E

■

Capital Cost Estimation of the Integrated Hydrotreating Approach

ACKNOWLEDGMENTS The authors gratefully acknowledge financial support from the U.S. Department of Energy (DOE) through Grant DEFE0009997 titled “Feasibilities of a Coal-Biomass to Liquids Plant in Sothern West Virginia”. The authors strongly acknowledge discussions with Eric Liese from the National Energy Technology Laboratory (NETL), Morgantown, WV, for the steam turbine design.

For a fair comparison between the novel integrated hydrotreating approach and the conventional separated hydrotreating approach, the capital costs of both technologies are evaluated in this study. The capital investment of the separated hydrotreating unit is reported by Bechtel,18 while the capital investment of the integrated hydrotreating approach is estimated in APEA, using Icarus database. All of the columns are sized in Aspen Plus; all of the heat exchangers are sized in Exchanger Design and Rating (EDR); and the remaining equipment items are sized in APEA. The materials of construction (MOC) for all of the equipment are selected on the basis of the operating temperature, service

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5448

NOMENCLATURE AGR = acid gas removal APEA = Aspen process economic analyzer DOI: 10.1021/acs.energyfuels.5b00360 Energy Fuels 2015, 29, 5434−5451

Article


Energy & Fuels

Conference of the Americas; Oakland, CA, Aug 29−Sept 2, 1999; Vol. 1, pp 843−853. (8) Wang, J.; Grushecky, S.; McNeel, J. Biomass Resources, Uses, and Opportunities in West Virginia; Division of Forestry and Natural Resources, West Virginia University: Morgantown, WV, 2007. (9) IEA Greenhouse Gas R&D Programme (IEAGHG). Potential for Biomass and Carbon Dioxide Capture and Storage; IEAGHG: Cheltenham, U.K., July 2011. (10) Bhattacharyya, D.; Turton, R.; Zitney, S. E. Steady-state simulation and optimization of an integrated gasification combined cycle power plant with CO2 capture. Ind. Eng. Chem. Res. 2011, 50, 1674−1690. (11) National Energy Technology Laboratory (NETL). Cost and Performance Baseline for Fossil Energy Plant Volume 1: Bituminous Coal and Natural Gas to Electricity; NETL: Pittsburgh, PA, Nov 2010; DOE/NETL-2010/1397, http://www.netl.doe.gov. (12) Hasan, M.; Baliban, R.; Elia, J.; Floudas, C. Modeling, simulation, and optimization of postcombustion CO2 capture for variable feed concentration and flow rate.1. Chemical absorption and membrane process. Ind. Eng. Chem. Res. 2012, 51 (48), 15642−15664. (13) IEA Greenhouse Gas R&D Programme (IEAGHG). CO2 Capture Ready Plants; IEAGHG: Cheltenham, U.K., May 2007; https://www.iea.org. (14) Hasan, M.; Baliban, R.; Elia, J.; Floudas, C. Modeling, simulation, and optimization of postcombustion CO2 capture for variable feed concentration and flow rate.2. Pressure swing adsorption and vacuum swing adsorption processes. Ind. Eng. Chem. Res. 2012, 51 (48), 15665−15682. (15) Liu, G.; Larson, E.; Williams, R. H.; Kreutz, T. G.; Guo, X. Making Fischer−Tropsch fuels and electricity from coal and biomass: performance and cost analysis. Energy Fuels 2011, 25, 415−437. (16) National Energy Technology Laboratory (NETL), Baseline Technical and Economic Assessment of a Commercial Scale Fischer− Tropsch Liquids Facility; NETL: Pittsburgh, PA, 2007; DOE/NETL2007/1260, http://www.netl.doe.gov. (17) Larson, E. D.; Fiorese, G.; Liu, G.; Williams, R. H.; Kreutz, T. G.; Consonni, S. Co-production of decarbonized synfuels and electricity from coal and biomass with CO2 capture and storage: an Illinois case study. Energy Environ. Sci. 2010, 3, 28−42. (18) Bechtel Corp. Baseline Design/Economics for Advanced Fischer− Tropsch Technology; Bechtel Corp.: San Francisco, CA, April 1998; DE-AC22-91PC90027, http://www.fischer-tropsch.org. (19) Bechtel Corp. Baseline Design/Economics for Advanced Fischer− Tropsch Technology; Bechtel Corp.: San Francisco, CA, Jan−March 1993; DE-AC22-91PC90027, http://www.fischer-tropsch.org. (20) Baliban, R. C.; Elia, J. A.; Floudas, C. A. Optimization framework for the simultaneous process synthesis, heat and power integration of a thermochemical hybrid biomass, coal and natural gas facility. Comput. Chem. Eng. 2011, 35, 1647−1690. (21) Rafiq, M. H.; Jakobsen, H. A.; Hustad, J. E. Modeling and simulation of catalytic partial oxidation of methane to synthesis gas by using a plasma-assisted gliding arc reactor. Fuel Process. Technol. 2012, 101, 44−57. (22) Steynberg, A. P.; Dry, M. E. Fischer−Tropsch Technology, 1st ed.; Elsevier Science: San Diego, CA, 2004. (23) Cavallo, E.; Januszweski, D.; Putek, S. Upgrade hydrocracked resid through integrated hydrotreating. Hydrocarbon Process. 2008, 87 (9), 83−92. (24) Jarullah, A. T.; Mujtaba, I. M.; Wood, A. S. Whole Crude Oil Hydrotreating from Small-Scale Laboratory Pilot Plant to Large-Scale trickle-Bed Reactor: Analysis of Operational Issues through Modeling. Energy Fuels 2012, 26, 629−641. (25) Comolli, A. G.; Lee, L. K.; Pradhan, V. R.; Stalzer, T. H.; Harris, E. C.; Mountainland, D. M.; Karolkiewicz, W. F.; Pablacio, R. M. Direct Liquefaction Proof-of-Concept Facility; Hydrocarbon Research, Inc.: Princeton, NJ, 1995; DE-AC22-92PC92148. (26) Fahim, M. A.; Al-Sahhaf, T. A.; Elkilani, A. Fundamentals of Petroleum Refining, 1st ed.; Elsevier Science: Oxford, U.K., 2010; pp 170−179.

ASU = air separation unit ATR = autothermal reformer BFD = block flow diagram CBTL = coal biomass to liquids CCS = CO2 capture and storage CEPCI = chemical engineering plant cost index GHG = greenhouse gas CI = cetane index CTL = coal to liquids EOS = equation of state EDR = exchanger design and rating FBP = final boiling point FG = fuel gas FT = Fischer−Tropsch FTL = Fischer−Tropsch liquid GT = gas turbine GTL = gas to liquids HHV = higher heating value HP = high pressure HPF = high pressure flash HRSG = heat recovery and steam generation IGCC = integrated gasification combined cycle LHV = lower heating value IP = intermediate pressure LG = light gases LP = low pressure LPF = low pressure flash LTFT = low-temperature Fischer−Tropsch MDEA = methyldiethanolamine MOC = materials of construction MON = motor octane number MP = medium pressure PFR = plug flow reactor PSA = pressure swing adsorption PZ = piperazine RFG = reformulated gasoline RON = research octane number RVP = Reid vapor pressure SCFB = standard cubic feet per barrel SHR = syngas heat recovery ST = steam turbine SWS = sour water stripper WGS = water-gas shift

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