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Technoeconomic Analysis for the Production of Mixed Alcohols via Indirect Gasification of Biomass Based on Demonstration Experiments Abhijit Dutta,*,† Jesse Hensley,† Richard Bain,† Kim Magrini,† Eric C. D. Tan,† George Apanel,‡ David Barton,§ Peter Groenendijk,§ Daniela Ferrari,§ Whitney Jablonski,† and Daniel Carpenter† †

National Renewable Energy Laboratory, 15013 Denver West Parkway, Golden, Colorado 80401, United States Rentech, 4150 East 60th Avenue, Commerce City, Colorado 80022, United States § The Dow Chemical Company, Midland, Michigan 48674, United States ‡

S Supporting Information *

ABSTRACT: An integrated biomass-to-mixed alcohol process was demonstrated at the pilot scale, including indirect gasification, tar and hydrocarbon reforming, gas conditioning, and gas to liquids operations. Additional bench-scale experiments were conducted to gather insights into reforming and fuel synthesis operations under more extensive conditions. Data from these experiments were combined with tools and knowledge owned by industrial partnersRentech and The Dow Chemical Companyto validate a conceptual commercial-scale cellulosic ethanol refinery. Results suggest that both reforming and mixed alcohol processes are scalable and economical, with a modeled mature plant ethanol production cost of $0.54/L ($0.82/L of gasoline equivalent). Sustainability metrics for the conversion process are also presented.

1. INTRODUCTION In 2006, the U.S. Department of Energy (DOE) set a 6 year goal to enable the production of cost-competitive lignocellulosic ethanol to help achieve the displacement of petroleumderived fuels as mandated by the Advanced Energy Initiative.1 Two conversion routes, thermochemical and biochemical, were adopted. This work details research and development outcomes under the thermochemical platform and is the culmination of a multiyear effort that included collaborative research between national laboratories and industrial partners. Conceptual designs are important and informative for providing direction and showing the practical implications of the research.2 At project onset, conceptual process designs for the conversion of biomass to ethanol were developed, and key economic barriers were identified.3 Within potential process configurations, many unit operations could be adapted from chemical and petroleum technology, but two areas required substantial innovation: (1) conversion of tars and light hydrocarbons (produced during biomass gasification) into H2 and CO (syngas) for adequate carbon utilization and syngas quality, and (2) syngas-to-ethanol catalyst(s) with improved performance and stability. In both areas, scalability was paramount. Thus, research and development were focused on tar reforming and ethanol synthesis, with time-staged performance targets leading up to 2012. Following initial design, a more detailed benchmark model was developed based on improved understanding, inclusive of required tar reforming and alcohol synthesis improvements, and peer reviewed by commercial, academic, and government entities.4 Full details of the process and financial assumptions were provided, along with the state of technology and projections of future improvements with impacts on a socalled minimum ethanol selling price (MESP) for a mature © XXXX American Chemical Society

commercial plant. In this work, the benchmark described previously is compared to bench- and pilot-scale experimental results as a means of process demonstration and modeled cost validation. Data from experiments and modeling are presented side by side where possible. For brevity, further details of experimental equipment and procedures are included in a sister publication,5 as well as other referenced publications throughout the work. 1.1. Process Model and Experimental Setup. Figure 1a shows a simplified flow diagram of the benchmark process model4,6 for the conversion of biomass to mixed alcohols. Biomass is gasified in an indirect gasifier (A), and the syngas is reformed to convert methane, light hydrocarbons, and tars to syngas in a steam reformer (B). Gases are cooled and scrubbed (C) and then compressed (D). The compressed clean syngas is mixed with unreacted gases and methanol from the synthesis reactor, spiked with H2S to maintain sulfide catalyst activity, and then sent to an alcohol synthesis reactor (E). CO2, a byproduct of alcohol synthesis, is removed with an acid gas removal system located within the syngas recycle loop (F). Alcohol product is cooled/condensed before fractionation. In total, the process is energy-sufficient; that is, all required process heat and electricity are ultimately generated from the biomass source. Capacities and flow rates of individual equipment in the 2,000 dry metric tons (or megagrams (Mg)) per day plant are detailed in Appendices B and E of the 2011 design report.4 Received: June 30, 2013 Revised: May 24, 2014 Accepted: June 24, 2014

A

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Figure 1. Simplified flow diagrams for (a) benchmark design process4 and (b) NREL pilot plant process. Diagrams are not to scale and do not include all equipment.

Some of the differences in pilot- and commercial-scale designs were born out of engineering and economic necessities. For example, multiple recycle streams and tight heat integration are not affordable at the smaller pilot scale. Second, the indoor location of the described system limits the use of some operations (such as very high pressures) due to safety considerations. Third, the use of a hydrodynamically similar tubular mixed alcohol reactor requires a greater syngas generation capacity than is presently available at NREL. Also, differences during the gasification and gas conditioning steps are discussed in detail in this work. The pilot process equipment and operations are described in further detail elsewhere.5

The National Renewable Energy Laboratory’s (NREL’s) pilot-scale demonstration plant is shown in Figure 1b. Included are an electrically heated bubbling fluidized bed gasifier and thermal cracker (A), electrically heated reformers (B), dodecane scrubber (C), pressure swing adsorber (PSA) for CO2 removal (F), syngas compressors (D), and a 10 L continuously stirred tank reactor (CSTR) for alcohol synthesis (E). Additional subunits, whose roles are discussed separately, were used to validate the process model. These include a recirculating regenerating reformer (R3) and a slip-stream microactivity reactor test system (SMARTS) reformer, a syngas generator, and a bench-scale alcohol synthesis fixed bed reactor (FBR). B

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1.2. Reforming. Hydrocarbons and tars produced during biomass gasification are reformed to boost carbon utilization and to improve process yields. The process design4 considers endothermic steam reforming that follows the stoichiometry [CnHmOp + (n−p) H2O → nCO + ((m/2)+(n−p)) H2]. For example, steam−methane reforming follows the general reaction [CH4 + H2O → CO + 3H2]. Methane is abundant in biomass-derived syngas and is also difficult to reform. Its conversion to CO and H2 is a key indicator of catalyst performance.7 Targets for the performance of steam reforming catalysts are shown in Table 1.8 Biomass syngas contains many

recently, a version of the catalyst was separately improved by The Dow Chemical Company and NREL. The NRELimproved catalyst was used for the pilot- and bench-scale demonstrations described in this publication. The K/Co/MoS2 catalyst was chosen over the rhodium-based one because of the availability of a detailed kinetic model, better compatibility with the syngas cleanup strategy being employed (especially with respect to sulfur removal), and the relative ease of producing pilot-scale quantities of material. The Dow Chemical Company provided a fundamental kinetic model (FKM) for mixed alcohol synthesis. This FKM was adjustable, allowing model predictions to follow variations in catalyst performance with process inputs. 1.4. Technoeconomic Analysis. The benchmark process used here is detailed elsewhere4 and includes calculation of and sensitivities around the base MESP. Projections assume mature nth plants, 2007 dollars, a 60:40 debt to equity ratio, debt financing for 10 years at 8%, a 10% return on equity financing, 7-year modified accelerated cost recovery system (MACRS) depreciation,12 35% income tax rate, and 30 year plant life. Details are provided in several tables in our design report:4 (1) financial assumptions are shown in Table 1; (2) cost factors for indirect costs are provided in Table 27; (3) details on the calculation of total capital investment ($516 million) are given in Table 28; (4) operating costs and their bases are provided in Tables 29, 30, and 31; and (5) the justification for co-product valuation at $0.50/L ($1.88/gallon) for higher alcohols is shown in Table 32.4 1.5. Sustainability Metrics. Key sustainability metrics for the conceptual process are presented to quantify environmental impacts. Direct CO2, NO2, and SO2 emissions, consumptive water use, and other process-related metrics were derived from the integrated process model.4 The boundary for all metrics is the biorefinery (or the fuel production stage) which was the focus of this research. Upstream processes (i.e., feedstock production and transportation) and downstream processes (i.e., fuel distribution and vehicle operation) are not considered in this study. Embodied emissions and resource consumption from the biorefinery related to construction and maintenance, material and energy inputs, and emission/consumption credits associated with co-products were also included. For life-cycle assessment, SimaPro v.7.313 software was used to develop and link units quantifying life-cycle impacts as previously documented by Hsu et al.14 with Ecoinvent v.2.015 and the U.S. Life Cycle Inventory (LCI)16 processes used to fill data gaps. The Ecoinvent processes were modified to reflect U.S. conditions, and the U.S. LCI processes were adapted to account for embodied emissions and energy usage. The LCI of the conversion step captures the impacts of input raw materials and outputs, such as emissions and waste as predicted by the process model, similar to those documented in Table 24 of the 2011 design report.4 In addition to the primary product ethanol, a co-product stream of higher alcohols is also available from the mixed alcohols process. These co-products can be used as stationary equipment fuel or can potentially be blended into gasoline following specifications of the Octamix waiver.17−19 Based on this premise, greenhouse gas (GHG) emission burdens are allocated between ethanol and the higher alcohols products by applying the direct energy allocation method20 using their lower heating values (LHV).

Table 1. Catalyst Improvement Targets Reforming Catalyst species methane benzene tars

conversion (%) 80 99 99

catalyst replenishment rate (for fluidizable 0.1% of inventory per day catalyst) Alcohol Synthesis Catalysta productivity (activity) improvement over initial state of the art catalyst selectivity for alcohols

20% maintained at the 20% higher activity

a

Specific performance improvements resulting from these targets, as manifested in a benchmark model, are shown in Appendix I of the 2011 design report4 and also in Table SI-5 in the Supporting Information.

contaminants such as H2S that can poison and/or deactivate steam reforming catalysts.7 In addition, coke deposits can block catalyst sites, rendering the catalyst ineffective. Thus, reforming catalysts must be regenerable upon deactivation. To that end, Rentech provided valuable insights into catalyst regenerability during steam and dry reforming of actual and simulated biomass syngas, as discussed later. 1.3. Alcohol Synthesis. Syngas is converted to a mixture of linear aliphatic alcohols on a sulfided cobalt molybdenum catalyst. Alcohol synthesis is an exothermic reaction, with an overall stoichiometry [nCO + 2nH2 → CnH2n+1OH + (n− 1)H2O]. For ethanol (n = 2) this reduces to [2CO + 4H2 → C2H5OH + H2O]. Alcohol synthesis reaction mechanisms, which vary by catalyst, are complicated. Intermediate and continuing reactions produce multiple byproducts, including methanol, C3+ alcohols, other oxygenates, methane, and light hydrocarbons. In addition, CO2 is produced from syngas via the water-gas shift reaction on sulfide-type catalysts. Therefore, detailed kinetic models are required to predict alcohol production within rigorous process models, especially when byproduct streams are recycled and when reactant concentrations can vary due to upstream changes. The primary research target was improvement of the catalysts’ ethanol productivity by increasing specific activity without sacrificing selectivity for ethanol (Table 1), which effectively allows one to produce the same amount of ethanol per mass of biomass while purchasing less catalyst and reducing reactor volume. Two different catalysts were considered, one of them a rhodiumbased catalyst developed by Pacific Northwest National Laboratory9,10 and the second a potassium-promoted cobalt molybdenum sulfide catalyst initially introduced and subsequently improved by The Dow Chemical Company.11 More C

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Figure 2. Pilot-scale reformers at NREL: (a) recirculating regenerating reformer (R3) circulating fluidized bed, (b) R600 fluidized bed, and (c) PBR (packed bed reformer) systems. (Diagrams not to scale.)

2.3. Reforming. Summaries of reforming catalysts and equipment used to develop, evaluate, and demonstrate those catalysts are shown in Table SI-2 in the Supporting Information. Many catalysts were tested in the bench-scale (SMARTS) reformer using biomass-derived syngas for the purpose of down-selection. An NREL-developed catalyst met all conversion targets (Table 1) at ≥925 °C, while a Johnson Matthey reforming (JMR) catalyst met the same targets at ≥800 °C. Other experimental and industrial catalysts did not meet target specifications. Hence, NREL and JMR catalysts were chosen for the validation exercises. The NREL catalyst, which is fluidizable by design, has previously been shown to maintain performance through reforming−regeneration cycles23 and was thus demonstrated in a recirculating regenerating system. 2.3.1. Recirculating Regenerating Reformer. 2.3.1.1. NREL Pilot Plant R3 System. Initially, demonstration activities were planned with utilization of only the R3 system for reforming (Figure 2a). The system was designed to demonstrate continuous reforming in a riser reformer where the catalyst is entrained in flowing raw syngas, allowing short-contact reforming in the riser, resulting in modest catalyst deactivation in each reforming cycle. Avoidance of deep deactivation associated with long contact times allows the catalyst to be regenerated more easily.23 Upon exiting the riser, the catalyst is separated from the reformed syngas using a cyclone. The separated catalyst is then exposed to regeneration gases, which can consist of H2/CO/N2/steam, as it falls via gravity through a separate vessel. The regenerated catalyst is then returned to the reformer, with motive nitrogen used to control catalyst flow. Upon commissioning, the system went through shakedown, where it experienced frequent plugging, preventing proper catalyst circulation. One of the runs demonstrated 83%−87% methane conversions at 900 °C for approximately 90 min, with nearly 35 reforming/regeneration cycles/hour and a 70 kg catalyst inventory. However, this steady-state operation was interrupted upon the introduction of some steam to study performance sensitivity, which caused excessive catalyst entrainment. While many additional short periods of successful continuous reforming were demonstrated, the system could not be fully balanced in time to complete scheduled activities, and

2. EXPERIMENTAL SECTION The pilot plant materials, equipment, and operations necessary for explanation of the technoeconomic analysis (TEA) are discussed below. Additional details are available in a sister publication.5 2.1. Feedstock. White oak pellets with average dimensions of 0.64 cm diameter and 1.27−2.54 cm length were used. Pellets were crushed prior to introduction for ease of feeding. Table SI-1 in the Supporting Information provides the ultimate (elemental) analysis and moisture content of the feedstock. A woody biomass feedstock cost of $67.87/(dry Mg) ($61.57/(dry US ton)) in 2007 dollars was used for the economic assessment, based on cost evaluation by researchers at the Idaho National Laboratory.4,8 The benchmark design was based on a 2,000 dry Mg/day feed rate. The sensitivity of the MESP to the variations in feedstock costs are shown in the 2011 design report.4 2.2. Biomass Gasification. Biomass gasification was performed in the NREL pilot-scale system21 as shown in Figure 1b. An electrically heated 20.3 cm i.d. olivine bed (23.25 kg load, 250 μm particle size) at 650 °C and 75 kPag was fluidized with superheated steam and CO2. Steam and biomass were fed at a ratio of 1.8 (mass/mass). This ratio is significantly higher than the ratio in the process model (0.4) and is due to the limitation of minimum steam required for fluidization. This resulted in nonoptimal gas compositions (high H2:CO) at the downstream alcohol synthesis reactor and adversely affected alcohol synthesis performance, as discussed later. The addition of CO2 to steam partially corrected this problem by forcing a reverse water-gas shift in the downstream reformer, but did not lower the H2:CO to modeled ratios. Product gas flowed into a thermal cracker (internal volume, 0.028 m3), maintained at 900 °C. The resulting gas/vapor/solid stream then entered two cyclone separators in series to remove entrained solids. Removed solids were cooled, collected, and weighed. While only oak gasification was considered in this work, a comprehensive study of the gasification of dif ferent feedstocks in the system described above can be found in a separate study.22 Gasifier operations are discussed in the Supporting Information and are shown in Figure SI-1. D

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Figure 3. Schematic of mixed alcohol synthesis (a) pilot-scale CSTR and (b) bench-scale FBR. (Drawings are not to scale.)

reforming was instead demonstrated using a fluidized bed reformer (R600) and NREL catalyst without regeneration,24 followed by a packed bed reformer (PBR) with the JMR catalyst (Figure 1b). 2.3.1.2. Pilot-Scale Validation by Rentech. The NREL reforming catalyst was tested separately in a recirculating regenerating configuration by Rentech.25 Several continuous reforming tests with synthetic syngas spiked with high levels of H2S and tar surrogates were conducted in 2009 (phase I), as with biomass-derived syngas and natural gas in 2011 (phase II) using a system similar (but not completely identical) to the one depicted in a patent.25 The catalyst regenerator uses hot flue gases with excess air to burn off carbon deposits. Additional gaseous fuel sources may be injected into the regenerator to maintain the endothermic reformer at a desired temperature via heat transfer from the hot regenerated catalyst. During the phase II tests, biomass-derived syngas was continuously provided from a 1 ton/day biomass gasifier. In both phases, reforming was demonstrated without complication, including a continuous test of 54 h in phase II, reusing the catalyst from phase I. Methane conversion averaged 95% with benzene and identifiable individual tar components reduced to below detection, giving a syngas with H2:CO of approximately 1.8. Online gas analysis with the aid of internal standards helped achieve >90% mass balance closure across the gasifier and reformers. Tars were separately analyzed using grab bag samples. Dry reforming experiments were conducted using a 1:1 mix of natural gas and CO2, with no added moisture. The mixture was spiked with H2S and tar surrogates, both at levels significantly elevated over wood-derived syngas, to simulate dry reforming of biomass syngas. Even under these conditions, methane conversion was above 80% and catalyst activity was maintained, showing that the fluidized catalyst was effectively regenerated in the recirculating system even under severely carbonizing conditions. Literature information suggests that the propensity of carbon deposition is reduced by operating close to 900 °C.26 High conversions during dry (CO2) reforming of methane using Ni-based catalysts have been previously reported,26,27 although long-term operations were not possible because of carbon deposits in a fixed bed. Proposed reaction mechanisms for CO2 reforming of methane are available in the literature.28 Tests by Rentech also showed that higher methane

conversions can be achieved by supplemental moisture during CO2 reforming. Given the preceding insights, the fluidized reformer and regenerator configuration was retained as the basis for technoeconomic analysis in this work, even though the shakedown of the NREL R3 system could not be completed in time for the integrated demonstration. 2.3.2. Steady-State Fluidized Bed Reformer (R600). A schematic of the 35.6 cm i.d. fluidized bed reformer (R600) is shown in Figure 2b. The reforming catalyst is fluidized by syngas and held at 900 °C by electrical heaters. An initial charge of 60 kg of the fluidizable catalyst was loaded into the reactor and activated by reduction with further additions to make up for process upsets and operational losses.5 2.3.3. Packed Bed Reformer. Because it was known from experience that the R600 reactor/catalyst configuration cannot maintain the 80% methane conversion target without regeneration,24 a subsequent PBR was installed to complete the conversion of tars and hydrocarbons to the desired levels in a continuous mode (Figure 1b). Syngas from the fluid bed reformer was sent to the PBR (Figure 2c) after passing through an alternating dual bed filter system.5 The 15.2 cm i.d. tubular reactor was loaded with 9 kg of JMR catalyst. The PBR was operated in a down-flow configuration and maintained between 770 and 780 °C using three bayonet-type heaters within the bed. 2.4. Syngas Conditioning. Reformed syngas was quenched and condensable vapors were removed using a dodecane scrubber system that has been described elsewhere.22 Syngas was compressed to 5.52 MPag and fed to a small PSA unit for CO2 removal.5 Syngas was then compressed to 20.7 MPa and stored in gas accumulators to ensure smooth and uninterrupted operation of the mixed alcohol synthesis reactor. An activated carbon bed was used to remove any iron carbonyls resulting from storage of the syngas in carbon steel accumulators. A slip stream of PSA effluent was compressed to 13.8 MPa and fed to a bench-scale FBR. Additional details regarding the use of a clean (nonbiomass) source of syngas for activation of the mixed alcohol catalyst and methods for syngas analysis are available in a sister publication.5 E

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Table 2. Performance of Reformer Systems system

design model: circulating reformer−regenerator dual bed system

NREL pilot: steady-state fluidized bed (R600) followed by PBR

catalyst catalyst loss (% inventory/day)

R&D target product 0.1

NREL in R600, JMR in PBR N/A

feed

gasifier products with recycled process gases, including CO2

Inlet gasifier products, including steam required for gasifier fluidization and added CO2

steam:carbona CO2:carbona

2.0 1.1

methane (CH4) benzene (C6H6) tars

80 99 99

a

6.2 2.3 Conversions (%) 86 97 99.9

Rentech-directed pilot runs: circulating reformer−regenerator dual bed system (NREL catalyst was one of several considered for the Rentech runs) NREL NREL 0.15 0.15

biomass-derived syngas from 1 ton/day pilot gasifier 1.8 1.1

natural gas reformed with CO2 only, spiked with elevated H2S and tar species

95 99.9b 99.9b

>80

0 1

99.9b b

Carbon calculation excludes CO and CO2 and subtracts oxygen already present in the species to be reformed. Below detection limits.

2.5. Mixed Alcohol Synthesis. Mixed alcohol synthesis was demonstrated in a gas-phase CSTR, shown schematically in Figure 3a, and concurrently at the bench scale using an FBR, shown in Figure 3b. Operation was limited to 13.8 MPa due to in-house safety standards, less than the 20.7 MPa in the benchmark design.4 An NREL-developed K-CoMoSx catalyst was used.5 H2S was maintained at ∼70−100 ppm at the reactor inlet to sustain activity of the sulfided catalyst. Reactor operations and analytical procedures are detailed separately.5

when converted at 97%. In an industrial PBR system, the mode of heat transfer will be different from a direct-contact electrical heater (such as the one used here), and the reformer can be operated at higher temperatures than the 780 °C set point for the direct-contact bayonet heaters. In previous experiments in the SMARTS reactor, benzene removal was essentially complete above 800 °C using a JMR catalyst. Effluent benzene was below detection limits during Rentech-directed reforming in a recirculating regenerating mode using the NREL catalyst, indicating >99% conversion. Methane conversion, a significant driver of process economics,4 was above the 80% conversion target in all of the configurations. The process implications of the successful dry reforming (Table 2) of simulated biomass syngas during the Rentech tests can be far reaching. The results indicate that the NREL reforming catalyst can be used in a recirculating regenerating system to produce a range of H2:CO ratios by using CO2 and H2O in various proportions and that the syngas composition can be tailored for the optimal performance of most fuel synthesis catalysts. This can be advantageously incorporated into the design of biomass and other gasification-based processes in a number of ways. Byproduct CO2 can be made available and possibly even recycled to extinction via conventional acid gas removal systems or direct recycle of CO2-rich gas streams from the process, and can be used for fluidization of biomass in the gasifier to reduce process steam requirements.31 For example, CO2 is a byproduct from the mixed alcohol synthesis4 and Fischer−Tropsch32 reactors using iron-based catalysts because of the water-gas shift reaction. Preliminary calculations from the Rentech phase II runs estimate a catalyst loss of 0.15% of inventory per day, which is close to the 0.1% target in Table 1. 3.2. Acid Gas Removal. The PSA system removed CO2 from the syngas to levels more suitable for mixed alcohol synthesis. (See Table SI-4 in the Supporting Information for further details.) The H2:CO ratio also increased slightly across the PSA due to partial removal of CO. As an example, a gas with 35% CO2 and 3:1 H2:CO entering the PSA exited at 28% CO2 and 3.5:1 H2:CO. 3.3. Mixed Alcohol Synthesis. The catalyst productivity improvement target of 20%, while maintaining selectivity for ethanol (Table 1), was exceeded and was demonstrated in comparative bench-scale experiments with the initial and improved catalysts. The catalyst was also shown to be robust

3. RESULTS AND DISCUSSION 3.1. Gasification and Reforming. Table SI-3 in the Supporting Information compares gas compositions from the NREL pilot plant and those anticipated in the 2011 design case.4 Differences in gas composition can be attributed to the following: (1) steam was introduced into the gasifier in excess to maintain fluidization and was not countered by additional feed of biomass due to plant limitations, resulting in an increase of the H2:CO ratio; (2) CO2 was introduced into the gasifier to lower the high H2:CO ratio due to high gasifier steam; (3) CO2 removal was limited in the PSA system; and (4) N2 was introduced into the syngas through the solids feed system and through various purges in the process. None of these pilot plant limitations should exist in a commercial-scale system because (1) gasifier steam can be reduced by appropriate design to meet the desired H2:CO ratios, as was demonstrated during operation of the Battelle Columbus Laboratory (BCL)entrained flow gasifier29 (the basis for the design case4); (2) recycled CO2 from an integrated process can be a desirable steam substitute for fluidization in the gasifier, based on Rentech dry reforming results; (3) different technologies would be used for CO2 removal from syngas;30 and (4) nitrogen purges can be replaced with CO2-rich process gas purges in the commercial design. Table 2 summarizes the performance of the NREL and JMR catalysts, including results for the NREL catalyst used in Rentech-directed pilot runs. Near-complete tar conversions were achieved in both configurations. The removal of tars is important not only for increasing process yields but also for preventing fouling of downstream equipment and should be considered imperative. Benzene was not as completely reformed but should not cause equipment fouling and was not observed to have a negative impact on alcohol synthesis F

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Figure 4. Performance of pelletized catalyst operated in the CSTR at approximately 30% CO conversion. Data from both biomass-derived syngas and decomposed methanol syngas at H2:CO ratios of 3−3.5, without methanol cofeed. Data points are averages of 4−15 repeat data collected over periods of 12−24 h. Design case metrics are shown in Supporting Information Table SI-5.

3.4. Technoeconomic Analysis. As discussed above, the tar reforming and alcohol synthesis catalysts met or exceeded the key improvement criteria shown in Table 1. Only the tar reforming catalyst attrition loss metric was not met. This was estimated at 0.15% of the reactor-regenerator inventory per day based on phase II pilot-scale experiments by Rentech. This is close to the 0.1% target in the design case. Even so, the economic impact of the additional loss of 0.05% is an increase of 0.37 cents/L (1.4 cents/gallon) in the MESP. This 0.37 cent/L cost can be offset by an increase in methane conversion to 84%, which was demonstrated during both the NREL and Rentech pilot runs (Table 2). The overall conclusions from the comparison between the 2011 benchmark design4 and the pilot-scale experimental results are (1) in the reforming section, demonstrated methane conversions of >84% exceed the 80% design criterion, which offsets the cost of higher than expected catalyst loss rates; and (2) for the scaled-up alcohol synthesis catalyst performance, extrapolation was necessary to the design conditions to show that the benchmark performance was achievable (Figure 4) because the same operating conditions could not be achieved in the pilot plant. Even so, design targets for alcohol productivity and selectivity were met or exceeded at the bench scale, and no adverse impacts of using biomass syngas vs bottled syngas were observed, adding confidence to commercial model predictions. The experimental results are thus consistent with the modeled $0.54/L ($2.05/gallon) MESP in the design case. 3.4.1. Additional TEA Scenarios with Improved Catalysts. The benchmark design case4 represents a single route to cost competiveness and was derived from catalyst improvements (Table 1) deemed achievable via research. Further analysis of the experimental results showed that it was possible to meet, and even improve upon, the MESP target of $0.54/L with other process variations. As mentioned earlier, dry reforming tests by Rentech showed that the NREL catalyst is versatile and capable of reforming hydrocarbons and tar species by utilizing (1) steam added during biomass gasification, (2) combinations of steam and CO2, or (3) CO2 alone as the source of oxygen for reforming. This allows for production of biomass syngas with a wide range of H2:CO ratios. This capability can help improve the economics of the integrated process. Further, during the process of improving mixed alcohol synthesis catalysts, we observed that, in many cases, ethanol productivity and selectivity comparable to the design criteria

during prolonged continuous tests at the bench scale. Further details are presented elsewhere.5 Results from bench- and pilot-scale mixed alcohol synthesis are shown in Supporting Information Table SI-5. Corresponding 2011 design criteria4 are shown alongside for comparison. Catalyst productivity was generally low, primarily due to dilution of the syngas (inerts, CO2) and resulting lower partial pressures of H2 and CO. A higher-than-designed H2:CO ratio of 3.5 (compared to design of 1.5) further contributed to a lowering of CO activity. In general, catalyst productivity was higher in the FBR than in the CSTR,5 as one would expect due to the operating principle of each (an FBR is kinetically equivalent to a large number of CSTRs in series). Selectivity to ethanol was similar in both systems, ranging from 17% to 38%, with the CSTR producing a larger quantity of byproduct hydrocarbons. Figure 4 shows the relationship between CO partial pressure and alcohol production in the CSTR at approximately 30% CO conversion and suggests that increasing CO partial pressure to the design criteria of 6.16 MPa, along with a reduction in the H2:CO ratio to minimize total reactor pressure, may result in the corresponding ethanol selectivity and productivity. (See Table SI-5 in the Supporting Information for reactor conditions and metrics.) While this is an extrapolation of CO partial pressure to the design case conditions using less than optimal (CSTR) reactor data and H2:CO ratios, it is later shown that the design cost targets can be met by operating a FBR at a significantly lower pressure. FBR3 and FBR9 in Supporting Information Table SI-5 show results from two runs with comparable (but not identical) CO partial pressures, space velocities, and CO conversions using biomass syngas and bottled syngas, respectively. A bottled syngas mixture identical to the composition of one of the biomass-syngas cases could not be prepared because of the difficulty in mixing high concentrations of CO2 at high pressure. The FBR3 and FBR9 data show no apparent detrimental effects of biomass syngas, suggesting that data collected with bottled syngas are applicable to designs with biomass syngas. The slightly lower ethanol selectivity in FBR3 compared to FBR9 can be explained qualitatively as the result of lower CO partial pressure, a trend suggested by Figure 4. The syngas source agnosticism of the catalyst was also otherwise observed when the source was switched between biomass syngas and either methanol-derived (CSTR) or bottled (FBR) syngas during pilot operations. G

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(see Supporting Information Table SI-5) can be achieved at lower than designed operating pressures. This suggests the potential for improved process economics, as lower operating pressures will translate to lower compression costs and could include lower capital costs for reactor equipment. To that end, the FKM provided by The Dow Chemical Company was used to capture these catalyst improvements in the integrated process model by making the assumption that isothermal reactor performance data are valid in large-scale nonisothermal operation when incorporated into a combined kinetic and hydrodynamic model. Scaling parameters were adjusted in the FKM to achieve performance parity (to the extent possible) between The Dow Chemical Company’s model and NREL’s catalyst data. Scaled operational factors such as nonisothermal reactor tubes, use of catalyst pellets with inert binder and resulting increases in reactor volume and catalyst cost, the presence of inert species, and changes in compression duty were included in the analysis. Furthermore, upstream model assumptions were changed to reflect experimentally verified performance (86% methane conversion, 0.15% loss of reforming catalyst per day, 1.4 steam:carbon ratio, and 1.0 CO2:carbon ratio) based on Table 2. The results of this analysis are shown in Table 3, and a detailed cost breakdown is shown in Figure 5. The table and figure show that an MESP of $0.54/

Figure 5. MESP breakdown for a cost-competitive case based on an integrated model at lower alcohol synthesis pressure with respect to the design case.4 Alcohol synthesis reactor conditions are shown in Table 3.

L ($2.05/gallon) can be achieved at a much lower alcohol synthesis reactor pressure (12.7 MPa versus 20.7 MPa) by taking advantage of more active alcohol synthesis catalysts, based on pilot plant and bench reactor data. Note that the case shown in Figure 5 was not optimized for operating conditions (especially pressure) and recycles, which can significantly improve process economics, because model optimization would tip the close parity between the integrated model and experimental results. Because this lower operating pressure is considerably less risky, both operationally and financially, this new process configuration and model should be more fully explored. Two additional possibilities for cost reduction were modeled without significant extrapolations beyond this experimentally based point.33 These cases leveraged results from the Rentech-directed pilot runs, where it was shown that the NREL catalyst can produce syngas with lower H2:CO ratios via dry reforming and is capable of higher methane conversions (Table 2). The first variation involved lowering the H2:CO ratio to 1.0 (from 1.23 in the modeled case shown in Table 3). This was accomplished by using less steam and more recycled process gases containing CO2 in the gasifier for fluidization. The modeled MESP for this case was reduced 2.4% to $0.53/L ($2.00/gallon).33 The second case considered the impact of the Rentech-demonstrated higher methane conversion of 95% (compared to 86% used for Figure 5), which lowered the MESP further by 1.5% to $0.52/L ($1.97/ gallon).33 One of the key impacts of the lowering of the H2:CO ratio is the effective increase in the CO partial pressure at the same total pressure. This has a significant positive impact on the alcohol productivity as shown earlier (Figure 4). In addition, using less steam in the process also increases energy efficiency. The increase in methane conversion to 95% increases the proportion of the desired syngas components CO and H2, which helps the alcohol synthesis reactor performance. 3.5. Sustainability Metrics of the Conversion Process. Direct CO2, SO2, and NO2 emissions and water usage for the conversion process are presented in Table 4. The emissions are primarily from the char combustor for the gasifier and the regenerator for the reformer, which also burns process gases to maintain reformer temperatures. Emissions of CO2 (3.2 kg/L ethanol equivalent) and SO2 (0.13 g/L ethanol equivalent)

Table 3. Experimental Data and Conditioned Kinetic Model To Replicate Performance in the Integrated Plant Model for Mixed Alcohol Synthesis integrated model inlet temperature (°C) max reactor temperature (°C) inlet pressure (MPa) H2:CO inlet CO partial pressure (MPa) inlet methanol (mol %) methanol (mol % of CO + H2 + methanol) space velocity (h−1, std vol basis)c single pass CO conversion (%)c ethanol selectivity (% CO2-free) total alcohols selectivity (% CO2-free) ethanol productivity ((g/(kg of cat))/h) ethanol prod. with binder factor ((g/(kg of cat))/h) total alcohol productivity ((g/(kg of cat))/h) total alcohol prodictivity with binder factor ((g/(kg of cat))/h)

experimentala

300 327.3 12.72 1.23 4.28 2.25 2.92b 5325 33.1 57.2 70.0 188 156d

330 330 11.53 1.26 ± 0.01 4.23 ± 0.02 2.89 ± 0.08 2.96 ± 0.08 6390 30.4 ± 0.2 57.4 ± 0.8 69.9 ± 0.6 218 ± 4

238 197d

259 ± 8

a

Isothermal tubular bench-scale reactor shown in Figure 3b, results using bottled syngas. Average of measurements taken over 5 h, after achieving steady state at this condition. These inlet conditions were set after the catalyst was online for 600 h. This reactor−catalyst combination was online for over 2,500 h at various inlet conditions with no noticeable catalyst degradation. bOther key constituents include CO2 (14 mol %) and CH4 (8 mol %). cComplete parity could not be achieved for space velocity and CO conversion. However, they are directionally consistent for the two cases (i.e., conversion is higher at the lower space velocity). d76% of experimental total alcohols productivity to account for 24% binder in the catalyst. Corresponding ethanol productivity was 72% of the experimental result. These assumptions of 24% and 28% reduction in productivity for total alcohols and ethanol, respectively, were conservative based on comparative experimental results for the scaled-up catalyst with material from the constituent batches prior to pelletization.33 H

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Table 4. Direct CO2, NO2 (Only Biogenic Shown), and SO2 Emissions and Water Usage per Unit of Ethanol and Higher Alcohol Products (Liters Are Ethanol Equivalent Based on LHV) CO2 SO2 NO2

g/L

g/MJ

3200 0.13 3.6

150 6.0 × 10−3 0.17 L/MJ

L/L water usage

2.3

Table 5. Summary of sustainability metrics for the conversion process GHG emissions (g of CO2e/MJ) consumptive water use (m3/day) consumptive water use (L/L ethanol equivalenta) total fuel efficiency (L/(dry Mg)) carbon-to-fuel efficiency (C in the fuel/(C in biomass), %) net fossil energy consumption (kJ/MJ)

0.27 1247 1.6 392 32 5.5

a

Ethanol equivalent of ethanol and higher alcohol products based on LHV.

0.11

determined to be 1.6 L/L ethanol equivalent or 1,247 m3/ day. With the carbon-to-fuel efficiency of 32%, the product yield or total fuel efficiency for the current process is 392 L/ (dry Mg of woody biomass). 3.6. Future Work. Based on the lessons learned, we propose the following to further advance biomass gasificationbased technologies, with respect to the catalysts developed and future demonstrations: (1) provide flexibility during pilot-scale gasification to produce syngas with a range of H2:CO ratios for downstream use, (2) perform additional pilot-scale recirculating-regenerating reformer system tests, (3) explore efficient CO2 removal technologies applicable to biomass-derived syngas at the pilot scale, (4) include practically feasible recycle capabilities during pilot-scale experiments, (5) work to decrease the attrition rate of fluidized reforming catalyst via operational and materials changes, and (6) perform additional pilot testing of the alcohol synthesis catalyst using more optimal syngas compositions. In addition, there remain significant opportunities for further cost reduction via process optimization, and pilot-scale tests can be used for performance validation of model-predicted optima. The involvement of industrial partners (e.g., Rentech and The Dow Chemical Company) helps to facilitate pathways to commercialization, market factors permitting. There also exist possibilities in using the catalysts described here in other process configurations and in converting different feedstocks.

from combustion can typically be estimated with a high degree of certainty with a process simulation tool (Aspen Plus in our case). Carbon and sulfur content are almost completely oxidized during combustion to CO2 and SO2. On the other hand, NOx emissions depend on complicated chemical and physical processes during the combustion. The process model is configured to estimate NO2 emissions, accounting only for the biological nitrogen. The direct water usage is attributed to makeup water for the cooling towers and boiler system. Estimates for gasoline and corn ethanol are shown in Table 25 of the 2011 design report.4 Details of contributions to GHG emissions and fossil energy consumption at the conversion stage are presented in Table SI6 in the Supporting Information, which corresponds to the information in Table 24 in the 2011 design report.4 Because almost all of the technical improvement criteria in the design case were met, the results in Supporting Information Table SI-6 are identical to those from the design case, with the exception of the contribution from the higher tar reformer catalyst replenishment rate. GHG emissions associated with the conversion stage are estimated to be 5.8 g of CO2e/L of neat ethanol (or 0.27 g of CO2e/MJ). It is evident that there is minimal contribution to GHG emissions directly from the biorefinery. Direct CO2 emission from the conversion process (shown in Table 4) is biogenic (i.e., CO2 absorbed from the atmosphere and incorporated as biomass during the feedstock production phase). With its biomass origin, biogenic CO2 does not contribute to the increase of GHG in the atmosphere and is not considered in the Intergovernmental Panel on Climate Change global warming methodology.34 Hence, the contributions to GHG at the conversion stage are solely from the associated underlying processes (e.g., material inputs/outputs to and from the facility to support process operations). Supporting Information Table SI-6 shows the fossil energy consumption for the thermochemical process. The current biorefinery does not require any significant direct fossil energy input such as natural gas or fossil-derived electricity from the grid, because of the design decision to make the conceptual process energy-sufficient even at the expense of process yields. Table 5 summarizes the key sustainability metrics for the process. The U.S. Geological Survey defined consumptive water use as “water that is evaporated, transpired, incorporated into products or crops, consumed by humans or livestock, or otherwise removed from an immediate water environment.”35 Biorefinery consumptive use may occur through incorporation into the product and evaporation from cooling and heating processes. Based on this definition, the consumptive water use for the current conversion platform is equal to water for ash wetting plus makeup water to the cooling tower and boiler system, less cooling tower blowdown and wastewater to an offsite treatment facility. The consumptive water use is

4. CONCLUSION We have coupled experimental results with predictions from a technoeconomic model to show that cost-competitive ethanol is feasible in a mature plant using the catalysts developed through research. Key achievements of this work include the following: (a) A reforming catalyst was developed that is capable of significantly improving the quality of biomass-derived syngas. (b) Pilot-scale tests by Rentech showed that the performance targets established at the beginning of the research were exceeded. (c) It was demonstrated that steam and/or CO2 can be used for the reforming of syngas, which enables better customization of H2:CO ratios to optimize the use of specific synthesis catalysts. The reforming catalyst can maintain its performance in carbonizing conditions when used in a recirculatingregenerating dual bed configuration. (d) A JMR catalyst was demonstrated to be suitable for biomass-syngas cleanup in a polishing packed bed reactor. In the pilot demonstration, the packed bed was preceded by a steady-state non-regenerated catalyst in a fluidized bed, which aided the conversion process. (e) A mixed-alcohols synthesis catalyst exceeded the targeted criteria for performance and was demonstrated at scale. I

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(7) Yung, M. M.; Jablonski, W. S.; Magrini-Bair, K. A. Review of Catalytic Conditioning of Biomass-Derived Syngas. Energy Fuels 2009, 23, 1874. (8) Biomass Multi-Year Program Plan; U.S. Department of Energy Office of Energy Efficiency and Renewable Energy, April 2011. http:// www1.eere.energy.gov/biomass/pdfs/mypp_april_2011.pdf. (9) Gerber, M. A.; White, J. F.; Gray, M. J.; Thompson, B. L.; Stevens, D. J. Evaluation of Promoters for Rhodium-Based Catalysts for Mixed Alcohol Synthesis2009 Progress Report; Pacific Northwest National Laboratory: Richland, WA, USA, 2010. http://www.pnl.gov/ main/publications/external/technical_reports/PNNL-20115.pdf. (10) Hensley, J.; Gerber, M. 2011 DOE Biomass Program Review RD&D PresentationAdvanced Thermochemical Biofuels. Presented at the Office of Biomass Program’s Thermochemical Platform Peer Review, Feb. 17, 2011. Available at http://obpreview2011.govtools.us/ thermochem/. (11) Stevens, R. R. Process for Producing Alcohols from Synthesis Gas. U.S. Patent 4882360, Nov. 21, 1989. (12) Internal Revenue Service. How to Depreciate Property, Internal Revenue Service Publication 946; United States Department of the Treasury: Washington, DC, USA, 2009. http://www.irs.gov/pub/irspdf/p946.pdf. (13) SimaPro, v.7.3; Product Ecology Consultants: Amersfoort, The Netherlands, 2011. (14) Hsu, D.; Heath, G.; Wolfrum, E.; Mann, M. K.; Aden, A. Life Cycle Environmental Impacts of Selected U.S. Ethanol Production and Use Pathways in 2022. Environ. Sci. Technol. 2010, 5289. (15) Ecoinvent, v.2.2; Swiss Center for Life Cycle Inventories: Duebendorf, Switzerland, 2010. (16) LCI, U.S. Life-Cycle Inventory, v. 1.6.0. National Renewable Energy Laboratory: Golden, CO, USA, 2008. (17) 53 Federal Register at 3636, 1988. (18) U.S. Environmental Protection Agency. EPA Announces Decision for Regulation of Fuel and Fuel Additives: Modification to Octamix Waiver Regarding TXCeed, EPA-420-F-12−020; Office of Transportation and Air Quality, U.S. Environmental Protection Agency: Washington, DC, USA, June 2012. http://www.epa.gov/otaq/fuels/registrationfuels/ documents/420f12020.pdf. (19) Hensley, J. E.; Lovestead, T. M.; Christensen, E.; Dutta, A.; Bruno, T. J.; McCormick, R. Compositional Analysis and Advanced Distillation Curve for Mixed Alcohols Produced via Syngas on a KCoMoSx Catalyst. Energy Fuels 2013, 27 (6), 3246. (20) Wang, M.; Huo, H.; Arora, S. Methods of Dealing with Coproducts of Biofuels in Life-Cycle Analysis and Consequent Results within the US Context. Energy Policy 2011, 39, 5726. (21) Bain, R. L.; Dayton, D. C.; Carpenter, D. L.; Czernik, S. R.; Feik, C. J.; French, R. J.; Magrini-Bair, K.; Phillips, S. D. Evaluation of Catalyst Deactivation during Catalytic Steam Reforming of BiomassDerived Syngas. Ind. Eng. Chem. Res. 2005, 44, 7945 http://dx.doi.org/ 10.1021/ie050098w. (22) Carpenter, D. L.; Bain, R. L.; Davis, R. E.; Dutta, A.; Feik, C. J.; Gaston, K. R.; Jablonski, W.; Phillips, S. D.; Nimlos, M. R. Pilot-Scale Gasification of Corn Stover, Switchgrass, Wheat Straw, and Wood: 1. Parametric Study and Comparison with Literature. Ind. Eng. Chem. Res. 2010, 49, 1859 http://dx.doi.org/10.1021/ie900595m. (23) Magrini-Bair, K. A.; Jablonski, W. S.; Parent, Y. O.; Yung, M. M. Bench- and Pilot-Scale Studies of Reaction and Regeneration of Ni− Mg−K/Al2O3 for Catalytic Conditioning of Biomass-Derived Syngas. Top. Catal. 2012, 55, 209 http://dx.doi.org/10.1007/s11244-0129789-z. (24) Yung, M. M.; Magrini-Bair, K. A.; Parent, Y. O.; Carpenter, D. L.; Feik, C. J.; Gaston, K. R.; Pomeroy, M. D.; Phillips, S. D. Demonstration and Characterization of Ni/Mg/K/AD90 Used for Pilot-Scale Conditioning of Biomass-Derived Syngas. Catal. Lett. 2010, 134, 242 http://dx.doi.org/10.1007/s10562-009-0246-y. (25) Apanel, G.; Wright, H. A. System and Method for Dual Fluidized Bed Gasification. U.S. Patent US 8,241,523 B2, Aug. 14, 2012.

ASSOCIATED CONTENT

S Supporting Information *

Tables listing additional information on feedstock composition, reforming catalysts and reactor configurations, syngas compositions at the gasifier outlet, reformer outlet, and alcohol synthesis reactor inlet; key alcohol synthesis reactor conditions and metrics, greenhouse gases and fossil fuel metrics, figure showing pilot plant gasifier run timeline, and brief explanatory text with accompanying references. This material is available free of charge via the Internet at http://pubs.acs.org.

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AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS We are grateful for funding provided under U.S. Department of Energy (DOE) Contract DE-AC36-08-GO28308, and to The Dow Chemical Company for providing alcohol synthesis catalysts and for the use of their kinetic models under CRADA CRD-08-292. We thank Rentech for providing pilot test results for the NREL catalyst, Johnson Matthey for use of their reforming catalyst, Pacific Northwest National Laboratory for their work on an alternate alcohol synthesis catalyst, and Idaho National Laboratory for feedstock information. Also thanks to Paul Grabowski for continued support during the entire project, Michael Talmadge for TEA work, and communications support from Sara Havig, Stephanie Price, and Kristi Theis. We thank the entire NREL biomass thermochemical conversion platform staff for their tireless efforts during the challenging integrated pilot plant demonstration.

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REFERENCES

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K

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