How-to, Tips and Articles: Life-Cycle Analysis of Methanol Production From Coal and Natural Gas

Executive Summary

This analysis seeks to begin addressing that challenge by focusing on one primary product, methanol, which also serves as a readily-transportable intermediate to many other products including olefins, gasoline, and dimethyl ether (DME). This report presents the design configuration, performance summaries, and cost estimates of three cases which generate crude methanol.

Case 1: Coal-to-crude methanol without carbon capture and sequestration (CCS) (i.e., representing a building block to other derivatives, not chemical-grade methanol)
Case 2: Coal-to-crude methanol with CCS
Case 3: Natural gas-to-crude methanol with CCS

The required selling price (RSP) of crude methanol was estimated for each case and the results are listed in Exhibit ES-1. The RSP for the natural gas case is significantly lower than the coal cases due primarily to the lower capital requirements, which more than offset the higher feedstock cost.

Exhibit ES-1 Cost estimation results

The sensitivity of the RSP to the coal and natural gas feedstock prices is illustrated in Exhibit ES-2 and Exhibit ES-3, respectively. The prices used in this study are based on the National Energy Technology Laboratory (NETL) Quality Guidelines for Energy System Studies (QGESS) Recommended Fuel Prices. [1] Cases 1 and 2 use coal from the Powder River Basin (PRB) region in Montana, priced at $36.57/ton ($2.1351/MMBtu) including delivery to the Midwestern site, as the feedstock. Case 3 uses natural gas, priced at $6.13/MMBtu. Cases 1 and 2 also burn some natural gas to provide additional power required by the process.1 The expected RSPs for the case using natural gas as feedstock are below those for the coal-based cases because of the lower capital and fixed operating charges. On an energy content basis, the RSPs for all cases increase at approximately the same rate with increases in the feedstock prices due to the similarity between the feedstock requirements per unit of product. Cases 1 and 2 require 9.75 MMBtu of PRB coal per gallon of methanol generated and Case 3 requires 9.27 MMBtu of natural gas per gallon of methanol generated.

Given the higher cost per MMBtu of natural gas relative to coal and the similar energy input requirements of the conversion process, a 100 percent increase in natural gas prices leads to a 50 to 55 percent increase in RSP, for commercial and loan guarantee financing respectively, versus a 100 percent increase in coal prices leading to a 13 to 16 percent increase in RSP consistent with the natural gas–based methanol production being the less capital intense but more operating margin dependent technology choice. Hence, the natural gas route is more exposed to feedstock price volatility.

1 A natural gas combined cycle plant was used to generate additional power in the coal gasification cases to make the plant approximately power neutral while maximizing the production of crude methanol (i.e., no syngas is diverted for power production).

Exhibit ES-2 Sensitivity of the required selling price to the coal price

Based on Methanol density of 332.6 gal/tonne or 301.73 gal/ton

Exhibit ES-3 Sensitivity of the required selling price to the natural gas price

Based on Methanol density of 332.6 gal/tonne or 301.73 gal/ton

1. Overview

This report presents the design configuration, performance, and cost of the three crude methanol cases.

1.1 Background

Gasification of coal, in addition to generating syngas for power production, has the potential to produce a diverse array of high-value products. It is a challenge to understand the optimal use of this domestic coal resource amidst the potential technology options, product slates (including co- production of power), and competing feedstocks (natural gas, petroleum). This analysis seeks to begin to address that challenge by focusing on one primary product, methanol, which also serves as a readily-transportable intermediate to many other products including olefins, gasoline, and di- methyl ether (DME). The information from this analysis and previous studies can be used, along with available data on the production of methanol derivatives, in developing a framework for evaluating and optimizing the utilization of domestic coal and natural gas resources.

1.2 Case Descriptions

For each of the cases listed in Exhibit 1-1, and also described below, a system study was completed in accordance with the May 2010 version of the report, “Scope and Reporting Requirements for NETL System Studies.” [2] The cases in this report are limited to the following:

Coal-to-crude methanol with and without carbon capture and sequestration (CCS) (i.e., representing a building block to other derivatives, not chemical-grade methanol)
Natural gas-to-crude methanol with CCS

All three cases are sized to produce approximately 10,000 metric tons of methanol per day (3,326,000 gal/day based on 332.6 gal/tonne [3]). This plant size is considered large scale but typical for the current design of new plants. [4] This report presents the results of the performance modeling and cost estimating for the crude methanol product cases (cases 1, 2, and 3).

Exhibit 1-1 Case descriptions

2. Design Criteria

2.1 Site Description

All plants in this study are assumed to be located at a generic plant site in Midwestern U.S., with ambient conditions and site characteristics as presented in Exhibit 2-1 and Exhibit 2-2. These assumptions are identical to the National Energy Technology Laboratory (NETL) baseline studies. [6,7] The ambient conditions are the same as International Organization for Standardization (ISO) conditions.

Exhibit 2-1 Site ambient conditions

Exhibit 2-2 Site characteristics

As assumed for gasification-based cases in the NETL baseline studies, the land area for these cases (including the natural gas reforming case) is estimated as 30 acres required for the plant proper with the balance providing a buffer of approximately 0.25 miles to the fence line. [6] The extra land could also provide for a rail loop, if required.

In all cases, it was assumed that the steam turbine is enclosed in a turbine building. The gasifiers, reformers, methanol synthesis reactors, and the combustion turbines are not enclosed.

Allowances for normal conditions and construction are included in the cost estimates. The following design parameters are considered site-specific, and are not quantified for this study. Costs associated with the site specific parameters can have significant impact on capital cost estimates.

Flood plain considerations
Existing soil/site conditions
Water discharges and reuse
Rainfall/snowfall criteria
Seismic design
Buildings/enclosures
Local code height requirements
Noise regulations – Impact on site and surrounding area

2.2 Design Feedstocks

2.2.1 Natural Gas Characteristics

The natural gas composition used in this analysis, representative of natural gas after going through standard midstream processing, is presented in Exhibit 2-3. [8]

Exhibit 2-3 Natural gas composition

2.2.2 Coal Characteristics

The coal properties from National Energy Technology Laboratory’s (NETL) Quality Guidelines for Energy System Studies (QGESS): Specifications for Selected Feedstocks, for a subbituminous Powder River Basin (PRB) coal from Montana are shown in Exhibit 2-4. [8]

Exhibit 2-4 Montana Rosebud PRB, Area D, Western Energy Co. Mine, subbituminous design coal analysis

2.3 Environmental Requirements

The environmental limits presented in this section refer to the gasification/power cycle only, because the environmental requirements for the methanol plant are considered beyond the scope of the study.

The environmental targets for the study were considered on a technology- and fuel-specific basis. Since all the cases are located at a greenfield site, permitting a new plant would involve the New Source Review (NSR) permitting process. The NSR process requires installation of emission control technology, meeting either the best available control technology (BACT) determinations for new sources located in areas meeting ambient air quality standards (attainment areas), or the lowest achievable emission rate (LAER) technology for sources located in areas that do not meet ambient air quality standards (non-attainment areas). This study uses the BACT guidelines, which are summarized in Exhibit 2-5.

Exhibit 2-5 BACT environmental design basis

Selection of the process technology accounts for obtaining minimum sulfur content in the syngas and final product.

The following regulatory assumptions are used for assessing environmental control technologies:

NOx Emission Reduction Credits (ERC) and allowances are not available for the project emission requirements when located in the ozone attainment area
Solid waste disposal is either offsite at a fixed $/ton fee or is classified as a byproduct for reuse, claiming no net revenue ($/ton) or cost
Raw water is available to meet technology needs
Wastewater discharge will meet effluent guidelines rather than water quality standards for this screening

Based on published vendor literature, it was assumed that low NOx burners (LNB) and nitrogen dilution can meet 15 ppmv (dry) limits on NOx emissions from the combustion turbines at 15 percent O2. This value was used for all cases.

The acid gas removal (AGR) process in the coal gasification cases must have a sulfur capture efficiency of about 99.7 percent to reach the environmental target and to also produce sulfur-free syngas to avoid poisoning of catalysts in methanol synthesis process. Vendor data on AGR processes used indicate that this level of sulfur removal is possible.

In the CO2 capture cases, the two-stage Rectisol process in the coal gasification cases was designed for 90 percent plant CO2 removal. The amine capture process for the natural gas case and the natural gas combined cycle (NGCC) system in the coal fired gasification case with capture was also designed for 90 percent plant CO2 removal. The 90 percent capture design for both CO2 point sources results in an overall 90 percent capture rate for the plant.

For the coal feedstock cases, most of the coal ash is removed from the gasifier as slag. Any ash that remains entrained in the syngas is captured in the downstream equipment, including the syngas scrubber and a cyclone followed by either ceramic or metallic candle filters. The environmental target can be achieved with the combination of particulate control devices.

Mercury capture is not required for the case with processed natural gas as the starting feedstock. For the coal feedstock cases, however, the environmental target for mercury capture is greater than 90 percent capture. Eastman Chemical’s operating experience at its coal-to-methanol plant in Kingsport, Tennessee has shown mercury removal efficiency of 95 percent. This value was used as the assumed performance level for this study. Sulfur-impregnated activated carbon is used by Eastman as the adsorbent in the packed beds operating at 30°C (86°F) and 6.2 MPa (900 psig). Mercury removal between 90 and 95 percent has been reported with a bed life of 18 to 24 months. Removal efficiencies may be even higher, but at 95 percent the measurement precision limit is reached. Mercury removals of greater than 99 percent can be achieved by the use of dual beds, i.e., two beds in series. However, this study assumes that the use of sulfur-impregnated carbon in a single carbon bed achieves 95 percent reduction of mercury emissions, which meets the environmental target limits and the New Source Performance Standards limit in all cases.

For the cases that feature carbon capture, CO2 transport and storage (T&S) was modeled based on the specifications in the NETL QGESS: Estimating Carbon Dioxide Transport and Storage Costs. [9] The CO2 is supplied to a 100 km (62 mi) pipeline at the plant fence line at a pressure of 15.3 MPa (2,215 psia). The CO2 product gas composition varies in the cases presented, but is expected to meet the specification described in Exhibit 2-6. [10] A glycol dryer located near the mid-point of the compression train is used to meet the moisture specification.

Exhibit 2-6 CO2 pipeline specification

2.4 Balance of Plant Requirements

Assumed balance of plant requirements are listed in Exhibit 2-7.

Exhibit 2-7 Balance of plant design requirements

2.5 Crude Methanol – Deviations from the C-MTG Study

This study is based on the coal-to-methanol-to-gasoline (C-MTG) cases in the Baseline Analysis of Subbituminous Coal and Biomass to Gasoline (Indirect Liquefaction by Methanol Synthesis)[11] with the methanol purification and gasoline synthesis processes removed. A natural gas to methanol case was added for comparison. The cases covered in this report simply consider a relaxation of the product purity specifications typically associated with solvent-grade (Grade A) and chemical-grade (Grade AA) methanol. Methanol is the raw material for many chemical products such as formaldehyde, dimethyl terephthalate, methylamines, methyl halides, methyl methacrylate, and acetic acid. Most importantly, methanol provides the foundation of the methanol-to-hydrocarbon platform which has commercially proven processes to convert methanol into gasoline and olefins (such as ethylene, the building block of the petrochemical industry).

Since the methanol-to-hydrocarbon mechanism is fairly forgiving with respect to oxygenated hydrocarbons and limited amounts of water, cases were considered where methanol purification (i.e. distillation columns) was removed. This is consistent with the methanol being fed to the commercial methanol-to-gasoline (MTG) process in New Zealand following mild equilibration over an acidic alumina bed (to bring the methanol-DME-water into equilibrium)2. Skipping this step (removal of methanol purification) allows for potential savings in capital investment, so it was eliminated from the cases in this report.

Use of natural gas instead of coal significantly reduces the amount of capital investment. Besides elimination of coal handling functions, the syngas production function of the Shell gasifier in the coal cases is replaced by steam methane reforming, partial oxidation, or autothermal reforming for natural-gas-based applications. Based on experience, autothermal reforming is considered the proper technology choice for any reasonable-scale, natural-gas-fed methanol plant. The autothermal reformer produces a H2/CO ratio of two, as preferred for methanol synthesis, similar to the case where coal is fed to a Shell gasifier and a water-gas-shift process.

The methanol synthesis is essentially the same following the production of syngas of an appropriate stoichiometric number.3 The designation crude merely conveys not purifying the methanol beyond the requirements for (essentially) immediate consumption in a methanol-to- products (gasoline or similar) facility. This study did investigate high level adjustments to the reactor that could lead to reduced capital expenditure, such as lowering the pressure, but the corresponding impact on process yields made the decision to not investigate such options further obvious. Long-range opportunities to make coal-based systems more competitive, such as revisiting non-copper-based catalytic systems, were not analyzed. Non-copper-based catalytic systems (i.e., zinc oxide and chromium oxide systems) are more tolerant to common coal contaminants and, therefore, may require less capital investment. Technology scoping of these opportunities was beyond the scope of this study.

2 An allocation of the cost of the equilibration bed should be included in evaluating this option; however, this cost is outside the scope of this study and would be part of a more complex integrated facility cost optimization if implemented.

3 The stoichiometric number or “S” is defined as [moles of hydrogen – moles of carbon dioxide]/[moles of carbon monoxide + moles of carbon dioxide].

3. Crude Methanol Model Performance Results

3.1 Technology Background and Readiness Assumptions

The current status of the major technology systems included in this study is highlighted below.

Coal gasification is a mature technology that has been deployed throughout the world, including in conjunction with methanol synthesis plants.
Natural gas reforming is a mature technology that has been deployed throughout the world, including in conjunction with methanol synthesis plants.
Reductions in the price of natural gas and increases in available quantities have made natural gas the preeminent feedstock for commercial methanol production since the mid- twentieth century
Methanol synthesis from syngas is a commercial process first used in 1923. The current catalytic process has been in widespread use since the late 1960s and is considered to be a proven and robust technology. The current catalytic process is optimized for the dominant natural gas feedstocks rather than coal, but is proven to be economically viable for either feedstock. [12]
Elemental sulfur is generated by the well-established Claus process for the coal cases. The sulfur is collected with a purity of 99 percent.

In Cases 2 and 3, carbon dioxide is captured, treated, and compressed to meet the CO2 pipeline specifications. CO2 specifications are negotiated and determined on a site-by-site basis. Drying was included in the CO2 capture and compression cost estimations, but additional cleanup that could be needed to meet more rigorous pipeline specifications was considered to be beyond the scope of this study. [13]

3.2 Modeled Performance Summary

The three cases analyzed in this report were modeled using Aspen Plus® (Aspen). Performance and process estimates were based upon published reports, presentations, information obtained from vendors, cost and performance data from design/build utility projects, and/or best engineering judgment.

The mass balances for the two coal-based cases are summarized in Exhibit 3-1. All three of the cases are sized to produce approximately 10,000 metric tons of methanol per day (3,326,000 gal/day based on 332.6 gal/tonne [14]). This plant size is considered large scale but typical for the current design of new plants. [4] The Case 2 mass balance is the same as Case 1, except that the CO2 recovered from the Rectisol unit is not vented, but rather compressed and sequestered. The mass balance for the natural-gas-based case is shown in Exhibit 3-2.

A comparison of the syngas feed to the methanol synthesis system as well as the raw methanol product stream is shown in Exhibit 3-3. The methanol synthesis systems for the three cases are virtually identical.

A summary of the auxiliary power requirements for the crude methanol plants in each of the three cases is given in Exhibit 3-4 and Exhibit 3-5. Some of the auxiliary power requirements for each case are met by adding a heat recovery steam generation (HRSG) system to power a steam turbine. Small gas turbines are combined with the HRSGs in each case to supply additional power, and the performance estimates for the additional gas turbine plants are shown in Exhibit 3-6.

In the natural gas case, the combustion turbine is fueled by the process exhaust gases solely to recover energy value from this stream. Additional natural gas is not necessary to supply sufficient power because the combustion of the tail gases generates excess power beyond auxiliary power load requirements. The excess power can be sold to the grid; however, the sale may be at a steep discount as entities that are negotiating a power purchase agreement will know the power production is an inherent by-product of core methanol production operations.

Consequently, the actual achieved transfer price for excess power will be a significant risk in natural gas feedstock projects and would be highly project dependent. The sensitivity of the results to the electricity selling price received for the excess power is included in the results section of this report and indicates that the impact on the methanol RSP is fairly small. High excess power is endemic to the production of methanol from natural gas because there is considerable heat recovery from the exothermic synthesis process, the production of significant amounts of tail gas without use such as coal drying, and lower energy requirements than coal- based processes.

In the coal cases, the exhaust gases are used in coal drying; so the combustion turbine is fueled by natural gas to generate additional auxiliary power while maximizing the production of crude methanol (i.e., no syngas is diverted for power production). The NGCCs are sized to meet the auxiliary power load requirements that are beyond the capabilities of the HRSG alone and eliminate the need for power from the local electrical grid for the coal cases. The performance estimates for the addition of the gas turbines is shown in Exhibit 3-6

Exhibit 3-1 Mass balance summary of coal to crude methanol plant

Exhibit 3-2 Mass balance summary of natural gas to crude methanol plant

Exhibit 3-3 Stream compositions to and from methanol synthesis process

Exhibit 3-4 Auxiliary power load summary of CTM plants

Exhibit 3-5 Auxiliary power load summary of NGTM plant

Exhibit 3-6 Performance summary of NGCC plants

3.3 Process Descriptions

This section describes the crude methanol production process, divided into the functions (as “blocks”) of each process step. The section includes process flow diagrams (PFD) and stream tables as referred to in the text. The diagrams and tables are grouped together at the end of this section.

3.3.1 General Process Descriptions

The conversion of coal-to-methanol is a two-step process: first conversion of the coal to the appropriate quality syngas via gasification and applying the water gas shift reaction, and second, catalytically converting the syngas to methanol.

For the coal-to-methanol (CTM) cases 1 and 2, the plant configuration is illustrated in Exhibit 3-7. Syngas is generated from the gasification of PRB coal in a high-pressure, oxygen- blown Shell quench-type gasifier. The high temperature entrained-bed gasifier uses a partial water quench and syngas cooler to cool the hot syngas stream and generate steam for the water gas shift reactors and power generation. Crude raw syngas (post quench) from the gasification unit is scrubbed and split into two streams. The first stream is fed to a sour water gas shift reactor (WGSR) to increase the hydrogen content so that a H2/CO molar ratio of 2:1 in the feed stream to the synthesis reactor can be achieved, and the second stream bypasses the WGSR. The streams are combined to achieve the desired composition. This partial bypass mode of operation allows the shift reactor to operate at a higher conversion ratio resulting in a smaller size. The syngas is cooled in a low-temperature heat-recovery system and then cleaned of mercury, sulfur, and CO2 in preparation for methanol synthesis. Clean syngas is fed into a fixed-bed process to generate methanol.

Exhibit 3-7 Simplified process flow diagram of coal to crude methanol

For the natural gas to methanol (NGTM) case, the plant configuration is illustrated in Exhibit 3-8. Syngas is generated by combining natural gas, steam, and oxygen in an autothermal reformer (ATR). The reformed syngas is cooled to recover heat at a useful temperature and fed into a fixed-bed process to generate methanol. For this study, no additional cleanup of the syngas was included based on the assumption that contaminants were removed from the natural gas feedstock by the supplier. Site specific circumstances may require sulfur polishing and other cleanup processes in a more detailed design basis, but they were considered beyond the scope and accuracy of this study.

Exhibit 3-8 Simplified process flow diagram of natural gas to crude methanol

3.3.2 Air Separation Unit

All cases include an air separation unit (ASU) for generating oxygen. The ASU is a conventional, cryogenic, pumped liquid oxygen (LOX) unit that provides oxygen for the gasification and reforming processes, as well as nitrogen for ancillary equipment. The ASU is designed to produce 95 mole percent O2 for use in the gasifier and Claus plant in the coal cases and for use in the reformer in the natural gas case. The air compressor is powered by an electric motor. Nitrogen is recovered and used as a diluent for coal drying in the coal cases and vented in the natural gas case.4

The ASU process for the coal fed cases is shown in Exhibit 3-11 with the gasification, quench, and dry solid removal processes. Cases 1 and 2 share a common PFD. The mass balances of this process for the two cases are presented by Exhibit 3-12. Since Case 2 uses the same specifications as Case 1, they both share a common mass balance in the exhibit. The stream numbers on the PFD correspond to the stream numbers in the mass balance tables. The ASU for the natural gas case was not modeled, but the performance was estimated from similar ASU systems.

4 In the natural gas cases, no disposition was assigned to the nitrogen from the ASU and no potential credit from its sale was applied to the process economics.

3.3.3 Coal Gasification Syngas Process Description

Cases 1 and 2 use coal gasification as the initial processing step. The gasifier block utilizes coal, oxygen, and steam to produce raw syngas. The following units in this subsection are included in this block.

3.3.3.1 Coal Milling, Grinding, and Drying

The coal drying PFD is shown in Exhibit 3-9. Case 1 and Case 2 share a common PFD. The mass balances of the coal drying process are presented in Exhibit 3-10. Since Case 2 uses the same specifications as Case 1, they both share a common mass balance in the exhibit. The stream numbers on the PFD correspond to the stream numbers in the mass balance tables.

The Shell process uses a dry feed system that is sensitive to the coal moisture content. Coal moisture consists of two parts: surface moisture and inherent moisture. For coal to flow smoothly through the lock hoppers, the surface moisture must be removed. The PRB coal used in this study contains 25.77 percent total moisture on an as-received basis. It was assumed that the coal must be dried to 6 percent moisture to allow for smooth flow through the dry feed system. The coal is simultaneously crushed and dried in the coal mill, then delivered to a surge hopper.

The drying heat is provided by burning the tail-gas from the Claus plant and the flash and purge gas from the methanol synthesis process in an incinerator. The hot incinerator flue gas mixes with N2 from the ASU and exhaust is recycled from the mill to maintain a drying gas temperature of less than 500°F with oxygen content lower than 8 vol%. The dried coal is drawn from the surge hoppers and fed through a pressurization lock hopper system to a dense-phase pneumatic conveyor, which uses CO2 to convey the coal to the gasifier. Using CO2 rather than N2 as the transport gas has the benefit that CO2 is removed in the downstream acid gas removal (Rectisol) process, minimizing the buildup of inert species in the methanol generation recycle loop, thus reducing the size of the equipment needed. For this study, it was assumed that there was no impact of hot, concentrated CO2 in the presence of moisture on standard materials of construction.

3.3.3.2 Gasification, Gas Quench, and Water Quench

The gasification and quench processes for the coal fed cases are shown in Exhibit 3-11. Cases 1 and 2 share a common PFD. The mass balances of this process for the two cases are presented by Exhibit 3-12. Since Case 2 uses the same specifications as Case 1, they both share a common mass balance in the exhibit. The stream numbers on the PFD correspond to the stream numbers in the mass balance tables.

Syngas is generated from the gasification of PRB coal in a high-pressure, oxygen-blown Shell quench-type gasifier. The high temperature entrained-bed gasifier uses a partial water quench and syngas cooler to cool the hot syngas stream and generate steam for the water gas shift reactors and power generation.

The coal feedstock is gasified in the presence of O2 and superheated process steam. The O2 requirement (O2 required/dried feedstock) depends on the moisture content of the dried feedstock feeding the Shell gasifier. The coal is gasified at 2,550°F and 650 psia in a membrane wall reactor installed inside a pressure vessel, forming syngas, fly ash, and slag. The reactor design includes entrained flow, high temperature, recycled ash particulates, and slagging gasification and achieves carbon conversion greater than 99 percent.

The syngas leaving the gasifier is quenched to 850°F before entering a cyclone system for initial particulate removal described in the next section. The majority of the slag leaves the gasifier via the bottom as molten slag and is quenched and scattered to small glassy granulates in the slag (water) bath. In addition to the syngas and slag, the gasifier produces medium pressure saturated steam. Steam is produced from water that is circulated over the membrane wall of the gasifier to remove the heat of reaction and maintain an operating temperature of 2,550°F in the gasifier.

For this study, the design size requires eight operating trains with one spare train for a total of nine gasifiers. The facility contains one spare gasifier train to allow operation at a 90 percent capacity factor and to generally improve availability. The spare gasifier train feeds into the same gas clean-up trains as the other gasifier trains so that start-up/operation is transparent to downstream processes. The gasifier costs were scaled from previous studies which assumed two trains plus one spare train to achieve a high availability. [6,7,15]

3.3.3.3 Dry Solids Removal and Wet Scrubbing

The dry solid removal processes are shown in Exhibit 3-11 with the gasification and quench processes. Cases 1 and 2 share a common PFD. The mass balances of this process for the two cases are presented by Exhibit 3-12. Since Case 2 uses the same specifications as Case 1, they both share a common mass balance in the exhibit. The stream numbers on the PFD correspond to the stream numbers in the mass balance tables.

After passing through the water quench system, the syngas passes through a cyclone and a raw gas candle filter where a majority of the fine particles are removed and returned to the gasifier with the coal feedstock. Fines produced by the gasification system are recirculated to extinction. Final dust removal is achieved in the wet scrubbing section, to lower the dust content of the syngas to <1 mg/Nm³, and to lower its halide content to <1 ppmv. The wet scrubbing system consists of a Venturi scrubber followed by a packed bed wash column. A three percent by weight slurry bleed is fed to the primary wastewater treatment. Syngas is water saturated and leaves the wet scrubbing system at a temperature of 425°F.

3.3.3.4 Water Gas Shift

The water gas shift process is shown in Exhibit 3-13, where both coal cases share a common PFD. The material balances for the two cases are presented in Exhibit 3-14. Since Case 2 uses the same specifications as Case 1, they both share a common mass balance in the exhibit. The stream numbers on the PFD correspond to the stream numbers in the mass balance tables.

Coal-derived syngas from the wet scrubber enters the sour shift and cooling section. In order to achieve a 2:1 ratio of H2 to CO in the final syngas, approximately 55 to 60 percent of the coal- derived syngas is shifted. The syngas to be shifted is heated in a countercurrent exchanger by the effluent of the first shift reactor to a temperature of 530°F. This syngas passes through the first stage shift reactor where the shift reaction exothermically converts CO and H2O into H2 and CO2, leading to an outlet syngas temperature of 940°F. The syngas is then cooled to 495°F by a heat recovery steam generating exchanger followed by the counter current exchanger with the first stage feed stream.

The syngas/steam mixture passes through the second stage shift reactor where the shift reaction converts additional CO and H2O into H2 and CO2, with a resulting second-stage shift reactor outlet syngas temperature of approximately 600°F. After cooling, the shifted syngas from the second-stage shift reactor outlet mixes with the bypass syngas and is sent to low-temperature gas cooling before being sent to the downstream Rectisol unit for removal of sulfur and CO2.

3.3.3.5 Low-Temperature Gas Cooling

The low-temperature gas cooling process is shown in Exhibit 3-15 where both coal cases share a common PFD. The material balance for this process is presented in Exhibit 3-16. Since Case 2 uses the same specifications as Case 1, they both share a common mass balance in the exhibit.

The stream numbers on the PFD correspond to the stream numbers in the mass balance tables.

Syngas is cooled in a number of steps to recover heat at useful temperatures. As the shifted syngas is cooled, process condensate, feed water, and a mix of clarified water and raw water are being heated.

The condensate from the air cooler and trim cooler knock-out drums absorbs nearly all of the NH3 from the syngas. This condensate is mixed and sent to the process condensate stripper where low pressure steam is used to strip NH3 and other absorbed gases from the condensate. The stripped gases are sent to the sulfur recovery unit to be treated with other sour gas streams.

The stripped condensate mixes with process condensate separated from the syngas. The mixed temperature is 325 °F. The process condensate is then pumped and heated to 390°F before being fed into the wet scrubbers.

3.3.3.6 Mercury Removal

The mercury removal process is shown in Exhibit 3-15 where both coal cases share a common PFD. The material balance for this process is presented by Exhibit 3-16. Since Case 2 uses the same specifications as Case 1, they both share a common mass balance in the exhibit. The stream numbers on the PFD correspond to the stream numbers in the mass balance tables.

Mercury removal is achieved by a packed bed of sulfur-impregnated activated carbon. The low- temperature syngas must be heated to 105°F prior to feeding it to the mercury-removal bed. The packed carbon bed vessels located upstream of the sulfur recovery unit with 20-second superficial gas residence time would achieve more than 90 percent removal of mercury (based on mercury content in the gasifier feedstock) in addition to removal of some portion of other volatile heavy metals such as arsenic. Mercury-removal systems using sulfur-impregnated activated carbon downstream of a coal gasifier have a reported bed life of 18 to 24 months, and usually replacement is required due to fouling of the bed rather than mercury saturation.

3.3.3.7 Acid Gas Removal and CO2 Compression

The acid gas removal and CO2 compression processes are shown in Exhibit 3-17, where Case 1 and Case 2 are shown in separate PFDs. Exhibit 3-19 represents the material balances of this process for the cases. Case 1 and Case 2 are shown in separate tables, as the recovered CO2 in Case 1 is not compressed and sequestered. The stream numbers on the PFD correspond to the stream numbers in the mass balance tables.

A feature of this plant configuration is that H2S and CO2 are removed within the same process, via the Rectisol unit. The Rectisol acid gas removal (AGR) process was specified primarily because the methanol synthesis catalyst requires an H2S level below 100 ppbv in order to maintain an adequate catalyst lifetime.

The Rectisol process uses chilled methanol as a solvent. Because of high vapor pressure of methanol, the process is operated at a temperature in the range of -30 to -100 °F. There are many possible process configurations for the Rectisol process depending on process requirements, product specifications, and scalability. In this study, the methanol solvent contacting the feed gas in the first stage of the absorber is stripped in two stages of flashing via pressure reduction.

The acid gas leaving the first stage solvent regenerator is suitable for processing in a Claus plant. The regenerated solvent is virtually free of sulfur compounds but contains some CO2. The second stage of absorption then removes the remaining CO2 present. The rich solvent from the bottom of the second stage of the absorber is stripped in a steam-heated regenerator and returned to the top of the absorption column after cooling and refrigeration.

In Case 1, sufficient CO2 for use in transporting the coal is compressed as needed; the remaining recovered CO2 is vented. In Case 2, the low-pressure CO2 stream recovered from the Rectisol unit is compressed to 2,200 psig in a multiple-stage, intercooled compressor to supercritical conditions, which is then ready for pipeline transport.

3.3.3.8 Sulfur Recovery Unit

The sulfur recovery unit is shown in Exhibit 3-18, where both coal cases share a common PFD. Exhibit 3-19 represents the material balances. Case 1 and Case 2 material balances are shown separately. The stream numbers on the PFD correspond to the stream numbers in the mass balance tables.

The purpose of the sulfur recovery unit is to treat the acid gas from the Rectisol unit and sour gas streams from the sour water strippers to make an effluent gas acceptable for venting to the atmosphere or burning.

Currently, the Claus process remains the mainstay for sulfur recovery. Conventional three-stage Claus plants, with indirect reheat and feeds with a high H2S content, can approach greater than 98 percent sulfur recovery.

The Claus process converts H2S to elemental sulfur via the following reactions:

H2S + 3/2 O2 → H2O + SO2

2H2S + SO2 → 2H2O + 3S

The second reaction, the Claus reaction, is equilibrium limited. The overall reaction is:

3H2S + 3/2 O2 → 3H2O + 3S

The sulfur in the vapor phase exists as S2, S6, and S8 molecular species, with the S2 predominant at higher temperatures, and S8 predominant at lower temperatures.

In this process, one-third of the H2S is burned in the furnace with oxygen from the ASU to give sufficient SO2 to react with the remaining H2S. Since these reactions are highly exothermic, a waste heat boiler that recovers high-pressure steam follows the furnace. Sulfur is recovered in a condenser that follows the high-pressure steam recovery section. Low-pressure steam is raised in the condenser. The tail gas from the first condenser then goes to three catalytic conversion stages, where the remaining sulfur is recovered via the Claus reaction. Each catalytic stage consists of gas preheat, a catalytic reactor, and a sulfur condenser. The liquid sulfur goes to the sulfur pit, while the tail gas proceeds to the incinerator for coal drying.

3.3.4 Natural Gas / Steam Reforming

In case 3, a methane reformer is used to convert natural gas to a syngas suitable for methanol synthesis. Natural gas is combined with oxygen in the autothermal reformer (ATR). The ATR represents a process intensification of syngas production where partial oxidation of the feedstock provides the energy to drive the endothermic reforming of the feedstock to syngas. The reforming is accomplished through contacting the reaction mixture with a nickel supported on alumina catalyst. The ATR is fed 95 percent pure oxygen from the ASU; a purified oxygen feed was chosen to minimize the amount of inert gases introduced into the methanol synthesis loop.

The ATR operates at 355.3 psia and 1,935ºF (24.5 bar and 1,057°C). Steam and oxygen feeds were adjusted to the reformer to obtain a H2/CO ratio of approximately 2/1 and to provide enough steam to mitigate the risk of excessive coking. The reformed syngas is cooled to recover heat at a useful temperature. As the syngas is cooled, process condensate, feed water, and raw water are being heated. For this study, it was assumed that no additional cleanup of the syngas was necessary before the production of crude methanol. Sulfur polishing and other cleanup processes may be required in a more detailed design, but they were considered beyond the scope and accuracy of this study.

3.3.5 Methanol Reactor and Synthesis Loop - All Cases

The Methanol Synthesis Loop PFD is shown in Exhibit 3-22, where all the cases share a common PFD and the material balances for the cases are shown in Exhibit 3-23. Since Case 2 uses the same specifications as Case 1, they both share a common mass balance in the exhibit while the natural gas-case balance is shown on the second page. The stream numbers on the PFD correspond to the stream numbers in the mass balance tables.

There are two routes for the production of methanol: vapor phase and liquid phase. A vapor- phase methanol process was chosen over a liquid-phase methanol process due to the breadth of operating experience with vapor-phase production units and the lack of commercial operating experience with liquid-phase methanol production. [16] However, the advent of mega- methanol projects in stranded gas fields may rapidly provide the experience curve to justify revisiting this design choice.

The methanol reactor converts hydrogen and carbon monoxide to methanol. The reactor is a catalytic packed-bed reactor. The primary side reactions produce ethanol, propanol, and formaldehyde. Acetone and acetaldehyde are also common impurities in the methanol product and are captured in this analysis.

CO2-lean syngas with a H2/CO ratio of 2:1 from the AGR process or from the natural gas reformer is compressed from 490 psia to the synthesis loop operating pressure of 755 psia in the syngas compressor. The compressed syngas is mixed with the recycled gas, heated to 400ºF, and routed to the methanol reactor. The reactor is steam cooled to facilitate near isothermal operation at 475ºF and 735 psia. In-line blowers, coolers, and knock-out drums are used within the synthesis loop to maintain pressure and remove crude methanol.

Because the methanol synthesis reaction is equilibrium limited, in order to promote continued production of products, methanol reactor effluent is cooled to condense out the product crude methanol that is removed in a phase separator. Ninety-six percent of the separated gas is compressed to reactor pressure and readmitted along with fresh syngas to the methanol reactor. This recycling elevates the overall conversion of carbon, overcoming the low per-pass conversion of CO (60 percent). A small purge-gas stream (approximately four percent) is taken to limit the build-up of inert gas species. In the coal cases, the exhaust gas is burned in the coal- drying process where some heat is recovered as steam for power generation. In the natural gas case the exhaust gas is combusted in the combustion turbine for power generation.

3.3.6 Heat Recovery and Power Generation

The power cycle heat and material balance diagrams for all cases are shown in Exhibit 3-24. Both the gasification and the methanol synthesis processes generate a large amount of heat that can be recovered. The recovered heat is used to generate steam for process requirements and power generation. The process steam is generated at three different pressure levels. Additional power is generated by adding a small (50 to 100 MW) gas turbine for combined cycle power generation in each of the cases. The gas turbine in the natural gas case is fed process tail gases only; but the coal cases, which use the tail cases for coal drying, are fed natural gas. The total power generated is designed to equal the total estimated auxiliary loads for the coal fed cases 1 and 2. Excess power is generated from the process heat recovered and tail gas combustion in the natural gas case. No additional natural gas is consumed in the power cycle for the natural gas case. The model performance results are discussed above in section 3.2.

3.3.7 Water Balance

Water required for the operation of the facility is obtained from a source such as a lake, river, or well. If the quality of the water is adequate, raw water is used directly as makeup to the cooling tower and the gasifier quench. To meet the rest of the plant’s water needs, makeup must be treated first by filtration to create service-quality water. This quality of water serves as makeup to the plant’s potable water, demineralized water, fire water, and service water systems. Higher quality boiler feedwater is treated by a typical reverse osmosis and electrodeionization package. Water rejected by the system is of an acceptable quality to be used as makeup to the cooling tower.

In addition to meeting the makeup water needs of the facility, water treatment systems must be capable of capturing and treating on-site waste streams. Wastewater created by the gasification process must pass through a number of pretreatment steps before being combined with other wastewater streams. Metals, ammonia, and suspended solids are removed from the stream through the use of a clarifier and a biological treatment unit. Once processed, the wastewater can be combined with the cooling tower blowdown as well as other plant waste streams in a final clarifier. Dechlorination and pH adjustment are performed as needed at this step of the process in order to meet all local discharge regulations. Solids separated out in this process are dried by means of a filter press and taken away for offsite disposal.

Exhibit 3-25 shows the water balances for the coal cases. Since Case 2 uses the same specifications as Case 1, they both share a common mass balance in the exhibit. The water usage for the natural gas case was not modeled, but the water consumption and cooling requirements were estimated from similar systems.

Exhibit 3-9 Coal drying process flow diagram Case 1 and Case 2

Exhibit 3-10 Mass balance of coal drying process Case 1 and Case 2 – coal drying

Exhibit 3-11 Shell gasification process flow diagram Case 1 and Case 2

Exhibit 3-12 Mass balance of gasification process Case 1 and Case 2

Exhibit 3-13 CO shift reaction process flow diagram Case 1 and Case 2

Exhibit 3-14 Mass balance of shift reaction process Case 1 and Case 2

Exhibit 3-15 Low temperature gas cooling process flow diagram Case 1 and Case 2

Exhibit 3-16 Mass balance of low temperature gas cooling Case 1 and Case 2

Exhibit 3-17 Rectisol process flow diagram Case 1

Exhibit 3-17 Rectisol process flow diagram (continued) Case 2

Exhibit 3-18 Claus process flow diagram Case 1 and Case 2

Exhibit 3-19 Mass balance of Rectisol and Claus processes Case 1

Exhibit 3-19 Mass balance of Rectisol and Claus processes (continued) Case 2

Exhibit 3-20 Natural gas reforming process flow diagram Case 3

Exhibit 3-21 Mass balance of natural gas reforming Case 3

Exhibit 3-22 Methanol synthesis process flow diagram Case 1, Case 2, and Case 3

Exhibit 3-23 Mass balance of methanol synthesis Case 1 and Case 2

Exhibit 3-23 Mass balance of methanol synthesis (continued) Case 3

Exhibit 3-24 Power cycle heat and material balances Case 1

Exhibit 3-24 Power cycle heat and material balances (continued) Case 2

Exhibit 3-24 Power cycle heat and material balances (continued) Case 3

Exhibit 3-25 Water balances Case 1 and 2

Exhibit 3-25 Water balances (continued) Case 3

4. Economic Analysis

4.1 Cost Estimating Methodology

Capital and operating cost estimates developed for these cases were based on adjusted vendor- furnished quotes, previous studies, actual cost data, or the best available information. All estimates are expressed in June 2011 dollars consistent with NETL methodology documented in Quality Guidelines and baseline reports. [17,18,19]

4.1.1 Capital Costs

The total overnight cost (TOC) for each plant was calculated by adding owner’s costs to the total plant cost (TPC). TPC includes all equipment (complete with initial chemical and catalyst loadings), materials, labor (direct and indirect), engineering and construction management, and contingencies (process and project).

The capital costs have an estimated accuracy of +30/-15 percent, consistent with the screening study level of design engineering applied to the various cases in the study. The value of the study lies not in the absolute accuracy of the individual cases, but in the fact that all cases were evaluated under the same set of technical and economic assumptions. The consistency of the approach allows meaningful comparisons among the cases evaluated.

Process contingency was added to cost accounts to compensate for uncertainty in cost estimates caused by performance uncertainties associated with the development status of a technology.

Project contingency was added to the engineering/procurement/construction management (EPCM) capital accounts to cover project uncertainty and the cost of any additional equipment that would result from a detailed design. The contingencies represent costs that are expected to occur. Each bare erected cost (BEC) account was evaluated against the level of estimate detail and field experience to determine project contingency.

TOC values are expressed in June 2011 dollars. The estimate represents current commercial offerings for the gasification and methanol synthesis and conversion processes. The estimates represent a complete fuels plant facility. The boundary limit is defined as the total plant facility within the fence line, including the coal receiving facilities and water supply system. Costs were grouped according to a process/system-oriented code of accounts; all reasonably allocable components of a system or process are included in the specific system account in contrast to a facility, area, or commodity account structure.

4.1.2 Feedstock prices

4.1.2.1 Coal Price

The coal type assumed for this study is a Powder River Basin (PRB) subbituminous coal supplied from the Montana Rosebud mine. The coal price was assumed to be $36.57 per short ton (year 2011 dollars) based on Montana Rosebud PRB Coal delivered to the Midwest as specified in the QGESS: Fuel Prices for Selected Feedstocks in NETL Studies. [1]

4.1.2.2 Natural Gas Price

The natural gas price was assumed to be $6.13/MMBtu (2011 dollars) based on natural gas prices delivered to Midwest power plants as specified in the QGESS: Fuel Prices for Selected Feedstocks in NETL Studies. [1]

4.1.3 Production Costs and Expenses

The production, or operations and maintenance (O&M), costs described in this section pertain to charges associated with operating and maintaining the methanol plant over its expected life.

O&M costs are determined an annual basis for the first year of operation. Quantities for major consumables, such as feedstock and fuel, were taken from the heat and mass balance developed for this application. Using reference data, other consumables were evaluated on the basis of the quantity required. Operating labor costs were determined on the basis of the number of operators. Maintenance costs were evaluated on the basis of requirements for each major plant section.

The O&M costs and expenses associated with the plant include the following:

Operating labor
Maintenance – material and labor
Administrative and support labor
Consumables
Fuel/Feedstock cost
Taxes and insurance

These costs and expenses are estimated on a reference basis and escalated to June 2011 dollars. The costs assume normal operation and do not include the initial startup costs. The operating labor, maintenance material and labor, and other labor-related costs were combined and then divided into two components: fixed O&M costs, which are independent of liquids production, and variable O&M costs, which are proportional to liquids production. The variable O&M cost estimate allocation is based on the plant capacity factor.

The other operating costs, consumables, and feedstock, are determined on a daily, 100-percent operating-capacity basis and are adjusted to an annual plant operation basis. The inputs for each category of operating costs and expenses are identified in the succeeding subsections, along with more specific discussion of the evaluation processes.

4.1.4 Required Selling Price

The figure-of-merit in this report is the required selling price (RSP) expressed in $/gal of crude methanol. The RSP values were calculated using the Power Systems Financial Model (PSFM)[20] and estimated to be the value calculated when the required return on equity (ROE) equals the internal rate of return (IRR) for 30 years of operation based on the assumed financial structure and escalations. RSP was assumed to escalate at three percent per year for the thirty- year economic life of the plant. All costs are expressed in June 2011 dollars.

For the natural gas case which generates excess electricity, the price of electricity was assumed to be $59.59/MWh expressed in June 2011 dollar. This is based on the lowest cost option from the Bituminous Baseline study results [19], which was considered to be typical of baseline 2011 plant designs. A sensitivity analysis of the RSP to this assumed value is included in the results discussion. While the excess power can be sold to the grid; the sale may be at a steep discount as entities that are negotiating a power purchase agreement will know the power production is an inherent by-product of core methanol production operations. Consequently, the actual achieved transfer price for excess power will be a significant risk in natural gas feedstock projects and would be highly project dependent. The impact of the price on the methanol RSP is reflected in the sensitivity analysis results.

The capital and operating costs for CO2 T&S were independently estimated by NETL at $11/metric ton of CO2 in 2011 dollars. [9]

The RSP was calculated for each case assuming (i) a financial structure representative of a commercial fuels project, and (ii) a financial structure with loan guarantees or other government subsidies. The financial assumptions and structures used to estimate the RSPs are shown in Exhibit 4-1, Exhibit 4-2, and Exhibit 4-3. [21] A sensitivity of the RSP to this assumed structure was included in the results discussion.

Exhibit 4-1 Financial assumptions for PSFM [17]

Exhibit 4-2 Financial structure for commercial fuels projects [22]

Exhibit 4-3 Financial structure for loan guarantee projects [23]

4.2 Cost Estimation results

The capital and O&M costs for each of the cases are shown in Exhibit 4-4 through Exhibit 4-9.

Exhibit 4-4 Capital cost summary for Case 1

Exhibit 4-5 Capital cost summary for Case 2

Exhibit 4-6 Capital cost summary for Case 3

Exhibit 4-7 O&M cost summary for Case 1

Exhibit 4-8 O&M cost summary for Case 2

Exhibit 4-9 O&M cost summary for Case 3

4.3 Summary Comparisons

The cost of product (crude methanol) was estimated for each case and the results are listed in Exhibit 4-10 and illustrated in Exhibit 4-11. The RSP for the natural gas case is significantly lower than the coal cases. While the coal case values are dominated by the capital costs, the natural gas case is dominated by the feedstock costs.

Exhibit 4-10 Cost estimation results

The RSP for the natural gas case is consistent with the information provided in Methanex’s 2011 annual report. [24] Specifically, the margin inferred from their financial statements is on par with the weighted average cost of capital used in the PSFM for this study. The average sales prices of methanol reported by Methanex and other vendors (~ 1.28 dollars per gallon or $426 dollars per metric ton in Methanex historical data reference [25]) are higher than the RSP of the natural gas cases in this study because their prices also cover other significant costs of methanol sales and distribution associated with transporting the methanol from production locations to markets and distribution within markets5. The Methanex price may also include the costs of purification into refined methanol.

Exhibit 4-11 Components of required selling price

The sensitivity of the RSP to the coal and natural gas feedstock prices is illustrated in Exhibit 4-12 and Exhibit 4-13, respectively. The prices used in this study are based on the NETL QGESS Recommended Fuel Prices. [1] Cases 1 and 2 use coal from the Powder River Basin (PRB) region in Montana, priced at $36.57/ton ($2.1351/MMBtu) including delivery to the Midwestern site, as the feedstock. Case 3 uses natural gas, priced at $6.13/MMBtu. Cases 1 and 2 also burn some natural gas to provide additional power required by the process.6 The cost of the natural gas used for the NGCCs in the coal cases is less than four percent of the RSP values, so it was assumed constant at $6.13/MMBtu for the coal cases in this chart. The expected RSPs for the case using natural gas as feedstock are below those for the coal-based cases because of the lower capital and fixed operating costs. On an energy content basis, the RSPs for all cases increase at approximately the same rate with increases in the feedstock prices due to the similarity between the feedstock requirements per unit of product. Cases 1 and 2 require 9.75 MMBtu of PRB coal per gallon of Methanol generated and Case 3 requires 9.27 MMBtu of natural gas per gallon of Methanol generated.

5 These costs are a key reason why chemicals and fuels projects derivative of potential mega-methanol projects are generally co-located with methanol production.

6 A natural gas combined cycle plant was used to generate additional power in the coal gasification cases to make the plant approximately power neutral while maximizing the production of crude methanol (i.e., no syngas is diverted for power production).

Given the relatively higher cost per MMBtu of natural gas compared to coal and the similar energy input requirements of the conversion process, a 100 percent increase in natural gas prices leads to a 50 to 55 percent increase in RSP, for commercial and loan guarantee financing respectively. However, the same increase of 100 percent in coal prices leads to a 13 to 16 percent increase in RSP for the same financing structures, respectively. These results are consistent with the natural gas–based methanol production being the less capital intense but more operating margin dependent technology choice. Hence, the natural gas route is more exposed to feedstock price volatility.

Exhibit 4-12 Sensitivity of the required selling price to the coal price

Exhibit 4-13 Sensitivity of the required selling price to the natural gas price

The heat recovery from the exothermic methanol synthesis reaction combined with burning the excess tail gases in a combustion turbine result in a substantial amount of excess electricity being generated in the natural gas case. If this electricity is sold, the revenue is applied to the cash flow and results in a lower RSP. The sensitivity of the RSPs to the selling price of electricity is illustrated in Exhibit 4-14. The RSPs for the coal feedstock cases remain constant but are shown for comparison. The RSPs for the natural gas feedstock case decease by approximately 0.08 cents per gallon for each dollar per MWh increase in electricity selling price.

While the excess power can be sold to the grid; the sale may be at a steep discount as entities that are negotiating a power purchase agreement will know the power production is an inherent by- product of core methanol production operations. Consequently, the actual achieved transfer price for excess power will be a significant risk in natural gas feedstock projects and would be highly project dependent. However, as shown in Exhibit 4-14, the impact of reduced electricity revenues has a fairly small impact on the methanol RSP. Decreasing the electricity sell price to $0/MWh only increases the methanol RSP by about $0.05/gal.

Exhibit 4-14 Sensitivity of the required selling price to the electric selling price

Options for carbon dioxide sequestration include both storage in a saline reservoir (i.e., carbon capture and sequestration (CCS)) and usage in enhanced oil recovery (EOR) (i.e., carbon capture, utilization and storage (CCUS)). The EOR option may be attractive even without the passage of carbon regulations. The impact of selling the captured CO2 for EOR or other uses at various plant gate sale prices was estimated, and the results shown in Exhibit 4-15. The horizontal lines represent the without-capture case RSP values. As the plant gate sale price increases, the RSP values decrease and approach the without-capture values. The plant gate sale price at the point where each capture case line crosses the corresponding without-capture line is equal to the cost of CO2 captured for that capture case. The cost of capture can be interpreted as the breakeven plant gate sale price where the cost of capture equals the revenue generated by selling the recovered CO2.

Exhibit 4-15 Sensitivity of RSP to CO2 plant gate sales price

The capacity factor was assumed to be 90 percent for the cases in this study. The sensitivity of the RSP to the capacity factor is illustrated in Exhibit 4-16. The RSPs for the coal cases increase at a faster rate than the natural gas cases as the capacity factor falls because the coal-based cases have more dollars of capital underutilized when the capacity factor is reduced.

Exhibit 4-16 Sensitivity of the required selling price to the capacity factor

Two financial structures were assumed for calculating RSPs in this activity as described in Section 4.1.4. These structures are based on typical values for fuel projects with and without loan guarantees or government subsidies. The assumed values were used in the PSFM to establish capital charge factors (CCF), the portion of the total overnight capital cost to include in the annual cost of producing a product, for each financial structure. The sensitivity of the RSP to the CCF is illustrated in Exhibit 4-17. The RSP values were calculated for CCFs ranging from 10 percent to 35 percent (the value estimated for a project assuming 100 percent equity and 20 percent internal rate of return on equity). The RSP values calculated for a CCF of 12.4 percent (the value estimated for a high risk investor-owned utility (IOU) project assuming 45 percent debt at 5.5 percent interest and 55 percent equity and 12 percent internal rate of return on equity) are also included in the chart. The RSPs for the coal cases increase at a faster rate than the natural gas cases as the capital charge factor increases because the coal-based cases have more dollars of capital to be included in the cost of production. The coal case values increase by approximately 5 cents per gallon for each one absolute percent increase in the capital charge factor. The natural gas case values increase by approximately 2 cents per gallon for each one absolute percent increase in the capital charge factor. Loan guarantees and/or government subsidies could reduce the RSPs for the coal-based cases closer to the values of commercial natural gas cases.

Exhibit 4-17 Sensitivity of the required selling price to the capital charge factor

Source: National Energy Technology Laboratory

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Life-Cycle Analysis of Methanol Production From Coal and Natural Gas