AICHE Coal Gasification Report

AIChE National Design Competition Competition

Coal Gasification Study Process Design Report By: Kaitlyn D. Kelly Kelly

March 7, 2008

Summary 

Rising oil prices and increased demand for chemical feedstock have added to the push to find alternative sources sources of energy. The gasification gasification of coal to produce methanol methanol has been proposed as a means for meeting the demand for energy and chemical feedstock. The 2008 AIChE National Student Design Contest poses the problem to determine if a company can capitalize on this opportunity through the design, construction, construction, and operation of a world-scale methanol production production facility. A technical and economic evaluation of the proposed project has been completed to help answer this question. The facility must be able to produce 5000 metric metric tons of methanol per day. The methanol must meet AA grade purity requirements. The key elements of the base case design can be summarized with the following block flow diagram.

A material and energy balance was created to provide for quick analysis of the flows in the process. This also allowed the major equipment in the process to be simulated separately without the need to combine the overall process in one simulation. The software used for simulation of the majority of the process was Chemcad. The gasifier was simulated using the software Gasify. The composition of the methanol product may be seen in the following table. Methanol Product (lb/hr)

Water Nitrogen Oxygen Hydrogen Carbon monoxide Carbon dioxide Argon hydrogen sulfide COS Methane Ammonia MDEA Carbon Methanol Ethanol Coal Ash

2149 0 0 0 0 495 0 0 0 0 0 0 0 454173 0 0 0

Total

456817

Table 1: Methanol product component flow rates.

A control scheme was developed for the process to help prevent excursions during operation. A limited safety analysis was conducted on the project. Hazard and operability studies were completed for one hazard for each of the major process units. For each of those identified hazards, one initiating event was considered for a layer of protection analysis. This brief analysis served to highlight major areas of concern in the process. A primary concern is that the process has very large flow rates and thus very large inventories of material. Additionally, there are areas of the process that operate at elevated temperatures. These present an inherent safety problem. The inventories of both material and energy need to be decreased as much as possible to increase the safety of the system. The major process equipment was sized and priced. This was done to determine the capital investment that would be required for the project. The pumps, drums, and control systems were not priced but were estimated. The pumps and drums were assumed to be 10% of the calculated capital costs. The control systems were estimated to cost an additional 10% of the total capital costs. These are summarized in the following table.

Capital Investment

Costed Equipment

$

440,593,371

Estimates for pumps/ drums Control Systems Total

$ $ $

44,059,337 48,465,271 533,117,979

Table 2: Summary of the required capital investment for the project.

The base case design is not economically feasible. The estimated costs of liquid nitrogen and waste processing drive up the cost of annual operation. A sensitivity analysis was conducted on the economics to determine the impact of errors in the estimates of the variables used in the analysis. It was found that it would take a combination of deviations in the values to substantially alter the economics and make the project profitable. A summary of the annual expenditures, revenues, and discounted cash flow may be seen below.  Annual Expenditures

Consumption Coal (lb) Oxygen (lb) Process Water (lb) LP Steam (lb) MP Steam (lb) HP Steam (lb) HHP Steam (lb) Electricity (KW) Cooling Water Makeup (lb) Waste Water Treatment (lb) Bulk Liquid Waste Proc. (lb) Vents/Vapors Processing (lb) MDEA (lb) N2

Hour 542591.807 276803.21 286000 90528.9728 0 0 157706.4 48299.7347 10000 1331705.53 110202.689 446203.307 11282.0691 5634512.5

Day Year 13022203.36 4753104226 6643277.043 2424796121 6864000 2505360000 2172695.348 793033802 0 0 0 0 3784953.6 1381508064 1159193.633 423105676 240000 87600000 31960932.71 1.1666E+10 2644864.539 965375557 10708879.36 3908740968 270769.6579 98830925.1 135228300 4.9358E+10 Total Expenditures

$ $ $ $ $ $ $ $ $ $ $ $ $ $ $

Yearly Value 144,494,368 75,774,879 300,804 5,310,494 16,035,361 29,617,397 11,569 1,400,638 176,698,205 575,841,303 74,123,194 1,628,824,874 2,728,433,087

Table 3: Annual expenditures and operating costs.

 Annual Revenues

Income/Credits Methanol (lb)

Hour 456817

Day 10963601.85

Year 4001714674

Yearly Value $ 571,673,524.91

Table 4: Annual income from the methanol production.

Discounted Cash Flow

Year 2008 2009 2010 2011 2012 2013 2014 2015 2016 2017 2018 2019 2020 2021 2022 2023 2024 2025 2026 2027 2028

Expenses $533,117,979 $2,826,279,619 $2,910,588,201 $2,997,426,041 $3,086,869,016 $3,178,995,280 $3,273,885,332 $3,371,622,086 $3,472,290,943 $3,575,979,865 $3,682,779,455 $3,792,783,032 $3,906,086,717 $4,022,789,512 $4,142,993,391 $4,266,803,387 $4,394,327,682 $4,525,677,707 $4,660,968,232 $4,800,317,472 $4,943,847,190

Revenues $0 $588,823,731 $606,488,443 $624,683,096 $643,423,589 $662,726,296 $682,608,085 $703,086,328 $724,178,918 $745,904,285 $768,281,414 $791,329,856 $815,069,752 $839,521,844 $864,707,500 $890,648,725 $917,368,186 $944,889,232 $973,235,909 $1,002,432,986 $1,032,505,976

Depreciation $0 $35,541,199 $35,541,199 $35,541,199 $35,541,199 $35,541,199 $35,541,199 $35,541,199 $35,541,199 $35,541,199 $35,541,199 $35,541,199 $35,541,199 $35,541,199 $35,541,199 $35,541,199 $0 $0 $0 $0 $0

Taxes ($213,247,192) ($909,198,835) ($935,856,383) ($963,313,657) ($991,594,650) ($1,020,724,073) ($1,050,727,378) ($1,081,630,783) ($1,113,461,289) ($1,146,246,711) ($1,180,015,696) ($1,214,797,750) ($1,250,623,265) ($1,287,523,547) ($1,325,530,836) ($1,364,678,344) ($1,390,783,798) ($1,432,315,390) ($1,475,092,929) ($1,519,153,794) ($1,564,536,486)

Profit ($319,870,788) ($1,328,257,053) ($1,368,243,376) ($1,409,429,288) ($1,451,850,777) ($1,495,544,911) ($1,540,549,869) ($1,586,904,976) ($1,634,650,736) ($1,683,828,868) ($1,734,482,345) ($1,786,655,426) ($1,840,393,699) ($1,895,744,121) ($1,952,755,055) ($2,011,476,318) ($2,086,175,697) ($2,148,473,085) ($2,212,639,393) ($2,278,730,692) ($2,346,804,729)

Table 5: Annual discounted cash flow analysis.

There is the potential to decrease some of the expenses through optimization of the design. The majority of the vapor waste that requires processing comes from the methanol refining process and primarily consists of carbon dioxide. Since carbon dioxide is a reactant in methanol synthesis, this can be used rather than treated as waste. The vapor may be sent to an additional reactor where hydrogen is added. The methanol produced from the reaction can be recycled to the purification step. Not only would this decrease the expense of vapor waste processing, but the increased methanol yield would decrease the feeds required to the process. The smaller feeds would result in smaller flow rates throughout the process, decreasing inventory and increasing the inherent safety. Also, the equipment could be smaller, reducing capital costs. Additionally, the lower feeds may result in lower utility usage, particularly of the liquid nitrogen which is by far the highest expense. In addition to this change in the process, I recommend integrating the heat exchange in section 400. There is currently a condenser which uses liquid nitrogen. If the recycled gas, which is at -28°F, were used to first cool the reactor effluent before it is sent to the condenser, this would decrease the nitrogen required. Additionally, since the recycled gas must be re-heated before entering the reactor, this would decrease the required amount of steam for the process. These two changes have the potential to make the process profitable.

It is my recommendation that the optimization schemes be further evaluated before the design proceeds. Additionally, before the design is complete, a more thorough safety analysis needs to be completed on the project including a full risk assessment. Inherent safety needs to be explored in detail. Finally, the prices of liquid nitrogen, MDEA, and waste processing need to be researched further to ensure that correct values are used in the economic analysis.

Introduction

The AIChE National Design Competition for 2008 poses the challenge of converting coal to methanol with a gasification process. The gasification of coal to produce methanol is proposed as a method for storing energy as well as an intermediate for the chemicals industry. A business opportunity exists due to the rapid rise in the price of crude oil in combination with the increased demand for chemical feedstock. Crude oil is the primary source of feedstock for the chemicals industry. These factors have further fueled the drive to find alternative sources of energy and chemical feedstock. Coal gasification presents an opportunity in the United States because of the vast supply of coal. According to the source, “Beyond Oil & Gas: The Methanol Economy,” the utilization of methanol may provide a means for meeting t he world’s future energy needs.

The goal of this study is to determine if a company can capitalize on this opportunity. The company would design, construct, and operate a world-scale methanol production facility. To evaluate the potential project, a complete technical and economic evaluation needed to be conducted. If the project is found to be viable, the plant will be built in the Texas Gulf Coast. A flat thirty day time limit has been imposed on the evaluation of the project. The evaluation includes a preliminary design of the project as well as a study of the economic feasibility. On a more detailed level, the analysis considers the objectives and limitations of the project. A strategy of approaching the problem was developed. The scope also includes the identification of important commercial sourcing options. A primary part of the analysis was to determine if the targeted performance is attainable. This was done through process simulation. A base case of the design was simulated using Chemcad software. From the base case, the major equipment was sized and the capital costs were estimated. These were used, along with the annual expenses and revenues, to assess the economics of the project. Since most of the economic variables are estimates, the sensitivity of the project on these variables was evaluated. Safety is a primary consideration in any design. The safety was very briefly evaluated. Hazards and operability studies were done on one hazard for each major unit operation. Layer of Protection Analyses were done for one initiating event for each identified hazard. This was by no means an adequate analysis of the safety of the entire process. But, it did provide some insight into the hazards that are inherent in the process as well as the additional layers of protection that need to be added to the design. Areas of concern have been identified for further analysis. This report serves to make recommendations regarding the continuation of the project. Aspects of the design that require further consideration are presented in this report. Safety aspects of the proposed design are also outlined. In addition to providing findings from the study, the report also serves to highlight what is still unknown about the process. Design Basis

The coal feed for the gasification was to be chosen from three provided sources. They are summarized in the following table with composition in weight percent:

Figure 2: Compositions of coal feeds available.

The other feeds to the process are oxygen from an off-site source and steam and water. The following prices were given for the feeds: Feed Prices

Coal TX Lignite WY Sub-Bituminous IL Bituminous Pre-Processing Oxygen

Cost $/short ton $ 15.20 $ 10.60 $ 32.00 $ 40.00 $ 70.00

Transportation $/short ton $ 3.90 $ 10.20 $ 6.90 /short ton /metric ton

Table 6: Prices of feeds to the process.

The synthesis gas from the gasifier will have sulfur compounds present. These need to be removed before further processing of the gas because sulfur is a known catalyst poison. The sulfur must be removed to 0.1ppmv or less. The technology chosen for removing the acid gas needs to have high selectivity for hydrogen sulfide to carbon dioxide. This is because carbon dioxide can be used as a reactant in methanol synthesis. The final methanol product must meet AA methanol grade purity. It must be greater than 99.85wt% methanol on a dry basis. Also, there must be less than 0.1wt% water. Additionally, the ethanol needs to be removed to 50ppmw or less. The selling price of the methanol is $320/metric ton. The plant must be able to produce methanol on a world scale. The production capacity must be 5000 metric tons per day. The plant will be built on a brown field site in the Texas Gulf Coast. Land purchase has been neglected in the economic analysis. Utilities are available for the plant, raw water is available from neighborhood resources, and electricity is available on the power grid. Off-site waste treatment is also available and is assumed to be in compliance with environmental regulations. Utility costs have been provided for most of the utilities that are needed. The following is a summary of the provided utility costs.

Utilities

HHP Steam HP Steam (400psig) MP Steam (100psig) LP Steam (50psig) Electricity Condensate Cooling Water makeup Process Water Demin Water Potable Water Waste Water Treatment TOC in Waste Water Instrument Air Bulk Liquids Waste Processing Bulk Solids Waste Processing Vents/Vapors Processing Inert Gas

$26.00 $22.00 $19.00 $15.00 $0.07 $0.75 $1.10 $1.00 $3.00 $2.50 $1.00 $0.70 $0.45 $410.00 $325.00 $330.00 $0.35

/metric ton /metric ton /metric ton /metric ton /KWH /metric ton /Mgals /Mgals /Mgals /Mgals /Mgals /lb TOC /MCF /metric ton /metric ton /metric ton /MCF

Table 7: Summary of utility costs

Calculation Basis:

Reaction data was provided for the water-gas shift reaction as well as the methanol synthesis reaction using carbon monoxide. The equilibrium data is the following: CO + H2O ↔ CO2 + H2 CO +2H2 ↔ CH3OH

lnKp = -4.33 + 4577.8/T(K) lnKp = -42.918 + 11284/T(K)

Additionally, I know that carbon dioxide can also be used to produce methanol. Supporting evidence of this may found in the articles by Graaf et al. on the kinetics of methanol synthesis (see references). These articles also provide equilibrium data for the three reactions that take place in the methanol synthesis reactor. To determine that the data provided in the problem statement was consistent with that in the article, I converted the units for the values of the equilibrium constant in the article and took the natural log of these values. I then compared these with the given values of the equilibrium constants for the water-gas shift reaction and the reaction of carbon monoxide to methanol from the problem statement. These were found to be generally consistent. So, the equilibrium constant for the reaction of carbon dioxide to methanol was taken from the article. This equilibrium data is the following: CO2 + 3H2 ↔ CH3OH + H2O

lnKp = -15.7 + 4095.9/T(K)

Ethanol is also created in the methanol synthesis reaction. It is made at a rate of one part ethanol per one hundred parts methanol. From this, the kinetics of the reaction could be estimated. However, since the effluent composition of ethanol is known, I just added that much to the feed of the tower separating water from methanol. This was to ensure that enough of the ethanol could be removed to meet the specifications. This was confirmed, so I decided that it wasn’t necessary to calculate the ethanol synthesis kinetics since the overall goal was to ensure that it was removed from the product. For simulations, the NRTL (Non-Random Two Liquid) model was used to determine K-values. The problem statement says that this is to be used for the liquid phase equilibria. Since most of the separations occur in the liquid phase, NRTL was used as the K-value model for the entire process. NRTL calculates activity coefficients to account for non-ideality. To determine which coal was the best value, I used the shipping, coal, and pre-processing costs that were given on a per ton of coal (as received) basis. Using the coal compositions, I determined what the costs of the coal were on a per ton dry basis as well as a per ton of carbon basis. Simulation of the gasifier was done using the program Gasify   which may be found on the companion website for the book Gasification by Higman and Burgt, http://www.gasification.higman.de/ . The program uses a basis of one hundred kilograms of coal. This was scaled up to the desired quantity of coal to determine the amount of the other feeds that were required and the amount of products. To simplify the calculations, I assumed that all the ash is separated. I know that some ash will be carried with the syngas. However most should be removed in the particulate filter following the gasifier so this assumption should be okay. It was recommended in the problem statement that an equilibrium reactor be used to simulate the water-gas shift reaction. However, I was having difficulty obtaining a converged solution. So, ultimately, I used a Gibb’s reactor for this simulation.

It found the conversion of the products by

minimizing Gibb’s fee energy. The ratio of the products of this reaction was two to one hydrogen to

carbon monoxide on a molar basis based on the amount of hydrogen that is required for the carbon monoxide reaction. However, it was later seen that due to the additional reaction of carbon dioxide with hydrogen to produce methanol, it would have been advantageous to produce much more hydrogen. Chemcad was used for sizing the towers and heat exchangers since it has a convenient utility that can size those pieces of equipment. The reactors were sized by hand. The gasifier was sized as a vertical, cylindrical furnace due to the combustion reaction and the high temperatures involved. The methanol synthesis and water-gas shift reactors were sized as vertical pressure vessels. The parameters determined from the hand calculations were the design temperature, design pressure, height of vessel, diameter of vessel, thickness, and allowable stress. To cost the equipment, chemical engineering price indices were used. Construction of the plant was slated to begin in 2009, so the price indices were projected to January 2009 based on trends from February 2003 to June 2007. The costing utility in Chemcad was used to price the same equipment that

was sized with the software along with the flash vessel, compressor, and expander. The reactors were priced by hand using design heuristics. These prices are all estimates and are only assumed to be accurate within thirty percent. Though there will be pressure drops in the lines, through the valves and equipment, and due to height changes, these were not accounted for as they would just be estimates. There are pumps in the process flow diagram, but it is likely these will either not be necessary or will need to be replaced with compressors since the material is mostly gas. So, to account for the uncertainty, 10% of the capital cost was added as stated above. This should be enough to add extra pumps or compressors as required. Technology Selection & Criteria

There are a variety of different technologies available for gasification.

The resource,

Gasification, provides a summary of those which are used in industry. The criteria I used in selecting the

technology was oxygen consumption, synthesis gas purity, carbon conversion, and the types of coal the process could handle. I chose entrained flow as the general gasification process. It provides the highest quality of synthesis gas and works with a variety of different coals. This is important so that if the coal source must change, the process will still operate. Of the various technologies available for an entrained flow process, I found the Noell/GSP gasifier to be the best suited for this project. The GSP process is owned by Future Energy GmbH. Its unique characteristics are that it is top-fired with a single mounted burner. So, all the reactants enter at the same place. This makes the equipment simpler and thus, inherently safer than the alternative technologies. The cylindrical design allows for lower equipment costs compared to the other available technologies. The slag (ash and particulates) leaves through the same outlet as the synthesis gas. This decreases the potential for blockages to occur in the slag outlet. A water quench is also incorporated into the design to quickly cool the gas to 900°C while minimizing contact with the walls of the reactor. The effluent gas from the reactor will be sour, containing sulfur compounds. Therefore, the gas must be treated in an acid gas removal process. Absorption is the most widely used process to remove acid gas when succeeding coal gasification. To choose the best method of absorption, gas purity, raw gas composition, and selectivity were considered. Chemical absorption onto a liquid solvent was chosen because it has a higher loading capacity than physical absorption. N-methyl-diethanolamine, or MDEA, was the solvent of choice because not only is it the most widely used in industry, but it is highly selective to remove more hydrogen sulfide than carbon dioxide. Again, this is important considering carbon dioxide is a reactant in the synthesis of methanol. As was previously stated, the prices of the three different types of coal were compared on a per carbon basis. The results may be seen in the following table. Coal Source As Received Dry Basis Carbon Basis

Martin Lake TX Lignite $59.10/ton $90.68/ton $140.28/ton C

PRB Sub-Bituminous $60.80/ton $67.94/ton $101.64/ton C

Illinois Bituminous $78.90/ton $90.69/ton $115.18/ton C

Table 8: Summary of coal prices on a basis on a short ton of Carbon

In addition to price, the coal sources were compared considering energy densities and composition. The PRB sub-bituminous coal was selected because it had the lowest sulfur content. Since the sulfur ultimately must be removed to 0.1ppmv, this was an important factor. Additionally the PRB was the best price for the quality of coal when compared on a per ton of carbon basis as well as on a dry basis. Either a two or three-phase reactor system could be used for the methanol synthesis. Due to the high concentration of carbon dioxide entering the reactor, I chose the three-phase model. The three-phase model consumes more carbon dioxide in the reaction than does the two-phase model. Following the concept of minimization for inherent safety, five reactors in series were actually used rather than one very large reactor. The flowsheet shows only one, but it can be viewed as the overall reaction system. Information comparing the two and three-phase models was found in the article “Comparison of two-phase and three-phase methanol synthesis processes”  by Graaf and Beenackers. Process Analysis and Discussion

The block flow diagram on the following page provides a simplified view of the overall process.

The coal is fed along with steam and oxygen to the gasifier. Inside the gasifier, the coal is combusted to produce synthesis gas. The gas is produced at 1500°C but is also quenched with water inside the gasifier. The outlet gas is at 900°C. Any particulates that are in the gas are removed in a filter following the gasifier. The hot gas is cooled in a synthesis gas cooler that simultaneously raises low pressure steam from the cooling water. The cooled synthesis gas is sent to the acid gas removal system. Absorption is used with MDEA as the solvent to remove the acid gas from the system. The sulfur compounds are targeted for removal. The rich MDEA is then regenerated in a second tower using steam as the stripping gas. A rich/lean heat exchanger is used to cool the lean MDEA that is recycled to the absorber and to heat the rich MDEA that is fed to the stripper. Along with water, the sweetened gas is then fed to the water-gas shift reactor. This reacts carbon monoxide and water to form carbon dioxide and hydrogen. Enough hydrogen should be created to provide a minimum ratio of two to one hydrogen to carbon monoxide. Three-phase methanol synthesis is used in the next step. The gas from the watergas shift reactor is combined with recycled un-reacted gas from the methanol synthesis. It is then heated to 530K and fed to the synthesis reactor. As mentioned previously, this is actually a system of five reactors in series. The reactor effluent is cooled and compressed to lower the vapor fraction. Following, it is sent to a flash vessel where the methanol and water are removed along with some carbon dioxide. A fraction of the remaining material is purged to prevent buildup and the rest is recycled back to the feed of the reactor. The methanol stream is sent to the refining and purification system. This consists of two distillation towers. In the first tower, water is removed from the methanol out the bottoms as waste. The distillate from this tower is primarily methanol and carbon dioxide. This stream is sent to the second tower where carbon dioxide is removed in the distillate as waste. The bottoms of this tower is the methanol product. This process can be seen in more detail in the process flow diagram on the following pages. A control scheme has been developed for the process to act as a safeguard against excursions. The controls may also be seen in the process flow diagram.

The compositions and properties of all the streams in the flowsheet may be seen in the following material balance. The streams shaded brown are removed from the process and the streams shaded blue are fed. Not all stream properties are known because the process wasn’t tied together in

the simulator and pumps were not simulated. The specifications made for the material balance were found through simulation and may also be seen on the following pages. Temperatures are in degrees Fahrenheit and pressures are in psia.

The energy balance summarizes the heat exchanger duties and the gasifier duty. It may be seen below. Energy Balance

Equipment # Equipment Name Utility Used Duty (MMBtu/hr)

D-101 Gasifier Gas 6083

E-101 Syngas cooler Cooling Water -125.2

E-201 Rich/lean exchanger N/A 1042

E-301 D301 Preheat Steam 21.5

E-401 D401 Preheat Steam 146.3

Equipment # Equipment Name Utility Used Duty (MMBtu/hr)

E-402 N2 Cooler N2 -583

E-501 T501 Condenser Cooling Water -33.5

E-502 T501 Reboiler Steam 20.1

E-503 T502 Condenser Cooling Water -21.9

E-504 T502 Reboiler Steam 45.7

Table 10: Summary of heat provided to or removed from the process.

The major equipment has been sized and priced. The pumps and drums were neglected at this point and just assumed to be ten percent of the calculated capital cost. Additionally, control systems were just assumed to be ten percent of the total capital costs. Equipment summaries and equipment costs may be seen below. Towers

trays

spacing (ft)

Passes

diameter op. pressure height (ft) (ft) (psia)

T-201 (2) AGR Absorber

40

3

3

19

136

120

T-202

Stripper

35

2

3

32.5

86

30

T-501

W Removal

44

2.5 / 2

1

22

114.5

30

T-502

CO2 Removal

10

2

1

14

36

30

#

Name

Table 11: Equipment summaries, towers.

Reactors

#

Name

D-101 D-301 D-401

Gasifier Water-Gas Shift MeOH Synthesis

volume (ft3) 10000 26750 5900

length (ft) 110 37.8 32

diameter (ft) 11 30 16

design P (psia) 508 95 80

Table 12: Equipment summaries, reactors.

design T (psia) 3000 710 600

Allowable Stress (psia) 21600 16920

Heat Exchangers

#

Name

E-101 E-201 E-301 E-401 E-402 E-501 E-502 E-503 E-504

Syngas cooler (2) Rich/Lean Exch. D301 Preheat D401 Preheat (2) N2 Cooler T-501 Cond T-501 Reboil T-502 Cond T-502 Reboil

req'd area (ft2) Utility 12018 CW 18635 Process 748.9 Steam 22223 Steam 45386 N2 1952 CW 899 Steam 1440 N2 780 Steam

Duty (MMBtu/hr) ΔT (°F) -125.2 350 1042 71 21.5 75 146.3 214 -583 900 -33.5 4 20.1 0.5 -21.9 80 45.7 0

Table 13: Equipment summaries, heat exchangers.

Pump/Compressor/Expander

#

Name

hp

J-102 J-203 (2) C-401

Expander Lean MDEA Pump Compressor

-35216 1086 96900

ΔP (psi)

-315 90 120

Table 14: Equipment summaries, pressure exchangers.

Equipment Costs

Equipment # Equipment Name D-101 Gasifier J-102 Expander T-201a AGR Absorber T-201b AGR Absorber E-101 Syngas cooler (2) E-201 Rich/Lean Exch. T-202 Stripper J-203 (2) Lean MDEA Pump E-301 D301 Preheat D-301 H2O-Gas Shift E-401 D401 Preheat (2) D-401 (5) MeOH Syn React. C-401 Compressor E-402 N2 Cooler F-401 Separator T-501 Refining E-501 T-501 Cond E-502 T-501 Reboil T-502 CO2 Removal E-503 T-502 Cond E-504 T-502 Reboil

Cost Index 754.4 941 754.4 754.4 806.7 806.7 754.4 941 806.7 754.4 806.7 754.4 941 806.7 754.4 754.4 806.7 806.7 754.5 806.7 806.7

Module Factor Base Cost 4.34 $5,816,961 3.21 $184,359,904 4.34 $4,919,996 4.34 $4,919,996 3.39 $649,216 3.39 $1,877,210 4.34 $32,078,802 3.48 $585,534 3.39 $90,228 4.34 $4,476,713 3.39 $2,866,160 4.34 $11,361,544 3.21 $67,835,016 3.39 $4,434,939 2.39 $3,464,757 4.34 $7,901,236 3.39 $192,804 3.39 $60,938 2.39 $1,305,816 3.39 $149,481 3.39 $92,943

Contingency Working Fee Capital Total Cost $1,047,053 $686,401 $7,550,415 $33,184,783 $21,754,469 $239,299,155 $885,599 $580,560 $6,386,155 $885,599 $580,560 $6,386,155 $116,859 $76,607 $842,682 $337,898 $221,511 $2,436,619 $5,774,184 $3,785,299 $41,638,285 $105,396 $69,093 $760,023 $16,241 $10,647 $117,116 $805,808 $528,252 $5,810,773 $515,909 $338,207 $3,720,276 $2,045,078 $1,340,662 $14,747,284 $12,210,303 $8,004,532 $88,049,851 $798,289 $523,323 $5,756,551 $623,656 $408,841 $4,497,255 $1,422,222 $932,346 $10,255,804 $34,705 $22,751 $250,260 $10,969 $7,191 $79,098 $235,047 $154,086 $1,694,949 $26,907 $17,639 $194,026 $16,730 $10,967 $120,640 Total Investment: $440,593,371

Table 15: Summary of equipment costs.

The utilities used in the process were steam, liquid nitrogen, and cooling water. Low pressure steam was created from the synthesis gas cooler so this was subtracted from the low pressure steam used in the process. A summary of utility usage and costs may be found in the economic analysis of annual expenditures on the following pages. The economic analysis shows that the base case design is not profitable. Many of the costs used were estimates or are likely to change over the life of the plant. Some of the economic variables were studied to determine the impact that changes would have on the economics. Alone, none of the variables have the weight to substantially alter the economic analysis. However, together if a group of prices or values of consumption is actually lower than predicted, this could possibly make the project economical. The nitrogen usage or price must substantially decrease. This will be discussed further with optimization schemes. A summary of the annual expenditures, revenues, and discounted cash flow analysis may be seen below.

 Annual Expenditures

Consumption Coal (lb) Oxygen (lb)

Hour 542591.807 276803.21

Day 13022203.36 6643277.043

Year 4753104226 2424796121

Process Water (lb) LP Steam (lb) MP Steam (lb) HP Steam (lb) HHP Steam (lb) Electricity (KW) Cooling Water Makeup (lb) Waste Water Treatment (lb) Bulk Liquid Waste Proc. (lb) Vents/Vapors Processing (lb) MDEA (lb) N2

286000 90528.9728 0 0 157706.4 48299.7347 10000 1331705.53 110202.689 446203.307 11282.0691 5634512.5

6864000 2505360000 2172695.348 793033802 0 0 0 0 3784953.6 1381508064 1159193.633 423105676 240000 87600000 31960932.71 1.1666E+10 2644864.539 965375557 10708879.36 3908740968 270769.6579 98830925.1 135228300 4.9358E+10 Sum

Yearly Value 144,494,368 75,774,879

$ $ $ 300,804 $ 5,310,494 $ $ $ 16,035,361 $ 29,617,397 $ 11,569 $ 1,400,638 $ 176,698,205 $ 575,841,303 $ 74,123,194 $ 1,628,824,874 $ 2,728,433,087

Table 16: Annual expenditures and operating costs.

 Annual Revenues

Income/Credits Methanol (lb)

Hour 456817

Day 10963601.85

Year 4001714674

Yearly Value $ 571,673,524.91

Table 17: Annual income from methanol production.

Capital Investment

Costed Equipment

$

440,593,371

Estimates for pumps/ drums Control Systems Total

$ $ $

44,059,337 48,465,271 533,117,979

Table 18: Estimated capital expenditure required for project.

Discounted Cash Flow

Year 2008 2009 2010 2011 2012 2013 2014 2015 2016 2017 2018 2019 2020 2021 2022 2023 2024 2025 2026 2027 2028

Expenses $533,117,979 $2,826,279,619 $2,910,588,201 $2,997,426,041 $3,086,869,016 $3,178,995,280 $3,273,885,332 $3,371,622,086 $3,472,290,943 $3,575,979,865 $3,682,779,455 $3,792,783,032 $3,906,086,717 $4,022,789,512 $4,142,993,391 $4,266,803,387 $4,394,327,682 $4,525,677,707 $4,660,968,232 $4,800,317,472 $4,943,847,190

Revenues $0 $588,823,731 $606,488,443 $624,683,096 $643,423,589 $662,726,296 $682,608,085 $703,086,328 $724,178,918 $745,904,285 $768,281,414 $791,329,856 $815,069,752 $839,521,844 $864,707,500 $890,648,725 $917,368,186 $944,889,232 $973,235,909 $1,002,432,986 $1,032,505,976

Depreciation $0 $35,541,199 $35,541,199 $35,541,199 $35,541,199 $35,541,199 $35,541,199 $35,541,199 $35,541,199 $35,541,199 $35,541,199 $35,541,199 $35,541,199 $35,541,199 $35,541,199 $35,541,199 $0 $0 $0 $0 $0

Taxes ($213,247,192) ($909,198,835) ($935,856,383) ($963,313,657) ($991,594,650) ($1,020,724,073) ($1,050,727,378) ($1,081,630,783) ($1,113,461,289) ($1,146,246,711) ($1,180,015,696) ($1,214,797,750) ($1,250,623,265) ($1,287,523,547) ($1,325,530,836) ($1,364,678,344) ($1,390,783,798) ($1,432,315,390) ($1,475,092,929) ($1,519,153,794) ($1,564,536,486)

Profit ($319,870,788) ($1,328,257,053) ($1,368,243,376) ($1,409,429,288) ($1,451,850,777) ($1,495,544,911) ($1,540,549,869) ($1,586,904,976) ($1,634,650,736) ($1,683,828,868) ($1,734,482,345) ($1,786,655,426) ($1,840,393,699) ($1,895,744,121) ($1,952,755,055) ($2,011,476,318) ($2,086,175,697) ($2,148,473,085) ($2,212,639,393) ($2,278,730,692) ($2,346,804,729)

Table 19: Annual discounted cash flow analysis.

As can be seen from the annual expenditures, one large cost is bulk vapor waste processing. This primarily comes from the waste of T-502 where methanol is separated from carbon dioxide. To substantially decrease this cost, I recommend using the waste carbon dioxide in production of methanol. Since little carbon monoxide is present, it would be difficult to generate hydrogen using the water-gas shift reaction. So, hydrogen will need to be purchased for this reaction. Using this carbon dioxide to produce methanol would decrease the amount of feed that is required for the process because more methanol could be made from the same amount of coal fed. The produced methanol can be recycled to towers 501 and 502 for refining and purification. The rest of the process, however, would see decreased flow rates due to the decreased feed. This would help to lower both capital and operating costs. Particularly, it might lower the amount of liquid nitrogen that is required to cool the methanol synthesis effluent. This change would decrease costs in multiple ways. Additionally, it would minimize the amount of material and energy that is contained within the process because of the smaller flow rates. This would make the process inherently safer because a smaller amount of hazardous material would be in the inventory.

It is unclear whether this change would be enough to make the project economical. An additional way to improve the process would be to integrate the heat in section 400 of the process flow diagram. Since the recycled gas is cold, about -28°F, and will eventually be heated, it makes sense to use it as a first step in cooling the reactor effluent. This would decrease the amount of liquid nitrogen that is required to cool the effluent as well as the amount of steam required to pre-heat the feed to the reactor. A brief and limited process safety analysis was conducted. The analysis was limited to one hazard for each major unit operation and one initiating event per hazard identified. Many additional safety systems will need to be added when a more detailed design is constructed. All of the major unit operations have the threat of over-pressure and vessel rupture. This is of particular concern with the reactors where runaway is a possibility.

Therefore, rupture disks or relief valves need to be

incorporated into the design. Since the reactions have the potential to runaway, high temperature and pressure alarms should be added. These should be on a separate system from the basic process control system so that they may serve as an independent layer of protection. Control valves have already been incorporated into the design so these can also act as a layer of protection in most cases. It is likely when a more thorough analysis is conducted, more control valves will be added to the process. Though it will not serve as an independent layer of protection because it is tied to the basic process control system, external cooling to the reactors should be added as an additional safeguard against runaway. Due to the large quantities of material in the process, any release could potentially be a category five. For this reason, I recommend invoking inherent safety measures wherever possible. This is why recovering and reusing the carbon dioxide waste is of particular interest. Reducing inventories in the process should be a top concern when further evaluating the project. As was stated, this was a very limited safety analysis. Further analysis should not only be conducted on the remaining equipment, but also on the hazards and initiating events that weren’t studied for the major process equipment. Conclusions & Recommendations

The base case design for the project is not profitable. However, the optimization schemes show potential for the project to become economically sound. These need to be explored further and in greater detail before a decision is made regarding the continuation of the project. The safety of the project should also be explored more thoroughly. A full risk assessment needs to be completed. The risk analysis that has been conducted shows that the unmitigated risks are high. This is primarily due to the quantity of both material and energy contained in the process. The inventory of hazardous material and energy needs to be minimized as much as possible to increase the inherent safety of the process. Some of the larger reactors and towers should be split, if possible, to reduce the consequence of an incident, should one occur. A study of the sensitivity of the economics shows that some values need to be determined more accurately. The MDEA and liquid nitrogen prices were estimated from online resources. These, in particular need to be determined more accurately through contact with a provider of these materials. Before a final decision is made on whether or not a company can capitalize on the opportunity, more analysis needs to be completed.

References 1. C. Higman and M. van der Burgt, Gasification, Elsevier, Amsterdam, 2003. 2. G.H. Graaf and A.A.C.M. Beenackers, “Comparison of two-phase and three-phase methanol synthesis processes,” Chemical Engineering and Processing 35 (1996) pages 413-427.

3. G.H. Graaf, E.J. Stamhuis, and A.A.C.M. Beenackers, “Kinetics of low -pressure methanol synthesis,” Chemical Engineering Science 43 (1988) pages 3185-3195. 4. G.M. Graaf, J.G.M. Winkelmand, E.J. Stamhuis, and A.A.C.M. Beenackers, “Kinetics of the three phase methanol synthesis,” Chemical Engineering Science 43 (1988) pages 2161-2168.

5. Green, Don W. and Robert H. Perry. 2007. Perry’s Chemical Engineers’ Handbook, 8th Ed., McGraw-Hill Companies, New York.

 Appendices

1. Problem Statement 2. Simulation Run Logs 3. Calculations