AIChE National Design Competition

Coal Gasification Study

Process Design Report

By: Kaitlyn D. Kelly

March 7, 2008

Summary

Rising oil prices and increased demand for chemical feedstock have added to the push to find

alternative sources of energy. The gasification of coal to produce 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, and operation of a world-scale methanol 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 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 2149

Nitrogen 0

Oxygen 0

Hydrogen 0

Carbon monoxide 0

Carbon dioxide 495

Argon 0

hydrogen sulfide 0

COS 0

Methane 0

Ammonia 0

MDEA 0

Carbon 0

Methanol 454173

Ethanol 0

Coal 0

Ash 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

$ 44,059,337

Control Systems $ 48,465,271

Total $ 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 Hour Day Year Yearly Value

Coal (lb) 542591.807 13022203.36 4753104226 $ 144,494,368

Oxygen (lb) 276803.21 6643277.043 2424796121 $ 75,774,879

Process Water (lb) 286000 6864000 2505360000 $ 300,804

LP Steam (lb) 90528.9728 2172695.348 793033802 $ 5,310,494

MP Steam (lb) 0 0 0 $ -

HP Steam (lb) 0 0 0 $ -

HHP Steam (lb) 157706.4 3784953.6 1381508064 $ 16,035,361

Electricity (KW) 48299.7347 1159193.633 423105676 $ 29,617,397

Cooling Water Makeup (lb) 10000 240000 87600000 $ 11,569

Waste Water Treatment (lb) 1331705.53 31960932.71 1.1666E+10 $ 1,400,638

Bulk Liquid Waste Proc. (lb) 110202.689 2644864.539 965375557 $ 176,698,205

Vents/Vapors Processing (lb) 446203.307 10708879.36 3908740968 $ 575,841,303

MDEA (lb) 11282.0691 270769.6579 98830925.1 $ 74,123,194

N2 5634512.5 135228300 4.9358E+10 $ 1,628,824,874 Total Expenditures $ 2,728,433,087

Table 3: Annual expenditures and operating costs.

Annual Revenues

Income/Credits Hour Day Year Yearly Value

Methanol (lb) 456817 10963601.85 4001714674 $ 571,673,524.91 Table 4: Annual income from the methanol production.

Discounted Cash Flow

Year Expenses Revenues Depreciation Taxes Profit

2008 $533,117,979 $0 $0 ($213,247,192) ($319,870,788)

2009 $2,826,279,619 $588,823,731 $35,541,199 ($909,198,835) ($1,328,257,053)

2010 $2,910,588,201 $606,488,443 $35,541,199 ($935,856,383) ($1,368,243,376)

2011 $2,997,426,041 $624,683,096 $35,541,199 ($963,313,657) ($1,409,429,288)

2012 $3,086,869,016 $643,423,589 $35,541,199 ($991,594,650) ($1,451,850,777)

2013 $3,178,995,280 $662,726,296 $35,541,199 ($1,020,724,073) ($1,495,544,911)

2014 $3,273,885,332 $682,608,085 $35,541,199 ($1,050,727,378) ($1,540,549,869)

2015 $3,371,622,086 $703,086,328 $35,541,199 ($1,081,630,783) ($1,586,904,976)

2016 $3,472,290,943 $724,178,918 $35,541,199 ($1,113,461,289) ($1,634,650,736)

2017 $3,575,979,865 $745,904,285 $35,541,199 ($1,146,246,711) ($1,683,828,868)

2018 $3,682,779,455 $768,281,414 $35,541,199 ($1,180,015,696) ($1,734,482,345)

2019 $3,792,783,032 $791,329,856 $35,541,199 ($1,214,797,750) ($1,786,655,426)

2020 $3,906,086,717 $815,069,752 $35,541,199 ($1,250,623,265) ($1,840,393,699)

2021 $4,022,789,512 $839,521,844 $35,541,199 ($1,287,523,547) ($1,895,744,121)

2022 $4,142,993,391 $864,707,500 $35,541,199 ($1,325,530,836) ($1,952,755,055)

2023 $4,266,803,387 $890,648,725 $35,541,199 ($1,364,678,344) ($2,011,476,318)

2024 $4,394,327,682 $917,368,186 $0 ($1,390,783,798) ($2,086,175,697)

2025 $4,525,677,707 $944,889,232 $0 ($1,432,315,390) ($2,148,473,085)

2026 $4,660,968,232 $973,235,909 $0 ($1,475,092,929) ($2,212,639,393)

2027 $4,800,317,472 $1,002,432,986 $0 ($1,519,153,794) ($2,278,730,692)

2028 $4,943,847,190 $1,032,505,976 $0 ($1,564,536,486) ($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 the 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 Cost $/short ton

Transportation $/short ton

TX Lignite $ 15.20 $ 3.90

WY Sub-Bituminous $ 10.60 $ 10.20

IL Bituminous $ 32.00 $ 6.90

Pre-Processing $ 40.00 /short ton

Oxygen $ 70.00 /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 $26.00 /metric ton

HP Steam (400psig) $22.00 /metric ton

MP Steam (100psig) $19.00 /metric ton

LP Steam (50psig) $15.00 /metric ton

Electricity $0.07 /KWH

Condensate $0.75 /metric ton

Cooling Water makeup $1.10 /Mgals

Process Water $1.00 /Mgals

Demin Water $3.00 /Mgals

Potable Water $2.50 /Mgals

Waste Water Treatment $1.00 /Mgals

TOC in Waste Water $0.70 /lb TOC

Instrument Air $0.45 /MCF

Bulk Liquids Waste Processing $410.00 /metric ton

Bulk Solids Waste Processing $325.00 /metric ton

Vents/Vapors Processing $330.00 /metric ton

Inert Gas $0.35 /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 lnKp = -4.33 + 4577.8/T(K)

CO +2H2 ↔ CH3OH 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 Martin Lake TX Lignite PRB Sub-Bituminous Illinois Bituminous

As Received $59.10/ton $60.80/ton $78.90/ton

Dry Basis $90.68/ton $67.94/ton $90.69/ton

Carbon Basis $140.28/ton C $101.64/ton C $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 water-

gas 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 # D-101 E-101 E-201 E-301 E-401

Equipment Name Gasifier Syngas cooler Rich/lean exchanger D301 Preheat D401 Preheat

Utility Used Gas Cooling Water N/A Steam Steam

Duty (MMBtu/hr) 6083 -125.2 1042 21.5 146.3

Equipment # E-402 E-501 E-502 E-503 E-504

Equipment Name N2 Cooler T501 Condenser T501 Reboiler T502 Condenser T502 Reboiler

Utility Used N2 Cooling Water Steam Cooling Water Steam

Duty (MMBtu/hr) -583 -33.5 20.1 -21.9 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

# Name trays spacing (ft)

Passes diameter (ft)

height (ft) op. pressure (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 Table 11: Equipment summaries, towers.

Reactors

# Name volume (ft3)

length (ft)

diameter (ft)

design P (psia)

design T (psia)

Allowable Stress (psia)

D-101 Gasifier 10000 110 11 508 3000

D-301 Water-Gas Shift 26750 37.8 30 95 710 21600

D-401 MeOH Synthesis 5900 32 16 80 600 16920 Table 12: Equipment summaries, reactors.

Heat Exchangers

# Name req'd area (ft2) Utility

Duty (MMBtu/hr) ΔT (°F)

E-101 Syngas cooler (2) 12018 CW -125.2 350

E-201 Rich/Lean Exch. 18635 Process 1042 71

E-301 D301 Preheat 748.9 Steam 21.5 75

E-401 D401 Preheat (2) 22223 Steam 146.3 214

E-402 N2 Cooler 45386 N2 -583 900

E-501 T-501 Cond 1952 CW -33.5 4

E-502 T-501 Reboil 899 Steam 20.1 0.5

E-503 T-502 Cond 1440 N2 -21.9 80

E-504 T-502 Reboil 780 Steam 45.7 0 Table 13: Equipment summaries, heat exchangers.

Pump/Compressor/Expander

# Name hp ΔP (psi)

J-102 Expander -35216 -315

J-203 (2) Lean MDEA Pump 1086 90

C-401 Compressor 96900 120 Table 14: Equipment summaries, pressure exchangers.

Equipment Costs

Equipment # Equipment Name

Cost Index

Module Factor Base Cost

Contingency Fee

Working Capital Total Cost

D-101 Gasifier 754.4 4.34 $5,816,961 $1,047,053 $686,401 $7,550,415

J-102 Expander 941 3.21 $184,359,904 $33,184,783 $21,754,469 $239,299,155

T-201a AGR Absorber 754.4 4.34 $4,919,996 $885,599 $580,560 $6,386,155

T-201b AGR Absorber 754.4 4.34 $4,919,996 $885,599 $580,560 $6,386,155

E-101 Syngas cooler (2) 806.7 3.39 $649,216 $116,859 $76,607 $842,682

E-201 Rich/Lean Exch. 806.7 3.39 $1,877,210 $337,898 $221,511 $2,436,619

T-202 Stripper 754.4 4.34 $32,078,802 $5,774,184 $3,785,299 $41,638,285

J-203 (2) Lean MDEA Pump 941 3.48 $585,534 $105,396 $69,093 $760,023

E-301 D301 Preheat 806.7 3.39 $90,228 $16,241 $10,647 $117,116

D-301 H2O-Gas Shift 754.4 4.34 $4,476,713 $805,808 $528,252 $5,810,773

E-401 D401 Preheat (2) 806.7 3.39 $2,866,160 $515,909 $338,207 $3,720,276

D-401 (5) MeOH Syn React. 754.4 4.34 $11,361,544 $2,045,078 $1,340,662 $14,747,284

C-401 Compressor 941 3.21 $67,835,016 $12,210,303 $8,004,532 $88,049,851

E-402 N2 Cooler 806.7 3.39 $4,434,939 $798,289 $523,323 $5,756,551

F-401 Separator 754.4 2.39 $3,464,757 $623,656 $408,841 $4,497,255

T-501 Refining 754.4 4.34 $7,901,236 $1,422,222 $932,346 $10,255,804

E-501 T-501 Cond 806.7 3.39 $192,804 $34,705 $22,751 $250,260

E-502 T-501 Reboil 806.7 3.39 $60,938 $10,969 $7,191 $79,098

T-502 CO2 Removal 754.5 2.39 $1,305,816 $235,047 $154,086 $1,694,949

E-503 T-502 Cond 806.7 3.39 $149,481 $26,907 $17,639 $194,026

E-504 T-502 Reboil 806.7 3.39 $92,943 $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 Hour Day Year Yearly Value

Coal (lb) 542591.807 13022203.36 4753104226 $ 144,494,368

Oxygen (lb) 276803.21 6643277.043 2424796121 $ 75,774,879

Process Water (lb) 286000 6864000 2505360000 $ 300,804

LP Steam (lb) 90528.9728 2172695.348 793033802 $ 5,310,494

MP Steam (lb) 0 0 0 $ -

HP Steam (lb) 0 0 0 $ -

HHP Steam (lb) 157706.4 3784953.6 1381508064 $ 16,035,361

Electricity (KW) 48299.7347 1159193.633 423105676 $ 29,617,397

Cooling Water Makeup (lb) 10000 240000 87600000 $ 11,569

Waste Water Treatment (lb) 1331705.53 31960932.71 1.1666E+10 $ 1,400,638

Bulk Liquid Waste Proc. (lb) 110202.689 2644864.539 965375557 $ 176,698,205

Vents/Vapors Processing (lb) 446203.307 10708879.36 3908740968 $ 575,841,303

MDEA (lb) 11282.0691 270769.6579 98830925.1 $ 74,123,194

N2 5634512.5 135228300 4.9358E+10 $ 1,628,824,874

Sum $ 2,728,433,087 Table 16: Annual expenditures and operating costs.

Annual Revenues

Income/Credits Hour Day Year Yearly Value

Methanol (lb) 456817 10963601.85 4001714674 $ 571,673,524.91 Table 17: Annual income from methanol production.

Capital Investment

Costed Equipment $ 440,593,371

Estimates for pumps/ drums

$ 44,059,337

Control Systems $ 48,465,271

Total $ 533,117,979 Table 18: Estimated capital expenditure required for project.

Discounted Cash Flow

Year Expenses Revenues Depreciation Taxes Profit

2008 $533,117,979 $0 $0 ($213,247,192) ($319,870,788)

2009 $2,826,279,619 $588,823,731 $35,541,199 ($909,198,835) ($1,328,257,053)

2010 $2,910,588,201 $606,488,443 $35,541,199 ($935,856,383) ($1,368,243,376)

2011 $2,997,426,041 $624,683,096 $35,541,199 ($963,313,657) ($1,409,429,288)

2012 $3,086,869,016 $643,423,589 $35,541,199 ($991,594,650) ($1,451,850,777)

2013 $3,178,995,280 $662,726,296 $35,541,199 ($1,020,724,073) ($1,495,544,911)

2014 $3,273,885,332 $682,608,085 $35,541,199 ($1,050,727,378) ($1,540,549,869)

2015 $3,371,622,086 $703,086,328 $35,541,199 ($1,081,630,783) ($1,586,904,976)

2016 $3,472,290,943 $724,178,918 $35,541,199 ($1,113,461,289) ($1,634,650,736)

2017 $3,575,979,865 $745,904,285 $35,541,199 ($1,146,246,711) ($1,683,828,868)

2018 $3,682,779,455 $768,281,414 $35,541,199 ($1,180,015,696) ($1,734,482,345)

2019 $3,792,783,032 $791,329,856 $35,541,199 ($1,214,797,750) ($1,786,655,426)

2020 $3,906,086,717 $815,069,752 $35,541,199 ($1,250,623,265) ($1,840,393,699)

2021 $4,022,789,512 $839,521,844 $35,541,199 ($1,287,523,547) ($1,895,744,121)

2022 $4,142,993,391 $864,707,500 $35,541,199 ($1,325,530,836) ($1,952,755,055)

2023 $4,266,803,387 $890,648,725 $35,541,199 ($1,364,678,344) ($2,011,476,318)

2024 $4,394,327,682 $917,368,186 $0 ($1,390,783,798) ($2,086,175,697)

2025 $4,525,677,707 $944,889,232 $0 ($1,432,315,390) ($2,148,473,085)

2026 $4,660,968,232 $973,235,909 $0 ($1,475,092,929) ($2,212,639,393)

2027 $4,800,317,472 $1,002,432,986 $0 ($1,519,153,794) ($2,278,730,692)

2028 $4,943,847,190 $1,032,505,976 $0 ($1,564,536,486) ($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