METHANOL PRODUCTIONIN

TRINIDAD & TOBAGO

Final Report: Phase II

University of California, Davis

Date of Report: June 07, 2006

Design Group One

Elton AmirkhasRaj Bedi

Steve HarleyTrevor Lango

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Executive SummaryThis report is the first phase of a final report designed to investigate the feasibility of methanol production in Trinidad and Tobago. Specifically, this report outlines a proposed four-stage process for producing methanol:

STAGE 1: Syngas production STAGE 2: Upstream processing STAGE 3: Methanol production STAGE 4: Downstream processing

The proposed design produces 5,116 MTPD of 99.85 wt% methanol. As designed, the total bare module cost of the plant is $372 million. The inside battery limit and outside battery limit costs are $349 million and $23 million respectively. Total capital investment includes the direct permanent investment of $512 million and is $779 million. The calculated BTROI is 42% with annual net earnings of roughly $203 million per year. The NPV is $1.2 billion in the last year of production and suggests a profitable venture.

The project managers expressed to us their concern over the current price of oxygen. Given that our plant consumes 1.78 trillion lbs per year, which costs a total of $41.5 million, they instructed us to lead an investigation into possible onsite production possibilities. In the course of our investigation we found that VPSA, PSA, and membrane separations were strongly lacking in both purity requirements and desirable flow rates. This then lead us to the conclusion that the Claude process, a highly energy-optimized cryogenic separations technique, would suit our needs. Not only will this process supply the required purity of 99.5% mole basis, but it is robust enough to supply the large flow rates needed. The estimated total capital investment is $72.2 million. Due to the difficulty encountered in costing cryogenic process units it is recommended that a more detailed capital cost analysis be performed. As designed, the on-site oxygen production plant is able to produce oxygen at $0.0194/lb which results in a $7.7 million annual savings in feedstock oxygen costs. Please find the on-site oxygen production plant report immediately following the methanol production plant report.

As designed it is worthwhile to invest in a methanol production plant with on-site oxygen production.

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Table of Contents

Executive Summary ........................................................................................................................ 21 Introduction............................................................................................................................. 6

1.1 Overview......................................................................................................................... 61.2 Current Manufacturing Methods..................................................................................... 6

1.2.1 Syngas Production .................................................................................................. 61.2.2 Methanol Production............................................................................................... 7

1.3 Selected Production Method ........................................................................................... 71.4 Production Level & Plant Location ................................................................................ 7

1.4.1 Production Level..................................................................................................... 71.4.2 Plant Location ......................................................................................................... 7

1.5 Market Considerations .................................................................................................... 71.6 Environmental Issues ...................................................................................................... 7

1.6.1 Chemical Toxicity................................................................................................... 71.6.2 Potential Safety Problems....................................................................................... 8

2 Process Description................................................................................................................. 92.1 Block Flow Diagram....................................................................................................... 9

2.1.1 Chemical Reactions ................................................................................................ 92.1.2 Separations.............................................................................................................. 9

2.2 Detailed Flow Diagram................................................................................................... 92.2.1 STAGE 1: Syngas Production ................................................................................ 9

2.2.1.1 Natural Gas Furnace (F-100) .............................................................................. 92.2.1.2 Steam Methane Reformer (R-100)...................................................................... 92.2.1.3 Oxygen Blown Reformer (R-200) ...................................................................... 9

2.2.2 STAGE 2: Upstream Processing........................................................................... 102.2.2.1 Steam Generator (E-100) .................................................................................. 102.2.2.2 Syngas Cooler (C-100) ..................................................................................... 102.2.2.3 Flash Unit (U-200)............................................................................................ 102.2.2.4 Water Mixer (M-200) ....................................................................................... 102.2.2.5 Water Make-up Pump (P-100).......................................................................... 102.2.2.6 Syngas Compressor (CMP-200) ....................................................................... 10

2.2.3 STAGE 3: Methanol Production........................................................................... 102.2.3.1 Feed Splitter (S-100)......................................................................................... 102.2.3.2 Methanol Synthesis Reactor (R-300)................................................................ 102.2.3.3 Product Mixer (M-300)..................................................................................... 10

2.2.4 STAGE 4: Downstream Processing...................................................................... 112.2.4.1 Product Cooler (C-200)..................................................................................... 112.2.4.2 Syngas Separator (U-200)................................................................................. 112.2.4.3 Depressurizer (V-100) ...................................................................................... 112.2.4.4 Distillation (D-100)........................................................................................... 112.2.4.5 Final Product Mixer (M-100)............................................................................ 112.2.4.6 Final Product Cooler (C-300) ........................................................................... 11

3 Energy Balance & Utility Requirements .............................................................................. 123.1 Energy Requirements.................................................................................................... 123.2 Process Integration........................................................................................................ 12

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4 Equipment List & Unit Descriptions .................................................................................... 134.1 Reactors......................................................................................................................... 13

4.1.1 Steam Methane Reformer (R-100)........................................................................ 134.1.2 Oxygen Blown Reformer (R-200) ........................................................................ 154.1.3 Methanol Synthesis Reactor (R-300).................................................................... 16

4.1.3.1 Kinetics ............................................................................................................. 164.1.3.2 Maximum Conversion ...................................................................................... 184.1.3.3 Catalyst ............................................................................................................. 194.1.3.4 Operating Temperature Sensitivity ................................................................... 194.1.3.5 Coolant.............................................................................................................. 204.1.3.6 Temperature Profile .......................................................................................... 20

4.2 Upstream Processing (H-100, V-200, CMP-200)......................................................... 214.2.1 Water Removal ..................................................................................................... 214.2.2 CMP-200............................................................................................................... 224.2.3 C-100..................................................................................................................... 224.2.4 V-200 .................................................................................................................... 22

4.3 Downstream Processing (H-200, V-300, G-100, D-100) ............................................. 234.3.1 CO2 Removal ........................................................................................................ 234.3.2 Recycle & Conversion .......................................................................................... 244.3.3 Flash Vessel (U-300) ............................................................................................ 244.3.4 Coolers (C-200, C-300) ........................................................................................ 244.3.5 Distillation (D-100)............................................................................................... 24

4.4 Methanol Storage .......................................................................................................... 245 Equipment Cost Summary .................................................................................................... 25

5.1 Pump Costs ................................................................................................................... 255.2 Compressor Costs ......................................................................................................... 255.3 Furnace Costs................................................................................................................ 255.4 Storage Tank Costs ....................................................................................................... 255.5 Reactor Costs ................................................................................................................ 255.6 Heat Exchanger Costs ................................................................................................... 255.7 Separation Vessel Costs................................................................................................ 25

6 Fixed Capital Investment Summary...................................................................................... 276.1 Bare Module Costs........................................................................................................ 276.2 Direct Permanent Investment & Total Capital Investment........................................... 27

7 Other Important Considerations............................................................................................ 287.1 Health & Safety............................................................................................................. 287.2 Process Control & Instrumentation............................................................................... 287.3 Environmental............................................................................................................... 28

7.3.1 Chemical Toxicity................................................................................................. 287.3.2 Potential Safety Problems..................................................................................... 287.3.3 Required Permits................................................................................................... 29

8 Operating Cost & Economic Analysis.................................................................................. 308.1 Cost Sheet ..................................................................................................................... 308.2 Working Capital............................................................................................................ 318.3 Total Capital Investment............................................................................................... 318.4 Profitability Measures................................................................................................... 32

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8.4.1 Return on Investment (ROI) ................................................................................. 328.4.2 Net Present Value (NPV)...................................................................................... 338.4.3 Cash Flows (CF) ................................................................................................... 338.4.4 Depreciation Schedule (MACRS)......................................................................... 348.4.5 Investors Rate of Return (IRR) ............................................................................. 34

9 Conclusions & Recommendations........................................................................................ 3610 Acknowledgements........................................................................................................... 3711 References......................................................................................................................... 3812 Appendices........................................................................................................................ 39

12.1 Appendix I: Detailed Equipment Costing..................................................................... 3912.1.1 I.1 Heat Exchanger Sizing Technique .................................................................. 3912.1.2 I.2 Flash Unit Sizing Procedure ............................................................................ 4512.1.3 II.3 Maximum Thermodynamically Attainable Conversion................................. 4612.1.4 I.4 General Reactor Sizing Techniques ................................................................ 47

12.2 Appendix II: Upstream Processing ............................................................................... 5312.3 Appendix III: Kinetic Models....................................................................................... 5512.4 Appendix IV: Example Detailed Equipment Costing................................................... 57

12.4.1 IV.1 Pumps ........................................................................................................... 5712.4.2 IV.2 Storage Tanks ............................................................................................... 5912.4.3 IV.3 Compressors ................................................................................................. 6012.4.4 IV.4 Reactors ........................................................................................................ 6112.4.5 IV.5 Furnaces........................................................................................................ 6412.4.6 IV.6 Heat Exchangers ........................................................................................... 6512.4.7 IV.7 Flash Vessels ................................................................................................ 6612.4.8 IV.8 Distillation Columns..................................................................................... 68

12.5 Appendix V: Direct Permanent Investment & Total Capital Investment ..................... 70

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1 Introduction

1.1 OverviewMethanol is a primary raw product for industries producing formaldehyde, methyl tertiary butyl ether (MTBE), and acetic acid9. Methanol is also consumed in the manufacture of chloromethanes, methylamines, and fuels. It is generally used as a solvent and as antifreeze, being a component in paint strippers, car windshield washer compounds, and as a deicer for natural gas pipelines.

In 2004, approximately 34% of global methanol production was used to produce formaldehyde, 21% for MTBE and other fuel additives, and 9% for acetic acid9. Worldwide consumption of methanol increased on the order of 1% from 2003 to 2008 and is projected to increase 2% from 2008 to 20133. These estimated growths do not reflect new demands associated with new technologies requiring methanol such as direct methanol fuel cells1. Therefore, there is a higher potential for profit.

1.2 Current Manufacturing MethodsContemporary production techniques convert natural gas (mostly methane) to syngas4, which is in turn converted to methanol. The general flowsheet is given in Figure 1.

Figure 1: General flowsheet for methanol production.

1.2.1 Syngas ProductionCurrent methods of syngas production include steam reforming, partial oxidation, carbon dioxide reforming, autothermal reforming, and coal gasification. The raw materials required are methane, steam, and oxygen. The primary byproduct is carbon dioxide. Eqns. (1) – (3) list the principal chemical reactions for syngas production.

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CH4 H2OCO 3H2

CH4 1

2O2 CO 2H2

1

2O2 H2 H2O

(1)

(2)

(3)

1.2.2 Methanol ProductionMethanol is produced from syngas via the Fischer-Tropsch Process, given by Eqn. (4). There are no significant byproducts or intermediates.

CO 2H2 CH3OH (4)

1.3 Selected Production MethodThe selected method for syngas production is a steam methane reformer (SMR) and an oxygen blown reformer (OBR) in series. The OBR was deemed necessary to completely consume the methane while minimizing the production of carbon dioxide. Syngas is converted to methanol in a parallel tube plug flow reactor (methanol synthesis reactor or MSR).

1.4 Production Level & Plant Location

1.4.1 Production LevelA production level of 5,000 MTPD was selected to reflect current anticipated market demand.

1.4.2 Plant LocationTrinidad and Tobago was selected as an ideal location for methanol production due to its large natural gas reserves.

1.5 Market ConsiderationsBecause methanol is easily transported, methanol production could become an important outlet for enhancing the value of natural gas. With growing worldwide natural gas reserves and recent advances in methanol technology, methanol may be poised to take on renewed importance in the fuels and petrochemical markets.

1.6 Environmental Issues

1.6.1 Chemical ToxicityMost commonly humans are exposed to methanol through skin contact and vapor inhalation. Although carcinogenicity of methanol has not been determined, exposure to methanol has been linked to reproductive defects in rats8. Methanol is known to cause headaches, dizziness, giddiness, insomnia, nausea, gastric disturbances, conjunctivitis, blurred vision, and blindness in humans. High doses of methanol may be fatal. OSHA’s regulatory concentration of methanol

for human exposure without adverse effects in an 8 hour day is 260mg

m3 or 198 ppm5.

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1.6.2 Potential Safety ProblemsMethanol is readily degraded in the environment by photo oxidation9 and biodegradation processes5. Half-lives of 7 – 18 days have been reported for the atmospheric reaction of methanol with hydroxyl radicals. Methanol is readily degradable under both aerobic and anaerobic conditions in a wide variety of environmental media including fresh and salt water, sediments and soils, ground water, aquifier material and industrial wastewater. Methanol is of low toxicity to aquatic and terrestrial organisms, and effects due to environmental exposure to methanol are unlikely to be observed except in the case of a spill5, 9.

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2 Process Description

2.1 Block Flow Diagram

Figure 3: Simplified natural gas-to-methanol flowsheet.

2.1.1 Chemical ReactionsMethane, steam, and oxygen are catalytically reacted in the syngas production stage to produce hydrogen and carbon monoxide. The resulting syngas is catalytically reacted in the methanol synthesis reactor block to produce methanol.

2.1.2 SeparationsUpstream processing removes water from the process; downstream processing removes methanol. Methanol is separated from the process via a two-stage separation. First light gases are removed in a flash unit. Secondly, methanol is separated from carbon dioxide and any remaining water in a distillation column.

2.2 Detailed Flow Diagram

2.2.1 STAGE 1: Syngas Production

2.2.1.1 Natural Gas Furnace (F-100)A fired furnace is used to preheat the natural gas being fed to the steam methane reformer (to maximize the rate of reaction).

2.2.1.2 Steam Methane Reformer (R-100)Steam methane reforming was selected for syngas production because it is a well understood process and produces syngas with the desired H2-to-CO ratio.

2.2.1.3 Oxygen Blown Reformer (R-200)By selectively operating the steam methane reformer at an optimized (reduced) conversion the production of carbon dioxide is minimized. The oxygen blown reformer is used to completely consume the methane fed to the steam methane reformer.

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2.2.2 STAGE 2: Upstream Processing

2.2.2.1 Steam Generator (E-100)The excess heat in the OBR effluent is used to produce steam by exchanging heat with the effluent with process water. This also serves to cool the syngas so that water can be flashed out downstream.

2.2.2.2 Syngas Cooler (C-100)The syngas is cooled to the optimum downstream flash conditions for water removal.

2.2.2.3 Flash Unit (U-200)Water is flashed out of the syngas stream to optimize downstream product separations. The liquid water is converted to steam by exchanging heat with the OBR effluent thereby reducing the amount of utility water required.

2.2.2.4 Water Mixer (M-200)Recovered liquid water and make-up utility water are mixed before being converted to steam.

2.2.2.5 Water Make-up Pump (P-100)Utility water must be pumped to match the operating conditions of the recovered water before being mixed.

2.2.2.6 Syngas Compressor (CMP-200)The syngas process stream is brought to the optimal operating temperature and pressure of the methanol synthesis reactor using an inter-stage compressor. The inter-stage compressor has the advantage of excellent temperature control and is an efficient method for compression of gases.

2.2.3 STAGE 3: Methanol Production

2.2.3.1 Feed Splitter (S-100)The syngas feed is equally split to each methanol synthesis reactor unit.

2.2.3.2 Methanol Synthesis Reactor (R-300)The methanol synthesis reactor stage is comprised of two parallel tube plug flow reactors operating in parallel (to optimize the residence time through each). These reactors are operated as heat exchangers to maximize heat transfer characteristics and ensure adequate temperature control.

2.2.3.3 Product Mixer (M-300)The effluent from each methanol synthesis reactor unit is mixed before being fed to the methanol processing stage.

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2.2.4 STAGE 4: Downstream Processing

2.2.4.1 Product Cooler (C-200)The methanol product stream must be cooled for optimal downstream separations.

2.2.4.2 Syngas Separator (U-200)Light gases are removed from the methanol product stream to reduce the required downstream separation equipment duties.

2.2.4.3 Depressurizer (V-100)The methanol product stream pressure must be decreased to the final product specification.

2.2.4.4 Distillation (D-100)The major contaminants (carbon dioxide and water) must be removed from the methanol product stream to produce methanol with the specified product purity.

2.2.4.5 Final Product Mixer (M-100)Methanol from the two distillation liquid effluent streams is combined to produce the final product.

2.2.4.6 Final Product Cooler (C-300)The mixed distillation effluent is cooled to the final product specification temperature.

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3 Energy Balance & Utility Requirements

3.1 Energy RequirementsThe energy requirements of the process are tabulated in Table 1.

Table 1: Energy duties for all process units with corresponding vehicles for satisfaction.

Process Unit Demand (in Watts) Vehicle For Satisfaction

ReactorsR-100 0.34576E+09 Combustion of natural gasR-300 -0.17666E+09 Dowtherm Q

FurnacesF-100 0.55597E+08 Combustion of natural gas

Heat ExchangersE-100 1.791577830176E+08 Exchange streams 6 & 31

enthalpy

CoolersC-100 -0.27290E+09 Cooling waterC-200 -0.84559E+08 Cooling waterC-300 -0.75508E+07 Cooling water

Pumps & CompressorsP-100 123,705 ElectricityCMP-200 (cooling/electrical) -0.116222+08/0.216511+08 Cooling water/Electricity

VesselsD-100 (reboiler/condenser) 0.268502+08/-0.105732+08 Combustion/Cooling water

3.2 Process IntegrationThe measures adopted to improve the plant economics by energy and mass conservation are:

1. Steam generation using OBR effluent energy2. Recovery and recycle of separated liquid water from syngas (as steam)3. Combustion of recovered light gases as furnace gases

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4 Equipment List & Unit Descriptions

4.1 Reactors

4.1.1 Steam Methane Reformer (R-100)The steam methane reformer is designed to generate syngas via Eqn. (5).

4 2 2 2980 0 3 H 206 /CH H C H kJ mole (5)

This reaction is endothermic; therefore, during its operation it will be heated via the combustion of natural gas. If we assume that 90% of the combusted energy is transferred to the optimized R-100, it will require 13.1 m3/s of combusted natural gas. Furthermore, catalyst (Raschig ring, 5/8" L x 5/8" D OD, with 3/16" hole) with a void fraction of 0.45 will be used to increase the reaction rate.

The optimization goals for the R-100 were to minimize carbon dioxide production (as it presented significant downstream separation issues and kinetic data was not available for modeling its conversion to methanol) and to maximize the production of syngas through the varying of temperature, pressure, and the steam-to-methane ratio. We preformed a sensitivity analysis by varying the aforementioned parameters. Figures 4 – 6 represent the results of this sensitivity analysis.

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0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

1400 1450 1500 1550 1600 1650 1700 1750 1800

SMR Temp [F]

0

1

2

3

4

5

6

7

H2

-to

-CO

Ra

tio

(S

yn

ga

s)

SYNCH4 SYNH2 SYNCO SYNCO2 H2CO

Figure 4: Effect of temperature on R-100.

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

200 250 300 350 400 450 500 550 600

SMR Pres [psig]

3.6

3.7

3.8

3.9

4

4.1

4.2

4.3

4.4

H2

-to

-CO

Ra

tio

(S

yn

ga

s)

SYNCH4 SYNH2 SYNCO SYNCO2 H2CO

Figure 5: Effect of pressure on R-100.

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0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

0 0.5 1 1.5 2 2.5 3 3.5 4

H20-to-CH4 Ratio (Feed)

0

1

2

3

4

5

6

H2

-to

-CO

Ra

tio

(S

yn

ga

s)

SYNCH4 SYNH2 SYNCO SYNCO2 H2CO

Figure 6: Effect of steam-to-methane ratio on R-100.

As a result of the sensitivity analysis, we chose to operate the R-100 at 3.55MPa and 1158.2K at a steam-to-methane ratio of 1.14.

The R-100 reactor was modeled as an RGIBBS unit in Aspen. The limitations of this model are that it only tells us what the reactor would produce if it were infinitely long. In reality the reactor length would be specified with appropriate kinetics data. A procedure on how this reactor would be sized if the kinetics were known may be found in Appendix I.4. Kinetic data should be obtained to properly size the R-100.

4.1.2 Oxygen Blown Reformer (R-200)The OBR is designed to:

Lower the H2-to-CO ratio via Eqn. (6):

OHOH 222 molekJH /242298 (6)

Partially oxidize methane via Eqn. (7):

4 2 2 298

10 2 H 36 /

2CH O C H kJ mole (7)

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In optimizing this reactor our goal was to adjust the hydrogen-to-carbon monoxide ratio produced in the SMR and consume remaining methane. As demonstrated previously, it is not beneficial to consume large amounts of methane in the R-100 reactor as this also results in significant carbon dioxide production.

After optimizing this unit, we determined that the cost of changing any of the parameters from the R-100 reactor to this reactor were large as compared to any conversion benefit we produced. For example, the cost of cooling the OBR feed was greater than the benefit produced by the small increase in conversion in the R-200. Thus we operated R-200 at the same parameters as the R-100.

The R-200 was modeled as an adiabatic RSTOIC reactor in Aspen. As with the R-100 reactor, the R-200 reactor has no kinetic data associated with its operation; thus, empirical sizing procedures can be found in Appendix I.4.

4.1.3 Methanol Synthesis Reactor (R-300)

4.1.3.1 Kinetics

OHCHHCO 322 molekJH /5.90298 (8)

Kinetic data for the proprietary catalyst was made available to us in the form of the dependence on the rate of production of methanol on hydrogen, carbon monoxide, and methanol partial pressures. Thus in modeling the MSR, only these components can be taken into account. In the process of fitting the rate data, 17 kinetic models were attempted assuming various rate determining steps (Appendix III). It was found that a Langmuir-Hinshelwood-Hougen-Watson (LHHW) model, which assumes a reversible reaction with all components adhering to the catalyst associatively, gave the best non-linear regression fit. Eqns. (9) – (14) are the result of this non-linear regression.

3))()()(1(

),(*)(

3322

3

2

3

OHCHOHCHHHCOCO

eq

OHCH

HCO

nOHCH yPTKPTKPTK

PTK

PPP

TTkr

(9)

RT

678-0.191)(ln Tk (10)

392.0TT n (11)

4.12ln5705

103)(ln T

TKCO (12)

4.19ln9496

155)(ln2

T

TK H (13)

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76.4ln1692

7.46)(ln3

T

TK OHCH (14)

This model yielded an r-squared value of 0.999 when fit to the given data.

While the regression may suggest a very good fit, it is simply that; a good fit to the data that was given. With this in mind we offer sever discretions as to the validity of this model:

Several literature articles suggest that both carbon dioxide and water play key roles in the catalysis of methanol15. Catalyst analysis showed that the catalyst surface must contain hydroxyl groups for the initiation of the advanced intermediate pathways13. These groups originate from water vapor in the synthesis stream. Furthermore the presence of carbon dioxide or water will compete for sites on the catalyst thus reducing the predicted rate of reaction. Consequently we suggest obtaining more data with respect to water and carbon dioxide partial pressures such that we may investigate the importance of this site competition mechanism.

It has been found that carbon dioxide may too be used to produce methanol via:

2 2 3 23CO H CH OH H O 14(15)

Given the fact that our aspen simulations have shown that the steam reformer produces significant amounts of carbon dioxide, ignoring this reaction pathway for methanol production would inefficiently use our syngas. Thus we suggest additional rate information be obtained for this reaction such that we may maximize the conversion of all syngas components.

Another possibility to remove carbon dioxide, but not waste the carbon would be to utilize a carbon dioxide reformer8. This reaction is as follows:

2 4 22 2CO CH CO H (16)

Given the fact that our methanol reactor cannot convert carbon dioxide to methanol, we can convert the carbon dioxide to more usable syngas via this reaction. There are catalysts available that selectively promote this reaction and we suggest this as a possibility to beneficially remove carbon dioxide.

The data set given was too small to really obtain a proper fit to a model. These models have anywhere from 3 to 9 unknown parameters in them. Thus finding some combination of these parameters to fit the 17 data points is a trivial task. But just because a fit was obtained, does not mean that the kinetic model physically describes which mechanism is dominant. To statistically fix this, more data points are needed.

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While our analysis has shown that the best kinetic models have very high r-squared values, we may only be certain they are accurate in the temperature range of 475-495°K. We can only assume that the rate law behaves similarly over larger temperature ranges. Furthermore given the large investment required to build this plant we feel that this uncertainty in our model will carry over to other parts of our simulation thereby entering uncertainty into our plant wide cost estimates. We highly recommend obtaining moretemperature data points for regression.

In Aspen, the MSR is modeled as a plug flow reactor (PFR). While ultimately Aspen is used to determine the exact size of the PFR, custom algorithms were used to obtain estimates of the required size. Eqn. (17)6 takes into account the volume change on reaction and uses an average temperature to account for the temperature profile through the reactor.

dX

ratioX

XTK

ratioX

XratioTK

ratioX

XTK

PTKratioX

X

ratioX

Xratio

ratioX

X

TTkFVX

OHCHHCO

eqnAO

0

3)21

)(21

2)(

21

1)(1(

),(21

21

2

21

1

*)(

32

(17)

Note: “ratio” denotes the molar H2-to-CO ratio.

4.1.3.2 Maximum ConversionEver present in the optimization of the MSR is the trade-off between the maximum thermodynamically attainable conversion and the kinetic reaction rate. Thermodynamically the maximum conversion is a function of temperature, pressure, and reactant ratios, which according to Le Châtelier’s principles will favor low temperature, high pressure, and the excess of any one reactant, while the reaction rate favors high temperatures. Thus to pick the proper operating conditions we need to know just how this equilibrium constant behaves as a function of its inputs. Literature equilibrium data was obtained7 and regressed. Using techniques outlined in Fogler we were able to write an algorithm (Appendix I.3) to find the maximum conversions listed in Table 2.

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Table 2: Maximum conversion as a function of H2-to-CO ratio and temperature.

Temperature (K)

Pressure (MPa)

H2:CO ratio

Maximum Conversion

400 7 2 0.95400 7 5 0.99400 4 2 0.92450 7 2 0.83450 7 5 0.99450 4 2 0.75500 7 2 0.61500 7 5 0.87500 4 2 0.46

It is evident that of the three variables, pressure has the lowest effect on the maximum conversion. However, pressure has a large effect on the cost of the reactor, thus a low operating pressure was chosen. What is most surprising about this analysis is the very large effect that the syngas ratio has. At 500K, a stoichiometric H2:CO ratio, and 7 MPa, the maximum conversion is 0.46. By changing this ratio to 5, the maximum conversion increases to 0.87. Thus to overcome a thermodynamic barrier, excess hydrogen should be used. This becomes important in the downstream processing section where the use of a recycle stream is considered.

4.1.3.3 CatalystThermal degradation of the catalyst occurs at 500K. Given that the reaction in the MSR is highly exothermic (Eqn. 8), the reactor requires strategic cooling to prevent the buildup of thermal energy inside the reactor. This will solve the problem of heat buildup along the length of the tubes, but the temperature profile across the diameter is another issue. Thus the task becomes determining the proper tube diameter such that the tube thermal resistance is negligible as compared to the fluid phase resistance. In such a condition the temperature gradient across the diameter of the tube may be considered zero. To do this we will test the Biot number, where the fluid phase resistance is approximated by a shell side heat transfer coefficient and the tube/catalyst thermal resistances by their thermal conductivities. When the Biot number is much less than one, we may assume no thermal gradient across the diameter. With a diameter of 2 inches we found the Biot number to be 0.3, thus this was the diameter used. In addition the industry standard for tube length was found to be 20 ft12, therefore the MSR pipes are specified as 20 ft in length and 2” inches in diameter. Steel pipes that can withstand a pressure of 7MPa were found in Seider to be Schedule 80 with a nominal pipe size of 2.38-in. O.D. and 1.939-in I.D.

4.1.3.4 Operating Temperature SensitivityThe MSR is very sensitive to its inlet temperature. As seen in Figure 7 if the temperature entering the reactor exceeds 335 K, the reactor temperature will runaway, resulting in catalyst degradation. Thus a large portion of the analysis is focused on determining the optimum inlet temperature. We found this temperature to be 330K.

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200

300

400

500

600

700

800

900

1000

1100

1200

270 290 310 330 350 370 390 410 430 450 470

Precooled Temperature [K]

Rea

cto

r T

emp

erat

ure

[K

]

Figure 7: Excessive MSR inlet temperature causes runaway temperature.

4.1.3.5 CoolantDowtherm Q was selected as an ideal coolant for its excellent heat transfer properties; its overall heat transfer coefficient being nearly five times that of water. As a result, Dowtherm Q prevents a runaway reaction from occurring (due to nearly isothermal operation) in the reactor (Fig. 8). As the reaction proceeds, Dowtherm Q effectively limits the rate of reaction and corresponding temperature increase within acceptable catalytic decomposition limits.

4.1.3.6 Temperature ProfileThe MSR is very sensitive to inlet temperature, coolant temperature, and coolant flow rate. Figure 8 demonstrates optimum temperature profiles through the reactor to achieve maximum conversion.

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Figure 8: Co-current cooling temperature profiles.

4.2 Upstream Processing (H-100, V-200, CMP-200)

Figure 9: STAGE 2 - Upstream processing.

4.2.1 Water RemovalThe OBR effluent is at 885°C and 2 MPa. The methanol reactor temperature cannot exceed 500K. The reactor pressure should be minimized and therefore is specified at the accepted lower range value of 7 MPa. Thus the OBR stream requires compression and cooling to achieve

300

320

340

360

380

400

420

440

460

480

500

0 2 4 6 8 10 12

Distance From Inlet [m]

Reactor Temperature Coolant Temperature

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22

optimum MSR operational conditions. Being that the compressibility of any gas is directly proportional to temperature12, the stream must first be cooled and then compressed in an effort to minimize compressor utility costs. As we implemented this design, it was noted that the cooled OBR exit stream contained condensed water. This natural partitioning of components presents a unique opportunity in separations design. We propose separating all water from the system before the stream enters the methanol reactor. This procedure has the advantage of:

1. Decreasing molar flow rates in our system, thus decreasing capital costs with respect to equipment size.

2. Allows for downstream separations to only be concerned with the separation of methanol from syngas rather than the separation of methanol from water. This is useful given that the separation of methanol from water was found to require a high a capital cost and high operating utilities (Appendix II).

3. The act of cooling the stream is sunk, in that the cost of cooling the stream is a necessity regardless of if we decide to separate the water or not. Thus any action that takes advantage of sunk costs will benefit the profitability of the plant.

4. The absence of water will reduce the competition for sites on the methanol catalyst thus increasing the reaction rate. (This would be a real world effect, as our kinetic model does not take into account water vapor concentration).

4.2.2 CMP-200Given that the feed temperature into the MSR is very sensitive (as described in the R-100 section), an inter-stage compressor (CMP-200) is used. The inter-stage compressor has the advantage of better temperature control, and it will also decrease the energy required to compress the gases. We designed to compressor as a five-stage compressor with a total cooling duty of 11.6 MW.

4.2.3 C-100As with all heat exchanges in this report, the C-100 will be sized according to the procedure outlined in Appendix I.1. In all our heat exchangers the pipes are 16 feet long, have 1 inch triangular spacing, ¾-inch O.D., 0.56-inch I.D., a 1 inch pitch, and are Schedule 80. Furthermore all heat exchangers have a 1 – 2 shell and tube configuration. Using this configuration we found that the C-100 exchanger would require 728 tubes with a 31-inch I.D. shell. The E-100 will require 302 tubes with a 21.75-inch I.D. shell.

4.2.4 V-200As with all the flash units, the V-200 is sized according to the procedure outlined in Appendix I.2. Using this technique we found that this unit will be 14.5 feet tall and have a diameter of 12 feet.

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4.3 Downstream Processing (H-200, V-300, G-100, D-100)

Figure 10: STAGE 4 - downstream processing.

The MSR effluent is 66-wt% methanol. The minimum product purity specification is 99.75-wt% methanol. Downstream separation processing is required to achieve the production quality target.

In reality, when higher alcohols, fusel oils and waxes are present, gases will first be separated from the crude methanol product by distillation in a topping column. Water, fusel oils and methanol will then be separated from methanol in a refining column2.

In our simulation the MSR effluent exits at 374K and 7MPa. It must first be cooled with the goal of causing methanol to liquefy, followed by a flash unit to separate it from the syngas. When this technique was implemented we found that not only does methanol liquefy, but so does carbon dioxide (this was verified upon checking the phase diagram for carbon dioxide). Because the next step of our separations involves the use of a flash unit to remove syngas from methanol, we initially thought this to be an inefficient separation train (as the carbon dioxide syngas was still mixed with the methanol stream).

4.3.1 CO2 RemovalTo rectify this issue we used an expander to drop the pressure and then attempted flashing the stream. Not only did carbon dioxide still appear in the liquid stream, possibly as a dissolved gas, the flash unit caused 25% of our methanol product to exit the vapor stream. We would suggest using a partial condenser in the vapor stream of the flash unit to recover this methanol, but no such Aspen unit exits. Thus counter-intuitively we found it very difficult to remove carbon dioxide from methanol, even in temperature and pressure ranges where carbon dioxide should be vapor, carbon dioxide was the major liquid contaminant in our product stream.

Several temperature and pressure variations were attempted to address the carbon dioxide issue, yet the only Aspen based unit that we were able to get to work was a 6-stage distillation column with a partial condenser. We would like you to note that in reality the need for a distillation column may not be necessary. A flash unit should be able to separate carbon dioxide from methanol at standard temperature and pressure. We believe that this difficulty in separation is a result of Aspen’s thermodynamic package and suggest further investigation into this issue.

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4.3.2 Recycle & ConversionThe next issue is how to handle the vapor stream in the distillation unit. We found that Aspen was very sensitive in the convergence of recycle loops. Flowsheets with recycle loops that had been working for days would suddenly stop working and not converge. Many days were spent trying to get the flowsheet to converge again with no success. The goal then became to design a flowsheet with high conversions thus negating the need for a recycle. As stated in the methanol synthesis reactor section, the only way to obtain acceptable conversions at high temperatures and high pressures is with high hydrogen-to-carbon monoxide ratios. Thus the SMR and OBR were optimized to obtain a hydrogen-to-carbon monoxide ratio of 3.5, which resulted in complete conversion of carbon monoxide within the methanol reactor.

While using this technique consumed most of our carbon monoxide, there was a significant amount of hydrogen in the vapor stream. This stream is then sent to a furnace to recover energy.

4.3.3 Flash Vessel (U-300)The U-300 was sized to be 16.45’ high and have a diameter of 9.5’.

4.3.4 Coolers (C-200, C-300)The C-200 will require 728 tubes with a 31-inch I.D. shell. The C-300 will require 82 tubes with a 12-inch I.D. shell.

4.3.5 Distillation (D-100)We were able to optimize the D-100 with 6 stages, 18-inch tray spacing, a distillate to feed ratio of 0.15, 10.6 MW condenser duty, and 26.9 MW reboiler duty.

4.4 Methanol StorageMethanol storage is needed for constant operation in adjoining facilities in the case of scheduled (or unscheduled) plant downtime. The project managers specified that our storage contingency needs to be 10 days. Based on this specification, we need to store 63,211m3 of methanol. Assuming this volume of methanol can be set in 20 tanks, we can size each tank using:

1

2063,211m3 r2h

h

D 3

(17)

When we solve these equations we find that each of our 20 storage tanks needs to have the following dimensions:

r 7.98m r 26.19 ft

h 47.88m h 157 ft

(18)

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5 Equipment Cost Summary

5.1 Pump CostsA centrifugal pump, of the radial type, was chosen to pump liquid water to the Steam Methane Reformer based upon the volumetric flow rate and head required12. The cost of the pump was obtained using the flow rate and head as the sizing factors for obtaining a base cost.Furthermore, cast iron was assumed to be the appropriate material for the construction of the pump.

5.2 Compressor CostsA centrifugal compressor was chosen based upon the horsepower required to compress the gas to the required pressure12. The cost of the compressor was obtained using the horsepower as the sizing factors for obtaining a base cost. Furthermore, carbon steel was assumed to be the appropriate material for the construction of the pump, and a steam turbine (80% efficiency) was used to take advantage of the utilities present at the plant. Also, the compressors were assumed to be 75% efficient12.

5.3 Furnace CostsThe furnace was assumed to be a fired heater, and its cost estimation is based upon heat duty as the sizing factor. Stainless steel construction is assumed to withstand a pressure of 500 psig.The furnace was assumed to be 75% efficient12.

5.4 Storage Tank CostsOperating specifications require storage of 10 days supply of methanol. Hence, storing 16.7 million gallons of methanol requires 17 tanks with a one million gallon capacity each.

5.5 Reactor CostsThe Steam Methane Reformer was sized as a heater, and a cost estimate was obtained based upon it heating value with a 75% efficiency12. Furthermore, vessel inside this reactor was also considered in the cost analysis. The Oxygen Blown reformer and Methanol Synthesis reactors were sized as pressure vessels12, with pressure being the sizing factor.

5.6 Heat Exchanger CostsAll heat exchangers in the design are shell and tube heat exchangers where the sizing factor is the surface area of heat transfer. The heat transfer area and heat duty were obtained from Aspen for the reactor E-100. From this, the heat transfer coefficient was calculated. Using this calculated heat transfer coefficient, and heat duties obtained from reactors, approximate heat transfer surface areas were found for other three heat exchangers for cost determination.

5.7 Separation Vessel CostsUnits U-200 and U-300 were sized as flash units, and cost was estimated based on the costing method for pressure vessels12. While the cost estimate for distillation column, unit D-100, was obtained using the same method, a slightly different procedure is followed based on Seider’s approach. Carbon dioxide is the main impurity in our methanol product, which can be removed

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by a flash unit. Implementation of this flash unit was difficult in Aspen, hence, a distillation was column was necessary for simulation purposes. The distillation column was sized as other flash units for costing purposes.

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6 Fixed Capital Investment Summary

6.1 Bare Module Costs

A detailed cost analysis for each unit in the process flow diagram was performed based upon methods presented by Seider. Assumptions made in the cost analysis are listed under the section for specific units, while the cost of each unit is presented in Table 3. A detailed cost analysis with specific procedures and correlations are presented in Appendix IV.

Table 3: Descriptions & estimated costs of specific units in the process flow diagram.

Unit Type Description Base Cost / unit No. of Units Total CostPumps $ $P-100 Pump 17,237,109 1 17,237,109

P-spare Spare Pump 17,237,109 1 17,237,109Compressors $ $

CMP-200 Compressor 46,311,255 1 46,311,255Furnaces $ $

F-100 Furnace 4,635,643 1 4,635,643Storage Tanks $ $Floating roof Storage Tanks 430,558 17 7,319,479

Reactors $ $R-100 SMR-furnace 17,237,109 1 17,237,109R-100 SMR-vessel 57,592,250 1 57,592,250R-200 OBR 40,389,195 1 40,389,195R-300 MSR 16,392,569 10 163,925,690HEX $ $E-100 Heat Exchanger 344,960 1 344,960C-100 Heat Exchanger 7,243,525 1 7,243,525C-200 Heat Exchanger 8,948,014 1 8,948,014C-300 Heat Exchanger 182,177 1 182,177

Separators $ $U-200 Flash unit 698,198 1 698,198U-300 Flash unit 276,663 1 276,663D-100 Distillation tower 581,086 1 581,086

Total $355,978,579

6.2 Direct Permanent Investment & Total Capital Investment

The initial estimate of Direct Permanent Investment (DPI) was calculated to be $511.5 million. Adding 30 percent contingency, site and facility preparation, waste removal cost, utility allocation cost, startup costs, land costs, and working capital, the Total Capital investment (TCI) will be $779.5 million. Detailed calculations were performed based on the Guthrie12 method, and can be found in Appendix V.

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7 Other Important Considerations

7.1 Health & SafetyHuman health and safety is an important consideration in the presence of flammable materials such as methane and methanol. Monitoring of fugitive emissions is especially important since methane is colorless and odorless and is therefore not readily identifiable. Although methane and methanol do not pose high hazards to health2, effective measures must be taken to ensure the integrity of plant personnel health and safety.

Methanol and methane are both flammable and present fire and explosive hazards. Methanol should be handled in a confined area, which must be well ventilated. Respirators must be used while working in an area where methanol vapor concentration is high2. Also, gloves and other protective equipment must be used while working in areas of high methanol concentration.

7.2 Process Control & InstrumentationProcess control equipment must be utilized to operate the process equipment within design specifications and to handle possible plant upsets. Process control equipment monitors the temperatures and pressures of reactors, the effectiveness of separation equipment, and all process streams throughout the plant to ensure safe and on-spec operation. Deviations from set-point should provide signals to the controller, and appropriate action(s) must be taken to prevent endangerment of human life, destruction and/or damage of process units, and production of off-spec product.

7.3 Environmental

7.3.1 Chemical ToxicityAs previously stated, humans are most commonly exposed to methanol through skin contact and vapor inhalation. Although carcinogenicity of methanol has not been determined, exposure to methanol has been linked to reproductive defects in rats8. Methanol is known to cause headaches, dizziness, giddiness, insomnia, nausea, gastric disturbances, conjunctivitis, blurred vision, and blindness in humans. High doses of methanol may be fatal. OSHA’s regulatory concentration of methanol for human exposure without adverse effects in an 8 hour day is

260mg

m3 or 198 ppm5.

7.3.2 Potential Safety ProblemsMethanol is readily degraded in the environment by photo oxidation9 and biodegradation processes5. Half-lives of 7 – 18 days have been reported for the atmospheric reaction of methanol with hydroxyl radicals. Methanol is readily degradable under both aerobic and anaerobic conditions in a wide variety of environmental media including fresh and salt water, sediments and soils, ground water, aquifier material and industrial wastewater. Methanol is of low toxicity to aquatic and terrestrial organisms, and effects due to environmental exposure to methanol are unlikely to be observed except in the case of a spill5, 9.

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7.3.3 Required PermitsThe Trinidad and Tobago Environmental Management Agency requires a $500 permit fee and a maximum environmental impact assessment fee of $600,000.

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8 Operating Cost & Economic Analysis

8.1 Cost SheetThe cost sheet was determined by allocating appropriate costs for each category. These categories encompassed utilities, operation overhead, maintenance, labor, property taxes and insurance, depreciation, and general expenses. The total cost of manufacture was determined by adding up all categories of manufacturing cost. The total production cost was determined by adding the total cost of manufacture with general expenses. The sales revenue was determined by knowing the output product flow rate and multiplying it by its selling price; unit conversions were used. The cost sheet is an annual economic analysis sheet.

Table 4: Summary of plant costs and operations.Cost Factor Annual Cost

$2,802,866$364,001

$42,395,868

Feedstocks (raw materials)Natural gas Boiler feed water make-upOxygen

Total $45,562,735

$7,634,955 $7,769,894 $90,009,785 $19,211,641

UtilitiesElectricityCooling water, 90F, 65psig (CW)Chilled cooling water, 60FNatural Gas (fuel), 90F, 75 psig, 1050 BTU/SCF

Total $124,626,275

Operations (labor-related) (O)$524,160 $104,832

$4,504,177 $78,624

Direct wages and benefits (DW&B)Direct salaries and benefitsOperating supplies and servicesControl laboratory

Total $5,211,793

Maintenance (M)$11,260,443 $18,767,405

Wages and benefits (MW&B)Materials and services

Total $30,027,488

Operating Overhead $11,968,059

Property Taxes and Insurance $225,208,854 (entire plant life)

Depreciation (D) $665,018,119 (entire plant life)

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COST OF MANUFACTURE (COM) $183,764,004

TOTAL GENERAL EXPENSES (GE) $10,764,720

TOTAL PRODUCTION COST (C) $194,528,724Sales

Methanol Product $538,236,002Total Sales $538,236,002

8.2 Working CapitalOur design managers provided the working capital equation as:

½(Product Storage)+30 days Mfg. Cost + Spare Parts (reference)

Only one spare part was provided that being a pump and the product storage of methanol was 10 days. The 30 days Mfg. Cost was determined by taking the COM (cost of manufacture) dividing it by 350 days of plant operation and multiplying it by 30 days. Sizing the appropriate storage tank and multiplying it by 10 days determined the product storage cost. The spare pump was sized and then its cost was determined.

8.3 Total Capital InvestmentPresented Total Capital Investment: $777,623,059

This was determined by the working capital and total permanent (fixed) investment.

Table 5: Categorized annual costs.Cost Factor Typical factor in American engineering units

UtilitiesSteam, 300psig $2.40/1000lbsElectricity $0.04/KWhCooling water $0.05/1000galNatural gas $1.50/million BTU HHVWaste treatment $3.00/1000galMSR catalyst price $6.00/lbSMR catalyst price %12.00/lbBoiler feed water make-up $1.50/1000galOxygen price $0.025/lb

Operations (labor-related) (O)Operating labor $10/hrOperator cost/y/(operator/shift) $87,360Operating supplies 0.6% of F.C.ISupervision 20% of operating laborLaboratory 15% of operating labor

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Maintenance (M)Maintenance materials 2.5% of F.C.I

Depreciation (D) 5 yr MARCS scheduleYear 1 20.00%Year 2 32.00%Year 3 19.20%Year 4 11.52%Year 5 11.52%

TaxesEffective US federal tax rate 38%

General Expenses (GE)General Expenses (SARE) 2% of methanol valuePlant overhead 60% of Op. Labor+superv+maintenance+lab

CreditsLight gas by-product credit $1.50/million BTU HHVHigher alcohols by-product

credit $0.12/lbExport steam credit $3.00/1000lbs

Inside battery limits (ISBL)Service facilities & buildings 25% of ISBLWaste treatment capital 6% of ISBLSite development 3% of ISBL

8.4 Profitability Measures

8.4.1 Return on Investment (ROI)

The return on investment calculation is as follows: TCIC

CStROI

))(1(

Where t = U.S. federal tax rate of 38%, S = total sales revenue on an annual basis, C= Cost of production on an annual basis, and CTCI = Total capital investment. All variables are in U.S. dollars. The following calculation was performed,

2740.059$777,623,0

24)$194,528,7-02$538,236,0)(38.01())(1(

TCIC

CStROI

and so the final ROI is roughly 27.4 %.

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Year MACRS fCTDC CWC D CExcl. Dep. S Net Earnings Cash Flow NPV

2009 20.00% ($665,018,119) ($26,926,877) $133,003,624 $183,764,004 $538,236,002 $137,310,392 ($421,630,980) ($421,630,980)2010 32.00% $212,805,798 $183,764,004 $538,236,002 $87,833,044 $300,638,842 $273,308,0382011 19.20% $127,683,479 $183,764,004 $538,236,002 $140,608,882 $268,292,361 $221,729,2242012 11.52% $76,610,087 $183,764,004 $538,236,002 $172,274,385 $248,884,472 $186,990,5882013 11.52% $76,610,087 $183,764,004 $538,236,002 $172,274,385 $248,884,472 $169,991,4432014 5.76% $38,305,044 $183,764,004 $538,236,002 $196,023,512 $234,328,555 $145,499,5972015 $183,764,004 $538,236,002 $219,772,639 $219,772,639 $124,055,9252016 $183,764,004 $538,236,002 $219,772,639 $219,772,639 $112,778,1142017 $183,764,004 $538,236,002 $219,772,639 $219,772,639 $102,525,5582018 $183,764,004 $538,236,002 $219,772,639 $219,772,639 $93,205,0532019 $183,764,004 $538,236,002 $219,772,639 $219,772,639 $84,731,8662020 $183,764,004 $538,236,002 $219,772,639 $219,772,639 $77,028,9692021 $183,764,004 $538,236,002 $219,772,639 $219,772,639 $70,026,3362022 $183,764,004 $538,236,002 $219,772,639 $219,772,639 $63,660,3052023 $183,764,004 $538,236,002 $219,772,639 $219,772,639 $57,873,005

8.4.2 Net Present Value (NPV)To evaluate the net present value of a proposed plant, its cash flows are computed for each year of the projected life of the plant along with construction and startup phases. The sum of all the discounted cash flows equals the net present value. The following table provides the NPV and CF values at 10% interest rate for the life of the plant, which was 15 years.

Table 6: Calculation of Cash Flows and NPV.

The NPV was $1,361,773,040 across the plant life of 15 years.

8.4.3 Cash Flows (CF)During the years of plant construction, the CF for a particular year is as follows:

landWCTDC CCfCCF (ref.)

For the after-tax earnings plus depreciation CF for a particular year the following equation was used:

DCStCF ))(1( (ref.)

The above equation is used for actual years of production not construction. The following table provides the CF for all 15 years of the plant. Notice that during the construction years the CF is negative meaning those were the years of mechanical design and plant construction.

Table 7: Annual cash flows.Year Year of operation Cash Flow

2009 1 ($421,630,980)

2010 2 $300,638,842

2011 3 $268,292,361

2012 4 $248,884,472

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2013 5 $248,884,472

2014 6 $234,328,555

2015 7 $219,772,639

2016 8 $219,772,639

2017 9 $219,772,639

2018 10 $219,772,639

2019 11 $219,772,639

2020 12 $219,772,639

2021 13 $219,772,639

2022 14 $219,772,639

2023 15 $219,772,639

8.4.4 Depreciation Schedule (MACRS)Our design managers provided the schedule of depreciation to us. The following table provides the total amount of depreciation with a class life of 5 years.

Table 8: Depreciation schedule.Year Year of operation MACRS D ($/yr) Taxes Saved ($/yr) 2009 1 20.00% $133,003,624 $50,541,3772010 2 32.00% $212,805,798 $80,866,2032011 3 19.20% $127,683,479 $48,519,7222012 4 11.52% $76,610,087 $29,111,8332013 5 11.52% $76,610,087 $29,111,8332014 6 5.76% $38,305,044 $14,555,9172015 7 - - -2016 8 - - -2017 9 - - -2018 10 - - -2019 11 - - -2020 12 - - -2021 13 - - -2022 14 - - -2023 15 - - -

Total Taxes Saved = $252,706,885Total Depreciation = $665,018,119Present Value of Income Tax Savings (Total) = $195,408,232

8.4.5 Investors Rate of Return (IRR)Using the provided spreadsheet the IRR was roughly 58%. The IRR is also known as the discounted cash flow rate of return (DCFRR). This interest rate or discounted rate gives a net present value of zero and since it is positive this means that building the plant will be profitable.

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The largest IRR is the most desirable, which is the case here. Recall that our NPV value was large and positive.

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9 Conclusions & RecommendationsThe proposed design produces 5,116 MTPD of 99.85 wt% methanol. As designed, the total bare module cost of the plant is $372 million. The inside battery limit and outside battery limit costs are $349 million and $23 million respectively. Total capital investment includes the direct permanent investment of $512 million and is $779 million. The calculated BTROI is 42% with annual net earnings of roughly $203 million per year. The NPV is $1.2 billion in the last year of production and suggests a profitable venture.

Methanol production is a high-risk venture and for such ventures the ROI should ideally be 20–40% in order to justify construction and operation of the plant. The calculated ROI of 26% with annual earnings of roughly $203 million per year suggest a worthwhile investment. The NPV of $1.2 billion in the last year of production also suggests a profitable venture.

The removal of water in upstream processing proved highly beneficial in reducing the total capital and operating costs. It should be noted that the MSR, being highly exothermic, is extremely sensitive to the inlet temperature. Any small perturbation to the inlet temperature could upset the process resulting in a runaway reaction. Therefore a large amount of the operating cost is focused on cooling of the reactor and its inlet stream to prevent emergency upsets.

Another misgiving of the simulation software package arose in the separation of carbon dioxide from methanol. Although intuition dictates that given the two species’ relative volatility, separation should effectively take place in a flash unit, implementation of a simple distillation column was required to effect the desired separation. Removal of the distillation column will also significantly reduce capital and operational costs.

As designed it is worthwhile to pursue investment in this production plant.

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10 AcknowledgementsWe would like to thank the following individuals for their assistance in preparing this report:

Professor Nael El Farra Dr. Jeff Feerer Mr. Richard Anderson

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11 References1. Apanel, G., Davenport, R. E. “Issues Facing Global Methanol Industry.” Chemical Week

Conference. Houston, Texas. October 25, 2004.2. Cheng, Wu-Hsun, Kung, Harold H. Methanol Production and Use. New York: Marcel

Dekker Inc. 1994.3. Davenport, R.E., 2004. Issues Facing Global Methanol Industry. SRI Consulting. Presented

to Chemical Week Conference.4. Dybkjær, I., Bøgild Hansen, J. 1997. Alternative Use of Natural Gas. Elsevier Science.

107:99-116.5. Environmental Protection Agency. “Methanol” Fact Sheet. 08 April, 2006.

<http://www.epa.gov/ttn/atw/hlthef/methanol.html>6. Fogler, H.S., Elements of Chemical Reaction Engineering, 3rd ed. New Jersey: Prentice Hall

PTR, 1999.7. Graaf, G.H., Sijtsema, J.M., Stamhuis, E.J., and Joosten, G.E.H. 1985.

Chemical Equilibria in Methanol Synthesis. Chemical Engineering Science,41:2883-2890.

8. Hu, Y.H. and Ruckenstein, E. 2002. Binary MgO-based Solid Solution Catalysis. Catalysis Reviews. 44:423-453.

9. IPSC INCHEM. “Environmental Health Criteria 196 – Methanol.” http://www.inchem.org/documents/ehc/ehc/ehc196.htm. World Health Organization. Geneva, 1997. 08 April, 2006.

10. McCabe, W. L., and Thiele, E.W. 1925. Graphical design of fractionating columns. Ind. Engr. Chem. 17:605-611.

11. Perry, R. H. Perry’s Chemical Engineers’ Handbook, 7th Ed. McGraw-Hill. 1997.12. Seader, J. D., and Seider, W. D. Product & Process Design Principles. Wiley & Sons, Inc.

2004.13. Stiles, A. B. 1977. Methanol, Past, Present, and Speculation on the Future. AIChE Journal,

23:362-376.14. Tijm, P.J.A., Waller, F.J., and Brown, D.M. 2001. Methanol technology developments for

the new millennium, Applied Catalysis. 221:275-282.15. Villa, P., Forgatti, G., Garone, G. and Pasquon, I. 1985. Syntheis of alcohols from carbon

oxides and hydrogen. 1. Kinetics of low-pressure methanol syntheis. Ind. Eng. Chem. Proc. Des. Dev., 24:2-10.

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12 Appendices

12.1Appendix I: Detailed Equipment Costing

12.1.1 I.1 Heat Exchanger Sizing Technique

This procedure was adapted from Seider:1. Calculate R and S:

hothot

coldcold

Cpm

CpmR

*

*

and

incoldinhot

incoldoutcold

TT

TTS

,,

,,

(#)

2. Calculate F for a 1 shell pass-2 tube pass heat exchanger:

112

112ln*)1(

1

1ln*1

2

2

2

RRS

RRSR

SR

SR

F (#)

3. Compute log mean temperature

incoldouth

outcoldinhot

incoldouthoutcoldinhotlm

TT

TT

TTTTT

,,

,,

,,,,

ln(#)

4. Calculate required Q and UA

outhotinhothothot TTCpmQ ,, and lmTF

QUA

*(#)

5. Now starts the process of iteration. Guess a U (typically 100hftF

BTU

** 2), and then

compute a required area.6. The pipes we chose to use are BWG 14, 16’ in length, with 1” triangular spacing, ¾” OD,

0.560” ID, and 80 schedule. Using these parameters then calculate how many tubes are required.

7. With this number of tubes, go to Table 13.6 in Seider and find the closest number of tubes with a corresponding shell diameter.

8. Then calculated the shell side and tube side heat transfer coefficient from the database of possible heat transfer coefficient correlations specific to your Reynolds number. This database may be found in the mathematica code for this section.

9. Then compute the overall heat transfer coefficient using:

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tubesidemw

iw

o

i

shellside hAk

At

A

A

h

U11

1

(#)

The Mathematica code used for this procedure is:

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12.1.2 I.2 Flash Unit Sizing Procedure

This procedure was adapted from the EnDeCor engineering reference library:1. Max Vapor Velocity

V

VLKU

max (#)

2. Compute Minimum Diameter

Max

VMin U

QD

4

(#)

3. Compute Liquid Height

2

4

D

QtH LL

L (#)

4. Compute Vapor Height4vH (#)

5. Compute overall Length

vL HHL (#)

6. Compute L/D ratio7. If L/D>6 increase D and repeat steps 3-6

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12.1.3 II.3 Maximum Thermodynamically Attainable Conversion

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12.1.4 I.4 General Reactor Sizing Techniques

Case #1: SMR/OBR

Given appropriate kinetic data the SMR/OBR can be sized to achieve certain production levels of synthesis gas (Syngas) from natural gas feedstocks. If the kinetic data is given as the partial pressures of both the components forming the synthesis gas than certain modeling techniques can be used. All the species in the reaction are in the gas phase.

(1) Reaction of Interest:

- Steam Reforming: )(2)()(2)(4 3 gggg HCOOHCH 206298 H kJ/mol

(2) Assumptions:

- Assume a 1:3 ratio of CO: H2 exists at the above conditions.- Endothermic reaction- Multi-tubular packed bed reactor- Catalyst types1: R-67-7H, KATALCO 57-4, and Ni-based- Sufficient heat has been produced from steam- Pure methane gas in the feed- Assume simple power law kinetics - Product species are independent of the reaction rate - No pressure drop in the reactor and no catalyst decay- Steady-state reactor- Assume, for convention, that the reaction rate will be with respect to natural gas.

(3) Operational Guidelines:

- Temperature range (outlet conditions) 3: 1400-1625 F - Pressure range (outlet conditions) 3: 300-450 psig - Inlet temperature3: 1000 F- Molar ratio range3

424

2 CH

OH

(4) SMR Sizing:

1.) Symbolic expression: dDcCbBaA (A=CH4, B=H2O, C=CO, and D=H2) 4

2.) The lowercase letters denote the moles of each corresponding species.3.) Proposed power law expression:

DCBAA CCCkCr 4.) Assume that the product species do not inhibit the forward reaction and that they are

in excess. So the above reaction is an irreversible one. This implies that and = 0.

5.) The new power law expression becomes: BAA CkCr (Fogler, eqn. 3-5)

6.) Now we need to determine Cj = hj(X) or concentration as a function of conversion, X.

REPORT

48

7.) Looking at the symbolic expression yields: Da

dC

a

cB

a

bA and from the above

reaction we know that 1,1,1 cba , and 3d . 8.) The new form of the reaction becomes the following: DCBA 3 . Since we have

a flow system in the gas phase v

FC j

j (Eqn. 3-45) where v

=volumetric flow rate of the inlet stream and j = species of interest and Fj = the molar flow rate of species j. We only

have two species so v

FC CH

CH4

4 and v

FC OH

OH2

2 .

9.) The design equation relating volumetric flow rates to molar flow rates is (Eqn. 3-41) where both temperature dependence and pressure dependence is shown.

10.) Now we can rewrite the above two concentration equations as:

a.

0

0

00

44

T

T

P

P

F

Fv

FC

T

T

CHCH

and

0

0

00

2

2

T

T

P

P

F

Fv

FC

T

T

OHOH

b. If we assume isothermal reactor conditions and negligible pressure drops in the reactor we obtain the following: 0TT and 0PP implying that

T

TCHCH Fv

FFC

0

044 and

T

TOHOH Fv

FFC

0

02

2 .

c. Recall that 0

0

0T

T Cv

F (Eqn. 3-40) and so

02

2 TT

OHOH C

F

FC along with

04

4 TT

CHCH C

F

FC .

d. So in terms of concentration as a function of molar flow rates we have:

i.0

44 T

T

CHCH C

F

FC

ii.0

2

2 TT

OHOH C

F

FC

iii. The above two expression can be plugged into the rate law giving:

1.

00

4 2

TT

OHT

T

CHA C

F

FC

F

Fkr

2. The above expression has reaction rate as a function of concentration of both reactant species in the gas phase.

11.) Next, we can assume that no phase changes occur in the reactor and that no semi-permeable is present. As a result, the design equation relating volumetric flow rate to conversion, neglecting pressure drops along with isothermal operation, is

)1(0 Xvv (Eqn. 3-44).

a. 0Ay (Eqn. 3-36) and =

1

a

b

a

c

a

d (Eqn. 3-23) = 21113

0

0

00 T

T

P

P

F

Fvv

T

T

REPORT

49

b. Also 0Ay =

0

0

T

A

C

C (Example 3-7) which implies that 0,4CHy0

0,4

T

CH

C

C .

c. And so = 2

0

0,4

T

CH

C

C. Since

0

00

T

AA F

Fy (Eqn. 3-39 or 3-40) the volumetric flow

rates cancel and 000 TAA CyC (Example 3-7).

d. Under isothermal and no pressure drop reactor conditions,

X

XvCC jjA

j

1

)(0 (Eqn. 3-46). For Species A we have the following:

X

XvCC AAA

A

1

)(0 where Av = -1 and A = 1.

e. Now that we have concentration as a function of conversion,X

XCC A

A

1

)1(0

Let’s do the same for the second species. Bv = 1a

b and0

0

T

BB F

F (Fogler,

Table 3-3).

Plugging in the following expressions yields: (E3-6.4)

f. We can further simplify this expression to:

g. Now that we have both species in terms of conversion we can plug this into the rate law. This implies that now we have the reaction rate as a function of conversion.

h. We can obtain values for all of the above parameters using ASPEN Tech. Again we are assuming that the reaction rate depends on natural gas and water.

i.

X

XC

X

XCkr BAA

A 11

)1( 00

ii. Given conversion values and knowing the reaction order one can determine thereaction rate as a function of conversion. One must also know the reaction rate constant to completely size the reactor.

X

Xa

bC

CBA

B

1

0

X

XCC BA

B

1

0

REPORT

50

iii. If the assumption of the reaction rate is poor recall that the proposed power law would be:

DCBAA CCCkCr

iv. Plugging in the similar expressions for the products yields:

1.

X

XC

X

XC

X

XC

X

XCkr DACABAA

A 1

3

111

)1( 0000

2. Where 0

0

T

CC F

F (Table 3-3) and

0

0

T

DD F

F (Table 3-3).

v. Knowing reaction rate as a function of conversion one can make a Levenspiel plot to determine the necessary reactor volume for any desired conversion.

(5) PFR case for the SMR:

1.) Assuming that the multi-tubular reactor behaves as a plug flow reactor, the design equation needed tosize the reactor is the following:

a. AA rdV

dXF 0 (Eqn. 2-15)

b. Integrating to find volume yields: X

AA r

dXFV

0

0 (Eqn. 2-16) where the upper

limit is the desired conversion. Numerical methods must be used to solve this integral given reaction rate and conversion data.

c. The final expression with the simplified power law would be as follows:

i.

X

BAA

A

X

XC

X

XCk

dXFV

0 00

0

11

)1(

d. The final expression for the more complicated power law would be as follows:

i.

X

DACABAA

A

X

XC

X

XC

X

XC

X

XCk

dXFV

0 0000

0

1

3

111

)1(

(6) Tube sizing and the number of tubes in the SMR:

REPORT

51

1) If the SMR is non-isothermal and at steady state than heat flow will vary along the length of the reactor. As a result, the heat flux equation must be integrated along the length of the reactor to obtain the total heat added to the reactor.

a. Heat flux equation: )( TTUadV

Qda

(Eqn. 8-44)

b. Now for a tubular reactor (such as the SMR) with heat gain or loss we have the following energy balance equation:

(Eqn. 8-56)

This implies that ),( TXgdV

dT must be coupled with the mole balance,

),(0

TXfF

r

dV

dX

A

A

.

c. Numerical integration of the above two coupled differential equations is required.

i. We can plug in the rate law determined earlier into the above equations and solve for how the temperature changes with volume (i.e., distance) down the reactor.

ii. Using the above equations one can determine the amount of heat added to the SMR by using the heat flux equation.

d. Recall that for a gaseous flow system:

0

0

00 T

TPP

FF

vvT

T (Eqn. 3-41) but this

time we cannot neglect the temperatures terms.

i. The equation becomes

000 T

T

F

Fvv

T

T .

ii. Also recall that

00 )1(

T

TXvv where the pressure drop terms

have been neglected.

iii. Going through the same procedure in step 11.) one obtains the following:

1.

T

T

X

XC

T

T

X

XCkr BAA

A0000

11

)1(

)(

)]()[()(

0 piiA

RxAa

CXCpF

THrTTUa

dV

dT

REPORT

52

2. The same applies for the more complex power law relationship.

e. Knowing Ua and

Q one can determine the number of pipes needed by calculating the area of heat transfer. Looking up tube data one can determine the number of tubes needed for the SMR.

f. Tube wall thickness, Tube outer diameter and inner diameter would have to be calculated or given. Also the tube pitch would be needed in order to determine the number of tubes from a tube sheet layout chart.

REPORT

53

12.2 Appendix II: Upstream Processing

The following analysis will demonstrate why there will be a high capital cost and utility requirement for the separation of water from methanol. Using modified Raoult’s law (Eqn. #),

sati i i iy P x P (#)

we were able to model vapor-liquid equilibrium behavior for a water methanol system. The saturation pressure was approximated by Antoinne’s equation and the activity coefficient by the Van-Laar equation (for excess Gibbs free energy). Note: all parameters for these equations may be found in the following Mathematica code section. Using these equations we were able to generate Figure 11.

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1

Dew Line Bubble Line

Figure 11: x-y diagram for water-methanol system.

We then used the McCabe-Thiele method10 to determine that under typical operating conditions anywhere from 30-60 (also verified in literature11) trays would be necessary to obtain methanol purities of 99.75%. The construction of such a tall tower would require a very large capital investment. Furthermore, this large tower coupled with high molar flow rates will require large heating and cooling duties for the reboiler and condenser respectively1. Thus we sought additional strategies for separations.

REPORT

54

Mathematica Code:

REPORT

55

12.3 Appendix III: Kinetic Models

Once a model was either developed from theory or extracted from literature, nonlinear regression based on a Quasi Newton search method was implemented. To minimize the number of parameters to fit, the regression was done at each temperature, i.e. the model was first fit as a function of concentration holding temperature constant. Then the sum of the squared residuals was computed and reported in Table 2 for each temperature. The model equation that yielded to smallest sum of squared residuals was then expanded, such that each of the temperature dependent terms was assumed to behave via an Arrhenius model. Then the entire data matrix was curve fit both as a function of temperature and concentration and the result may be found in equation #.

Model EquationEquation

Description

Sum of Squared Residuals for

Temp1/Temp 2

3

2 2

0.5 1.3( )CH OHCO H CO H

eq

Pk P P P P

K

Agny & Takoudis31.3036 / 9.00813

3

2

2

2

31 2 3( )

CH OHCO H

eq

CO H

PP P

K

K K P K P

Villa et al. 4.2568 / 1.1053

3 3

2 2

2 2 3 3

2

3

( )

(1 )

CH OH CH OHH H CO CO

eq

H H CO CO CH OH CH OH

K PK P K P

Kk

K P K P K P

Langmuir: Surface Reaction

is RDS and all species adhere associatively

4.42868 / 1.8636

3 3

2 2

2 2 3 3

2

5

( )

(1 )

CH OH CH OHH H CO CO

eq

H H CO CO CH OH CH OH

K PK P K P

Kk

K P K P K P

Langmuir: Surface Reaction

is RDS and all species adhere

disassociatively4.26236 / 1.16466

3

3 2

1

1 2 3(1 )CH OH

CH OH CO H

K Pk

K P K P K P

Langmuir: Methanol

Desorption is RDS

4.46794 / 1.46263

3 21 2 3(1 )CO

CH OH CO H

Pk

K P K P K P

Langmuir: Carbon Monoxide

Associative Adsorption is

RDS8.60895 / 2.69909

3 21 2 3(1 )CO

CH OH CO H

Pk

K P K P K P

Langmuir: Carbon Monoxide

Disassociative Adsorption is

9.33096 / 2.93866

REPORT

56

RDS

3 3

2 2

2 2 3 3

2

3

( )

(1 )

CH OH CH OHH H CO CO

eq

H H CO CO CH OH CH OH

K PK P K P

Kk

K P K P K P

Langmuir: Surface Reaction

is RDS and Carbon Monoxide

adheres disassociatively

53.5992 / 1.09881

2

3 21 2 3(1 )H

CH OH CO H

Pk

K P K P K P

Langmuir: Hydrogen

Associative Adsorption is

RDS5.28849 / 1.56053

2

3 21 2 3(1 )

H

CH OH CO H

Pk

K P K P K P

Langmuir: Hydrogen

Disassociative Adsorption is

RDS543.565 / 133.225

3

2

2 3

2 42 3

51 2 3 4( )

CH OHCO H

eq

CO H CH OH

K PK P K P

Kk

K K P K P K P

Langmuir: Surface Reaction

is RDS and Hydrogen adheres disassociatively 5.3448 / 1.5872

32

3 2

1 20.5

1 0.5 3

cCH OHCO H

cCH OH CO H

PP Pk

P P P c

Leonov et al1410.8735 / 2.76844

3

2

2 3

2

31 2 3 4( )

CH OHCO H

eq

CO H CH OH

PP P

K

K K P K P K P

Natta et al104.25568 / 1.15103

3

2

2 3

2

31 2 3 4( )

CH OH

CO Heq

CO H CH OH

PP P

K

K K P K P K P

Pasquon14 4.2086 / 1.0470

2 3

1 2 3C C CH CO CH OHkP P P

Simplified Power Law 3.818 / 1.0475

3

2

31 2

CCH OHC C

H COeq

Pk P P

K

True Power Law 5.5417 / 1.0530

3

2

2 2 3 3

2

1

CH OHCO H

eq

CO CO H H CH OH CH OH

PP P

Kk

K P K P K P

3

Equation Provided by

Project Managers 0.252 / 0.351

REPORT

57

12.4Appendix IV: Example Detailed Equipment Costing

12.4.1 IV.1 Pumps

Cost of the centrifugal pump is based upon volumetric flowrate and head required12. The base cost of the centrifugal pump and the motor was determined using equations # and # respectively12 where Pc is the power consumption, and S is the sizing factor for the pump.

)(ln(S))0.0519+(ln(S))0.6019-(9.2951 2

e BC Eqn #

)))(ln(P0.0035549-))(ln(P0.028628+))(ln(P0.053255+))(ln(P0.13141+(5.4866B

4c

3c

2cceC Eqn #

The Sizing Factor and the Power Consumption were estimated as follows:

5.0HQS Eqn #

mpc

HQP

33000Eqn #

where

Symbol Name Units

Pc Power consumption Horsepower Q Flowrate Gallons per minuteH Pump head Ft Density Lbs/gallon

2(lnQ)0.01199-(lnQ)0.24051+-0.316 p Fractional efficiency of the pump

Dimensionless

2cc )(lnP0.00182-)(lnP0.0319+0.80 m Fractional efficiency

of the motorDimensionless

Now the purchasing cost, Cp, can be determined by the following equations:

Pump BMTp CFFC Eqn #

Motor BTp CFC Eqn #

Where

Symbol NameSeider Value

REPORT

58

FT Type Factor 2.7FM Material Factor 1.0

REPORT

59

12.4.2 IV.2 Storage Tanks

Unit Type Size factor, S Range of S Cost equationSeider

Floating roof Volume, gal 10,000 - 1,000,000 gal Cp = 375*V^0.51

No. of gallons needed to be stored: 16684322 ten days' production

No. of tanks required: 17 at 1,000,000 gallons/tank

Unit Type Cost / Unit No. of Units Total Cost$ $

Floating roof 430,558 17 7,319,479

REPORT

60

12.4.3 IV.3 Compressors

Purchase Cost, Cp

Cp = Fd * Fm * Cb

I. For electric motor drive, cast iron or carbon-steel constructionpurchase cost is obtained directly from Garrett and Walas (1988) [Table 16.19]12:

Cp = Cb

II. For other drives and materials of construction:

Cp = Fd * Fm * Cb

Table 1: Factors for material construction and motor drive for Compressors12.Drive Fd Material Fm

Steam turbine 1.15 Stainless steel 2.5Gas turbine 1.25 Nickel Alloy 5

III. Cost of different Compressors:

Table 7: Base and Purchasing cost for various types of compressors using Fd =1.15 and Fm = 1 for CMP-200 unit.

Compressor Horsepower Cb ($) Cp ($)

Centrifugal 29,035 5,092,634 5,856,530 Cb = EXP(7.2223+0.8*LN(Hp))

Reciprocating 29,035 7,492,446 8,616,312 Cb = EXP(7.6084+0.8*LN(Hp))

Screw 29,035 4,029,662 4,634,112 Cb = EXP(7.7661+0.7243*LN(Hp))

REPORT

61

12.4.4 IV.4 Reactors

Steam Methane Reformer

Type of Heater Size factor, Q Range of Q Q value CpBtu/hr Btu/hr Btu/hr $

Reformer Heat absorbed 10 - 500 Million 1,179,779,845 15,092,023 Cp = 0.677*(Q)^0.81

Oxygen Blown Reformer

Oxygen Blown reformer is treated as a Pressurized vessel for cost purposes.

Mulet, Corripio and Evans method (Seider and Seader, Page 527)

carbon steel construction and includes an allowance for platforms, ladders, and a nominal number of nozzles and manholes

Cp = Fm* Cv + CplCv = f.o.b cost of the empty vessel including nozzles, manholes and supports based on weight in pounds, W, of shell and two elliptical regions

Horrizontal vessels for 1,000 < W < 920,000 pounds

Cv = exp(8.717-0.2330(ln(W))+0.4333(ln(W))^2)

W (lbs) Cv ($)

8831 26294

Cpl = the added cost for platforms and ladders depends on vessel diameter in feet and for vertical vessel, the length L.

Horrizontal vessels for 3 < Di < 12 ft

Cpl = 1580*(Di)^0.20294

Di (ft) L (ft) Cpl ($)

3.28 N/A 2011

Cp = Fm* Cv + Cpl

Vessel Fm Cv ($) Cpl ($) Cp ($)

Horizontal 1 26294 2011 28304

REPORT

62

Weight Calculation: W = pi()*(Di+ts)*(L+0.8Di)*ts*rho

Pd = exp(6.60608+0.91615(ln(Po))+0.0015655(ln(Po))^2)Po minimum = 10 psig. For pressures greater than 1000 psig, use Po = 1.1*Po and not the above equation.

Tp = (Pd*Di)/(2S*E-1.2*Pd)tp must be greater than a minimum value for rigidity based on the diameter (Table on Page 530)

S = maximum allowable stress of the shell material at the design temperature in pounds per square inchE = fractional weld efficiency

find minimum wall thickness, tmin, and ts must be greater than tmin

OR

FOR Vertical vessels take into account effects of wind and earth quake

Tv = tp(0.75+0.22*E*(((L/Di)^2)/Pd))

Then use either tp or tv:

Add corrosion allowance of 1/8 inch

Then find ts = tp or tv + corrosion allowance

Find Weight, W (lbs)

Unit type L Di Op. Po Pd S E Tp tv ts W(ft) (ft) (psig) (psig) (psi) Ft ft ft lbs

Drum h. 9.84 3.28 500 578 13100 1 0.07 0.06 0.13 8831

Methanol Synthesis Reactor (MSR)

I. Shell and Tube Heat Exchanger Approach

Fixed Head

Cb = exp(11.0545-0.9228(ln(A))+0.09861(ln(A))^2)

A is the area for heat exchange in sq ft.

Calculating Cp

REPORT

63

Fm = a+((A/100)^b)

Fl: can be looked up in Seider and Seader,page 523, table at the bottom

Fp = 0.9803+0.018*(P/100)+0.0017*(P/100)^2

Cp = Fp*Fm*Fl*Cb

Table 8: Base and Purchasing cost of the Methanol Synthesis Reactor.HEX type P a b A Fm Fl Fp Cb Cp

Psig sqft $ $Fixed 1000 0 0 418,879 1 1 1.33 5,302,257 7,053,593

REPORT

64

12.4.5 IV.5 Furnaces

Furnace was assumed to be a fired heater for cost determination

Cb = exp(0.08505+0.766*(ln(Q)))

Fp = 0.986-0.0035*(P/500)+0.0175((P/500)^2)

Fm = 1.7 for stainless steel

Cp =Fp*Fm*Cb

Unit P (psig)Q value (Btu/hr) Fp Fm Cb ($) Cp ($)

F-100 500 189,703,402* 1 1.7 2,387,494 4,058,740*Heating value assumes 85% efficiency of the furnace.

REPORT

65

12.4.6 IV.6 Heat Exchangers

I. Shell and Tube

Floating Head

Cb = exp(11.667-0.8709(ln(A))+0.09005(ln(A))^2)

Fixed Head

Cb = exp(11.0545-0.9228(ln(A))+0.09861(ln(A))^2)

Calculating Cp

Fm = a+((A/100)^b)

Fl: can be looked up in Seider and Seader, page 523

Fp = 0.9803+0.018*(P/100)+0.0017*(P/100)^2

Cp = Fp*Fm*Fl*Cb

Heat duty and heat transfer surface area for unit E-100 were obtained from Aspen. This was used to calculate U, the heat transfer coefficient. This coefficient was used to approximate the heat transfer surface area for other heat exchangers in this network. Results of these calculations are shown below.

Unit Heat duty Cp water Tout Tin mflow rate Area req.Btu/hr Btu/lbs/ºF ºF ºF lbs/hr sqft

E-100 611311704 0.9990924 110 90 30593352 3165C-100 931173410 0.9990924 110 90 46600966 4821C-200 288527271 0.9990924 110 90 14439469 1494C-300 25764398 0.9990924 110 90 1289390.2 133

A = Q/ (U*(∆Tlm))

Ucalc = 9657.4

Unit HEX type P a* b* A Fm Fl Fp Cb Cppsig sq ft. $ $

E-100 Floating 500 2.7 0.07 3,176 3.97 1 1.11 36,297 160,509C-100 Floating 500 2.7 0.07 4,822 4.01 1 1.1128 47,001 209,820C-200 Floating 1000 2.7 0.07 1,494 3.91 1 1.3303 24,644 128,132C-300 Floating 5 2.7 0.07 134 3.72 1 0.981204 14,209 51,876

REPORT

66

12.4.7 IV.7 Flash Vessels

Separators modeled as pressure vessels

Mulet, Corripio and Evans method (Seider and Seader, Page 527)

carbon steel construction and includes an allowance for platforms, ladders, and a nominal number of nozzles and manholes

Cp = Fm* Cv + Cpl

Cv = f.o.b cost of the empty vessel including nozzles, manholes and supports based on weight in pounds, W, of shell and two elliptical regions

Vertical vessels for 4,200 < W < 1,000,000 pounds

Cv = exp(6.775+0.18255(ln(W))+0.02297(ln(W))^2)

Unit W (lbs) Cv ($)

U-200 150577 202066U-300 163161 214295

Cpl = the added cost for platforms and ladders depends on vessel diameter in feet and for vertical vessel, the length L.

Vertical vessels for 3 < Di < 12 ft and 12 < L < 40 ft

Cpl = 285.1* (Di)^0.73960 * (L)^0.7684

Unit Di (ft) L (ft) Cpl ($)

U-200 12 14.5 13982U-300 9.5 16.5 12991

Weight Calculation: W = pi()*(Di+ts)*(L+0.8Di)*ts*rho

Pd = exp(6.60608+0.91615(ln(Po))+0.0015655(ln(Po))^2)Po minimum = 10 psig. For pressures greater than 1000 psig, use Po = 1.1*Po and not the above equation.

tp = (Pd*Di)/(2S*E-1.2*Pd)tp must be greater than a minimum value for rigidity based on the diameter (Table on Page 530)

S = maximum allowable stress of the shell material at the design temperature in pounds per

REPORT

67

square inch

E = fractional weld efficiency

find minimum wall thickness, tmin, and ts must be greater than tmin

OR

FOR Vertical vessels take into account effects of wind and earth quake

tv = tp(0.75+0.22*E*(((L/Di)^2)/Pd))

Then use either tp or tv:

Add corrosion allowance of 1/8 inch

Then find ts = tp or tv + corrosion allowance

Find Weight, W (lbs)

Unit Unit type L Di Op. Po Pd S E tp tv ts Wft ft (psig) (psig) (psi) ft ft ft lbs

U-200 Flash v. 14.5 12 500 578 13100 1 0.27 0.20 0.33 150577U-300 Flash v. 16.5 9.5 1000 1107 13100 1 0.42 0.32 0.44 163161

Cp = Fm* Cv + Cpl

Vessel Fm Cv ($) Cpl ($) Cp ($)

U-200 1 202,066 13982 216,047U-300 1 214,295 12991 227,286

REPORT

68

12.4.8 IV.8 Distillation Columns

Towers for 9,000 < W < 2,500,000 lb

Cv = exp(7.0374+0.18255(ln(W))+0.02297(ln(W))^2)

W (lbs) Cv ($)

6901 34399

Towers for 3 < Di < 24 ft and 27 < L < 170 ft

Cpl = 237.1* (Di)^0.63316 * (L)^0.80161

Weight Calculation: W = pi()*(Di+ts)*(L+0.8Di)*ts*rho

Pd = exp(6.60608+0.91615(ln(Po))+0.0015655(ln(Po))^2)Po minimum = 10 psig. For pressures greater than 1000 psig, use Po = 1.1*Po and not the above equation.

tp = (Pd*Di)/(2S*E-1.2*Pd)tp must be greater than a minimum value for rigidity based on the diameter (Table on Page 530)

S = maximum allowable stress of the shell material at the design temperature in pounds per square inchE = fractional weld efficiency

find minimum wall thickness, tmin, and ts must be greater than tmin

OR

FOR Vertical vessels take into account effects of wind and earth quake

Tv = tp(0.75+0.22*E*(((L/Di)^2)/Pd))

Then use either tp or tv:

Add corrosion allowance of 1/8 inch

Then find ts = tp or tv + corrosion allowance

Find W

Unit type L Di Op. Po Pd S E Tp tv ts Wft ft (psig) (psig) (psi) Ft ft ft Lbs

distill v. 9 3 5 8 15000 1 0.00 0.00 0.13 6901

Di (ft) L (ft) Cpl ($)

3 9 2767

REPORT

69

Cpl = 237.1* (Di)^0.63316 * (L)^0.80161

Vessel Fm Cv ($) Cpl ($) Cp ($)

Tower 2.1 34399 2767 75005

Unit Type Base Cost / unit No. of Units Total CostSeparators $ $

U-200 216,047 1 216,047U-300 227,286 1 227,286D-100 75,005 1 75,005

Total $518,338

REPORT

70

12.5Appendix V: Direct Permanent Investment & Total Capital Investment

WC $ TPI $ Total $TCI = SUM 28,829,285 750,696,181 779,525,466

½ Prod. Stor. 30 day Mft. $ Spare $ Total $WC = SUM 7,319,479 18,023,215 3,486,591 28,829,285

TPI = SUM TDC Startup Total $665,018,119 17,432,955 682,451,074

Startup = 5% TDC 17,432,955

Contingency DPI Total $TDC = SUM 153,465,719 511,552,399 665,018,119

Contingency = 30% DPI, $ 153,465,719

Site prep Srvc. Fac. TBM Total $30% TBM 25% TBM

DPI 52,298,865 69,731,820 372,088,759 511,552,399

Fab. Equip.Proc.

Machin. Storage Spare Catalyst Total $TBM 40,000,000 300,659,100 7,319,479 3,486,591 12,623,589 372,088,759

Cost Density Volume Total$/lb lb/ft^3 ft^3 $

MSR Catalyst 6 90 17500 9,450,000SMR Catalyst 12 54 1000 3,173,589

Total All Catalyst $ 12,623,589