METHANOL PRODUCTION IN TRINIDAD & TOBAGO Final Report: Phase II
University of California, Davis
Date of Report: June 07, 2006
Design Group One Elton Amirkhas Raj Bedi Steve Harley Trevor Lango
REPORT
Executive Summary This 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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REPORT Table of Contents Executive Summary ........................................................................................................................ 2 1 Introduction............................................................................................................................. 6 1.1 Overview......................................................................................................................... 6 1.2 Current Manufacturing Methods..................................................................................... 6 1.2.1 Syngas Production .................................................................................................. 6 1.2.2 Methanol Production............................................................................................... 7 1.3 Selected Production Method ........................................................................................... 7 1.4 Production Level & Plant Location ................................................................................ 7 1.4.1 Production Level..................................................................................................... 7 1.4.2 Plant Location ......................................................................................................... 7 1.5 Market Considerations .................................................................................................... 7 1.6 Environmental Issues ...................................................................................................... 7 1.6.1 Chemical Toxicity................................................................................................... 7 1.6.2 Potential Safety Problems ....................................................................................... 8 2 Process Description................................................................................................................. 9 2.1 Block Flow Diagram....................................................................................................... 9 2.1.1 Chemical Reactions ................................................................................................ 9 2.1.2 Separations.............................................................................................................. 9 2.2 Detailed Flow Diagram................................................................................................... 9 2.2.1 STAGE 1: Syngas Production ................................................................................ 9 2.2.1.1 Natural Gas Furnace (F-100) .............................................................................. 9 2.2.1.2 Steam Methane Reformer (R-100)...................................................................... 9 2.2.1.3 Oxygen Blown Reformer (R-200) ...................................................................... 9 2.2.2 STAGE 2: Upstream Processing........................................................................... 10 2.2.2.1 Steam Generator (E-100) .................................................................................. 10 2.2.2.2 Syngas Cooler (C-100) ..................................................................................... 10 2.2.2.3 Flash Unit (U-200)............................................................................................ 10 2.2.2.4 Water Mixer (M-200) ....................................................................................... 10 2.2.2.5 Water Make-up Pump (P-100).......................................................................... 10 2.2.2.6 Syngas Compressor (CMP-200) ....................................................................... 10 2.2.3 STAGE 3: Methanol Production........................................................................... 10 2.2.3.1 Feed Splitter (S-100)......................................................................................... 10 2.2.3.2 Methanol Synthesis Reactor (R-300)................................................................ 10 2.2.3.3 Product Mixer (M-300)..................................................................................... 10 2.2.4 STAGE 4: Downstream Processing...................................................................... 11 2.2.4.1 Product Cooler (C-200)..................................................................................... 11 2.2.4.2 Syngas Separator (U-200)................................................................................. 11 2.2.4.3 Depressurizer (V-100) ...................................................................................... 11 2.2.4.4 Distillation (D-100)........................................................................................... 11 2.2.4.5 Final Product Mixer (M-100)............................................................................ 11 2.2.4.6 Final Product Cooler (C-300) ........................................................................... 11 3 Energy Balance & Utility Requirements .............................................................................. 12 3.1 Energy Requirements.................................................................................................... 12 3.2 Process Integration........................................................................................................ 12
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Equipment List & Unit Descriptions .................................................................................... 13 4.1 Reactors......................................................................................................................... 13 4.1.1 Steam Methane Reformer (R-100)........................................................................ 13 4.1.2 Oxygen Blown Reformer (R-200) ........................................................................ 15 4.1.3 Methanol Synthesis Reactor (R-300).................................................................... 16 4.1.3.1 Kinetics ............................................................................................................. 16 4.1.3.2 Maximum Conversion ...................................................................................... 18 4.1.3.3 Catalyst ............................................................................................................. 19 4.1.3.4 Operating Temperature Sensitivity ................................................................... 19 4.1.3.5 Coolant.............................................................................................................. 20 4.1.3.6 Temperature Profile .......................................................................................... 20 4.2 Upstream Processing (H-100, V-200, CMP-200)......................................................... 21 4.2.1 Water Removal ..................................................................................................... 21 4.2.2 CMP-200............................................................................................................... 22 4.2.3 C-100..................................................................................................................... 22 4.2.4 V-200 .................................................................................................................... 22 4.3 Downstream Processing (H-200, V-300, G-100, D-100) ............................................. 23 4.3.1 CO2 Removal ........................................................................................................ 23 4.3.2 Recycle & Conversion .......................................................................................... 24 4.3.3 Flash Vessel (U-300) ............................................................................................ 24 4.3.4 Coolers (C-200, C-300) ........................................................................................ 24 4.3.5 Distillation (D-100)............................................................................................... 24 4.4 Methanol Storage .......................................................................................................... 24 Equipment Cost Summary .................................................................................................... 25 5.1 Pump Costs ................................................................................................................... 25 5.2 Compressor Costs ......................................................................................................... 25 5.3 Furnace Costs................................................................................................................ 25 5.4 Storage Tank Costs ....................................................................................................... 25 5.5 Reactor Costs ................................................................................................................ 25 5.6 Heat Exchanger Costs ................................................................................................... 25 5.7 Separation Vessel Costs................................................................................................ 25 Fixed Capital Investment Summary...................................................................................... 27 6.1 Bare Module Costs........................................................................................................ 27 6.2 Direct Permanent Investment & Total Capital Investment........................................... 27 Other Important Considerations............................................................................................ 28 7.1 Health & Safety............................................................................................................. 28 7.2 Process Control & Instrumentation............................................................................... 28 7.3 Environmental............................................................................................................... 28 7.3.1 Chemical Toxicity................................................................................................. 28 7.3.2 Potential Safety Problems ..................................................................................... 28 7.3.3 Required Permits................................................................................................... 29 Operating Cost & Economic Analysis.................................................................................. 30 8.1 Cost Sheet ..................................................................................................................... 30 8.2 Working Capital............................................................................................................ 31 8.3 Total Capital Investment............................................................................................... 31 8.4 Profitability Measures................................................................................................... 32
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REPORT 8.4.1 Return on Investment (ROI) ................................................................................. 32 8.4.2 Net Present Value (NPV)...................................................................................... 33 8.4.3 Cash Flows (CF) ................................................................................................... 33 8.4.4 Depreciation Schedule (MACRS)......................................................................... 34 8.4.5 Investors Rate of Return (IRR) ............................................................................. 34 9 Conclusions & Recommendations........................................................................................ 36 10 Acknowledgements........................................................................................................... 37 11 References......................................................................................................................... 38 12 Appendices........................................................................................................................ 39 12.1 Appendix I: Detailed Equipment Costing..................................................................... 39 12.1.1 I.1 Heat Exchanger Sizing Technique .................................................................. 39 12.1.2 I.2 Flash Unit Sizing Procedure ............................................................................ 45 12.1.3 II.3 Maximum Thermodynamically Attainable Conversion................................. 46 12.1.4 I.4 General Reactor Sizing Techniques ................................................................ 47 12.2 Appendix II: Upstream Processing ............................................................................... 53 12.3 Appendix III: Kinetic Models....................................................................................... 55 12.4 Appendix IV: Example Detailed Equipment Costing................................................... 57 12.4.1 IV.1 Pumps ........................................................................................................... 57 12.4.2 IV.2 Storage Tanks ............................................................................................... 59 12.4.3 IV.3 Compressors ................................................................................................. 60 12.4.4 IV.4 Reactors ........................................................................................................ 61 12.4.5 IV.5 Furnaces........................................................................................................ 64 12.4.6 IV.6 Heat Exchangers ........................................................................................... 65 12.4.7 IV.7 Flash Vessels ................................................................................................ 66 12.4.8 IV.8 Distillation Columns..................................................................................... 68 12.5 Appendix V: Direct Permanent Investment & Total Capital Investment ..................... 70
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REPORT
1 Introduction 1.1 Overview Methanol 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 Methods Contemporary 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 Production Current 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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CH 4 H 2O CO 3H 2 1 CH 4 O2 CO 2H 2 2 1 O2 H 2 H 2O 2
(1) (2) (3)
1.2.2 Methanol Production Methanol is produced from syngas via the Fischer-Tropsch Process, given by Eqn. (4). There are no significant byproducts or intermediates. CO 2H 2 CH 3OH
(4)
1.3 Selected Production Method The 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 Level A production level of 5,000 MTPD was selected to reflect current anticipated market demand.
1.4.2 Plant Location Trinidad and Tobago was selected as an ideal location for methanol production due to its large natural gas reserves.
1.5 Market Considerations Because 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 Toxicity Most 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 mg for human exposure without adverse effects in an 8 hour day is 260 3 or 198 ppm5. m
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1.6.2 Potential Safety Problems Methanol 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 Reactions Methane, 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 Separations Upstream 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 Requirements The 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
Reactors R-100 R-300
0.34576E+09 -0.17666E+09
Combustion of natural gas Dowtherm Q
Furnaces F-100
0.55597E+08
Combustion of natural gas
Heat Exchangers E-100
1.791577830176E+08
Exchange streams 6 & 31 enthalpy
Coolers C-100 C-200 C-300
-0.27290E+09 -0.84559E+08 -0.75508E+07
Cooling water Cooling water Cooling water
Pumps & Compressors P-100 CMP-200 (cooling/electrical)
123,705 -0.116222+08/0.216511+08
Electricity Cooling water/Electricity
Vessels D-100 (reboiler/condenser)
0.268502+08/-0.105732+08
Combustion/Cooling water
3.2 Process Integration The measures adopted to improve the plant economics by energy and mass conservation are: 1. Steam generation using OBR effluent energy 2. 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). CH 4 H 2 0 C 0 3H 2
H 298 206kJ / 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 R100, 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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0.7
7
0.6
6
0.5
5
0.4
4
0.3
3
0.2
2
0.1
1
0 1400
1450
1500
1550
1600
1650
1700
1750
H2-to-CO Ratio (Syngas)
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0 1800
SMR Temp [F] SYNCH4
SYNH2
SYNCO
SYNCO2
H2CO
Figure 4: Effect of temperature on R-100. 0.7
4.4
4.3
0.6
4.2
4.1 0.4 4 0.3 3.9
H2-to-CO Ratio (Syngas)
0.5
0.2 3.8
0.1
0 200
3.7
250
300
350
400
450
500
550
3.6 600
SMR Pres [psig] SYNCH4
SYNH2
SYNCO
SYNCO2
H2CO
Figure 5: Effect of pressure on R-100.
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1
0.9 5 0.8
4 0.6
3
0.5
0.4 2
H2-to-CO Ratio (Syngas)
0.7
0.3
0.2 1 0.1
0
0 0
0.5
1
1.5
2
2.5
3
SYNCO2
H2CO
3.5
4
H20-to-CH4 Ratio (Feed) SYNCH4
SYNH2
SYNCO
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): H 2 O2 H 2 O H 298 242kJ / mole
(6)
Partially oxidize methane via Eqn. (7): 1 CH 4 O2 C 0 2 H 2 2
H 298 36kJ / mole
(7)
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REPORT 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 CO 2 H 2 CH 3 OH H 298 90.5kJ / mole
(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. PCO PH 2 rCH 3OH k (T ) * T n
PCH 3OH K eq (T , P)
(1 K CO (T ) PCO K H 2 (T ) PH 2 K CH 3OH (T ) yPCH 3OH ) ln k (T ) 0.191 -
678 RT
T n T 0.392 lnK CO (T ) 103
ln K H 2 (T ) 155
(9) 3
(10) (11)
5705 ln 12.4 T
(12)
9496 ln19.4 T
(13)
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ln K CH 3OH (T ) 46.7
1692 ln 4.76 T
(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: CO2 3H 2 CH 3OH H 2O 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: CO2 CH 4 2CO 2 H 2
(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 more temperature 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.
X 1 2 X ratio X 1 2 X ratio X K eq (T , P ) 1 2 X ratio 1 2 X ratio V FAO k (T ) * T n dX 0 1 X X 3 ratio 2 X (1 K CO (T ) ) K CH 3OH (T ) K H 2 (T ) 1 2 X ratio 1 2 X ratio 1 2 X ratio
(17)
Note: “ratio” denotes the molar H2-to-CO ratio.
4.1.3.2 Maximum Conversion Ever 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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REPORT Table 2: Maximum conversion as a function of H2-to-CO ratio and temperature. Temperature Pressure (K) (MPa) 400 7 400 7 400 4 450 7 450 7 450 4 500 7 500 7 500 4
H2:CO ratio 2 5 2 2 5 2 2 5 2
Maximum Conversion 0.95 0.99 0.92 0.83 0.99 0.75 0.61 0.87 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 Catalyst Thermal 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 Sensitivity The 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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1200
1100
1000
Reactor Temperature [K]
900
800
700
600
500
400
300
200 270
290
310
330
350
370
390
410
430
450
470
Precooled Temperature [K]
Figure 7: Excessive MSR inlet temperature causes runaway temperature.
4.1.3.5 Coolant Dowtherm 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 Profile The 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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500
480 460
440
420 400
380
360
340 320
300 0
2
4
6
8
10
12
Distance From Inlet [m] Reactor Temperature
Coolant Temperature
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 Removal The 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 21
REPORT 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-200 Given 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-100 As 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-200 As 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 Removal To 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 & Conversion The 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 Storage Methanol 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 63,211m 3 r 2 h 20 h 3 D
(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
(18)
h 47.88m h 157 ft
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5 Equipment Cost Summary 5.1 Pump Costs A 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 Costs A 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 Costs The 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 Costs Operating 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 Costs The 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 Costs All 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 Costs Units 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 25
REPORT 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 Pumps P-100 P-spare Compressors CMP-200 Furnaces F-100 Storage Tanks Floating roof Reactors R-100 R-100 R-200 R-300 HEX E-100 C-100 C-200 C-300 Separators U-200 U-300 D-100
Description Pump Spare Pump Compressor Furnace Storage Tanks SMR-furnace SMR-vessel OBR MSR Heat Exchanger Heat Exchanger Heat Exchanger Heat Exchanger Flash unit Flash unit Distillation tower
Base Cost / unit $ 17,237,109 17,237,109 $ 46,311,255 $ 4,635,643 $ 430,558 $ 17,237,109 57,592,250 40,389,195 16,392,569 $ 344,960 7,243,525 8,948,014 182,177 $ 698,198 276,663 581,086
No. of Units 1 1 1 1 17 1 1 1 10 1 1 1 1 1 1 1 Total
Total Cost $ 17,237,109 17,237,109 $ 46,311,255 $ 4,635,643 $ 7,319,479 $ 17,237,109 57,592,250 40,389,195 163,925,690 $ 344,960 7,243,525 8,948,014 182,177 $ 698,198 276,663 581,086 $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. 27
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7 Other Important Considerations 7.1 Health & Safety Human 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 & Instrumentation Process 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 offspec product.
7.3 Environmental 7.3.1 Chemical Toxicity As 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 mg 260 3 or 198 ppm5. m
7.3.2 Potential Safety Problems Methanol 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 Permits The 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 Sheet The 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 Feedstocks (raw materials) Natural gas Boiler feed water make-up Oxygen Total Utilities Electricity Cooling water, 90F, 65psig (CW) Chilled cooling water, 60F Natural Gas (fuel), 90F, 75 psig, 1050 BTU/SCF Total
Annual Cost $2,802,866 $364,001 $42,395,868 $45,562,735 $7,634,955 $7,769,894 $90,009,785 $19,211,641 $124,626,275
Operations (labor-related) (O) Direct wages and benefits (DW&B) Direct salaries and benefits Operating supplies and services Control laboratory Total
$524,160 $104,832 $4,504,177 $78,624 $5,211,793
Maintenance (M) Wages and benefits (MW&B) Materials and services Total
$11,260,443 $18,767,405 $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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REPORT COST OF MANUFACTURE (COM)
$183,764,004
TOTAL GENERAL EXPENSES (GE)
$10,764,720
TOTAL PRODUCTION COST (C) Sales Methanol Product Total Sales
$194,528,724 $538,236,002 $538,236,002
8.2 Working Capital Our 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 Investment Presented 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 Utilities Steam, 300psig Electricity Cooling water Natural gas Waste treatment MSR catalyst price SMR catalyst price Boiler feed water make-up Oxygen price Operations (labor-related) (O) Operating labor Operator cost/y/(operator/shift) Operating supplies Supervision Laboratory
Typical factor in American engineering units $2.40/1000lbs $0.04/KWh $0.05/1000gal $1.50/million BTU HHV $3.00/1000gal $6.00/lb %12.00/lb $1.50/1000gal $0.025/lb
$10/hr $87,360 0.6% of F.C.I 20% of operating labor 15% of operating labor
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Maintenance (M) Maintenance materials
2.5% of F.C.I
Depreciation (D)
5 yr MARCS schedule Year 1 20.00% Year 2 32.00% Year 3 19.20% Year 4 11.52% Year 5 11.52%
Taxes Effective US federal tax rate General Expenses (GE) General Expenses (SARE) Plant overhead
38%
2% of methanol value 60% of Op. Labor+superv+maintenance+lab
Credits Light gas by-product credit Higher alcohols by-product credit Export steam credit Inside battery limits (ISBL) Service facilities & buildings Waste treatment capital Site development
$1.50/million BTU HHV $0.12/lb $3.00/1000lbs
25% of ISBL 6% of ISBL 3% of ISBL
8.4 Profitability Measures 8.4.1 Return on Investment (ROI) (1 t )( S C ) CTCI 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,
The return on investment calculation is as follows: ROI
ROI
(1 t )( S C ) (1 0.38)($538,236,002 - $194,528,724) 0.2740 $777,623,059 CTCI
and so the final ROI is roughly 27.4 %.
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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. Year 2009 2010 2011 2012 2013 2014 2015 2016 2017 2018 2019 2020 2021 2022 2023
MACRS 20.00% 32.00% 19.20% 11.52% 11.52% 5.76%
fCTDC ($665,018,119)
CWC ($26,926,877)
D $133,003,624 $212,805,798 $127,683,479 $76,610,087 $76,610,087 $38,305,044
CExcl. Dep. $183,764,004 $183,764,004 $183,764,004 $183,764,004 $183,764,004 $183,764,004 $183,764,004 $183,764,004 $183,764,004 $183,764,004 $183,764,004 $183,764,004 $183,764,004 $183,764,004 $183,764,004
S $538,236,002 $538,236,002 $538,236,002 $538,236,002 $538,236,002 $538,236,002 $538,236,002 $538,236,002 $538,236,002 $538,236,002 $538,236,002 $538,236,002 $538,236,002 $538,236,002 $538,236,002
Net Earnings $137,310,392 $87,833,044 $140,608,882 $172,274,385 $172,274,385 $196,023,512 $219,772,639 $219,772,639 $219,772,639 $219,772,639 $219,772,639 $219,772,639 $219,772,639 $219,772,639 $219,772,639
Cash Flow ($421,630,980) $300,638,842 $268,292,361 $248,884,472 $248,884,472 $234,328,555 $219,772,639 $219,772,639 $219,772,639 $219,772,639 $219,772,639 $219,772,639 $219,772,639 $219,772,639 $219,772,639
NPV ($421,630,980) $273,308,038 $221,729,224 $186,990,588 $169,991,443 $145,499,597 $124,055,925 $112,778,114 $102,525,558 $93,205,053 $84,731,866 $77,028,969 $70,026,336 $63,660,305 $57,873,005
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: CF fCTDC CWC Cland (ref.) For the after-tax earnings plus depreciation CF for a particular year the following equation was used: CF (1 t )( S C ) D (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 2009 1 2010 2 2011 3 2012 4
Cash Flow ($421,630,980) $300,638,842 $268,292,361 $248,884,472
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REPORT 2013 2014 2015 2016 2017 2018 2019 2020 2021 2022 2023
5 6 7 8 9 10 11 12 13 14 15
$248,884,472 $234,328,555 $219,772,639 $219,772,639 $219,772,639 $219,772,639 $219,772,639 $219,772,639 $219,772,639 $219,772,639 $219,772,639
8.4.4 Depreciation Schedule (MACRS) Our design managers provided the schedule of depreciation to us. The following table provides the total amount of depreciation with a class life of 5 years. Table 8: Depreciation schedule. Year Year of operation MACRS 2009 1 20.00% 2010 2 32.00% 2011 3 19.20% 2012 4 11.52% 2013 5 11.52% 2014 6 5.76% 2015 7 2016 8 2017 9 2018 10 2019 11 2020 12 2021 13 2022 14 2023 15 -
D ($/yr) $133,003,624 $212,805,798 $127,683,479 $76,610,087 $76,610,087 $38,305,044 -
Taxes Saved ($/yr) $50,541,377 $80,866,203 $48,519,722 $29,111,833 $29,111,833 $14,555,917 -
Total Taxes Saved = $252,706,885 Total Depreciation = $665,018,119 Present 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. 34
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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 & Recommendations 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. 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 Acknowledgements We 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 References 1. 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. 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 12.1 Appendix I: Detailed Equipment Costing 12.1.1
I.1 Heat Exchanger Sizing Technique
This procedure was adapted from Seider: 1. Calculate R and S: Tcold ,out Tcold ,in m * Cp cold R cold and S m hot * Cp hot Thot ,in Tcold ,in 2. Calculate F for a 1 shell pass-2 tube pass heat exchanger: 1 S R 2 1 * ln 1 R S F 2 S R 1 R2 1 ( R 1) * ln 2 S R 1 R2 1
(#)
(#)
3. Compute log mean temperature T Tcold ,out Th,out Tcold ,in Tlm hot ,in Thot ,in Tcold ,out ln T T cold ,in h ,out
(#)
4. Calculate required Q and UA Q mhot Cp hot Thot ,in Thot ,out and UA
Q F * Tlm
5. Now starts the process of iteration. Guess a U (typically 100
(#)
BTU ), and then F * ft 2 * h
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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U
1 1 Ai t w Ai 1 hshellside Ao k w Am htubeside
(#)
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 L V U max K V
(#)
2. Compute Minimum Diameter
DMin
4QV U Max
(#)
3. Compute Liquid Height HL
t L QL 4 D 2
(#)
4. Compute Vapor Height Hv 4
(#)
L HL Hv
(#)
5. Compute overall Length
6. Compute L/D ratio 7. 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: CH 4 ( g ) H 2O( g ) CO( g ) 3H 2 ( g )
H 298 206 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 H 2O 2 4 CH 4
(4) SMR Sizing: Symbolic expression: aA bB cC dD (A=CH4, B=H2O, C=CO, and D=H2) 4 The lowercase letters denote the moles of each corresponding species. Proposed power law expression: rA kC A C B CC C D 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. 1.) 2.) 3.)
The new power law expression becomes: rA kC A C B (Fogler, eqn. 3-5) 6.) Now we need to determine Cj = hj(X) or concentration as a function of conversion, X. 5.)
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REPORT
7.)
Looking at the symbolic expression yields: A b B c C d D and from the above a
a
a
reaction we know that a 1,b 1,c 1, and d 3 . 8.) The new form of the reaction becomes the following: A B C 3D . Since we have a flow system in the gas phase C j F j (Eqn. 3-45) where v 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 9.)
10.)
CCH 4
FCH 4 v
F P T v v0 T 0 FT 0 P T0
and C H O FH O . 2
v
2
The design equation relating volumetric flow rates to molar flow rates is (Eqn. 3-41) where both temperature dependence and pressure dependence is shown. Now we can rewrite the above two concentration equations as: a. C CH 4
FH 2O FCH 4 and C H 2O F P T F P T v0 T 0 v0 T 0 FT 0 P T0 FT 0 P T0
b. If we assume isothermal reactor conditions and negligible pressure drops in the reactor we obtain the following: T T0 and P P0 implying that CCH 4
FH O FT 0 FCH 4 FT 0 and . C H 2O 2 v0 FT v0 FT
F c. Recall that FT 0 CT 0 (Eqn. 3-40) and so CH O H O CT 0 along with 2
v0
CCH 4
2
FT
FCH 4 CT 0 . FT
d. So in terms of concentration as a function of molar flow rates we have: i. CCH 4 FCH 4 CT 0 FT
ii. CH O 2
FH 2O FT
CT 0
iii. The above two expression can be plugged into the rate law giving: FCH 4 FH 2O 1. rA k CT 0 CT 0 FT FT 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 semipermeable is present. As a result, the design equation relating volumetric flow rate to conversion, neglecting pressure drops along with isothermal operation, is v v0 (1 X ) (Eqn. 3-44). a. y A0 (Eqn. 3-36) and = d c b 1 (Eqn. 3-23) = 3 1 1 1 2 a
a
a
48
REPORT b. Also
y A0 =
CCH , 0 . C A 0 (Example 3-7) which implies that y CH 4 , 0 CT 0 CT 0 4
CCH 4 ,0 . Since y A0 FA0 (Eqn. 3-39 or 3-40) the volumetric flow C FT 0 T0
c. And so = 2
rates cancel and C A0 y A0CT 0 (Example 3-7). d. Under isothermal and no pressure drop reactor conditions, C ( v j X ) (Eqn. 3-46). For Species A we have the following: C j A0 j 1 X CA
C A 0 ( A v A X ) where v A = -1 and A = 1. 1 X
e. Now that we have concentration as a function of conversion, C A C A0 (1 X ) 1 X
Let’s do the same for the second species. v B = b 1 and B FB 0 (Fogler, a
FT 0
Table 3-3).
Plugging in the following expressions (E3-6.4)
f. We can further simplify this expression to:
CB
CB
b C A 0 B X a 1 X
C A 0 B X 1 X
yields:
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. C (1 X ) C A0 B X i. rA k A0 1 X 1 X ii. Given conversion values and knowing the reaction order one can determine the reaction rate as a function of conversion. One must also know the reaction rate constant to completely size the reactor.
49
REPORT iii. If the assumption of the reaction rate is poor recall that the proposed power law would be: rA kC A C B CC C D iv. Plugging in the similar expressions for the products yields:
C (1 X) CA0 B X CA0 C X CA0 D 3X 1. rA k A0 1 X 1 X 1 X 1 X 2. Where C FC 0 (Table 3-3) and D F D 0 (Table 3-3).
FT 0
FT 0
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 to size the reactor is the following: a. FA0
dX rA dV
(Eqn. 2-15)
X b. Integrating to find volume yields: V FA0 dX (Eqn. 2-16) where the upper 0
rA
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: dX i. V F A0 0 C A0 (1 X ) C A0 B X k 1 X 1 X X
d. The final expression for the more complicated power law would be as follows: dX i. V F A0 0 CA0 (1 X ) CA0 B X CA0 C X CA0 D 3X k 1 X 1 X 1 X 1 X X
(6) Tube sizing and the number of tubes in the SMR:
50
REPORT 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.
dQ a. Heat flux equation: Ua(Ta T ) (Eqn. 8-44) dV b. Now for a tubular reactor (such as the SMR) with heat gain or loss we have the following energy balance equation:
dT Ua (Ta T ) ( rA )[H Rx (T )] (Eqn. 8-56) dV FA0 (i Cpi XC p )
This implies that dT g ( X , T ) must be coupled with the mole balance, dV
dX rA f ( X ,T ) . dV FA0
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:
F P T v v0 T 0 (Eqn. FT 0 P T0
3-41) but this
time we cannot neglect the temperatures terms. i. The equation becomes v v0 FT T . FT 0 T0
ii. Also recall that v v0 (1 X ) T where the pressure drop terms T0
have been neglected. iii. Going through the same procedure in step 11.) one obtains the following:
C (1 X ) T0 1. rA k A0 1 X T
C A0 B X T0 T 1 X
51
REPORT
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.
52
REPORT
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. #), yi P i xi Pi sat (#) 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. 1
0.9
0.8
0.7
0.6
0.5
0.4
0.3
0.2
0.1
0 0
0.1
0.2
0.3
0.4
0.5 Dew Line
0.6
0.7
0.8
0.9
1
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.
53
REPORT
Mathematica Code:
54
REPORT
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 Equation PCH3OH 0.5 1.3 k PCO PH 2 ( PCO PH 2 ) K eq PCH3OH PCO PH22 K eq
( K1 K 2 PCO K 3 PH 2 )3 ( K H 2 PH 2 ) 2 K CO PCO k
K CH 3OH PCH 3OH
(1 K H 2 PH 2 K CO PCO K CH3OH PCH3OH )3 ( K H 2 PH 2 )2 K CO PCO
k
K eq
K CH3OH PCH3OH K eq
(1 K H 2 PH 2 K CO PCO K CH 3OH PCH3OH )5
k
k
k
K1 PCH3OH (1 K1PCH3OH K 2 PCO K 3 PH 2 )
(1 K1PCH3OH
PCO K 2 PCO K 3 PH 2 ) PCO
(1 K1 PCH3OH K 2 PCO K 3 PH 2 )
Equation Description
Sum of Squared Residuals for Temp1/Temp 2
Agny & Takoudis
31.3036 / 9.00813
Villa et al. 4.2568 / 1.1053 Langmuir: Surface Reaction is RDS and all species adhere 4.42868 / associatively 1.8636 Langmuir: Surface Reaction is RDS and all species adhere 4.26236 / disassociatively 1.16466 Langmuir: Methanol Desorption is 4.46794 / RDS 1.46263 Langmuir: Carbon Monoxide Associative Adsorption is 8.60895 / RDS 2.69909 Langmuir: Carbon Monoxide Disassociative 9.33096 / Adsorption is 2.93866 55
REPORT
( K H 2 PH 2 ) 2 K CO PCO k
K CH3OH PCH3OH K eq
(1 K H 2 PH 2 K CO PCO K CH3OH PCH3OH )3
k
k
PH 2 (1 K1PCH3OH K 2 PCO K 3 PH 2 )
PH 2 (1 K1 PCH3OH K 2 PCO K 3 PH 2 )
K 2 PCO K 3 PH 2 k
2
K 4 PCH 3OH K eq
( K1 K 2 PCO K 3 PH 2 K 4 PCH3OH )5 1 c 2 0.5 PCO PCH PH 2 OH k c1 0.5 3 PCH OH PCO PH c3 3 2 PCH 3OH PCO PH22 K eq
( K1 K 2 PCO K 3 PH 2 K 4 PCH 3OH )3 PCO PH22
C3 C1 C 2 PCH 3OH k PH 2 PCO K eq PCH3OH PCO PH22 K eq
Natta et al10
4.25568 / 1.15103
Pasquon14 4.2086 / 1.0470 Simplified Power Law 3.818 / 1.0475
P K H 2 PH 2 K CH3OH PCH3OH
CO CO
10.8735 / 2.76844
K eq
C 2 C3 kPHC21PCO PCH3OH
1 K
Leonov et al14
PCH 3OH
( K1 K 2 PCO K 3 PH 2 K 4 PCH3OH )3
k
RDS Langmuir: Surface Reaction is RDS and Carbon Monoxide adheres 53.5992 / disassociatively 1.09881 Langmuir: Hydrogen Associative Adsorption is 5.28849 / RDS 1.56053 Langmuir: Hydrogen Disassociative Adsorption is 543.565 / RDS 133.225 Langmuir: Surface Reaction is RDS and Hydrogen adheres disassociatively 5.3448 / 1.5872
True Power Law 5.5417 / 1.0530
3
Equation Provided by Project Managers
0.252 / 0.351
56
REPORT
12.4 Appendix 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.
C B e (9.2951-0.6019(ln(S)) +0.0519(ln(S))
CB e(5.4866+ 0.13141(ln(Pc )) + 0.053255(ln(Pc ))
2
2
)
Eqn #
+ 0.028628 (ln(Pc )) 3 - 0.0035549 (ln(Pc )) 4 )
Eqn #
The Sizing Factor and the Power Consumption were estimated as follows: S Q H 0.5 Q H Pc 33000 p m
Eqn # Eqn #
where Symbol
Name
Units
Pc Q H
Power consumption Flowrate Pump head Density Fractional efficiency of the pump Fractional efficiency of the motor
Horsepower Gallons per minute Ft Lbs/gallon Dimensionless
p -0.316 + 0.24051 (lnQ) - 0.01199 (lnQ) 2 m 0.80 + 0.0319 (lnPc ) - 0.00182 (lnPc ) 2
Dimensionless
Now the purchasing cost, Cp, can be determined by the following equations: Pump
C p FT FM C B
Eqn #
Motor
C p FT C B
Eqn #
Where
Symbol
NameSeider
Value
57
REPORT
FT FM
Type Factor Material Factor
2.7 1.0
58
REPORT
12.4.2
IV.2 Storage Tanks
Unit Type
Size factor, S
Range of S
Floating roof
Volume, gal
10,000 - 1,000,000 gal
No. of gallons needed to be stored: No. of tanks required:
Unit Type Floating roof
Cost / Unit $ 430,558
16684322 17
No. of Units 17
Cost equation
Seider
Cp = 375*V^0.51
ten days' production at 1,000,000 gallons/tank
Total Cost $ 7,319,479
59
REPORT
12.4.3
IV.3 Compressors Purchase Cost, Cp Cp = Fd * Fm * Cb
I.
For electric motor drive, cast iron or carbon-steel construction purchase cost is obtained directly from Garrett and Walas (1988) 12 [Table 16.19] : 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 Gas turbine III.
1.15 1.25
Stainless steel Nickel Alloy
2.5 5
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))
60
REPORT
12.4.4
IV.4 Reactors
Steam Methane Reformer Type of Heater Reformer
Size factor, Q Btu/hr Heat absorbed
Range of Q Btu/hr 10 - 500 Million
Q value Btu/hr 1,179,779,845
Cp $ 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 + 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
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
61
REPORT
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 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 type Drum h.
L (ft) 9.84
Di (ft) 3.28
Op. Po (psig) 500
Pd (psig) 578
S (psi) 13100
E 1
Tp Ft 0.07
tv ft 0.06
ts ft 0.13
W lbs 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
62
REPORT
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 Psig sqft Fixed 1000 0 0 418,879 1 1 1.33
Cb $ 5,302,257
Cp $ 7,053,593
63
REPORT
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.
64
REPORT
12.4.6 I.
IV.6 Heat Exchangers 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 E-100 C-100 C-200 C-300
Heat duty Btu/hr 611311704 931173410 288527271 25764398
Cp water Btu/lbs/ºF 0.9990924 0.9990924 0.9990924 0.9990924
Tout ºF 110 110 110 110
Tin ºF 90 90 90 90
mflow rate lbs/hr 30593352 46600966 14439469 1289390.2
Area req. sqft 3165 4821 1494 133
A = Q/ (U*(∆Tlm)) Ucalc = Unit E-100 C-100 C-200 C-300
9657.4 HEX type Floating Floating Floating Floating
P psig 500 500 1000 5
a*
b*
2.7 2.7 2.7 2.7
0.07 0.07 0.07 0.07
A sq ft. 3,176 4,822 1,494 134
Fm
Fl
Fp
3.97 4.01 3.91 3.72
1 1 1 1
1.11 1.1128 1.3303 0.981204
Cb $ 36,297 47,001 24,644 14,209
Cp $ 160,509 209,820 128,132 51,876
65
REPORT
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 ($)
U-200 U-300
150577 163161
202066 214295
Cpl = the added cost for platforms and ladders depends on vessel diameter in feet and for vertical vessel, the length L. Vertical vessels for 3 < Di < 12 ft and 12 < L < 40 ft Cpl = 285.1* (Di)^0.73960 * (L)^0.7684 Unit
Di (ft)
L (ft)
Cpl ($)
U-200 U-300
12 9.5
14.5 16.5
13982 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
66
REPORT 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
U-200 U-300
Unit type
L ft
Di ft
Op. Po (psig)
Pd (psig)
S (psi)
E
tp ft
tv ft
ts ft
W lbs
Flash v. Flash v.
14.5 16.5
12 9.5
500 1000
578 1107
13100 13100
1 1
0.27 0.42
0.20 0.32
0.33 0.44
150577 163161
Cp = Fm* Cv + Cpl Vessel
Fm
Cv ($)
Cpl ($)
Cp ($)
U-200 U-300
1 1
202,066 214,295
13982 12991
216,047 227,286
67
REPORT
12.4.8
IV.8 Distillation Columns
Towers for 9,000 < W < 2,500,000 lb Cv = exp(7.0374+0.18255(ln(W))+0.02297(ln(W))^2) W (lbs)
Cv ($)
Di (ft)
L (ft)
Cpl ($)
6901
34399
3
9
2767
Towers for 3 < Di < 24 ft and 27 < L < 170 ft Cpl = 237.1* (Di)^0.63316 * (L)^0.80161
Weight Calculation:
W = pi()*(Di+ts)*(L+0.8Di)*ts*rho
Pd = exp(6.60608+0.91615(ln(Po))+0.0015655(ln(Po))^2) Po minimum = 10 psig. For pressures greater than 1000 psig, use Po = 1.1*Po and not the above equation. tp = (Pd*Di)/(2S*E-1.2*Pd) tp must be greater than a minimum value for rigidity based on the diameter (Table on Page 530) S = maximum allowable stress of the shell material at the design temperature in pounds per square inch E = fractional weld efficiency find minimum wall thickness, tmin, and ts must be greater than tmin OR FOR Vertical vessels take into account effects of wind and earth quake Tv = tp(0.75+0.22*E*(((L/Di)^2)/Pd)) Then use either tp or tv: Add corrosion allowance of 1/8 inch Then find ts = tp or tv + corrosion allowance Find W Unit type distill v.
L ft 9
Di ft 3
Op. Po (psig) 5
Pd (psig) 8
S (psi) 15000
E 1
Tp Ft 0.00
tv ft 0.00
ts ft 0.13
W Lbs 6901
68
REPORT
Cpl = 237.1* (Di)^0.63316 * (L)^0.80161 Vessel
Fm
Cv ($)
Cpl ($)
Cp ($)
Tower
2.1
34399
2767
75005
Unit Type Separators U-200 U-300 D-100
Base Cost / unit $
No. of Units
216,047 227,286 75,005
Total Cost $ 1 1 1
Total
216,047 227,286 75,005 $518,338
69
REPORT
12.5 Appendix V: Direct Permanent Investment & Total Capital Investment TCI = SUM
WC $ 28,829,285
TPI $ 750,696,181
Total $ 779,525,466
WC = SUM
½ Prod. Stor. 7,319,479
30 day Mft. $ 18,023,215
Spare $ 3,486,591
TDC 665,018,119
Startup 17,432,955
Total $ 682,451,074
5% TDC
17,432,955
TDC = SUM
Contingency 153,465,719
DPI 511,552,399
Contingency =
30% DPI, $
153,465,719
DPI
Site prep 30% TBM 52,298,865
Srvc. Fac. 25% TBM 69,731,820
TBM
Fab. Equip. 40,000,000
Proc. Machin. 300,659,100
MSR Catalyst SMR Catalyst
Cost $/lb 6 12
TPI = SUM
Startup =
Total $ 28,829,285
Total $ 665,018,119
TBM
Total $
372,088,759
511,552,399
Storage 7,319,479
Density Volume lb/ft^3 ft^3 90 17500 54 1000 Total All Catalyst $
Spare 3,486,591
Catalyst 12,623,589
Total $ 372,088,759
Total $ 9,450,000 3,173,589 12,623,589
70
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