DESIGN OF METHANOL PLANT

8/10/2019 DESIGN OF METHANOL PLANT

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DESIGN OF METHANOL PLANTEURECHA Student Contest Problem Competition 2013

Chan Wei Nian, [email protected] You, [email protected]

Department of Chemical & Biomolecular Engineering, National University of Singapore, Engineering

Drive 4, Singapore 117576, Republic of Singapore

EXECUTIVE SUMMARY

This report presents a techno-economic analysis of setting up a methanol plant with a capacity of50,000 MT/yr from natural gas feedstock of 31,000 MT/yr and 90% methane. This is performed in theEuropean context with due considerations for environmental sustainability. The plant is designedassuming 10-year plant life and 8000 hours of operation per annum.

The methanol process was developed according to open literature and industry standard, as described inSection 2; it consists of 5 units/sections: steam methane reformer, compression trains, methanol

synthesis reactor, hydrogen separation and methanol distillation. These were grouped into two mainsegments of steam methane reforming and methanol synthesis, and were independently optimized inSections 3.1 and 3.2 using Aspen HYSYS v7.2 by varying important process parameters withinreported ranges. This is followed by a plant-wide heat integration study with Aspen Energy Analyzerv7.2, described in Sections 3.3 and 4.3. Operating conditions favouring better overall process profit ofthe order of USD$10 6 are reported in Sections 4.1 and 4.2 for the ideal case ignoring operatingconstraints, and the base case for economic analysis was chosen based on practical industry constraints;the reforming section utilizes an excess steam to carbon ratio of 2.7, inlet pressure of 2000 kPa andfurnace outlet temperature of 900C and while the methanol synthesis loop employs a reactortemperature of 255C, compressor discharge pressure of 8010 kPa and a stoichiometric ratio of 9.035 ,as summarized in Section 4.4.

The fixed capital investment, cost of manufacturing and revenue for the base case is 31,562,300 , 29, 520,600 and 43,952,670 respectively. The payback period is 3.2 years and the net present value(NPV) of the investment is 32,840,270 . The breakeven price for methanol is 0. 3285/kg and thediscounted cash flow rate of return is 26.02% . The former is most sensitive to natural gas, hydrogenand methanol prices, as shown in Section 4.5. In Section 4.6, Monte Carlo simulations performedaccording to expected price fluctuations of natural gas and methanol obtained in the past 10 yearsreported a 7.5% probability of making an overall investment loss.

The feasibility of substituting biogas for natural gas was analyzed for the given feedstocks in Section4.7. Several options were explored, and the highest savings in net present value of 1,299,770 as wellas reduction of 2074 MT/yr of CO 2 emissions were obtained for the case of utilizing the non-liquidfeeds. Options to capture carbon emissions from the plant as well as other carbon sources wereexplored, and the most suitable technology for the plant was found to be absorption using mono-ethanol amines, as elaborated in Section 4.8.


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1. INTRODUCTION

Methanol demand in the global market is burgeoning with a 23% increase from 2010 to 2012 to 61million tons and an expected increase to 137 million tons in 2022 [1]. However, its raw material:natural gas faces both supply limitation and volatility; as a non-renewable resource, natural gas isexpected to deplete as early as 37 years depending on the level of conservativeness [2] while gasinfrastructure in Europe is fragmented and inconsistent [3]. At the same time, environmental concerns

for carbon emissions necessitate the integration of carbon capturing technologies. These usher a newage of sustainable design in this age-old process, and it will be reviewed in this report. A base case forthe methanol synthesis was developed and optimized, followed by case studies on biogas substitutionand alternative carbon capture opportunities.

2. PROCESS DESCRIPTION

For the base case, the methanol synthesis is generally split into steam reforming, methanol synthesis,hydrogen separation and methanol purification, which are described in this section.

Steam reforming: This is an overall endothermic reaction involving catalytic conversion of methaneand steam into syngas at high temperatures above 800 oC [4], medium pressures of 15-30 bar [5] andsteam to carbon (S/C) ratio of 2 to 4 using nickel catalysts. Steam reforming is limited by equilibrium[6], and higher conversion is achieved by increasing temperature, lowering pressure and higher S/Cratio. It usually occurs in nickel catalyst-packed tubes located in the radiant section of a furnace.

Steam Reforming: (H25C = 206kJ/mol) ----------------- (1)Water Gas Shift: (H25C = -41kJ/mol) ----------------- (2)

Methanol Synthesis: Syngas can be catalytically converted to methanol via an overall exothermicreaction at medium temperatures of 210- 270C and high pressures of 50 -100 bar [7], over copper-alumina catalysts. The presence of the water-gas shift reaction necessitates the use of a modified

stoichiometric ratio defined below, and it is typically in the range of 9-10 [4].Methanol Synthesis: (H25C = -91kJ/mol) --------------- (3)

Modified Stoichiometric Ratio: -------------------------------------------------------- (4)

Methanol synthesis reactors are designed to remove reaction heat via cooling service fluids such asunheated reactants or boiler feed water as well as feed quenching [8]. Isothermal cooling in shell andreactor tube setups was modeled under conditions known to approach equilibrium [4], and it canachieve higher conversions with lower temperatures, higher pressures and stoichiometric ratios.

Methanol Purification: Methanol synthesis products containing methanol and syngas are flashed to

separate unconverted light ends and crude methanol. Crude methanol is distilled to separate remaininglight-ends, methanol at 98%wt and water, in an atmospheric column with a partial condenser.

Hydrogen Separation: The flashed light-ends contain excess hydrogen; to avoid recycleaccumulation, these must be purged for furnace fueling or purified for hydrogen credit via pressureswing adsorption, cryogenic distillation or membrane separation [9] with the latter being cheaper at thelow recoveries and purities required [10]. For this, polyimide membranes are fabricated as hollow fibertubes in a shell, and separation is achieved by partial pressure differences.


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3. METHODOLOGY

The process was simulated using Aspen HYSYS v7.2, and individual units were designed to sufficientdetail for costing based on the module factor approach [11]. Expected revenue from methanol( 0.413/kg [12]) and hydrogen ( 8.04/kg [13]) sales was factored in for a plant-wide optimization. Themethodology for design and costing for each unit for the optimization is elaborated below.

3.1 Steam Reforming Section OptimizationThis section focuses on the reformer furnace with the reaction modeled to reach equilibrium and thecosting based on energy used. This is independent of downstream synthesis and is optimizedseparately. The variables affecting this section are the pressure, furnace temperatures (i.e., pre-heat andoutlet temperatures) and S/C ratio; these affect costs of individual sections/units, and their expectedeffect on profit is shown in Table 1.

Table 1: Expected effect of i ncrease in design vari ables on profi t due to costs of individual sections/uni ts

Design Variable Energy Recoveryfrom Natural Gas

ConvectionSection

RadiantSection

Compression Conversion

Pressure Uncertain Uncertain

Pre-heat Temperature NA Uncertain Outlet Temperature NA NA Uncertain S/C Ratio NA

En ergy Recovery: Natural gas is expected to be delivered at 75 bar within European pipelines [3] andmust be expanded via a valve or turbine for energy recovery to the desired pressure; the latter can beachieved with radial gas turbines at 75-88% efficiency [14] together with a minimum motor generatorefficiency of 95% mandated by EU Directive 640/2009 [15]. The turbine was priced based on energyrecovery rate and capital costs.

Steam Refor mer F ur nace: Natural gas is mixed with steam and preheated in the convection section

before reaction in the radiant section. The duties are summed, and used for furnace capital and energycosts. Catalyst life is assumed to be within industrial norm, and is priced based on USD$0.55 per kilomole of natural gas processed [16]. A 95% approach to equilibrium was found to best match industrialdata from [17], and this was used in the reactor simulations

Water L et-D own Vessel: Excess steam is condensed and removed from the reformer products, in aknock-out vessel to ensure dryness for the compression train. The design is based on the waterthroughput to allow for 7.5 min of residence time , and the vessel volume is used for capital costing.The condensate is pumped up to process pressure and reheated as recycle steam to the reformer.

M ake-up Compressor: A separate compression train for syngas products (i.e., fresh feed) is designed

because it has less variability than the recycle compressor and so requires less capacity allowance. Forthe discharge pressure of 50-100 bar and required flow of about 0.410 m 3/s, reciprocating compressorsare suitable and are used for costing at 75% efficiency. To avoid adiabatic temperature rise beyond themaximum temperature of 480 K [14], two-stage compression is used with inter-stage cooling. Thiscooler is designed using heat transfer area obtained using typical overall heat transfer coefficients andlog mean temperature difference for cooling water service [18].


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3.2 Methanol Synthesis Section OptimizationDue to low equilibrium conversion tomethanol, recycle of syngas is necessary. Theeffect of reactor temperature, pressure andstoichiometric ratio of reactants on

profitability of this section is shown in Figure1. For modeling and optimization purposes,common reaction conditions of 210-270C and5-10 MPa were used. Costs involved forcooling reactant products in the crudemethanol flash, pre-heating for distillation andcolumn duties are calculated based on a

preliminary setup, and optimized later using pinch analysis together with the steamreforming heat exchangers.

Figur e 1: Schematic f or i nter-l inked eff ects of varyingparameters i ndependent vari ables (black), parameters

aff ectin g cost (shaded pink), calcu lated parameters (blue)

M ethanol Synthesis Reactor : The recycle and fresh syngas streams are heated to the reactiontemperature with reaction heat and an optional pre-heating or cooling; this is possible in shell and tubereactors [19]. The reactor was sized based on heat transfer area [20]. Catalyst cost, based on itsreplacement norms, is USD$0.0513 per kilo mole methanol, for isothermal reactors [4].

M ethanol Pur if ication: The crude methanol flash vessel was sized similar to the water let-downvessel, and the distillation column was optimized based on typical feed of 78-80%mol methanol at70oC. Trade-off between reflux ratio and number of stages was analysed by costing the column,condenser, reboiler, reflux drum and reflux pump, and the optimum number stages was found to be 27for reflux ratio from 0.4 to 0.6 (Figure 2) . This design is deemed optimal and used for all runs in theoptimization study.

Figure 2: L ocal optimi zation for distil lationcolumn

Figur e 3: H ollow F iber M embrane Separation M oduleSource: [ 21]

H ollow F iber M embrane M odule: Hydrogen separation is as shown in Figure 3 to recover hydrogenfrom the purge stream. Feed to permeate pressure ratio of 6 was used to minimize recompression ofhydrogen [9], and this sets the limit for hydrogen recovery to 95%. The annualized capital cost increaseis calculated to be in the range of USD$2,442.9 for a 0.01 fraction increase in recovery, whereas thegain in revenue was USD$365,600 for the same increase. As such, hydrogen recovery was maximizedat 95% for each optimization run. With the hydrogen flux constant, only the membrane cost will vary

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20 25 30 35 40 45 50 A n n u a l i z e d C o s t s

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Number of Stages


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according to required membrane area (USD$21/m 2 [22]) obtained from changes in log mean pressuredifference as a result of changing reactor pressure levels for each run.

Recycle Compressor: The recycle compressor was designed for approximately 0.495 m 3/s of flow withreciprocating compressors as the suitable choice [14]. Due to a smaller pressure difference, single stagecompression is sufficient, and the compressor was priced using duty and 75% efficiency.

3.3 Heat IntegrationIn order to improve the energy efficiency of the methanol plant, heat integration is performed torecover process heat, using Aspen Energy Analyzer (AEA) v7.2 coupled with the Aspen HYSYS v7.2simulation of the methanol process. The Utility Composite Curve from pinch analysis is shown inFigure 4 with minimum approach temperature (dT min) contribution for various streams listed in Table2. From Figure 4, the shifted process pinch temperature is 135.3C and the overall heating and coolingtargets are 15.9 MW and 13.3 MW respectively. The area target returned by AEA is 2345 m 2 for 1 shell

pass and 2-tube pass heat exchangers.

Due to the lower price of natural gas compared to steam, AEA program recommends generating highamounts of steam from the fired heater as seen in Figure 4. However, the credit from exporting steam

may be diminished if there is lack of demand for steam. Hence, the plant will not generate excess steamfrom natural gas.

Figur e 4: Util ity Composite Curve for M ethanol Pl ant Process

Table 2: dTmin Contri bution f or Each of the Stream Types

Stream Type dT min contribution (K)Condensing/vaporizing 2.5

Liquid 5.0Gas 7.5

Subsequently, a Heat Exchanger Network (HEN) is designed with the following operating constraints:(1) HP steam for SMR is to be generated in the furnace to ensure a steady supply of feed for the

process; (2) the reforming heat of reaction must be supplied by a furnace to maintain optimal reactiontemperature as process heat exchange potentially introduces fluctuations; and (3) crude methanoldistillation column condenser and reboiler are to be serviced by utilities to ensure controllability ofcolumn operations.


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4. RESULTS AND DISCUSSION

Simulations were run for variations in operating parameters beyond the practical constraints applied inindustry; this was to explore the potential cost savings from surpassing operating constraints. Data wasextracted and costed according to Section 3 with detailed calculations in Section A4.2. The optimal

process was then chosen after applying the appropriate constraints and the heat integration study wasconducted to obtain the base case.

4.1 Optimization of Steam Methane ReformingMethanol profit excluding methanol synthesis costs was found to increase with higher temperature,lower S/C ratio and pressure while pre-heat temperature was found to have negligible impact (Figure5). The results show that significant savings are possible in the magnitude of USD 10 6/yr if equilibriumconversion can be achieved at the extreme ends of the optimization range. Constraints of temperature at900 oC due to mechanical strength limits of tubes [5], pressure at 20 bar and S/C at 2.7 to preventexcessive coking [4] constrains the reforming section to operate at these limits and these are chosen asthe base case conditions.

Figur e 5: Var iati on of expected profi t with S/Cand temperatur e

F igur e 6: Vari ation of methanol profi t with pressure andR: Red sur face plot for r eaction temperatur e of 211 oC;

Green dots for 222 oC, Red for 233 oC, M agenta for 244 oCand multi -colour surf ace plot for 255 oC

4.2 Optimization of Methanol Synthesis LoopFigure 6 shows that optimal methanol profitwithout reforming cost increases with lowertemperature and higher S/C ratio. However, toensure that the equilibrium can be achieved,temperature was constrained to 255 oC [23] andS/C from 9 to 10 [4]. At these conditions, effect of

pressure on cost reaches a plateau at 8900 kPa, asshown in Figure 7, and this is chosen as theoperating pressure.

Figur e 7: Var iati on of expected profi t with Pressureat reactor temperature of 255C and R=9.063

$3,360.00

$3,370.00

$3,380.00

$3,390.00

$3,400.00

6500 7000 7500 8000 8500 9000 9500 A n n u a l i z e

d P r o

f i t p e r

M e t

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l , U S D

$ / k g

Pressure, kPa


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4.3 Heat Integration ResultsFollowing the practical considerations for heat integration, the HEN shown in Appendix A1.1 is

proposed. The heat recovered from the hot syngas is used to preheat methanol reactor feed as well as tovaporize recycled water/condensate for steam reforming reaction. However, natural gas preheating forthe SMR reactor is not subjected to process heat exchange as this will cause difficulty in achievingsteady state during start up; hence the duty will be provided by LPS. Crude methanol feed to thedistillation column is preheated using LPS as the methanol reactor effluent stream is used to generateMPS and LPS for plant-wide heating purposes and there are no other hot process streams in close

proximity for process heat exchange. The performance of the proposed HEN is summarized in Table 3. The targets are not fully met due to the operating constraints stated in Section 3.3 as well as practicalconsiderations for process heat exchange due to the physical location of the streams, in which caseutility steam is generated for energy recovery.

Table 3: H EN Performance Summary

Overall targets Target CurrentHeating (MW) 15.9 26.8Cooling (MW) 13.3 24.2Total area (m 2) 2345 2078

Utility targetsFired Heater (MW) 32.7 23.4

HPS (MW) 0 0MPS (MW) 0 3.5LPS (MW) 0 0.3

HPS generation (MW) 7.3 0MPS generation (MW) 8.1 7.1LPS generation (MW) 1.3 0.9Cooling Water (MW) 13.3 16.2

4.4 Base Case PlantThe integrated results from the process optimization and pinch analysis are summarized in Table 4 andthe equipment schedule is shown in Section A4 with the PFD and stream data for the base case inAppendix A1.1.

Table 4: Base Case Conditi ons for M ethanol Plant

Steam Methane Reforming Methanol Synthesis LoopTemperature 900C 255C

Pressure 2000kPa8010kPa

(adjusted after heat integration)S/C Ratio / R Value 2.7 9.035

UtilitiesFurnace Duty: 22,998kWCooling Water: 5493.5kW

MPS Generation: 3,672kWCooling Water: 10,732kW

Compression Duty: 4,128kWConversion 85.0% Methane Conversion 93.7% Syngas Conversion

With the capital and utility costs, additional direct, general and fixed manufacturing costs were factoredusing multiplication factors on the process data [24]. Grassroots cost were used to account for off-siteand auxiliary costs. A straight line depreciation method typical of Europe [25] was used in generatingthe cash flow diagram. Detailed calculations can be found in Section A4.2.


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Figur e 8: Cumul ative discounted cash f low f or 10%discount, 30% tax and 10 years plant li fe

F igur e 9: Br eakeven methanol pr ice for dif ferentpri ce of natural gas

The fixed capital investment required is 31,562,300, and cost of manufacturing is 29, 520,590 fornatural gas price of 0.361/kg [26] . Taking methanol price to be 0.413/kg and hydrogen price

8.04/kg, the revenue is 43,952,670 . The payback period for the base case is 3.2 years from the timeof plant start-up after 2 years of allocated construction time. The net present value (NPV) of the

investment is 32,840,270 , as shown in the cash flow diagram in Figure 8. The breakeven price formethanol is 0. 3285/kg holding other prices constant and discount rate at 10%.

4.5 Sensitivity AnalysisTo understand the susceptibility of profits to the various parameters affecting the plant costs, asensitivity analysis was conducted. 10% changes in key parameters were made, and the change in NPVwas measured as a percentage of the NPV reported above; these results are reported in Table 5.

Table 5: Sensiti vity of NPV to 10% changes in key parameters

10% Change in % Change in NPVMethanol Price 25.8

Natural Gas Price -24.3Hydrogen Price 27.2

Boiler Feed Water Price -0.6Capital Cost -7.98

As observed, the production costs are most susceptible to changes in the methanol, hydrogen andnatural gas prices as expected. It is much less sensitive to the raw material price of boiler feed watersince it represents a smaller portion of overall costs, and capital cost changes have a smaller effect

because of its diluted effect over the plant life.

4.6 Economic Analysis

From cost calculations of the base case, a range of breakeven prices for the given discount rate can becalculated for each variation in natural gas price. This is summarized in Figure 9. To understand theimpact of market volatility in natural gas and methanol, price trends were obtained from BP andMethanex market surveys for the past 10 years [27] [28]. The price fluctuations were quantified by

performing a polynomial regression on the data points and determining the deviations from theexpected trend. Figure 10 and Figure 11 show the regression data:

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Figur e 10: N atural gas pri ces for 2001 to 2011 fr omBP Report

F igur e 11: M ethanol pr ices for 2002 to 2013 fromM ethanex Report

Although, price estimation involves a more complex interaction of supply and demand, these plots provide a statistical estimate for the expected volatility for initial projections. From these, natural gaswas determined to fluctuate from 16.6% to -19.7%; for methanol, 17% to -22%. To estimate the rangeof variation in sales volume, the world methanol capacity was taken as an index to obtain a fluctuation

Figure 12: F requency Distri bution of NPVsgenerated fr om random permutations of volati le

parameters

of 40% in production volume [29]. Using thesevariations, Monte Carlo simulations were conductedassuming triangular distribution of parameterswithin the assumed fluctuations, to calculate NPVfor each random permutation. Figure 12 displays the

plot of NPV frequency distribution. It can be seenthat, despite the price and volume fluctuations, thereis only a 7.5% chance of making a negative NPV.This indicates that the cash position of the plant issafe from the volatility of the market for the given

price fluctuations. Standard deviation of thedistribution in Figure 12 is 22.8mil.

4.7 Case 1: Biogas substitutionTo investigate the impact of biogas, yield, productivity and throughput data for the available foodwaste were obtained from the literature, and these are summarized in Table 6:

Table 6: Bi ogas Kinetic Data

Organic Load Methane Yield Productivity,m 3.CH 4/m

3/dayThroughput ofOrganic Waste

Reference

Corn Silage 181 L/kg.TS 0.73 5400 MT/yr [30] Corn Grain 300 L/kg.TS 5.4 2600 MT/yr [31] Wastewater

(Slaughterhouse)240 L/kg.COD 0.72 150,000 m

3/yr5000 mg.COD/L [32]

Animal Waste(Chicken Dropping)

207 L/kg.VS 0.37 30,000 MT/yr40% TS, 85% VS [33]

Pre-treated WheatStraw

349L/kg.TS 3.11 1257.8 MT/yr40% TS, 85% VS [34]

y = -0.0691x 2 + 1.7619x + 0.7863R = 0.841

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U S D

$ / G J

Index Time

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- 0.1022x2

+ 7.0285x + 120.99R = 0.3264 0

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These data represent an average of the relative crop rates; actual values are scattered about theseestimates [34] [35]. Corn grain is clearly more productive, which allows the reactor size to be smaller.This effect on capital cost was calculated with costing exponential factor of 0.7 taken from [36] for thereactor and auxiliary facility costs as well as compression requirements and storage. The operating costwould vary with the amount of load transported and the travel duration [37] while heating costs wereneglected due to surplus heat streams from the main process. The costs were annualized across the

plant life and an average cost for each kg of biomethane is calculated. From this, it was concluded thatwastewater is unprofitable to use in isolation because the transportation of the dilute organic load isuneconomical; the final cost of biomethane from such production is 3 times the natural gas price( 10.09/kg vs 0.361/kg) . For solid waste, the savings in transportation allows biogas to be produced ata lowest possible productivity value of 0.118 m 3 of CH 4/m3/day while maintaining biogas cost to becomparable to natural gas price; this productivity is easily achieved in industry and literature.

Economic Anal ysis: From Table 6, several options to utilize the available waste for biogas productionwere analysed to maintain positive NPV as shown in Table 7. Option 1 is the base case without any

biogas, Option 2 maximizes NPV by only utilizing the corn grain, Option 3 maximizes the productionof biogas, Option 4 maximizes NPV with a substitute for corn of plant origin, and Option 5 is thescenario where all solid waste was processed to take advantage of the cheaper transport. The effect ofeach biogas substitution option on the Monte Carlo simulated probability of earning a negative NPV isalso reported. This was done by reducing the variability and price of natural gas as a result of partialsubstitution by biogas and assuming biogas does not fluctuate in price. As seen in Table 2, the biogassubstitution is economically justified, and the best improvement in NPV is from Option 5.

Table 7: Biogas Production Scenar ios

Description Option 1 Option 2 Option 3 Option 4 Option 5Bi omethane Produced, KT/yr 0 0.439 2.027 0.439 1.995Bi omethane Cost, USD$/kg 0.4689 0.0981 0.4689 0.0900 0.2714

kg.CO2eq/kg.methanol 0.6587 0.6504 0.6174 0.6180 0.6173NPV Change, USD$ 0.00 259,732.00 0.00 28,153.00 1,688,303.00

% of -ve NPV 7.52 7.33 7.99 7.72 6.53

Envir onmental Analysis: The carbon emissions data for each of the above options was also analyzedto consider their environmental impact. This was developed by defining the system boundary to includethe natural gas production facility from fossil fuels, the treatment facilities of the respective organicwastes, the alternative biogas production facilities for each respective organic wastes and the effect oftransportation. The effect of post-digestion sludge as well as the methanol plant was deemed to stayconstant for all scenarios. Greenhouse gas emission data were adapted from life cycle assessment

studies from [38] and [39] with global warming potential equivalency taken from [40]. Appendix A1.3 provides an elaborate write-up on the environmental analysis. As seen in Table 7, Options 3 and 5 arethe scenarios with the most carbon reduction.

Choice of Biogas Substitu tion Opti on: The best option is clearly Option 5. It yields the highest NPVincrease, has a lower probability of making a loss and has superior levels of carbon reduction. This is

because the transportation of solid wastes is both economical and environmentally sustainable, andhence it should be capitalized.


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Food for Energy: Corn has the best productivity and yield, and maximizing its use in biogas production for Option 2 also yields some carbon reduction. However, its competition for arable landmay cause overall poorer performance in other life cycle parameters such as human health, ecosystemquality and resource depletion [41]. This also spills over to market pressures to increase food price.This should and can be avoided with readily available alternatives such as straw, fruits and vegetablewaste or plant residues [42]. For example, wheat straw pre-treated with alkaline can achieve very high

productivity similar to that of corn [34]; accordingly, Options 2 and 4 clearly show a similar performance to the high productivity of corn grain. Alternatively, co-digestion of fruits and vegetablewaste with manure or slaughterhouse wastewater can yield up to 51.5% increase in biogas yield [43].Co-digestion of wheat straw and chicken manure is also possible [44]. Therefore, these options ofdirect substitution and co-digestion are viable alternatives.

4.8 Case 2: Carbon Capture TechnologyThe methanol production process leads to carbon dioxide effluents from its purge stream (which can berouted to flare systems or into the furnace for additional fuel) and from furnace flue gas stream. Thesemust be effectively captured to the desired flow of 200 MT/day required for the production rateenhancement. Table 8 summarizes the effluents specifications for re-use.

Table 8: F lue Gas and Target Composit ion

Available Flue Gas Composition Target Feed CompositionMass Flow 138 MT/day 200 MT/dayPressure ~ 1 bar 22.2 bar

Temperature ~ 200 C ~ 200 CComposition in Mole Fraction unless otherwise stated

Water 0.171 < 600 ppmCO2 0.0893 > 0.95

Oxygen 0.0175 < 10 ppmNitrogen 0.723 < 0.04

An additional 62 MT/day of CO 2 is necessary, and this can be obtained from fossil-fuel firing units;these include power plants or other chemical plants in neighbouring complexes depending on transportconstraints. The CO 2 effluent stream would need to be purified to prevent introducing excess inertgases such as nitrogen that would strain the recycle loop; to conform to certain standards, the IPCCstandard will be used [45]. The purification can be setup as a pre-combustion, post-combustion or oxy-recycle [46]. The post-combustion method was chosen due to its flexibility of installation, suitabilityover the oxy-recycle method specifically for natural gas fuels in the furnace as well as the need to

preserve carbon within the syngas flow for process purposes. For this method, CO 2 is removed after itleaves the furnace and options for performing the purification can be chosen as shown in Table 9.

Table 9: Summary of Di ff erent Carbon Dioxide Puri fi cation Technologies; See Text f or References

Absorption(MEA)

Adsorption MembraneSeparation

CryogenicDistillation

Energy Usage, GJ/kg.CO 2 3.5 3.23 0.605 High VacuumHigh

RefrigerationCost, /kg.CO 2 28 44 80

Plant Life and Discount Rate 20 yrs, 10% 25 yrs, 7% 25 yrs, 5%


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Cryogenic distillation is generally taken to be too expensive due to its extreme conditions [46]. Data formembrane separation from [47] reports the lowest energy, unlike the excessive heating or re-compression in liquid absorption and pressure swing adsorption (PSA); however, it is most expensivedue to the need for multi-stage membrane modules and compression for the low pressure carbondioxide feed. PSA energy usage reported in [48] is slightly better than that from absorption via mono-ethanol amine (MEA) [49] despite the reported cost for the latter [50] to be cheaper than the former asshown by [51]. From this, it is evident that there is a trade-off between profitability and energy use.

In consideration of economics and environmental energy concerns, liquid absorption with MEA isideal. This is due to it being the cheapest technology and most convenient to integrate in terms ofenergy use; it only requires once-through compression as opposed to cycles in adsorption, and itsheating duty can be easily provided from the excess heat in the plant, hence mitigating its energy

penalty. In addition, membrane operation is inherently more prone to aging and poisoning issues suchas membrane wetting or plugging [52], therefore it is also superior in technical robustness. With this inmind, the liquid absorption method was chosen.

5. CONCLUSIONS

As shown in this preliminary process development, the base case for methanol production is profitableto a company setting up operations within the European context. Furthermore, the traditional process ofmethanol production holds potential for increased cost savings if operating constraints in reactor designcan be overcome to achieve equilibrium at more extreme conditions. In the context of sustainabledevelopment, options to substitute the traditional natural gas feed with renewable biogas as well ascarbon capture were evaluated and shown to be economically justified while reducing carbon emissionsfor the plant. This supports the relevance of the traditional methanol production process in a moderntechnical, economic and environmental perspective.

In this study on methanol process development and design, steam methane reforming, methanolsynthesis, distillation and absorption of carbon dioxide are well-established technologies that can bedesigned and implemented readily. On the other hand, hydrogen separation using membranetechnology is a relatively newer technology, and biogas generation requires further testing on thespecific feedstock to affirm the suitability of the feed. Further work on these is recommended beforeimplementing the proposed methanol process with biogas substitution for part of the natural gas feed.

Figure 13: L iquid Absorption via M EASystem

Figure 13 shows the schematic representation of the

chosen carbon dioxide purification system. Flue gas is pressurized as required, CO 2 is absorbed by lean MEAsolvent, and then it is boiled off in the regenerator columnto release the high pressure CO 2 as feed to the steamreforming section. Steam required for regeneration/reboiler is readily available from the main process steamfacilities.


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A1. APPENDICES

A1.1 Appendix A: Process Flow Diagram and Stream Data TableA1.2 Appendix B: HYSYS Simulation and Optimization DetailsA1.3 Appendix C: Biogas Life Cycle System Boundary and Details

A2. METHANOL PLANT SIMULATION FILE

A3. AEA HEAT INTEGRATION FILE

A4. MS EXCEL SPREADSHEETS

A4.1 Optimization SpreadsheetA4.2 Cost Spreadsheet

REFERENCES

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[14] H. Silla, Chemical Process Engineering, New York: Marcel Dekker Inc, 2003.

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[17] A. Acuna, C. Fuentesa and C. A. Smith, "Dynamic Simulation of a Furnace Steam Reforming of Natural Gas,"Tecnologia y Futuro , 1999.

[18] G. Towler and R. Sinnott, Chemical Engineering Design, Burlington: Elsevier, 2008.

[19] W.-H. Cheng and H. H. Kung, Methanol Production and Use, New York: Marcel Dekker Inc, 1994.

[20] S. H. Fogler and N. M. Gurmen, "Essentials of Chemical Reaction Engineering," University of Michigan, 2007.[Online]. Available: http://www.umich.edu/~essen/. [Accessed 12 April 2013].

[21] Federal Aviation Administration Fire Safety, "Hollow Fiber Membrane Gas Separation," 8 January 2013. [Online].

Available: http://www.fire.tc.faa.gov/systems/fueltank/hfm.stm. [Accessed 23 April 2013].[22] A. A. Al-Rabiah, K. D. Timmerhaus and R. D. Noble, "Membrane Technology for Hydrogen Separation in Ethylene

Plants," in 6th World Congress of Chemical Engineering , Melbourne, 2001.

[23] M. Shahrohki and G. R. Baghmisheh, "Modeling, simulation and control of a methanol synthesis fixed-bed reactor,"Chemical Engineering Science, pp. 4275-4285, 2005.

[24] R. Turton, R. C. Bailie, W. B. Whiting, J. A. Shaeiwitz and D. Bhattacharyya, Analysis, Synthesis, and Design ofChemical Processes, New Jersey: Pearson Education, 2012.

[25] B. W. Mast, "CRS Report for Congress: Selected International Depreciation Rates by Asset and Country,"Congressional Research Service, 2007.

[26] Europe's Energy Portal, "Natural Gas Prices," 2013. [Online]. Available: http://www.energy.eu/. [Accessed 23 April

2013].[27] BP, "Natural Gas Prices," 2013. [Online]. Available:

http://www.bp.com/sectiongenericarticle800.do?categoryId=9037181&contentId=7068643. [Accessed 23 April 2013].

[28] Methanex, "Methanol Price," 28 March 2013. [Online]. Available:http://www.methanex.com/products/methanolprice.html. [Accessed 23 April 2013].

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[30] E. Zauner and U. Kuntzer, "Methane Production from Ensiled Plant Material," Biomass, pp. 207-223, 1986.

[31] B. K. Richards, R. J. Cummings and W. J. Jewell, "High Rate Low Solids Methane Fermentation," Biomass and Bioenergy, pp. 249-260, 1991.

[32] M. R. Johns, "Developments in Wastewater Treatment in the Meat Processing Industry: A Review," BioresourceTechnology, pp. 203-216, 1995.

[33] S. Sakar, K. Yetilmezsoy and E. Kocak, "Anaerobic digestion technology in poultry and livestock waste treatment - aliterature review," Waste Management and Research, pp. 3-18, 2009.

[34] N. V. Gunaseelan, "Anaerobic Digestion of Biomass for Methane Production: A Review," Biomass and Bioenergy, pp.83-114, 1997.

[35] M. Demuynck and E. J. Nyns, Biogas Plants in Europe, Dordrecht: D Reidel Publishing Company, 1984.

[36] B. G. Yeoh, "A Technical and Economic Analysis of Heat and Power Generation from Biomethanation of Palm OilMill Effluent," in Electricity Supply Industry in Transition: Issues and Prospect for Asia , Selangor, 2004.

[37] R. Palm, "The economic potential for production of upgraded biogas used as vehicle fuel in Sweden," ChalmersUniversity of Technology, Gteborg, 2010.

[38] P. Borjesson and M. Berglund, "Environmental systems analysis of biogas systems Part I: Fuel-cycle emissions," Biomass and Bioenergy, pp. 469-485, 2006.

[39] J. Bacenetti, A. Mena, M. Negri, P. Cantarella, S. Bocchi and M. Fiala, "Energetic and Environmental Balance of aBiogas Plant in Northern Italy," Regione Lombardia, Milano, 2010.

[40] IPCC, "TS.2.5 Net Global Radiative Forcing, Global Warming Potentials and Patterns of Forcing," 2007. [Online].Available: http://www.ipcc.ch/publications_and_data/ar4/wg1/en/tssts-2-5.html. [Accessed 26 April 2013].

[41] C. Jury, E. Benetto, D. Koster, B. Schmitt and J. Welfring, "Life Cycle Assessment of biogas production by


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A1.1 APPENDIX A


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Stream Number 1 2 3 4 5 6 7 8 9 1Temperature, C 41.369 76.831 11.995 219.137 218.284 219.951 900.000 484.962 136.754 35.00Pressure, kPa 7500.000 6750.000 2222.000 2222.000 2222.000 2222.000 1799.820 1619.838 1457.854 956.49Molar Flow, kmol/h 222.222 222.222 222.222 540.039 263.500 276.539 1177.615 1177.615 1177.615 1177.61Mass Flow, kg/hr 3876.800 3876.800 3876.800 9730.254 4746.979 4983.275 13607.224 13607.224 13607.224 13607.224Composition, mol fraction Methane 9.000E-01 9.000E-01 9.000E-01 1.915E-09 1.915E-09 3.739E-09 3.122E-02 3.122E-02 3.122E-02 3.122E-02Ethane 1.000E-01 1.000E-01 1.000E-01 1.214E-15 1.214E-15 2.370E-15 7.889E-07 7.889E-07 7.889E-07 7.889E-07Water 0.000E+00 0.000E+00 0.000E+00 9.999E-01 9.999E-01 9.998E-01 2.395E-01 2.395E-01 2.395E-01 2.395E-01Hydrogen 0.000E+00 0.000E+00 0.000E+00 5.156E-06 5.156E-06 1.007E-05 5.529E-01 5.529E-01 5.529E-01 5.529E-01Carbon Monoxide 0.000E+00 0.000E+00 0.000E+00 1.580E-06 1.580E-06 3.086E-06 1.337E-01 1.337E-01 1.337E-01 1.337E-01Carbon Dioxide 0.000E+00 0.000E+00 0.000E+00 1.026E-04 1.026E-04 2.003E-04 4.273E-02 4.273E-02 4.273E-02 4.273E-02Methanol 0.000E+00 0.000E+00 0.000E+00 0.000E+00 0.000E+00 0.000E+00 0.000E+00 0.000E+00 0.000E+00 0.000E+00Ethanol 0.000E+00 0.000E+00 0.000E+00 0.000E+00 0.000E+00 0.000E+00 0.000E+00 0.000E+00 0.000E+00 0.000E+00

Stream Number 11 12 13 14 15 16 17 18 19 2Temperature, C 35.000 206.850 53.000 187.445 105.530 35.000 35.122 255.000 255.000 195.02Pressure, kPa 956.498 3359.884 3023.896 8010.000 8010.000 956.498 2282.000 7209.000 6488.100 5839.29Molar Flow, kmol/h 901.077 901.077 901.077 901.077 3253.539 276.539 276.539 4154.626 3765.392 3765.39Mass Flow, kg/hr 8623.949 8623.949 8623.949 8623.949 18420.662 4983.275 4983.275 27044.836 27044.661 27044.66Composition, mol fraction

Methane 4.080E-02 4.080E-02 4.080E-02 4.080E-02 1.347E-01 3.739E-09 3.739E-09 1.143E-01 1.261E-01 1.261E-01Ethane 1.031E-06 1.031E-06 1.031E-06 1.031E-06 3.350E-06 2.370E-15 2.370E-15 2.847E-06 3.142E-06 3.142E-06Water 6.164E-03 6.164E-03 6.164E-03 6.164E-03 2.563E-04 9.998E-01 9.998E-01 1.538E-03 1.293E-02 1.293E-02Hydrogen 7.226E-01 7.226E-01 7.226E-01 7.226E-01 8.141E-01 1.007E-05 1.007E-05 7.943E-01 7.618E-01 7.618E-01Carbon Monoxide 1.747E-01 1.747E-01 1.747E-01 1.747E-01 1.893E-02 3.086E-06 3.086E-06 5.272E-02 1.772E-02 1.772E-02Carbon Dioxide 5.579E-02 5.579E-02 5.579E-02 5.579E-02 2.501E-02 2.003E-04 2.003E-04 3.168E-02 2.373E-02 2.373E-02Methanol 0.000E+00 0.000E+00 0.000E+00 0.000E+00 6.986E-03 0.000E+00 0.000E+00 5.472E-03 5.772E-02 5.772E-02Ethanol 0.000E+00 0.000E+00 0.000E+00 0.000E+00 0.000E+00 0.000E+00 0.000E+00 0.000E+00 0.000E+00 0.000E+00


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Stream Number 21 22 23 24 25 26 27 28 29 3Temperature, C 169.832 40.000 40.000 40.000 70.000 134.053 64.607 45.923 48.978 42.27Pressure, kPa 5255.361 4729.825 4729.825 4729.825 365.000 304.955 101.325 101.325 5416.000 5400.00Molar Flow, kmol/h 3765.392 3765.392 243.102 3522.290 243.102 40.989 198.034 4.079 175.036 93.71Mass Flow, kg/hr 27044.661 27044.661 7102.405 19942.256 7102.405 738.424 6250.544 113.437 352.873 1168.72Composition, mol fraction Methane 1.261E-01 1.261E-01 2.330E-03 1.347E-01 2.330E-03 1.203E-30 4.174E-05 1.369E-01 0.000E+00 3.863E-01Ethane 3.142E-06 3.142E-06 1.179E-07 3.350E-06 1.179E-07 1.656E-73 5.156E-10 7.001E-06 0.000E+00 9.608E-06Water 1.293E-02 1.293E-02 1.966E-01 2.563E-04 1.966E-01 1.000E+00 3.421E-02 5.244E-03 0.000E+00 7.351E-04Hydrogen 7.618E-01 7.618E-01 3.064E-03 8.141E-01 3.064E-03 1.433E-30 2.901E-05 1.812E-01 1.000E+00 4.669E-01Carbon Monoxide 1.772E-02 1.772E-02 1.070E-04 1.893E-02 1.070E-04 3.262E-50 2.260E-06 6.265E-03 0.000E+00 5.429E-02Carbon Dioxide 2.373E-02 2.373E-02 5.079E-03 2.501E-02 5.079E-03 1.633E-30 2.138E-04 2.923E-01 0.000E+00 7.173E-02Methanol 5.772E-02 5.772E-02 7.929E-01 6.986E-03 7.929E-01 2.198E-06 9.655E-01 3.781E-01 0.000E+00 2.003E-02Ethanol 0.000E+00 0.000E+00 0.000E+00 0.000E+00 0.000E+00 0.000E+00 0.000E+00 0.000E+00 0.000E+00 0.000E+00


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A1.2 APPENDIX B

This appendix provides additional details and assumptions of the simulations in Section A2 to supportthe main features discussed in the main report regarding process development and optimization

procedures in Sections 3.1, 3.2, 4.1 and 4.2.

General Simulation and Optimization Assumptions

For the simulations, HYSYS files used were developed according to a generic heat exchange network.The core assumption is that the effect of heat integration presents a similar effect to all simulations. Inaddition, HYSYS files used for optimization in each segment were separately developed to minimizeconvergence errors across segments. The core assumption in this was that minor process parameters,apart from the main factors discussed in Section 3, have marginal effect on the overall process

performance.

In the process development, pressure drops across heat exchangers and reactors was taken to be 10% ofinlet pressure for gases, 60 kPa for liquids of viscosity 1-10 mN.s/m 2 and 35 kPa for


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Operating Costs: The utilities used are cooling water, boiler feed water, medium pressure steam,natural gas and electricity. The costs and streams used are shown in A4.2 Cost Spreadsheet.xlsx .

Capital Costs: The general formula for capital costing is as such:

= 10^ 1+ 2lg( A) + 3lg(A)2 ------------------------------- (1)

1 = 1+ 2 ---------------------------------------------- (2) 2 = --------------------------------------------------------- (3)

= -------------------------- (4)

, =(P )D

944 0.9 . (P )+CAtmin

------------------------------------------------------------- (5)

Where,K 1, K 2, K 3 Cost correlation factorsB1, B2 Bare module correlation factorsFM Material FactorFP Pressure Factor

A Dependent variable

FP,vessel Pressure Factor for VesselsP Operating pressure, bargCA Corrosion Allowance (assumed to be 0.00315), mD Vessel Diameter, m

Pressure factors for all units except vessels are calculated with a formula of similar form in Equation(1) for the case of high pressures and using pressure in barg as the dependent variable; the value isassumed to be 1 for low pressures or vacuum conditions. For vessels, the pressure factor used isdescribed in Equation (5) using a L/D ratio of 3 and assuming an ellipsoidal head of 2:1 ratio;horizontal vessels are used for volumes larger than 3.8m 3. Raw material factors are chosen based on

process requirement. Stainless steel is used for any stream containing methanol, alloy steel is used forthe furnace, while carbon steel is used for the make-up compressor and steam generation streams.Other constants are obtained from [25] and are summarized in the costing spreadsheet of Section A4.The factor used for annualization of capital costs is:

= (1+ )(1+ ) 1 ----------------------------------------------- (5)Where,r Discount rate n Plant life in number of years


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A1.3 APPENDIX C

This appendix provides a clarification for the system boundaries of the options in the biogassubstitution case study. These are shown in the diagrams below. The annotation and superscript notesused in these figures are summarized in Table 1.

Table 1: Summary of Annotation and Notes

Superscript Number Description1 Expressed in terms of potential methane that can be produced from reported kinetics2 Proportion of feed converted to edible substance3 Mass of waste 4 Proportion of waste converted to sludge

Figure 1: Option 1 System Boundary

In the base case of Option 1 without biogas substitution, the methanol plant procures natural gas fromfossil fuel sources while the potential feedstock for biogas remains in its default treatment procedures.Corn is taken to be used for animal feed and ultimately converted to food; wheat straw and the otherorganic waste of wastewater and animal waste are assumed to be treated in nearby facilities. Theemissions for corn used in animal feed is ignored while the rest were calculated from [39] for the caseof energy generation from anaerobic digestion. Methanol production and sludge treatment are ignoredas they remain constant across the different options.


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Accordingly, routing corn from its default option of animal feed usage to maximize NPV from the high productivity of corn grains would substitute part of the natural gas feedstock from fossil fuels. Thisresults in reduced emissions from fossil fuels, but results in increased emission from transportation ofthe feedstocks and the biogas treatment procedures. The magnitude of the change in emissions aresummed and compared across the options. The subsequent figures elaborate on the last three options.


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DESIGN OF METHANOL PLANT

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