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University of Calgary
PRISM: University of Calgary's Digital Repository
Graduate Studies The Vault: Electronic Theses and Dissertations
2014-04-30
Feasibility Study for a Small Scale Integrated
On-Farm Ethanol Plant
Andrade, Virginia
Andrade, V. (2014). Feasibility Study for a Small Scale Integrated On-Farm Ethanol Plant
(Unpublished master's thesis). University of Calgary, Calgary, AB. doi:10.11575/PRISM/26168
http://hdl.handle.net/11023/1458
master thesis
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UNIVERSITY OF CALGARY
Feasibility Study for a Small Scale Integrated On-Farm Ethanol Plant
by
Virginia Andrade
A THESIS
SUBMITTED TO THE FACULTY OF GRADUATE STUDIES
IN PARTIAL FULFILMENT OF THE REQUIREMENTS FOR THE
DEGREE OF MASTER OF SCIENCE
DEPARTMENT OF CHEMICAL AND PETROLEUM ENGINEERING
CALGARY, ALBERTA
APRIL, 2014
© Virginia Andrade 2014
1
Abstract
The availability of wheat, along with the large amounts of manure produced by
the livestock industry in the province of Alberta, presents an interesting
opportunity for integration of farm operation with biofuels production in western
Canada. Anaerobic co-digestion of manure and the wastewater from ethanol
production generates biogas which could be used to produce steam and
electricity, as well as digestate for fertilizer. Animal feed is obtained as co-
product from ethanol production. In the present study, a conceptual design of a
small scale integrated on-farm ethanol plant for the production of 112 L/h of
anhydrous ethanol was developed. The process was designed and optimized
for low energy consumption using a commercial process simulator. Optimization
through heat integration lowered the steam consumption to 3.71 kg per liter of
ethanol. The total capital investment for the project was estimated at 4,100,000
CD with a payback period of 5 years and a return on investment of 25%.
2
Acknowledgements
I would like to acknowledge my supervisor, Dr. Michael Foley, for his continuous
support and guidance in the development of this project.
I also acknowledge Whitefox Technologies for funding my Master’s research and
thank all Whitefox team members for their support and fellowship throughout this
period.
I would like to thank the examiners of this thesis: Dr. Edwin Nowicki, Dr. Matthew
Clarke, and Dr. Alex De Visscher, for their time in reviewing this document.
My sincere gratitude to my dear family for lifting my arms when I felt like giving
up, for their constant prayers, love and concern, and for being an example of
courage, integrity, commitment and effort. Thank you for celebrating my
successes and sharing my tears. Despite the physical distance between us, your
love and care has covered me and given me strength. I would not be here if it
were not for each of you.
Thanks to Daniel, for his support, for making me smile in the hardest moments,
and for always having the right words to say.
Finally I wish to thanks all my dear friends who have become my family here in
Canada.
3
Dedication
A mi fiel Dios, el eterno, mi fortaleza y refugio
seguro, el honor y la gloria son para ti
A mi más grande bendición, mi amada familia
4
Table of Contents
1. Introduction ............................................................................................................... 18
1.1 General Biofuels Overview ............................................................................... 18
1.2 Wheat as Feedstock for Biofuels ..................................................................... 22
1.3 Cattle Operation in Alberta ............................................................................. 24
1.4 Bioethanol Perspective, Market and Policies in Canada ........................... 24
1.5 Process Overview: Conventional vs. Proposed ............................................ 25
1.5.1 Conventional Process ................................................................................ 25
1.5.2 Proposed Process: IDF ................................................................................ 27
1.6 Importance of Integrated Systems ................................................................. 29
1.6.1 Energy Savings and GHG Emissions Reduction ..................................... 31
1.6.2 Production based on local farming yields .............................................. 36
1.7 Project Objectives ............................................................................................. 37
2. Literature Review ...................................................................................................... 40
2.1 Process and Equipment Selection .................................................................. 40
2.1.1 Feedstock preparation – Milling ............................................................... 42
2.1.2 Cooking and Fermentation ...................................................................... 45
2.1.2.1 Cooking ................................................................................................. 45
2.1.2.2 Fermentation ........................................................................................ 48
2.1.3 Distillation ..................................................................................................... 52
2.1.4 Dehydration ................................................................................................. 58
2.1.5 Anaerobic Digestion .................................................................................. 64
2.1.6 Co-Generation System .............................................................................. 74
2.1.7 Plant Co-products ...................................................................................... 82
2.1.7.1 Stillage ................................................................................................... 82
2.1.7.2 Bio-fertilizer ............................................................................................ 84
2.1.7.3 Biogas .................................................................................................... 84
3. Process Selection ...................................................................................................... 86
5
3.1 Feedstock Preparation – Milling ...................................................................... 86
3.2 Cooking and Fermentation ............................................................................. 87
3.2.1 Simultaneous Saccharification and Fermentation Reaction Kinetics ... 88
3.3 Distillation ............................................................................................................ 90
3.4 Dehydration ....................................................................................................... 92
3.5 Anaerobic Digestion ......................................................................................... 93
3.6 Co-generation System ...................................................................................... 94
4. Methodology ............................................................................................................ 96
4.1 Process Design ................................................................................................... 96
4.1.1 Software and Property Package Selection ............................................ 96
4.1.2 Simulation of Unit Operations ................................................................... 99
4.1.2.1 SSF Reactor ........................................................................................... 99
4.1.2.2 CO2 Scrubber ..................................................................................... 100
4.1.2.3 Distillation............................................................................................. 100
4.1.2.4 Dehydration ........................................................................................ 101
4.1.2.5 Anaerobic Digester ........................................................................... 101
4.1.2.6 Co-generation .................................................................................... 101
4.1.2.7 Heat exchangers ............................................................................... 102
4.1.2.8 Pumps .................................................................................................. 102
4.1.2.9 Control valves ..................................................................................... 103
4.1.2.10 Utilities .................................................................................................. 103
4.1.3 Pinch Analysis ............................................................................................ 103
4.2 Process Description ......................................................................................... 107
4.2.1 Base Design Case ..................................................................................... 107
4.2.1.1 Base-Case Design Process Flow Diagram ...................................... 107
4.2.1.1.1 Feedstock Preparation................................................................... 114
4.2.1.1.2 Fermentation ................................................................................... 114
4.2.1.1.3 Distillation and Dehydration .......................................................... 115
4.2.1.1.4 Anaerobic Digestion ...................................................................... 116
6
4.2.1.1.5 Co-Generation ................................................................................ 116
4.2.2 Improved Process ..................................................................................... 117
4.2.2.1 Improved Process Flow Diagram ..................................................... 120
4.3 Equipment Sizing .............................................................................................. 128
4.3.1 Milling .......................................................................................................... 128
4.3.2 Saccharification and Fermentation Set ................................................ 128
4.3.3 CO2 Absorber ............................................................................................ 129
4.3.4 Distillation Section ..................................................................................... 131
4.3.5 Dehydration ............................................................................................... 132
4.3.6 Anaerobic Digestion ................................................................................ 133
4.3.7 Co-Generation System ............................................................................ 134
4.3.8 Vessels and Tanks ..................................................................................... 134
4.3.9 Heat Exchangers....................................................................................... 136
4.3.10 Pumps ......................................................................................................... 137
4.3.11 Miscellaneous ............................................................................................ 139
4.4 Economic Evaluation ...................................................................................... 140
4.4.1 Equipment Costing ................................................................................... 140
4.5 Capital Investment .......................................................................................... 143
4.5.1 Economic Model ...................................................................................... 143
4.5.2 Economic Indicators ................................................................................ 151
4.5.2.1 EBITDA .................................................................................................. 151
4.5.2.2 ROI ........................................................................................................ 152
4.5.2.3 ROS ....................................................................................................... 152
4.5.2.4 Payback Period .................................................................................. 152
4.5.3 Sensitivity Analysis ..................................................................................... 153
5. Process Design ........................................................................................................ 154
5.1. Base Design Case Simulation Work .............................................................. 154
5.1.1. SSF Reactor ................................................................................................ 154
5.1.2. CO2 Scrubber ............................................................................................ 156
7
5.1.3. Distillation ................................................................................................... 156
5.1.4. Dehydration ............................................................................................... 159
5.1.5. Anaerobic Digestion ................................................................................ 159
5.1.6. Co-generation System ............................................................................. 160
5.2. Pinch Analysis ................................................................................................... 161
5.3. Improved Design ............................................................................................. 166
5.3.1 Improved Design Simulation Work ......................................................... 167
5.3.2 Heat integration ........................................................................................ 169
5.3.3 Mass Balance ............................................................................................ 171
5.3.4 Control Strategy ........................................................................................ 173
5.3.4.1 Wheat Hydration ................................................................................ 174
5.3.4.2 Meal Saccharification and Fermentation ..................................... 175
5.3.4.3 CO2 Absorption .................................................................................. 176
5.3.4.4 Distillation............................................................................................. 177
5.3.4.5 Dehydration ........................................................................................ 178
5.3.4.6 Anaerobic Digestion ......................................................................... 178
5.3.4.7 Co-generation System ...................................................................... 179
5.4. Occupational Health and Safety Analysis .................................................. 179
6. Economic Evaluation ............................................................................................. 182
6.1. Equipment Sizing .............................................................................................. 182
6.1.1. Milling .......................................................................................................... 182
6.1.2. Simultaneous Saccharification and Fermentation .............................. 182
6.1.3. CO2 Absorber ........................................................................................... 182
6.1.4. Distillation Section ..................................................................................... 183
6.1.5. Dehydration Section ................................................................................ 183
6.1.6. Anaerobic Digestion ................................................................................ 184
6.1.7. Co-Generation System ............................................................................ 185
6.1.8. Vessels and Tanks ..................................................................................... 185
6.1.9. Heat Exchangers....................................................................................... 186
8
6.1.10. Pumps ...................................................................................................... 188
6.1.11. Miscellaneous Equipment .................................................................... 189
6.2. CAPEX Evaluation ............................................................................................ 190
6.2.1 Equipment Costing ................................................................................... 190
6.2.2 Capital Investment ................................................................................... 193
6.3. Farmer’s Financial Evaluation ........................................................................ 194
6.3.1 Profitability Analysis ................................................................................... 199
6.3.2 Sensitivity Analysis ..................................................................................... 199
7. Conclusions ............................................................................................................. 203
Appendices ................................................................................................................... 229
Appendix 1: McCabe Thiele diagrams for distillation column design for the base design case ...................................................................................................... 229
Appendix 2: McCabe Thiele diagrams for distillation column B for the optimized design ....................................................................................................... 232
Appendix 3: Operation schedule of the IDF ethanol plant ................................ 235
Appendix 4: Distillation columns and absorber sizing .......................................... 236
Appendix 5: Stripper distillate flash drum sizing .................................................... 238
Appendix 6: Knock-out drum sizing ........................................................................ 240
Appendix 7: SSF set sizing ......................................................................................... 241
Appendix 8: Vessels and tanks sizing ...................................................................... 243
Appendix 9: Heat exchangers sizing ...................................................................... 244
Appendix 10: Pumps sizing ....................................................................................... 245
Appendix 9: Vacuum pump P-244 performance curve [158] ........................... 247
Appendix 10: Hydrated Meal Pump (P-209) Performance Curve [162] ........... 248
Appendix 11: SSF Reactor Pump (P-213) Performance Curve [162] ................. 249
Appendix 12: CO2 Absorber Pump (P-219) Performance Curve [162] ............. 250
Appendix 13: Yeas Slurry Pump (P-222) Performance Curve [162] .................... 251
Appendix 14: Enzyme Pump (P-224) Performance Curve [162] ......................... 252
Appendix 15: Permeate Pump (P-245) Performance Curve [162] .................... 252
Appendix 16: Thin Stillage Pump (P-249) Performance Curve [162] .................. 253
9
Appendix 17: Equipment costing calculations ..................................................... 255
Appendix 17: Cumulative cash flows for 10 years of project life time .............. 257
Appendix 18: Sensitivity analysis results .................................................................. 258
Appendix 19: MSDS of anhydrous ethanol [165] .................................................. 259
10
List of Tables
Table 1: Fuel properties comparison for gasoline and ethanol ............................... 21 Table 2: Prices for gasoline/ethanol blends in the US [15] ........................................ 22 Table 3: Energy balance for ethanol production using wheat as feedstock [31] 35 Table 4: Greenhouse gas emissions for ethanol production using wheat as feedstock through conventional process and IDF configuration [31] ................... 36 Table 5: Exemplary table for evaluating the different process alternatives using the screening method [37] ............................................................................................ 40 Table 6: Characteristics of the milling processes considered for the IDF Ethanol Plant .................................................................................................................................. 45 Table 7: Characteristics of the cooking and fermentation processes considered for the IDF Ethanol Plant ................................................................................................. 51 Table 8: Characteristics of the packing types considered for the IDF Ethanol Plant ........................................................................................................................................... 57 Table 9: Characteristics of the dehydration technologies considered for the IDF Ethanol Plant .................................................................................................................... 63 Table 10: Characteristics of the anaerobic digestion processes considered for the IDF Ethanol Plant ...................................................................................................... 67 Table 11: Characteristics of the co-generation systems considered for the IDF Ethanol Plant [90] ............................................................................................................ 81 Table 12: Composition of biogas and natural gas .................................................... 85 Table 13: Criteria grading for feedstock preparation - milling ................................. 86 Table 14: Criteria grading for cooking and fermentation ........................................ 87 Table 15: SSF model parameters [100] ......................................................................... 90 Table 16: Criteria grading for distillation ...................................................................... 91 Table 17: Criteria grading for dehydration ................................................................. 92 Table 18: Criteria grading for anaerobic digestion ................................................... 93 Table 19: Criteria grading for co-generation system. ............................................... 95 Table 20: Design specifications for the vapor permeation membrane module.132 Table 21: Comparison of the characteristics of the anaerobic digester of the IDF plant and an AD reported in the literature. .............................................................. 134 Table 22: Sizing parameters for tanks and surge vessels. ........................................ 136 Table 23: Equipment specifications for different pieces of equipment of the IDF plant. ............................................................................................................................... 140 Table 24: Scaling exponents employed in equipment costing. ............................ 141 Table 25: Chemical Engineering Plant Cost Index ................................................... 142 Table 26: Percentages for determining the capital investment ............................ 143
11
Table 27: Factors used in the economic model ....................................................... 146 Table 28: Income statement format. ......................................................................... 149 Table 29: Initial conditions used for solving differential equations describing the SSF reactor...................................................................................................................... 154 Table 30: Final results obtained from the SSF model. ............................................... 155 Table 31: Conversions used in the VMGSim™ model of the SSF reactor. ............ 156 Table 32: Design parameters and results obtained using the McCabe-Thiele method base-case design of the distillation column.............................................. 157 Table 33: Values of the variables input to the WFX membrane analysis tool. ..... 159 Table 34: WFX membrane analysis tool outputs for calculation of the splits for membrane dehydration module simulation ............................................................ 159 Table 35: Factors used for the anaerobic digester modelling. .............................. 160 Table 36: Heat stream inputs for pinch analysis in the base-case design. .......... 161 Table 37: Results obtained for alternative distillation section configurations. ..... 163 Table 38: Optimized design of distillation column in Configuration 2. ................. 168 Table 39: Hot and cold streams in the process for heat integration analysis. ..... 169 Table 40: Steam and cooling water consumption in the different process designs. ......................................................................................................................................... 171 Table 41: Feed and product characteristics for the continuous zone of the IDF plant. ............................................................................................................................... 172 Table 42: Fermentation section sizing results. ........................................................... 182 Table 43: Absorber sizing results. ................................................................................. 183 Table 44: Distillation section sizing results. .................................................................. 183 Table 45: Sizing parameters for the vacuum pump. ............................................... 184 Table 46: Results for the co-generation system sizing. ............................................. 185 Table 47: Equipment models and suppliers selected for the Co-Generation System ............................................................................................................................. 185 Table 48: Results of the vessel and tank sizing calculations. .................................. 186 Table 49: Heat exchanger sizing results. .................................................................... 187 Table 50: Pump sizing results ........................................................................................ 188 Table 51: Pump model selection ................................................................................ 188 Table 52: Pumps scaling values and exponents. ..................................................... 189 Table 53: Equipment models and suppliers selected. ............................................. 189 Table 54: Miscellaneous equipment scaling values and exponents .................... 190 Table 55: Equipment costing ....................................................................................... 191 Table 56: Total capital investment and percentage breakdown ......................... 194 Table 57: Income statement to compare farmer’s financial position before and after installation of the IDF ethanol plant .................................................................. 195
12
Table 58: Independent parameter variations for sensitivity analysis. ................... 200 Table 61: Parameters used for the distillation section sizing ................................... 236
13
List of Figures
Figure 1: Fuel ethanol production for the USA in million gallons per year [8] ........ 19 Figure 2: BFD for conventional ethanol production process. ................................... 26 Figure 3: BFD for proposed ethanol production process .......................................... 28 Figure 4: BFD for the conventional ethanol production process ............................. 34 Figure 5: Dry milling ethanol production [44] .............................................................. 43 Figure 6: Wet milling ethanol production [44] ............................................................ 44 Figure 7: Amylose structure ............................................................................................ 46 Figure 8: Amylopectin structure .................................................................................... 46 Figure 9: Vapour-liquid equilibrium curve for ethanol water mixtures at atmospheric pressure [38, 68] ....................................................................................... 54 Figure 10: Disc and donuts trays [38] ........................................................................... 57 Figure 11: Schematic of vapour permeation. P1 is greater than P2. ...................... 59 Figure 12: Schematic of membrane vapour permeation module for ethanol dehydration ..................................................................................................................... 60 Figure 13: Anaerobic Digestion Process [79] .............................................................. 65 Figure 14: Complete-mix anaerobic digestion system schematic [85] .................. 70 Figure 15: Installed complete-mix anaerobic digestion system [85] ....................... 71 Figure 16: Plug-flow anaerobic digester schematic [85] .......................................... 71 Figure 17: Installed plug-flow anaerobic digester [85] .............................................. 72 Figure 18: Upward-flow anaerobic sludge blanket systems schematic [86] ......... 73 Figure 19: Installed upward-flow anaerobic sludge blanket systems [87] ............. 73 Figure 20: Covered lagoon for anaerobic digestion schematic [85] ..................... 74 Figure 21: Installed covered lagoon for anaerobic digestion [85] .......................... 74 Figure 22: Steam turbine CHP system [90] ................................................................... 77 Figure 23: Gas turbine CHP system [90] ....................................................................... 78 Figure 24: Reciprocating internal combustion engine CHP system [90] ................ 79 Figure 25: Simulation software physical property package selection tree [104]. . 97 Figure 26: Selection tree for polar non-electrolyte mixtures [104]. .......................... 97 Figure 27: Vapour-Liquid equilibrium data for ethanol-water mixture. .................. 98 Figure 28: Liquid density for ethanol-water mixture. .................................................. 99 Figure 29: Sample Composite Curves Diagram [118]. ............................................ 104 Figure 30: Base design case distillation configuration - one column with steam preheat exchanger. ..................................................................................................... 118 Figure 31: First configuration for process improvement analysis - two towers (stripper + distillation column). .................................................................................... 119
14
Figure 32: Second configuration for process improvement analysis - two towers at different pressures. ................................................................................................... 120 Figure 33: SSF conical reactor [121]. .......................................................................... 129 Figure 34: Conical vessel schematic. ......................................................................... 135 Figure 35: Schematic of static head for pump calculations [128]. ....................... 138 Figure 36: Concentration profile for ethanol production in the SSF reactor. ...... 155 Figure 37: Optimization of number of stages for the distillation tower based on the reflux ratio rule-of-thumb. ..................................................................................... 158 Figure 38: Optimization of number of stages for the distillation tower based on minimizing reboiler duty. .............................................................................................. 158 Figure 39: Hot and cold composite curves for base-case design. ....................... 162 Figure 40: Hot and cold composite curves diagram for Configuration 1 (stripper + distillation tower – same pressure). ............................................................................. 164 Figure 41: Hot and cold composite curves for Configuration 2 (stripper – 6 bara, distillation column – 2.6 bara). .................................................................................... 165 Figure 42: Hot and cold composite curves for Configuration 2 (stripper – 4.8 bara, distillation column – 2.6 bara). .................................................................................... 166 Figure 43: Optimizing the number of stages for the stripper in Configuration 2 based on steam injection requirements. .................................................................. 168 Figure 44: Bucket elevator meal conveyor considered for MC-206 [160]. .......... 189 Figure 45: Screw press considered for SP-246 [161]. ................................................ 190 Figure 46: EBITDA analysis before and after the installation of the IDF plant. ..... 196 Figure 47: Gross profit analysis before and after installation of the IDF plant...... 197 Figure 48: Results of the economic analysis. ............................................................. 198 Figure 49: Cumulative net cash inflows/outflows. .................................................... 199 Figure 50: Sensitivity analysis for the difference in the gross profit over sales before and after installation of the IDF ethanol plant. ........................................... 200 Figure 51: Sensitivity analysis for the difference in the EBITDA over sales before and after installation of the IDF ethanol plant. ........................................................ 201 Figure 52: Sensitivity analysis for the difference in the net profit over sales before and after installation of the IDF ethanol plant. ........................................................ 201
15
List of Symbols, Abbreviations, Nomenclatures
A Area AAFC Agriculture and Agri-food Canada AARD Alberta Agriculture and Rural Development AD Anaerobic digestion AFSC Agriculture Financial Services Corporation AUC Alberta Utilities Commission BFD Block Flow Diagram bu Bushel C Purchased cost CAPEX Capital expenses CD Canadian Dollars CDS Condensed distillers' solubles CEPCI Chemical engineering plant cost index CHP Combined heat and power COD Chemical oxygen demand ∆TLM Log mean temperature difference D Diameter DDGS Dried distillers’ grains with solubles DG Distillers’ grain DGS Distillers’ grain with solubles DOE United States Department of Energy E Ethanol concentration E0 Tray efficiency E10 Ethanol - gasoline blend: 10% ethanol – 90% gasoline E20 Ethanol - gasoline blend: 20% ethanol – 80% gasoline E5 Ethanol - gasoline blend: 5% ethanol – 95% gasoline E85 Ethanol - gasoline blend: 85% ethanol – 15% gasoline EBITDA Earnings before interest, taxes, depreciation, amortization ED Extractive distillation Eq Equation ERoEI Energy return over energy invested F Feed to the column G Glucose concentration Gal Gallon GHG Greenhouse gas
16
GWP Global warming potential HRT Hydraulic retention time I Cost index IC Internal combustion IDF Integrated decentralised on-farm J Joules Ks Substrate (glucose) constant Lc Column height LCA Life cycle assessment LF Liquid Feed Flow mp Product maintenance rate constant ms Substrate maintenance rate constant NT Number of trays OPEX Operating expenses P Pressure PFD Process Flow Diagram q heat duty R Calculated reflux ratio Rmin Minimum reflux ratio ROI return on investment ROS Return on sales Rxn Reaction SG Specific gravity SHF Separated hydrolysis and fermentation SSF Simultaneous saccharification and fermentation U Heat transfer coefficient UASB Upward-flow anaerobic sludge blanket system USD American Dollars USDA United States Department of Agriculture V Volume V̇ volumetric flow vmax Maximum allowable vapour velocity VP Vapour permeation WDG Wet distillers' grain WDGS Wet distillers' grain with solubles WFX Whitefox Technologies Canada Ltd.
17
wt Weight X Biomass concentration xB Tower bottoms composition, ethanol concentration xD Distillate composition, ethanol concentration YPX Yield coefficient on ethanol YXS Yield coefficient on substrate zF Feed concentration ρ Density τ Residence time TCI Total capital investment μmax Maximum specific growth rate of biomass
18
1. Introduction
1.1 General Biofuels Overview
As the environmental concern related to greenhouse gas emissions caused by
burning fossil fuels increases, biofuels like ethanol and biodiesel appear to be an
viable alternative to gasoline and diesel fuels. Additional factors such as rising oil
prices, the depletion of global oil reserves, and the provision of alternative
outlets for agricultural producers reinforce the efforts taken in order to further
investigate these alternatives [1, 2].
Ethanol is the most economically significant biofuel. It offers a sustainable source
of energy since it is largely derived from renewable sources. It address issues of
global warming, energy independence (if the farming and fuel production are
domestic), and support for a farm economy [3]. The environmental benefit
provided by ethanol is the reduction of production and transportation-related
greenhouse gas (GHG) emissions. Recent life-cycle assessment studies have
reported direct-effect greenhouse gas emissions equivalent to a reduction of
48% to 59% for corn-ethanol systems in the US, as compared to gasoline [4]. Life
cycle assessment analysis performed for Canadian biofuel plants in the period
2008 - 2009 showed a reduction of 62% per MJ for ethanol as compared with
gasoline [5]. Moreover, the energy balance is favorable for ethanol as
compared to fossil fuels. Liska [4] found an ethanol-petroleum output/input ratio
ranging from 10:1 to 13:1 for corn-ethanol. Whiten and Reyes [6] reported that
19
ethanol reduces tailpipe carbon monoxide emissions by as much as 30%.
Reductions in toxic content (mainly benzene) of 13 wt% and 50% in tailpipe fine
particulate matter emissions were also observed.
Ethanol production worldwide has been rapidly expanding since the oil crises in
the 1970s. Its market grew from less than a billion liters in 1975 to more than 39
billion liters in 2006 [7]. Ethanol production in the US reached 50 billion liters in
2012 [8]. Ethanol production is projected to reach 100 billion liters per year by
2015 [7]. In 1999, there were 50 plants operating in the US with an ethanol
production capacity of 6,441 million liters per year. By January 2013, the number
of ethanol plants was 211 with a capacity of 55,686 million liters per year [8]. The
chart below shows the historic US ethanol production from 1980 to 2012.
Figure 1: Fuel ethanol production for the USA in million gallons per year [8]
0
10,000
20,000
30,000
40,000
50,000
60,000
1980
1981
1982
1983
1984
1985
1986
1987
1988
1989
1990
1991
1992
1993
1994
1995
1996
1997
1998
1999
2000
2001
2002
2003
2004
2005
2006
2007
2008
2009
2010
2011
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2013
Fuel
Eth
anol
Pro
duct
ion
(Mill
ion
of li
ters
)
Year
20
The market penetration of ethanol as a source of transportation fuel will
reportedly attain 53% of US gasoline consumption in 2030 [7]. One of the current
uses of ethanol is in blended gasoline (gasoline mixed with certain amounts of
ethanol). 3.4 billion gallons of ethanol were blended into gasoline in 2004 [9].
When ethanol is blended with gasoline, it improves some of its properties, such
as octane, oxygen content and volatility [10]. Researchers have found that the
high octane rating of ethanol reduces engine knock, thus improving engine
performance [11]. The table below shows the fuel properties of gasoline and
ethanol [12]:
21
Table 1: Fuel properties comparison for gasoline and ethanol
Property Fuels Gasoline Ethanol
Chemical Structure C4 to C12 CH3CH2OH
Feedstocks Crude Oil Corn, grains, or agricultural waste (cellulose)
Gasoline Gallon Equivalent 100%
1 gallon of E85 has 73% to 83% of the energy of one gallon gasoline(1) .1 gallon of E10 has 96.7% of the energy of one gallon of gasoline
Energy Content (lower heating value) 116,090 Btu/gal 76,330 Btu/gal for E100
Energy Content (higher heating value) 124,340 Btu/gal 84,530 Btu/gal for E100
Physical State Liquid Liquid Cetane Number N/A 0-54 Pump Octane Number 84-93 110 Flash Point -43 ºC 13 ºC Freezing point -40 ºC -114 ºC Autoignition Temperature 257 ºC 423 ºC Reid Vapor Pressure [13] 0.76 bar at 38 °C 0.14 bar at 38 °C
Energy Security Impacts Manufactured using oil, of which nearly 2/3 is imported.
Ethanol is produced domestically. E85 reduces lifecycle petroleum use by 70% and E10 reduces petroleum use by 6.3%.
1. Ethanol content in E85 varies depending on geography and season. During winter months in cold climates, ethanol content is lower to ensure vehicle starts. The content of ethanol in E85 varies from 51 to 83% corresponding to lower heating values from 83,950 to 95,450 Btu/gal. This is equivalent to 73% to 83% the content of gasoline
As mentioned in a study by the Oak Ridge National Laboratory [14]
gasoline/ethanol blends can be used in all gasoline-powered automobiles and
light trucks on the road today and are compatible with the existing service
station infrastructure. Moreover, gasoline/ethanol blends show a competitive
price in the fuel market. Biofuels Digest [15, 16] published a comparison for
22
different prices (adjusted for fuel economy) of gasoline/ethanol blends with fuel
economy (distance traveled vs. amount of fuel consumed in miles per gallon.
The publication showed E20 (ethanol 20%, gasoline 80%) as the blend with the
lowest price (adjusted for fuel economy), with a value of 3.72 USD as compared
with 4.07 USD for pure gasoline [17]. The table below shows the results obtained:
Table 2: Prices for gasoline/ethanol blends in the US [15]
Ethanol Blend [%]
Price Adjusted for Fuel Economy [USD/gallon]
0 4.07 10 4.08 20 3.72 30 3.89 40 4.21 50 3.80 85 3.97
The implications of ethanol in the North American economy are notable as well.
The Bioenergy Technologies Office (BETO), a unit from the U.S. Energy
Department, has reported that since 2004, the US biofuels industry has grown an
average of 8.9% annually and accounted for 87,000 direct jobs in 2011. Ethanol
in the US is estimated to have displaced about $47.2 billion worth of imported
crude in 2012, thus improving the trade deficit of the country [18].
1.2 Wheat as Feedstock for Biofuels
For decades, ethanol has been produced mainly from two feedstocks:
sugarcane in Brazil and starch-rich grains, principally corn in the US, and cereals
23
such as wheat in Canada [19]. Sugarcane’s main component is glucose, a
simple sugar that can be readily converted to ethanol [20]. The main
component of corn kernels and cereal grains is starch, a polysaccharide
constituted by alpha-linked glucose units that can be broken down to glucose
monomers and fermented to ethanol [21]. Lignocellulose feedstocks contain
cellulose (their main component), hemicellulose and lignin [22]. Cellulose is a
beta-linked glucose polymer that requires a breakdown of its sugar chains prior
to fermentation. The breakdown of the polysaccharides contained in
lignocellulose feedstocks is more difficult than that for starch feedstock. A pre-
treatment is required for this type of feedstocks [20, 21].
Canada produces 22 to 24 million tonnes of wheat per year. The majority of
Canadian wheat is grown in the Prairie Provinces. 46% of total production is in
Saskatchewan, 30% in Alberta and 14% in Manitoba. The remaining production is
in Ontario and Quebec [23]. There are 7 classes of wheat in Western Canada:
Canada Western Red Spring, Canada Western Amber Durum, Canada Western
Soft White Spring, Canada Western Extra Strong Red Spring, Canada Prairie
Spring Red, Canada Prairie Spring White, and Canada Western Red Winter. In
order to maintain wheat quality, the Canadian Grain Commission sets specific
standards and regulates the grading of wheat in Canada. When any of the
seven varieties mentioned above does not meet the standards specified for the
lowest scheduled grade within each class, the grain is graded as Canada
Western Feed (CW Feed), which is not used for human consumption. The ready
24
availability of wheat as feedstock for biofuels production, especially, CW Feed,
is a motivation for the production of ethanol in western Canada [24].
1.3 Cattle Operation in Alberta
There are 34,625 cattle ranching and farming operators and 5.7 million cattle in
the province of Alberta. The large number of cattle produces a large quantity of
manure that needs to be disposed and often limits the expansion of livestock
industry [25]. The main input for cattle operation is animal feed. In 2011, cattle
feed accounted for $5.6 billion, corresponding to 15% of total farm operating
expenses in Canada [26].
The feed requirements and the generation of manure during cattle operation
offer great opportunities for the integration of livestock operation with
anaerobic digestion to treat manure, and with ethanol production, which
produces distillers’ grain that can replace some of the feed requirements.
1.4 Bioethanol Perspective, Market and Policies in Canada
On September 1, 2010, the government of Canada announced the finalization
of Federal Renewable Fuel Regulations requiring an average of 5% renewable
content in gasoline across Canada. This Renewable Fuel Mandate came into
effect on December 15th, 2010 [27].
According to the Canadian Renewable Fuels Association, the renewable fuels
sector in Canada has provided economic growth through the creation of 14,177
25
new direct and indirect jobs. The organization reported the generation of $2.949
billion in economic activity from plant construction, as well as 1.473 billion in
economic activity generated annually from these facilities, mostly in rural
Canada [28].
The development of a substantial ethanol industry means new markets for
Canadian biomass, agriculture and forestry. The diversification of rural
economies is also an important contribution of bioethanol production. New
sources of revenue for Canada's forest industry would also be possible with
further developments in cellulosic ethanol production from forest feedstock,
such as wood waste [27].
1.5 Process Overview: Conventional vs. Proposed
1.5.1 Conventional Process
Conventional ethanol production consists of 2 main sections as depicted in the
block flow diagram below: farm operation and ethanol plant.
26
Figure 2: BFD for conventional ethanol production process.
Initially, the feedstock is obtained at the farm, where two operations are
distinguished: crop operation and cattle operation. The principal inputs for crop
operation are seeds, fertilizer and pesticide, while for cattle operation, the main
requirement is animal feed. Wheat is harvested and transported to a centralised
ethanol production facility in which the grain is milled, fermented and distilled in
order to obtain hydrous ethanol. At the base of the distillation tower, the solids,
known as stillage or whole stillage, are recovered and separated into wet grain
(also called wet cake or thick stillage) and thin stillage. The thin stillage is then
evaporated to obtain syrup that is mixed with the wet grain in order to obtain
DGS (distillers’ grain with solubles). DGS is dried to get DDGS (dried distillers’
grains with solubles). The ethanol that is recovered at the top of the distillation
column passes through a dehydration process in which water is removed from
27
the hydrous ethanol to obtain “anhydrous” – fuel grade ethanol (minimum
concentration: 92.1 volume % [29]). The anhydrous product is sent to a blending
facility in which ethanol is mixed with gasoline in order to obtain different
ethanol-gasoline blends (E5: 5% ethanol – 95% gasoline, E10: 10% ethanol – 90%
gasoline, E85: 85% ethanol – 15% gasoline). The fuel is finally transported to gas
stations for retail to the consumer.
1.5.2 Proposed Process: IDF
The proposed process involves an integrated decentralised on-farm (IDF)
ethanol production, in which a synergy between farm operation and ethanol
production is intended. The block flow diagram for the proposed system is
presented in Figure 3:
28
Figure 3: BFD for proposed ethanol production process
In the IDF configuration, the farm operation is merged with ethanol production.
The feedstock (wheat in the case of the present study) obtained in the farm
operation passes through milling, fermentation and distillation processes in order
to obtain hydrous ethanol and whole stillage in the base of the distillation tower.
Whole stillage is separated into wet grain that is recycled as animal feed for
cattle operation, and thin stillage that is sent to an anaerobic digestion (AD)
process. The thin stillage is co-digested in the AD reactor with the manure
obtained from cattle operation. The digestate from the AD is recovered as bio-
fertilizer that is sent back to the crop operation. The biogas obtained as the
other product of the AD is sent to a co-generation system for electricity and
29
steam generation. The ethanol recovered at the top of the distillation column
follows a dehydration process similar to the one described for the conventional
process.
1.6 Importance of Integrated Systems
Even though the concept of decentralised ethanol production in an integrated
on-farm concept is not a new idea, there are some political, socioeconomic
and technical factors that are making of this approach a feasible and
interesting option nowadays.
i) Environmental concerns: Greenhouse gas emissions can be reduced
by using this type of configuration. These reductions are possible by:
- Lower energy consumption though the recycle of co-products to farm
operation (fertilizer + animal feed), thus eliminating energy intensive steps
in the conventional production process as drying of DDGS.
- Reducing transportation steps. In the integrated configuration, which
combines the farming operation with ethanol production, transportation
of wheat to a centralised ethanol facility can be avoided, as well as
transportation of fertilizer to the farm.
- Co-generation of utilities (steam and electricity) to avoid emissions
associated with their generation and transportation to the ethanol plant.
ii) Bio-economy development: Governments are promoting the
formation of bioclusters (biorefinery concept) to encourage the rural
30
and agricultural sector development. Efforts are being made to
enhance agriculture based communities
iii) Feedstock enhancement: Feedstock has been developing through
genomics (improving fiber and crop characteristics via molecular
technologies), breeding (development of new cultivars for better
performance) and agronomy (optimizing crop operation practices –
seeding rates, fertilizer and water usage, harvesting practices) [30]. The
feedstock developments allow for better grain/land yields and
therefore better ethanol yields.
iv) Technology improvements: Progress has been made both in the
upstream and downstream process of ethanol production processes:
- Improved enzymes and yeast for mash preparation: Development of
hydrolyzing enzymes for low temperature hydrolysis which obviates the
need for cooking at elevated temperatures
- Simultaneous saccharification and fermentation process for combined
saccharification and fermentation which leads to easier operation and
lower costs.
- Modular dehydration technology which offers the modularity that enables
easy scale-up and scale-down of the dehydration process and
management of small size mobile units.
v) Environmental assessment tools: Life cycle assessment (LCA)
development has been applied recently to accurately assess biofuel
31
production in a comprehensive way. It has revealed the limitations that
can be found in large scale centralised ethanol production.
A decentralised and integrated approach for the production of ethanol in
an on-farm concept (IDF process) offers the following benefits compared to
its conventional production in a large scale centralised facility:
a) Energy savings and GHG emissions reduction by
i) elimination of transportation stages
ii) co-product generation and recycle (fertilizer + wet animal feed)
iii) co-generation of utilities (electricity + steam)
b) Production based on local farming yields (better yields than average)
1.6.1 Energy Savings and GHG Emissions Reduction
A decentralised on-farm ethanol plant configuration implies reducing
transportation, obtaining valuable co-products as animal feed and fertilizer,
elimination of co-product drying (DDGS) and a virtually self-sufficient energy
supply through co-generation of utilities. This in turn leads to:
i) Reduction of GHG emissions
ii) Better process energy balance (higher energy return over energy
invested - ERoEI or lower energy ratio measured as Joules consumed
per Joule delivered)
iii) Lower OPEX
32
To substantiate these statements, a life cycle approach can be used. For the
specific case of Alberta, S&T Consultants Inc. has evaluated ethanol production
from wheat in regards to GHG emissions and the energy life cycle implications
associated with the production of ethanol [31]. The tool used for the study was
GHGenius, a free publicly available Canadian life cycle model, in which GHG
emissions as well as energy balances are considered [32].
The following life cycle stages were incorporated in the model [31]:
- Fuel dispensing at the retail level: includes the emissions and energy use
associated with the transfer of ethanol from storage into vehicles at the
service station. It includes electricity for pumping, fugitive emissions and
spills.
- Fuel storage and distribution at all stages: related to the emissions and
energy use associated with storage and handling of ethanol. Includes all
emissions from the ethanol plant gate to the delivery of the ethanol at the
service station.
- Ethanol production: includes direct and indirect emissions and energy use
associated with conversion of the wheat into ethanol.
- Feedstock transport: considers direct and indirect emissions, as well as
energy use for transport of wheat from the producer’s storage to the
ethanol plant.
33
- Feedstock production and recovery: includes direct and indirect emissions
and energy use in the production of wheat. Includes all fuel used in the
farming process and transportation from the field to the farmers storage.
- Fertilizer and chemicals manufacture: Considers direct and indirect life
cycle emissions and energy use for fertilizers and pesticides used for
feedstock production, including raw material recovery, transport and
manufacturing of chemicals.
- Co-product credits: considers the GHG emissions and energy associated
with the material (animal feed) displaced by the co-product (distillers’
grain) [31, 33].
- Land use changes and cultivation associated with biomass derived fuels:
related to emissions associated with the change in the land use in
cultivation of crops, including N2O from application of fertilizer and
decomposition of crop residues, changes in soil carbon and biomass
resulting from cultivation, and increased methane emissions from soil.
- Carbon in fuel from air: carbon dioxide that was drawn from the
atmosphere to produce the ethanol feedstock. It is net of CO2 vented in
fermentation and carbon in the co-product.
- Emissions displaced by co-products of alternative fuels: includes the
emissions displaced by co-products of wheat ethanol (distillers’ grains).
System expansion is used to determine displacement ratios for distillers’
grains in animal feed rations.
34
- Vehicle Operation: the emissions from the combustion of the ethanol in a
vehicle.
Fig. 4 represents the transportation stages that are eliminated, and the process
sections where elimination in energy input are possible, when an IDF ethanol
production approach is considered:
Figure 4: BFD for the conventional ethanol production process
A comparison of the energy balance for the conventional process (Fig. 2) and
IDF configuration (Fig. 3) for production of wheat ethanol is shown in Table 3. The
energy balance is measured in Joules consumed per Joule delivered. The lower
this value, the more energy efficient the process is. This value can also be
reported as EROEI (energy return on energy invested) in units of J delivered / J
consumed, in which case, a more energy efficient process will show greater
35
values. The life cycle stages in which the main differences are found between
the two processes are:
- Feedstock transport: Elimination of transportation of wheat from the farm to the
centralised ethanol facility because of the installation of the ethanol plant on
the farm for the IDF configuration.
- Fertilizer manufacture: Since in the IDF configuration the fertilizer is obtained
from the digestate resulted from anaerobic digestion of manure and thin
stillage, there is no transportation of fertilizer to the farm or any energy required
for fertilizer manufacture.
- Ethanol production: In the IDF configuration, utilities as steam and electricity
are obtained from the biogas produced in the anaerobic digester. Moreover,
the animal feed is obtained as wet grains, eliminating the need of DGS drying as
is the case in centralised ethanol facilities.
Table 3: Energy balance for ethanol production using wheat as feedstock [31]
Scenario Conventional IDF Life Cycle Stage Joules consumed / Joule delivered Ethanol dispensing 0.0075 0.0075 Ethanol distribution and storage 0.0223 0.0223 Ethanol Production 0.7737 0.3470 Wheat transport 0.0205 0.0000 Wheat recovery 0.0305 0.0305 Agrochemical manufacture (fertilizer + pesticide)
0.1526 0.0101(1)
Co-product credits -0.1014 -0.1014 TOTAL [J consumed / J delivered] 0.9056 0.3160 EROEI [J delivered / J consumed] 1.1042 3.1646
36
1. 93.37% of the agrochemicals corresponds to fertilizer (value obtained from GHGeinus 4. 03a [32])
A comparison of the GHG emissions for the conventional and IDF configurations
is shown in Table 4. The GHG emissions are measured in g CO2 equivalent per
GJ. The grams of CO2 equivalent are the amount of greenhouse gas (GHG)
multiplied by the global warming potential (GWP) of the GHG over 100 years
timescale [34]. The life cycle stages in which the main differences are found
between the two processes are in the feedstock transport and the fertilizer
manufacture.
Table 4: Greenhouse gas emissions for ethanol production using wheat as feedstock through conventional process and IDF configuration [31]
Scenario Conventional IDF Life Cycle Stage GHG Emissions [g CO2eq / GJ] Ethanol dispensing 771 771 Ethanol distribution and storage 1,754 1,754 Ethanol Production 48,604 19,729 Wheat transport 1,622 0 Wheat recovery 2,713 2,713 Agrochemical manufacture (fertilizer) 9,444 0 Co-product credits -24,524 -24,524 Land-use changes, cultivation 7,705 7,705 TOTAL [g CO2eq / GJ] 48,089 8,148
1.6.2 Production based on local farming yields
The wheat-land yield plays a significant role in the production of ethanol. If the
yield is higher, more feedstock for ethanol production can be obtained from
specific land acreage. This fact is of importance in the food vs. fuel debate,
37
since it results in the possibility of producing enough wheat to satisfy all market
demands, including the food and biofuels sector.
In the case of centralised large scale ethanol plants, the amount of wheat
received from the farmers will be based on an average yield. On the contrary,
when the ethanol plant is installed on the farm, the raw material supplied for the
production will be based on the local farm yield. There are several farming
communities with highly efficient farming, which can obtain better yields than
average. For instance, in the case of Alberta, the average yield is 50 bushels of
wheat per acre of land [35]; however, specific farming communities in the
region have reported higher yields of around 80-95 bushels per acre [36].
1.7 Project Objectives
In order to improve the production of biofuels, concerns are raised regarding
the optimization of the process to make the technology more energy and cost
efficient, as well as to mitigate the environmental impact caused by GHG
emissions. The production of ethanol from wheat in an integrated and
decentralised configuration, in which the farm operation is combined with the
ethanol facility with the objective of maximizing the use of co-products
(recycling of animal feed and fertilizer to the farm operation) and the co-
generation of utilities (steam and electricity from biogas produced in the AD
system), allows a biofuel production that benefits the farmer’s economy,
improves the energy balance and reduces GHG emissions.
38
The objective of the present study is to evaluate the feasibility of installing an
ethanol plant within an integrated decentralised on-farm (IDF) concept. The
technical assessment is performed by:
- Conceptual level design and simulation of the ethanol plant including the
following sections: saccharification and fermentation, distillation and
dehydration, anaerobic digestion, and co-generation.
- Improvement of the process through heat integration using pinch analysis.
The economic assessment includes:
- Evaluating the farmer’s economic balance before and after the
installation of the IDF plant.
- Capital expenditures (CAPEX) evaluation through equipment costing.
- Consideration of economic indicators: earnings before interests, taxes,
depreciation and amortization (EBITDA), return on investment (ROI), and
payback period.
- Analysis of sensitivity to changes in ethanol, fertilizer and wheat prices.
The plant to be designed has the following characteristics:
- Location: 4,500 acres farm located in the province of Alberta.
- Feedstock: Wheat, class: CW Feed (sub-graded wheat)
- Configuration: Small scale ethanol production facility with anaerobic
digestion and co-generation system included.
39
- Product: Fuel grade ethanol with a minimum ethanol content of 99.5 wt%
- Co-products:
i) Biogas containing 60% methane and 40% CO2.
ii) Wet distillers’ grain to be used as animal feed
iii) Digestate to be used as fertilizer
- Capacity: 112 LPH / 954,000 LPY
- Days of operation: 355 days / year; 24 h / d
- Operation mode: Batch for saccharification and fermentation and
continuous from the distillation stage onwards.
40
2. Literature Review
2.1 Process and Equipment Selection
Different types of processes were evaluated for each section of the design. The
process selection was performed using the concept screening method [37]. In
this method, a selection matrix, which contains a list of criteria used to assess the
process alternatives, is completed. For each process, values of +1 are assigned
for each criterion in which the process has an advantage; while values of –1 are
assigned to the alternatives that do not show benefit for the specific category.
When there is no information available for a specific category, or the criterion is
not applicable for the evaluated process, a value of zero is assigned for all the
alternatives considered. The total or net score is then obtained and the
alternative processes are ranked. A table similar to the one presented below will
be used for selection purposes:
Table 5: Exemplary table for evaluating the different process alternatives using the screening method [37]
Alternative Process Equipment Criterion Process 1 Process 2 Process 3
1 + - + 2 - + + 3 - + + 4 + - - 5 + - - 6 - - +
Total score 0 -2 2 Rank 2 3 1
41
The criteria considered are described below:
- Ease of Operation: Ideally, the operation of the plant should be as easy
and straightforward as possible. The ethanol plant is intended for
operation by the personnel of the farm so that the operation should be
simple.
- Size: In terms of the size, small size equipment is preferred over large size.
Since the IDF plant is planned for installation on the farm, there is limitation
in terms of space available.
- Economic: The purpose of the IDF concept is to offer the farmer an
integrated facility that allows for improved economics from the OPEX
point of view with a low capital investment. In order to achieve this goal,
low cost equipment is the preferred option.
- Safety: Safe operation is the primary aim of the IDF ethanol plant. The
equipment selected must be safe to operate under various on-farm
installation conditions.
- Controllability: The selected equipment should be easy to control in order
for the farm operators to be able to run the plant without any difficulties.
- Practical Applications: It is important that the processes and equipment
selected are reliable over the long term. Processes already proven
successful in the industry are preferred over processes in the developing
stages.
42
- Environmental Impact: The process and equipment selected should
generate the lowest environmental impact possible with low GHG
emissions and energy consumption.
2.1.1 Feedstock preparation – Milling
The purpose of milling is to break down the cereal grains to an appropriate
particle size in order to enable the penetration of water for hydrolysis to take
place. There are two types of milling processes, wet milling and dry milling. Most
new distilleries use dry milling processes [38]. Wet milling is usually used by
processors that are more interested in by-products of ethanol production such
as dried starch, syrup, high-fructose sweetener, oil, or gluten meal for livestock,
than in fuel ethanol itself. On the other hand, fuel ethanol production is the
primary aim of dry-milling processes [39]. Dry mill plants are designed to convert,
through hydrolysis and fermentation, the sugars contained in powdered grain
into ethanol. The remaining flow containing fiber, oil, and protein is converted
into the co-product known as distillers’ grain (DG), used mainly as animal feed
[40]. The main difference between wet and dry milling is that in the former, the
germ, endosperm and bran of the grain kernel are mechanically separated
prior to fermentation, while in the second, the entire kernel is ground before
being treated with enzymes, fermented and distilled to obtain fuel ethanol. The
separation of the protein-rich germen and the starch containing endosperm is
achieved by soaking the grain in hot water prior to mechanical action [41].
43
Most of the ethanol produced in the US is obtained from dry milling plants. More
than 80% of currently operating facilities use this type of milling, the main reason
being the lower capital investment required for dry milling plants [42]. Dry-grind
processes are not only less capital intensive, they are also less energy intensive
than wet mill processes [43]. While wet milling facilities offer a more diverse
range of products, they are slightly less efficient in ethanol production (2.5 gal
per bushel compared to 2.8 gal per bushel) [43].
The following diagrams show the dry and wet milling processes[44]:
Figure 5: Dry milling ethanol production [44]
44
Figure 6: Wet milling ethanol production [44]
Table 6 collects the characteristics of the dry and wet milling processes. There
are mainly two types of mills available for grinding of cereal grains, the hammer
mill and the roller mill. Hammer mills are the preferred option in most distilleries
[38]. Clean grains should be fed to roller mills, which is not necessary with
hammer mills. The maintenance of hammer mills is easier. Even though the
replacement of the screen and hammers is more frequent, the time required for
their replacement is lower than that needed for rollers on a roller mill [45].
45
Table 6: Characteristics of the milling processes considered for the IDF Ethanol Plant
Alternative Process Criterion Dry milling Wet milling
Ease of Operation
Less complex operation because only fuel ethanol and 1 co-product (DG) are obtained
More complex operation because of additional processing to obtain several co-products such as oil, gluten meal, starch, syrup, among others
Size Less space requirement More space requirement Economic Less capital expenditures More capital costs due to
equipment required to produce co-products
Safety Basic safety considerations Basic safety considerations Controllability Easier to control because less
equipment is required More controls required due to process complexity
Practical Application
The majority of ethanol plants currently installed use dry milling processing. Better yields are obtained per bushel of grain processed
Only producers interested in co-product use wet milling processing. Lower yields are obtained per bushel of grain processed
Environmental Less energy intensive More energy intensive
2.1.2 Cooking and Fermentation
2.1.2.1 Cooking
Cooking is the process of converting starch into fermentable sugars (glucose).
Starch is a polysaccharide consisting of thousands of glucose units. It exists in 2
forms: amylose and amylopectin. Amylose is typically formed by more than 1000
units of D-glucopyranose (glucose) that are connected through α-1,4 linkages.
Amylopectin is similar to amylose with branches each of 20 – 25 glucose units
connected through α-1,6 linkages [46] as shown in Fig. 7 and 8.
46
Figure 7: Amylose structure
Figure 8: Amylopectin structure
Wheat starch is roughly three quarters amylopectin and one quarter amylose
[47]. In order to release glucose units from starch, enzymatic action is required to
break the amylose α-1,4, and the amylopectin α-1,6, linkages.
O
CH2OH
O
OH
OH
O
1
23
4
5
6
O
CH2OH
OH
OH
O
1
23
4
5
6
α-1,4
n
O
CH2OH
OH
OH
O
1
23
4
5
6
O
CH2OH
OH
OH
O
1
23
4
5
6
O
CH2OH
OH
OH
1
23
4
5
6
O
CH2
4
O
OH
OH
1
23
5O
CH2OH
OH
OH
O
1
23
4
5
6
O
CH2OH
OH
OH
O
1
23
4
5
6
α-1,6
6
α-1,4
47
Cooking consist of 3 basic steps: gelatinization, liquefaction and
saccharification.
a) Gelatinization: Gelatinization is the breakdown of starch granular
structure. During gelatinization, the granules absorb water and swell,
gradually losing their crystalline structure. They become large, gel-filled
sacs that break with abrasion and agitation preparing the material for the
enzymatic action. At the gelatinization point the mash reaches the
highest viscosity. Gelatinization is achieved at different temperatures, for
wheat the recommended temperature is between 58 and 64 °C [38].
b) Liquefaction: Liquefaction is the hydrolysis of starch polymers to obtain
shorter straight chains called dextrins. The addition of the enzyme α-
amylase on the exposed starch molecules breaks the α-1,4 linkages of
amylose and amylopectin to obtain dextrins. As a consequence of the
breakdown of the polymer, a reduction in viscosity is achieved.
Liquefaction is generally performed at high temperature (80 to 90 °C) and
pH of 6 to 6.5 [47].
c) Saccharification: Saccharification is the process that releases glucose
units from dextrins. The α-1,6 linkages are broken by the action of
glucoamylase. Saccharification produces glucose fermentable
monomers ready for fermentation [38]. The optimal pH for glucoamylase is
around 4 to 5.5 [40].
48
2.1.2.2 Fermentation
Fermentation is the process by which glucose is transformed into ethyl alcohol
by the action of yeast. The reaction that takes place produces two molecules of
ethanol and two molecules of carbon dioxide from 1 unit of glucose as follows:
C6H12O6 2 C2H5OH + 2 CO2 Rxn. 1
Yeast converts glucose to ethanol through a series of biosynthetic reactions, in
which, glucose is phosphorylated, split in glyceraldehyde-3-phosphate and
dihydroxy-acetone phosphate, which are finally converted into pyruvic acid
that yields ethanol and carbon dioxide under anaerobic conditions [38]. The
most commonly used microorganism for fermentation of glucose to ethanol is
Saccharomyces cerevisiae. This yeast strain proves to be resistant to stressful
conditions like variations in temperature and pH [38]. It can produce ethanol to
concentrations as high as 18% of the fermentation broth [48]. In batch
processes, yeast undergoes a series of phases: lag phase, accelerating phase,
log phase (exponential growth), deceleration phase, stationary phase and
declining phase. During the lag phase, the yeast adapts to its new environment,
synthetizing enzymes and activating its metabolism. With the first cell division, the
accelerating phase starts, in which the rate of division increases continuously. A
maximum and constant growth rate is achieved in the log phase. Under optimal
conditions, the time required to double the cell number is between 90 and 120
minutes. The deceleration phase is characterized by the decrease of growth
49
rate due to the presence of inhibitors or the reduction in nutrients. When the
number of yeast cells stays constant, the stationary phase is reached. There is a
balance between the number of new cells formed and the ones that die.
During the declining phase, the death rate of yeast cells exceeds their birth rate,
thus the total number of cells decreases [38]. Fermentation is usually achieved
within 48 hours and is of critical importance since it is important to obtain the
most ethanol possible without sacrificing the time efficiency of the process [49].
Conventionally, cooking and fermentation have been carried out as isolated
steps in a process known as separate hydrolysis and fermentation (SHF).
However, in recent years, simultaneous saccharification and fermentation (SSF)
processes have been developed. In SSF, the enzymatic degradation of starch is
combined with the fermentation of the glucose monomers obtained from
hydrolysis of starch. One of the advantages of SSF is that inhibition caused by
sugars is avoided because of the ability of the process in converting sugars into
ethanol as soon as they are formed. SSF is considered as an optimal process for
ethanol production because it requires less investment and lower enzyme
loading. Yeast fermentation helps to reduce end product inhibition from the
sugar monomers formed during hydrolysis [50-53]. The sugar produced during
starch breakdown slows downs the α-amylase action; this inhibition effect is
reduced in SSF since the presence of yeast or bacteria along with enzymes
minimizes the sugars accumulation in the vessel [54]. According to Wyman, SSF
can reach higher rates, yields and ethanol concentrations in comparison with
50
SHF [55]. Neves states that, at lower enzyme loading, higher rates, yields and
concentrations of ethanol are possible during SSF compared to SHF [54].
SSF has an easier operation and a lower equipment requirement since no
hydrolysis reactors are needed [56]. Other benefit of SSF compared to SHF is that
the action of undesirable microorganisms can be avoided by the presence of
ethanol in the broth [57]. One inconvenient characteristic of SSF is that it is
complex to control and to optimize, since the optimal operating parameters for
hydrolysis and fermentation, mainly temperature and pH, are different [56].
Some research studies have been successfully conducted for the SSF of wheat.
Neves et al. showed the suitability of low-grade wheat flour as substrate for
ethanol production through SSF [58]. Hofvendahl et al. performed SSF
experiments for the conversion of starch from whole-wheat flour. The researcher
developed a kinetic model that was verified experimentally [59]. Öhgren et al.
demonstrated higher overall ethanol yields for SSF of steam-preheated corn
stover in batch experiments. The study presented a 72.4% of stoichiometric /
theoretical yield for SSF compared to 59.3% obtained for SHF [60]. In terms of
economic analysis, Wingren et al. found that the lower capital cost and the
higher overall ethanol yield attained from SSF, resulted in an ethanol production
cost of 0.43 USD / L as compared with 0.63 USD / L for SHF [61].
The following table presents a comparison of the SSF and SHF processes:
51
Table 7: Characteristics of the cooking and fermentation processes considered for the IDF Ethanol Plant
Alternative Process Criterion Simultaneous Saccharification
and Fermentation SSF Separate Hydrolysis and
Fermentation SHF Ease of Operation
Simpler operation because saccharification step is avoided and the presence of ethanol in the broth inhibits the action of undesired microorganisms.
More complex operation because saccharification and fermentation is performed as 2 separated steps
Size Less space requirement More space requirement Economic Less capital expenditures
because saccharification equipment is not required
More capital costs due to equipment required for saccharification step
Safety Basic safety considerations Basic safety considerations Controllability Harder to control and optimize
because of the different operating conditions of the saccharification and fermentation steps.
Easier to control because fermentation and saccharification are performed as separate steps.
Practical Application
Relatively new technology More extensively used in the industry
Environmental Less energy intensive More energy intensive
As described above, mash cooking involves a high temperature enzymatic
liquefaction and saccharification step. New enzymatic developments have
made it possible to achieve hydrolysis of starch granules at low temperatures.
The enzymes used are called cold enzymes because they operate at
fermentation temperature of around 30 °C, as compared to the conventional
hydrolysis process in which temperatures greater than 90 °C are required.
Operating costs are reduced by decreasing the thermal energy input
requirements when cold enzymes are used [62]. Moreover, the lower
temperatures avoid the denaturisation of the enzymes that is sometimes
52
observed when high temperatures are used. Purohit and Mishra [63]
demonstrated a maximum efficiency in ethanol production of 96.7% when
sweet potato powder was treated with α-amylase in a SSF configuration.
There are commercially available cold enzymes that can be used together with
fermentation yeast in SSF processes for ethanol production. Stargen 002 [64] is a
combination of α-amylase and glucoamylase that maintains its activity under
the pH and temperature conditions of SSF. The operating temperatures of the
enzyme are between 20 and 40 °C with a pH of around 3.3 to 4.5. The SSF
process is achieved in 24 to 76 hours. Thorough agitation and a pre-treatment
tank for hydration are recommended. For wheat, a temperature of 56 to 57 °C is
optimal for the hydration process. The Stargen 002 enzyme dosage required is of
0.8 to 1.6 kg per metric ton of grain [64].
2.1.3 Distillation
After fermentation of grains, the beer obtained is an aqueous solution
containing 5 to 12 weight % ethanol [65]. Studies performed on wheat showed a
concentration of ethanol in beer after fermentation of around 12 to 14 volume
% [66]. To separate the ethanol from the other components (mainly water),
distillation is most often used. Distillation allows a separation of water and
ethanol up to 95 ethanol weight % (89.5 mole %). This is known as the azeotropic
point, when a constant boiling mixture of ethanol-water is formed, meaning that
the liquid will have the same temperature as the equilibrium vapour being
53
boiled from itself. When the azeotrope is formed, the composition of the vapour
is identical to the composition of the liquid solution from which it originated. Thus,
further purification of the ethanol product when this point is reached is not
possible through conventional distillation techniques. In order to concentrate
the ethanol solution beyond 95 wt %, dehydration technology is required [38]. In
general, distillation has proven effective for solutions containing 10 to 85 ethanol
wt %. For mixtures containing more than 85 wt %, distillation becomes expensive
because the feed ethanol concentration is near the azeotropic point (95.6
weight %), thus, increasing reflux ratios and creating the need for additional
equipment [67]. The figure below shows the vapour– liquid equilibrium curve for
ethanol – water mixtures at atmospheric pressure:
54
Figure 9: Vapour-liquid equilibrium curve for ethanol water mixtures at atmospheric pressure [38, 68]
The typical configuration of a complete distillation tower consists of a stripping
and a rectifying section. The mixture of components to be separated is fed to
the stripping section. The source of energy to drive the process is generally
steam that is provided through direct injection or an indirect heat exchanger
called a reboiler. The energy is supplied at the base of the column, where the
highest temperature will occur. The purified product, consisting mainly of the
component with the lowest boiling point (ethanol in this case), is recovered at
the top of the column as a vapour that is condensed in an overhead heat
exchanger called a condenser. The condensed vapour is split in two streams:
0
10
20
30
40
50
60
70
80
90
100
0 10 20 30 40 50 60 70 80 90 100
Etha
nol i
n va
por [
mol
e %
]
Ethanol in liquid [mole %]
Equilibrium curve
Azeotrope (89.5 mole %)
Distillation Dehydration
55
the distillate or tops product and the reflux which is sent back to the tower as
liquid downflow for the top portion of the tower.
As mentioned previously, there are two main configurations for the energy
supply in the bottom of the tower: direct steam injection and the use of a
reboiler. Direct steam injection has been the common practice in beer
distillation because of its simplicity [38]. In this configuration, the steam is directly
injected into the column and mixes with the product. The main benefits of direct
steam injection are the simple operation and the low investment and
maintenance costs. With the use of a reboiler, problems with high temperature
impacting the by-product quality, the loss of the steam condensate for the
boiler system and the increase in the stillage volume and its dilution can be
avoided [49, 69].
To accomplish the vapour-liquid mass transfer required for separation, proper
design of the tower internals is of the utmost importance. Distillation column trays
facilitate the intimate contact between the rising vapours and the descending
liquids in the tower [38, 70]. There are basically two types of tower internals that
can be used: trays or packing. For applications where solids are present and
fouling is a concern, trays are generally used [70]. Trays are also used for liquid
rates of 30 m3/m2.h or above. For applications in which minimizing the column
pressure drop is important, and for lower liquid rates (<50 m3/m2.h), structured
56
packing is recommended. Random packing is generally used for higher liquid
rates in which low pressure drop is the goal.
Pressure drop is lower in packed towers than tray columns because of their
larger open area. Packing is particularly appropriate for vacuum applications.
Packed towers are shorter in height. They offer additional benefits as
mechanical simplicity, ease of installation, and the possibility of using cost-
effective corrosion-resistant materials for their fabrication. Rigorous maintenance
is required for packed internals to stay reliable in the long term. Further, solid
handling is not efficient in packed towers. At high operating pressures, the
pressure drop is similar for packed and tray columns. Because more material per
unit area is required for packings, it is usually more expensive than trays. Liquid
distribution is a critical factor for the good operation of packed towers. The most
prevalent cause of operation failure in these distillation columns, is the mal
distribution of the liquid downflow [70].
Usually, dry mill grain beer is distilled using baffle or disc and donut trays because
they present good characteristics to handle solids and prevent fouling. The
figures below present schematics of these types of trays [38].
57
Figure 10: Disc and donuts trays [38]
The following table collects the characteristics of the two main types of
packings available as distillation tower internals.
Table 8: Characteristics of the packing types considered for the IDF Ethanol Plant
Alternative Process Equipment Criterion Trays Packings
Ease of Operation
Less complex operation because solids are effectively handled and fouling is avoided. Less maintenance is required.
More complex operation because of the possibility of fouling of the packing. Constant maintenance is required.
Size Larger distillation column Shorter distillation column Economic Less capital expenditures More capital costs due to
material needed to create large surface area per unit volume
Safety Basic safety considerations Basic safety considerations Controllability Conventional control strategy Less attenuation of load
disturbances because of smaller liquid holdup
Practical Application
The majority of ethanol plants use trays because of the solids present in the beer
Packed columns are not the most common choice for dry mill grain beer because of the solid handling
Environmental Basic environmental concerns Basic environmental concerns
58
2.1.4 Dehydration
In order to produce anhydrous fuel grade ethanol (>99 wt %), an additional step
[71] after distillation is required. Dehydration allows separation of ethanol and
water over the azeotropic point. There are three main technologies available for
ethanol dehydration: modified distillation techniques, including azeotropic and
extractive distillation (ED); pressure swing absorption or molecular sieve (MS)
technology; and vapour permeation (VP) [38, 72]. In azeotropic distillation, a
third volatile component or entrainer is added so as to form a ternary azeotrope
that changes the relative volatilities in the distillation system allowing for
separation of the intended components. The most commonly used entrainers for
ethanol dehydration are benzene, toluene and cyclohexane. Azeotropic
dehydration systems consist of 2 columns, one dehydration or azeotropic
column and an entrainer recovery column. This type of ethanol recovery has
high energy requirements, large capital cost and there are health and safety
concerns related to the entrainers employed: benzene because of its
carcinogenic nature and cyclohexane because of its flammability [65].
Extractive distillation is similar to azeotropic distillation in that a third component
is added to increase the relative volatility of the components to be separated.
In ED, a high boiling point solvent, called extractant, is used to achieve
separation. The solvents used can be liquids, ionic liquids, dissolved salts,
polymers, among others. One of the most common ED solvents is ethylene
glycol. Gasoline is sometimes used for extractive distillation of beer solution to
59
obtain motor fuel ethanol [65]. Pressure swing adsorption or vapour-phase
adsorption of water from ethanol-water mixtures generally uses inorganic
adsorbents such as molecular sieves to achieve dehydration. Zeolites molecular
sieves are extensively used in ethanol dehydration. The most commonly used
molecular sieves for separation of ethanol-water mixtures are type 3A and 4A
zeolites [73]. In molecular sieve applications the separation is carried out by the
penetration of water molecules in the zeolite pores while the ethanol molecules
are retained in the vapour [65]. Another process for ethanol dehydration is
vapour permeation. This process is carried out by membranes. The feed to the
membrane is in the vapour phase as well. The permeate and the purified fluid
are recovered as vapours that are generally subjected to condensation. Since
in this process the fluids are vapours, the driving force for vapour permeation is
the pressure difference between the feed and the permeate side of the
membrane as shown in the diagram below.
Feed vapour A + B P1 Permeate vapour, A
P2
Purified vapour, B
E-1 Condensed Permeate
E-1
Condensed purified fluid
Figure 11: Schematic of vapour permeation. P1 is greater than P2.
60
In the ethanol dehydration process, ethanol and water vapour are supplied to
the feed side of the membrane module at a certain pressure (P > Patm). When
the ethanol/water mixture gets in contact with the nonporous polymer
membrane fiber, water diffuses to the outside of the fibers (permeate side of the
membrane) through the polymer walls. The permeate side of the fibers is
maintained at vacuum. The driving force for the process is the pressure
difference between the feed and permeate side of the module. Ethanol is
retained in the fiber lumen (hollow fiber space inside the walls) and recovered
as anhydrous product in the other side of the membrane module, as shown in
the following figure:
Vapour feed – hydrous ethanol
Purified retentate – anhydrous ethanol
Permeate – water
Non porous membranes fibers
H2O - water
C2H5OH Anhydrous ethanol
C2H5OH + H2O Hydrous ethanol
Figure 12: Schematic of membrane vapour permeation module for ethanol dehydration
61
Vapour permeation through polymeric membranes offers energy savings, long
membrane life-time, simple operation, ease of scaling the process up and
down, among other benefits over other technologies [74].
Moura reported lower water consumption for ethylene glycol extractive
distillation and molecular sieve dehydration compared to cyclohexane
azeotropic distillation. It was also found that molecular sieve process requires a
recycle representing 15% of the total alcohol production, which is usually sent
back to the rectifying column [71].
In terms of safety, health and the environment, the best options are vapour
permeation and molecular sieves since cyclohexane and ethylene glycol are
flammable solvents, and because of the carcinogenic nature of some
entrainers used in azeotropic distillation like benzene [65, 71].
Regarding the economics of the process, cyclohexane azeotropic distillation
has the lowest capital investment since no vacuum or high pressures are
required lessening construction complexity, as well as piping and automation
costs. Ethylene glycol ED has a modest initial investment, while molecular sieves
and vapour permeation membrane modules represent the higher investment
among the dehydration processes assessed. In the case of molecular sieves, the
reason for the higher costs are the equipment that is designed to operate under
vacuum, the instrumentation required and the fact that larger plants are
necessary for the same capacity [71].
62
In terms of size, vapour permeation membrane modules, because of their
modularity configuration, represent the smallest size as compared with
molecular sieves or distillation columns in azeotropic distillation and extractive
distillation.
Azeotropic distillation processes require the simplest automation among the
dehydration technologies presented, while molecular sieves require a good
control system especially because it works in virtually unsteady state with
frequent shifting between the adsorption and regenerating cycles [71]. Vapour
permeation modules are not as complex as mole sieves since the process
operates in steady state with the degree of superheat being the most important
parameter to control [36].
The following table collects the characteristics of the main dehydration
technologies available for separation of water-ethanol mixtures above the
azeotropic point.
63
Table 9: Characteristics of the dehydration technologies considered for the IDF Ethanol Plant
Alternative Process Criterion Modified Distillation
Techniques: Azeotropic Distillation and ED
Molecular Sieves Vapour Permeation
Ease of Operation
Complex operation because of the configuration of the process requiring a combination of distillation towers.
Complex operation because of the unsteady-state operation due to shifting of the adsorption and regeneration cycles
Simpler operation
Size Process configuration requires several distillation towers
Process requires adsorption and regeneration beds
Small size because of the modularity concept, compact nature membrane fibers
Economic Less capital investment because no vacuum or high pressure equipment is required
High capital investment
High capital investment
Safety Safety concerns due to health and flammability issues of the solvents
Basic safety considerations
Basic safety considerations
Controllability Conventional control strategy
Complex control because of unsteady-state conditions
Conventional control strategy
Practical Application
Less used because of the health and safety concerns related to the solvents employed
Most ethanol plants currently use this technology
Relatively new technology
Environmental Energy intensive process
Energy efficient technology
Most energy efficient technology
64
2.1.5 Anaerobic Digestion Anaerobic digestion (AD) is the microbiological decomposition of organic
feedstocks under anaerobic or oxygen-free conditions. The main product
obtained from AD is biogas, a combustible gas that has many uses including the
production of electricity or as a natural gas substitute after it is upgraded [75]. In
the absence of oxygen, anaerobic microorganisms digest organic materials to
produce mainly methane and carbon dioxide. Some trace compounds such as
hydrogen sulfide (H2S) and ammonia (NH3) are also part of the biogas
produced during AD [76]. Biogas typically contains 60% methane (CH4) and 40%
carbon dioxide (CO2). Methane content in natural gas is usually 97 to 99% [77].
A complex mixture of microorganisms catalyzes the conversion of biomass to
biogas. There are basically 4 steps in biogas formation: hydrolysis, which consists
in the degradation of organic macromolecules into monomers, is followed by
an acidogenesis step in which the monomers produced are transformed into
volatile fatty acids. The next step is the formation of acetic acid in a process
called acetogenesis. Finally, through methanogenesis, the acetic acid is
transformed into methane [78]. The four phases of AD are described in the
diagram below [79]:
65
Figure 13: Anaerobic Digestion Process [79]
AD systems offer several benefits, especially in terms of environmental concerns.
Pathogen and odour reduction of over 90% can be achieved by the addition of
AD processes. This is attained through the microbiological breakdown of volatile
organic compounds and pathogenic organisms. As well, greenhouse gases
(GHG) can be captured and destroyed. Important co-products of AD include
livestock bedding, thermal energy and nutrients [75].
There are two main types of anaerobic digestion processes: thermophilic and
mesophilic. Mesophilic digesters operate in the range of 20 to 45 °C and usually
at 35 °C. Thermophilic digester temperatures range from 50 to 65 °C, with 55 °C
being the most used one. Thermophilic digesters present lower retention time
compared to mesophilic digesters. Although they require higher heat input and
are sensitive to operating and environmental variables, they are efficient in
terms of organic loading rate and gas production. The higher operating
66
temperature of thermophilic AD provides the sterilization required to eliminate
pathogens, viruses and seeds [76].
One of the benefits of thermophilic AD is that the organic loading rates are
almost twice those of mesophilic systems, allowing the use of smaller digesters
and thereby improving the process economics [49]. Nevertheless, mesophilic
digesters are more commonly used because of the greater availability of
mesophilic inoculum compared to thermophilic inoculum and the less sensitive
operation of mesophilic digestion. Moreover, mesophilic operation is easier to
start-up, and readily achieves optimal growth conditions. Despite the limitations
of thermophilic AD, Wilkie et al. reported the existence of 149 facilities treating
stillage from ethanol distillation [49]. The time needed to achieve the complete
degradation of organic matter (or retention time) for mesophilic reactors ranges
from 15 to 30 days, while thermophilic digesters require 12 to 14 days [76].
The following table summarizes the characteristics of these anaerobic digestion
processes.
67
Table 10: Characteristics of the anaerobic digestion processes considered for the IDF Ethanol Plant
Alternative Process Criterion Thermophilic Digestion Mesophilic Digestion
Ease of Operation
Sensitive to operating and environmental variables. Faster operation with retention time of 12 to 14 days.
Easy to start up and to stabilize to optimal growth conditions. Slower operation with retention time of 15 to 30 days.
Size Small because of high organic loading rates
Large because of low organic loading rates
Economic Less capital investment because of smaller size
High capital investment because of larger size
Safety Basic safety considerations Basic safety considerations Controllability Easier to control because of
the elimination of pathogens achieved due to high operating temperature
Rigid temperature control because microorganism might lose activity at temperatures greater than 45 °C
Practical Application
Not as commonly used as mesophilic digestion
Commonly used
Environmental High energy input to reach operating temperature
Low energy input to reach operating temperature
Some studies have been conducted for anaerobic digestion of ethanol
production by-products and farm residues. In current ethanol production from
wheat, 65% of the energy in the biomass feedstock ends up in the fuel ethanol
produced, while the remaining 35% is found in the distiller’s grain co-product
[78]. Stillage represents a good potential substrate for biogas production. As
well, from farm operation, manure represents an important substrate for AD, that
otherwise is considered as a residue that requires additional effort to be
managed and handled. The effluent leaving the digester is known as digestate.
It has nutrient value and can be applied to the crops like manure but it is less
odorous, reducing odours from livestock facilities by around 80% [25]. When AD
68
systems are installed on farms, a better management of manure is possible,
therefore a reduction of the risk of ground or surface water contamination due
to pathogens and nutrient leaching run-off is achieved [80]. According to Wu-
haan, manure provides buffering capacity and a wide range of nutrients, while
plant material (stillage) provides a high carbon content that balances the
carbon to nitrogen ratio (C/N) of the substrate, decreasing the risk of ammonia
inhibition and increasing biogas production [81].
According to Börjesson and Mattiasson, an important benefit of biogas systems
lies in their ability to reduce GHG emissions especially when manure is used as
source material [78].
Westerholm carried out experiments in which whole stillage was co-digested
with cattle manure for obtaining biogas through AD. The study demonstrated a
higher efficiency when stillage was used in combination with manure as
substrate for the AD process. Biogas productivity and process stability was
improved when whole stillage was used along with manure in an 85:15 ratio
based on volatile solids (85% whole stillage and 15% manure). The study
demonstrated stable operation over 640 days. Although the total methane yield
decreased when manure was co-digested with stillage, under the conditions of
the study, when stillage was used as sole substrate, the operation was not
sustainable in the long term [82].
69
Annand performed a study in which thin stillage was co-digested with manure in
manure-to-thin stillage ratios of 1:1 and 2:1 (in volatile solids). The experiments
achieved 92% and 88% methane yields obtained when stillage was used as the
sole substrate for AD, respectively. A methane yield of 446 mL CH4 per gram of
volatile solids added was achieved with a kinetic constant of 0.11 day-1. Other
ethanol by-products such as whole stillage and wet cake were tested as well in
combination with manure. However, thin stillage generated the best biogas
production [83].
One of the disadvantages of thermophilic AD is that, in comparison with
mesophilic operation, it requires an additional energy input to reach the
necessary temperature. However, when using ethanol production co-products
as substrate for AD, there is the possibility of using stillage which exits the
distillation column at temperatures above 55°C. This would decrease the
heating demand and improve the metabolic rates, efficiency and economics
of the process [83].
There are various types of AD reactors, including complete-mix systems, plug-
flow systems, upward-flow anaerobic sludge blanket systems (UASB) and
covered lagoons.
Complete-mix systems or continuous stirred tank reactors (CSTR) are designed to
feed the material continuously and are completely mixed. The mixing in the
reactor is achieved by mechanical agitation, effluent recirculation or biogas
70
recirculation [84]. The effluent of the round insulated vessel is produced at the
same rate at which the feed flow is introduced in the reactor. The biogas
produced is collected in the roof of the vessel. The typical influent contains 5 to
12% total solids and 20 to 50 days of retention time is common. The majority of
on-farm AD systems in Canada are complete-mix systems because of the ability
of this configuration to produce biomethane for electricity generation [75].
Figure 14: Complete-mix anaerobic digestion system schematic [85]
71
Figure 15: Installed complete-mix anaerobic digestion system [85]
In plug-flow reactors, the material is fed semi-continuously to a horizontal
reactor. The flow is maintained by a feeding schedule and control of the influent
total solids. The biogas in these systems is captured in a gas-tight cover with
external gas storage. The total solids for this application are between 11 and
14%. Typical retention times for plug-flow systems are 15 to 20 days. Generally,
the productivity in terms of biogas is not as high in these type of systems as in
other anaerobic digesters, so its use is focussed primarily on odour and
pathogen reduction [75, 84].
Figure 16: Plug-flow anaerobic digester schematic [85]
72
Figure 17: Installed plug-flow anaerobic digester [85]
Upward-flow anaerobic sludge blanket systems (UASB) have a floating blanket
of granular material which represents a constant medium for bacterial growth.
In general the reactors are tall, vertical tanks. The material is fed at the bottom
of the tank and removed after the digestion on the top of the reactor. 30 to 90
days are required to allow the sludge to develop the proper AD conditions in
the reactor [75]. UASB digesters are suited for dilute streams containing around
3% total suspended solids [86]. These systems are usually difficult to manage and
they are labour-intensive, thus making them inappropriate for on-farm use [75].
73
Figure 18: Upward-flow anaerobic sludge blanket systems schematic [86]
Figure 19: Installed upward-flow anaerobic sludge blanket systems [87]
Covered lagoons are a low maintenance anaerobic digestion alternative in
which biogas is captured under an impermeable gas-tight cover [84]. The
74
system consists of 2 cells. The first one is covered, and the second cell is
uncovered [86]. Covered lagoons are designed for solids content below 2% and
to be operated at ambient temperatures. Retention times of 25 to 60 days are
common. Due to seasonal temperature fluctuations, biogas production in
covered lagoons is variable [84].
Figure 20: Covered lagoon for anaerobic digestion schematic [85]
Figure 21: Installed covered lagoon for anaerobic digestion [85]
2.1.6 Co-Generation System
Co-generation or combined heat and power systems (CHP) use biogas
obtained during AD processes to generate heat and electricity simultaneously.
75
Biogas can be fed to boilers in order to produce heat or used in a CHP system to
produce heat and electricity [76]. According to Eriksson, the addition of CHP
systems to existing ethanol plants, can reduce production costs and mitigate
CO2 emissions because of the use of residual energy from the production
process [88]. CHP is the most common application of biogas produced by
anaerobic digestion systems at farm scale projects since it is the most
economical [75]. Cogeneration can achieve energy savings of approximately
40% of the energy required when a gas boiler is used along with the purchase of
electricity from the grid [89]. Boilers do not have a high gas quality requirement.
However, the removal of H2S and water vapour is recommended because of
corrosion issues. Gas engines, used in CHP systems, have similar gas requirements
as boilers, but the H2S content should be lower [76].
CHP systems consist of different components, among which the prime mover or
heat engine drives the overall system. Steam turbines, gas turbines or
combustion turbines, spark ignition engines, diesel engines, microturbines, and
fuel cells typically serve as heat engines in CHP systems. The purpose of these
prime movers is to convert a variety of fuels into shaft power or mechanical
energy [90]. Other components of CHP systems are boilers, absorption chiller,
desiccants, and gasifiers [90]. Steam turbines convert steam energy from a
boiler into mechanical power. Gas or combustion turbines, use heat to move
turbine blades that generate electricity. Reciprocating internal combustion (IC)
76
engines use liquid and gaseous fuels to produce electricity through the
reciprocating shaft power and by the use of a generator.
For biogas applications, the fuel required for gas turbines and reciprocating IC
engines is biogas; while steam turbines work with biomass or steam from a
biogas-fueled boiler. In terms of the preparation of the fuel, gas turbines and
reciprocating IC engines require particulate matter filters for the biogas. As well,
the moisture should be removed from the fuel prior to its use. In steam turbines,
the fuel can be directly used without any preparation procedure. Reciprocating
IC engines can achieve the highest electric efficiency of up to 45%, while up to
36% can be achieved through gas turbines, and 30% with steam turbines.
Steam turbines are highly reliable. Even though they require a slow start-up,
steam turbines have long life and the maintenance infrastructure is readily
available with the lowest operating and maintenance costs (less than 0.4
¢/kWh) among the described alternatives [90]. In steam turbines, the energy is
transferred from the boiler to the turbine through high-pressure steam, which
powers the turbine and generator. Steam turbines work according to Rankine
cycle, in which liquid water is converted to high-pressure steam in a boiler and
fed into the steam turbine. Power is created through the rotation of the blades
within the turbine caused by the steam obtained from the boiler. This power is
converted to electricity with a generator. The steam exiting the turbine is
77
collected and condensed in a heat exchanger and then pumped back to the
boiler to complete the cycle, as shown in Fig. 22.
Figure 22: Steam turbine CHP system [90]
There are three types of steam turbines: condensing steam turbines, extraction
turbines, and back pressure turbines. Condensing steam turbines are used to
generate power exclusively (Fig. 18). Extraction turbines allow for electricity
generation and steam usage. They have openings in their casings through
which a portion of the steam can be extracted at some intermediate pressure
for use in the process. Back-pressure turbines are used for steam generation
exclusively. In these applications, the entire flow of steam is delivered to the
process or facility at the required pressure [90].
Gas turbines are reliable prime movers. Gas turbines are internal combustion
engines that operate with rotational motion. Gas turbines use the Brayton cycle
in which the air is compressed to high pressure in a compressor. Combustion
78
takes place in a chamber operating at high pressure. The pressurized hot
combustion gas products pass through the turbine itself generating the energy
required to operate the generator and the upstream compressor [90]. Figure 23
illustrates a gas turbine system.
Figure 23: Gas turbine CHP system [90]
There are three types of gas turbines: simple-cycle operation, CHP operations
and combined-cycle operation. For producing power only, simple cycle
operation using a single gas turbine is used. A simple-cycle gas turbine is used
along with a heat recovery/heat exchange unit to produce thermal energy in
the form of steam or hot water using the turbine exhausts heat. This
configuration is known as CHP gas turbine system. In combined-cycle
operations, high-pressure steam is generated from recovered exhaust heat and
used in a steam turbine to create additional power.
Gas turbines require a gas compressor, but have no cooling requirements are
presented. Usually gas turbines are large utility units. Maintenance infrastructure
79
is also readily available for this type of turbines. The operating and maintenance
costs vary between 0.6 and 1.1 ¢/kWh [90].
Reciprocating internal combustion engines recover heat not only from the
exhaust, as in the case of gas turbines, but also from the jacket water and the
engine oil that are part of the system. The heat from the engine exhaust is used
to heat the jacket water before being sent to the heat exchanger. Then, the
jacket water is pumped through the oil cooler heat exchanger and back into
the engine. The engine lube oil is pumped through a cooler as well and sent
back to the engine. Figure 24 is a schematic diagram a closed-loop heat
recovery system for a reciprocating internal combustion engine.
Figure 24: Reciprocating internal combustion engine CHP system [90]
Between 1,000 and 2,200 Btu can be recovered from the exhaust per kilowatt of
engine shaft power. Moreover, the jacket water allows for 4,000 Btu/kWh of heat
recovery. The engine lube oil allows for the recuperation of 300 to 900 Btu / kWh.
80
Reciprocating IC engines have a fast start-up. One of its disadvantages is that
these systems generate a lot of noise. Operating and maintenance costs range
from 0.8 to 2.5 ¢/kWh [90]. The characteristics of the main co-generation
systems are available in Table 11.
81
Table 11: Characteristics of the co-generation systems considered for the IDF Ethanol Plant [90]
Alternative Process
Criterion Steam Turbine CHP System
Gas Turbine CHP System
Reciprocating Internal
Combustion Engines CHP
Systems Ease of Operation Biogas can be used
as fuel without previous treatment. Up to 30% of electric efficiency. Slow start-up. Long life. Maintenance infrastructure readily available.
Particulate matter filter is needed. Up to 36% of electric efficiency. High-grade heat available. No cooling required. Gas compressor required. Maintenance infrastructure readily available
Particulate matter filter is needed. Up to 45% of electric efficiency. Fast start-up. Cooling required. Maintenance infrastructure readily available.
Size Various sizes Generally large size Various sizes Economic Low operating and
maintenance cost: < 0.4 ¢ / kWh. Installed cost: $350 - $750 / kW (without boiler)
Higher operating and maintenance cost: 0.8 - 2 ¢ / kWh. Installed cost: $700 - $2000 / kW
Higher operating and maintenance cost: 0.8 – 2.5 ¢ / kWh Installed cost: $800 – $1500 / kW
Safety Basic safety considerations
Basic safety considerations
Noisy system
Controllability Basic control strategy Basic control strategy More complicated due to process complexity
Practical Application
Highly reliable. The CHP system with the larger history of use.
Highly reliable. Less biogas applications
Environmental Higher GHG emissions
Low GHG emissions Low GHG emissions
82
2.1.7 Plant Co-products
2.1.7.1 Stillage
Stillage is the aqueous by-product remaining from the distillation of ethanol
following the fermentation of carbohydrates. Stillage is also known as distillery
wastewater, distillery pot ale, distillery slops, distillery spent wash, dunder, mosto
and vinasse [49].
During saccharification and fermentation starch is utilized in the transformation
of sugars into ethanol. Other nutrients such as protein and fibre are
concentrated during the process [91]. Whole stillage contains the fiber, oil,
protein, and other unfermented components of the grain, and yeast cells [92].
For each liter of ethanol produced, up to 20 L of stillage can be generated [93].
Stillage has high chemical oxygen demand (COD) and biological oxygen
demand (BOD) values which carry a considerable pollution potential [94]. COD
values greater than 100 g/L are typical for stillage. An ethanol facility producing
10 million liters of ethanol per year generates stillage with a pollution potential
equivalent to the sewage of a city with a population of 500,000 [95]. There are
several alternatives available for the treatment of stillage including
physical/mechanical separation, evaporation and membrane separation,
single cell protein production, calcium magnesium acetate, algae production,
and anaerobic digestion. After an extensive review of the scientific literature,
Wilkie et al. [49] found that anaerobic digestion is an economic and sustainable
stillage treatment scheme. During anaerobic digestion treatment, COD in
83
stillage is removed and converted to biogas that can be used as fuel for the
ethanol facility. Stillage anaerobic digestion can be improved by the addition of
cattle manure as co-substrate as demonstrated by Westerholm et al. [82].
The nutrients contained in stillage are conserved through anaerobic digestion.
While the organic content of stillage is removed by AD, plant macro-nutrients (N,
P and K) and micro-nutrients (Fe, Zn, Mn, Cu, Mg) remain in the sludge. The
application of the effluents from the anaerobic digester to croplands as fertilizer
allows for a productive nutrient cycle [49].
Whole stillage is fractionated into a solid fraction known as wet distillers’ grains
(WDG) or wet cake, and a liquid fraction or thin stillage. Thin stillage is usually
concentrated through multiple effect evaporators to produce syrup known as
condensed distillers’ solubles (CDS). Wet distillers’ grain with solubles (WDGS) are
obtained when CDS are combined with WDG. WDS, CDS and WDGS can be
sold as animal feed; however, WDGS are typically dried in order to lengthen the
product’s shelf-life, and marketed as dried distillers’ grains plus solubles (DDGS)
for animal feed [92, 96].
Wet distillers’ grains contain 50 to 75% of carbohydrates on a dry basis. WDG
derived from wheat usually present higher protein content compared to other
ethanol production feedstocks. According to Mustafa [91], distillers’ grains are a
source of both energy and protein when used as cattle feed. It was concluded
that wet distillers’ grains have a higher energy value than dried distillers’ grains
because they are not subjected to heat for drying as in the case of DDGS,
84
avoiding the heat damage of specific amino acids such as lysine, threonine,
arginine and alanine [91].
2.1.7.2 Bio-fertilizer
Anaerobic digestion generates two main products, biogas used as source of
renewable energy and digestate which is considered as a highly valuable bio-
fertilizer. Digestate can substitute mineral fertilizers and its fertilizer value depends
on the nutrients present in the feedstock. The amount of nutrients in the
digestate is the same as the ones contained in the feedstock used in the
anaerobic digestion process. The relative quantity of mineral fertilizer nitrogen
needed to obtain the same yield of crop as the quantity of total nitrogen
supplied in digestate is known as the utilization percentage. It is used to measure
the fertilizer value of nitrogen in digestate. The fertilizer value of digestate
increases with increasing utilization percentage [97]. Birkmose investigated the
percentage of nitrogen utilization of digestate compared with manure when it is
use for direct application in the fields. Digestate showed 80% nitrogen utilization;
the corresponding value for cattle manure was only 32% [98].
2.1.7.3 Biogas
Biogas is obtained during anaerobic digestion of organic matter. Biogas is
mainly composed of methane and carbon dioxide with traces of hydrogen
sulphide, ammonia, hydrogen, nitrogen and carbon monoxide. At standard
conditions, biogas has a calorific value of 21.48 MJ/m3 compared to 36.14
85
MJ/m3 for natural gas. The specific composition of biogas varies according to
the waste composition used in the anaerobic digester [76].
The following table presents the composition of biogas and natural gas [76]:
Table 12: Composition of biogas and natural gas
Component Units Biogas Natural Gas Methane, CH4 vol% 55-70 91 Ethane, C2H6 vol% 0 5.1 Propane, C3H8 vol% 0 1.8 Butane, C4H10 vol% 0 0.9 Pentane, C5H10 vol% 0 0.3 Carbon Dioxide, CO2 vol% 30-45 0.61 Nitrogen, N2 vol% 0 - 2 0.32 Hydrogen Sulphide, H2S ppm ~ 500 ~1 Ammonia, NH3 ppm ~ 100 0
86
3. Process Selection
In the literature review section of Chapter 2, the characteristics of the processes
considered were compared in tables according to different criteria. In this
chapter, process alternatives are graded according to these criteria in order to
facilitate process/equipment selection for the IDF ethanol plant. The evaluation
is based on the modified concept screening method explained in Chapter 2.
3.1 Feedstock Preparation – Milling
As is evident from Table 13, the best alternative for wheat grinding is the dry
milling process. It satisfies the requirement that the IDF plant produce fuel
ethanol together with animal feed.
Table 13: Criteria grading for feedstock preparation - milling
Alternative Process Criterion Dry milling Wet milling
Ease of Operation +1 -1 Size +1 -1 Economic +1 -1 Safety 0 0 Controllability +1 -1 Practical Application +1 -1 Environmental +1 -1 Total score 6 -6 Rank 1 2
For the feedstock handling component of the proposed IDF ethanol plant, dry
milling processing using a hammer mill was chosen. Farm organizations tend to
87
promote dry milling processes since the capital investment to build them is
lower, and they require less staff to run [99].
3.2 Cooking and Fermentation
The analysis presented in Table 14 suggests that the best alternative for the mash
preparation (cooking) and fermentation steps is simultaneous saccharification
and fermentation. This process fulfills the requirements of the IDF plant in terms of
good ethanol yields, low capital expenditures and easy operation. To overcome
the controllability issues related to the different operating temperatures of
saccharification and fermentation steps, the use of cold enzymes was adopted.
This type of enzyme is also beneficial in terms of energy requirements due to the
lower temperature required. The commercially available Stargen 002 from
Genecor [64] is considered well-suited to the process.
Table 14: Criteria grading for cooking and fermentation
Alternative Process Criterion Simultaneous Saccharification
and Fermentation, SSF Separated Hydrolysis
and Fermentation, SHF Ease of Operation
+1 -1
Size +1 -1 Economic +1 -1 Safety 0 0 Controllability -1 +1 Practical Application
-1 +1
Environmental +1 -1 Total score 2 -2 Rank 1 2
88
Because of the benefits ascribed above, the mash preparation and
fermentation of the IDF ethanol plant will be carried out by simultaneous
saccharification and fermentation configuration in batch operation with the use
of cold enzymes (Stargen 002 [64]) and yeast (Saccharomyces cerevisiae). In
order to provide a continuous flow of material to the downstream distillation and
dehydration section, parallel batch reactors will be used.
3.2.1 Simultaneous Saccharification and Fermentation Reaction Kinetics
In order to model the SSF reactor, a kinetic model obtained from the literature
will be employed. Lantz and Lee provided a mathematical model for SSF
process from starch to ethanol [100]. The model uses a known concentration of
the glucose obtained from the saccharification step and includes a set of three
differential equations describing the transformation of glucose to ethanol. The
model was used to predict the ethanol concentration in the SSF reactor effluent.
The glucose concentration used in the model was taken from another SSF study
conducted by Davis [101] in which Stargen and Saccharomyces cerevisiae are
used for the experimental validation of the model.
The model proposed by Lantz and Lee was chosen for the present study
because it predicts reactor behaviour in the same process configuration (SSF)
and using the same enzyme and yeast constraints. The model considers three
variables:
89
G: glucose (substrate, S) concentration
X: biomass concentration
E: ethanol concentration
The reactions which need to be considered in the SSF process are the following:
Starch (amylose + amylopectin) Dextrins
Rxn. 2
Dextrins Glucose
Rxn. 3
C6H12O6 2 C2H5OH + 2 CO2
Rxn. 4
For modelling the reactor, the following simplified reactions were used:
Starch + 800 H2O 800 C6H12O6 [100] Rxn. 5
C6H12O6 2 C2H5OH + 2 CO2
Rxn. 6
The stoichiometric coefficients for the reaction of starch to glucose were
obtained considering that starch is composed of 800 glucose units [100].
The system of differential equations contained in the model is given below:
α-amylase
glucoamylase
yeast
yeast
enzyme
90
Glucose Utilization:
dSdt = - �
μmaxSYXS(Ks+S) +ms�X Eq. 1
Ethanol Production:
dEdt = - �YPX
μmaxSKs+S +mp�X Eq. 2
Biomass Growth:
dXdt = �
μmaxSKs+S �X Eq. 3
Definitions and values of the model constants are provided in Table 15.
Table 15: SSF model parameters [100]
𝜇𝑚𝑎𝑥 Maximum specific growth rate of biomass 0.045 h–1 Ks Substrate (glucose) constant 0.025 g/L YXS Yield coefficient on substrate 0.15 g/g YPX Yield coefficient on ethanol 5.33 g/g ms Substrate maintenance rate constant 0.036 g substrate/
g biomass/h mp Product maintenance rate constant
0 g product / g biomass/h
3.3 Distillation
It is apparent from Table 16 that the best alternative for distillation of the beer
obtained from the fermentation process is a trayed column, particularly due to
91
its ability to cope with solids in the feed, as well as its lower capital cost when
compared to packed columns.
Table 16: Criteria grading for distillation
Alternative Process Equipment Criterion Tray Column Packed Column Ease of Operation +1 -1 Size -1 +1 Economic +1 -1 Safety 0 0 Controllability 0 0 Practical Application +1 -1 Environmental 0 0 Total score 2 -2 Rank 1 2
Two columns are present in the IDF plant design: a stripper and a conventional
distillation tower. The tray type selected for these units was disc and donut trays
because of their reliable operation in the industry [38]. For the stripper, direct
steam injection will be employed so as to avoid the problems which arise when
a slurry is passed through a reboiler. Another motivation for direct steam
injection is its lower capital cost. The stillage will be used for the production of
biogas so problems related to the treatment of higher loads of stillage are not a
concern for this section of the IDF plant. As well, the dilution of the stillage is
considered beneficial because it is used as the heating medium for the
hydration water for the grain in the pre-treatment tank. A more dilute slurry will
be passed easily through the heat exchanger. Concerning the conventional
distillation tower, the vapour distillate from the stripper supplies the heat to an
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external reboiler. In this configuration, the beer and the heating media flow in
independent closed loops, allowing the recirculation of the distillate to the
stripper for enhanced ethanol recovery.
3.4 Dehydration
As per Table 17, the best alternative for dehydration of the ethanol-water
mixture obtained from the distillation section of the IDF plant is membrane
vapour permeation. This process fulfills the requirements of the IDF plant based
upon ease of operation, size, and health and safety considerations.
Table 17: Criteria grading for dehydration
Alternative Processes Criterion Modified Distillation
Techniques: AD and ED Molecular Sieves Membrane
Vapour Permeation
Ease of Operation
-1 -1 +1
Size -1 -1 +1 Economic +1 -1 -1 Safety -1 +1 +1 Controllability +1 -1 +1 Practical Application
-1 +1 -1
Environmental -1 +1 +1 Total score -3 -1 3 Rank 3 2 1
Membrane vapour permeation modules will be used in the IDF ethanol plant.
93
3.5 Anaerobic Digestion
From Table 18, one would conclude that the best alternative for anaerobic
digestion of the manure-thin stillage mixture for the IDF plant is thermophilic
digestion. This process fulfills the requirements of the IDF plant, especially
regarding size and capital investment.
Table 18: Criteria grading for anaerobic digestion
Alternative Processes Criterion Thermophilic Digestion Mesophilic Digestion
Ease of Operation +1 +1 Size +1 -1 Economic +1 -1 Safety 0 0 Controllability +1 -1 Practical Application -1 +1 Environmental -1 +1 Total Score 2 0 Rank 1 2
Because of the benefits endorsed above, the anaerobic digestion of the IDF
ethanol plant will be carried out through a thermophilic process in a well-mixed
reactor. The feed to the anaerobic digestion is a mixture of manure and thin
stillage. Stillage exits the distillation tower at a temperature of around 90 °C [49].
Therefore, the high-temperature limitation of thermophilic operation is overcome
[102]. A mix tank was selected because of its simple and reliable operation with
total solids capacity up to 12% and its efficient biogas production. The United
States Environmental Protection Agency reported 201 operational anaerobic
94
digestion systems across US by the end of 2012; 31 % of which were well-mixed
reactors [102]. Even though plug flow reactors also allow for high solids content
between 11 and 14%, they are not as efficient in biogas production. Other
applications such as UASB and covered lagoons only allow for low solids content
[75]. Therefore, they are not options for the IDF plant in which the manure-thin
stillage mixture fed to the anaerobic digester contains solids mainly from the
manure. The hydraulic retention time for a thermophilic well-mixed anaerobic
digester is between 12 and 14 days. In the proposed design, the digester is
operated for 14 days at a time.
3.6 Co-generation System
The best alternative for the co-generation system is the steam turbine
combined-heat-and-power system (see Table 19). This setup will permit the on-
farm ethanol plant to self-generate the steam and electricity it requires. It is
particularly advantageous because of its low capital and operational costs,
high reliability and straightforward operation. Furthermore, it accepts untreated
biogas.
Since the ethanol plant requires steam as well as electricity, an extraction steam
turbine will be used. Steam for the turbine will be generated in a steam boiler
and a generator following the turbine will generate electricity.
95
Table 19: Criteria grading for co-generation system.
Alternative Processes
Criterion Steam Turbine CHP System
Gas Turbine CHP System
Reciprocating Internal
Combustion Engines CHP
Systems Ease of Operation
+1 +1 +1
Size +1 -1 +1 Economic +1 -1 -1 Safety +1 +1 -1 Controllability +1 +1 -1 Practical Application
+1 +1 -1
Environmental -1 +1 +1 Total Score 5 3 -1 Ranking 1 2 3
96
4. Methodology
4.1 Process Design
An integrated decentralised on-farm (IDF) ethanol plant was designed and
simulated in order to perform base-case material and energy balances. After
examining the system through pinch analysis, an improved process was
proposed with heat integration in order to reduce the utility steam consumption.
4.1.1 Software and Property Package Selection
Process simulation is a tool for designing and understanding chemical processes.
Simulation models are simplified mathematical representations of chemical
processes. The mathematical equations are typically coded into computer
software that also contains databases of physical and thermodynamic
properties of different chemical compounds [103]. There are a number of
commercially-available process simulation programs. VMGSim™, developed by
Virtual Materials Group Inc., was used for the simulation of the IDF ethanol plant.
The property package for the simulation was selected using the decision trees
provided by Carlson [104] and reproduced in Figs. 25 and 26.
97
Figure 25: Simulation software physical property package selection tree [104].
Figure 26: Selection tree for polar non-electrolyte mixtures [104].
98
The main components for the process are ethanol and water. The mixture of
these components is polar and non-electrolytic with an operating pressure lower
than 10 bars. Therefore, NRTL and UNIQUAC are the property packages most
appropriate for the simulation. For this study, NRTL was selected.
In order to validate this choice, densities and vapour-liquid equilibrium data
from the literature [105, 106] were compared against VMGSim™ predictions with
NRTL as the selected property package:
Figure 27: Vapour-Liquid equilibrium data for ethanol-water mixture.
0
10
20
30
40
50
60
70
80
90
100
0 10 20 30 40 50 60 70 80 90 100
Etha
nol i
n va
por [
mol
e %
]
Ethanol in liquid [mole %]
Experimental DataVMGSim - NRTL
99
Figure 28: Liquid density for ethanol-water mixture.
As depicted in these graphs, the values for density and vapour-liquid equilibrium
obtained from the simulator using NRTL as property package are similar to the
experimental data. These results justify the choice of NRTL for simulation of the
proposed IDF plant.
4.1.2 Simulation of Unit Operations
VMGSim™ contains an extensive library of unit operation models which may be
utilized for process simulation. In the following sections, a description of the
approaches adopted for modelling different sections of the plant is presented.
4.1.2.1 SSF Reactor
The concentration of ethanol obtained in the simultaneous saccharification and
fermentation (SSF) reactor was found using the mathematical model presented
700
800
900
1000
1100
1200
0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1
Dens
ity [K
g/m
3 ]
Mass Fraction of Ethanol
Experimental DataVMGSim - NTRL
100
by Lantz and Li [100] which was described in Chapter 3. The initial conditions for
solving the model were obtained from the supplier of the enzyme [64] and yeast
[107]. The resultant system of differential equations was then solved using
Polymath 6.1™ [108].
Simulation of the batch SSF reactor in the steady-state VMGSim™ [68] model
was then accomplished using conversion reactors in which the conversions were
specified so as to match the final ethanol concentrations generated by the
transient Polymath models.
4.1.2.2 CO2 Scrubber
The CO2 scrubber was modelled in VMGSim™ as an absorber. The number of
stages and water flow rate were designed in order to obtain 90% ethanol
recovery in the bottom of the absorber.
4.1.2.3 Distillation
Different configurations for the distillation section were simulated in VMGSim™ as
described in Section 4.2.2. Feeds to the distillation columns were modelled as
saturated (bubble point) liquids. To determine the number of theoretical stages,
the McCabe-Thiele graphic method was used [109]. The number of stages was
optimized by finding the optimal reflux ratio and reboiler duty.
101
4.1.2.4 Dehydration
The membrane dehydration section of the plant was simulated in VMGSim™
[68] as a component splitter. The water-rich permeate and the anhydrous
ethanol retentate splits were obtained using a tool provided by Whitefox
Technologies Canada [110]. The mathematical model is able to predict the
performance of candidate vapour permeation membrane modules for ethanol
dehydration. The concentrations and flows of permeate and retentate are
derived from process conditions such as feed flow, composition and pressure,
vacuum applied, and operating temperature.
4.1.2.5 Anaerobic Digester
The anaerobic digester was incorporated as an MS-Excel unit operation. The
concentration and composition of biogas obtained from the co-digestion of
manure and thin stillage were calculated based on factors found in the
literature [25].
4.1.2.6 Co-generation
The cogeneration system consisting of a boiler and a steam turbine was
simulated in VMGSim™ [68] using a Gibbs reactor for the biogas combustion, a
heat exchanger for the boiler and an expander for the steam turbine. The CH4
combustion is simulated as an adiabatic combustion at constant pressure. The
adiabatic reaction of biogas (mainly CH4) with the oxygen in air at constant
pressure generates high temperatures (>1000 °C), at which stable species at
102
ambient temperatures can dissociate. This is demonstrated in the equilibrium
reactions below [111]:
CO2 ↔ CO + ½ O2 Rxn. 7
H2O ↔ H2 + ½ O2 Rxn. 8
Assuming that chemical reactors achieve equilibrium is a reasonable first
approximation. For an adiabatic equilibrium combustion system, the VMGSim™
Gibbs reactor computes reactant and product concentrations so that the Gibbs
free energy of the system is minimized (recall that ∆G = 0 for equilibrium systems).
This obviates the need for reaction rate expressions and kinetics [112].
4.1.2.7 Heat exchangers
Heat exchangers were designed to be counter-current. The pressure drop
selected for the heat exchangers was 20 kPa for the shell and the tube side
according to recommended values [113].
4.1.2.8 Pumps
Pumps were simulated with pump efficiencies of 75% [114]. Discharge pressures
were set according to the process requirements accounting for pressure drops
resulting from heat exchangers, control valves, etc.
103
4.1.2.9 Control valves
Control valves were included in the simulation with an estimated pressure drop
of 150 kPa for liquid lines. For the vapour lines, 100 kPa was selected because of
the higher costs related to compress gases compared to pumping liquids [115,
116].
4.1.2.10 Utilities
Steam was provided in the simulation at 600 kPa. Cooling water was supplied at
15 °C and 800 kPa. The cooling water return was set mainly to 40 °C, never
exceeding 50 °C as recommended [113].
4.1.3 Pinch Analysis
Pinch analysis is a tool intended to assist design engineers in configuring a heat
exchange network which will minimize external heating and cooling
requirements.
Some of the concepts used in pinch analysis are presented below [117]:
i) Streams flow with constant composition and must be heated or
cooled.
ii) Cold streams need to be heated up.
iii) Hot streams need to be cooled down.
104
iv) Hot composite curve: Single curve representing all the hot streams.
v) Cold composite curve: Single curve representing all the cold streams.
vi) Cold and hot composite curves diagram: This is a temperature vs. heat
flow diagram in which the hot and cold composite curves are drawn
together. The maximum amount of heat recovery is represented by the
section of overlap between the composite curves. The minimum
amount of external cooling is represented by the remaining portion to
the left of the cold stream. The one to the right of the hot stream
represents the minimum amount of external heating as illustrated in Fig.
29:
Figure 29: Sample Composite Curves Diagram [118].
105
The composite curves are drawn in a temperature-enthalpy diagram.
When a process stream receives a differential heat flow dQ (measured
in kW), its enthalpy H increases by CP dT, where:
CP = heat capacity flow rate [kW/K]
dT = differential temperature change [K]
The heat capacity flow rate is obtained as the product of the mass
flow W and the specific heat Cp:
CP = WCp Eq. 4
Assuming a constant Cp, the total heat added to a cold stream from a
supply temperature Ts to a target temperature TT is equal to its
enthalpy change:
Q =� CPdT
TT
Ts
= CP(TT-TS) = ∆H Eq. 5
Therefore, when representing this equation graphically, the slope of the
line is:
106
dT
dQ=
1
CP Eq. 6
vii) Pinch Point: The closest approach between the hot and the cold
composite curves. It represents the minimum temperature difference
between the hot and cold streams (∆Tmin). For a set ∆Tmin, the hot and
cold utility predicted by pinch analysis are the minima required to solve
the heat recovery problem.
Pinch analysis was performed on the base-case simulation in order to assess the
possibility of process improvement and heat integration. A spreadsheet created
by Norwood [118] was used to plot the cold and hot composite curves. Inputs
include the heat duties from the different streams, the supply and target
temperatures, and ∆Tmin. The composite curves diagram, the hot and cold
pinch, and the minimum hot and cold utility could then be obtained.
Heat duties were taken from the VMGSim™ simulation for all unit operations
except the reactors. For the simultaneous saccharification and fermentation
reactor, the duty of the reactor cooler was calculated using the heat of
reaction. In the case of the anaerobic digester (AD), an estimated energy
requirement for the process obtained from the literature [119] was used to
compute the duty of the AD heater.
107
4.2 Process Description
The IDF ethanol plant can be divided into two main zones, the first operating in
batch mode and the second continuously. Zone 1 corresponds to the
preparation of the beer which generates feed for the next zone. Zone 2
comprises the equipment which produces fuel grade ethanol, animal feed, bio-
fertilizer, steam and electricity for the plant. The beer well marks the division
between the two.
4.2.1 Base Design Case
A process description for the different sections of the IDF ethanol plant will now
be presented along with the process flow diagram for the facility. The process
summarized below corresponds to the base-case design prior to optimization.
4.2.1.1 Base-Case Design Process Flow Diagram
The basis process flow diagrams (PFDs) for the IDF ethanol plant appear on the
following six pages.
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REVISIONS
Process Flow Diagram Sheet 1 of 6
MC-206
Meal Conveyor
TK-201
Feedstock Storage Tank
MT-202
Magnetic Trap
MI-203
Hammer Mill
TK-204
Meal Storage Tank
U-205
Hopper Scale
Pretreatment Water Heater/Stillage Cooler
E-207 TK-208 A/B/C/D
Pretreatment Tank
Hydrated Meal Pump
P-209 A/B/C/D
Hydrated Meal Cooler
E-211
Grain Steam Supply
Cooling Water Supply
Stillage toSP-246
Sheet 005Process Water
Pre-treatment Tank Side Heater
E-210 A/B/C/D
Hydrated meal to R-212 A/B/C/D
Sheet 002
TK-208 A
P-209 A
TK-201
U-205
MC-206
TK-204
Acid Base
E-207
E-210 A
E-211
TK-208 B
P-209 B
E-210 B
MT-202
MI-203
Acid Base
Stillage fromT-227
Sheet 004
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REVISIONS
Process Flow Diagram Sheet 2 of 6
R-212 A/B/C/DSSF
Reactor
Hydrated meal fromE-211
Sheet 001
Cooling Water Supply
R-212 A
R-212 B
TK-215
Beer to E-225
Sheet 004
CO2 to T-218
Sheet 003
Permeate from P-241
Sheet 005
E-214 A/B/C/DSSF Reactor
Cooler
P-213 A/B/C/DSSF Reactor
Pump
TK-215Beer Well
P-213 A
P-213 B
E-214 A
E-214 B
P-217
P-217Beer
Pump
AG-216
AG-216Beer Well Agitator
Yeast slurry fromP-222
Sheet 003
Enzyme from P-224
Sheet 003
Ethanol from P-219
Sheet 003
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REVISIONS
Process Flow Diagram Sheet 3 of 6
E-221
Conditioning Vessel Water Heater
CO2
T-218
CO2 fromR-212 A/B/C/D
Sheet 002
Process Water Supply
V-220
Yeast
P-222
Process Water
Nutrients
Steam Supply
Yeast to R-212 A/B/C/D
Sheet 002
V-223
P-224
Enzyme to R-212 A/B/C/D
Sheet 002
Ethanol to TK-215
Sheet 002
V-220
Yeast Conditioning Vessel
P-222
Yeast Slurry Pump
V-223
Enzyme Storage Vessel
P-224
Enzyme Pump
T-218
CO2 Absorber
P-219
P-219
CO2 AbsorberPump
E-221
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Process Flow Diagram Sheet 4 of 6
E-225
Distillation Column Preheater
T-226
Distillation Column
E-227
Reflux Condenser
Reflux Knock Out Drum
D-228
Cooling Water Supply
Reflux Pump
P-229
Beer fromP-217
Sheet 002
E-225
T-226
Steam Supply
P-229
D-228
Membrane feed to F-233
Sheet 005
Stillage to SP-242
Sheet 005
E-230
E-227 E-231
E-232
Stillage Cooler
E-232
Distillate Superheater
E-231
Distillation Column Reboiler
E-230
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REVISIONS
Process Flow Diagram Sheet 5 of 6
E-238
Permeate Condenser
F-233
Membrane Module Feed Filter
M-234
Dehydration Membrane Module
E-235
Retentate Condenser
E-236
Retentate Cooler
TK-237
Retentate Product Tank
Permeate Vessel
V-239 P-240
Vacuum Pump
Permeate Pump
P-241
Thin Stillage Tank
TK-243
Whole Stillage Screw Press
SP-242
M-234F-233
E-235
E-238
P-240
P-241
V-239
Membrane feed from E-231Sheet 004
Cooling Water Supply
SP-242
Permeate toTK-215
Sheet 002
TK-237 Fuel Grade Ethanol to Blending Facility
TK-243 TK-244
Thin Stillage to R-246
Sheet 006
E-236
P-245
Wet Grain Tank
TK-244Thin Stillage
Pump
P-245
Stillage fromE-232
Sheet 004
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R-246
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REVISIONS
Process Flow Diagram Sheet 6 of 6
TB-251
Steam Turbine
R-246
Anaerobic Digester
P-247
Anaerobic Digester Pump
E-248
Anaerobic Digester Heater
D-249
Biogas Dewatering Knock Out Drum
BO-250
Biogas Boiler Electricity Generator
G-252
Manure from Dairy Operation
Thin Stillage fromP-245
Sheet 005
To Digestate Separation
E-248
P-247
Steam SupplyD-249
Electricity
Utility Steam
Boiler Water
BO-250
TB-251 G-252
114
4.2.1.1.1 Feedstock Preparation
The wheat stored in tank TK-201 passes through a magnetic trap (MT-202) in
which any ferrous or metallic particles are captured. The clean wheat is ground
in a hammer mill (MI-203) and stored in tank TK-204. In a hopper scale (U-205),
the milled grain is weighed and the desired amount is sent through a meal
conveyor (MC-206) to the fermentation process.
4.2.1.1.2 Fermentation
The meal is hydrated in one of the four pre-treatment tanks (TK-208 A/B/C/D)
operated at 56 – 57 °C and acidic pH (between 3.5 and 4). The concentration
of solids in the tank is around 20 to 34 wt% [64]. The pH can be adjusted by the
addition of acid or base compounds. The water used in the pre-treatment tank
is heated in E-207 before being sent to the tank. In order to maintain the
temperature and agitation of TK-208, a side heater (E-210 A/B/C/D) is used. The
hydrated wheat is cooled to 32 °C (E-211) and pumped (P-209) to one of the
four simultaneous saccharification and fermentation (SSF) reactors (R-212
A/B/C/D), in which the enzyme and yeast are added. The liquid enzyme is
stored in vessel V-223 and pumped to the reactor by P-224. The active dry yeast
is conditioned by the addition of water and nutrients in vessel V-220. Each SSF
reactor is equipped with an external cooler (E-214 A/B/C/D) and pump (P-213
A/B/C/D). The fermentation reaction is exothermic, so the reaction mixture is
115
sent (P-213 A/B/C/D) through the external cooler (E-214 A/B/C/D) in order to
control temperature in the reactor. The recirculation of the cooled stream also
promotes mixing in the SSF reactor.
The CO2 offgas stream, containing around 1 vol% of ethanol, is sent to a
scrubber (T-218) to recover the ethanol which is then collected in the beer well
(TK-215) together with the beer obtained from the SSF reactor as liquid product.
The beer stored in the well contains around 10 ethanol by volume [66].
4.2.1.1.3 Distillation and Dehydration
The beer obtained from the upstream batch fermentation process is preheated
(E-225) and pumped (P-217) at around 3 bars absolute pressure (bara) to the
distillation tower (T-226). Heat energy is supplied by a steam reboiler (E-230). A
partial condenser (E-227) connected to the top of the distillation tower
produces the liquid reflux that is sent back to the tower and the vapour distillate
(80 wt% ethanol – 20% water) passes through a knock out drum (D-228) to be
superheated (E-231) (by 15 °C). This vapour flows to the membrane module (M-
234) which performs the final step in the dehydration process. It is first filtered (F-
233) to avoid clogging the membrane fibers with solids. Water separated from
the ethanol-water vapour feed is called the permeate stream. The permeate is
condensed (E-238) and collected in vessel V-239 before being pumped (P-241)
back to the beer well (TK-215) to recover any remaining ethanol. A vacuum
pump (P-240) creates the pressure differential that is the driving force for
116
separation in the membrane module. The retentate product, containing around
99 wt% ethanol, is condensed (E-235), cooled (E-236) and routed to the
anhydrous ethanol product tank (TK-237).
4.2.1.1.4 Anaerobic Digestion
The stillage obtained at the bottom of the distillation tower is cooled (E-232) and
passed through a screw press (SP-242) that separates the solids fraction (called
wet grains or wet cake) from the thin stillage liquor. The wet grains are then
stored in tank TK-244. The thin stillage is recovered in tank TK-243 and pumped
(P-245) to the anaerobic digester (R-246) operating at approx. 55 °C. The
digester is fed with manure from the cattle operation and the thin stillage. A
portion of the reaction mixture is circulated (P-247) through a heater (E-248) for
temperature control purposes. The digestate product is further treated to obtain
fertilizer to be used in the farm operation.
4.2.1.1.5 Co-Generation
The biogas produced in R-245 is sent to a steam boiler (BO-250) after passing
through a knock out drum (D-249) to eliminate any liquid droplets from the
biogas. The steam generated actuates a steam turbine (TB-251) which outputs
utility steam for the plant as well as enough electricity (via G-252) to satisfy the
power requirements of the ethanol facility.
117
4.2.2 Improved Process
As described above, pinch analysis was used as tool for minimizing energy
consumption by the IDF plant. After analyzing those results, it was decided to
focus on the distillation section because its energy demand is greatest. Different
configurations for the distillation were simulated to find the one which yielded
the lowest steam and cooling water requirements together with the most waste
heat recovery via heat integration.
The configuration used for the distillation section in the base design case of the
present work consisted of one complete distillation tower in which vapour
distillate was obtained as the vent gas from a partial condenser. Steam was
used in a reboiler to generate the energy required for the separation, as can be
seen in Fig. 30.
118
Beer fromSSF
Cooling Water Supply
Distillation Column Preheater
Reflux Pump
Distillation Column
Distillation Tower Reboiler
Reflux Condenser
Steam Supply
Distillate to Dehydration
Figure 30: Base design case distillation configuration - one column with steam preheat exchanger.
The first alternative to be considered was a combination of a stripper and a
conventional column operating at the same pressure. Stripper column A has a
steam reboiler. The stripper overhead, containing around 50 vol% ethanol, is sent
to distillation column B which includes a steam reboiler and a partial condenser
where the distillate is recovered as vapour (Fig. 31). All condensed liquid is
returned to the tower as reflux as was also true of the column in Fig. 30.
119
Beer fromSSF
Cooling Water Supply
Stripper Preheater Reflux Pump
Distiillation Column B
Rectifier Reboiler
Reflux Condenser
Steam Supply
Distillate to Dehydration
Stripper Column A
Stripper Reboiler
Figure 31: First configuration for process improvement analysis - two towers (stripper + distillation column).
The next scenario for the distillation section involved a combination of a stripper
and a complete distillation column operating at different pressures. Half of the
beer is fed to each column. The stripper column A is run at higher pressure (4.8
bara) with direct steam injection. Its overhead vapour becomes the heat
medium for the tower B reboiler in which it is condensed to generate boilup in
distillation column B. This condensate is fed back to the stripper. A concentrated
ethanol vapour from the partial condenser on distillation column B represents
120
the distillation system product stream. The second distillation tower operates at
2.6 bara.
Beer fromSSF
Cooling Water Supply
Stripper Column A Preheater
Reflux Pump
Distillation Column BDistillation
Column B Reboiler
Reflux Condenser
Steam Supply
Distillate to Dehydration
Stripper Column A
Distillation Column B Preheater
Figure 32: Second configuration for process improvement analysis - two towers at different pressures.
4.2.2.1 Improved Process Flow Diagram
The final configuration described above, with two towers operating at different
pressures, was found to be the best alternative for the distillation section. Revised
121
process flow diagrams of the IDF ethanol plant design with the improved
process are presented below.
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Process Flow Diagram Sheet 1 of 6
MC-206
Meal Conveyor
TK-201
Feedstock Storage Tank
MT-202
Magnetic Trap
MI-203
Hammer Mill
TK-204
Meal Storage Tank
U-205
Hopper Scale
Pretreatment Water Heater/Stillage Cooler
E-207 TK-208 A/B/C/D
Pretreatment Tank
Hydrated Meal Pump
P-209 A/B/C/D
Hydrated Meal Cooler
E-211
Grain Steam Supply
Cooling Water Supply
Stillage toSP-246
Sheet 005Process Water
Pre-treatment Tank Side Heater
E-210 A/B/C/D
Hydrated meal to R-212 A/B/C/D
Sheet 002
TC
TC
TC
FT
FT
FRC
LC
VSD
TK-208 A
P-209 A
TK-201
U-205
MC-206
TK-204
Acid Base
E-207
E-210 A
E-211
TK-208 B
P-209 B
E-210 B
MT-202
MI-203
AC
Acid Base
AC
TC
FT
FT
FRC
LC
Stillage fromT-227
Sheet 004
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Process Flow Diagram Sheet 2 of 6
R-212 A/B/C/DSSF
Reactor
Hydrated meal fromE-211
Sheet 001
TC
Cooling Water Supply
R-212 A
R-212 BTC
AIA
AIA
LC
LC
TK-215LC
Low level override
controller
LS
Beer to E-225 & E-226
Sheet 004
CO2 to T-218
Sheet 003
Permeate from P-245
Sheet 005
E-214 A/B/C/DSSF Reactor
Cooler
P-213 A/B/C/DSSF Reactor
Pump
TK-215Beer Well
P-213 A
P-213 B
E-214 A
E-214 B
P-217
P-217Beer
Pump
AG-216
AG-216Beer Well Agitator
FC
Yeast slurry fromP-222
Sheet 003
Enzyme from P-224
Sheet 003
Ethanol from P-219
Sheet 003
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REVISIONS
Process Flow Diagram Sheet 3 of 6
E-221
Conditioning Vessel Water Heater
CO2
T-218
CO2 fromR-212 A/B/C/D
Sheet 002
Process Water Supply
FT
FT
FRC
LC
V-220
Yeast
P-222
Process Water
Nutrients
Steam Supply
Yeast to R-212 A/B/C/D
Sheet 002
V-223
P-224
Enzyme to R-212 A/B/C/D
Sheet 002
Ethanol to TK-215
Sheet 002
V-220
Yeast Conditioning Vessel
P-222
Yeast Slurry Pump
V-223
Enzyme Storage Vessel
P-224
Enzyme Pump
T-218
CO2 Absorber
P-219
P-219
CO2 AbsorberPump
E-221
1
2
3
Name 1 2 3VapFrac 1 0 0T [C] 32 15 33.7P [kPa] 100 100 100Std Liq Volume Flow [m3/hr] 0.113 0.03 0.033Mass Flow [kg/h] 92.13 30 32.12MassFraction [Fraction] METHANE 0 0 0 ETHANOL 0.0281 0 0.0616 WATER 0.0193 1 0.9374 CARBON DIOXIDE 0.9527 0 0.001
LI
LI
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Process Flow Diagram Sheet 4 of 6
D-230
Stripper Distillate Flashing Drum
E-225
Stripper Column A Preheater
E-226
Distillation Column B Preheater
T-227
Stripper Column A
T-228
Distillation Column B
E-229
Distillation Column B Reboiler
Stripper Distillate Pump
P-231 E-232
Reflux Condenser
Reflux Knock Out Drum
D-233
Distillate Superheater
E-235
Cooling Water Supply
Stillage toE-207
Sheet 001
Reflux Pump
P-234
Beer fromP-217
Sheet 002
T-227
E-225
FI
T-228
D-230
PC
LC
FRC
FC
FC
Steam Supply
TC
FCFRC
TC
P-234
D-233 LC
FC
PCTC
Membrane feed to F-237
Sheet 005
LC
Stillage to SP-246
Sheet 005
E-226
E-229
P-231
E-232 E-235
LC
4
5 6
7 8
9
10
11
12
13
14
15
161718
19
21
22
23 24
25
26
27
20
Level Signal to40
Sheet 005
E-236
Stilage from Distillation Column B Cooler
E-236
FT
Name 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27VapFrac 0 0 0 0 0 1 0 1 0 1 1 0.22009 0 0 0 0 0 1 1 1 1 0 0 0T [C] 32.6 32.6 141.9 32.6 120.5 158.9 150.4 158.9 146.3 157.6 133.8 118.2 118.2 118.3 150.2 128.2 89 118.2 118.2 108.3 113.3 138.8 15 50P [kPa] 650 500 480 280 260 500 480 450 430 480 480 360 360 480 480 260 235 360 260 260 260 410 650 630Std Liq Volume Flow [m3/hr] 1.093 0.546 0.546 0.546 0.546 0.118 0.118 0.092 0.092 0.16 0.304 0.304 0.225 0.225 0.628 0.468 0.468 0.079 0.079 0.078 0.157 0.092 1.282 1.282Mass Flow [kg/h] 1064.48 532.24 532.24 532.24 532.24 118.05 118.05 91.5 91.5 160 259.5 259.5 193.9 193.9 626.65 467.31 467.31 65.6 65.6 64.93 130.54 91.5 1280.58 1280.58MassFraction [Fraction] METHANE 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 ETHANOL 0.0995 0.0995 0.0995 0.0995 0.0995 0 0 0 0 0 0.6611 0.6611 0.6181 0.6181 0.0021 0.005 0.005 0.7881 0.7881 0.78 0.784 0 0 0 WATER 0.9005 0.9005 0.9005 0.9005 0.9005 1 1 1 1 1 0.3389 0.3389 0.3819 0.3819 0.9979 0.995 0.995 0.2119 0.2119 0.22 0.216 1 1 1
CARBON DIOXIDE 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0
PC
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REVISIONS
Process Flow Diagram Sheet 5 of 6
E-242
Permeate Condenser
LI
F-237
Membrane Module Feed Filter
M-238
Dehydration Membrane Module
E-239
Retentate Condenser
E-240
Retentate Cooler
TK-241
Retentate Product Tank
Permeate Vessel
V-243 P-244
Vacuum Pump
Permeate Pump
P-245
Thin Stillage Tank
TK-247
Whole Stillage Screw Press
SP-246
M-238F-237
E-239
E-242
P-244
P-245
V-243
Membrane feed from E-235Sheet 004
Cooling Water Supply
TC
LC
S
TC TC
SP-246
Permeate toTK-215
Sheet 002
TK-241 Fuel Grade Ethanol to Blending Facility
Stillage fromE-236
Sheet 004
TK-247 TK-248
Thin Stillage to R-250
Sheet 006
LC
E-240
P-249
Wet Grain Tank
TK-248Thin Stillage
Pump
P-249
28
29
30
31 32 33
3435
36
37
38
39
40
42
43Stillage fromE-207
Sheet 001
Level Signal fromT-227
Sheet 004
41
Name 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43
VapFrac 1 1 0 1 0 0 0 0 0 0 0 0 0 0 0 0T [C] 128.3 128.3 39.9 128.3 82.8 40 15 40 15 40 15 40 62.1 73.8 73.8 73.8P [kPa] 240 15 15 140 120 100 650 630 650 630 650 630 310 235 235 235Std Liq Volume Flow [m3/hr] 0.157 0.045 0.045 0.112 0.112 0.112 0.887 0.887 0.136 0.136 0.826 0.826 0.628 1.096 0.626 0.47Mass Flow [kg/h] 130.54 41.51 41.51 89.02 89.02 89.02 886.5 886.5 136.31 136.31 825.19 825.19 626.65 1094 625.12 468.84MassFraction [Fraction] METHANE 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 ETHANOL 0.784 0.3277 0.3277 0.9969 0.9969 0.9969 0 0 0 0 0 0 0.0021 0.0033 0.0033 0.0033 WATER 0.216 0.6723 0.6723 0.0031 0.0031 0.0031 1 1 1 1 1 1 0.9979 0.9967 0.9967 0.9967 CARBON DIOXIDE 0 4E-100 4E-100 0 0 0 0 0 0 0 0 0 0 0 0 0
LI
PC
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R-250
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Process Flow Diagram Sheet 6 of 6
TB-255
Steam Turbine
R-250
Anaerobic Digester
P-251
Anaerobic Digester Pump
E-252
Anaerobic Digester Heater
D-253
Biogas Dewatering Knock Out Drum
BO-254
Biogas Boiler Electricity Generator
G-256
Manure from Dairy Operation
Thin Stillage fromP-249
Sheet 005
To Digestate Separation
E-252
P-251
LC
Steam SupplyD-253
FC
PC
Electricity
Utility Steam
Boiler Water
BO-254
TB-255 G-256
44
49
50
51
TC
45
46
47
48
Name 44 45 46 47 48 49 50 51VapFrac 0 0 1 0 1 1 1 0T [C] 44.1 55 158.9 50 55 365 162.9 99.6P [kPa] 100 80 500 480 100 5880 600 100Std Liq Volume Flow [m3/hr] 0.357 0.357 0.006 0.006 0.093 0.252 0.252 0.252Mass Flow [kg/h] 356.83 356.83 6.09 6.09 46.91 251.34 251.34 251.34MassFraction [Fraction] METHANE 0 0 0 0 0.3553 0 0 0 ETHANOL 0.0011 0.0011 0 0 0 0 0 0 WATER 0.9989 0.9989 1 1 0 1 1 1 CARBON DIOXIDE 0 0 0 0 0.6447 0 0 0
128
4.3 Equipment Sizing
For the most part, the IDF plant equipment was sized using short-cut techniques
consistent with the level of detail one would expect to find in a design basis
memorandum. Sample calculations can be found in the Appendices.
4.3.1 Milling
For the sizing of the hammer milled, the angular velocity or hammer tip speeds
was calculated from [120]:
v = 1
60 πds
where:
v = hammer tip speed [m/s]
d = hammer tip arc diameter [m]
s = rotor speed [rpm]
A hammer tip arc between 80 and 120 m/s was recommended in [120].
4.3.2 Saccharification and Fermentation Set
Each simultaneous saccharification and fermentation (SSF) unit, consisting of a
pre-treatment tank (TK-208) and a SSF reactor (R-212), was sized based on the
required volume of beer to be fed to the continuous zone of the plant.
The four batch processes must be scheduled so as to have feed in the beer well
at all times to feed the distillation/dehydration section of the plant which
Eq. 7
129
operates continuously. The amount of beer produced in each batch equals that
required for 24 hours of operation in the continuous section of the plant. The
beer feedrate to the distillation column is 25.4 m3/d.
The fermenters and pre-treatment vessels were designed as conical tanks as
shown in the Fig. 33, with the conical portion corresponding to 1/3 of the total
height. An aspect ratio of 1.5 was employed when sizing the fermenter.
Figure 33: SSF conical reactor [121].
4.3.3 CO2 Absorber
The number of trays in the absorber was chosen so that 90% of the feed ethanol
would exit in the bottom of the absorber. To determine its length, Luyben’s
130
method was used, in which the tray spacing is fixed at 2 ft (0.61 m). The height
of the tower is inflated by 20% to allow for liquid disengagement at the top and
adequate surge volume in the column sump [116]:
Lc = 0.7315 (NT)theoretical / E0 Eq. 8
Where:
Lc = Absorber Height [m]
(NT)theoretical = Theoretical number of trays
E0 = Overall column efficiency
The tray efficiency selected was 70% according to recommended values [122].
The diameter of the absorber was calculated using the F-factor approach
which sets the column diameter based on the maximum vapour velocity. The
value considered for the F-factor was 1.22 as recommended by Luyben. The
following equations were used in the calculations [116]:
F-factor = vmax�ρv Eq. 9
Area = maximum vapor rate
maximum allowable velocity
Eq. 10
A =πD2
4
Eq. 11
131
where:
F-factor = 1.22
vmax = maximum allowable vapour velocity [m/s]
ρv = vapour density [m/s]
A = column cross sectional area [m2]
D = column diameter [m]
4.3.4 Distillation Section
As previously mentioned, the number of theoretical trays, NT, for the distillation
tower was obtained using McCabe Thiele graphical method. The number of
trays in the stripper was obtained through optimization of the steam injection
requirements as explained in Chapter 5. Once the number of theoretical trays
was obtained, the length of the distillation tower was calculated using Luyben’s
method, as outlined in Section 4.3.3.
Luyben’s F-factor method was also used to size the flash drum D-230 which
separates the liquid and vapor condensate of E-229 using a recommended F-
factor of 0.61 [116]. The drum was taken to be a vertical cylindrical vessel with a
length-to-diameter ratio of two. The selected liquid holdup time in the vessel was
five minutes and a liquid level of 50% was assumed. The reflux drum D-233 was
sized in similar fashion.
132
4.3.5 Dehydration
The parameters used for sizing the vapor permeation membrane module (M-
238) are listed in Table 20 below.
Table 20: Design specifications for the vapor permeation membrane module.
Parameter Value Units Membrane performance 2.08 L/h.m2 Plant capacity 112 L ethanol/h Operating pressure 2.40 bara Operating temperature 128.30 °C Feed ethanol concentration 0.78 EtOH wt% Feed flow rate 130.54 kg/h
The membrane performance and the required product flow were used to
estimate the membrane area and to select the membrane module according
to the supplier’s recommendations [123]. The other parameters summarized in
the table above were entered into a spreadsheet tool provided by the supplier,
Whitefox Technologies Canada [110], to verify that the calculated membrane
area would be sufficient to satisfy process operating specifications. The
permeate vessel (V-243) was sized in the same way as the reflux accumulator
(D-233).
The vacuum pump was sized according to the vacuum pressure required and
the air flow capacity. The air flow capacity was estimated considering an
evacuating time of 5 min and the volume of the lines connected to the vacuum
pump. This volume is the sum of the dehydration membrane module (M-238)
housing and the piping including the permeate condenser (E-242). The volume
133
of the membrane housing was estimated with information provided by the
membrane technology provider [123]. The design housing volume was inflated
by 20% to account for piping.
4.3.6 Anaerobic Digestion
The anaerobic digester (R-250) was sized according to its biogas production
capacity. Twelve days of hydraulic retention time were assumed for the well-
mixed anaerobic digester incorporated in this conceptual design. With this
residence time, a theoretical volume for the reactor was calculated as follows:
V = HRT V̇ Eq. 12
where:
V = reactor volume [m3]
HRT = hydraulic retention time [h]
V̇ = volumetric flow [m3/h]
The characteristics of the anaerobic digester (AD) were used to obtain the
actual volume according to an anaerobic digestion benchmark study
performed on twelve farm sites in British Columbia [75]. The AD employed in this
study is compared with that proposed in this thesis in Table 21.
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Table 21: Comparison of the characteristics of the anaerobic digester of the IDF plant and an AD reported in the literature.
Characteristics Units IDF BC Benchmark Study Production of biogas m3/h 47 49 Manure supplied m3/h 1.77 0.59 Dairy cattle 160 200 AD volume m3 ~370 500
4.3.7 Co-Generation System
The co-generation system consists of a biogas-fired steam boiler (BO-254), a
steam turbine (TB-255) and a generator (G-256). The boiler was sized base upon
the heat output or absorbed heat in the steam-water circuit at the specified
operating conditions [124]. The steam turbine and the generator were sized
according to the electricity obtained from the steam generated in the biogas-
fired boiler. Values of the heat output and the electricity in the co-generation
system were obtained from the VMGSim™ simulation of the IDF plant.
The biogas dewatering knock out drum (D-253) was sized in a similar way as the
reflux accumulator (D-233) assuming a water content in the biogas of 3 wt%
[125].
4.3.8 Vessels and Tanks
The residence time was used for sizing the tanks and vessels, as follows:
V = τ V̇ Eq. 13
where:
V = vessel volume [m3]
135
τ = residence time [h]
V̇ = volumetric flow [m3/h]
The residence times were selected according to the process requirements and
batch schedule. The vessels were designed as vertical tanks with cylindrical or
conical shapes as appropriate. For the pieces of equipment in which solids are
present, conical vessels were selected over cylinders. The dimensions of the
conical tanks were calculated based on the relation of the heights of the
cylindrical and conical portion of the tanks and the aspect ratio (length to
diameter) as shown in Fig. 34.
H = 2/3 L
h = 1/3 L
L
D
Figure 34: Conical vessel schematic.
136
The parameters used for the sizing of the different tanks and vessels are
summarized in the table below:
Table 22: Sizing parameters for tanks and surge vessels.
Equipment Description Residence Time
Aspect ratio
Additional information
required
TK-204 Meal Storage Tank
48 hours H = 2/3 L h = 1/3 L H = 1.5 D
Bulk density: 529 kg/m3 [126] Conical shape
TK-215 Beer Well 30 hours H = 2/3 L h = 1/3 L H = 1.5 D
Conical shape
V-220 Yeast conditioning vessel
24 hours H = 2/3 L h = 1/3 L H = 1.5 D
Conical shape
V-223 Enzyme Storage Vessel
168 hours L = 1.5 D Cylindrical shape
TK-241 Retentate Product Tank
168 hours L = 1.5 D Cylindrical shape
TK-247 Thin Stillage Tank 8 hours L = 1.5 D Cylindrical shape
TK-248 Wet Grain Tank 168 hours L = 1.5 D Cylindrical shape
4.3.9 Heat Exchangers
To obtain the heat duty in a shell and tube heat exchanger, the following
formula is used:
q = UA∆TLM Eq. 14
where:
137
q = heat duty [W]
U = heat transfer coefficient [W/m2 °C]
A = heat transfer area [m2]
∆TLM = log mean temperature difference [°C]
The log mean temperature difference is calculated as follows:
∆TLM=∆T1-∆T2
ln (∆T1/∆T2) Eq. 15
with ∆T1 and ∆T2 defined as the temperature differences at the two ends of the
heat exchanger [122].
Sizing of the heat exchangers was performed by finding the heat transfer area.
The heat load and the UA values were obtained from the simulation model. The
overall heat transfer coefficients were obtained from the literature considering
the fluids circulating in the heat exchangers [127].
4.3.10 Pumps
The pumps were sized according to the total head and the volumetric flow. The
total head was calculated as follows:
Total Head = Static Head + Friction Loss Head Eq. 16
The static head refers to the difference between the liquid level in the suction
tank and the point of discharge as shown in Fig. 35 [128]:
138
Figure 35: Schematic of static head for pump calculations [128].
The static head was estimated according to the tank height. If the liquid was
pumped from a tank at a pressure higher than atmospheric, that pressure was
converted to units of head and added to the static head.
To obtain the total head in feet, the following equation was used:
Head = P x 2.31
SG Eq. 17
where:
P = Pressure [psi]
SG = Specific gravity
Head = total head [ft]
The friction losses caused by the piping and fittings were considered to be 25%
of the static head [128]. The pressure drops caused by control valves and heat
139
exchangers were considered according to the values previously presented in
Section 4.1.2.
The calculated total head and the volumetric flows obtained from the
simulation were used to select the proper pumps. Performance or characteristic
curves provided by pump suppliers were used to select the pump and for
costing purposes.
4.3.11 Miscellaneous
Other pieces of equipment in the IDF plant were sized according to information
obtained from potential suppliers of the equipment or scaling variables that are
used in an exponential approach for costing the equipment. Table 23 presents
the specifications considered for the different pieces of equipment.
140
Table 23: Equipment specifications for different pieces of equipment of the IDF plant.
Equipment Description Specification MT-202 Magnetic Trap Included with the hammer mill U-205
Hopper Scale Capacity: 500 kg / h, 0.95 m3 / h Residence time: 6 hours Material: Wheat meal
MC-206
Meal Conveyor
Type: Bucket elevator conveyor Capacity: 4,306 kg/h (based on a filling time of 2 hours) Conveyor elevation: 7 m
SP-246
Screw Press
Capacity: 1,094 kg / h, 1.12 m3 / h Material: Whole stillage Separation required: 43% liquid fraction 57% solid fraction
AG-216 Beer Well Agitator Scaling variable: 34.09 m3 (tank volume) F-237 Membrane
Module Feed Filter Scaling variable: 130.54 kg/h (inlet mass flow)
BO-254 Biogas Boiler Heat output: 250 HP Pressure rating: 60 bara
TB-255 Steam Turbine
28 KW Inlet pressure: 58 bara Inlet temperature: 365 °C
G-256 Electricity Generator 28 KW
4.4 Economic Evaluation
4.4.1 Equipment Costing
For costing of the equipment an exponential scaling method was taken from
Humbird et al. [129] The following equation was utilized for this purpose:
New scaled purchase cost = Purchase cost × �New scaling value
Scaling value �n
Eq. 18
141
where:
New scaled purchase cost = Calculated cost of the piece of equipment to be
used in assessing the economic feasibility of the proposed IDF plant.
Purchase cost = Cost of the equipment obtained from the literature [127].
New scaling value = The value of the scaling variable selected according to the
working principles of the IDF equipment. Sample sizing calculations can be
found in the Appendices.
Scaling value = The value for the scaling variable for the equipment in question
obtained from literature sources [127].
n = The characteristic scaling exponent. This value is typically in the range of 0.6
to 0.7. The values used here are listed in Table 24 [129]:
Table 24: Scaling exponents employed in equipment costing.
Equipment Exponent Mills 0.6 Agitators 0.5 Distillation Columns 0.6 Heat exchangers 0.7 Pumps 0.8 Tanks, atmospheric 0.7 Solids handling equipment 0.8
When the purchase cost was not for the year of the analysis (2013), to account
for changing economic conditions (inflation), the following expression was used
[37]:
142
C2 = C1 × �I2I1� Eq. 19
where:
C = Purchased cost
I = Cost index = CEPCI (Chemical Engineering Plant Cost Index)
The subscripts: 1 = base time when cost is known
2 = time when cost is desired (2013)
Table 25 presents the CEPCI values [130] used for costing equipment in this
thesis.
Table 25: Chemical Engineering Plant Cost Index
Year CEPCI 1999 390.6 2000 394.3 2001 394.3 2002 395.6 2003 402 2004 444.2 2005 468.2 2006 499.6 2007 525.4 2008 575.4 2009 521.9 2010 550.8 2011 585.7 2012 584.6 2013 638.5
All equipment costs were adjusted to 2013 Canadian dollars (CD).
143
4.5 Capital Investment
After costing the equipment, the total capital investment was estimated by
assigning percentages to additional expenses not considered above [131]:
Table 26: Percentages for determining the capital investment
Category Percentage Equipment 54.7 Piping 10.9 Structural 6.3 Insulation 2.9 Electrical 3.0 Instrumentation 3.1 Automation 3.1 Engineering 16.0
4.5.1 Economic Model
In order to perform the economic evaluation, a model was developed to assess
the economic impact of implementing an on-farm IDF ethanol plant. Some of
the data used in this analysis was obtained from a representative of the southern
Alberta farming community [132].
The following assumptions were made in constructing the economic model:
a) Based on farm community information
- The analysis is pertinent to crop, dairy cattle and beef cattle farm
operations.
- Data used is from 2012 production year.
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- Total land under combine is 4,500 acres.
- 28% of the land is used for wheat production.
- Two types of feed are considered: CWRS #1 with 13% protein content
(34% of wheat production) and CW Feed (66% of wheat production).
- Other crops grown on the land: CWRS red spring, CW Feed, Feed Barley,
Canola, Hay, Straw.
- The total number of dairy cattle is 160 and 90 are producing milk with a
lactation period of 255 days per year.
- The total number of beef cattle is 450.
- Wheat-land yield is 90 bushels/acre.
b) Regarding ethanol production
- 100% of the wheat available is used for ethanol production
- The sales product is anhydrous fuel grade ethanol
- The co-products of the plant are used within the farm. Only the excess
wet grain is sold as animal feed.
- Ethanol price assumed: 0.634 CD / L (ethanol price December 2013 [133])
- Biogas is obtained through anaerobic digestion using all the dairy manure
available and all the thin stillage produced in the ethanol plant
- The digestate from the anaerobic digester is the sum of the solids in the
manure and the thin stillage liquor.
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- The nutrient content (N, K, P) of the fertilizer obtained from the digestate is
considered as the sum of the nitrogen, potassium and phosphorus in the
manure and the thin stillage.
- 40% of the animal feed is replaced by the wet grains (wet cake)
generated by the ethanol plant [25].
Under these assumptions, a factor approach was adopted in developing the
model. Information and factors were gathered from literature publications,
governmental institutions (Alberta Agriculture and Rural Development - AARD,
Statistics Canada, Agriculture and Agri-food Canada - AAFC, United States
Department of Agriculture - USDA, Agriculture Financial Services Corporation –
AFSC, Alberta Energy, Alberta Utilities Commission – AUC, US Department of
Energy – DOE) and the private sector. Table 27 displays the values of these
parameters.
146
Table 27: Factors used in the economic model
Process Factors Description Factor Units
Wheat land yield [132]: 90.00 bushels / acre Solids percentage in cattle manure [25] 0.12 tonne solids / tonne manure Biogas produced by dairy manure [134] 32.00 m3 biogas / tonne manure Biogas produced by thin stillage [134] 58.00 m3 biogas / tonne thin stillage Percentage of thin stillage in stillage [25] 0.25 kg thin stillage / kg stillage Dissolved and suspended solids in thin stillage [25] 0.07 kg solids / kg thin stillage Wet grain that can be used for cattle's feed 0.40 kg wet grains / kg cattle feed N for wheat production [135] 111.00 kg / ha P2O5 for wheat production [135] 39.00 kg / ha K2O for wheat production [135] 85.00 kg / ha Nitrogen content in thin stillage [136] 26.10 g/kg Phosphorus - P content in thin stillage [136] 12.10 g/kg Potassium - K content in thin stillage [136] 1.60 mg / kg Solids in wet grain or thick stillage [25] 0.35 kg solids / kg wet grain Make-up water requirement to overcome evaporation in cooling towers [137] 1.20 % of CW Make-up water requirement to overcome drift losses in cooling towers [137] 0.13 % of CW Make-up water requirement to overcome blowdown in cooling towers [137] 0.75 % of CW Wheat seeding rate [132] 15.90 kg / tonne of wheat Animal Feed [132] 10.00 lbs / day / animal Manure produced by dairy cattle [132] 90,72 kg manure / y / cow Nitrogen Content in Manure [138] 0.0050 lb N / lb manure Phosphorus Content in Manure [138] 0.00075 lb P / lb manure Potassium Content in Manure [138] 0.0028 lb K / lb manure
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Table 27: Factors used in the economic model …(con)
Process Factors … (con) Description Factor Units
Hay requirement for animal feed [132] 458.00 lbs / animal / year Hay requirement for beef cattle feed [132] 864.00 lbs / animal / year Solids in whole stillage [25] 0.12 kg solids / kg stillage
Market Factors Description Factor Units
No. 1 CW Red Spring Wheat Price (13%) [139] 6.91 CD/bushel Fresh Water Cost Alberta [140] 1.58 CD / m3 Average cooling tower water cost [99, 141] 7.14E-05 USD / L H2O Wheat seed cost [142] 0.54 CD / kg seed Average enzyme cost [99, 141, 143-145] 0.015 USD / L EtOH Average yeast cost [99, 141, 143-145] 0.0042 USD / L EtOH Average denaturant cost [99, 141] 0.0074 USD / L EtOH Average chemical cost [141, 144, 145] 0.0037 USD / L EtOH Wet grain price [146] 65.40 CD/ton Barley price [139] 5.50 CD/bushel Canola #2 Price [139] 13.0 CD/ton Hay Price [139] 0.050 CD/lb Straw Price [139] 0.029 CD/lb N fertilizer cost [142] 0.60 CD / lbs N P2O5 fertilizer cost [142] 0.63 CD / lbs P2O5 K2O fertilizer cost [142] 0.54 CD / lbs K2O Ethanol price [133] 2.40 CD / gal Industrial electricity cost [147] 0.093 CD cents / KWh Industrial steam cost [148] 0.0078 CD / kg steam Operator salary 90,000 CD / year
148
149
The economic model was used to create an income statement in the form of
Table 28.
Table 28: Income statement format.
INCOME STATEMENT Before IDF Ethanol Plant After IDF Ethanol Plant Monetary Value % Monetary Value % Sales Revenue Costs of Goods Sold GROSS PROFIT General Expenses Utility Expenses Labour Expenses Operating Expenses EBITDA (Earnings before interest, taxes, depreciation and amortization) Capital Investment Interest rate Financial expenses (interests) EARNINGS BEFORE TAXES Tax rate Taxes NET PROFIT
.
The sales revenue includes:
- Crop sales
- Dairy cattle and products sales
- Beef cattle sales
150
- Anhydrous ethanol sales
- Wet distillers’ grain (animal feed) sales
The costs of goods sold comprise direct material and labor expenses incurred
in making the product, including:
- Chemical
- Fertilizer
- Seeds
- Livestock purchases (beef cattle operation + dairy cattle operation)
- Supplies (beef cattle operation + dairy cattle operation)
- Veterinarian (beef cattle operation + dairy cattle operation)
- Hay (beef cattle operation + dairy cattle operation)
- Animal feed (beef cattle operation + dairy cattle operation)
- Process water
- Enzyme
- Yeast
The operating expenses include general expenses to run the farm ethanol
plant and the utility costs, including:
- Labour
- Cooling water
- Make-up water
- Custom work (crop operation + beef cattle operation + dairy cattle
operation)
151
- Business expenses
Considering that the farmer will need a loan to cover the initial investment, the
interest rate used was 5% according to Agriculture Financial Services
Corporation AFSC [149]. The tax rate considered for the project was 15% [150].
Income statements were prepared to compare the farmer’s financial position
before and after the installation of the IDF ethanol plant.
4.5.2 Economic Indicators
The economic indicators used for evaluating the results obtained from the
model are:
- EBITDA (Earnings Before Interest, Taxes, Depreciation and Amortization)
- ROI (Net Return on Investment)
- ROS (Return on Sales)
- Payback Period
4.5.2.1 EBITDA
The EBITDA is calculated as follows:
EBITDA = Gross Profit - Operating Expenses Eq. 20
EBITDA is essentially net income with interest, taxes, depreciation, and
amortization added back to it. This metric is useful in comparing profitability
between companies and industries because the effects of financing and
152
accounting decisions are eliminated [151]. Higher EBITDA values imply better
profitability.
4.5.2.2 ROI
The ROI can be calculated from:
ROI = Net Profit
Capital Investment Eq. 21
ROI is a performance measure that is used to evaluate the efficiency of an
investment or to compare the efficiency of different investments [152]. The
greater the ROI value the higher the profitability.
4.5.2.3 ROS
The ROS is given by:
ROS = Net Profit
Sales Revenue Eq. 22
The ROS indicates how much profit is being produced per dollar of sales. Higher
ROS values are associated with higher profitability. A business having a ROS of
30% is generally considered profitable [152].
4.5.2.4 Payback Period
The payback period is the number of years it takes before the cumulative cash
flow equals the initial capital investment. The payback period was computed
using a ten-year project lifetime.
153
4.5.3 Sensitivity Analysis
Sensitivity analysis is a technique used to assess the impact of different
independent parameters on a specific dependent variable under a given set of
assumptions. Independent variables one might consider for the sensitivity
analysis are factors such as ethanol price, wheat price and fertilizer price while
the dependent variables would indicate project profitability (EBITDA, ROI, etc.).
154
5. Process Design
5.1. Base Design Case Simulation Work
A description of the simulation output generated for various sections of the
base-case IDF plant design is presented below. These results were obtained
before process improvement was performed. The results after performing pinch
analysis and heat integration will be shown in the next section.
5.1.1. SSF Reactor
The ethanol concentration of the beer produced via fermentation in the SSF
reactor was computed using the mathematical model described in Chapter 3.
The initial conditions are presented in Table 29.
Table 29: Initial conditions used for solving differential equations describing the SSF reactor.
Glucose initial concentration 197.21 g/L Biomass initial concentration 0.55 g/L Residence time 76 hours
The stoichiometric relationship proposed by Lantz and Li [100] was used to
calculate the initial glucose concentration based on the starch concentration
present in the wheat feedstock. A starch content of 60 wt% for wheat was taken
from the literature [62, 66]. The biomass initial concentration was based on
recommended doses of commercial yeast [107]. The reaction time used in the
model was based on vendor information applicable to simultaneous
155
saccharification and fermentation performed using Stargen enzyme [64]. The
simulation results are displayed in Fig. 36 and Table 30.
Figure 36: Concentration profile for ethanol production in the SSF reactor.
Table 30: Final results obtained from the SSF model.
Variable Initial Value Final Value Ethanol 0 g/L 86.63 g/L Glucose 197.21 g/L 75.84 g/L Biomass 0.55 g/L 16.80 g/L
The final ethanol concentration was approximately 9 wt%. This result was used in
the steady-state VMGSim™ simulation when modelling the batch SSF reactor as
a conversion reactor. Conversions specified for the saccharification and
fermentation stages of the reactor were adjusted until the correct beer
concentration of 9 wt% (10 vol%) was obtained (see Table 31).
156
Table 31: Conversions used in the VMGSim™ model of the SSF reactor.
Reaction Conversion enzyme Starch + 800 H2O 800 glucose 95%
yeast C6H12O6 2 C2H5OH + 2 CO2 80%
These numbers were found to be consistent with literature values [100].
5.1.2. CO2 Scrubber
The scrubber was optimized by varying the number of stages and the scrubbing
water flow. The target ethanol recovery of 90% was attained with 4 stages and a
water flow rate of 30 kg/h.
5.1.3. Distillation
The distillation system was designed to operate at 2.6 bara producing a vapour
distillate of 80 ethanol wt% from a saturated liquid feed of ethanol purity 8.6 wt%.
The design parameters and results of the McCabe-Thiele analysis are provided in
Table 32.
The reflux ratio of 1.2 times the minimum was specified according to Douglas
[153]; that is, using the rule of thumb that the best trade-off between capital
and operating expenses occurs when columns are designed for a reflux ration
between 1.2 and 1.5 times the minimum value. The feed to the distillation tower
is passed through a heat exchanger which preheats the beer to its bubble point.
The relevant McCabe-Thiele diagrams can be found in Appendix 1.
157
Table 32: Design parameters and results obtained using the McCabe-Thiele method base-case design of the distillation column.
Design Parameter Symbol Value Units Feed to the column F 54.00 kmol/h Liquid Feed Flow LF 54.00 kmol/h
Feed concentration zF 0.037 Mole fraction 0.086 Mass fraction
Tower bottoms composition, ethanol concentration xB
0.0021 Mole fraction 0.0050 Mass fraction
Distillate composition, ethanol concentration xD
0.62 Mole fraction 0.80 Mass fraction
Reflux ratio factor 1.20 R=1.2 Rmin q-line LF/F 1 Rectifying section slope 0.65 Minimum reflux ratio Rmin 1.87 Calculated reflux ratio R 2.24 Minimum number of stages 4 Theoretical stages 9 Feed stage 4
The number of stages determined using McCabe Thiele graphic method was 9.
After developing the graphic method, VMGSim™ was used to define the
optimal number of stages as shown in Figs. 37 and 38. The number of stages was
varied and the reflux ratio and reboiler duty were obtained from the simulation.
The optimal number of stages was determined as the stage prior to the
minimum reflux ratio and the minimum reboiler duty.
158
Figure 37: Optimization of number of stages for the distillation tower based on the reflux ratio rule-of-thumb.
Figure 38: Optimization of number of stages for the distillation tower based on minimizing reboiler duty.
It was found that the optimal number of ideal equilibrium stages was nine in
both figures. Stage 4 was found to be the best stage at which to introduce the
feed.
0.5
1.5
2.5
3.5
4.5
5.5
6.5
7.5
8.5
4 6 8 10 12 14 16 18 20 22
Reflu
x Ra
tio
Number of Stages
Minimum reflux ratio: 1.1
Optimal reflux ratio: 1.4Number of Stages: 9
1.5E+05
2.0E+05
2.5E+05
3.0E+05
3.5E+05
4.0E+05
4 6 8 10 12 14 16 18 20 22
Rebo
iler D
uty
(W)
Number of Stages
Minimum reboiler duty: 1.6 x 105 W
Optimum reboiler duty: 1.7 x 105 W
159
5.1.4. Dehydration
The membrane dehydration module was simulated as a component splitter in
VMGSim™. The splits were calculated using the Whitefox Technologies
membrane analysis spreadsheet [110] with the inputs as given in Table 33.
Table 33: Values of the variables input to the WFX membrane analysis tool.
Parameter Value Membrane module feed flow 138.30 kg/h Membrane module feed temperature 122.06 °C Membrane module feed pressure 2.4 bara Ethanol composition in the feed to the membrane module 80 wt% Vacuum 150 mbar
The results (Table 34) obtained from this program were used to set the splits in
the membrane simulation model.
Table 34: WFX membrane analysis tool outputs for calculation of the splits for membrane dehydration module simulation
Parameter Value Membrane module permeate flow 41.11 kg/h Membrane module permeate ethanol composition 33.53 wt% Membrane module retentate flow 97.19 kg/h Membrane module retentate ethanol composition 99.66 wt%
5.1.5. Anaerobic Digestion
In order to estimate the amount of wheat grain and thin stillage produced as
well as the biogas obtained from the anaerobic digestion process, several
factors obtained from the literature were utilized (see Table 35).
160
Table 35: Factors used for the anaerobic digester modelling.
Factor Value Solids in whole stillage 12 wt % [25] Solids in thin stillage 7 wt % [25] Ratio of solids in thin stillage and solids in whole stillage
0.25 [25]
Manure generated by dairy cattle 90,718 kg manure / y / cow Number of dairy cattle 160 CH4 content in biogas 60 vol% [134] CO2 content in biogas 40 vol% [134] Biogas potential of manure 32 m3 / tonne [134] Biogas potential of thin stillage 58 m3 / tonne [134]
The amount of biogas produced from the thin stillage generated in the bottoms
of the distillation tower and manure from the dairy cattle operation is 44.54 m3/h
(at 55 °C).
5.1.6. Co-generation System
The biogas combustion was simulated in a Gibbs equilibrium reactor. The
amount of air fed to the system was 375.06 kg/h. 421.96 kg/h of hot exhaust gas
at 1550.1 °C was obtained from the boiler which generated steam that was in
turn superheated before entering the turbine. The turbine generated 251.34
kg/h of steam to be recycled to the process and 28 kW of energy. The turbine
adiabatic efficiency was set to 72% based on data found in the literature. The
adiabatic efficiency refers to the ratio between the actual work output and the
work done in the reversible adiabatic cycle [154].
161
5.2. Pinch Analysis
Pinch analysis was performed for the design base-case simulation in order to
assess potential energy savings through heat integration. The process streams
listed in Table 36 were considered in the analysis.
Table 36: Heat stream inputs for pinch analysis in the base-case design.
Stream Name Supplied T [°C]
Target T [°C]
Energy [kW]
Pre-treatment Water Heater E-207 20.00 87.00 70.18 Pre-treatment Tank Side Heater, E-210 45 55 24.96 Hydrated Meal Cooler E-211 55.14 32.00 51.78 Conditioning Vessel Water Heater, E-221 20.00 32.00 0.18 SSF Reactor Cooler E-214 35.00 32.00 0.0050 Distillation Column Preheater E-225 32.42 61.06 36.20 Distillation Tower Reboiler E-230 127.28 128.38 187.49 Reflux Condenser E-227 110.09 107.06 50.54 Stillage Cooler E-232 128.38 94.00 48.67 Distillate Superheater E-231 107.06 122.06 1.00 Retentate Condenser E-235 122.06 82.80 23.97 Retentate Cooler E-236 82.80 40.00 4.01 Permeate Condenser E-238 122.06 35.00 22.82 Anaerobic Digester Heater E-248 45.00 55.00 4.40
Another input for the pinch analysis spreadsheet is the minimum temperature
difference between the hot and cold streams. A value of ∆Tmin equal to 10 °C
was taken from Kemp [117].
162
The hot and cold composite curves diagram obtained using the IChemE pinch
analysis datasheet [155] is presented as Fig. 39. It can be seen that the
maximum possible heat recovery is around 137 kW, while the external hot and
cold utility requirements are 188 and 65 kW, respectively.
Figure 39: Hot and cold composite curves for base-case design.
It may also be observed that the main consumer of the hot utility is the distillation
column reboiler E-230 (vaporization heat is represented by the blue line
between the 127.28 and 128.38 °C). In an effort to reduce the energy
Hot stream
Heat recovery: 137 kW
External hot utility requirement: 188 kW
External cold utility: 65 kW
Cold steam
163
requirements of the proposed IDF ethanol plant, three design alternatives were
considered for the distillation section.
As summarized in Chapter 4 and Table 37 below, the first configuration consisted
of a stripper column and a distillation column both operating at the same
pressure, each equipped with a steam reboiler (Configuration 1: see Fig. 31 and
Fig. 40). The second configuration contained a stripper followed by a distillation
tower operating at different pressures (Configuration 2: see Fig. 32, Figs. 41 and
42). In the latter setup, the beer is sent to each column in equal amounts. The
stripper operates at a higher pressure with direct steam injection. The overhead
vapour from the stripper condenses in the reboiler of the second complete
distillation tower and is recycled to the stripper. Two sets of pressures were
studied for this configuration.
Table 37: Results obtained for alternative distillation section configurations.
Design Base- Case
Config. 1:
Config. 2 A:
Config. 2 B:
Operating Pressure Stripper [bara] - 2.60 6.00 4.80 Operating Pressure distillation column [bara] 2.60 2.6 2.60 2.60
Steam Consumption [kg/L ethanol] 6.28 375.95 4.79 4.60 Cooling Water Consumption [kg/L ethanol] 71.90 65.70 66.08 62.31 Minimum hot utility [kW] 187.74 134.64 94.92 97.41 Minimum cold utility [kW] 65.12 14.77 14.42 13.92 Heat recovery possibility [kW] 137.00 204.72 190.11 179.28
164
The steam and cooling water requirements were extracted from the three
VMGSim™ models.
Figure 40: Hot and cold composite curves diagram for Configuration 1 (stripper + distillation tower – same pressure).
Heat recovery: 204 kW
External hot utility requirement: 135 kW
External cold utility: 15 kW
165
Figure 41: Hot and cold composite curves for Configuration 2 (stripper – 6 bara, distillation column – 2.6 bara).
Heat recovery: 190 kW
External hot utility requirement: 95 kW
External cold utility: 14 kW
166
Figure 42: Hot and cold composite curves for Configuration 2 (stripper – 4.8 bara, distillation column – 2.6 bara).
5.3. Improved Design
After analyzing the different distillation schemes and the results of their
corresponding pinch analyses, it was decided to use the second configuration,
in which a stripper working at 4.8 bara receives half of the beer and a complete
distillation tower working at 2.6 bara treats the other half of the beer. This
configuration proved superior in terms of steam and cooling water consumption.
The steam requirement was 4.60 kg/L ethanol and the cooling water
Heat recovery: 179 kW
External hot utility requirement: 97 kW
External cold utility: 14 kW
167
requirement was 62.31 kg/L ethanol. The steam consumption was reduced by
27% compared to the base-case. The cooling water consumption was reduced
by 13%, and the potential for waste heat recovery is 1.3 times as high as that of
the base-case design.
5.3.1 Improved Design Simulation Work
In terms of simulation, the only section of the plant in which changes were made
from the base-case was the distillation unit. The McCabe-Thiele graphical
method was used to obtain the theoretical number of stages for distillation
column B (T-228) using the same approach as described earlier for the base-
case distillation column (see Appendix 2). The improved distillation column
(Table 38) operates at 2.6 bara, generating a vapour distillate with 78 wt%
ethanol from a saturated liquid feed of 9.9 ethanol wt %.
The number of stages for the stripper was obtained by varying the steam
injection rate while keeping the composition of the vapour stream (stream 24)
fed to the membrane module at 78 wt%. The required volumetric flow of
retentate product was maintained at 112 L/h. Figure 43 demonstrates that seven
stages are required when the steam injection is set at 160 kg/h, marginally
higher than its absolute minimum of 159 kg/h.
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Table 38: Optimized design of distillation column in Configuration 2.
Design Parameter Symbol Value Units Feed to the column F 27.81 kmol/h Liquid Feed Flow LF 27.81 kmol/h
Feed concentration zF 0.041 Mole fraction 0.099 Mass fraction
Tower bottoms composition, ethanol concentration xB
0.0021 Mole fraction 0.0050 Mass fraction
Distillate composition, ethanol concentration xD
0.58 Mole fraction 0.78 Mass fraction
Reflux ratio factor 1.20 R=1.2 Rmin q-line LF/F 1 Rectifying section slope 0.63 Minimum reflux ratio Rmin 1.71 Calculated reflux ratio R 2.06 Minimum number of stages 4 Theoretical stages 8 Feed stage 4
Figure 43: Optimizing the number of stages for the stripper in Configuration 2 based on steam injection requirements.
158
159
160
161
162
163
164
165
166
4 6 8 10 12
Stea
m In
ject
ion
[Kg/
h]
Number of Stages
Minimum steam injection: 159 Kg/h
Optimal steam injection: 160 Kg/hNumber of Stages: 7
169
5.3.2 Heat integration
Once the distillation section of the plant was designed, an analysis was
performed to define the hot streams which might be used to heat cold streams
in the process. These are identified in Table 39.
Table 39: Hot and cold streams in the process for heat integration analysis.
Stream Name Supply Temperature [C]
Target Temperature [C]
Heat Flow [kW]
Stream Type
Pre-treatment Water Heater 25.00 87.00 70.18 COLD Pre-treatment Tank Side Heater 45 55 24.96 COLD Hydrated Meal Cooler 55.14 32.00 51.78 HOT Conditioning Vessel Water Heater 20.00 32.00 0.18 COLD SSF Reactor Cooler 35.00 32.00 0.005 HOT Stripper Column A Preheater 32.54 142.13 69.69 COLD Distillation Column B Preheater 32.54 120.68 54.56 COLD Distillation Column B Reboiler 127.53 128.24 58.26 COLD Reflux Condenser 114.72 108.68 35.12 HOT Stripper Column A Stillage Cooler 150.24 95.00 43.90 HOT Distillation Column B Stillage Cooler 128.24 90.00 22.68 HOT Distillate Superheater 116.47 131.47 0.91 COLD Retentate Condenser 131.47 82.78 22.78 HOT Retentate Cooler 82.78 40.00 3.71 HOT Permeate Condenser 131.47 40.36 19.68 HOT Anaerobic Digester Heater 45.00 55.00 4.40 COLD
In order to increase waste heat recovery, two heat exchangers were integrated.
The stillage from the stripper was used as the heating medium for the heater
through which water for the hydration of wheat passes on its way to the pre-
treatment tank. This eliminated the need for a stillage cooler for the stripper
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column bottoms. The condensate from the preheater of the distillation column
was used as the heating media of the superheater. The process flow diagrams
for the complete IDF plant with the heat integration arrangement were
presented in Chapter 4.
The steam consumption for the heat integrated process is 3.71 kg/L of ethanol -
a reduction of 18% compared to the new process without heat integration.
When compared to the base-case design, the steam savings rise to 40%. A
similar trend is observed for the cooling water. With the optimized process, the
cooling water requirement is of 53.02 kg/L ethanol, which represents a decrease
of 14% compared to the process without heat integration and 26% when
compared to the base-case (refer to Table 40).
The heat integration strategy adopted for the optimized design was developed
considering operational feasibility. Other opportunities for heat integration do
exist; however, adopting them could overcomplicate the operation and control
of the IDF ethanol plant. One of the objectives of the design is to offer the
farmer a plant that is easy to operate so excessive heat integration should be
avoided.
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Table 40: Steam and cooling water consumption in the different process designs.
Design base- case
Improved distillation without heat integration
Improved distillation with
heat integration Steam Consumption [kg/L ethanol]
6.28 4.60 3.71
Cooling Water Consumption [kg/L ethanol] 71.90 62.31 53.02
With the improved design including heat integration, the total amount of steam
required is 418 kg/h. As previously mentioned, the amount of steam produced
by the plant is 251 kg/h. Hence the anaerobic digestion generates sufficient
biogas to be used in the combined heat and power co-generation system to
satisfy 60 % of the facility’s steam requirement.
5.3.3 Mass Balance
The mass balance for the IDF ethanol plant was obtained for the optimized
process design including heat integration. It was summarized in the process flow
diagrams presented earlier in Section 4.2.2.1. The mass balance is presented for
zone 2 of the plant, in which the operation is continuous. It should be pointed
out that when carrying out this mass balance that starch, glucose and carbon
dioxide concentrations were considered negligible downstream of the
fermentation and CO2 absorption units (R-212 and T-218). The feed and product
compositions and flows in the continuous zone are summarized in Table 41.
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Table 41: Feed and product characteristics for the continuous zone of the IDF plant.
Stream number Description Mass Flow Composition 4 Feed to Distillation 1064.48 kg/h 9.95 EtOH wt%
33 Anhydrous Ethanol Product 89.02 kg/h 99.7 EtOH wt%
In the batch section of the plant, the following considerations were made:
- The amount of wheat used per batch is 8611.92 kg. It was assumed that
starch content in the wheat is 60% [62, 66].
- The concentration of yeast in the simultaneous saccharification and
fermentation reactor was set to 0.55 g/L as recommended by the yeast
vendors [107]. This corresponds to 14.18 kg per batch. Water is added to
the yeast conditioning tank in order to prepare a 50 wt% yeast solution.
- 7.86 kg of enzyme is fed to the SSF reactor. This amount gives the dosage
recommended by the supplier of 1.1 kg of enzyme per tonne of wheat
[64].
The schedule for operation of the IDF plant is presented in Appendix 3. The
batch unit operations and their required operating times are listed below:
1. The pre-treatment tank (TK-208) is filled with water and the ground wheat
in order to hydrate the milled grain (24 hours).
2. The hydrated meal is emptied from the pre-treatment tank and sent to
the simultaneous saccharification and fermentation reactor (R-212). On its
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way from the pre-treatment tank to the SSF reactor, the meal passes
through a cooler to reach fermentation temperature (24 hours).
3. Yeast and enzyme are added to the SSF reactor in order to begin the
fermentation process (1 hour).
4. Fermentation takes place in the SSF reactor R-212 (76 hours).
5. The fermented beer is evacuated from the SSF reactor into the beer tank
TK-215 (10 hours).
6. In parallel, the yeast conditioning vessel (V-220) is filled with heated water
(1 hour) and nutrients. The yeast is mixed in the vessel so that it can be fed
to the SSF reactor as required in step 3 (24 hours).
These steps are repeated for each batch and the timing is adjusted to have
continuous distillation and dehydration processing.
5.3.4 Control Strategy
In the IDF ethanol plant, the batch section is operated mainly manually. The
operator is responsible for switching between each sets of batch equipment as
necessary. Each of the four independent sets of batch fermentation equipment
includes:
- 1 pre-treatment tank (TK-208)
- 1 hydrated meal pump (P-209)
- 1 pre-treatment tank side heater (E-210)
- 1 SSF reactor (R-212)
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- 1 SSF reactor pump (P-213)
- 1 SSF reactor cooler (E-214)
Manual valves are used for isolating batch fermentation sets when out of
service.
Control strategies for the different sections of the plant are described below.
5.3.4.1 Wheat Hydration
The objective of the pre-treatment tank (TK-208) is to mix the wheat with warm
water to obtain a hydrated meal. The selected operating temperature for good
hydration in TK-208 is 55 °C [64]. The water entering the pre-treatment tank is
preheated in E-207 using the stillage from the stripper (T-227) in the downstream
operation, which temperature is 150 °C. To avoid excessive heating of the
water, a bypass line is installed and a temperature controller in the line going to
TK-208 after the heat exchanger E-207 is used to achieve the desired
temperature.
To adjust the amount of ground wheat and water filling TK-208, a flow ratio
controller receives the signal from flow transmitters located in the water line and
in the meal conveyor transporting the wheat to the tank. The ratio is adjusted to
ensure 28 wt% of dry solids content in the pre-treatment tank.
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The pH of TK-208 should remain in the range 3.5 – 4 [64]. This is achieved by a pH
controller that activates the control valves located in the acid and base lines
connected to the tank.
When the pre-treatment tank is filled, a level controller sends a signal to the
variable speed drive of the meal conveyor (MC-206) to stop it and a solenoid
valve closes on the water line.
The side heater (E-210) of the pre-treatment tank is used to keep the
temperature at 55 °C during the 24 hours in which the tank is in operation. The
steam supplied to the heat exchanger (E-210) is controlled with a temperature
controller that uses the temperature inside the tank as the controlled variable.
5.3.4.2 Meal Saccharification and Fermentation
After the hydration of the meal, it needs to be cooled down to the fermentation
temperature (32 °C). The meal is passed through a cooling water exchanger (E-
211) the flow of which is manipulated by the cooler outlet temperature
controller.
Fermentation is an exothermic reaction; therefore, the temperature inside the
reactor is regulated using an external cooler (E-214). The temperature inside the
reactor is the controlled variable and the cooling water flowrate is manipulated.
The level in the reactor (E-212) is controlled by a control valve positioned after
the pump (P-213 B) that transports the reaction mixture to the effluent cooler (E-
214) and beer tank (TK-215) after the fermentation is completed. An alarm is
176
generated when the ethanol content in the SSF reactor (R-212) reaches its
desired value (8.6 wt% - 10.8 vol%).
When the fermentation step is accomplished, the beer is sent to the beer tank
(TK-215) which can be said to divide the batch and continuous zones of the
plant. The flow to the distillation section is controlled by manipulating the control
valve downstream of the beer tank pump P-217. An override output from the
beer tank level controller is compared with the flow controller output in a low
selector; the lesser of these signals is sent to the control valve. This strategy was
adopted to prevent the beer tank from emptying completely and damaging P-
217.
5.3.4.3 CO2 Absorption
The objective of the CO2 scrubber (T-218) is to recover the remaining ethanol in
the CO2 stream from the SSF reactor offgas. A 90% recovery is expected. The
amount of water to attain the CO2 absorption required is achieved through a
flow ratio controller triggering a control valve in the water line. The ratio
controller receives signals from flow transmitters in the water line and in the
vapour line entering the absorber (T-218). The level in the scrubber is controlled
by a control valve after the pump (P-219) that sends the ethanol recovered in
the absorber back to the beer tank.
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5.3.4.4 Distillation
Half of the beer fed to the distillation section is sent to the stripper column A (T-
227) and half to distillation column B (T-228). A flow ratio controller receives the
signal from a flow meter in the beer line and it sends the appropriate setpoints to
each of the two feed flow controllers. Before entering the towers, the beer is
preheated (E-225 and E-226) in steam heat exchangers. The flow of steam to
each preheater is set by temperature controllers located downstream of the
exchangers.
The composition of the stripper distillate (70 wt%) is controlled by maintaining the
pressure and the temperature in the stripper column A (T-227). The temperature
in the stripper is controlled with a flow controller that triggers a control valve in
the steam line with the signal that it receives from a flow controller in the steam
line. The pressure in the column is regulated by a back pressure regulator on the
stripper overhead (stream 21). Liquid level in the stripper sump is controlled by
manipulating the stillage flow to the cooler (E-207). The overhead vapour
obtained condensed in the distillation column reboiler is sent to a flash drum (D-
230). A level controller in the drum adjusts a control valve after the pump that
sends the liquid back to the stripper to recover the remaining ethanol.
The composition of the vapour leaving the distillation column (T-228) condenser
is indirectly regulated by controlling overhead temperature and pressure. A
temperature controller sends its signal to a flow controller in line 14 which
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manipulates the stripper overhead flowrate through the reboiler (E-229) of the
distillation column (T-228). The pressure in the column is controlled by a pressure
controller that triggers a control valve in the cooling water supply to the reflux
condenser (E-232). A flow controller is used in the liquid reflux line after the reflux
pump (P-234) to attain the desired reflux flowrate. The level in the reflux knock-
out drum (D-233) is controlled by the reflux flow controller. A control valve after
the stillage cooler of the distillation column B (T-228) is used to control the level in
the base of the distillation tower.
5.3.4.5 Dehydration
The pressure in the membrane module fibers lumen is controlled by a back
pressure regulator. The vacuum outside the fibers, in the membrane module
housing, is maintained by the vacuum pump (P-244). The temperatures of the
streams emerging from the retentate condenser (E-239) and cooler (E-240) as
well as the permeate condenser (E-242) are controlled by changing the cooling
water flowrates to these heat exchangers. The level in the permeate vessel (V-
243) is regulated by the permeate pump (P-245) using an on-off control strategy.
5.3.4.6 Anaerobic Digestion
The thin stillage that is sent to the anaerobic digester (R-250) is collected in tank
TK-247. The level in this tank is controlled by a valve downstream of pump (P-249)
that transports the thin stillage to the AD. The temperature in the AD is controlled
by a side heater (E-252) with steam on the shell side. The steam flow through the
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heat exchanger is adjusted by the AD temperature controller. The output of the
anaerobic digester level controller is sent to a control valve after the digester
pump (P-251).
5.3.4.7 Co-generation System
The steam turbine is used to simultaneously generate low-pressure steam and
electricity. To obtain the required amount of low pressure steam for the plant, a
bypass line was installed. A pressure controller in the steam line serves to make
sure that the demand of low-pressure steam in the plant is satisfied, because if
the low-pressure steam pressure drops, it opens up the turbine bypass to the high
pressure steam header. A flow controller in the high pressure steam line before
the turbine is used to make sure that the low pressure steam demand of the
plant is satisfied mostly by exhaust steam from which excess energy has already
been recovered as electricity. The flow controller keeps the flow in the bypass at
a minimum rate and increases the high pressure flow to the turbine whenever
the bypass flow increases.
5.4. Occupational Health and Safety Analysis
In order to have a safe operation of the IDF ethanol plant, it is important to
perform a hazard analysis. A hazard is considered as a potential for harm, often
associated with conditions or activity that can result in an injury or illness if it is not
properly controlled [156]. There are governmental units like the Labour Program
in Canada and the US Department of Labour in the United States, which are
180
charged with recognizing and eliminating hazards in the workplace. The
Occupational Safety & Health Administration unit of the U.S. Department of
Labour classifies hazardous locations as areas "where fire or explosion hazards
may exist due to flammable gases or vapors, flammable liquids, combustible
dust, or ignitable fibers or flyings" [157]. Hazardous areas are categorized into
three classes. Class I locations are characterized by the presence of flammable
gases or vapors in the air in concentrations sufficient to be explosive. The
presence of these materials in the atmosphere presents a potential for
explosion, and could be ignited if an electrical or other source of ignition is
present. Class II locations are characterized by the presence of combustible
dust or pulverized material that may explode when suspended in the
atmosphere. Class III locations are characterized by the presence of easily-
ignitable fibers or flyings [157].
Within these classes there exist two divisions. Division 1 conditions are those in
which ignitable concentrations of flammable gases, vapors, or liquids are likely
to exist under normal operating conditions or frequently exist because of
maintenance or recurrent equipment failure. Situations where the flammable
fluids are not likely to exist under normal operating conditions or are present in
closed containers where the hazard can only escape through accidental
rupture or breakdown of such containers or in case of abnormal operation of
equipment are considered to be ‘Division 2’ [157]
181
Since ethanol is a flammable compound, the IDF ethanol plant must be
categorized as a Class I location. Since ethanol vapors are normally enclosed in
vessels and equipment, the facility falls within Division 2.
All of the equipment and instrumentation of the plant should be manufactured
for a Class I, Division 2 location. Lower explosive limit (LEL) ethanol vapor
detectors will be installed at strategic locations within the plant.
Additionally, by law, each piece of equipment in the IDF plant should have a
CRN (Canadian Registration Number). This number confirms that all equipment
working under pressure has been designed according to the required standards,
has been pressure tested and has been approved by a certified engineer. The
pressure equipment safety authority for Alberta is called ABSA (Alberta Boilers
Safety Association).
The material safety data sheet (MSDS) for ethanol, listing all contingency
measures for safe operation is found in the Appendices.
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6. Economic Evaluation
6.1. Equipment Sizing
6.1.1. Milling
A hammer tip speed of 100 m/s and a rotor speed of 1800 rpm were selected for
the hammer mill [120]. The hammer tip arc diameter calculated using Eq. 7 was
1.06 m. The hammer mill cost was obtained with an exponential scaling
approach. The capacity of the mill (358.83 kg/h) was used as the scaling
variable.
6.1.2. Simultaneous Saccharification and Fermentation
The results of the sizing of the batch equipment in the saccharification and
fermentation section are summarized in Table 42.
Table 42: Fermentation section sizing results.
Equipment Description Total
Height, L (m)
Cylindrical Portion
Height, H (m)
Conical Portion
Height, h (m)
Diameter, D
(m)
TK-208 A/B/C/D Pre-treatment tank 6 4 2 2.7
R-212 A/B/C/D SSF reactor 6 4 2 2.7
6.1.3. CO2 Absorber
The results obtained when sizing of the CO2 absorber are given in Table 43.
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Table 43: Absorber sizing results.
Equipment Description Height (m)
Diameter (m)
Actual Number of Trays
T-218 CO2 Absorber 4.2 0.14 6
6.1.4. Distillation Section
The results obtained for the sizing of the distillation section are given in Table 44.
Table 44: Distillation section sizing results.
Equipment Description Height (m)
Diameter (m)
Actual Number of Trays
T-227 Stripper Column A 7.3 0.17 10 T-227 Distillation Column B 8.4 0.17 12
D-230 Stripper Distillate Flash Drum 0.6 0.3 -
D-233 Reflux Accumulator 0.6 0.3 -
6.1.5. Dehydration Section
For the dehydration requirements of the process, the vapor permeation
membrane module selected is a Whitefox Technologies MB2 membrane
module.
The permeate vessel (V-243) dimensions calculated were:
Height, L = 0.34 m
Diameter, D = 0.17 m
184
The sizing parameters for the vacuum pump (P-244) used to create the pressure
difference acting as the driving force for the separation in the membrane
module are provided in Table 45.
Table 45: Sizing parameters for the vacuum pump.
Parameter Value Units Membrane module housing volume
0.036 m3
Piping Volume 0.0073 m3 Volume to evacuate 0.044 m3 Evacuating time 5 min Airflow capability 0.0087
0.31 m3/min ft3/min
Vacuum required 150 4.4
mbar in Hg
Using these values and a performance curve obtained from the vendor [158]
(see Appendix 9), the following vacuum pump was selected:
Manufacturer: Air Cadet – Cole-Parmer
Model: RK-07532-40, 07531-40
6.1.6. Anaerobic Digestion
The anaerobic digester (R-250) was sized according to its characteristics in terms
of biogas production rate and manure capacity. After comparing with the
results obtained in a benchmark study performed in the province of British
Columbia, which compiled information from 12 on-farm digesters [75], a volume
of 500 m3 was selected for the AD in the proposed ethanol plant.
185
6.1.7. Co-Generation System
The heat output and the electricity in the co-generation system obtained from
the VMGSim™ simulation of the IDF plant are recorded in Table 46.
Table 46: Results for the co-generation system sizing.
Parameter Value Units Heat Output 181,447.31 W 243.32 hp Electricity 27.89 kW
With these specifications in hand, different suppliers were contacted for costing
purposes. The equipment listed in Table 47 was finally selected.
Table 47: Equipment models and suppliers selected for the Co-Generation System
Equipment Description Supplier Model BO-254 TB-255 G-256
Biogas Boiler + Steam Turbine Generator Set
Wayne Adams [159]
Dresser-Rand: RLA Model
The dimensions obtained for the biogas dewatering knock out drum (D-253)
were:
Height, L = 0.10 m
Diameter, D = 0.05 m
6.1.8. Vessels and Tanks
The results of the sizing of the vessels and tanks are summarized in Table 48.
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Table 48: Results of the vessel and tank sizing calculations.
Equipment Description
Total Height,
L (m)
Cylindrical Portion
Height, H (m)
Conical Portion
Height, h (m)
Diameter, D
(m)
TK-204 Meal Storage Tank 6.5 4.3 2.2 3 TK-215 Beer Well 6.6 4.4 2.2 3
V-220 Yeast Conditioning Vessel 1.5 1 0.5 0.7
V-223 Enzyme Storage Vessel 1.1 - - 0.8
TK-241 Retentate Product Tank 4 3
TK-247 Thin Stillage Tank 2.3 1.5 TK-248 Wet Grain Tank 7 4.5
6.1.9. Heat Exchangers
Similarly, Table 49 was created for the heat exchangers distributed throughout
the IDF plant.
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Table 49: Heat exchanger sizing results.
Equipment Description Heat Duty (kW)
UA (kW/K)
Heat transfer coefficient (kW/K m2)
Area (m2)
E-207 Pre-treatment Water Heater / Stillage Cooler
70.18 1.33 3.97 [122] 0.33
E-210 A/B/C/D
Pre-treatment Tank Side Heater 24.96 0.25 3.97 [122] 6.18E-02
E-211 Hydrated Meal Cooler 51.80 3.10 1.28 [122] 2.42
E-214 A/B/C/D
SSF Reactor Cooler 0.01 3.37E-04 1.14 [122] 2.97E-04
E-221 Conditioning Vessel Water Heater
0.18 1.38E-03 3.97 [122] 3.48E-04
E-225 Stripper Column A Preheater 70.04 1.52 0.31 [127] 4.87
E-226 Distillation Column B Preheater
54.81 0.88 0.31[127] 2.81
E-229 Distillation Column B Reboiler
72.22 10.62 1.26 [122, 127] 8.40
E-232 Reflux Condenser 48.73 0.63 0.85 [122] 0.74
E-235 Distillate Superheater 0.96 0.04 0.52 [122,
127] 8.51E-02
E-236
Stillage from Distillation Column B Cooler
23.37 0.29 1.28 [122] 0.23
E-239 Retentate Condenser 22.23 0.37 0.85 [122] 0.43
E-240 Retentate Cooler 3.67 0.11 0.57 [122] 0.19
E-242 Permeate Condenser 23.88 1.05 0.65 [122] 1.63
E-252 Anaerobic Digester Heater 4.40 59.98 3.97 [122] 0.02
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6.1.10. Pumps
Tables 50 and 51 pertain to the sizing and selection of pumps. Performance
curves for the various models can be found in the Appendices.
Table 50: Pump sizing results
Flow (m3/h)
Flow (gpm)
Head (ft)
P-209 A/B/C/D Hydrated Meal Pump 1.09 4.78 27.38 P-213 A/B/C/D SSF Reactor Pump 0.011 0.047 52.72 P-217 Beer Pump 1.09 4.84 248.99 P-219 CO2 Absorber Pump 0.033 0.15 54.66 P-222 Yeast Slurry Pump 0.014 0.063 21.08 P-224 Enzyme Pump 0.008 0.035 21.91 P-231 Stripper Distillate Pump 0.25 1.11 238.68 P-234 Reflux Pump 0.21 0.91 169.37 P-245 Permeate Pump 0.046 0.20 26.58 P-249 Thin Stillage Pump 0.48 2.11 65.83 P-251 Anaerobic Digester Pump 0.36 1.59 82.51
Table 51: Pump model selection
Equipment Description Supplier Model P-209 A/B/C/D Hydrated Meal Pump Cole-Parmer RK-72010-15 / 72010-70 P-213 A/B/C/D SSF Reactor Pump Cole-Parmer RK-72010-15 / 72012-20
P-219 CO2 Absorber Pump Cole-Parmer RK-72010-15 / 72012-20 P-222 Yeast Slurry Pump Cole-Parmer RK-72010-15 / 72010-50 P-224 Enzyme Pump Cole-Parmer RK-72010-15 / 72010-50 P-245 Permeate Pump Cole-Parmer RK-72010-15 / 72012-10 P-249 Thin Stillage Pump Cole-Parmer RK-72010-15 / 72012-20
The pump costing variables are displayed in Table 52. The scaling variable used
for the pumps was the total head.
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Table 52: Pumps scaling values and exponents.
Equipment Description Scaling Variable Scaling Exponent P-217 Beer Pump 248.99 ft 0.8 [129] P-231 Stripper Distillate Pump 238.68 ft 0.8 [129] P-234 Reflux Pump 169.37 ft 0.8 [129] P-251 Anaerobic Digestion Pump 82.506 ft 0.8 [129]
6.1.11. Miscellaneous Equipment
Different suppliers were considered in an effort to identify the equipment that
would best satisfy the process requirements of the IDF plant; they are identified
in Table 53. This is followed by some illustrative figures of the equipment.
Table 53: Equipment models and suppliers selected.
Equipment Description Supplier Model U-205 Hopper Scale Horizon Systems SH-846 MC-206 Meal Conveyor ANEX MFG Company SBL-09 SP-246 Screw Press Press Technology &
Mfg., Inc. AGRI-PRESS® AGP-800
Figure 44: Bucket elevator meal conveyor considered for MC-206 [160].
190
Figure 45: Screw press considered for SP-246 [161].
Some other pieces of equipment were cost based on a characteristic
parameter using the exponential rule. The relevant scaling variables and
exponents appear in Table 54.
Table 54: Miscellaneous equipment scaling values and exponents
Equipment Description Scaling Variable Scaling Exponent
AG-216 Beer Well Agitator Tank volume: 34.09 m3 0.51 [129] F-237 Membrane Module Feed Filter Mass flow: 130.54 kg/h 0.6
6.2. CAPEX Evaluation
6.2.1 Equipment Costing
The equipment was cost using a scaling exponential approach as previously
described in Chapter 4. Table 55 gives the equipment list with the cost assigned
191
to each item. A table with the corresponding exponents and reference prices
employed is presented in the Appendices.
Table 55: Equipment costing
Equip. ID Description Cost in project
year [CD] MI-203 Hammer Mill $ 8,000 TK-204 Meal Storage Tank $ 90,000 U-205 Hopper Scale $ 10,000 MC-206 Meal Conveyor $ 20,000
E-207 Pre-treatment Water Heater / Stillage Cooler $ 1,000
TK-208 A Pre-treatment Tank A $ 77,000 TK-208 B Pre-treatment Tank B $ 77,000 TK-208 C Pre-treatment Tank C $ 77,000 TK-208 D Pre-treatment Tank D $ 77,000 P-209 A Hydrated Meal Pump A $ 200 P-209 B Hydrated Meal Pump B $ 200 P-209 C Hydrated Meal Pump C $ 200 P-209 D Hydrated Meal Pump D $ 200 E-210 Pre-treatment Tank Side Heater $ 500 E-211 Hydrated Meal Cooler $ 4,000 R-212 A SSF Reactor A $ 76,000 R-212 B SSF Reactor B $ 76,000 R-212 C SSF Reactor C $ 76,000 R-212 D SSF Reactor D $ 76,000 P-213 A SSF Reactor Pump A $ 200 P-213 B SSF Reactor Pump B $ 200 P-213 C SSF Reactor Pump A $ 200 P-213 D SSF Reactor Pump A $ 200 E-214 A SSF Reactor Cooler A $ 70 E-214 B SSF Reactor Cooler B $ 70 E-214 C SSF Reactor Cooler C $ 70 E-214 D SSF Reactor Cooler D $ 70 TK-215 Beer Well $ 93,130
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Table 55: Equipment costing ... (con)
Equip. ID Description Cost in project
year [CD] AG-216 Beer Well Agitator $ 42,000 P-217 Beer Pump $ 5,000 T-218 CO2 Absorber $ 25,000 P-219 CO2 Absorber Pump $ 200 V-220 Yeast Conditioning Tank $ 2,000 E-221 Conditioning Vessel Water Heater $ 140 P-222 Yeast Slurry Pump $ 200 V-223 Enzyme Storage Vessel $ 2,000 P-224 Enzyme Pump $ 200 E-225 Stripper Column A Preheater $ 21,000 E-226 Distillation Column B Preheater $ 14,000 T-227 Stripper Column A $ 35,000 T-228 Distillation Column B $ 39,000 E-229 Distillation Column B Reboiler $ 21,000 D-230 Stripper Distillate Flash Drum $ 8,000 P-231 Stripper Distillate Pump $ 5,000 E-232 Reflux Condenser $ 3,000 D-233 Reflux Knock-Out Drum $ 6,000 P-234 Reflux Pump $ 4,000
Dehydration Package $ 142,700
E-235 Distillate Superheater
E-236 Stillage from Distillation Column B Cooler
F-237 Membrane Module Feed Filter M-238 Dehydration Membrane Module E-239 Retentate Condenser E-240 Retentate Cooler TK-241 Retentate Product Tank E-242 Permeate Condenser V-243 Permeate Vessel P-244 Vacuum Pump P-245 Permeate Pump
193
Table 55: Equipment costing ... (con)
Equip. ID Description Cost in project
year [CD] SP-246 Whole Stillage Screw Press $ 60,000 TK-247 Thin Stillage Tank $ 20,000 TK-248 Wet Grain Tank $ 207,000 P-249 Thin Stillage Pump $ 200 R-250 Anaerobic Digester $ 500,000 P-251 Anaerobic Digester Pump $ 9,000 E-252 Anaerobic Digester Heater $ 400
D-253 Biogas Dewatering Knock-Out Drum $ 200
BO-254 Biogas Boiler
$ 220,000 TB-255 Steam Turbine G-256 Electricity Generator
TOTAL EQUIPMENT COST $ 2,200,000
6.2.2 Capital Investment The equipment cost obtained for the IDF plant was used to calculate the total
capital investment (TCI). The categories and percentages considered, and the
amounts allocated to each are found in Table 56, from which we conclude that
the TCI for the small scale integrated on-farm ethanol plant is 4,100,000 CD.
194
Table 56: Total capital investment and percentage breakdown
Category Percentage Value Equipment 54.7 $ 2,200,000 Piping 10.9 $ 400,000 Structural 6.3 $ 300,000 Insulation 2.9 $ 100,000 Electrical 3.0 $ 100,000 Instrumentation 3.1 $ 100,000 Automation 3.1 $ 100,000 Engineering 16.0 $ 700,000
Total Capital Investment $ 4,100,000
6.3. Farmer’s Financial Evaluation
The economic model was used to create an income statement, Table 57.
195
Table 57: Income statement to compare farmer’s financial position before and after installation of the IDF ethanol plant
INCOME STATEMENT Before IDF Ethanol Plant After IDF Ethanol Plant Monetary Value % Monetary Value % Sales Revenue $2,544,000 100% $3,153,000 100% Costs of Goods Sold $1,161,000 46% $759,000 24% GROSS PROFIT $1,383,000 54% $2,394,000 76% General Expenses $308,000 $308,000 Utility Expenses $4,000 $14,000 Labour $360,000 $630,000 Operating Expenses $672,000 26% $952,000 30%
EBITDA (Earnings before interest, taxes, depreciation,
amortization $711,000 28% $1,441,000 46%
Capital Investment $4,100,000 Interest rate 5% [149] Financial expenses (interests) $204,000 EARNINGS BEFORE TAXES $711,000 28% $1,237,000 39% Tax rate 15% [150] 15% [150] Taxes $107,000 4% $216,000 7% NET PROFIT $604,000 24% $1,021,000 32%
The motive for installation of the small scale integrated on-farm ethanol plant is
primarily economic, i.e. improvement of the farmer’s earnings, measured in this
analysis using the EBITDA (earnings before interest, taxes, depreciation and
196
amortization). The results of the income statement of Table 57 show that the
EBITDA for the farmers after the ethanol facility is integrated with the crop and
cattle operation is 49% higher than before its installation (see Fig. 46). When
considering the earnings EBITDA over the revenues coming from the sales of the
products (ethanol + animal feed + farm products), a 46% increment is observed
when ethanol production is integrated to the farm compared to 28% for the
only-farm scenario. In other words, if the farmer were to construct and operate
the complete IDF ethanol plant, s/he would earn 46 Canadian dollars for each
100 Canadian dollar of sales.
Figure 46: EBITDA analysis before and after the installation of the IDF plant.
Another conclusion to be drawn from the income statement is that the cost of
goods sold (costs directly related to the production of the farm and ethanol
$ 711,000.00
$ 1,441,000.00
Before IDFEthanol Plant
After IDFEthanol Plant
EBITDA (Earnings before interest, taxes, depreciation, amortization)
197
operations products) over the sales revenue after implementing the IDF ethanol
plant are 22% lower than before installing the plant. Therefore, the gross profit
over sales is 22% higher as illustrated in Fig. 47. The reduction in the cost of goods
sold over sales revenue is caused by decreased costs related to the fertilizer and
animal feed that are generated as co-products in the ethanol plant scenario.
Figure 47: Gross profit analysis before and after installation of the IDF plant.
The operating expenses over the sales are similar before and after the
installation of the ethanol plant. There is an increase in the costs attached to
utilities when the ethanol plant is installed as well as an increase in the labor
costs. The operating expenses over sales are 4% higher in the IDF ethanol
scenario.
It is important to mention that the IDF ethanol plant requires capital investment,
therefore, to install the plant there are extra financial considerations. Even after
46%
24%
54%
76%
Before IDF Ethanol Plant After IDF Ethanol Plant
Costs of Goods Sold GROSS PROFIT
22%
22%
198
financial expenses, though, the earnings before taxes over sales are 11% higher
than before the plant were installed on the farm.
After removing the expenses related to the interest and taxes, the final result of
implementing this project nets a 32% profit over sales. This means that when
integrating ethanol production to the farm operation, for each 100 dollars of
sales, 32 dollars return to the farmer as net profit. If the ethanol plant is not
installed on the farm, the net profit over sales is 24%. The results of this economic
analysis are summarized in Fig. 48.
Figure 48: Results of the economic analysis.
100%
54%
28% 28%24%
46%
26%
0%4%
100%
76%
46%39%
32%24%
30%
6% 7%
SALESREVENUE
GROSSPROFIT
EBITDA EARNINGSBEFORETAXES
NET PROFIT Costs ofGoods Sold
OperatingExpenses
FinancialExpenses
Taxes
Before IDF Ethanol Plant After IDF Ethanol Plant
199
6.3.1 Profitability Analysis
The return on investment calculated for this project is 25%. For each Canadian
dollar that the farmer will invest s/he will receive $0.25 Canadian dollars back.
The cash flow analysis demonstrated that the total amount projected over the
ten-year project lifetime was $ 6,110,000 Canadian dollars. Figure 49 shows that
by the fifth year, the farmer will start to recover the investment as it is expressed
in the cash flows presented in the Appendices.
Figure 49: Cumulative net cash inflows/outflows.
6.3.2 Sensitivity Analysis
Table 58 lists the price changes considered in the sensitivity analysis.
$ (6,000,000.00)
$ (4,000,000.00)
$ (2,000,000.00)
$ -
$ 2,000,000.00
$ 4,000,000.00
$ 6,000,000.00
$ 8,000,000.00
0 1 2 3 4 5 6 7 8 9 10
Net
inflo
ws
/ ou
tflow
s
Project Year
200
Table 58: Independent parameter variations for sensitivity analysis.
% change Ethanol Price
[CD/gal] Fertilizer Price
[CD/lb] Wheat Price
[CD/lb] -50 1.2 N: 0.302 / P2O5: 0.316 / K2O: 0.269 10.37 -25 1.8 N: 0.452 / P2O5: 0.474 / K2O: 0.403 8.64 0 2.4 N: 0.603 / P2O5: 0.632 / K2O: 0.537 6.91
25 3.0 N: 0.754 / P2O5: 0.790 / K2O: 0.672 5.18 50 3.6 N: 0.905 / P2O5: 0.948 / K2O: 0.806 3.46
In Figs. 50-52, the effects of these price variations on the economic indicators
are plotted. The complete set of results (numerical values) can be found in the
Appendices. No graph was made for the payback period since the payback
period remained constant at 5 years of payback period in the presence of all
price deviations.
Figure 50: Sensitivity analysis for the difference in the gross profit over sales before and after installation of the IDF ethanol plant.
10%
15%
20%
25%
30%
-50 -30 -10 10 30 50
Perc
enta
ge o
ver s
ales
[%
]
Variation Percentage [%]
∆ Gross Profit - Fertilizer Price
∆ Gross Profit - Ethanol Price
∆ Gross Profit - Wheat Price
201
Figure 51: Sensitivity analysis for the difference in the EBITDA over sales before and after installation of the IDF ethanol plant.
Figure 52: Sensitivity analysis for the difference in the net profit over sales before and after installation of the IDF ethanol plant.
10%
15%
20%
25%
30%
35%
-50 -30 -10 10 30 50
Perc
enta
ge o
ver s
ales
[%]
Variation Percentage [%]
∆ EBITDA - Fertilizer Price
∆ EBITDA - Ethanol Price
∆ EBITDA - Wheat Price
5%
10%
15%
20%
-50 -30 -10 10 30 50
Perc
enta
ge o
ver s
ales
[%]
Variation Percentage [%]
∆ Net Profit - Fertilizer Price
∆ Net Profit - Ethanol Price
∆ Net Profit - Wheat Price
202
The sensitivity analysis shows that any variation of wheat, fertilizer or ethanol
prices within a range of ± 50%, does not affect the profit improvement when
ethanol production is integrated with the crop and cattle operations of the
farm.
In terms of profit, the main variable affecting the profits (gross profit, EBITDA, net
profit) is the fertilizer. The range of change in the difference of the net profit
before and after the installation of the IDF ethanol plant for varying fertilizer cost
is 12% compared to 8% for wheat and 7% for ethanol. It is observed that as
fertilizer cost rises, so does the net profit difference (between the farm without
ethanol production and the farm after the installation of the ethanol facility). For
each 25% change in the fertilizer price, the net profit difference changes by 3%.
At higher fertilizer costs, the scenario in which the ethanol plant is integrated with
farm operations improves net profit.
In the case of wheat and ethanol, a 25% change in their prices induces a 2%
change in the net profit difference between the only-farm scenario and the
farming-ethanol production scenario. At greater ethanol prices, the net profit
difference increases, while the opposite trend is observed for increasing wheat
prices.
203
7. Conclusions
This thesis has evaluated the feasibility of installing a small-scale ethanol plant
within an integrated decentralised on-farm (IDF) concept. This model
incorporates ethanol production with crop and cattle operation using wheat as
feedstock for ethanol production and cattle manure as a co-substrate for
anaerobic digestion. The co-products of ethanol production are recycles to the
farm as fertilizer and animal feed. The ethanol plant was designed in two main
sections, one operating in batch mode and the other continuously for 355 days
per year, 24 hours a day. The plant was designed for a production capacity of
112 liters of anhydrous ethanol (purity 99.7 wt%) per hour.
The ethanol plant incorporates a batch simultaneous saccharification and
fermentation section producing beer (~9 EtOH wt%), after which the ethanol is
recovered in the distillation (~78 EtOH wt%) and dehydration sections (~99.7
EtOH wt%). An anaerobic digestion (AD) section was included to convert the
liquid fraction from the distillation column bottoms and manure from the cattle
operation to biogas. The process also contains a co-generation system which
uses the biogas from the AD to generate electric power (27 kW) and steam (60%
of facility requirements).
This conceptual design work was achieved using the VMGSim™ process simulator
[68]. Two design cases were evaluated. The base design case did not include
mechanisms for waste heat recovery. In the alternative process, heat
204
integration was used to generate an optimized design. The main differences
between the two design cases are to be found in the distillation section. In the
base case, a single distillation column was used, while the improved process
involved two towers operating at different pressures. The first tower was a
stripper, the overhead vapour of which serves as heating medium for the
reboiler of the distillation tower. Using this configuration and after performing
heat integration following a pinch analysis, reductions in steam consumption of
40% and of cooling water of 26% were achieved.
After completing the process design for the IDF plant, a mass balance was
obtained in order to size and cost the equipment. Costing was performed using
a scaling exponential approach in which reference prices of comparable
equipment were used and scaled up or down based on scaling variables
selected according to the operating principles of the equipment. The total
equipment cost estimated for the project was 2,200,000 CD. Using percentages
for additional items (piping, structural, insulation, electrical, instrumentation,
automation and engineering), the total capital investment (TCI) required for the
IDF plant was found to be 4,100,000 CD.
An economic analysis was then performed to investigate the economic benefit
which the farmer could expect following installation of the IDF ethanol plant. A
model was developed for a farm located in southern Alberta with 4500 acres of
205
land, of which 28% is used for wheat production. The farm also supports both
dairy (160 cows) and beef (450 head) cattle operation.
The results from this analysis showed that installing the IDF ethanol plant would
significantly improve the farmer’s financial position. If s/he integrates ethanol
production with farm operations, for each 100 Canadian dollars of sales, 18
more Canadian dollars would return. In terms of costs related to obtaining the
products, before the installation of the IDF ethanol plant, for each 100 Canadian
dollars of sales, the farmer requires 22 Canadian dollars more than when the
ethanol plant is installed on the farm. This reduction in the costs of goods sold is
mainly due to the fertilizer and animal feed that is produced in the IDF plant.
That represents an important portion of farm expenses when they are not
produced within the farm. Taking into consideration that the farmer would
require a loan to make the initial investment for the IDF ethanol plant, a 5%
interest rate was used in computing financial expenses. The earnings after the
financial expenses show that the farmer could expect an improvement in
revenue of 11 more Canadian dollars for each 100 Canadian dollars of sales
when coupling ethanol production with the farm activities.
The integration of ethanol production with farm operation resulted in a net profit
over sales of 32%. If the farmer adopts this project, s/he would earn 8 more
Canadian dollars for each 100 Canadian dollars of sales.
206
The return on investment of the project is 25% with a payback period of 5 years.
For every 100 Canadian dollars invested, the farmer will receive 25 Canadian
dollars back. The total capital investment of 4,100,000 CD will be paid after 5
years of ethanol plant operation.
A sensitivity analysis was performed on the economic indicators used to
evaluate the feasibility of the project in order to see the effect of different
ethanol, fertilizer and wheat prices on the economic results obtained. The
sensitivity analysis showed that the improvement of farm profitability through the
installation of an on-farm integrated ethanol production is not affected by ± 50%
change in fertilizer, wheat or ethanol prices. Changes in fertilizer cost affect the
net profit to a greater extent than the wheat or ethanol prices.
The present project is considered feasible from the technical and economical
points of view. It was concluded that the integration of ethanol production to a
farm operation is a promising approach which improves the farmer’s financial
position by reducing expenses related to fertilizer and animal feed, due to their
co-production in the ethanol facility, together with the improvement of the
energy balance of bioethanol production though the co-generation of utilities.
The process designed in this work has not been previously reported in the
literature. The use of SSF with cold enzymes, the heat integrated distillation
system and the use of membrane vapor permeation is unique to this work.
207
Although studies had been conducted for the main individual sections of the IDF
plant, the contribution of this work is the complete highly integrated system
designed, simulated and the economic assessment attached to the technical
evaluation .
The use of pinch analysis and heat integration as tools for process improvement
for the IDF plant is another contribution of the project. Some recommendations
for future work include the modelling of the anaerobic digester together with
detailed equipment design for the different unit operations of the IDF ethanol
plant. In terms of the economic analysis, other indicators, such as the net
present value (NPV) and the internal rate of return (IRR) can be calculated in
order to have a more comprehensive economic evaluation. The examination of
the co-products, mainly the digestate and the wet grain or thick stillage, is an
important assessment to have a better understanding of the different
integration possibilities. The future goal for this project would be the installation
of an on-farm demonstration facility and the evaluation of the funding
opportunities that the farmer can use in order to invest in this type of projects.
208
References
[1] S. I. Mussatto, G. Dragone, P. M. R. Guimarães, J. P. A. Silva, L. M. Carneiro,
I. C. Roberto, et al., "Technological trends, global market, and challenges
of bio-ethanol production," Biotechnology Advances, vol. 28, pp. 817-830,
2010.
[2] C. Piccolo and F. Bezzo, "A techno-economic comparison between two
technologies for bioethanol production from lignocellulose," Biomass and
Bioenergy, vol. 33, pp. 478-491, 2009.
[3] H. L. MacLean, L. B. Lave, and R. Lankey, "A life-cycle comparison of
alternative automobile fuels," Journal of the Air & Waste Management
Association (Air & Waste Management Association), vol. 50, pp. 1769-
1779, 2000.
[4] A. J. Liska, H. S. Yang, V. R. Bremer, T. J. Klopfenstein, D. T. Walters, G. E.
Erickson, et al., "Improvements in Life Cycle Energy Efficiency and
Greenhouse Gas Emissions of Corn-Ethanol," Journal of Industrial Ecology,
vol. 13, pp. 58-74, 2009.
[5] Cheminfo Services Inc., "Life Cycle Assessment of Renewable Fuel
Production from Canadian Biofuel Plants for 2008-2009," Canadian
Renewable Fuels Association, OttawaNovember 25, 2009 2009.
[6] Whitten and Reyes, "Air Quality and Ethanol in Gasoline " presented at the
9th Annual National Ethanol Conference: Policy & Marketing, 2004.
209
[7] K. R. Szulczyk, B. A. McCarl, and G. Cornforth, "Market penetration of
ethanol," Renewable and Sustainable Energy Reviews, vol. 14, pp. 394-
403, 2010.
[8] Renewable Fuel Association. (2013, June 25, 2013). Historic U.S. Fuel
Ethanol Production. Available: http://www.ethanolrfa.org/pages/statistics
[9] A. E. Farrell, R. J. Plevin, B. T. Turner, A. D. Jones, M. O'Hare, and D. M.
Kammen, "Ethanol Can Contribute to Energy and Environmental Goals,"
Science, vol. 311, pp. 506-508, 2006.
[10] Renewable Fuels Association, "Fuel Ethanol - Industry Guidelines,
Specifications and Procedures," 2010.
[11] L. Bromberg, D. R. Cohn, and J. B. Heywood, "Calculations Of Knock
Suppression In Highly Turbocharged Gasoline/Ethanol Engines Using Direct
Ethanol Injection," Massachusetts Institute of Technology2006.
[12] AFDC, "Fuel Properties Comparison Results," U. D. o. Energy, Ed., ed:
Alternative Fuels Data Center, 2013.
[13] Bailey and Russell, "Title," unpublished|.
[14] Hadder, "Ethanol Demand in United States Regional Production of
Oxygenate-limited Gasoline," U.S. Department of Energy, Ed., ed. Oak
Ridge: Oak Ridge National Laboratory, 2000.
[15] J. Lane. (2013, June 20, 2013). Freedom of choice: E20 ethanol blends
take the value crown from gasoline, E10. Biofuels Digest.
210
Available: http://www.biofuelsdigest.com/bdigest/2013/06/20/freedom-
of-choice-e20-ethanol-blends-take-the-value-crown-from-gasoline-e10/
[16] (2013, June 20, 2013). Reported E85 Prices.
Available: http://www.e85prices.com/
[17] R. Shockey, T. Aulich, B. Jones, G. Mead, and P. Steevens, "Optimal
Ethanol Blend-Level Investigation," Energy & Environmental Research
Center - University of North Dakota Minnesota Center for Automotive
Research - Minnesota State UniversityDecember 31, 2007 2007.
[18] Bioenergy Technologies Office / Energy Efficiency & Renewable Energy,
"Replacing the Whole Barrel To Reduce U.S. Dependence on Oil," US
Department of Energy, Ed., ed, 2013, pp. 1-21.
[19] L. R. Lynd, J. H. Cushman, R. J. Nichols, and C. E. Wyman, "Fuel Ethanol
from Cellulosic Biomass," Science, vol. 251, pp. 1318-1323, March 15, 1991
1991.
[20] E. Felix and D. R. Tilley, "Integrated energy, environmental and financial
analysis of ethanol production from cellulosic switchgrass," Energy, vol. 34,
pp. 410-436, 2009.
[21] A. McAloon, Frank Taylor, W. Yee, K. Ibsen, and R. Wooley, "Determining
the Cost of Producing Ethanol from Corn Starch and Lignocellulosic
Feedstocks," U.S. Department of Agriculture and U.S. Department of
Energy, Ed., ed. Golden, 2000.
211
[22] B. Schuster and M. Chinn, "Consolidated Bioprocessing of Lignocellulosic
Feedstocks for Ethanol Fuel Production," BioEnergy Research, vol. 6, pp.
416-435, 2013/06/01 2013.
[23] AAFC, "Canada Wheat Sector Profile Part One: Overview," vol. 2, F. I. a. I.
Section, F. E. A. Division, R. A. Directorate, S. P. Branch, and A. A.-F.
Canada, Eds., ed: Market Analysis Group Grains and Oilseeds Division
Food Value Chain Bureau Market and Industry Services Branch Agriculture
& Agri-Food Canada, 2010.
[24] U. o. Saskatchewan, "Winter Wheat Production Manual: Grading and
Classes," in Winter Wheat Production Manual, U. o. Saskatchewan, Ed., ed,
2014.
[25] AARD 2008, Integrating Biogas, Confined Feedlot Operations and Ethanol
Production. Agri-Facts.
Available: http://www1.agric.gov.ab.ca/$department/deptdocs.nsf/all/a
gdex11839
[26] AAFC, "Canadian Farm Fuel and Fertilizer: Prices and Expenses " vol. 5, F. I.
a. I. Section, F. E. A. Division, R. A. Directorate, S. P. Branch, and A. A.-F.
Canada, Eds., ed: Market Analysis Group Grains and Oilseeds Division
Food Value Chain Bureau Market and Industry Services Branch Agriculture
& Agri-Food Canada, 2013.
212
[27] "Alternative Fuels - Ethanol," in Natural Resources Canada. Office of
Energy Efficiency. vol. http://oee.nrcan.gc.ca/transportation/alternative-
fuels/fuel-facts/ethanol/2070, ed, 2013.
[28] Canadian Renewable Fuels Association. (2011, June 25, 2013). Ethanol
Key Issues - Fact Sheets / Responses to Key Issues for Canada’s Ethanol
Industry. 23.
Available: http://www.greenfuels.org/uploads/documents/ethanol-fact-
sheet-august102011.pdf
[29] ASTM, "Standard Specification for Denatured Fuel Ethanol for Blending
with Gasolines for Use as Automotive Spark-Ignition Engine Fuel," vol.
D4806, ed. United States: ASTM International, 2013, p. 8.
[30] Harfield, "Biomass Conversion Technologies & Bio-Products," presented at
the 2013 Ag Biomass National Conference, Ottawa, Ontario 2013.
[31] S. T. C. Inc. and C. S. Inc., "Life Cycle Assessment of Selected Biofuel
Pathways in Alberta " Alberta Innovates – Energy and Environment
Solutions (AI-EES) AI-EES RFP No. AI-EES 11-01, 2013.
[32] S&T, "GHGenius - Model Version 4.03a," S. T. C. Inc., Ed., ed. Delta, British
Columbia, 2012.
[33] Edwards and Anex, "Co-product Allocation in Life Cycle Assessment: A
Case Study," Stillwater, MNJune 18, 2009 2009
213
[34] L. K. Gohar and K. P. Shine, "Equivalent CO2 and its use in understanding
the climate effects of increased greenhouse gas concentrations,"
Weather, vol. 62, pp. 307-311, 2007.
[35] S. Canada, "Production of principal field crops, July 2013," Statistics
Canada Catalogue no. 11-001-X, 2013.
[36] W. T. C. Ltd., "Internal Communication," V. Andrade, Ed., ed, 2013.
[37] R. Turton, R. C. Bailie, and W. B. Whiting, Analysis, synthesis, and design of
chemical processes: Prentice Hall, 2009.
[38] Jacques, Lyons, and Kelsall, The Alcohol Textbook, 4th ed. Nottingham,
UK, 2003.
[39] Smith. (2007, Wet vs. Dry. HEREFORD WORLD.
Available: http://www.hereford.org/
[40] Pandey. (2008). Handbook of Plant-Based Biofuels.
Available: http://www.crcnetbase.com.ezproxy.lib.ucalgary.ca/isbn/978-
1-56022-175-3
[41] Dale and Tyner, "Economic and Technical Analysis of Ethanol Dry Milling:
Model Description," Purdue UniversityApril 24, 2006 2006.
[42] U. D. o. Energy. (2014). Ethanol Production and Distribution.
Available: http://www.afdc.energy.gov/fuels/ethanol_production.html
[43] L. Tao and A. Aden, "The Economics of Current and Future Biofuels," In
Vitro Cellular & Developmental Biology. Plant, vol. 45, pp. 199-217, 2009.
214
[44] RFA. (2014, January 14, 2014). How Ethanol is Made.
Available: http://www.ethanolrfa.org/pages/how-ethanol-is-made
[45] L. Audet, "Emerging feed mill technology: keeping competitive," Animal
Feed Science and Technology, vol. 53, pp. 157-170, 1995.
[46] Wade, Organic Chemistry, 7th ed.: Prentice-Hall, 2010.
[47] S. D. Textor, G. A. Hill, D. G. Macdonald, and E. S. Denis, "Cold enzyme
hydrolysis of wheat starch granules," The Canadian Journal of Chemical
Engineering, vol. 76, pp. 87-93, 1998.
[48] Y. Lin and S. Tanaka, "Ethanol fermentation from biomass resources:
current state and prospects," Applied Microbiology and Biotechnology,
vol. 69, pp. 627-642, 2006/02/01 2006.
[49] A. C. Wilkie, K. J. Riedesel, and J. M. Owens, "Stillage characterization and
anaerobic treatment of ethanol stillage from conventional and cellulosic
feedstocks," Biomass and Bioenergy, vol. 19, pp. 63-102, 2000.
[50] H.-W. Kang, Y. Kim, S.-W. Kim, and G.-W. Choi, "Cellulosic ethanol
production on temperature-shift simultaneous saccharification and
fermentation using the thermostable yeast Kluyveromyces marxianus
CHY1612," Bioprocess and Biosystems Engineering, vol. 35, pp. 115-122,
2012/01/01 2012.
[51] I. Watanabe, N. Miyata, A. Ando, R. Shiroma, K. Tokuyasu, and T.
Nakamura, "Ethanol production by repeated-batch simultaneous
saccharification and fermentation (SSF) of alkali-treated rice straw using
215
immobilized Saccharomyces cerevisiae cells," Bioresource Technology,
vol. 123, pp. 695-698, 2012.
[52] A. Chandel, G. Chandrasekhar, M. Lakshmi Narasu, and L. Venkateswar
Rao, "Simultaneous saccharification and fermentation (SSF) of aqueous
ammonia pretreated Saccharum spontaneum (wild sugarcane) for
second generation ethanol production," Sugar Tech, vol. 12, pp. 125-132,
2010/06/01 2010.
[53] W. Zhang, Y. Lin, Q. Zhang, X. Wang, D. Wu, and H. Kong, "Optimisation of
simultaneous saccharification and fermentation of wheat straw for
ethanol production," Fuel, vol. 112, pp. 331-337, 2013.
[54] M. d. Neves, T. Kimura, N. Shimizu, and M. Nakajima, "State of the Art and
Future Trends of Bioethanol Production " Dynamic Biochemistry, Process
Biotechnology and Molecular Biology vol. 1, p. 14, 2007.
[55] C. E. Wyman, D. D. Spindler, and K. Grohmann, "Simultaneous
saccharification and fermentation of several lignocellulosic feedstocks to
fuel ethanol," Biomass and Bioenergy, vol. 3, pp. 301-307, 1992.
[56] C. A. Cardona and Ó. J. Sánchez, "Fuel ethanol production: Process
design trends and integration opportunities," Bioresource Technology, vol.
98, pp. 2415-2457, 2007.
[57] C. E. Wyman, "Ethanol from lignocellulosic biomass: Technology,
economics, and opportunities," Bioresource Technology, vol. 50, pp. 3-15,
1994.
216
[58] M. A. d. Neves, T. Kimura, N. Shimizu, and K. Shiiba, "Production of alcohol
by simultaneous saccharification and fermentation of low-grade wheat
flour," Brazilian Archives of Biology and Technology, vol. 49, pp. 481-490,
2006.
[59] K. Hofvendahl, C. Åkerberg, G. Zacchi, and B. Hahn-hägerdal,
"Simultaneous enzymatic wheat starch saccharification and fermentation
to lactic acid by Lactococcus lactis," Applied Microbiology and
Biotechnology, vol. 52, pp. 163-169, Aug 1999 1999.
[60] K. Öhgren, R. Bura, G. Lesnicki, J. Saddler, and G. Zacchi, "A comparison
between simultaneous saccharification and fermentation and separate
hydrolysis and fermentation using steam-pretreated corn stover," Process
Biochemistry, vol. 42, pp. 834-839, 2007.
[61] A. Wingren, M. Galbe, and G. Zacchi, "Techno-Economic Evaluation of
Producing Ethanol from Softwood: Comparison of SSF and SHF and
Identification of Bottlenecks," Biotechnology Progress, vol. 19, pp. 1109-
1117, 2003.
[62] G. A. H. D. G. M. Xiaosu Lang, "Recycle Bioreactor for Bioethanol
Production from Wheat Starch I. Cold Enzyme Hydrolysis," Energy Sources,
vol. 23, pp. 417-425, 2001/06/01 2001.
[63] Purohit and Mishra, "Simultaneous Saccharification and Fermentation of
Overnight Soaked Sweet Potato for Ethyl Alcohol Fermentation," Advance
Journal of Food Science and Technology, vol. 4, pp. 56-59, 2012.
217
[64] Genencor, "STARGEN™ 002. Granular Starch Hydrolyzing Enzyme for
Ethanol Production," D. U. Inc., Ed., ed, 2009.
[65] H.-J. Huang, S. Ramaswamy, U. W. Tschirner, and B. V. Ramarao, "A review
of separation technologies in current and future biorefineries," Separation
and Purification Technology, vol. 62, pp. 1-21, 2008.
[66] K. Sosulski and F. Sosulski, "Wheat as a feedstock for fuel ethanol," Applied
Biochemistry and Biotechnology, vol. 45-46, pp. 169-180, 1994/03/01 1994.
[67] J. P. Crawshaw and J. H. Hills, "Sorption of ethanol and water by starchy
materials," Industrial & Engineering Chemistry Research, vol. 29, pp. 307-
309, 1990/02/01 1990.
[68] V. M. Group, "VMG Process Simulator," 7.0 ed. Canada, 2012.
[69] Stoffers. (2008) Rethinking Reboilers. Biofuels International.
Available: http://www.alfalaval.com/industries/sugar/Documents/biofuels
_international_rethinking_reboilers_PPI00281EN.pdf
[70] Pilling and Holden. (2009, September, 2009) Choosing Trays and Packings
for Distillation CEP Magazine 7.
[71] Moura and Avila, "Comparative analysis of the main commercially
available alcohol dehydration processes-azeotropic, extractive and
molecular sieve-a technical and economical approach," presented at
the International Society of Sugar Cane Technologists 25th, 2005.
[72] R. Baker, Membrane Technology and Applications, 3rd Edition ed.
Somerset, NJ, USA: Wiley, 2012.
218
[73] M. J. Carmo, M. G. Adeodato, A. M. Moreira, E. J. S. Parente, Jr., and R. S.
Vieira, "Kinetic and Thermodynamic Study on the Liquid Phase Adsorption
by Starchy Materials in the Alcohol-Water System," Adsorption, vol. 10, pp.
211-218, 2004/09/01 2004.
[74] T. Pettersen and K. M. Lien, "Design of hybrid distillation and vapor
permeation processes," Journal of Membrane Science, vol. 102, pp. 21-30,
1995.
[75] Werner and Strehler, "British Columbia On-Farm Anaerobic Digestion
Benchmark Study," B.C. Agricultural Research and Development
Corporation.
[76] Monnet, "An Introduction to Anaerobic Digestion of Organic Wastes,"
Remade Scotland2003.
[77] Gas Processors Association, Engineering Data Book: Gas Processors
Suppliers Association, 2004.
[78] P. Börjesson and B. Mattiasson, "Biogas as a resource-efficient vehicle fuel,"
Trends in Biotechnology, vol. 26, pp. 7-13, 2008.
[79] I. McNeil Technologies, "Colorado Agriculture IOF Technology
Assessments: Anaerobic Digestion " Governor’s Office of Energy
Conservation and Management 2005.
[80] A. A. a. R. Development. (2008, Anaerobic Digesters. Agri-Facts.
Available: http://www1.agric.gov.ab.ca/$department/deptdocs.nsf/all/a
gdex10945
219
[81] Wu-haan, "Evaluation of ultrasonic pretreatment on anaerobic digestion
of biomass for methane production," Master of Science, Environmental
Science - Biorenewable Resources and Technology, Iowa State University,
Ames, Iowa, 2008.
[82] M. Westerholm, M. Hansson, and A. Schnürer, "Improved biogas
production from whole stillage by co-digestion with cattle manure,"
Bioresource Technology, vol. 114, pp. 314-319, 2012.
[83] H. Annand, "Biochemical Methane Potential for Wheat-Based Fuel Ethanol
and Beef Feedlot Integration," Master of Science Department of Chemical
and Biological Engineering University of Saskatchewan, Saskatoon,
Saskatchewan, 2011.
[84] A. C. Wilkie, "Anaerobic Digestion of Dairy Manure: Design and Process
Considerations " Cornell University NRAES - 176, 2005.
[85] AgSTAR. (2012). Anaerobic Digesters.
Available: http://www.epa.gov/agstar/anaerobic/ad101/anaerobic-
digesters.html
[86] Hamilton. (2012, January 19, 2014). Types of Anaerobic Digesters.
Available: http://www.extension.org/pages/30307/types-of-anaerobic-
digesters#.UtwvSfTn8nJ
[87] ETPC. (2011, January 19, 2014). Upflow Anaerobic Sludge Blanket.
Available: http://www.etpcwater.com/products/Upflow_Anaerobic_Slud
ge_Blanket
220
[88] G. Eriksson and B. Kjellström, "Assessment of combined heat and power
(CHP) integrated with wood-based ethanol production," Applied Energy,
vol. 87, pp. 3632-3641, 2010.
[89] C. Energy. (2014, January 20, 2014). Cogeneration & CHP.
Available: http://www.clarke-energy.com/chp-cogeneration/
[90] E. a. E. A. Inc, I. I. Company, and E. R. G. Inc, "Biomass Combined Heat
and Power Catalog of Technologies," U. S. Environmental Protection
Agency, Combined Heat and Power Partnership2007.
[91] Mustafa, McKinnon, and Christensen, "The Nutritive Value of Thin Stillage
and Wet Distillers’ Grains for Ruminants: A Review," University of
Saskatchewan2000.
[92] Y. Kim, N. S. Mosier, R. Hendrickson, T. Ezeji, H. Blaschek, B. Dien, et al.,
"Composition of corn dry-grind ethanol by-products: DDGS, wet cake,
and thin stillage," Bioresource Technology, vol. 99, pp. 5165-5176, 2008.
[93] Haandel and Frassinetti, "Profitability Increase of Alcohol Distilleries by the
Rational Use of Byproducts," Water Science & Technology, vol. 29, pp.
117–124 1994.
[94] G. J. Sheehan and P. F. Greenfield, "Utilisation, treatment and disposal of
distillery wastewater," Water Research, vol. 14, pp. 257-277, 1980.
[95] S. S. Lele, P. J. Rajadhyaksha, and J. B. Joshi, "Effluent treatment for
alcohol distillery: Thermal pretreatment with energy recovery,"
Environmental Progress, vol. 8, pp. 245-252, 1989.
221
[96] E. M. Larson, R. A. Stock, T. J. Klopfenstein, M. H. Sindt, and R. P. Huffman,
"Feeding value of wet distillers byproducts for finishing ruminants," Journal
of Animal Science, vol. 71, pp. 2228-36, August 1, 1993 1993.
[97] Lukehurst, Frost, and Al-Seadi, "Utilisation of digestate from biogas plants
as biofertiliser," I. Bioenergy, Ed., ed, 2010.
[98] Birkmose, "Nitrogen recovery from organic manures: improved slurry
application techniques and treatment: the Danish scenario.," in
International Fertiliser Society, York, 2009.
[99] McAloon, Taylor, Yee, Ibsen, and Wooley, "Determining the Cost of
Producing Ethanol from Corn Starch and Lignocellulosic Feedstocks," U. S.
D. o. A. a. U. S. D. o. Energy, Ed., ed. Golden, Colorado: National
Renewable Energy Laboratory, 2000.
[100] Lantz and Li, "Bioethanol from Starch Process Simulation Tutorial on PRO
II/8.2," Washington State University.
[101] R. Davis, "Parameter Estimation for Simultaneous Saccharification and
Fermentation of Food Waste Into Ethanol Using Matlab Simulink," Applied
Biochemistry and Biotechnology, vol. 147, pp. 11-21, 2008/03/01 2008.
[102] AgSTAR, "AD System Designs and End Uses," E.-U. S. E. P. Agency, Ed., ed,
2012.
[103] R. L. Motard, M. Shacham, and E. M. Rosen, "Steady state chemical
process simulation," AIChE Journal, vol. 21, pp. 417-436, 1975.
222
[104] Carlson, "Donèt Gamble with Physical Properties For Simulations,"
Chemical Engineering Progress, 1996.
[105] J. Barthel and R. Buchner, Dechema Chemistry Data Series: Deutsche
Gesellschaft für Chemisches Aparatewesen., 1995.
[106] D. R. Lide, CRC Handbook of Chemistry and Physics, 85th Edition: Taylor &
Francis, 2004.
[107] L. B. D. Spirits, "Bio-Ferm® XR High Performance Active Dry Yeast," L. B. D.
Spirits, Ed., ed. Duluth 2013.
[108] Elly, "POLYMATH 6.1," vol. 6.1, C. D. Corporation, Ed., ed, 2006.
[109] C. Geankoplis, Transport processes and separation process principles
(includes unit operations) fourth edition: Prentice Hall Press, 2003.
[110] W. T. C. Ltd., "Internal Communcation - Membrane Tool," W. T. Canada,
Ed., ed. Calgary, Canada, 2012.
[111] F. M. El-Mahallawy and S. E.-D. Habik, Fundamentals and Technology of
Combustion. Amsterdam, NLD: Elsevier Science & Technology, 2002.
[112] McAllister, Chen, and Fernandez-Pello, Fundamentals of Combustion
Processes: Springer New York, 2011.
[113] J. R. Couper, W. R. Penney, and J. R. Fair, Chemical Process Equipment :
Selection and Design (3rd Edition). St. Louis, MO, USA: Butterworth-
Heinemann, 2012.
[114] Viguri, "Chemical Process Design. Subject 7. Equipment Sizing and
Costing," Chemical Engineering and Inorganic Chemistry Department.
223
[115] M. Foley, V. Andrade, Ed., ed, 2014.
[116] W. L. Luyben, Plantwide dynamic simulators in chemical processing and
control New York: Marcel Dekker Inc., 2002.
[117] I. C. Kemp, "Pinch Analysis and Process Integration - A User Guide on
Process Integration for the Efficient Use of Energy (2nd Edition)," ed:
Elsevier.
[118] I. C. Kemp, Pinch Analysis and Process Integration - A User Guide on
Process Integration for the Efficient Use of Energy (2nd Edition): Elsevier,
2007.
[119] G. D. Zupančič and M. Roš, "Heat and energy requirements in
thermophilic anaerobic sludge digestion," Renewable Energy, vol. 28, pp.
2255-2267, 2003.
[120] K. Koch, "Hammermills and Roller Mills," North Dakota State University
Kansas State University Agricultural Experiment Station and Cooperative
Extension Service Northern Crops Institute Fargo, ND 2002.
[121] L. Shanghai Zonx International Trading Co. (2014, February 27, 2014). Beer
fermenter tank. Available: http://zonx-china.en.made-in-
china.com/offer/CqZmMEsoyPWO/Sell-Beer-Fermenter-Tank.html
[122] Seider, Seader, and Lewin, Product & Process Design Principles. Synthesis,
Analysis and Evaluation. New York: John Wiley & Sons, Inc., 2004.
[123] Andrade, "Whitefox Membrane Performance. Internal Report," Whitefox
Technologies Canada2013.
224
[124] Kaupp, "Title," unpublished|.
[125] E. T. Inc., "Feasibility Study – Biogas upgrading and grid injection in the
Fraser Valley, British Columbia " BC Innovation Council2008.
[126] PowderHandling. (2014). PowderHandling - Bulk Density Chart.
Available: http://www.powderhandling.com.au/bulk-density-chart
[127] Andrade, "Whitefox Customer Project. Internal Report," Whitefox
Technologies Canada2013.
[128] Chaurette. (2005, The influence of Specific Gravity on Total Head.
Available: http://www.pumpfundamentals.com/yahoo/The%20influence
%20of%20specific%20gravity%20on%20total%20head.pdf
[129] D. Humbird, R. Davis, L. Tao, C. Kinchin, D. Hsu, A. Aden, et al., "Process
Design and Economics for Biochemical Conversion of Lignocellulosic
Biomass to Ethanol: Dilute-Acid Pretreatment and Enzymatic Hydrolysis of
Corn Stover," National Renewable Energy Laboratory, U.S. Department of
Energy, Office of Energy Efficiency & Renewable Energy, Golden,
Colorado2011.
[130] (2013, 2013/07//) Chemical engineering plant cost index (CEPCI).
Chemical Engineering [Brief article
Statistical table]. 60.
Available: http://go.galegroup.com/ps/i.do?id=GALE%7CA336605689&v=
2.1&u=ucalgary&it=r&p=AONE&sw=w
225
[131] Andrade, "Whitefox Percentage Analysis. Internal Report," Whitefox
Technologies Canada2013.
[132] S. Blum, "Summary of Meeting with John Hofer," V. Andrade, Ed., ed, 2013.
[133] USDA, "Agricultural Marketing Resource Center - Weekly Ethanol, Distillers
Grain and Corn Prices," U. S. D. o. Agriculture, Ed., ed, 2013.
[134] A. A. a. R. Development. (2011, Biogas Energy Potential in Alberta. Agri-
Facts.
Available: http://www1.agric.gov.ab.ca/$department/deptdocs.nsf/all/a
gdex11397/$file/768-3.pdf?OpenElement
[135] AARD, "Alberta Fertilizer Guide," A. A. a. R. Development, Ed., ed, 2004.
[136] A. F. Mustafa, J. J. McKinnon, and D. A. Christensen, "Chemical
characterization and in vitro crude protein degradability of thin stillage
derived from barley- and wheat-based ethanol production," Animal Feed
Science and Technology, vol. 80, pp. 247-256, 1999.
[137] NCDENR, "Water Management Options," N. C. D. o. E. a. N. R. D. o. P. P. a.
E. Assistance, Ed., ed.
[138] Hart, Gangwer, Graham, and Marx, "Dairy manure as a fertilizer source," U.
S. D. o. Agriculture, Ed., ed: Oregon State University, 1997.
[139] A. Government. (2013, October 11, 2013). Agriculture Financial Services
Corporation (AFSC). Available: http://www.afsc.ca/home.aspx
[140] AUC, "Rate Schedule 2 Commercial Water Cost," A. U. Commission, Ed.,
ed. Edmonton, 2013, p. 5.
226
[141] J. R. Kwiatkowski, A. J. McAloon, F. Taylor, and D. B. Johnston, "Modeling
the process and costs of fuel ethanol production by the corn dry-grind
process," Industrial Crops and Products, vol. 23, pp. 288-296, 2006.
[142] AARD. (2013). Average Farm Input Prices for Alberta.
Available: http://www.agric.gov.ab.ca/app21/farminputprices
[143] Urbanchuk, "Economic Impacts On The Farm Community Of Cooperative
Ownership Of Ethanol Production," National Corn Growers
AssociationSeptember 8, 2006 2006.
[144] Tiffany, "Dry-Grind Ethanol Production: Economic Sensitivity," University of
Minnesota2003.
[145] Butterfield, "Farm-scale ethanol fuel production plant," C. D. o. F. a.
Agriculture, Ed., ed: Ethanol Fuels Program - Long Range Planning Unit.
[146] L. Farms.com. (2014 Distillers Grain Prices Don’t Decline As Fast
Available: http://www.farms.com/news/distillers-grain-prices-don-t-
decline-as-fast-68124.aspx
[147] AUC, "Electricity - Monthly Regulated Retail Option Rates," A. U.
Commission, Ed., ed. Edmonton, 2013, p. 5.
[148] CARD. (2013, Historical Ethanol Operating Margins. ARD is a center
located within the Department of Economics in the College of Agriculture
and Life Sciences at Iowa State University.
Available: http://www.card.iastate.edu/research/bio/tools/hist_eth_gm.as
px
227
[149] ASFC, "Weekly Interest Rates ", G. o. Alberta, Ed., ed, 2014.
[150] Simmons, Swales, Magnan, and Seyer. (2013). Insurance industry: key tax
rates and updates. Available: http://www.pwc.com/en_CA/ca/tax-
insights/publications/pwc-insurance-industry-key-tax-rates-and-updates-
canada-2013-09-en.pdf
[151] F. Frezatti and A. Andson Braga de, "Ebitda: Possible Impacts On Business
Management," Revista Universo Contabil, vol. 3, p. 7, 2007 2007.
[152] Brealey, Myers, and Merton, Principles f of Corporate Finance: McGraw-
Hill Companies, Inc, 2011.
[153] J. M. Douglas, Conceptual Design of Chemical Processes: McGraw-Hill,
New York, 1988.
[154] Nelson, "Why condensing steam turbines are more efficient than gas
turbines," in Tenth Annual Industrial Energy Technology Conference,
Houston, TX, 1988.
[155] Norwood, "Pinch Analysis Spreadsheet," IChemE, Ed., ed: Institution of
Chemical Engineers, 2009.
[156] OSHA, "Job Hazard Analysis ", U. S. D. o. Labor, Ed., ed, 2002.
[157] OSHA, "Hazardous (Classified) Locations," U. D. o. Labor, Ed., ed, 1996.
[158] Cole-Parmer, "Vacuum Pumps," C.-P. Canada, Ed., ed, 2014.
[159] Adams, "Co-generation system quote," Andrade, Ed., ed. Calgary, 2014.
[160] Anex, "Bucket Elevator Conveyors," A. M. Company, Ed., ed, 2014.
228
[161] D. Berner, "AGRI-PRESS® AGP-800," I. Press Technology & Mfg., Ed., ed,
2014.
[162] Cole-Parmer, "Centrifugal Pumps," C.-P. Canada, Ed., ed, 2014.
[163] N. Arifeen, R. Wang, I. Kookos, C. Webb, and A. A. Koutinas, "Optimization
and Cost Estimation of Novel Wheat Biorefining for Continuous Production
of Fermentation Feedstock," Biotechnology Progress, vol. 23, pp. 872-880,
2007.
[164] Andrade, "Whitefox Equipment Cross Reference. Internal Report," 2013.
[165] C. Alcohols. Material Safety Data Sheets [Online].
Available: http://www.comalc.com/sheets.php
229
Appendices
Appendix 1: McCabe Thiele diagrams for distillation column design for the base
design case
0
0.1
0.2
0.3
0.4
0.5
0.6
0.7
0.8
0.9
1
0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1
Etha
nol in
vap
or, y
EtO
H[m
ole
fract
ion]
Ethanol in liquid, xEtOH [mole fraction]
zF xD
slope =Rmin / (Rmin + 1)
xB
230
0
0.1
0.2
0.3
0.4
0.5
0.6
0.7
0.8
0.9
1
0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1
Etha
nol i
n va
por,
y EtO
H[m
ole
fract
ion]
Ethanol in liquid, xEtOH [mole fraction]
zF xDxB
Rectifying operating line
Equilibrium Curve
Stripping operating line
Feed Stage
231
0
0.1
0.2
0.3
0.4
0.5
0.6
0.7
0.8
0.9
1
0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1
Etha
nol i
n va
por,
y EtO
H[m
ole
fract
ion]
Ethanol in liquid, xEtOH [mole fraction]
xDxB zF
Minimum number of stages: 4
232
Appendix 2: McCabe Thiele diagrams for distillation column B for the optimized
design
0
0.1
0.2
0.3
0.4
0.5
0.6
0.7
0.8
0.9
1
0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1
Etha
nol i
n va
por,
y EtO
H[m
ole
fract
ion]
Ethanol in liquid, xEtOH [mole fraction]
zF xD
slope =Rmin / (Rmin + 1)
xB
233
0
0.1
0.2
0.3
0.4
0.5
0.6
0.7
0.8
0.9
1
0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1
Etha
nol i
n va
por,
y EtO
H[m
ole
fract
ion]
Ethanol in liquid, xEtOH [mole fraction]
zF xDxB
234
0
0.1
0.2
0.3
0.4
0.5
0.6
0.7
0.8
0.9
1
0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1
Etha
nol i
n va
por,
y EtO
H[m
ole
fract
ion]
Ethanol in liquid, xEtOH [mole fraction]
xDxB zF
Appendix 3: Operation Schedule of the IDF Ethanol Plant
1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29
Filling of pretreatment tank TK-208 A for batch 1
Filling of yeast conditioning vessel V-220 for batch 1
Yeast conditioning for batch 1
Evacuation of pretreatment tank TK-208 A from batch 1 / Filling of SSF reactor R-212 A for batch 1
Filling of pretreatment tank TK-208 B for batch 2
Filling of yeast conditioning vessel V-220 for batch 2
Yeast conditioning for batch 2
Evacuation of pretreatment tank TK-208 B from batch 2 / Filling of SSF reactor R-212 B for batch 2
Filling of pretreatment tank TK-208 C for batch 3
Addition of yeast and enzyme to SSF reactor R-212 A for batch 1
Fermentation of batch 1 in R-212 A
Filling of yeast conditioning vessel V-220 for batch 3
Yeast conditioning for batch 3
Evacuation of pretreatment tank TK-208 C from batch 3 / Filling of SSF reactor R-212 C for batch 3
Filling of pretreatment tank TK-208 D for batch 4
Addition of yeast and enzyme to SSF reactor R-212 B for batch 2
Fermentation of batch 2 in R-212 B
Filling of yeast conditioning vessel V-220 for batch 4
Yeast conditioning for batch 4
Evacuation of pretreatment tank TK-208 D from batch 4 / Filling of SSF reactor R-212 D for batch 4
Filling of pretreatment tank TK-208 A for batch 5
Addition of yeast and enzyme to SSF reactor R-212 C for batch 3
Fermentation of batch 3 in R-212 C
Filling of yeast conditioning vessel V-220 for batch 5
Yeast conditioning for batch 5
Evacuation of pretreatment tank TK-208 A from batch 5 / Filling of SSF reactor R-212 A for batch 5
Filling of pretreatment tank TK-208 B for batch 6
Addition of yeast and enzyme to SSF reactor R-212 D for batch 4
Fermentation of batch 4 in R-212 D
Evacuation of SSF reactor R-212 A from batch 1 / Filling of beer tank TK-215 with beer obtained from batch 1
Distillation and dehydration of batch 1
Filling of yeast conditioning vessel V-220 for batch 6
Yeast conditioning for batch 6
Evacuation of pretreatment tank TK-208 B from batch 6 / Filling of SSF reactor R-212 B for batch 6
Filling of pretreatment tank TK-208 C for batch 7
Addition of yeast and enzyme to SSF reactor R-212 A for batch 5
Fermentation of batch 5 in R-212 A
Evacuation of SSF reactor R-212 B from batch 2 / Filling of beer tank TK-215 with beer obtained from batch 2
Distillation and dehydration of batch 2
Filling of yeast conditioning vessel V-220 for batch 7
Yeast conditioning for batch 7
Evacuation of pretreatment tank TK-208 C from batch 7 / Filling of SSF reactor R-212 C for batch 7
Filling of pretreatment tank TK-208 D for batch 8
Addition of yeast and enzyme to SSF reactor R-212 B for batch 6
Fermentation of batch 6 in R-212 B
Evacuation of SSF reactor R-212 C from batch 3 / Filling of beer tank TK-215 with beer obtained from batch 3
Distillation and dehydration of batch 3
Filling of yeast conditioning vessel V-220 for batch 8
30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60 61 62 63 64 65 66 67 68 69 70 71 72 73 74 75 76 77 78 79 80 81
82 83 84 85 86 87 88 89 90 91 92 93 94 95 96 97 98 99 100 101 102 103 104 105 106 107 108 109 110 111 112 113 114 115 116 117 118 119 120 121 122 123 124 125 126 127 128 129 130 131 132 133
245
134 135 136 137 138 139 140 141 142 143 144 145 146 147 148 149 150 151 152 153 154 155 156 157 158 159 160 161 162 163 164 165 166 167 168 169 170 171 172 173 174 175 176 177 178 179 180 181 182 183 184 185
186 187 188 189 190 191 192
236
Appendix 4: Distillation columns and absorber sizing
The calculations shown below correspond to the distillation column B (T-228). The
stripper column A (T-227) and the CO2 absorber (T-218) were sized following the
same procedure. The parameters used in the calculations are summarized in the
table below:
Table 59: Parameters used for the distillation section sizing
Parameter Equipment Symbol Value Units
Theoretical number of stages
T-227 NT
7 T-228 8 T-218 4
Tray efficiency T-227
E0 70 % T-228 T-218
F-factor T-227
1.22 (kg/ms2)1/2 T-228 T-218
Maximum vapor density T-227
ρv 2.50
kg/m3 T-228 1.44 T-218 1.67
Maximum vapor rate T-227 65.21
m3/h T-228 84.16 T-218 54.89
The column height was obtained as follows:
Lc = 0.7315 (NT)theoretical / E0
Lc = 0.7315 x 8 / 0.7
Lc = 8.4 m
237
The diameter of the column was calculated using the maximum allowable
velocity:
vmax = F-factor�ρv
vmax = 1.22 Kg1/2
m1/2s
�1.44 Kgm3
vmax = 1.02 m/s
Area = maximum vapor rate
maximum allowable velocity = 84.16 m3
h × h3600 s
1.02 ms
= 0.023 m2
D = �4Aπ
D = �4×0.023 m2
π = 0.17 m
238
Appendix 5: Stripper distillate flash drum sizing
The calculations performed for the sizing of the flashing drum (D-230) are
presented below:
vmax = F-factor�ρv
vmax = 0.61 Kg1/2
m1/2s� 3.83 Kg/m3
vmax = 0.31 m/s
Area = vapour volumetric flow ratemaximum vapour velocity
Area = 17.12 m3
h × h3600 s
0.31 ms
=0.015 m2
Dmin = �4×0.015 m2
π =0.14 m
Once the minimum diameter of the vessel is obtained, the liquid holdup
requirements were considered. In order to have a vessel half full and a holdup
time of 5 min, the volume of the vessel is:
V = 2 ×5 min×h
60 min × 0.253 m3
h = 0.042 m3
239
The diameter was then calculated with an aspect ratio (length to diameter) of
2:
L = 2D
V = A×L = πD2
4 ×L = πD2
4 ×2D = πD3
2
0.042 m3 = πD3
2
D = �2×0.042 m3
π
3
= 0.30 m
The diameter obtained was larger than the one calculated through the F-factor,
so the dimensions specified for the flash drum were:
D = 0.30 m
L = 0.60 m
240
Appendix 6: Knock-out drum sizing
The calculations performed for the sizing of the reflux knock-out drum (D-233)
and the biogas dewatering knock-out drum (D-253) were based on 5 minutes of
residence time. The reflux knock-out drum calculations are presented below as
an example:
V = 2 ×5 min×h
60 min × 0.16 m3
h = 0.026 m3
The diameter was then calculated with an aspect ratio (length to diameter) of
2:
L = 2D
V = A×L = πD2
4 ×L = πD2
4 ×2D = πD3
2
0.026 m3 = πD3
2
D = �2×0.026 m3
π
3
= 0.26 m
The dimensions specified for the knock out drum are:
D = 0.30 m
L = 0.60 m
241
Appendix 7: SSF set sizing
The pre-treatment tanks (TK-208) and the simultaneous saccharification and
fermentation reactors (R-212) were sized based on the beer volume fed to the
distillation tower each 24 hours:
1.058 m3 beer
h ×24 h = 25.4 m3 beer
Therefore, the volume of each fermenter and pre-treatment tank was 25.4 m3.
The dimensions of the fermenter and pre-treatment tank were calculated as
follows:
Vconical portion = 13 π
D2
4 h
Vcylindrical portion = πD2
4 H
Vtotal = πD2
4 H + 13 π
D2
4 h
H = 23 L
h = 13 L
H = 1.5D
242
Vtotal = π×� 2 L3 x1.5�
2
4 ×23 L +
13 π×
� 2 L3 x1.5�
2
4 ×13 L = 0.1206 L3
25.4 m3 = 0.1206 L3
L = 6 m
H = 4 m
h = 2 m
D = 2.7 m
243
Appendix 8: Vessels and tanks sizing
For vessels and tanks, the residence time was used to obtain the volume of the
equipment. The calculations for the anaerobic digester (R-250) are presented as
example:
V = τV̇
V = 12 days ×24 h
1 day ×1.103 m3
h = 318 m3
For the conical vessels, the dimensions were obtained the same way as in the
case of the fermenters.
244
Appendix 9: Heat exchangers sizing
For heat exchanger, the heat transfer coefficient, U, collected from research
and industrial sources, was used to obtain the heat transfer areas. The UA values
were obtained from the steady state simulation. The calculations for the
distillation column B preheater (E-226) are presented as example:
UA = 0.88 KW / K
U = 0.31 KW / m2K
A = 2.81 m2
245
Appendix 10: Pumps sizing
The pumps were sized according to the total head and the capacity in terms of
volumetric flow. The calculations performed for the enzyme pump (P-224) are
presented as example:
The static head was calculated as the difference between the surface of the
suction tank and the point of discharge:
The friction losses were estimated as 25% of the static head:
Total Head = 5.45 m + 0.25 × 5.45 m = 6.8 m = 22 ft
To transform 22 ft to pressure units:
Head = P x 2.31
SG
0.55
m
V-223
6 m
1
2
5.45
m
R-212
246
22 ft = P x 2.310.998
P = 9.5 psi
The energy requirement of the pumps was calculated as follows:
HP = GPM x Head
1715 x E
HP = 0.0346 x 9.51715 x 0.75 = 2.55 x 10-4
247
Appendix 9: Vacuum pump P-244 performance curve [158]
248
Appendix 10: Hydrated Meal Pump (P-209) Performance Curve [162]
249
Appendix 11: SSF Reactor Pump (P-213) Performance Curve [162]
250
Appendix 12: CO2 Absorber Pump (P-219) Performance Curve [162]
251
Appendix 13: Yeas Slurry Pump (P-222) Performance Curve [162]
252
Appendix 14: Enzyme Pump (P-224) Performance Curve [162]
253
Appendix 15: Permeate Pump (P-245) Performance Curve [162]
254
Appendix 16: Thin Stillage Pump (P-249) Performance Curve [162]
255
Appendix 17: Equipment costing calculations
EQUIPMENT INFORMATION QUOTE ECONOMIC VALUES COST SCALING EQUIPMENT COST EQ
PT ID
EQUIPMENT TITLE
NUM
. REQ
UIRE
D
YEA
R O
F Q
UOTE
CEPCI IN
QUOTE YEAR
PURCHASE COST [USD]
SCALING VARIABLE
SCA
LIN
G V
ALU
E
SCA
LIN
G E
XP
NEW
SC
ALI
NG
VA
RIA
BLE
VALU
E
SIZE
RA
TIO
SCALED PURCHASE COST [USD]
SCALED COST IN PROJECT YEAR [USD]
ROUNDED SCALED COST IN PROJECT
YEAR [USD]
MI-203 Hammer Mill 1 2007 525.4 103,000 [163] Feed flow [kg/h] 37,669.00 0.6 358.83 0.010 $ 6,312.19 $ 7,670.98 $ 8,000 TK-204 Meal Storage Tank 1 2012 584.6 24,679 [164] Volume [m3] 5.80 0.7 32.55 5.6 $ 82,548.98 $ 90,159.98 $ 90,000 U-205 Hopper Scale 1 $ 10,300.00 $ 10,000
MC-206 Meal Conveyor 1 $ 19,910.00 $ 20,000
E-207 Pre-treatment Water Heater / Stillage Cooler 1 2013 638.5 1,402 [164] Area [m2] 0.30 0.7 0.33 1.1 $ 1,498.73 $ 1,498.73 $ 1,000
TK-208 A Pre-treatment Tank A 1 2012 584.6 24,679 [164] Volume [m3] 5.80 0.7 26.00 4.5 $ 70,535.44 $ 77,038.79 $ 77,000 TK-208 B Pre-treatment Tank B 1 2012 584.6 24,679 [164] Volume [m3] 5.80 0.7 26.00 4.5 $ 70,535.44 $ 77,038.79 $ 77,000 TK-208 C Pre-treatment Tank C 1 2012 584.6 24,679 [164] Volume [m3] 5.80 0.7 26.00 4.5 $ 70,535.44 $ 77,038.79 $ 77,000 TK-208 D Pre-treatment Tank D 1 2012 584.6 24,679 [164] Volume [m3] 5.80 0.7 26.00 4.5 $ 70,535.44 $ 77,038.79 $ 77,000 P-209 A Hydrated Meal Pump A 1 $ 174.00 $ 200 P-209 B Hydrated Meal Pump B 1 $ 174.00 $ 200 P-209 C Hydrated Meal Pump C 1 $ 174.00 $ 200 P-209 D Hydrated Meal Pump D 1 $ 174.00 $ 200
E-210 Pre-treatment Tank Side Heater 1 2013 638.5 1,402 [164] Area [m2] 0.30 0.7 0.06 0.2 $ 463.93 $ 463.93 $ 500
E-211 Hydrated Meal Cooler 1 2012 584.6 4,000 [164] Area [m2] 2.42 0.7 2.42 1.0 $ 4,000.00 $ 4,368.80 $ 4,000 R-212 A SSF Reactor A 1 2012 584.6 24,679 [164] Volume [m3] 5.80 0.7 25.40 4.4 $ 69,392.04 $ 75,789.97 $ 76,000 R-212 B SSF Reactor B 1 2012 584.6 24,679 [164] Volume [m3] 5.80 0.7 25.40 4.4 $ 69,392.04 $ 75,789.97 $ 76,000 R-212 C SSF Reactor C 1 2012 584.6 24,679 [164] Volume [m3] 5.80 0.7 25.40 4.4 $ 69,392.04 $ 75,789.97 $ 76,000 R-212 D SSF Reactor D 1 2012 584.6 24,679 [164] Volume [m3] 5.80 0.7 25.40 4.4 $ 69,392.04 $ 75,789.97 $ 76,000 P-213 A SSF Reactor Pump A 1 $ 174.00 $ 200 P-213 B SSF Reactor Pump B 1 $ 174.00 $ 200 P-213 C SSF Reactor Pump A 1 $ 174.00 $ 200 P-213 D SSF Reactor Pump A 1 $ 174.00 $ 200 E-214 A SSF Reactor Cooler A 1 2012 584.6 1,125 [164] Area [m2] 0.10 0.5 0.000297 0.003 $ 61.31 $ 66.96 $ 70 E-214 B SSF Reactor Cooler B 1 2012 584.6 1,125 [164] Area [m2] 0.10 0.5 0.000297 0.003 $ 61.31 $ 66.96 $ 70 E-214 C SSF Reactor Cooler C 1 2012 584.6 1,125 [164] Area [m2] 0.10 0.5 0.000297 0.003 $ 61.31 $ 66.96 $ 70 E-214 D SSF Reactor Cooler D 1 2012 584.6 1,125 [164] Area [m2] 0.10 0.5 0.000297 0.003 $ 61.31 $ 66.96 $ 70 TK-215 Beer Well 1 2012 584.6 24,679 [164] Volume [m3] 5.80 0.7 34.09000 5.9 $ 85,263.85 $ 93,125.16 $ 93,130
AG-216 Beer Well Agitator 1 2012 584.6 8,056 [164] Tank volume [m3] 1.53 0.5 34.09 22.3 $ 38,003.30 $ 41,507.20 $ 42,000 P-217 Beer Pump 1 2012 584.6 6,625 [164] Total head [ft] 380.00 0.8 249.00 0.7 $ 4,724.09 $ 5,159.65 $ 5,000 T-218 CO2 Absorber 1 2012 584.6 33,486 [164] Height [m] 7.70 0.6 4.20 0.5 $ 23,276.54 $ 25,422.63 $ 25,000 P-219 CO2 Absorber Pump 1 $ 174.00 $ 200 V-220 Yeast Conditioning Tank 1 2012 584.6 2,175 [164] Volume [m3] 0.64 0.7 0.41 0.6 $ 1,592.51 $ 1,739.34 $ 2,000
E-221 Conditioning Vessel Water Heater 1 2012 584.6 1,125 [164] Area [m2] 0.10 0.5 0.00138 0.0 $ 132.16 $ 144.34 $ 140
P-222 Yeast Slurry Pump 1 $ 174.00 $ 200 V-223 Enzyme Storage Vessel 1 2012 584.6 2,175 [164] Volume [m3] 0.64 0.7 0.48 0.8 $ 1,778.29 $ 1,942.25 $ 2,000 P-224 Enzyme Pump 1 $ 174.00 $ 200 E-225 Stripper Column A Preheater 1 2012 584.6 13,449 [164] Area [m2] 3.00 0.7 4.9 1.6 $ 18,878.69 $ 20,619.30 $ 21,000
256
EQUIPMENT INFORMATION QUOTE ECONOMIC VALUES COST SCALING EQUIPMENT COST EQ
PT ID
EQUIPMENT TITLE
NUM
. REQ
UIRE
D
YEA
R O
F Q
UOTE
CEPCI IN
QUOTE YEAR
PURCHASE COST [USD]
SCALING VARIABLE
SCA
LIN
G V
ALU
E
SCA
LIN
G E
XP
NEW
SC
ALI
NG
VA
RIA
BLE
VALU
E
SIZE
RA
TIO
SCALED PURCHASE COST [USD]
SCALED COST IN PROJECT YEAR [USD]
ROUNDED SCALED COST IN PROJECT
YEAR [USD]
E-226 Distillation Column B Preheater 1 2012 584.6 13,449 [164] Area [m2] 3.00 0.7 2.8 0.9 $ 12,846.83 $ 14,031.31 $ 14,000
T-227 Stripper Column A 1 2012 584.6 33,486 [164] Heigth [m] 7.70 0.6 7.3 0.9 $ 32,431.17 $ 35,421.31 $ 35,000 T-228 Distillation Column B 1 2012 584.6 33,486 [164] Heigth [m] 7.70 0.6 8.4 1.1 $ 35,280.64 $ 38,533.50 $ 39,000 E-229 Distillation Column B Reboiler 1 2012 584.6 33,656 [164] Area [m2] 19.00 0.7 8.4 0.4 $ 19,007.74 $ 20,760.25 $ 21,000
D-230 Stripper Distillate Flashing Drum 1 2012 584.6 2,715 [164] Volume [m3] 0.01 0.7 0.0 4.1 $ 7,311.80 $ 7,985.94 $ 8,000
P-231 Stripper Distillate Pump 1 2012 584.6 6,625 [164] Total head [ft] 380.00 0.8 239.0 0.6 $ 4,571.70 $ 4,993.21 $ 5,000 E-232 Reflux Condenser 1 2012 584.6 1,700 [164] Area [m2] 0.40 0.7 0.7 1.9 $ 2,614.98 $ 2,856.09 $ 3,000 D-233 Reflux Knock-Out Drum 1 2012 584.6 2,715 [164] Volume [m3] 0.01 0.7 0.026 2.6 $ 5,287.10 $ 5,774.57 $ 6,000 P-234 Reflux Pump 1 2012 584.6 6,625 [164] Total head [ft] 380.00 0.8 169.4 0.4 $ 3,470.79 $ 3,790.79 $ 4,000
Dehydration Package 1 $ 142,700 E-235 Distillate Superheater 1
E-236 Stillage from Distillation Column B Cooler 1
F-237 Membrane Module Feed Filter 1
M-238 Dehydration Membrane Module 1
E-239 Retentate Condenser 1 E-240 Retentate Cooler 1 TK-241 Retentate Product Tank 1 E-242 Permeate Condenser 1 V-243 Permeate Vessel 1 P-244 Vacuum Pump 1 P-245 Permeate Pump 1 SP-246 Whole Stillage Screw Press 1 $ 60,000.00 $ 60,000 TK-247 Thin Stillage Tank 1 2012 584.6 24,679 [164] Volume [m3] 5.80 0.7 3.9 0.7 $ 18,524.59 $ 20,232.55 $ 20,000 TK-248 Wet Grain Tank 1 2012 584.6 24,679 [164] Volume [m3] 5.80 0.7 107.0 18.4 $ 189,886.75 $ 207,394.27 $ 207,000 P-249 Thin Stillage Pump 1 $ 174.00 $ 200 R-250 Anaerobic Digester 1 $ 500,000 P-251 Anaerobic Digester Pump 1 2012 584.6 6,295 [164] Total head [ft] 61.00 0.8 82.5 1.4 $ 8,016.04 $ 8,755.12 $ 9,000 E-252 Anaerobic Digester Heater 1 2012 584.6 1,125 [164] Area [m2] 0.10 0.7 0.0 0.2 $ 364.65 $ 398.27 $ 400
D-253 Biogas Dewatering Knock-Out Drum 1 2012 584.6 2,715 [164] Volume [m3] 0.01 0.7 0.0 0.0 $ 145.52 $ 158.94 $ 200
BO-254 Biogas Boiler 1 $ 220,000 TB-255 Steam Turbine 1
G-256 Electricity Generator 1 Total Equipment Cost $ 2,232,750
257
Appendix 17: Cumulative cash flows for 10 years of project life time
Year Outflows Inflows Net inflows / outflows
Cumulative net inflows / outflows
0 $ (4,100,000.00) $ - $ (4,100,000.00) $ (4,100,000.00) 1 $ - $ 1,021,000.00 $ 1,021,000.00 $ (3,079,000.00) 2 $ - $ 1,021,000.00 $ 1,021,000.00 $ (2,058,000.00) 3 $ - $ 1,021,000.00 $ 1,021,000.00 $ (1,037,000.00) 4 $ - $ 1,021,000.00 $ 1,021,000.00 $ (16,000.00) 5 $ - $ 1,021,000.00 $ 1,021,000.00 $ 1,005,000.00 6 $ - $ 1,021,000.00 $ 1,021,000.00 $ 2,026,000.00 7 $ - $ 1,021,000.00 $ 1,021,000.00 $ 3,047,000.00 8 $ - $ 1,021,000.00 $ 1,021,000.00 $ 4,068,000.00 9 $ - $ 1,021,000.00 $ 1,021,000.00 $ 5,089,000.00
10 $ - $ 1,021,000.00 $ 1,021,000.00 $ 6,110,000.00
258
Appendix 18: Sensitivity analysis results
Ethanol Price Change ∆ Gross Profit ∆ EBITDA ∆ Net Profit ROI Payback Period
-50 % 19% 20% 10% 32% 4 -25 % 20% 22% 12% 35% 3 0 % 22% 24% 14% 38% 3
25 % 23% 25% 15% 41% 3 50 % 24% 27% 17% 44% 3
Fertilizer Price Change ∆ Gross Profit ∆ EBITDA ∆ Net Profit ROI Payback Period
-50 % 14% 16% 8% 40% 3 -25 % 18% 20% 11% 39% 3 0 % 22% 24% 14% 38% 3 25 % 25% 27% 17% 37% 3 50 % 29% 31% 20% 37% 3
Wheat Price Change ∆ Gross Profit ∆ EBITDA ∆ Net Profit ROI Payback Period
-50 25% 29% 18% 38% 3 -25 23% 26% 16% 38% 3 0 22% 24% 14% 38% 3
25 20% 21% 12% 38% 3 50 18% 19% 10% 38% 3
259
Appendix 19: MSDS of anhydrous ethanol [165]
260
261
262
263
