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RETROFITTING ANALYSIS TO INTEGRATE THE MIXALCO PROCESS
TO THE CRUDE OIL DISTILLATION PROCESS
Thesis
By
LAURA PATRICIA PRADA VILLAMIZAR
Submitted to the Office of Graduate Studies of
Universidad de Los Andes
In partial fulfillment of the requirements for the degree of
M.SC. CHEMICAL ENGINEERING
August 2013
Major Subject: Chemical Engineering
Retrofitting analysis to integrate the MixAlco process to the crude oil distillation
process
Copyright 2013 Laura Patricia Prada Villamizar
RETROFITTING ANALYSIS TO INTEGRATE THE MIXALCO PROCESS
TO THE CRUDE OIL DISTILLATION PROCESS
Thesis
By
LAURA PATRICIA PRADA VILLAMIZAR
Submitted to the Office of Graduate Studies of
Universidad de Los Andes
In partial fulfillment of the requirements for the degree of
M.SC. CHEMICAL ENGINEERING
Approved by:
Chair of committee, Roco Sierra Ramrez, PhD.
Committee Members, Jorge Mario Gmez, Phd.
Head of Department, Oscar Alvarez Solano, PhD.
August 2013
Major Subject: Chemical Engineering
i
ABSTRACT
Retrofitting analysis to integrate the MixAlco process to the crude oil distillation
process (August 2013)
Laura Patricia Prada Villamizar, Universidad de los Andes, Colombia
Advisor: Roco Sierra Ramrez, Ph.D.
The MixAlco technology comprises a processing facility to produce liquid
transportation fuels and/or value-added chemicals from biomass resources; however,
build and run a new MixAlco plant may be very costly. On the other hand, high quality
and easily exploitable fossil-fuels resources inevitably dwindle worldwide. Both the
preservation of high quality fossil-fuel resources and the feasibility of a MixAlco plant
can be importantly enhanced by retrofitting the MixAlco process into an existing
fossil-fuel processing facility. This retrofitting is attainable because both processes have
similar products (bio-gasoline, gasoline, bio-jet, jet). This work assesses a retrofitting
analysis to integrate the MixAlco process to a selected case of crude oil distillation
process (CODP). The proposed methodology suggests a hierarchy cost using the
following tools: process simulations, mass and energy integrations, and economic
evaluations. The work starts by assessing improvements for a base case of each of the
two involved plants separately. Then, comparisons between base cases and the
retrofitting of both processes (the resulting plant is regarded here as integrated bio-
refinery) is made. The most remarkable result was a Net Present Value (NPV)
ii
increment from MM USD 7.30 to MM USD15.7, and Return On Investment (ROI)
increment from 11.1% to 12.4% for MixAlco process.
iii
RESUMEN
Retrofitting analysis to integrate the MixAlco process to the crude oil distillation
process (August 2013)
Laura Patricia Prada Villamizar, Universidad de los Andes, Colombia
Advisor: Roco Sierra Ramrez, Ph.D.
MixAlco es una tecnologa donde se producen combustibles lquidos de
transporte y / o productos qumicos de valor agregado a partir de fuentes de biomasa, sin
embargo, construir y operar una planta nueva de MixAlco puede ser muy costoso. Por
otro lado, los recursos combustibles fsiles de alta calidad y fcilmente explotables
disminuyen en todo el mundo. La preservacin de los combustibles fsiles de alta
calidad y la viabilidad de una planta MixAlco, pueden mejorarse mediante la
integracin del proceso MixAlco en una instalacin existente de procesamiento de
combustibles fsiles.
Esta integracin es posible gracias a que ambos procesos tienen productos
similares (bio-gasolina, gasolina, bio-jet, jet). Este trabajo evala un anlisis de
integracin entre el proceso de MixAlco con el proceso de destilacin de crudo de
petrleo (PDCP) como caso seleccionado. La metodologa propuesta sugiere una
jerarqua de costos con las siguientes herramientas: simulacin de procesos,
integraciones de masa y energa, y evaluaciones econmicas. Este trabajo inicia
evaluando mejoras para un caso base en cada una de las plantas involucradas por
iv
separado. Despus, se hacen comparaciones entre los casos base y la integracin de los
dos procesos (la planta resultante se considera como "bio-refinera integrada"). El
resultado ms importante para el proceso de MixAlco presenta un incremento en el
Valor Presente Neto (VPN) de MM USD 7.30 a MM USD 15.7 y un incremento en la
tasa interna de retorno de la inversin (TIR) de 11.1% a 12.4%
v
ACKNOWLEDGEMENTS
I would like to thank my family for the love, belief, and support they have
provided me throughout my life, especially to my mother, Laura Villamizar. She gave
me much love and support, and thanks to my two brothers Dany and Sergio. I would like
to thank my two big loves Guillermo and Santiago, for their compression, support and
company all the time, especially during this work.
I would to express my deepest gratitude to Dr. Roco Sierra, for her guidance,
and for her patience throughout this work. Thank you for all support during my graduate
study. I would like also to thank to all her group members for all their support and help.
I would like also to thank Cesar Mahecha for the support that he provide me
during this work.
vi
NOMENCLATURE
AEA: Aspen Energy Analyzer
AFC: Annualized Fixed Cost
AGO: Atmospheric Gas Oil
APEA: Aspen Process Economic Analyzer
API: Standard API gravity
BPD: Barrel Per Day
C: Cooler
CE: Chemical Engineering Plant Cost Index
CM: Compressor
CODP: Crude Oil Distillation Process
CON: Conveyor
CSTR: Continually Stirred Tank Reactors
DHFORM: Formation Enthalpy
DW&B: Direct Wage and Benefits
E: Heat Exchanger
EIA: US Energy Information Administration
FCI: Fixed Capital Investment
FOB: Free On Board
FOC: Fixed Operating Cost
vii
GAL: U.S liquid gallon, (231 in3)
GCC: Grand Composite Curve
H: Heater
HEN: Heat Exchange Network
HVGO: Heavy Vacuum Gas Oil
IRR: Internal Rate of Return
INHSPCD: In-house Pure Component Database
LVGO: Light Vacuum Gas Oil
M: Mixer
MACRS: Modified Accelerated Cost Recovery System
MOC: Minimum Operating Cost
MM: Million
MR: Cumulative Mass Lost
MS: Marshall and Swift Cost Index
MSA: Mass-Separating Agent
MTAC: Minimizing Total Annualized Cost
MW: Molecular weight
MW&B: Maintenance Wages and Benefits
Net_G: Net Generation
NF: Nelson-Farrer Refinery Construction Index
viii
NPV: Net Present Value
NREL: National Renewable Energy Laboratory
NRTL: Non-random-two-liquid
P: Pump
PBP: Payback Period
PSA: Pressure Swing Adsorption
R: Reactor
ROI: Return On Investment
RKS: Redlich-Kwong-Soave
S: Splitter
SCF: Standard Cubic Foot
SG: Standard specific gravity at 60F
SP: Separator
T: Distillation tower
TBP: True Normal boiling point
TCI: Total Capital Investment
TEHL: Table of Exchangeable Heat Loads
TID: Temperature-Interval Diagram
TK: Tank
TR: Turbine
ix
TON: Metric ton (1,000kg)
USD: United States dollars
VFAs: Volatile Fatty Acids
VLSTD: Standard Liquid MolarVolume at 60F
VOC: Variable Operating Cost
VP: Venture Profit
VS: Volatile Solids
WCI: Working Capital Investment
WWT: Waste Water Treatment
ZC: Critical Compressibility Factor
x
TABLE OF CONTENTS
Page
ABSTRACT .............................................................................................................. i
RESUMEN ................................................................................................................ iii
ACKNOWLEDGEMENTS ...................................................................................... v
NOMENCLATURE .................................................................................................. vi
TABLE OF CONTENTS .......................................................................................... x
LIST OF FIGURES ................................................................................................... xiii
LIST OF TABLES .................................................................................................... xv
1. INTRODUCTION ............................................................................................... 1
2. OBJECTIVES ..................................................................................................... 5
2.1 General objective ..................................................................................... 5
2.2 Specific objectives ................................................................................... 5
3. METHODOLOGY .............................................................................................. 6
3.1 Description of the proposed methodology .............................................. 6
3.1.1 Define needs ................................................................................. 6
3.1.2 Process arrangements .................................................................. 7
3.1.3 Feasibility .................................................................................... 7
3.2 Simulation Tools ........................................................................................... 9
3.2.1 MixAlco process base case ........................................................ 9
3.2.2 CODP base case ............................................................................ 11
3.2.3 MixAlco process and CODP retrofitted plant ............................ 14
xi
Page
3.3 Process integration ........................................................................................ 14
3.3.1 Material rerouting .......................................................................... 15
3.3.2 Heat Exchanger Network (HEN) .................................................. 16
3.3.3 Cost Analysis................................................................................. 16
4. RESULTS AND DISCUSSION ................................................................... 19
4.1 Simulation results .................................................................................... 19
4.1.1 Simulation builds up and results for MixAlco base case ......... 19
4.1.1.1 MixAlco block description .......................................... 21
4.1.1.2 MixAlco overall mass balance results ......................... 54
4.1.1.3 MixAlco overall heat balance results .......................... 56
4.1.2 Simulation builds up and results for CODP base case ................ 57
4.1.2.1 CODP Block description ................................................ 58
4.1.2.2 CODP overall mass balance results ............................... 68
4.1.2.3 CODP overall heat balance results ................................. 69
4.2 Define needs ............................................................................................ 69
4.3 Retrofitting procedure applied: Process arrangements ............................ 70
4.3.1 Internal rearrangements ............................................................... 70
4.3.1.1 MixAlco process ......................................................... 70
4.3.1.2 CODP ............................................................................. 79
4.3.2 Internal modifications ................................................................. 84
4.3.2.1 MixAlco process ......................................................... 84
xii
Page
4.3.2.2 CODP ............................................................................................ 95
4.3.3 External modifications ................................................................ 103
4.3.3.1 Case 1 ............................................................................. 105
4.3.3.2 Case 2 ............................................................................. 110
4.3.3.3 Comparison between cases ............................................. 123
4.4 Sensitivity analysis .................................................................................. 125
4.4.1 Variation of gasoline prices ......................................................... 125
4.4.2 Variation of Jet prices ................................................................. 126
4.4.3 Variation of Biomass prices ........................................................ 127
4.4.4 Variation of MixAlco plant capacity ........................................ 127
CONCLUSIONS ....................................................................................................... 129
RECOMMENDATIONS AND FUTURE WORK ................................................... 132
REFERENCES .......................................................................................................... 134
APPENDIX A ........................................................................................................... 137
APPENDIX B ........................................................................................................... 153
APPENDIX C ........................................................................................................... 157
APPENDIX D ........................................................................................................... 159
APPENDIX E ............................................................................................................ 166
VITA ................................................................................................................ 177
xiii
LIST OF FIGURES
FIGURE Page
1-1 Pathways for converting biomass to hydrocarbon fuels .......................... 2
3-1 Flowchart of the proposed methodology ................................................. 8
3-2 Crude oil distillation TPB ....................................................................... 13
4-1 Blocks of MixAlco process simulation. ............................................... 20
4-2 Feed handling simulation ........................................................................ 23
4-3 Pretreatment simulation ............................................................................ 25
4-4 Fermentation simulation .......................................................................... 29
4-5 Dewatering simulation ............................................................................ 34
4-6 Ketonization simulation .......................................................................... 39
4-7 Lime kiln simulation ............................................................................... 43
4-8 Final simulation ....................................................................................... 46
4-9 Distillation curve for gasoline ................................................................. 49
4-10 Distillation curve for Jet ............................................................................. 50
4-11 Gasification simulation... ........................................................................... 51
4-12 Blocks of CODP simulation ....................................................................... 57
4-13 First pre-heating train ................................................................................. 58
4-14 Second pre-heating train ............................................................................. 61
4-15 Atmospheric distillation column ............................................................... 63
4-16 Vacuum distillation column ....................................................................... 66
4-17 Mass integration for MixAlco process .................................................... 73
xiv
FIGURE Page
4-18 Power integration for MixAlco process .................................................. 73
4-19 Heat integration in Reactors for MixAlco process .................................. 74
4-20 Cash flow for MixAlco process in the base case .................................... 79
4-21 Hot and Cold composite for MixAlco HEN............................................ 87
4-22 Grand composite curve for MixAlco HEN ............................................. 87
4-23 Grid diagram for MixAlco HEN ............................................................. 89
4-24 Cash flow for MixAlco process with HEN ............................................. 95
4-25 Hot and Cold composite for CODP HEN .................................................. 97
4-26 Grand composite curve for CODP HEN .................................................... 97
4-27 Grid diagram for CODP HEN .................................................................... 99
4-28 MixAlco and CODP simulation integrated ............................................. 104
4-29 Cash flow of MixAlco process in case 1 ................................................. 109
4-30 Hot and Cold composite for case 2 ............................................................ 112
4-31 Grand composite curve for case 2 .............................................................. 113
4-32 Grid diagram for case 2 .............................................................................. 115
4-33 Cash flow of MixAlco process in case 2 ................................................. 123
4-34 Variation of gasoline price for MixAlco process .................................... 125
4-35 Variation of Jet price for MixAlco process ............................................. 126
4-36 Variation of Biomass price ......................................................................... 127
4-37 Variation of MixAlco plant capacity ....................................................... 128
xv
LIST OF TABLES
TABLE Page
3-1 Biomass feed composition for MixAlco process .................................... 10
3-2 MixAlco operating conditions ................................................................. 11
3-3 Assay data for crude oil .............................................................................. 12
3-4 Assay data for crude oil Light ends ............................................................ 12
3-5 CODP operating conditions ....................................................................... 13
3-6 Feedstock, utilities and product prices ....................................................... 17
4-1 Feed handling mass and heat balance ........................................................ 24
4-2 Heat balances for Feed Handling equipment ............................................. 24
4-3 Pretreatment mass and heat balance ........................................................... 27
4-4 Fermentation mass and heat balance .......................................................... 30
4-5 Heat balances for Pretreatment and Fermentation equipments .................. 33
4-6 Dewatering mass and heat balance ............................................................. 35
4-7 Heat balances for Dewatering equipments ................................................. 38
4-8 Ketonization mass and heat balance ........................................................... 40
4-9 Heat balances for Ketonization equipments ............................................... 42
4-10 Heat balances for Lime kiln equipments .................................................... 43
4-11 Lime kiln mass and heat balance ................................................................ 44
4-12 Final mass and heat balance ....................................................................... 46
4-13 Heat balances for Final equipments ........................................................... 49
xvi
TABLE Page
4-14 Gasification mass and heat balance ............................................................ 52
4-15 Heat balances for Gasification equipments ................................................ 53
4-16 MixAlco yields ........................................................................................ 54
4-17 Summary fo heat balances for MixAlco processs ................................... 56
4-18 First train preheating mass and heat balance .............................................. 59
4-19 Heat balances for equipments in 1st preheating train ................................ 60
4-20 Second train preheating mass and heat balance ......................................... 62
4-21 Heat balances for equipments in 2nd preheating train ............................... 62
4-22 Atmospheric distillation mass and heat balance ........................................ 64
4-23 Heat balances for equipments in atmospheric distillation unit .................. 65
4-24 Vacuum distillation mass and heat balance ................................................ 67
4-25 Heat balances for equipments in vacuum distillation unit ......................... 68
4-26 CODP Yields .............................................................................................. 68
4-27 Overall heat balances for CODP ................................................................ 69
4-28 Fresh MixAlco streams ........................................................................... 71
4-29 Waste MixAlco streams .......................................................................... 72
4-30 VOC of MixAlco process in base case.................................................... 75
4-31 FOC of MixAlco process in base case .................................................... 76
4-32 FIC of MixAlco process in base case ...................................................... 77
4-33 Summary MixAlco economic results in base case .................................. 77
4-34 Fresh CODP streams .................................................................................. 80
xvii
TABLE Page
4-35 Waste CODP Streams ................................................................................ 81
4-36 FCI for CODP in base case ........................................................................ 82
4-37 VOC for CODP in base case ...................................................................... 82
4-38 FOC for CODP in base case ....................................................................... 83
4-39 MixAlco process streams for HEN ......................................................... 85
4-40 Heat integration for MixAlco process ..................................................... 88
4-41 Heat exchangers for MixAlco HEN ........................................................ 90
4-42 Coolers for MixAlco HEN ...................................................................... 90
4-43 Heaters for MixAlco HEN ...................................................................... 91
4-44 VOC of MixAlco process with HEN ...................................................... 92
4-45 FOC of MixAlco process with HEN ....................................................... 93
4-46 Summary MixAlco economic results with HEN..................................... 94
4-47 Process streams for CODP ......................................................................... 96
4-48 Summary of HEN cases for CODP ............................................................ 98
4-49 Heat exchangers in the best CODP case .................................................... 100
4-50 Coolers in the best CODP case .................................................................. 101
4-51 Heaters in the best CODP case ................................................................... 102
4-52 VOC of CODP with HEN .......................................................................... 102
4-53 Heat integration for MixAlco and CODP in case 1 ................................ 106
4-54 VOC of MixAlco process in case 1 ......................................................... 106
4-55 FOC of MixAlco process in case 1 ......................................................... 107
xviii
TABLE Page
4-56 Summary MixAlco economic results in case 1 ....................................... 108
4-57 Process streams for case 2 .......................................................................... 110
4-58 Heat integration for MixAlco and CODP case 2..................................... 114
4-59 Heat exchangers for case 2 ......................................................................... 115
4-60 Coolers for case 2 ....................................................................................... 118
4-61 Heaters for case 2 ....................................................................................... 119
4-62 VOC of MixAlco process in case 2 ......................................................... 120
4-63 VOC of CODP in case 2 ............................................................................ 120
4-64 FOC of MixAlco and CODP in case 2 .................................................... 121
4-65 Summary MixAlco and CODP economic results in case 2 .................... 122
4-66 Heat integration comparison of cases ......................................................... 124
4-67 Economic comparison of cases .................................................................. 124
1
1. INTRODUCTION
World liquid fuels consumption grew by 0.800 MM bpd in 2012. US Energy
Information Administration (EIA) expects consumption growth will be higher over the
next two years, at 0.900 million bpd in 2013 and 1.20 MM bpd in 2014 (EIA, 2013).
However, the liquid fuel production is estimated to be decline. Furthermore, the price of
crude oil is very sensitive to international politic issues. Clearly, new alternatives for
renewable fuels are necessary.
The MixAlco technology, invented by Professor M. Holtzapple at Texas A&M
University (Holtzapple, 2009), comprises a processing facility to produce liquid
transportation fuels and/or value-added chemicals from sustainable resources.
MixAlco converts materials such as municipal solid waste (MSW), sewage sludge,
forest product residues, and non-edible energy crops such as sweet sorghum into a wide
array of chemicals and secondary alcohols that can be further refined through separate,
well-established processes to produce renewable gasoline, jet fuel or diesel. The bio-
gasoline produced through the MixAlco technology is not ethanol. In fact, it has a
higher energy value than ethanol and can be blended directly with gasoline produced
from hydrocarbons. (Terrabon, 2010).
MixAlco process comprises a fermentation stage, which employs a mixed culture
of acid-forming microorganisms that convert biomass components (carbohydrates,
proteins, and fats) to carboxylate salts. Depending on the choice of buffer, the salts may
2
be ammonium carboxylates (buffered by NH4HCO3) or calcium carboxylates (buffered
by CaCO3) among others. Via pathway C in Figure 1-1, calcium carboxylates are
thermally converted into ketones, which are subsequently hydrogenated into a mixture of
secondary alcohols. Finally, these alcohols are chemically converted into hydrocarbon
fuels (gasoline, jet fuel, and diesel) (Pham, Holtzapple, & El-Halwagi, 2010). In
Appendix A, details on the MixAlco process are briefly discussed.
Figure 1-1. Pathways for converting biomass to hydrocarbon fuels (Pham et al., 2010)
Build and run a new plant, such as the one required for the MixAlco process, is
very costly; however, its economic performance may be greatly enhanced by retrofitting
analysis. This strategy comprises adding a bio-fuels plant like MixAlco process to an
existing fossil fuel plant like a crude oil distillation process (CODP).
Juan Camilo
3
This mechanism is beneficial for both parties because an inexpensive increase of
the production capacity of the refinery may be obtained, while economic matters for the
biofuels-producing plant are resolved. The systems obtained by integrating a biomass
fuel plant to the fossil fuel plant is regarded here as bio-refineries or integrated fossil
bio-refineries.
Basically, CODP comprises a preheating train where crude oil is fed from the
holding tank; then vaporized in the furnace where the combustion of a fuel is taking
place. Finally, it is fed to the bottom of the distillation column. The distillation column is
considered the master unit since all different cuts like light and heavy naphtha, kerosene,
light and heavy gas oils, and atmospheric residue are separated and purified. The
vacuum distillation unit further distills residual bottoms from the atmospheric tower,
where different cuts can be obtained like atmospheric, light vacuum, and heavy vacuum
gas oil. A large amount of heat is transported out to the preheating train from the
condenser, the end products, the strippers and the pumparounds.
The substantial energy requirement of crude oil distillation columns is met partly
by costly utilities, such as steam and fuel for fired heaters, and partly by heat recovered
from the process, using process-to-process heat exchange. Energy savings, therefore,
demand not only a distillation column that is energy-efficient, but also a heat exchanger
network (HEN) which minimizes utility costs by maximizing heat recovery. (Benali &
Tondeur, 2011). The CODP in this work corresponds to a modified version of a
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distillation unit of Ecopetrol. Modifications were necessary to protect intellectual
property rights.
This work assesses a retrofitting analysis to integrate the MixAlco process to
the crude oil distillation process. Focus is given to the problem of process modification
to the crude oil distillation system by considering increase the process profitability and
material substitution with biomass feedstocks. The approach proposed for this analysis
was developed by B. Cormier under the advisory of Dr. M. El-Halwagi at Texas A&M
University (Cormier, 2005). The proposed hierarchy is based on costs analysis and
involves internal process modification, operating-condition adjustment, and feedstock
substitution. If is needed, new units are added followed by the incorporation of new
production lines. Then, heat and mass integration techniques are used to link the units
and streams. (Cormier, 2005)
The competitiveness of markets nowadays and the focus on energy efficiency
requires improved heat-integrated process designs. Aspen Energy Analyzer (AEA)
working in concert with flowsheet simulators such as Aspen Plus provides an easy
environment to perform optimal heat exchanger network design and pinch analysis.
For the economic evaluation, an Aspen Process Economic Analyzer (APEA)
provides benefits that will be explored in this study.
5
2. OBJECTIVES
2.1 General Objective
Apply a retrofitting analysis to integrate into a crude oil distillation system, the
MixAlco process using the kenitonization route.
2.2 Specific Objectives
Conduct process integration studies to determine cost-effective strategies for
enhancing production incorporating the MixAlco process into the crude oil
distillation system.
Develop several energy and mass integration approaches and use them to induce
synergism and to reduce cost by exchanging heat, material utilities, and by
sharing equipment.
Develop cost-benefit analysis to guide the decision-making process and to
compare various production routes.
6
3. METHODOLOGY
Results from a previous work on MixAlco process economics were used as a base
case (Pham et al., 2012). In that work, the economics of the calcium carboxylate
platform (pathway C in Figure 3-1) using municipal solid waste or sugarcane bagasse as
feedstock were estimated. On this basis, the following MixAlco process features were
used: no requirement for sterility or any external enzymes, low capital cost, and cost-
effective dewatering, which comprise the use of an effective evaporation system, briefly
explained in Appendix A. In the previous work, the minimum selling prices of
hydrocarbon fuels reported can be around 1.57 USD /gal if municipal solid waste is
available at the US average tipping fee of 45 USD/dry ton (40 ton/h plant, with internal
hydrogen production). (Pham et al., 2012)
Retrofitting analysis was performed using the methodology developed by Cormier &
El-Halwagi developed on the framework of mass and energy process integration. An
overview of the methodology is shown in Figure 3-1 (Cormier, 2005). An explanation is
found in Section 3.1.
3.1 Description of the proposed methodology
3.1.1 Define needs
In the first step for the retrofitting analysis, it is possible to define the
opportunities in the processes that would result in an increased profitability.
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7
3.1.2 Process arrangements
Figure 3-1 shows three building blocks that are arranged in order of increasing cost.
The definition of each block is explained below (Cormier, 2005):
- Internal rearrangements: The goal is to reach the production target using low cost
strategies. These include process reconfiguration (e.g., stream rerouting) and
modification of operating conditions.
- Internal modification by adding new units: it is aimed to pursue medium-cost
modifications. These include required addition of new units, and/or replacement of
the existing units with new ones.
- External modification by adding new lines: Capital-intensive strategies are
invoked. These include the addition of new production lines.
3.1.3 Feasibility
Once two candidates are integrated into the current plant by heat and mass
integration, the ROI can be calculated. The decision to go deeper into the analysis
depends on the obtained value. (Cormier, 2005)
8
Figure 3-1. Flowchart of the proposed methodology (Cormier, 2005).
Define needs
Internal
rearrangements
only
Simulation
improvement
Cost analysis
Feasible?
Internal
modification by
adding new units
Simulation
improvement
Cost analysis
Feasible?
External
modification by
adding separated
lines
Simulation
improvement
Cost analysis
Proposal to the
company
Feasible?
Medium Cost
Low Cost
High Cost
Redefine needs
Internal rearrangements:
Process Reconfiguration
Modification of Operating Conditions
Internal Modifications:
Add New units
Modification of Operating Conditions
External Modifications:
Add new lines
Modification of Operating Conditions
Yes
Yes
Yes
9
3.2 Simulation Tools
Regardless to existence of specialized software for petroleum industry (Aspen
Hysys and PRO II), Aspen Plus software was used according a specific database
that estimates most of the desired properties of biomass.
Initially, each one of these two base cases was simulated separately. Then, both
plants were put together in a single worksheet to make integrations possible.
Specificities of each of these simulation cases are given below. For all simulations, the
following three steps were necessary:
- Flow sheet definition: All inlet and outlet streams to the different stages in both
the MixAlco and the CODP systems, as well as all unit operations and their
interconnecting streams were defined.
- Chemical components: All chemical components in the system, from reagents
to intermediates and products were specified during simulations. Appendix B
shows the properties for each of the substances used in the simulations..
- Operating conditions: The operating conditions, such as temperature, pressure,
heat duties etc., for each unit operation were specified for each process in next
sections.
3.2.1 MixAlco process base case.
The simulation was made using the National Renewable Energy Laboratory
(NREL) database In-house Pure Component Database (INHSPCD). Within this tool,
10
estimation of most of the properties of biomass components such as glucose, xylose,
cellulose, xylan, lignin were possible. Components other than the ones listed above, are
identified within this database as solslds. (Wooley & Putsche, 1996).
As Wooley and Putsche, (1996) suggested, the thermodynamic package used in
this simulation was non-random-two-liquid (NTRL), where NRTL liquid include
activity coefficient model, Henrys law for the dissolved gases, and Redlich-Kwong-
Soave (RKS) equation of state for the vapor phase.
The MixAlco process simulation was made for a capacity of 40.0 ton/h of
biomass. The biomass feedstock in the simulation was a mixture 80.0% - 20.0% w/w of
sugarcane bagasse (like carbohydrate source) and chicken manure (like secondary
nutrient source) respectively. Table 3-1 shows the biomass feed composition. Other
details on the feedstock stream (e.g., properties) are shown with simulation results in
Section 4.1.1.1.
Table 3-1. Biomass feed composition for MixAlco process
Component Feedstock
ton/h %w/w
Cellulose 16.8 43.0
Hemicellulose (xylan) 7.50 19.0
Lignin 10.0 25.0
Solslds 5.20 13.0
Total 40.0 100
Operating conditions used for simulation of all MixAlco unit operations are
shown in Table 3-2.
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Table 3-2. MixAlco operating conditions
Process Parameter Value
Feed handling Temperature (C) 55.0
Pressure (bar_a) 1.00
Pretreatment Temperature (C) 55.0
Pressure (bar_a) 1.00
Fermentation Temperature (C) 55.0
Pressure (bar_a) 1.00
Ketonization Temperature (C) 430
Pressure (bar_a) 0.0400
Ketone hydrogenation Temperature (C) 130
Pressure (bar_a) 55.0
Lime kiln Temperature (C) 500
Pressure (bar_a) 1.00
Alcohol dehydration Temperature (C) 300
Pressure (bar_a) 3.00
Oligomerization Temperature (C) 300
Pressure (bar_a) 3.00
Olefin hydrogenation Temperature (C) 130
Pressure (bar_a) 55.0
Gasification Temperature (C) 760
Pressure (bar_a) 1.00
Steam-gas shift Temperature (C) 254
Pressure (bar_a) 1.00
3.2.2 CODP base case.
The CODP simulation was made using Grayson Streed and Braun k-10 (BK10)
as thermodynamic packages, for a crude oil load of 27,012 bpd (162 ton/h) and 22.8
API. Because, Grayson property methods were developed for systems containing
hydrocarbons and light gases and BK10 property method is suited for vacuum and low
pressure applications. (Aspen plus, 2013). Table 3-3 and 3-4 shows the assay data.
The crude oil distillation curve is presented in Figure 3-2.
12
Table 3-3. Assay data for crude oil****
%Distilled* Molecular Weight Specific
Gravity
Sulfur
curve**
Viscosity***
(80C)
Viscosity***
(100C)
0.929 71.0 0.646
2.54 97.0 0.725
5.46 116 0.759 5x10-3
8.71 144 0.788 0.0190
11.5 153 0.810 0.0500
13.7 171 0.823 0.0790
16.6 186 0.838
21.9 206 0.862 0.322
28.8 250 0.872
36.6 280 0.897 0.808
41.3 323 0.920 9.13 5.80
44.9 357 0.921 0.995 10.3 5.79
53.8 378 0.932 18.0 10.3
62.1 436 0.946 1.19 43.3 22.5
67.3 514 0.956 1.42 77.8 34.5
85.0 1,247 1.03 289,150 31,971
Molecular weight:
288
Bulk value:
0.0488
*Mid-percent distilled (same basis as distillation data, i.e., volume or weight)
**Given in %w/w
***Given in centistokes
****Data supply by Ecopetrol
Table 3-4. Assay data for crude oil Light ends**
Component* % Mass
C2 0.212
C3 3.60
IC4 5.40
NC4 14.5
IC5 24.4
NC5 23.5
Hexane 28.4
Total % light ends in the assay 0.940
*The number after C refers to the number of carbons in the alkane molecule. I is for iso (non-linear)
structures and N for straight chains
**Data supply by Ecopetrol
13
Figure 3-2. Crude oil distillation TPB
Operating conditions used for the atmospheric and vacuum columns are shown in
Table 3-5.
Table 3-5. CODP operating conditions
Process Parameter Value
Column T-204 - Atmospheric tower
Heat duty condenser (kJ/s) 7,206
Tray crude feed 17.0
Tray steam feed 19.0
Tray number 19.0
Condenser Partial
Temperature tray 1 (C) 98.0
Temperature tray 19 (C) 366
Pressure tray 1 (bar_a) 1.70
Pressure tray 19 (bar_a) 2.20
Heat duty pumparound MPA (kJ/s) 5,113
Heat duty pumparound MPACAL (kJ/s) 563
Column T-205- Vaccum tower
Heat duty condenser (kJ/s) 308
Tray crude feed 7.00
Tray steam feed 8.00
Tray number 8.00
14
Process Parameter Value
Condenser Partial
Temperature tray 1 (C) 60.0
Temperature tray 8 (C) 389
Pressure tray 1 (bar_a) 0.0400
Pressure tray 8 (bar_a) 0.140
Heat duty pumparound UPA (kJ/s) 2,117
Heat duty pumparound MPA (kJ/s) 5,874
Heat duty pumparound MPACAL (kJ/s) 219
3.2.3 MixAlco process and CODP retrofitted plant
The simulation of MixAlco process and CODP integrated was made from the base
case of each plant, so the operating conditions and feedstock properties were the same
shown in sections 3.2.1 and 3.2.2; regarding the thermodynamic package: NTRL and
Grayson Streed.
3.3 Process integration
The design of any industrial process relies on process simulators and programs for
unit operation design. The core of process design rests on two important dimensions:
mass and energy. Mass involves the creation and routing of chemical species in reaction,
separation, and byproduct/waste-processing systems. These constitute the heart of the
process and define a companys technology base. Energy provides the necessary heating,
cooling, and shaftwork for those systems.
Because most industrial processes are complicated, performance and economics
depend not only on proper selection and design of individual components but also on
15
proper assembling of building blocks. Fundamental principles can guide this assembly.
Process integration comprises all means to achieve the goals of optimal assembly and
performance. In process integration, the unity of the entire process is emphasized.
Pinch analysis is the most successful way to achieve energy integration; which impacts
mainly in process economics. Mass integration, on the other hand, has received great
attention and development, because it directly impacts process performance. (El-
Halwagi & Spriggs, 1998).
For this, material rerouting and heat exchanger network (HEN) were considered
(Cormier, 2005).
3.3.1. Material rerouting
Mass integration is a systematic methodology that provides a fundamental
understanding of the global flow of mass within the process and employs this
understanding. To apply this integration the following steps were used:
Analysis: Detect the minimum fresh resource consumption and minimum waste
discharge streams.
Retrofit: Modify an existing water-using network to maximize water reuse and
minimize wastewater generation through effective process changes (Cormier,
2005).
For additional details about this type of integration see Appendix C.
16
3.3.2 Heat Exchanger Network (HEN)
In the plant, heating and cooling represent an important operating cost. In order to
minimize the operating cost for the heat utilities, heat integration is needed. The
following multiple design objectives are pursued:
Minimize the investment cost of the units (i.e., surface area of exchanger, heater
and/or cooler).
Minimize the operating cost of utilities (steam, cooling water, etc).
Minimize the number of units (i.e., heat exchanger). (Cormier, 2005)
In this work, energy integration was performed using Aspen Energy Analyzer
(AEA) in compliance to all license agreements. For additional details about this type of
integration see Appendix D.
3.3.3 Cost Analysis
Many technical and environmental decisions during process design are strongly
impacted by economic factors; therefore, an essential component of any sustainable
design is an economic analysis, which is performed on the basis of total investment and
operating costs. (El-Halwagi, 2012). Typically, a minimum of 15.0% for the ROI is
pursued. If this is not achievable, ROIs of 5.00 to 10.0% may be acceptable under
17
current market conditions. (Cormier, 2005) Additional details about economic analysis
concepts are presented in Appendix E.
In this work, the prices for feedstocks, chemicals and material disposal for
MixAlco process were taken from Pham et al., (2012). The prices of crude oil were
taken from the EIA official website. The utilities costs were taken from the database of
AEA except for the steam cost, which was taken from Seider, (2004). Finally, the prices
for refinery products were found in the EIA official website. Appendix D shows the
price profile for these products. For Atmospheric Gas Oil (AGO) and Light Vacuum Gas
Oil (LVGO) prices, a factor of 10% from the price of Heavy Vacuum Gas Oil (HVGO)
was used. All these prices are shown in Table 3-6.
Table 3-6. Feedstock, utilities and product prices
Item Costs and Prices
(USD / unit)
Fee
dst
ock
cost
s
Sugarcane bagasse (USD/ton) 60.0
Chicken manure (USD/ton) 10.0
Crude oil (USD/ton) 643
Crude oil (USD/barrel) 93.4
Quick Lime (USD/ton) 70.0
Flocculant (USD/ton) 991
Iodoform (USD/kg) 25.0
CaCO3 (USD/ton) 50.0
Material disposal (USD/ton) 18.0
Uti
liti
es c
ost
Fired Heat (USD/ton) 2.55
MP Steam (USD/ton) 4.36
Cooling Water (USD/m3) 4x10-3
LP Steam (USD/ton) 4.17
Refrigerant (USD/m3) 0.0130
Electricity (USD/kWh) 0.0620
Steam @ 353C (USD/ton) 10.0
Steam @ 454C (USD/ton) 10.0
18
Item Costs and Prices
(USD / unit) P
rod
uct
s se
llin
g
pri
ces
Gasoline (USD/gal) 3.28
Jet (USD/gal) 2.88
Diesel (USD/gal) 3.75
AGO (USD/gal) 2.40
LVGO (USD/gal) 2.30
HVGO (USD/gal) 2.20
Asphalt (USD/gal) 1.30
19
4 RESULTS AND DISCUSSION
Simulation results are presented for each one of the two base cases (MixAlco and
COPD). This simulations were necessary in order to be able to compare MixAlco
alone vs MixAlco retrofitted within a COPD plant. Then, the methodology for
retrofitting (shown in Figure 3-1) will be followed, which eventually (third loop)
conduces to the results obtained for the combined MixAlco-COPD integrated plant.
An economic analysis is presented for each one of the integration possibilities and for
the combined MixAlco-COPD. Extensive comparisons are presented at the end. The
Enthalpy reported by Aspen Plus is in their standard states at 1 atm and 298.15K.
4.1 Simulation results
In this section, simulation building procedures as well as relevant results of each one
of the base cases are discussed next (MixAlco base case and COPD base case).
4.1.1 Simulation builds up and results for MixAlco base case
The MixAlco simulation was divided in seven blocks to build up a simulation,
as shown in Figure 4-1.
These blocks are listed and explained below:
1. FEED-HAN: Feed handling process
2. PRET -FER: Pretreatment and fermentation process
3. DEWATER: Dewatering process.
4. KETONIZA: Ketonization and ketone hydrogenation processes.
20
5. LIME-KIL: Lime kiln process.
6. FINAL: Dehydratation, oligomerization and saturation processes.
7. GASIFICA: Gasification reactor, steam gas shift reactor, and adsorption process
Figure 4-1. Blocks of MixAlco process simulation
21
4.1.1.1 MixAlco Block description
- Feed Handling (Unit 1)
The Feed Handling block exists only for simulation purposes and it is meant: (i) to mix
the reacting substances (biomass, water, and lime) to prepare them for pretreatment and (ii)
to obtain lime (Ca(OH)2) from quick lime (CaO) as shown in Equation 4-1.
(4-1)
Quick lime Lime
In the actual MixAlco process, feed handling would occur simultaneously (in the
same unit) with pretreatment. This is because the reaction in Eq. 4-1 is exothermic
(Enthalpy of reaction obtained was 1.960kJ/s shown in Table 4-2); thus, it is advantageous
to use the reaction heat to obtain an increase of temperature necessary for pretreatment to
occur at a measurable rate. Because in the simulation feed handling and pretreatment were
not put in the same unit, this fact could not be considered. Instead, an external source of
heat was implemented for the pretreatment stage.
Two sources for quick lime were considered: The first is CaO produced in-site and
the second is make-up fresh quick lime. The quick lime that is produced in site comes from
the LIME-KIL block explained in a later section. The stream that carries this reactant has
been labeled as CAO-RECY in Figure 4-2. As shown in Table 4-1, this stream contains
4.06 ton/h of CO2 which corresponds to 44.0% w/w of the stream composition. This gas is
22
a reaction byproduct which in the actual process is expelled as it is produced, but in this
simulation has to be carried all the way to the end gasification block in the SP-115. The
unit operation CON-101 is a conveyor set up to transport this recycled stream. On the other
hand, the fresh, make-up CaO (labeled as CAO-MAKE in Figure 4-2) is purchased with a
cost of 70.0USD/ton. The mass ratio CAO:CAO-MAKE is 1:10 which clearly shows that a
lime recovery process is represented in a saving operating cost. In addition a water fed at a
flow rate of 2ton/h, stream labeled as H2O-LIME in Figure 4-2 was considered.
For the reaction (Eq. 4-1, occurring in R-101), a conversion factor of 1 was
employed, although the reactants (i.e., water and quick lime) were fed in exact
stoichiometric amounts (i.e., no reactant was fed in excess). (Gosseaume, 2011).
Two streams leave this block: (i) Stream 1(OUT) in Figure 4-2 required for the
reactor convergence and after a mixing unit (TK-101) (ii) Stream (BIOM-LIM(OUT))
which is the stream that contains biomass mixed with water and lime and goes to
pretreatment.
Results from mass and heat balance in the simulation for this block are shown per
stream in Table 4-1. On the other hand, the heat balance for the equipment in this block is
shown in Table 4-2, where the conveyor power consumption is very low.
23
Figure 4-2. Feed handling simulation
24
Table 4-1. Feed handling mass and heat balance
BIOMASS CA-BIO CA-BIOM CAO CAO-MAKE CAO-RECY H20-LIME
Temperature (C) 25.0 55.0 55.0 55.0 55.0 55.0 25.0
Pressure (bar_a) 1.00 1.00 1.00 1.00 1.00 1.00 1.00
Mass vapor fraction 0 0.0800 0.0800 0.440 0 0.440 0
Mass solid fraction 0.870 0.820 0.820 0.560 1.00 0.560 0
Mass flow (ton/h) 39.5 51.6 51.6 9.24 0.900 9.24 2.00
Enthalpy (kJ/s) 80,720 120,205 120,205 26,310 2,824 26,310 8,809
Component mass flow (ton/h)
CELLU-01 16.8 16.8 16.8 0 0 0 0
XYLAN 7.50 7.50 7.50 0 0 0 0
LIGNI-01 10.0 10.0 10.0 0 0 0 0
SOLSL-01 5.17 5.17 5.17 0 0 0 0
SOLUN-01 0 0 0 0 0 0 0
WATER 0 0.0500 0.0500 0 0 0 2.00
CO2 0 4.06 4.06 4.06 0 4.06 0
CA(OH)2 0 8.03 8.03 0 0 0 0
CAO 0 0 0 5.18 0.900 5.18 0
Table 4-2. Heat balances for Feed Handling equipment
Type Unit Heat duty [kJ/sec] Heat duty [kJ/h]
Reactor R-101 -1,541 -5.55x106
Conveyor CON-101 0.750 2.7 x103
25
- Pretreatment and Fermentation (Unit 2)
Lignocellulosic materials are resistant to the enzymatic degradation, because cellulose
and hemicelluloses (carbohydrates) are encapsulated by lignin, which keeps the enzymes
secreted by the microorganisms from reaching it. Pretreatment is necessary to remove
lignin and enable the fermentation step. (Gosseaume, 2011). The pretreatment simulation
is shown in Figure 4-3. For pretreatment conditions 400 ton/h of fresh water stream
(labeled as H2O-PRET in Figure 4-3) was required. Also a blower (CM-101) was
simulated for bring the air in the pretreatment slurry (6.70ton/h).
Figure 4-3. Pretreatment simulation
26
Due to the complex reaction in pretreatment stage, the reactor R-102 was simulated
in two ways: (i) the first way (for mass balance) was through Ryield, based on known yield
of the exit current. And (ii) the second way (for heat balance) was through Rstoic in order
to calculate the endothermic heat of reaction that was 14,044kJ/s shown in Table 4-5.
For the Rstoic reactor was assumed a conversion factor of 15.0% (Eq. 4-2), 35.0%
(Eq. 4-3) and 30.0% (Eq. 4-4) for cellulose, xylan and lignin in the undigested biomass
(Mixed), respectively (Sierra, Garca, & Holtzapple, 2010). These conversions were based
on a study of lime pretreatment of poplar wood at laboratory scale. Based on previous
studies of MixAlco process at different capacities as Holtzapple, (2004), it is assumed
that yields are not affected by the scaling capacity. The biomass undigested conversion is
0.200 ton per ton of biomass VS, (in stream BIOM-LVS, 34.9 ton/h is biomass VS)
resulting in 8.80 ton/h of undigested biomass, that is directed to gasification process
(labeled as BIOM-LV in Figure 4-3). The remaining biomass is digested (Cisolid) (labeled
as BIOM-S in Figure 4-3) and continued to fermentation process.
In this stream the theorical conversion of 0.800 ton of digested biomass per ton of
biomass VS is satisfied, resulting in 26.1 ton/h.
CELLU-01(Cisolid) --> CELLU-01(Mixed) (4-2)
XYLAN (Cisolid) --> XYLAN (Mixed) (4-3)
LIGNI-01(Cisolid) --> LIGNI-01(Mixed) (4-4)
In fermentation process the biomass digested (labeled as BIOM-LV in Figure 4-3
and 4-4) is converted in carboxylates salts using as a buffer CaCO3. The fermentation
simulation is shown in Figure 4-4.
27
Table 4-3. Pretreatment mass and heat balance
AIR AIR2 AIRR BIOM-LIM BIOM-LV BIOM-LVS BIOM-S H20-PRET
Temperature (C) 25.0 30.5 30.5 55.0 55.0 55.0 55.0 25.0
Pressure (bar_a) 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00
Mass vapor fraction 1.00 1.00 1.00 0.0770 0 0 0 0
Mass solid fraction 0 0 0 0.820 0.0200 0.0760 1.00 0
Mass flow (ton/h) 6.70 6.70 6.70 51.6 416 442 26.1 400
Enthalpy (kJ/s) 3.46x10-13 10.4 10.4 120,205 1,854,269 1,910,655 56,386 1,761,927
Component mass flow (ton/h)
CELLU-01 0 0 0 16.8 2.80 17.0 14.2 0.0
XYLAN 0 0 0 7.50 2.80 7.80 5.00 0.0
LIGNI-01 0 0 0 10.0 3.20 10.1 6.90 0.0
SOLSL-01 0 0 0 5.2 5.00 5.00 0 0
WATER 0 0 0 0.0 402 402 0 400
CO2 0 0 0 4.10 0 0 0 0
CA(OH)2 0 0 0 8.00 0 0 0 0
NITROGEN 5.30 5.30 5.30 0 0 0 0 0
O2 1.40 1.40 1.40 0 0 0 0 0
Two sources for calcium carbonate were considered: The first is CaCO3 recycled from KETONIZA block and the second is
make-up fresh CaCO3. The flow rate of CaCO3 recycled is 6.30 ton/h as shown in Table 4-4; this stream is labeled as CACO3REC in
Figure 4-4. The make-up flow rate is 9.30 ton/h (labeled as MK-CACO3 in Figure 4-4) and is purchased with a cost of 50.0 USD/ton.
The CaCO3 recycled represent 40.0% of CaCO3 consumption resulting in a saving operating cost.
28
The conversion factors for the serial reactions performed in the fermentation train
(R-103 to R-105) are shown in Table A-1 (Gosseaume, 2011). Besides, the reactions in
fermentation process are shown in Equations A-1 to A-11. The salts in solution are
obtained in the stream called SALTS shown in Figure 4-4, with a total flow rate of 25.1
ton/h as shown in Table 4-4. A theorical conversion is getting for 0.600 ton of carboxylate
salts per ton of biomass feed. The stream residue from fermentation (BIOMASS) with a
flow of 16.1 ton/h goes to a gasification process. Table 4-4 shows the material balance for
this stage. In addition a water fed at a flow rate of 200 ton/h, stream labeled as H2O-FERM
in Figure 4-4 was considered. The global heats of reaction are exothermic for reactor R-
103, R-104, R-105 (with enthalpies 1,051kJ/s; 768 kJ/s; 278kJ/s respectively); and
endothermic for reactor R-106 (with enthalpy 6,480kJ/s). Table 4-5 shows the summary of
heat equipment loads.
Finally, a water cooling circuit is simulated in order to quantify the cost of those
equipment for improve the cost analysis of this process.
Results from mass and heat simulation for this block are shown per stream in Table
4-4. On the other hand, heat balances for this block shows a power consumption of 40.4kW
(Table 4-5). For a heat integration study the heat exchangers simulated in this block were
assumed as coolers for count the cooling water utility in the operating cost, that why the
total cooling required in this block is 7,957 kJ/s.
29
Figure 4-4. Fermentation simulation
30
Table 4-41. Fermentation mass and heat balance
BIO-SAL BIO-SAL1 BIO-SAL2 BIO-SAL3 BIOM1 BIOM2 BIOM3 BIOMASS
Temperature (C) 55.0 55.0 55.0 55.0 55.0 55.0 55.0 55.0
Pressure (bar_a) 1.01 1.01 1.01 1.01 1.01 1.01 1.01 1.01
Mass vapor fraction 0 0.0100 0.0170 0.0210 0.147 0.302 0.418 0.486
Mass solid fraction 0.177 0.129 0.0890 0.0590 0.853 0.698 0.582 0.514
Mass flow (ton/h) 247.6 236.4 227.5 220.9 22.0 18.8 17.0 16.1
Entalphy (kJ/s) -991,492 970,009 951,047 934,782 52,274 49.090 47,282 46,334
Component mass flow (ton/h)
CELLU-01 8.80 4.60 2.20 1.00 8.80 4.60 2.20 1.00
XYLAN 3.10 1.60 0.80 0.400 3.10 1.60 0.800 0.400
LIGNI-01 6.90 6.90 6.90 6.90 6.90 6.90 6.90 6.90
WATER 200.5 200.3 200.1 200.1 0 0 0 0
CO2 3.20 5.70 7.10 7.80 3.20 5.70 7.10 7.80
CA(OH)2 0 0 0 0 0 0 0 0
CACO3 1.30 3.20 3.90 2.50 0 0 0 0
CA(CH-01 19.2 11.4 5.40 1.90 0 0 0 0
CA(CH-02 1.40 1.00 0.30 0.100 0 0 0 0
CA(CH-03 3.30 1.80 0.90 0.200 0 0 0 0
(Continued Table 4-4)
CACO3 CACO3-1 CACO3-2 CACO3-4 CACO3-5 CACO3REC CW-1 CW2 CW3 CW4
Temperature (C) 55.0 37.3 37.3 37.3 37.3 130 25.0 31.1 25.0 31.1
Pressure (bar_a) 1.00 1.00 1.00 1.00 1.00 7.60 1.00 0.800 1.00 0.800
Mass vapor fraction 0 0 0 0 0 0 0 0 0 0
Mass solid fraction 1.00 1.00 1.00 1.00 1.00 1.00 0 0 0 0
Mass flow (ton/h) 6.30 3.90 3.90 3.90 3.90 6.30 301.8 301.8 301.8 301.8
31
CACO3 CACO3-1 CACO3-2 CACO3-4 CACO3-5 CACO3REC CW-1 CW2 CW3 CW4
Entalphy (kJ/s) 21,066 13,018 13,018 13,018 13,018 20,949 1.33x106
Component mass flow (ton/h)
CELLU-01 0 0 0 0 0 0 0 0 0 0
XYLAN 0 0 0 0 0 0 0 0 0 0
LIGNI-01 0 0 0 0 0 0 0 0 0 0
WATER 0 0 0 0 0 0 301.8 301.8 301.8 301.8
CO2 0 0 0 0 0 0 0 0 0 0
CA(OH)2 0 0 0 0 0 0 0 0 0 0
CACO3 6.30 3.90 3.90 3.90 3.90 6.30 0 0 0 0
CA(CH-01 0 0 0 0 0 0 0 0 0 0
CA(CH-02 0 0 0 0 0 0 0 0 0 0
CA(CH-03 0 0 0 0 0 0 0 0 0 0
(Continued Table 4-4)
CW5 CW6 CW6 CW7 CW8 H2O H20-FERM H20-PRET MK-CACO3 SAL-H2O SALT3
Temperature (C) 25.0 31.1 31.1 25 31.1 41.1 50.0 25.0 25.0 55.0 55.0
Pressure (bar_a) 1.00 0.800 0.800 1.00 0.800 0.800 1.00 1.00 1.00 2.06 1.01
Mass vapor fraction 0 0 0 0 0 0 0 0 0 0 0
Mass solid fraction 0 0 0 0 0 0 0 0 1.00 0.112 0.0800
Mass flow (ton/h) 302 302 302 302 302 200 200 400.0 9.30 225.6 217.6
Entalphy (kJ/s) 1.33x106 877,480 875,520 1,761,927 31,006 939,210 921,027
Component mass flow (ton/h)
CELLU-01 0 0 0 0 0 0 0 0 0 0 0
XYLAN 0 0 0 0 0 0 0 0 0 0 0
LIGNI-01 0 0 0 0 0 0 0 0 0 0 0
SOLSL-01 0 0 0 0 0 0 0 0 0 0 0
WATER 302 302 302 302 302 200 200 400 0 200 200
32
CW5 CW6 CW6 CW7 CW8 H2O H20-FERM H20-PRET MK-CACO3 SAL-H2O SALT3
CO2 0 0 0 0 0 0 0 0 0 0 0
CA(OH)2 0 0 0 0 0 0 0 0 0 0 0
CACO3 0 0 0 0 0 0 0 0 9.30 1.30 3.20
CA(CH-01 0 0 0 0 0 0 0 0 0 19.2 11.4
CA(CH-02 0 0 0 0 0 0 0 0 0 1.40 1.00
CA(CH-03 0 0 0 0 0 0 0 0 0 3.30 1.80
(Continued Table 4-4)
SALT4 SALT5 SALTS SALW1 SALW2 SALW3 SALW4 SALW5 SALW6
Temperature (C) 55.0 55.0 55.0 46.3 55.0 46.3 55.0 46.3 55.0
Pressure (bar_a) 1.01 1.01 1.01 1.86 2.06 1.86 2.06 1.86 2.06
Mass vapor fraction 0 0 0 0 0 0 0 0 0
Mass solid fraction 0.0490 0.0230 0.112 0.0800 0.0800 0.0490 0.0490 0.0230 0.0230
Mass flow (ton/h) 210 205 226 218 218 210 210 205 205
Entalphy (kJ/s) 903,943 888,661 939,217 922,979 921,020 905,896 903,936 890,613 888,653
Component mass flow (ton/h)
CELLU-01 0 0 0 0 0 0 0 0 0
XYLAN 0 0 0 0 0 0 0 0 0
LIGNI-01 0 0 0 0 0 0 0 0 0
SOLSL-01 0 0 0 0 0 0 0 0 0
WATER 200 200 200 200 200 200 200 200 200
CO2 0 0 0 0 0 0 0 0 0
CA(OH)2 0 0 0 0 0 0 0 0 0
CACO3 3.90 2.50 1.30 3.20 3.20 3.90 3.90 2.50 2.50
CA(CH-01 5.40 1.90 19.2 11.4 11.4 5.40 5.40 1.90 1.90
CA(CH-02 0.300 0.100 1.40 1.00 1.00 0.300 0.300 0.100 0.100
CA(CH-03 0.900 0.200 3.30 1.80 1.80 0.900 0.900 0.200 0.200
33
Table 4-5. Heat balances for Pretreatment and Fermentation equipments
Type Unit Heat duty [kJ/sec] Heat duty [kJ/h]
Cooler C-101 117 4.21x105
Heat Exchanger E-101 1,960 7.06x106
Heat Exchanger E-102 1,960 7.06x106
Heat Exchanger E-103 1,960 7.06x106
Heat Exchanger E-104 1,960 7.06x106
Pumps P-101 7.50 2.70x104
Pumps P-102 7.50 2.70x104
Pumps P-103 7.50 2.70x104
Pumps P-104 7.60 2.70x104
Compresor CM-101 10.4 3.74x104
Reactor R-102 14,044 5.06x107
Reactor R-103 -1,051 -3.78x106
Reactor R-104 -768 -2.76x106
Reactor R-105 -278 -1.00x106
Reactor R-106 6,480 2.33x107
- Dewatering (Unit 3)
Dewatering block exits only for simulated the water separation from the produced
fermentation broth, using a vapor compression. Figure 4-5 shows the block simulation.
The fermentation broth labeled as SALT-H20 comes 25.1 ton/h of salt plus 200 ton/h of
water. A six train of heat exchangers and separators are used to simulate the vapor
compression system, where the steam separated in the first train is compressed for recycling
in the process. The separated water (labeled as WATDISTI in Figure 4-5) is a waste water
stream. The separated salts labeled as SALTDES continued to ketonization process.
Others packing units are simulated in order to quantify the cost of that equipment for
improve the cost analysis of this block.
34
Results from mass and heat simulation for this block are shown per stream in Table 4-6. On the other hand, heat balances for this block
shows power consumption for compressor CM-102 of 1,214 kW. A heating load required in this block is 113,763 kJ/s (Table 4-7)
Figure 4-5. Dewatering simulation
35
Table 4-62. Dewatering mass and heat balance
SAL-DESC SAL-H20 SAL1 SAL2 SAL3 SAL4 SAL5 SAL6 SALT SALT-H20
SALT-
WAT
Temperature (C) 55.0 55.0 162 162 163 165 165 162 163 150 55.0
Pressure (bar_a) 2.06 2.06 6.00 6.50 6.60 6.90 7.00 6.50 6.00 1.90 2.10
Mass vapor fraction 0 0 0 0 0 0 0 0 0 0.900 0
Mass solid fraction 0.112 0.112 1.00 1.00 1.00 1.00 1.00 1.00 1.00 0.100 0.100
Mass flow (ton/h) 225.6 225.6 4.20 4.20 4.20 4.20 4.20 4.20 25.2 225.6 225.6
Enthalpy (kJ/s) 939,210 939,210 10,431 10,431 10,431 10,431 10,431 10,431 62,572 796,918 939,210
Component mass flow (ton/h)
WATER 200 200 0 0 0 0 0 0 0 200 200
CACO3 1.30 1.30 0.200 0.200 0.200 0.200 0.200 0.200 1.30 1.30 1.30
CA(CH-01 19.2 19.2 3.20 3.20 3.20 3.20 3.20 3.20 19.2 19.2 19.2
CA(CH-02 1.40 1.40 0.200 0.200 0.200 0.200 0.200 0.200 1.40 1.40 1.40
CA(CH-03 3.30 3.30 0.500 0.500 0.500 0.500 0.500 0.500 3.30 3.30 3.30
(Continued Table 4-6)
SALTDE
S
SALWR
1
SALWR
2
SALWR
3
SALWR
4
SALWR
5
SALWR
6
SALWR
7
SALWR
8
SALWR
9
SALWR1
0
SALWR1
1
Temperature (C) 163 150 150 150 150 150 150 165 165 165 165 165
Pressure (bar_a) 6.00 1.90 1.90 1.90 1.90 1.90 1.90 7.00 7.00 7.10 7.40 7.50
Mass vapor
fraction 0 0.900 0.900 0.900 0.900 0.900 0.900 0 0 0 0 0
Mass solid
fraction 1.00 0.100 0.100 0.100 0.100 0.100 0.100 0.100 0.100 0.100 0.100 0.100
Mass flow
(ton/h) 25.2 37.6 37.6 37.6 37.6 37.6 37.6 37.6 37.6 37.6 37.6 37.6
Enthalpy (kJ/s) 62,572 132,846 132,846 132,846 132,846 132,846 132,846 132,846 132,846 132,846 132,846 132,846
36
SALTDE
S
SALWR
1
SALWR
2
SALWR
3
SALWR
4
SALWR
5
SALWR
6
SALWR
7
SALWR
8
SALWR
9
SALWR1
0
SALWR1
1
Component mass flow (ton/h)
WATER 0 33.4 33.4 33.4 33.4 33.4 33.4 33.4 33.4 33.4 33.4 33.4
CACO3 1.30 0.200 0.200 0.200 0.200 0.200 0.200 0.200 0.200 0.200 0.200 0.200
CA(CH-01 19.2 3.20 3.20 3.20 3.20 3.20 3.20 3.20 3.20 3.20 3.20 3.20
CA(CH-02 1.40 0.200 0.200 0.200 0.200 0.200 0.200 0.200 0.200 0.200 0.200 0.200
CA(CH-03 3.30 0.500 0.500 0.500 0.500 0.500 0.500 0.500 0.500 0.500 0.500 0.500
(Continued Table 4-6)
SALWR12 SALWR13 SALWR14 SALWR15 SALWR16 SALWR17 SALWR18 ST2 ST3 ST4 ST5
Temperature (C) 164 162 162 163 165 165 162 177 175 172 170
Pressure (bar_a) 7.00 6.50 6.50 6.60 6.90 7.00 6.50 9.30 8.80 8.30 7.80
Mass vapor fraction 0 0.1 0 0 0 0 0 1 1 1.00 1.00
Mass solid fraction 0.100 0.100 0.100 0.100 0.100 0.100 0.100 0 0 0 0
Mass flow (ton/h) 37.6 37.6 37.6 37.6 37.6 37.6 37.6 33.4 33.4 33.4 33.4
Enthalpy (kJ/s) 151,669 150,843 151,649 151,649 151,649 151,649 151,649 121,929 121,971 122,015 122,061
Component mass flow (ton/h)
WATER 33.4 33.4 33.4 33.4 33.4 33.4 33.4 33.4 33.4 33.4 33.4
CACO3 0.200 0.200 0.200 0.200 0.200 0.200 0.200 0 0 0 0
CA(CH-01 3.20 3.20 3.20 3.20 3.20 3.20 3.20 0 0 0 0
CA(CH-02 0.200 0.200 0.200 0.200 0.200 0.200 0.200 0 0 0 0
CA(CH-03 0.500 0.500 0.500 0.500 0.500 0.500 0.500 0 0 0 0
37
(Continued Table 4-6)
ST6 ST7 STEAMM WAT1 WAT2 WAT3 WAT4 WAT5 WAT6 WATDISTI WATER
Temperature (C) 167 166 230 163 162 163 164 165 162 60.0 162
Pressure (bar_a) 7.30 7.10 9.80 6.00 6.50 6.60 6.90 7.00 6.50 5.50 6.50
Mass vapor fraction 1.00 1.00 1.00 1.00 1.00 1.00 0.0 1.00 1.00 0 1.00
Mass solid fraction 0 0 0 0 0 0 0 0 0 0 0
Mass flow (ton/h) 33.4 33.4 33.4 33.4 33.4 33.4 33.4 33.4 33.4 200 200
Enthalpy (kJ/s) 122,110 122,130 120,971 122,185 122,193 122,185 122,185 122,185 122,185 875,450 733,155
Component mass flow (ton/h)
WATER 33.4 33.4 33.4 33.4 33.4 33.4 33.4 33.4 33.4 200 200
CACO3 0 0 0 0 0 0 0 0 0 0 0
CA(CH-01 0 0 0 0 0 0 0 0 0 0 0
CA(CH-02 0 0 0 0 0 0 0 0 0 0 0
CA(CH-03 0 0 0 0 0 0 0 0 0 0 0
- Ketonization (Unit 4)
Ketonization simulation is shown in Figure 4-6; and Table 4-8 shows the material balance. In ketonization block the carboxylate
salts (labeled as SALDEH in Figure 4-6) are converted into ketones (labeled as KET-CACO in Figure 4-6) by a thermal conversion at
high temperatures (430C), and vacuum pressure (30 mmHg); producing 9.60 ton/h of ketones. The conversion factor for the serial
reactions performed in the reactor R-107 was 0.99 (Gosseaume, 2011). The reactions in ketonization are shown in Equations A-12 to
A-16 (Appendix A).
38
Table 4-7. Heat balances for Dewatering equipments
Type Unit Heat duty [kJ/sec] Heat duty [kJ/h]
Heater H-101 18,956 6.82x107
Heater H-102 18,956 6.82x107
Heater H-103 18,956 6.82x107
Heater H-104 18,956 6.82x107
Heater H-105 18,956 6.82x107
Heater H-106 18,983 6.83x107
Heat Exchanger E-105 142,292 5.12x108
Heat Exchanger E-106 958 3.45x106
Heat Exchanger E-107 42.0 1.51x105
Heat Exchanger E-108 44.0 1.58x105
Heat Exchanger E-109 46.0 1.66x105
Heat Exchanger E-110 48.0 1.73x105
Heat Exchanger E-111 20.0 7.20x104
Compressor CM-102 1,214 4.37x106
By the thermal conversion 14.2 ton/h of calcium carbonate was produced. The
carbonate produced (labeled as CACO3 in Figure 4-6) leaves this block to a LIME-KIL
block explained in the next section.
Followed by ketonization, a ketone hydrogenation process continued to produce
alcohols. In reactor R-108 the conversion factor for serial reactions was 1 (Gosseaume,
2011). The equations for hydrogenation are shown in (Appendix A) Eq. A-17 to A-22. The
hydrogenation conditions are high pressure (55 bar) and isothermal (130C). The net
demand of hydrogen is 0.0290 ton H2/ton mixed alcohol, and it is produced in gasification
block explained in last section.
39
The reaction for R-107 is endothermic with enthalpy 4,580kJ/s, but the reaction for R-108 is exothermic with enthalpy -2,615
kJ/s, then heat integration is possible to study. Table 4-9 shows the summary of heat equipment loads. On the other hand, heat balances
for this block shows a power consumption of 1,309 kW for pumps and compressor (Table 4-9). The cooling demand in this block is
4,792kJ/s and the heating demand is 4,003 kJ/s.
Figure 4-6. Ketonization simulation
Table 4-8. Ketonization mass and heat balance
ALCOHOL CACO3 H2 H2-1 H21 KET KET-CACO KETO KETONES
40
ALCOHOL CACO3 H2 H2-1 H21 KET KET-CACO KETO KETONES
Temperature (C) 130 130 43 130 961 130 430 133 130
Pressure (bar_a) 55.0 7.60 0.900 54.8 55.0 7.60 0.0400 55.0 7.60
Mass vapor fraction 0.0190 0 1.00 1.00 1.00 0 0.38 0 0
Mass solid fraction 3x10-3 1.00 0 0 0 0.619 0.619 3x10-3 3x10-3
Mass flow (ton/h) 10 15.5 0.340 0.340 0.340 25.2 25.2 9.6 9.6
Enthalpy (kJ/s) 12,226 51,675 24 142 1,640 61,455 57,409 9,754 9,779
Component mass flow (ton/h)
CACO3 0 15.5 0 0 0 15.5 15.5 0 0
CA(CH-03) 0.0300 0 0 0 0 0.0300 0.0300 0.0300 0.0300
HYDROGEN 0.0500 0 0.340 0.340 0.340 0 0 0 0
ACETONE 0 0 0 0 0 7.00 7.00 7.00 7.00
BUTANONE 0 0 0 0 0 0.200 0.200 0.200 0.200
HEXANONE 0 0 0 0 0 0 0 0 0
PENTANON 0 0 0 0 0 0.500 0.500 0.500 0.500
HEPTANON 0 0 0 0 0 0 0 0 0
NONANONE 0 0 0 0 0 1.90 1.90 1.90 1.90
ISOPROPANOL 7.20 0 0 0 0 0 0 0 0
BUTANOL 0.180 0 0 0 0 0 0 0 0
HEXANOL 2x10-3 0 0 0 0 0 0 0 0
PENTANOL 0.560 0 0 0 0 0 0 0 0
HEPTANOL 0.0100 0 0 0 0 0 0 0 0
NONANOL 1.90 0 0 0 0 0 0 0 0
(Continued Table 4-8)
41
KETS KT-CACO3 OH SAL-DEH SALT SALTS
Temperature (C) -14.6 -15 300 163.1 430 430
Pressure (bar_a) 7.80 0.0400 3.00 6.00 5.50 5.50
Mass vapor fraction 0 0 0.997 0 0 0
Mass solid fraction 0.619 0.619 3x10-3 1.00 1.00 1.00
Mass flow (ton/h) 25.2 25.2 10.0 25.2 25.2 25.2
Enthalpy (kJ/s) 62,893 62,896 9,711 62,572 61,989 61,989
Component mass flow (ton/h)
CACO3 15.5 15.5 0 1.30 1.30 1.30
CA(CH-01 0 0 0 19.2 19.2 19.2
CA(CH-02 0 0 0 1.40 1.40 1.40
CA(CH-03 0.0300 0.0300 0.0300 3.30 3.30 3.30
HYDROGEN 0 0 0.0500 0 0 0
ACETONE 7.00 7.00 0 0 0 0
BUTANONE 0.200 0.200 0 0 0 0
PENTANON 0.500 0.500 0 0 0 0
NONANONE 1.90 1.90 0 0 0 0
ISOPROPANOL 0 0 7.22 0 0 0
BUTANOL 0 0 0.180 0 0 0
PENTANOL 0 0 0.560 0 0 0
HEPTANOL 0 0 0.0200 0 0 0
NONANOL 0 0 1.91 0 0 0
42
Table 4-9. Heat balances for Ketonization equipments
Type Unit Heat duty [kJ/sec] Heat duty [kJ/h]
Heater H-107 583 2.10x106
Heater H-108 905 3.26x106
Heater H-109 2,515 9.05x106
Cooler C-102 3,629 1.31x107
Cooler C-103 1,163 4.19x106
Pumps P-105 3.10 1.12x104
Pumps P-106 25.2 9.07x104
Compressor CM-103 1,281 4.61x106
Reactor R-107 4,580 1.65x107
Reactor R-108 -2,615 -9.41x106
- Lime kiln (Unit 5)
In LIME KIL block the calcium carbonate labeled as CACO3 that come from
KETONIZA block is divided in two streams: (i) the stream labeled as CACO3-2 with a
flow of 9.20 ton/h is converted into quick lime (CaO). And (ii) the second stream labeled as
CACO3-1 with a flow of 6.30 ton/h is recycled to a PRET-FER block for Fermentation
process as was explained in that block before. The lime kiln simulation is shown in Figure
4-7. The conversion factor for Equation 4-5 in the reactor R-109 is 1, with an
endohothermic enthalpy of 4,529 kJ/ (Gosseaume, 2011). Table 4-11 shows the mass and
heat balance of this process.
(Eq. 4-5)
43
Figure 4-7. Lime kiln simulation
Table 4-10. Heat balances for Lime kiln equipments
Type Unit Heat duty [kJ/sec] Heat duty [kJ/h]
Cooler C-104 521 1.88x106
Heater H-110 987 3.55x106
Reactor R-109 4,529 1.63x107
44
Table 4-11. Lime kiln mass and heat balance
CACO3 CACO3-1 CACO3-2 CACO3-3 CAO CAO-CO2
Temperature (C) 130 130 130 500 500 55.0
Pressure (bar_a) 7.60 7.60 7.60 1.00 1.00 1.00
Mass vapor fraction 0 0 0 0 0.440 0.440
Mass solid fraction 1.00 1.00 1.00 1.00 0.560 0.560
Mass flow (ton/h) 15.54 6.3 9.24 9.24 9.24 9.24
Enthalpy (kJ/s) 51,675 20,949 30,725 29,738 25,209 26,310
Component mass flow (ton/h)
CO2 0 0 0 0 4.06 4.06
CACO3 15.54 6.30 9.24 9.24 0 0
CAO 0 0 0 0 5.18 5.18
- Final (Unit 6)
The final block includes the mixed alcohols (stream labeled as OH in Figure 4-8)
conversion produce hydrocarbon fuels by alcohols dehydration olefins oligomerization
(product stream labeled as OLF-C9-12 in Figure 4-8) and olefin hydrogenation (product
stream labeled as PARAFIN in Figure 4-8). The final block simulation is shown in Figure
4-8.
The alcohols dehydration from stream labeled OH to produced 7.30 ton/h of olefins C3
to C9 stream labeled as OLF-C3-9 in Figure 4-8, occurred in reactor R-110, where the
conversion factor is 1 for the reactions shown in Appendix A (Eq. A-23 to A-28)
(Gosseaume, 2011). The heat duty for an endothermic reaction is 1,892kJ/s (Table 4-13).
The olefins produced in R-110 goes to a Oligomerization process to produced 7.30
ton/h of olefins C3 to C12 stream labeled as OLF-C3-12 in Figure 4-8, these reactions were
present in reactor R-111. In Table A-2 (Appendix A) are shown the conversion factors for
45
reactions by the Equations A-29 to A-36.The heat duty for an exothermic reaction is -
1,645kJ/s (Table 4-13).
To improve fuel quality, the olefins labeled as OLEFIN in Figure 4-8 were
hydrogenated to make 7.30 ton/h of corresponding paraffins (stream labeled as PARAFIN
in Figure 4-8) in reactor R-112, where the conversion factor is 1 (Gosseaume, 2011).
Olefin hydrogenation reactions are presented in Equations A-37 to A-45 (Appendix A).
And the heat duty for an exothermic reaction is -2,421kJ/s (Table 4-13). In this block, the
net demand of hydrogen is 0.0190 ton H2/ton hydrocarbon fuels; this hydrogen is produced
in gasification block explained in the next section.
Finally, the hydrocarbon fuel labeled as HC in Figure 4-8 is distilled into C8- and C9+
fractions. The light fraction and the heavy fraction can be used as blending components for
gasoline and jet fuel, respectively, as Pham et al., (2012) mentioned. A ratio of 53 gallons
of light fraction per ton of biomass is obtained for a total of 2,127 gallon/h of gasoline. And
for heavy fraction the ratio is 19 gallons per ton of biomass for a total of 762 gallon/h of jet.
Table 4-12 shows the material balance of the entire block. On the other hand, heat
balances for this block shows a power consumption of 997 kW for compressor CM-104 and
CM-105 (Table 4-13). The cooling demand in this block is 4,470kJ/s and the heating
demand is 1,536kJ/s.
Figure 4-9 shows a comparison between the gasoline obtained by MixAlco and a
mixture of light naphtha (LVN) and gases fossil fuel consulted in an article of Cartagena
refinery (Fernndez, 2007). The gasoline curve obtained by MixAlco had a similar
behavior of LVN except for gas fraction.
46
Figure 4-8. Final simulation
Table 4-12. Final mass and heat balance
C3 H2 H2- H2-1 H20 HC HC-C4--8 HC-C9-12 HEAVY LIGHT
Temperature (C) 130 43 961 130 300 130 266 408 25 25
Pressure (bar_a) 55.0 0.900 55.0 55.0 3.00 55.0 50.0 53.0 1.00 1.00
Mass vapor fraction 0.770 1.00 1.00 1.00 1.00 0 0 0 0 0
Mass solid fraction 0.230 0 0 0 0 0 0 0 0 0
Mass flow (ton/h) 0.130 0.100 0.100 0.100 2.60 7.40 5.30 2.10 2.10 5.30
Enthalpy (kJ/s) 32.0 7.97 537 537 9.17 4.04 2,081 481 1,228 3,382
Component mass flow (ton/h)
47
C3 H2 H2- H2-1 H20 HC HC-C4--8 HC-C9-12 HEAVY LIGHT
WATER 0 0 0 0 2.60 0 0 0 0 0
CA(CH-03 0.0300 0 0 0 0 0 0 0 0 0
HYDROGEN 0 0.100 0.100 0.100 0 0 0 0 0 0
C3H6 0.100 0 0 0 0 0 0 0 0 0
C4H10 0 0 0 0 0 0.0900 0.0900 0 0 0.0900
C5H12 0 0 0 0 0 0.200 0.20 0 0 0.200
C6H14 0 0 0 0 0 4.73 4.73 0 0 4.73
C7H16 0 0 0 0 0 0.0300 0.0300 0 0 0.0300
C8H18 0 0 0 0 0 0.230 0.230 0 0 0.230
C9H20 0 0 0 0 0 1.41 0 1.41 1.41 0
C10H22 0 0 0 0 0 0.0400 0 0.0400 0.0400 0
C11H24 0 0 0 0 0 0.190 0 0.190 0.190 0
C12H26 0 0 0 0 0 0.450 0 0.450 0.450 0
(Continued Table 4-12)
OH OLEFIN OLF OLF-C3-9 OLF-DEH OLFC3-12 PARAFIN
Temperature (C) 300 130 399.9 300 300 300 130
Pressure (bar_a) 3.00 54.5 55.0 3.00 3.00 3.00 55.0
Mass vapor fraction 0.997 0.0390 0.996 0.997 0.996 0.997 7x10-3
Mass solid fraction 3x10-3 4x10-3 4x10-3 3x10-3 4x10-3 3x10-3 4x10-3
Mass flow (ton/h) 10.0 7.40 7.40 10.0 7.40 10.0 7.50
Enthalpy (kJ/s) 9,711 1,699 287 7,819 289 9,464 4,073
Component mass flow (ton/h)
WATER 0 0 0 2.56 0 2.56 0
CA(CH-03 0.030 0.030 0.030 0.030 0.030 0.030 0.030
48
OH OLEFIN OLF OLF-C3-9 OLF-DEH OLFC3-12 PARAFIN
HYDROGEN 0.050 0.050 0.050 0.050 0.050 0.050 0.010
ISOPROPANOL 7.22 0 0 0 0 0 0
BUTANOL 0.180 0 0 0 0 0 0
HEXANOL 0 0 0 0 0 0 0
PENTANOL 0.560 0 0 0 0 0 0
HEPTANOL 0.0200 0 0 0 0 0 0
NONANOL 1.91 0 0 0 0 0 0
C3H6 0 0.100 0.100 5.05 0.100 0.100 0.101
C4H8 0 0.0900 0.0900 0.140 0.0900 0.0900 0
C5H10 0 0.190 0.190 0.450 0.190 0.190 0
C6H12 0 4.62 4.62 0.00 4.62 4.62 0
C7H14 0 0.0300 0.0300 0.0100 0.0300 0.0300 0
C9H18 0 1.38 1.38 1.67 1.38 1.38 0
C8H16 0 0.230 0.230 0 0.230 0.230 0
C10H20 0 0.0400 0.0400 0 0.0400 0.0400 0
C11H22 0 0.180 0.180 0 0.180 0.180 0
C12H24 0 0.450 0.450 0 0.450 0.450 0
C4H10 0 0 0 0 0 0 0.0900
C5H12 0 0 0 0 0 0 0.200
C6H14 0 0 0 0 0 0 4.73
C7H16 0 0 0 0 0 0 0.0300
C8H18 0 0 0 0 0 0 0.230
C9H20 0 0 0 0 0 0 1.41
C10H22 0 0 0 0 0 0 0.0400
C11H24 0 0 0 0 0 0 0.190
C12H26 0 0 0 0 0 0 0.450
49
Table 4-13. Heat balances for Final equipments
Type Unit Heat duty [kJ/sec] Heat duty [kJ/h]
Heater Reboiler T-101 1,536 5.53x106
Cooler C-105 381 1.37x106