Retrofitting analysis to integrate the MixAlco® process to the crude oil distillation process

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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

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Retrofitting analysis to integrate the MixAlco® process to the crude oil distillation

process

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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

Approved by:

Chair of committee, Rocío Sierra Ramírez, PhD. Committee Members, Jorge Mario Gómez, Phd. Head of Department, Oscar Alvarez Solano, PhD.

August 2013

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ABSTRACT

process (August 2013)

Laura Patricia Prada Villamizar, Universidad de los Andes, Colombia

Advisor: Rocío Sierra Ramírez, 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

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increment from MM USD 7.30 to MM USD15.7, and Return On Investment (ROI)

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RESUMEN

process (August 2013)

Laura Patricia Prada Villamizar, Universidad de los Andes, Colombia

Advisor: Rocío Sierra Ramírez, Ph.D.

MixAlco® es una tecnología donde se producen combustibles líquidos de

transporte y / o productos químicos 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 fósiles de alta calidad y fácilmente explotables

disminuyen en todo el mundo. La preservación de los combustibles fósiles de alta

calidad y la viabilidad de una planta MixAlco®, pueden mejorarse mediante la

integración del proceso MixAlco ® en una instalación existente de procesamiento de

combustibles fósiles.

Esta integración es posible gracias a que ambos procesos tienen productos

similares (bio-gasolina, gasolina, bio-jet, jet). Este trabajo evalúa un análisis de

integración entre el proceso de MixAlco® con el proceso de destilación de crudo de

petróleo (PDCP) como caso seleccionado. La metodología propuesta sugiere una

jerarquía de costos con las siguientes herramientas: simulación de procesos,

integraciones de masa y energía, y evaluaciones económicas. Este trabajo inicia

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separado. Después, se hacen comparaciones entre los casos base y la integración de los

dos procesos (la planta resultante se considera como "bio-refinería integrada"). El

resultado más 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

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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. Rocío 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

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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

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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

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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 60°F

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

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TON: Metric ton (1,000kg)

USD: United States dollars

VFAs: Volatile Fatty Acids

VLSTD: Standard Liquid MolarVolume at 60°F

VOC: Variable Operating Cost

VP: Venture Profit

VS: Volatile Solids

WCI: Working Capital Investment

WWT: Waste Water Treatment

ZC: Critical Compressibility Factor

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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

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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

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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

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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

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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

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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

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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

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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

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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

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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,

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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

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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 &

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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)

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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

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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 cost-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.

3.1Description of the proposed methodology

3.1.1 Define needs

In the first step for the retrofitting analysis, it is possible to define the

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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

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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

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3.2Simulation 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

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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, Henry’s 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

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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

Fermentation Temperature (°C) 55.0

Ketonization Temperature (°C) 430

Ketone hydrogenation Temperature (°C) 130

Lime kiln Temperature (°C) 500

Alcohol dehydration Temperature (°C) 300

Oligomerization Temperature (°C) 300

Olefin hydrogenation Temperature (°C) 130

Gasification Temperature (°C) 760

Steam-gas shift Temperature (°C) 254

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.

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Table 3-3. Assay data for crude oil****

%Distilled* Molecular Weight Specific

Gravity Sulfur curve** Viscosity*** (80ºC) Viscosity*** (100ºC)

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

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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

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

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

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Condenser Partial

Temperature tray 1 (°C) 60.0

Temperature tray 8 (°C) 389

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.

3.3Process 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 company´s technology base. Energy provides the necessary heating,

cooling, and shaftwork for those systems.

Because most industrial processes are complicated, performance and economics

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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).

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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

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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 andPrices

(USD / unit)

Fe ed stoc k c osts

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

Ut

il

itie

s 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 @ 353°C (USD/ton) 10.0

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Item Costs andPrices (USD / unit)

Pr

od

u

cts s

ell

in

g

p

rice

s

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

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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 298.15°K.

4.1Simulation 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.

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5. LIME-KIL: Lime kiln process.

6. FINAL: Dehydratation, oligomerization and saturation processes.

7. GASIFICA: Gasification reactor, steam gas shift reactor, and adsorption process

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4.1.1.1MixAlco® 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

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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

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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

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- 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).

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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, García, & 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.

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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

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.

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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

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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

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

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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

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

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

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

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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

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

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

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Table 4-5. Heat balances for Pretreatment and Fermentation equipments

Cooler C-101 117 4.21x105

Heat Exchanger E-101 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

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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)

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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

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

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

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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

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

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

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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

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 (430°C), 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

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Table 4-7. Heat balances for Dewatering equipments

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-106 958 3.45x106

Heat Exchanger E-107 42.0 1.51x105

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 (130°C). The net

demand of hydrogen is 0.0290 ton H2/ton mixed alcohol, and it is produced in gasification block explained in last section.

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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

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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

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

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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

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

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Table 4-9. Heat balances for Ketonization equipments

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.

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Figure 4-7. Lime kiln simulation

Table 4-10. Heat balances for Lime kiln equipments

Cooler C-104 521 1.88x106

Heater H-110 987 3.55x106