palm oil research

REACTOR DESIGN AND COST FOR PRODUCING BIODIESEL FROM CANOLA OIL FOR lOMILLION GALLONS PER YEAR CONCEPTUAL PLANT

A Thesis

Submitted to the Graduate Faculty of the University of South Alabama

in partial fulfillment of the requirements for the degree of

Master of Science

in

Chemical Engineering

by Shali Vemparala

B.TECH, BRECW (Affiliated with JNT University), 2007 May 2010

UMI Number: 1484491

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THE UNIVERSITY OF SOUTH ALABAMA COLLEGE OF ENGINEERING

REACTOR DESIGN AND COST FOR PRODUCING BIODIESEL FROM CANOLA OIL FOR 10 MILLION GALLONS PER YEAR CONCEPTUAL PLANT

BY

Shali Vemparala

A Thesis

Submitted to the Graduate Faculty of the University of South Alabama

in partial fulfillment of the requirements for the degree of

Master of Science

in

Chemical Engineering

May 2010

Date: Q - 2 ) d 4 / 10

oijerhj to

z/y/o

Approved:

Chair of Thesis Com] Jagdhish C. Dhawan

Committee Member: Dr. Srinivas Palanki k l H i ^-.

K^ang-Tii Committee Member: Dr. Kwang-Ting Hsiao CCU ._..

Chair of Department: Dr. Srinivas Balanki

Dean of the Graduate School: Dr. B. Keith Harrison

ACKNOWLEDGEMENTS

This work would not have been possible without the support and encouragement of

Dr. Jagdhish C. Dhawan, under whose supervision I chose this topic and began this

thesis. I appreciate the assistance that he has given me throughout my research at

University of South Alabama. I would like to thank Dr. Srinivas Palanki and Dr. Kuang-

Ting Hsiao for serving on my committee. I would also like to thank the rest of the

academic staff of Department of Chemical and Biomolecular Engineering for their

cooperation. I cannot end up without thanking my family and friends Mohan, Chaitanya,

Shwetha and Vijyanthi on whose constant encouragement and love, I have relied

throughout my time at the university.

ii

TABLE OF CONTENTS

Page

LIST OF TABLES vi

LIST OF FIGURES viii

LIST OF ABBREVIATIONS x

NOMENCLATURE xi

ABSTRACT xiv

CHAPTER 1: INTRODUCTION 1

1.1 Energy Consumption in World 1 1.2 Biodiesel 2 1.3 Scope and Objectives 3 1.4 Significance of the Research 4

CHAPTER 2: LITERATURE SURVEY 5

2.1 History of Biodiesel 5 2.2 Aliphatic Fatty Acid Chains 6 2.3 Transesterification of Oils 8

2.3.1 Base Catalyzed Transesterification Mechanism 9

2.4 Conceptual Biodiesel Process Block Flow Diagram (BFD) 10 2.5 Byproduct (Glycerol) Utilization Potential 12 2.6 Literature Summary - Rate Constants 12

2.6.1 Reaction Mechanism and Kinetic Rate Data 16

iii

CHAPTER 3: BIODIESEL REACTOR DESIGN 20

3.1 Plug Flow Reactor (PFR) Design 20 3.2 Performance Equations for a Plug Flow Reactor (PFR) 22 3.3 MathCAD Solution to PFR Model 25

3.3.1 Effect of Methanol to Canola Oil Feed Ratio on Conversion 34

3.4 Aspen Plus Simulation of PFR 34

3.4.1 Plug Flow Reactor Size for Diameter 0.038m (1.5 inch OD) and 0.051m (2.0 inch OD) Tubes 36

3.4.2 Effect of Methanol to Canola Oil Molar Feed Ratio on Conversion 43 3.4.3 Plug Flow Reactor Length Required for 1.5 inch OD and 2.0 inch OD Tubes43 3.4.4 Pressure Drop Across Plug Flow Reactor 44 3.4.5 Plug Flow Reactor Cost 46

CHAPTER 4: CONTINUOUS STIRRED TANK REACTOR (CSTR) 48

4.1 Performance Equations Model for a Continuous Stirred Tank Reactor (CSTR)... 48 4.2 MathCAD Solution to CSTR Model 50 4.3 Aspen Plus Simulation of CSTR Model 51

4.3.1 Effect of Residence Time on Conversion 55 4.3.2 Effect of Reactor Volume on Conversion 56

4.4 Design Dimensions of the CSTR at Methanol/Oil Molar Feed Ratio of 12 56 4.5 Continuous Stirred Tank Reactor Cost 57 4.6 Continuous Stirred Tank Reactor in Series 57

4.6.1 A System of Three CSTR Biodiesel Reactors in Series 57 4.6.2 Economic Analysis of Three CSTR in Series 62 4.6.3 Design of Methanol Recovery Column 65

4.6.3.1 Column Diameter 65 4.6.3.2 Tray Hydraulics 65 4.6.3.3 Dry Pressure Drop 66 4.6.3.4 Check for Down Comer Residence Time 69 4.6.3.5 Weeping Check 69

CHAPTER 5: CONCLUSIONS AND RECOMMENDATIONS FOR FUTURE RESEARCH 71

5.1 Conclusions 71 5.2 Recommendations for Future Research 74

iv

REFERENCES 75

General References 76

APPENDICES 77

Appendix A: MathCAD Program for Solution of PFR for Diameter 0.038m (1.5 inch OD) and 0.051m (2.0 inch OD) Tubes 77 Appendix B: PFR Sensitivity Results from Aspen Plus Simulation for Diameter 0.038m (1.5 inch OD) and 0.051m (2.0 inch OD) Tubes 82 Appendix C: Shell/ Tube Configuration of a PFR 92 Appendix D: PFR Pressure Drop in 0.038m (1.5 inch OD) and 0.051m (2.0 inch OD) Tubes 93 Appendix E: Plug Flow Reactor Cost for Diameter 0.038m (1.5 inch OD) and 0.051m (2.0 inch OD) Tubes 95 Appendix F: MathCAD Program for a Solution to CSTR 99 Appendix G: CSTR Sensitivity Results from Aspen Plus Simulation 102 Appendix H: Sizing of Single CSTR 110 Appendix I: A Graphical Representation of Purchased Cost of Jacketed and Stirred Reactors 111 Appendix J: AspenPlus Input File for Three CSTR in Series with Pump Around System 112 Appendix K: Design of Methanol Recovery Column Calculations 127

BIOGRAPHICAL SKETCH 139

v

LIST OF TABLES

Page

Table 1: Fatty acid composition (%wt) in different types of oils 6

Table 2: Literature summary of biodiesel production from various types of oils 13

Table 3: k values for soybean oil are reported by Noureddini and Zhu, 1997 16

Table 4: k values for vegetable oil are reported by Sharma, 2008 17

Table 5: k values for palm oil are reported by Leevijit, 2004 17

Table 6: k values for soybean oil are reported by Marchetti, 2007 17

Table 7: k values for vegetable oil are reported by Komers, 2002 (based on particular data regression) 18

Table 8: Reaction kinetics relating rate constants, activation energy and Arrhenius constant (Noureddini et al., 1997) 19

Table 9: MathCAD results for diameter 0.038m (1.5 inch OD) tube 26

Table 10: MathCAD results for diameter 0.051m (2.0 inch OD) tube 27

Table 11: Summary of ASPEN results of PFR tube OD 0.038m (1.5 inch) 36

Table 12: Summary of ASPEN results of PFR tube OD 0.051m (2.0 inch) 36

Table 13: Cost of plug flow reactor for different diameters 47

Table 14: Summary of MathCAD results for different reactor volume and methanol to oil molar feed ratio: 51

Table 15: Summary of ASPEN results of CSTR volume 7m3 52

Table 16: Material balance data for three CSTR's in series to produce biodiesel from canola oil (Plant capacity: 10 million gallons/year of biodiesel product) 60

vi

Table 17: Below summarizes the results of overall conversion as a function of methanol to canola oil feed ratio 62

Table 18: Total utility cost ($/hr) as a function of methanol to canola oil mole feed ratio 62

Table 19: Cost of methanol recovery distillation column 63

Appendix

Table 20: Aspen stream results for various methanol to canola oil feed ratios as a the function of reactor length for diameter 0.038m (1.5 inch OD) 90

Table 21: Aspen stream results for various methanol to canola oil feed ratios as a the function of reactor length for diameter 0.051m (2.0 inch OD) 91

Table 22: Aspen steam results for various methanol to canola oil molar feed ratio as function reactor volume of CSTR 109

vii

LIST OF FIGURES

Page

Figure 1: World fuel consumption in year 2008 1

Figure 2: Schematic diagram of biodiesel production 3

Figure 3: Schematic diagram of biodiesel production process 11

Figure 4: Canola oil reacts with methanol in PFR to generate biodiesel and glycerol 22

Figure 5: Plug flow reactor length (m) versus conversion of canola oil to biodiesel for diameter 0.38m (1.5 inch OD) at 323K (50°C) and 4.053 x 105 Pa (4atms) 28

-J Figure 6: Length (m) of plug flow reactor versus component concentrations (kmol/m ) for diameter 0.038m (1.5 inch OD) at 323K (50°C) and 4.053 x 105Pa (4atms) 29



Figure 9: Length of reactor (m) versus component concentrations (kmol/m ) for diameter 0.051m (2.0 inch OD) at 323K (50°C) and 4.053 x 105 Pa (4atms) 32


Figure 11: ASPEN simulation of PFR diameter-0.038m (1.5 inch) and 0.051m (2.0 inch) OD 35

Figure 12: Plug flow reactor residence time versus conversion of oil to ester for diameter 0.038m (1.5 inch OD) at 323K (50°C) and 4.053 x 105Pa (4atms) 37

Figure 13: Length of plug flow reactor versus component flow rate for diameter 0.038m( 1.5 inch OD) at 323K (50°C) and 4.053 x 105 Pa (4atms) 38

viii

Figure 14 : Plug flow reactor residence time versus conversion of oil to ester for diameter 0.38m (1.5 inch OD), 323K (50°C) and 4.053 x 105 Pa (4atms) 39

Figure 15: Plug flow reactor residence time versus conversion of canola oil to biodiesel for diameter 0.051m (2.0 inch OD) at 323K (50°C) and 4.053 x 105 Pa (4atms) 40

Figure 16: Length of plug flow reactor versus component flow rate for diameter 0.038m (2.0 inch OD) at 323K (50°C) and 4.053 x 105Pa (4atms) 41

Figure 17: Plug flow reactor residence time versus conversion of canola oil to biodiesel for diameter 0.051m (2.0 inch OD), at 323K (50°C) and 4.053 x 105 Pa (4atms) 42

Figure 18: 180 degree tube bends of a PFR 44

Figure 19: Canola oil reacts with methanol in a CSTR to produce biodiesel and glycerol 48

Figure 20: ASPEN Simulation of CSTR volume 7m3 52

Figure 21: Continuous stirred tank reactor residence time versus conversion of canola oil to biodiesel for methanol to oil ratio 3, 6, 9, 12 varying along the volume of reactor at 323K (50°C) and 4.053 x 105Pa (4tams) 53

Figure 22: Volume of Continuous stirred tank reactor versus component flow rate for methanol to oil ratio 12 at 323K (50°C) and 4.053 x 105Pa (4tams) 54

Figure 23: Continuous stirred tank reactor residence time versus conversion of canola oil to biodiesel for methanol to oil ratio 12 varying along the volume of reactor at 323K (50°C) and 4.053 x 105Pa (4tams) 55

Figure 24: ASPEN simulation of three CSTR's in series 59

Figure 25: Methanol to oil mole ratio versus total cost 64

Figure 26: Schematic diagram of distillation column 65

Appendix

Figure 27: Conversion of canola oil to biodiesel versus length of reactor for diameter 0.038m (1.5 inch OD) tube 79

Figure 28: Conversion of canola oil to biodiesel versus length of reactor for diameter 0.051m (2.0 inch OD) tube 81

Figure 29: Purchased cost of jacketed and stirred reactors (Peters et al., 2003) 111

ix

LIST OF ABBREVIATIONS

FAME Fatty Acid Methyl Esters

PFR Plug Flow Reactor

CSTR Constant Stirred Tank Reactor

LHSV Liquid Hourly Space Velocity

x

NOMENCLATURE

English Letters

Symbol Description Units

A Arrhenius constant L/mol*min

Ar Area of the reactor m

Area Area of plug flow reactor m2

BD Biodiesel unitless

C concentration mol/m

CONV Conversion of canola oil to biodiesel unitless

Cost Cost of Plug Flow Reactor (2008) $

DG Diglyceride unitless

DSheii Shell inside diameter m

E Activation Energy J/mol

F Friction factor unitless

F flow rate mol/s

Fm Material Factor for a 316 stainless steel pipe unitless

GY Glycerol unitless

ID Tube inside diameter m

L Assumed length of a shell m

xi

Lreactor Total length of a PFR m

J TG, DG, MG, A, BD, GY. Unitless

k Rate constant m3/mol.min

Kf Frictional loss in a 180 degree bend unitless

mins minutes minutes

MG Monoglyceride unitless

ME Methanol unitless

NtUbes Total number of tubes in a PFR unitless

OD Tube out side diametr m

APrcactor Pressure drop in PFR Pa

APtotal Pressure drop in all the tubes of a PFR Pa

r Triglyceride of canola oil unitless

R Rate of reaction min"1

R' Alcohol group unitless

TG Triglyceride unitless

V Volume of reactor m3

Wj Conversion of canola oil to biodiesel for a methanol unitless to oil feed ratio of 3

Xj Conversion of canola oil to biodiesel for a methanol unitless to oil feed ratio of 6

Yj Conversion of canola oil to biodiesel for a methanol unitless to oil feed ratio of 9

Zj Conversion of canola oil to biodiesel for a methanol unitless to oil feed ratio of 12

xii

Greek Letters

Symbol Description Units

P Density of liquid in side a tube kg/ m3

v volumetric flow rate m /s

v velocity in a tube m/s

0 residence time s

r residence time s

xiii

ABSTRACT

Vemparala, Shali, M.S., University of South Alabama, May 2010, Reactor Design and Cost for Producing Biodiesel from Canola Oil for 10 Million Gallons Per Year Conceptual Plant. Chair of Committee: Dr. Jagdhish C. Dhawan.

Biodiesel can be produced from many natural renewable sources (vegetable oils,

animal fats, algae etc). The present study concentrates on production of biodiesel using

canola oil and methanol as reactants at 298K (25°C) and 4.053 x 105 Pa (4atms) in the

presence of sodium methylate acting as catalyst. A plug flow (PFR) and continuous

stirred tank (CSTR) reactors are designed using the rate expressions available from the

literature. Reactor performance was evaluated with respect to conversion versus reactor

volume and the effect of methanol to oil molar feed ratio on conversion at 323K (50°C)

and 4.053 x 105 Pa (4atms) was also evaluated. A plug flow reactor of 1.50 inch OD

requires a total length of 2580 meters. The reactor length can be decreased to 1140 meters

when the diameter is 2.0 inch OD. The reactor pressure drop is significantly high and

ranges from 1.317 x 106 Pa (13atms) to 3.445 x 106 Pa (34atms) depending upon the

reactor length. A single CSTR of 7m volume provides 90.3% conversion. However, if

three 5m3 volume CSTR reactors are used in series, an overall conversion of 99.9% can

be achieved. Three CSTR reactors in series with a pump around system for thorough

mixing are recommended. A conversion of 99.9% eliminates the product purification

step to recover the un-reacted material from the product stream.

xiv

CHAPTER 1: INTRODUCTION

1.1 Energy Consumption in World

Energy is an integrated part of human life. Energy can be obtained from two different

sources, renewable and non renewable. The primary source of energy has been from non

renewable sources which are fossil fuels. These fuels include coal, oil and natural gas.

Fuel consumption in the world is shown in the Figure 1 (BP Statistical Review, 2008).

Hydro ^^.Nuclem

6.50%

Figurel: World fuel consumption in year 2008.

The fossil fuel reserves in the world are so unevenly distributed that many countries

have to depend on other countries for their requirements to be fulfilled. The recovery and

1

processing of fossil fuels is known to damage the environment we live in. When fossil

fuels undergo combustion acids like carbonic, sulfuric and nitric are released, which are

the main cause of acid rains. Small amounts of radioactive materials like uranium and

thorium are also present in fossil fuels; hence these together harm the environment when

released into the atmosphere (Gabbard, 1993). All these factors necessitate continued

search and sustainable development of renewable energy sources such as biofuels that are

environmentally friendly.

1.2 Biodiesel

Energy derived from biological sources is Bio-energy. Bio-energy sources are

biomass (e.g. forest residue), biogas (e.g. methanol), biofuel (biodiesel, bio-ethanol and

bio-methanol). Vegetable oils are used as raw material for biofuels to serve as an

alternative source of transportation fuel.

Biodiesel is a sulfur-free clean burning alternative fuel, which can be produced from

domestic renewable resources. Biodiesel contains no petroleum. It can be blended at any

concentration with petroleum-derived diesel to make a biodiesel blend. It can be used in

diesel engines with little or no modifications. Biodiesel can easily be adapted to reduce

gasoline consumption and is free of sulfur and aromatics.

These biofuels are non toxic, renewable and they are not associated with adverse

effects on the environment because they emit less harmful emissions and green house

gases.

2

1.3 Scope and Objectives

In this thesis a Plug Flow Reactor (PFR) and a Continous Stirred Tank Reactor

(CSTR) are designed to produce 10 million gallons of biodiesel (Haas et al., 2006 and

Noureddini and Zhu, 1997) from Canola oil.

Figure 2: Schematic diagram of biodiesel production.

The specific objectives are as follows:

1) Design a PFR and a CSTR; make recommendations for reactor choice to produce

biodiesel from Canola oil.

2) Evaluate the impact of methanol to canola oil mole feed ratio on the reactor conversion

3) Conduct Aspen simulations for reactor effluent separation into biodiesel product,

glycerol co-product and recovery and recycle of excess methanol

4) Carry out preliminary equipment sizing of PFR, CSTR and, a sieve tray distillation

column for methanol recovery.

5) Estimate and compare the costs of a PFR and a CSTR.

3

1.4 Significance of the Research

Biodiesel is a liquid transportation fuel that can be produced from renewable raw

material such as Canola oil. Reaction kinetics of transesterification of vegetable oils into

Fatty Acid Methyl Esters (FAME) called Biodiesel, have been experimentally established

and are well documented in the open literature. However, to produce 10 million gallons

of biodiesel per year in a large capacity commercial-size plant, a detailed design and

analysis of biodiesel reactor is required. The main significance of this research is the

design of a suitable PFR or CSTR under most economical conditions.

An important property of biodiesel is that it can be directly blended in any petroleum

based liquid fuels such as diesel. Biodiesel and petroleum diesel blends are designated

with letter ' B \ For example B20 indicates 20% of biodiesel and 80% of petroleum diesel

and B100 indicates pure biodiesel. Biodiesel blends perform better than petroleum

diesel. Biodiesel can also be used in automobiles without any blending with petroleum

derived diesel (National Biodiesel Board, 2001).

Lubricity is an important property of diesel fuel. The National Biodiesel Board found

that one half of samples of petroleum based diesel sold in the Unites States did not meet

the current standard for lubricity (National Biodiesel Board, 2001). Biodiesel has better

lubricity than low sulfur petroleum diesel (500 ppm of sulfur by wt). In 2006 ultra-low

sulfur petroleum (15pp of sulfur by wt) was introduced, which has even less lubricity. A

small percentage (1-2%) of biodiesel in low sulfur petroleum improves lubricity.

In this research we investigate the feasibility of production of biodiesel in a PFR and a

CSTR. The reaction kinetics and rate constants used are those that have been

experimentally determined by Noureddini and Zhu, 1997.

4

CHAPTER 2: LITERATURE SURVEY

2.1 History of Biodiesel

In 1853 Duffy and Patrick conducted transesterifiction of vegetable oil, before diesel

engine came into existence. Diesel's model, an iron cylinder 1 Oft (3m) long with a

flywheel as its base, was able to run on its own power for the first time in Augsburg,

Germany on Aug 10, 1893. This day is now acknowledged as "International biodiesel

day" (Knothe, 2001).

A diesel engine running on peanut oil built by French Otto Company was

demonstrated by Diesel inl900 at the World Fair in Paris, where it received the Grand

Prix award.

The petroleum industry entered into fuel markets because petroleum oil-based fuel

was much cheaper to produce than biomass-derived fuel. Hence in the 1920s, diesel

engine manufacturers were altering their engines to accommodate a lower viscosity

petroleum fuel rather than vegetable oils. As a result biomass based fuels have been

neglected (Knothe, 2001).

A Beligian Patent granted in 1937 to Chavanne, constitutes the first report on what is

today known as biodiesel (Chavanne, 1938 and Knothe, 2001). It describes the use of

ethyl esters of palm oil as a diesel engine fuel. These esters were obtained by acid

catalyzed transesterification of the oil.

5

2.2 Aliphatic Fatty Acid Chains

Table 1 shows the typical range of saturated and unsaturated fatty acids concentration

(mass %) of vegetable oils suitable as raw materials to produce biodiesel fuel (McCance

etal., 1991).

Table 1: Fatty acid composition (%wt) in different types of oils.

Saturated Unsaturated Mono -

unsaturated Poly -unsaturated

Capric acid

Laurie acid

Myristi c acid

Palmitic acid

Stearic acid

Oleic acid Linoleic acid

Alpha linoleic

acid Soybean

oil - - - 11% 4% 24% 54% 7%

Canola oil - - - 4% 2% 62% 22% 10%

Palm oil - - 1% 45% 4% 40% 10% -

Vegetable oil

- - - 15% 5% 25-30% 45-50% 2-3%

Chemical structures of typical fatty acids are given below:

1) Myristic acid (C14H28O2)

O

H3C H2C H2C H2C H2C H2C H2C OH

2) Palmitic acid (C, 6H3202)

H3C H2C H2C H2C H2C H2C H2C H2C C OH

6

3) Stearic acid (CibH3602)

0

H3C H2C H2C H2C H2C H2C H2C H2C H2C OH

4) Oleic acid (Ci8H3402)

P

H3C H2C H2C H2C HC H2C H2C H2C H2C OH

5) Linoleic acid (Ci8H3202)

0

H3C H2C H2C HC HC H2C H2C H2C H2C OH

6) Alpha Linoleic acid (Ci8H30O2)

H3C H2C HC H2C HC H2C H2C H2C H2C OH

7

2.3 Transesterification of Oils

The transesterification process is the most common chemical approach used for the

production of biodiesel. This process is also called an alcoholysis process. It is a reaction

between an ester and an alcohol in which the -O -R group of the ester and the -O-R'

group of the alcohol trade places, as shown below (Schuchard et al., 1998):

X O + H O R • X O + H O R (2.1)

There are several methods for carrying out the transesterification reaction including

the supercritical processes (higher than critical temperature of methanol is considered,

thus eliminating the requirement of catalyst), ultrasonic methods (influence of different

sound frequencies versus traditional stirring) and microwave method (microwave

irradiation than conventional heating for synthesis of FAME from triglycerides).

The most commonly used catalysts for above process are acid catalyst (H2SO4 and

HC1) base catalyst (NaOH and KOH) and enzyme catalyst (lipase). Sodium methylate

(CHsONa) can also be used directly without having to use sodium or potassium

hydroxide crystals. Sodium methylate is soluble in the reacting system at all

compositions.

8

Transesterfication reaction chemistry: O

£H2 - O - C - Ri O

0

CH3 - O - C - Ri + 0

C H - O - C - R 2 +3CH3OH 0

• C H 3 - O - C - R 2 (KOHorNaOH) + Q

+

CH2 - O - € - R3 CH3 - O - C - R3

Triglyceride Methanol Mixture of fatty esters

Transesterification reaction stoichiometry:

Vegetable Oil + 3 MeOH -(C57H,O406) (CH3OH)

3 Biodiesel + Glycerol (C19H3602) (C3H803)

CH2-OH

C H - O H

:H2 - OH

Glycerol (2.2)

(2.3)

2.3.1 Base Catalyzed Transesterification Mechanism

In the first step base reacts with alcohol, producing an alkoxide and the protonated

catalyst. A tetrahedral intermediate is generated when nucleophile of alkoxide attacks the

carbonyl group of triglyceride, from which the alkyl ester and anion of the diglyceride are

produced. "The latter deprotonates acts as a catalyst, thus regenerating the active species,

which is now able to react with a second molecule of the alcohol, which starts another

catalytic cycle". By using the same mechanism, diglycerides and monoglycerides can be

converted to alkyl esters and glycerol (Schuchard et al., 1998).

R - O H + Base <« • RO" + BaseH+ (2.4)

9

R,COO-CH 2

R2COO - CH

H2C - O"

R" <-

BH

RiCOO - <pH

—• R 2 COO-CH OR (2.5)

H2C - O - c - R3

O"

R,COO-CH 2

-*R2COO - CH + RCOOR3 (2.6)

H2C - O"

RiCOO-CH 2

R 2 C O O - J H + B (2.7)

H2C - OH

2.4 Conceptual Biodiesel Process Block Flow Diagram (BFD)

A conceptual biodiesel process Block Flow Diagram is shown in Figure 3. Methanol

and sodium methylate solution are mixed and fed to a mixer. This mixture along with

vegetable oil is sent to a reactor R-l with a steam jacket to heat the reacting mixture to

about 333K. Transesterification reaction takes place between oil and methanol to form

FAME and glycerol. FAME and glycerol along with excess methanol are sequentially

separated in two distillation columns D-l and D-2. Recovered methanol is recycled back

to the reactor.

10

j St-1 - Methanol Storage tank

I St-2 - Catalyst Storage I tank

j St-3 - Oil Storage tank M-l - Methanol/Catalyst Mixer

R-l- Transesterification Reactor D-l -Methanol/FAME

Figure 3: Schematic diagram of biodiesel production process.

11

2.5 Byproduct (Glycerol) Utilization Potential

As can be seen from the reaction stoichiometry, glycerol is the only byproduct in the

biodiesel manufacturing process. Approximately 1 kg of glycerol is produced for every 9

kg of biodiesel. Currently, glycerol is used in the making of many products which

include personal care products, food, oral care, tobacco, polymers, cosmetics, soap and

pharmaceuticals. Glycerol is also used in production of monoglycerides, diglycerides

and pyritic acid. Other applications include production of 1, 2-propanediol and 1, 3-

propanediol by hydrogenation of glycerol replacing the use of petroleum feedstock.

Catalyzed liquid-phase etherification of glycerol to oligomers is another application. The

preparation of the alkyl ethers of glycerol by etherification with isobutylene or other

olefins is one of the possibilities of glycerol utilization. Glycerol pyrolysis followed by

water gas shift reaction may provide a route to generate hydrogen required to operate fuel

cells (Byrd et al., 2008 and Klepacova et al., 2007).

2.6 Literature Summary - Rate Constants

Different oils and catalyst used for transesterification reactions, oil to alcohol molar

ratio, reaction temperature and time as reported in the literature are summarized in the

chronological order in Table 2 (Sharma et al., 2008).

12

Table 2: Literature summary of biodiesel production from various types of oils.

Year Feed stock Solvent Used

Molar ratio

(methanol :oil)

catalyst Reaction temp(K)

Duration Coversion/ Yeild

2004 Sunflower

oil Supercritical

methanol 40:1 No catalyst

473-673 (P= 200

bar) 10-40 mins

78-96% conversion

with increase in

temperature Supercritical

ethanol 23%

coversion Methanol or

Ethanol 5:1 Supercritical

C02+lipase (Novozym4 35) 30% wt

of oil

318 6 hrs 27% conversion

2005

Pongamia pinnata

Methanol 10:1 Different catalysts

378 1.5 hrs 92% coversion

with KOH, 83% with ZnO, 59% withHb

Zeolite, 47% with

Montmorilno nite

2005 Madhuca indica

Methanol 0.30-0.35v/v -0.25 v/v

l%v/v H2S04, 0.7 wt% KOH

333 1 hr 98% yield

2005 Rubber-seed oil

Methanol 6:1 H2S04 0.5% by volume

318 20-30 mins

9:1 NaOH 0.5% by volume 318 30 mins

13

Table 2: continued

Year Feed stock Solvent Used

Molar ratio

(methano l:oil)

catalyst Reaction temp(K)

Duratio n

Coversion/ Yeild

2004 Sunflower

oil Supercritical

methanol 40:1 No catalyst

473-673 (P= 200

bar) 10-40 mins

78-96% conversion

with increase in

temperature Supercritical

ethanol 23%

coversion Methanol or

Ethanol 5:1 Supercritical

C02+lipase (Novozym4 35) 30% wt

of oil

318 6 hrs 27% conversion

2005

Pongamia pinnata

Methanol 10:1 Different catalysts

378 1.5 hrs 92% coversion

with KOH, 83% with ZnO, 59% with Hb

Zeolite, 47% with

Montmorilno nite

2005 Madhuca indica

Methanol 0.30-0.35v/v -0.25 v/v

l%v/v H2S04, 0.7 wt% KOH

333 1 hr 98% yield

2005 Rubber-seed oil

Methanol 6:1 H2S04 0.5% by volume

318 20-30 mins

9:1 NaOH 0.5% by volume 318 30 mins

14

Table 2: continued

2006 Chlorella

protothec-oides Methanol 56:1 Acid catalyst 303

coversion > 80%

2006 Chlorella

protothec-oides Methanol 56:1

H2S04 (100%) on the basis of

oil wt 303 4 hrs 63% yield

2006 Neat

Cannola oil Methanol 6:1

NaOH 1.0 %wt 318 15 mins

Ester content 98wt%

Used Frying oil 7:1

NaOH 1.1 %wt 333 20 mins

Ester content 94.6wt%

2006 Nicotiana tabacum Methanol 18:1

H2S04(1% with low

molar ratio) 333 25 mins

Yield 91% in 30min

6:1

KOH(l% based on oil

wt) 39 mins

2006 Pongamia pinnata Methanol 6:1

KOH(" 1% by wt) 338 2 hrs

Yield 97-98%

2006

Soybean oil Methanol 4.5:1 Tio2/Zr02( 1 lwt% Ti)

A1203/Zr0 2(2.6%wt

Al) K20/Zr02( 3.3 wt% K)

448 2 hrs conver over 95%

coversion over 100%

15

2.6.1 Reaction Mechanism and Kinetic Rate Data

K, C57H104O6 + CH3OH > C39H72O5 + C19H36O2 AH°r =-4752 KJ/Kmole (2.7) Canola Oil Methanol < Diglyceride Biodiesel

K2

K3 C39H72O5 + CH3OH > C21H40O4 + C19H36O2 AH°r = 91060 KJ/Kmole (2.8) Diglyceride Methanol < Monoglyceride Biodiesel

K4

K5 C21H40O4 + CH3OH > C 3 H 8 0 3 + C i 9 H 3 6 0 2 AH°R = -95340 KJ/Kmole (2.9) MonoGlyceride Methanol< Glycerol Biodiesel

K6

Overall Reaction:

C57Hi04O6 + 3 CH3OH > 3 C,9H3602 + C3H803 AH°r = -4770 KJ/Kmole (2.10) Canola Oil Methanol < Biodiesel Glycerol

The second order reversible reaction kinetic data reported by various investigators is

listed in the following tables:

Table 3: k values for soybean oil are reported by Noureddini and Zhu, 1997.

Rate constant lit/mol*min m3/mol*hr kl 0.05 0.003 k2 0.11 0.0066 k3 0.22 0.0129 k4 1.23 0.07368 k5 0.24 0.01452 k6 0.01 0.0042

16

Table 4: k values for vegetable oil are reported by Sharma et al., 2008.

Rate constant m3/mol*sec m3/mol*hr kl 8.33 x 10"5 0.2998 k2 8.217 x 10"5 0.2958 k3 4.945 x 10"4 1.7802 k4 5.9 x 10"6 0.0212 k5 4.9838 x 10"5 0.1794 k6 1.317 x 10"5 0.0474

Table 5: k values for palm oil are reported by Leevijit et al., 2004.

Rate constant m3/mol*sec m3/mol*hr kl 1.057 x 10"5 0.0381 k2 0 0 k3 1.184 x 10"4 0.4262 k4 8.187 x 10"5 0.2947 k5 1.31 x 10"4 0.4716 k6 2.011 x 10"6 0.0072

Table 6: k values for soybean oil are reported by Marchetti et al., 2007.

Rate constants (L/mol*min) kl 0.049 k2 0.102 k3 0.218 k4 1.28 k5 0.239 k6 0.007 k7 7.84 x 10"5

k8 1.58 xlO"5

17

In the reference Marchetti et al, 2007, k7 and k8 are rate constants of overall

reversible reaction.

Table 7: k values for vegetable oil are reported by Komers et al., 2002 (based on particular data regression).

Method Runge Kutta Gespi Averaging mehods k(m3/mol*sec)

k2 1.297 x 10"4 8.415 x 10"5 8.342 x 10"5

k2r 6.405 x 10"5 4.967 x 10"5 5.908 x 10"5

k4 1.432 x 10"4 1.567 x 10"4 8.215 x 10"5

k4r 4.213 x 10"5 1.219 x 10"4 4.978 x 10"5

k6 3.942 x 10"4 2.53 x 10"4 4.945 x 10"4

k6r 1.172 x 10"5 1.164 x 10"4 1.322 x 10"5

k8 3.9 x 10"7 2.217 x 10"7 2.117 x 10"6

k9 2.27 x 10"6 5.727 x 10"5 3.233 x 10"6

klO 5.84 x 10"6 1.246 x 10"5 5.985 x 10"6

k l l 3.208 x 10"6 7.757 x 10"5 1.212 x 10"4

In reference Komers et al., 2002, the first six rate constants are related to the

transesterification reaction (klis k2, k2 is k2r, k3 is k4, k4 is k4r, k5 is k6 and k6 is k6r).

The last four rate constants k8, k9, klO and kl 1 are related to saponification reactions. If

free fatty acids are present in the feed oil, these undergo a saponification reaction.

From these data the forward reaction rates are controlling when compared with

reverse reactions, as the activation energy is high for the forward reactions. Hence the

equilibrium lies in the formation of products.

Reaction rate constants for soybean oil transesterifications have been experimentally

determined by Noureddini and Zhu and are well accepted by other researchers (Sharma,

18

Komers and Marchetti etc). They employed a methanolic solution of NaOH as catalyst

(concentration of catalyst is 0.02% by wt of canola oil) and the reaction temperature was

323K. The results of kinetic parameters were related to Arrhenius equations expressed

in the form of a power law as shown below.

k ( T ) = A . f rp \

VToy

- E

R T (2 .11)

Numerical values of constants for transesterification of soybean oil are listed in Table 8.

Table 8: Reaction kinetics relating rate constants, activation energy and Arrhenius constant (Noureddini and Zou., 1997).

Reaction K E A (L/mol*min) (J/mol) x 104 (L/mol*min)

1 TG + ME—> DG + BD 0.05 5.504 4.008 x 10' 2 DG + BD—> TG + ME 0.11 4.158 5.874 x 105

3 DG + ME—> MG + BD 0.215 8.315 6.093 x 1012

4 MG + BD—> DG + ME 1.228 6.129 1.012 x 1010

5 MG + ME—> BD + GY 0.242 2.688 5.41 x 10j

6 BD + GY—> MG + ME 0.007 4.014 2.186 x 104

19

CHAPTER 3: BIODIESEL REACTOR DESIGN

Alkali catalyzed transesterification reactions are essentially liquid phase reactions.

The volume change is negligible. Liquid phase reactions are easily carried out either in a

Plug Flow (PFR) or in a Continuous Stirred Tank (CSTR) reactor. As discussed earlier,

there are three series and reversible reactions associated with transesterification of canola

oil. The overall heat of reaction is -4770 KJ/Kmol, which indicates a small increase in

temperature. The temperature rises from 323 K (50°C) to 343 K (70°C) at about 90%

conversion. This chapter deals with the design of a PFR. The reactants are canola oil

and methanol. The catalyst employed is Sodium methylate (CHsONa), which is highly

soluble in methanol. The reactor products and co-products are biodiesel and glycerol.

Since the reaction chemistry involves three reversible reactions in series, trace amounts

of diglyceride and monoglycerides are also formed as byproducts. The esterification

reaction temperature is around 323 K (50°C) to 333 K (60°C). Depending upon the

pressure drop, reactor pressure ranges between 2 x 105 Pa (1.97 atms) and 4 x 105 Pa

(3.94 atms). Under these conditions, reactants and products are in liquid phase.

3.1 Plug Flow Reactor (PFR) Design

All non reacting molecules have equal residence time in a PFR. Any back mixing is

considered incidental. Natural turbulence provides mixing of reacting material. As a

20

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result of chemical reactions, concentration gradients are developed in the axial direction.

PFRs are mostly pipe reactors with pipe diameters ranging from 0.01 to 0.14 m. A single

long tube with several 180 degree return bends can be placed in a shell to make one

reactor unit where the shell side can be used for heating or cooling of the reactor as

shown in Figure 4. The selection of tube diameter depends upon providing a turbulent

flow and a compromise between construction cost, pumping cost and surface area for the

required heat transfer.

Reaction time is the most significant design parameter. Selection of a reaction time

depends upon the reactor volume and volumetric flow rate of reactants. Thus the ratio of

reactor volume to inlet volumetric flow rate defines 'apparent' residence time. The true

residence time (x) is found by integration of reactor volume to volumetric flow rate as

shown below:

r= rdK= r dn_

o J r * u (3.1)

In equation (3.1), ° is volumetric flow rate, r is rate of reaction and n is the local

molal flow rate of canola oil, the key component of the reacting mixture. Often a reactor

design is defined by a related concept of the term called 'space velocity'. Space velocity

in terms of LHSV is defined as the ratio of a flow rate at standard condition (288.7K

(15.7°C) and 1.013 x 105 Pa (latm)) to the volume of the reactor. Thus:

LHSV= (Initial liquid volumetric flow rate at 288.7K (15.7°C) per hour)/(reactor volume)

A PFR provides higher conversion without external mixing. However, conversion is

directly proportional to the length of the reactor. It should be noted that since conversion

21

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increases with length, so does the pressure drop. Thus higher conversion is associated

with higher pressure drop across the reactor.

3.2 Performance Equations for a Plug Flow Reactor (PFR)

The reaction chemistry for the conversion of canola oil into biodiesel involves three

successive liquid phase bimolecular second order reversible reactions in series as

described below:

Canola oil + Methanol-1=

PFR Biodiesel + Glycerol

Figure 4: Canola oil reacts with methanol in PFR to generate biodiesel and glycerol.

K, Triglyceride (TG) + Methanol (ME) ^^Diglyceride (DG) + Biodiesel (BD)

K2 (3.2)

K3 Diglyceride (DG) + Methanol (ME) ^^Monoglyceride (MG) + Biodiesel (BD)

K4 (3.3)

K5 Monoglyceride (MG) + Methanol (ME) Glycerol (GY) +Biodiesel (BD)

K6 (3.4)

The general material balance for component j in a PFR can be defined as:

22

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r

<

Rate of Accumulation of component j^-

r

v.

<

Rate of inflow of component j > <

v.

r ^ Rate of

outflow of component j

r Rate of Generation of

V J component j y by chemical reaction

The rate of appearance or disappearance of species 'j ' can be expressed in terms of its

molar flow rates Fj (mol/s) and its molar concentration Cj (mol/m3) as follows

(Fogler, 2006):

dF, r, = — -J dV

FJ=CJo„

d(CjU0) r , = 7 dV

(3.5)

(3.6)

(3.7)

The second order reaction rate laws rj for each species of the reactions 3.2 to 3.4 are

related to the rate constant kj and molar concentration Cj as follows:

r - - k C C + k C C ' TG 1 ^ TG ME T 2 BD ^ DG

r - k C C - k C C - k C C + k r c ' DG 1 ^ TG ^ ME 2 DG BD 3 DG ME T "" 4 MG ^ BD

' MC, k C C - k C C + k C C + k C C 3 DG ME 4 MG BD T ^ 5 ^ MG ME T 6 ^ GY ^ BD

(3.8)

(3.9)

(3.10)

_ r — _]r C C + k C C — k C C 4- k C C — k C C 'ME ^ I v TG ME T DG BD 113 ^ DG ME

T "" 4 MG ^ BD 5 MG ME

+ k C C t n, 6 Gr BD

(3.11)

rBD k\CTGCME k2 C DG C BD + k3C OG C ME kA C MG C BD + k5CMGCMr;

- k C C 6 cy BD

(3.12)

23

- rGY - k5CMGCm k6CGYCBD (3-13)

Substituting equations 3.8 to 3.13 in equation 3.7 yields six first order differential

equations given below:

d{CTGv())_ + t r r - o G14) aK

d(CDGoo) , r r _k r f _ir r C + k C C = 0 (3 15) dV

d{CMGv0) , r c -k C C +k C C +k C C = 0 f3 16) + K-J^JXJ^ME 4 MG BD 5 MG ME 6 GY BD u V-""1^ dv

- kxCTGCm + k2CDGCHD - k3CDGCMI, + kACMGCBO - k5CMGCm + k6CGYCBD = 0 d(CMUun)

dV (3.17)

BD^O ) + kxCTGC MB — k2C DGC BD + k2C DGC MIS k4CMGCBO + k5CMGCME k6CGYCBD 0 dV

(3.18)

+ k5CMGCMB k6CGYCBD - 0 d{CGYo„) dV J MU u (3.19)

The solution of six ordinary differential equations 3.14 to 3.19 with initial condition

for reactants provides the concentration of all the species as a function of length.

For liquid phase reactions concentrations change can be related to reactor length (L) as

follows:

From the reaction 3.19

ArL = v0d (3.20)

Differentiating on both sides

AdL = vndO (3.21)

24

dO = ^-dL (3.22) »o

dt = dL (3.23)

dC, dC, dt A^dL

Where:

j - TG, MG, DG, BD, ME and GY.

Cj- Molar Concentration (mol/m3)

Ar- Area of reactor (m )

no - Intial velocity (m2/s)

L - Length of reactor (m)

k - Rate constant (m /mol.min)

ry - Rate of reaction (min~)

F - Flow rate (mol/s) -2

V - Volume of reactor (m )

0 - Residence time (s)

3.3 MathCAD Solution to PFR Model

A MathCAD program to solve these differential equations is presented in Appendix A.

For a commercial plant capacity of 10 million gallons per year of biodiesel plant, the

reactor molar flow rates required are 6.575 kmol/hr of canola oil and 19.725 kmol/hr of

MeOH (1 mole oil to 3 moles of methanol). The amount of catalyst added is 116.3 kg

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(0.02 wt% of canola oil). Using the rate constant values listed in Table 3, the results of

the length versus conversion for different methanol to oil molar feed ratio for 0.038 m

(1.5 inch OD) and 0.051 m (2.0 inch OD) tubes are presented in the Tables 9 and 10

respectively. Graphical representations of these results are shown in Figures 5 through

10.

Table 9: MathCAD results for diameter 0.038m (1.5 inch OD) tube.

Length of reactor MeOH/Oil- 3 MeOH/Oil- 6 MeOH/Oil- 9 MeOH/Oil- 12

Conversion

(m) (%)

400 21.63 38.02 50.75 60.80

800 35.695 56.95 71.02 80.72

1200 44.57 66.83 80.51 89.05

1600 50.17 72.54 85.77 93.23

2000 53.74 76.21 89.06 95.54

2400 56.06 78.78 91.26 96.87

2580 56.84 79.71 92.02 97.27

26

Table 10: MathCAD results for diameter 0.051m (2.0 inch OD) tube.

Length of reactor MeOH/Oil- 3 MeOH/Oil- 6 MeOH/Oil- 9 MeOH/Oil-12

Conversion

(m) (%)

200 20.80 36.77 49.29 59.26

400 34.65 55.70 69.78 79.57

600 43.59 65.80 79.54 88.24

800 49.34 71.70 85.01 92.65

1000 53.06 75.50 88.44 95.12

1140 54.89 77.45 90.14 96.21

1200 55.52 78.17 90.75 96.57

27

Figure 5: Plug flow reactor length (m) versus conversion of canola oil to biodiesel for diameter 0.38m (1.5 inch OD) at 323K (50°C) and 4.053 x 10s Pa (4atms).

28

12

10

TGj 8

DQ

MGj 6

MEj

BDj 4

GY;

2

0

Figure 6: Length (m) of plug flow reactor versus component concentrations (kmol/m3) for diameter 0.038m (1.5 inch OD) at 323K (50°C) and 4.053 x 105Pa

(4atms).

29

Li

Figure 7: Plug flow reactor length (m) versus conversion of canola oil to biodiesel for diameter 0.38m (1.5 inch OD) at 323K (50°C) and 4.053 x 105 Pa (4atms).

30

Figure 8: Plug flow reactor length (m) versus conversion of canola oil to biodiesel for diameter 0.051m (2.0 inch OD) at 323K (50°C) and 4.053 x 10s Pa (4atms).

31

Figure 9: Length of reactor (m) versus component concentrations (kmol/m ) for diameter 0.051m (2.0 inch OD) at 323K (50°C) and 4.053 x 105 Pa (4atms).

32

Figure 10: Plug flow reactor length (m) versus conversion of canola oil to biodiesel for diameter 0.051m (2.0 inch OD) at 323K (50°C) and 4.053 x 105 Pa (4atms).

33

3.3.1 Effect of Methanol to Canola Oil Feed Ratio on Conversion

The reversible transesterification reaction requires high concentration of methanol for

the production of Biodiesel. Desired methanol to oil ratio is 3 but at this concentration

separation is not possible, hence methanol to oil ratio is increased to 6, 9 and 12 but at

higher concentrations separation of methanol from biodiesel is tedious. For the practical

purposes methanol concentration is maintained at optimum levels i.e 12. In the reactor,

the concentration of compounds changes with the length of the reactor as the

transesterification reaction progresses. Reactants (TG, ME, DG and MG) concentrations

decrease along the length of reactor. Products (BD and GY) concentration increase along

the length of reactor. Figures 6 and 9 are the concentration profile for methanol/oil feed

ratio-12. From the Figures 5 and 8, it is observed that as the molar ratio of methanol is

increased from 3 to 12, the conversion of canola oil changes from 54% to 97%.

3.4 Aspen Plus Simulation of PFR

The performance of a PFR was evaluated using, an AspenPlus simulation under the

following operating conditions:

Reactor feed: 298K (25°C) and 4.050 x 105Pa (4atms).

Reaction temperature as 323K (50°C) and pressure 4.050 x 105Pa (4atms)

Reactor Tube (inside diameter): 0.038m (1.5 inch) OD and 0.051m (2.0 inch) OD.

Reactor length: 2580m and 1140m.

Sensitivity Blockl: vary methanol to canola oil mole ration as 3, 6, 9, 10 and 12.

Sensitivity Block2: vary reactor inside tube diameter as 0.03 and 0.045 meters.

34

A schematic of the process under consideration is shown in Figure 11. The details of

Aspen input file are given in Appendix (B).

Figure 11: ASPEN simulation of PFR diameter-0.038m (1.5 inch) and 0.051m (2.0 inch) OD.

35

3.4.1 Plug Flow Reactor Size for Diameter 0.038m (1.5 inch OD) and 0.051m

(2.0 inch OD) Tubes

A summary of Aspen simulation results for reactor tube OD of 0.038m (1.5 inch) and

0.05 lm (2.0 inch) is given in Table 11:

Table 11: Summary of ASPEN results of PFR tube OD 0.038m (1.5 inch).

Lengt Ratio CONV T BD DG MG ME GY TG h OUT OUT OUT OUT OUT

(m) mins kg/hr kg/hr kg/hr kg/hr kg/hr kg/hr 2580 3 52.12 10.85 2053.94 920.48 132.82 409.15 144.38 2782.72 2580 6 72.87 10.32 3539.83 650.82 112.50 879.72 315.17 1576.66 2580 9 84.59 9.855 4483.67 403.25 84.18 1408.87 430.08 895.85 2580 12 90.83 9.343 5012.18 250.42 62.56 1982.94 496.12 532.89

Table 12: Summary of ASPEN results of PFR tube OD 0.051m (2.0 inch).

Length Ratio CONV T BD DG MG ME GY TG OUT OUT OUT OUT OUT


Graphical representations of results are shown in Figures 12 through 17.

36

Residence time vs Conversion

—MeOH/Oi l -3 MeOH/Oil-12

• MeOH/Oil-6 • Length of reactor

MeOH/Oil-9

s 'vi > s o U • mm o « "3 c « U

100

90 80

g > 0 f e o 0> 3 50

a 4 0

30 20 10

0

s-o u « e*

61) C -J

10 15 20

Residence time (minutes)

Figure 12: Plug flow reactor residence time versus conversion of oil to ester for diameter 0.038m (1.5 inch OD) at 323K (50°C) and 4.053 x 105Pa (4atms).

37

Length of reactor vs Component Flow Rates —•—BDOUT - " - D G O U T MGOUT -*~GYOUT MEOUT —•—TGOUT

O O &

cs pfj £ o

80.000

70.000

60.000

50.000

40.000

30.000

20.000

10.000

0.000 500

•O 0 O O o o o — o o o o o o o o

-X—*- -K X K X X X X X X *

1500 — • — * — • —

2500 Length (m)

— m —

3500 — » —

4500

Figure 13: Length of plug flow reactor versus component flow rate for diameter 0.038m(1.5 inch OD) at 323K (50°C) and 4.053 x 105 Pa (4atms).

38

Figure 14 : Plug flow reactor residence time versus conversion of oil to ester for diameter 0.38m (1.5 inch OD), 323K (50°C) and 4.053 x 105 Pa (4atms).

39

Residence time vs Conversion • MeOH/Oil-3 MeOH/Oil-12

MeOH/Oil-6 Length of reactor

MeOH/Oil-9

10 20 30 40 Residence time (minutes)

m

5000 4500 4000 3500 3000 2500 2000 1500 1000

500

9 -w w «

o xi % c V J

50

Figure 15: Plug flow reactor residence time versus conversion of canola oil to biodiesel for diameter 0.051m (2.0 inch OD) at 323K (50°C) and 4.053 x 105 Pa

(4atms).

40

Length of reactor vs Component Flow Rates —•—BDOUT — D G O U T MGOUT —*— GYOUT -e -MEOUT - • - T G O U T

80.000

70.000

Ja 60.000

50.000

,40.000

o o s a

30.000

20.000

0> -M «

£ E 10.000

0.000 ^ 500

n o—&- o o o o o o—o o o o o p

-X— -X X X X X X X X * »—«.... » — m—e~e

1500 2500 3500 4500 Length (m)

Figure 16: Length of plug flow reactor versus component flow rate for diameter 0.038m (2.0 inch OD) at 323K (50°C) and 4.053 x 105Pa (4atms).

41

Residence time vs Conversion

-•—MeOH/Oil-12 Length of reactor

SO ia > e o U .2 o

e « u

-a o s

100 90 80 70 60 50 40 30 20 10 0

T 5000

10

4500 4000 J 3500 © 3000 w «

0) 2500 0£

<M 2000 O

JS 1500 W) G 1000 Ol

J 500

20 30 40 Residence time (minutes)

50

Figure 17: Plug flow reactor residence time versus conversion of canola oil to biodiesel for diameter 0.051m (2.0 inch OD), at 323K (50°C) and 4.053 x 10s Pa

(4atms).

42

3.4.2 Effect of Methanol to Canola Oil Molar Feed Ratio on Conversion

Figures 13 and 16 are the flow rate profiles of TG, DG, MG, ME, BD, GY varying

along the length of the reactor for the two reactors with dimensions of 0.038 m OD, 2580

m long and 0.051 m OD, 1140 m long. In reactor flow rates of compounds vary along the

length of reactor as the transesterification reactions progresses with time. As expected

reactants (TG, ME, DG and MG) flow rates decrease along the length of reactor, whereas

products (BD and GY) flow rates increase along the length of reactor. Residence time

varies with conversion of oil to biodiesel for different methanol to oil molar feed ratios

are shown in Figures 12 and 15. From Figures 12 and 15 it can be concluded that for a

methanol to oil feed ratio of 12, 90.83% overall conversion is achieved under the

residence time of 9.3 minutes.

3.4.3 Plug Flow Reactor Length Required for 1.5 inch OD and 2.0 inch OD Tubes

Two different AspenPlus simulations were carried out. In the first case PFR input

parameters were: 0.38 m (1.5 inch) OD tube diameter, 2580 meters length and methanol

to canola oil molar feed ratio ofl2. In the second case input parameters are: 0.51 m (2.0

inch) OD tube diameter and 1140 meters long and methanol to canola oil feed ratio of 12.

Under these conditions, an overall conversion of canola oil was 90.88% in both the cases.

Then calculations for shell diameter:

N t u h e s = ^ - (3.31)

L = 4m

N t u b e S = 645 for 0.038m (1.5 inch) OD tube

N t u b e s = 285 for 0.051m (2.0 inch) OD tube

43

Dshe„ = 1.25 * Tubepitch * ^Nluhes

DSheii =3.225 m (10.6ft) for 0.038 m (1.5 inch) OD tube

Dsheii =2.144 m (7.0ft) for 0.051 m (2.0 inch) OD tube

Detailed calculations can be found in Appendix C.

3.4.4 Pressure Drop Across Plug Flow Reactor

For a 90.8% conversion, reactor length was found to be 2580 m for a 0.038 m (1.5

inch OD) tube and at the same value of conversion; a 2.0 inch OD reactor should be 1140

meters long. Assuming a shell-tube configuration, for a shell length of 4-meters, an S-

type tube length configuration would require Nbends of 180 degree tube bends as shown

below:

Figure 18: 180 degree tube bends of a PFR.

Pressure drop can be calculated from the following equations (Peters and Timmerhaus,

2003):

44

Pressure drop across straight length tube

2 a jjj r j

AP = — - ^SL.) (3.33) 2 D

Pressure drop across 180 degree bend

2

AP = p*^-*(Kf) (3.34)

Total pressure drop per tube length with one 180 degree bend 2 A jjs /* jJ> J

AP = — J - H2«2!- + K f ) (3.35) 2 D

This pressure drop is for one 4m long one-bend tube

NB e n d s=N l u b e s- l (3.36)

APlnlal=AP*Nliemh (3.37)

From the equation 3.37, NreactorCan be assumed and the total reactor pressure drop will

be:

^ a c l nr=^P l„ l a l*N r e a c , o r (3.38)

Where:

Lreactor ~ Total length of PFR (m)

p - Density (kg/m )

D - Diameter of tube (m)

/ - Friction factor

ve - Velocity in a tube (m/s)

K f - Frictional loss in a 180 degree bend

AP - Pressure drop (Pa)

45

Neends - Total number of 180 degree bends in a PFR

Ntubes - Total number of tubes in a PFR

Nreactor- Total number of reactors required

A P r e a c t o r - Pressure drop in a PFR (Pa)

A P t o t a i - Pressure drop in all tube of a PFR (Pa)

Detailed pressure drop calculations are summarized in Appendix D. For case 1 (0.038

m (1.5 inch OD) and L = 2580 m) the total pressure drop is found to be 3.55 x 106 Pa (35

atms). In case 2 (0.051 m (2.0 inch OD) and L= 1140), the total pressure reduces to 1.32

x 106 Pa (13 atms). It should be noted that the pressure drop can be further reduced by

increasing the tube diameter. However, tube diameters larger than 0.051 meters (2

inches) may not provide the uniform concentration in axial and radial direction of the

reactor.

The cost for the year 2008 of the designed PFR is estimated from the following

correlation (Seider et al., 2004):

3.4.5 Plug Flow Reactor Cost

Cost = Exp(11.0545 - 0.9228 * In (Area) + 0.0979(ln (Area)2)) * Fh M (3.39)

(3.40) 4

(3.41)

The results are shown in Table 13

46

Table 13: Cost of Plug flow reactor for different diameters.

Case Tube OD Reactor Gear Pump Reactor Total Length cost cost cost

1 0.038 m (1.5 inch) 2580 meters $22000 $97,000 $166,700

2 0.051 m (2.0 inch) 1140 meters $9,400 $169,000 $250,000

As the diameter is increased from 0.038 m (1.5 inch) OD to 0.051 m (2.0 inch) OD the

cost of a PFR is increased by 66.7%. Details of plug flow reactor cost calculations are

summarized in Appendix E.

47

CHAPTER 4: CONTINUOUS STIRRED TANK REACTOR (CSTR)

The continuous stirred tank reactor (CSTR) is also known as an ideal reactor or an

agitated tank reactor. This reactor can be used for liquids and slurries as reactants.

4.1 Performance Equations Model for a Continuous Stirred Tank Reactor (CSTR)

A general material balance relationship for liquid phase chemical reactions a CSTR is

presented in Figure 19:

Canola oil + Methanol- CSTR -•Biodiesel + Glycerol

Figure 19: Canola oil reacts with methanol in a CSTR to produce biodiesel and glycerol.

r

< Rate of Accumulation of component j/"

r

v.

Rate of inflow of component j

v.

> <

Rate of outflow of component j r

r Rate of

+ J Generation of a Component j

by chemical reaction

v.

>

48

The rate of appearance or disappearance of species 'j' can be expressed in terms of its

molar flow rates Fj (mol/s), molar concentration Cj (mol/m3) as follows (Fogler, 2006):

F (4-1)

F]=CJv0 (4.2)

(C -C,)*o0 j y v /

V 9 = — (4.4)

The second order reaction rate laws for each specie can be related to the rate

constant kj and the molar concentration C7 are described below.

Substituting equations 3.8 to 3.13 for r, in equation 4.3 yield:

Cm -CTGa + {-kxC1GC ME + k2CBDCDG) *0 = 0 (4.5)

Cdg ~ Cdg„ + (kiCTGCME — k2CDGCBD — k3C/x;CME + kACMGCBD) * 6 = 0 (4.6)

Cmg ~ CMg„ + (k3CDGCME — k4CMGCljn + k5CMGCME + k6CGrC BD) * (9 = 0 (4.7)

Cmi^m/:, k]CTGCME + k2CDGCBD k3CDGCME + k4CMGCBD k5CMGCME

+ k6CGyCBD)*0 = 0

Cbd ~ c b d 0 +(kiCTGCME ~k2CDGCME + k2CDGCME — k4CMGCBD + k5CMGCME

- k6CGyCBD)* 0 = 0

(4.8)

(4.9)

Q,y _ + (k5CMG CME - k6CGY CBD )* 0 = 0 (4.10)

A solution of these six equations (Equations 4.5-4.10) with initial condition for

reactants provides the concentration of each compound as a function of time (0).

49

Where:

j - TG, MG, DG, BD, ME and GY.

Cj- Molar Conentration (mol/m3)

Ar- Area of reactor (m )

Do - Intial velocity (m /s)

L - Length of reactor (m)

k - Rate constant (m /mol.min)

Xj - Rate of reaction (min1)

F - Flow rate (mol/s)

V - Volume of reactor (m )

0 - Residence time (s)

4.2 MathCAD Solution to CSTR Model

A MathCAD program to solve these differential equations 4.5 to 4.10 is presented in

Appendix F. The reactor feed conditions and flow rates are same as described earlier for

the design of PFR.

50

Table 14: Summary of MathCAD results for different reactor volume and methanol to oil molar feed ratio.

Reactor Methanol to Conversion Residence Biodiesel Volume Oil Ratio time Production rate

(mJ) (%) (minutes) (kg/hr) 6 12 88.08 34.01 4888.21 7 11 88.46 40.93 4910.84 7 12 89.48 40.00 4991.20 8 10 87.8 47.90 5038.54 8 11 89.59 47.08 4992.37 8 12 90.56 46.00 5071.00 9 10 89.39 55.01 4969.66 9 11 90.42 53.00 5050.63 9 12 91.41 52.02 5131.80 10 9 88.34 62.03 4896.52 10 10 90.07 61.01 5021.54 10 11 91.21 60.03 5194.21 10 12 92.11 58.03 5187.00

4.3 Aspen Plus Simulation of CSTR Model

The performance of a CSTR was also evaluated using an AspenPlus simulation under

the following operating conditions:

The details of Aspen input file are given in Appendix (G).

Reactor feed: 298K (25°C) and 4.053 x 105 Pa (4atms).

Reaction temperature: 323K (50°C)

Pressure: 4.053 x 105 Pa (4atms)

Reactor volume: 7m3.

Sensitivity Blockl: vary methanol to canola oil mole ration as 3, 6, 9 and 12.

Sensitivity Block2: vary reactor volume 1 to 10 m .

51

The details of Aspen input file are given in Appendix (G). Figure 20 shows a

schematic diagram of the process under consideration. A summary of Aspen simulation

results for a CSTR volume of 7m3 is given in the Table 15.

Table 15: Summary of ASPEN results of CSTR volume 7m3.

Lengt Ratio CONV X BD DG MG ME GY TG h Product OUT OUT OUT OUT OUT


52

Graphical representations of results are shown in Figures 21 to 23 and analysis of

these results is presented below in 4.3.1 and 4.3.2.

Details of all simulation results are given in Appendix G.

Figure 21: Continuous stirred tank reactor residence time versus conversion of canola oil to biodiesel for methanol to oil ratio 3, 6, 9,12 varying along the volume of

reactor at 323K (50°C) and 4.053 x 105Pa (4tams).

53

Volume vs Component Flow rates

BDOUT -®~TGOUT DGOUT -m-MGOUT -©-AOUT —•-GYOUT

Volume (m3)

Figure 22: Volume of Continuous stirred tank reactor versus component flow rate for methanol to oil ratio 12 at 323K (50°C) and 4.053 x 105Pa (4tams).

54

Figure 23: Continuous stirred tank reactor residence time versus conversion of canola oil to biodiesel for methanol to oil ratio 12 varying along the volume of

reactor at 323K (50°C) and 4.053 x 10sPa (4tams).

4.3.1 Effect of Residence Time on Conversion

These results show that if the reactions are carried out in a single reactor, conversion

depends upon the methanol to canola oil molar feed ratio under the same residence time

in the range of 40 to 45 minutes. As it can be seen from Figure 21, at a given reactant

feed ratio, increasing residence time beyond 40 minutes, there is no significant increase

in the overall conversion. Over 90% conversion can be achieved under 40 minutes of

residence time, if the methanol to oil feed ratio is increased to 12.

55

4.3.2 Effect of Reactor Volume on Conversion

At reactant feed ratio of 12, a single reactor with 7 m3 volume would be required for

90% conversion. Increasing the reactor volume from 4 to 10 m3, there is no significant

increase the production of bio diesel as shown in Figure 22. The reactor performance

with respect to residence time, volume and conversion is shown in Figure 23.

4.4 Design Dimensions of the CSTR at Methanol/Oil Molar Feed Ratio of 12

Input= Fresh feed

Volumetric flow rate = (Fresh feed)/ (pavg)

Assuming

Design Residence time =1.5* Reactor residence time

Design Volume of reactor = Volumetric flow rate* Residence time

Area of reactor = 3.14* (diameter)2 / 4

Height of reactor = Design Volume of reactor / Area

Actual feed volume = Volumetric flow rate*Reactor residence time

Reactor Design Volume = Area of reactor * Height of reactor

Percent over design = Reactor Design Volume/ Actual feed volume

Percent over reactor volume (based on rules of thumb). Calculations of CSTR sizing

can be found in Appendix design for CSTR of volume 7 m3 is 1.5. The author

recommends a 50% over design in H.

56

4.5 Continuous Stirred Tank Reactor Cost

Using M and S index (Chemical Engineering Magazine, 2009) and Figure H, cost of

single CSTR with material of construction as stainless steel is $77,880.

Details of utility cost of CSTR can be found in Appendix I. This represents over

324% lower reactor cost when compared with the PFR cost of $250,000.

4.6 Continuous Stirred Tank Reactor in Series

In a CSTR reactor design, it is important that reactor contents are thoroughly mixed

since perfect mixing is the fundamental basis of reactor performance for desired

conversion. Therefore reactor residence time should be greater than the mechanical

mixing time. Typically, the reactor residence time should be at least 5 times the

mechanical mixing time. It is economically beneficial to operate several CSTR units in

series. The multiple reactors in series require a smaller reactor volume and offer higher

overall conversion. In the production of biodiesel, reactants, catalyst and products are all

in liquid phase, a simple pump around configuration ensures vigorous mixing. This

approach for mixing offers better economics since pumping energy required is much less

than the energy requirements for impeller type mechanical mixing.

4.6.1 A System of Three CSTR Biodiesel Reactors in Series

Figure 24 shows an Aspen simulation diagram of three reactors in series. The material

balance results are presented in Table 16. Each reactor volume is 3 m3. The feed to the

first reactor consists of canola oil, methanol at 323 K (50 °C) and 4 x 105 Pa (4atms) (abs)

pressure. The effluent stream from each reactor is subjected to a pump around system

57

where 90 % of the reactor effluent stream is recycled back to the reactor. The recycle

stream is cooled to 313 K (40 °C) before entering the first reactor. As a result of inter

cooling; feed to the second reactor is at 322 K (49 °C). Under adiabatic conditions, the

temperature in the second reactor increases from 322 K to 343 K (49 °C to 70 °C). In the

third reactor, temperature rises from 343 K to 382 K (70 °C to 79 °C). It should be

pointed out that the reaction rate constants used in this study were determined

experimentally at 333 K (60 °C). If these rate constants are assumed to be valid up to

353 K (80 °C), then no inter cooling for the second and the third reactor will is required.

The Aspen simulation of methanol recovery from reactor effluent stream using direct

distillation reveals that sodium methoxide catalyst will boil-off with methanol in the

overhead stream to be recycled. This is mainly due to lack of liquid phase dissociation of

methanol- sodium methylate system. Thus reactor effluent may be cooled and separated

into two liquid phases. The recovery of various compounds for each liquid phase can be

accomplished by approaches similar to those found in literature (Apostolakou et al., 2009

and Haas et al., 2006).

58

Figure 24: ASPEN simulation of three CSTR's in series.

59

Table 16: Material balance data for three CSTR's in series to produce biodiesel from canola oil (Plant capacity: 10 million gallons/year of biodiesel product).

Stream Name TG Methanol Rl-IN Rl-OUT Rl-RCY R2-IN R2-OUT

Phase Liquid Liquid Liquid Liquid Liquid Liquid Liquid Mass Flow

(kg/hr) TG 6144 0 6144 10115 9103 1012 1974 BD 0 0 0 48396 43556 4840 58538

Methanol 0 3336 3336 28126 25313 2813 27030 GY 0 0 0 4757 4282 476 5965 DG 0 0 0 2653 2388 265 976 MG 0 0 0 734 661 73 305

Mass fraction

TG 1 0 0.648 0.107 0.107 0.107 0.021 BD 0 0 0 0.511 0.511 0.511 0.681

Methanol 0 1 0.352 0.297 0.297 0.297 0.285 GY 0 0 0 0.05 0.05 0.05 0.063 DG 0 0 0 0.028 0.028 0.028 0.01 MG 0 0 0 0.008 0.008 0.008 0.003

Total Flow (kmol/hr)

7 104.3 111.2 1111.9 1000.7 111.2 1111.9

Total Flow (kg/hr)

6144 3336 9480 9478.2 85302 9478 94789

Temperature (°C)

15 15 50 50 50 50 50

Pressure (bar) 1 1 4 1 5 5 1

Density (kg/cum)

909.4 803.6 588 778.8 778.7 778.7 826.4

Average (MW)

884 32 85.3 85.2 85.2 85.2 85.2

60

Table 16: continued

Stream Name R3-IN R3-OUT RAW1 M-RCY RAW2 RAW3 GY

Phase Liquid Liquid Liquid Liquid Liquid Liquid Liquid Mass Flow

(kg/hr) TG 197 611 61 0 61 61 0 BD 5854 60418 6042 0 6042 6042 0

Methanol 2703 26824 2682 2682 1 0 1 GY 597 6201 620 0 620 1 620 DG 98 529 53 0 53 53 0 MG 31 198 20 0 20 20 0

Mass fraction

TG 0.021 0.006 0.006 0 0.009 0.01 0 BD 0.618 0.637 0.637 0 0.889 0.978 0

Methanol 0.285 0.283 0.283 1 0 0 0.001 GY 0.063 0.065 0.065 0 0.091 0 0.999 DG 0.01 0.006 0.006 0 0.008 0.009 0 MG 0.003 0.002 0.002 0 0.003 0.003 0

Total Flow (kmol/hr)

111.2 1111.9 111.2 83.8 27.4 20.6 6.8

Total Flow (kg/hr)

9479 94782 9478 2682 6797 6176 620

Temperature

(°C)

50 50 50 67 224 25 25

Pressure (bar) 5 1 5 1.1 1.1 2 2 Density

(kg/cum) 826.2 835.2 835 740.3 723.6 863.6 1271

Average (MW)

85.2 85.2 85.2 32 248.2 299.3 91.9

61

Table 17: Below summarizes the results of overall conversion as a function of methanol to canola oil feed ratio.

Methanol to Canola Oil mole ratio

%Conversion CSTR -1

%Conversion CSTR-2

%Conversion CSTR-3

%Conversion Over all

3 48.1 21.5 7.8 62.4 6 67.4 45.6 24.7 86.7 9 77.1 65.2 44.1 95.5 12 80.8 75.2 59.6 98.1 15 83.5 80.5 69.1 99.0

4.6.2 Economic Analysis of Three CSTR in Series

The utility requirements for heat exchangers, methanol recycle distillation column and

transfer pumps are summarized in Table 18. The size (diameter and height) and cost of

methanol recycle distillation column at various methanol to canola oil ratios is given in

Table 18. The utility prices used in Table 18 are: $0.59/1000 kg steam; $0,045/1000 kg

cooling water; $0.085/kWh electricity.

Table 18: Total utility cost ($/hr) as a function of methanol to canola oil mole feed ratio.

Methanol/Canola Oil Ratio Ratio Ratio Ratio Ratio 3 6 9 12 15

Distillation cost($/hr) cost($/hr) cost($/hr) cost($/sec) cost($/sec) Reboiler steam cost $2.16 $5.40 $9.72 $14.40 $19.08

Condenser water cost $6.12 $10.08 $16.20 $23.04 $29.88

Power cost for Pump -1 $1.44 $1.44 $1.44 $1.44 $1.80 Power cost for Pump -1 $1.44 $1.44 $1.44 $1.44 $1.80 Power cost for Pump -1 $1.44 $1.44 $1.44 $1.44 $1.80

Total Utility Cost ($/hr) $12.60 $19.80 $30.30 $41.90 $53.80 Overall Conversion 62.40% 86.70% 95.50% 98.10% 99.00%

62

For each value of methanol to canola oil mole feed ratio, the amortized cost of

methanol distillation column is shown in Table 19

Table 19: Cost of methanol recovery distillation column.

Methanol/Canola 3 6 9 12 15 Oil ratio

Column diameter 0.5 0.7 0.8 1.2 1.4 (m)*

Methanol Column $252,000 $331,200 $388,800 $504,000 $684,000 Cost**

Amortized Cost $32,508.0 $42,724.8 $50,155.2 $65,016.0 $88,236 ($/yr)***

Column Cost $/hr $4.5 $5.9 $7.0 $9.0 $12.3 * Column height 11 meters, **Cost includes installation and auxiliaries *** Amortization factor (A/P, 10, 5) = 0.129

In order to carryout an economic analysis of incremental change in methanol feed

ratio to incremental increase in canola oil conversion, the equipment cost of reactors,

pumps and heat exchanger will not have any impact on the outcome of results since these

costs will be the same for all cases. Therefore the total cost can be expressed as the sum

of the Methanol distillation column cost (Fixed cost) plus the utility cost (variable cost).

The results of total cost form the data given in Table 18 and 19 are presented in Figure 25

as shown below:

63

$86.3

Cost ($/hr) v's Biodiesel Product Purity as a function of Methanol to Oil mole Ratio

Utility -Conversion —A— eqpcost Total-Cost

$69.0

t" $51.8

U) o $34.5

$17.3

$0.0

$17.1 $T2.6

• $19.8

- Jr$30.3 • $37.2

* 62.4%, -A-5t9

• 86.7%r nA-7.0 • 95.5%r • 98.1%, 9 12

Methanol to Oil mole Ratio

12.3

99.0% 15

Figure 25: Methanol to oil mole ratio versus total cost.

64

4.6.3 Design of Methanol Recovery Column

Methanol recovery column was designed using following rigorous correlations (Wankat, 2007):

Figure 26: Schematic diagram of distillation column.

4.6.3.1 Column Diameter: Fair Flow parameter (Csb) is used determine column

diameter at top and bottom column conditions. The calculated results are

Djop = 1.361 m

DBottom= 1.614 m

Use 1.614 m diameter column

4.6.3.2 Tray Hydraulics: To carry out tray hydraulic calculations, the following tray

layout was assumed:

Tray Thickness = T t r a y = 0.0019 m (0.078 inch)

Hole diameter = 0.0047 m (3/16 inch)

Column Area = AT

65

Down comer Area = Ad = 12% of AT

Net Area = An = 88% of AT (Vapor Flow Area)

Active Area = Aa = (AT - 2Ad)

Hole Area = Ah = 10% AT

Weir Length = lw = 77% of Diameter (Top or Bottom)

Using the above tray specifications, the following results are obtained:

i) Bottom column area = 2.055 m ,

ii) Net tray area = 1.808 m2,

iii) Down comer flow area = 0.247 m2,

iv) Active tray area = 1.561 m2,

v) Tray hole area = 0.156 m2,

vi)Weir Length = 1.243 m

Pressure drop across the tray consists of pressure drop caused by the various

hydrodynamic heads which include dry pressure drop (hdC) plus static head by the weir

( h w e i r ) plus head caused by the liquid crest (hcrest) plus static head of the liquid under the

down comer (hdU).

hdc W Dry + Kveir+ ^crest + hdu

Calculation results of each of these terms are presented below. Detailed calculations are

presented in Appendix K.

4.6.3.3 Dry Pressure Drop: It should be noted that all pressure drop results in inches (m)

of water.

66

Calculate orifice coefficient (C0) for a tray hole using the following Kesseler and

Wankat correlation:

C 0 := 0 . 8 5 0 3 2 - 0 . 0 4 2 3 1 -dhole N

V tray ) 0.0017954-

r dhole V

v W ) C0 = 0.759

hAP_Dry := 0.003-' f t V

I V s -

P VBottom jb

ft3

P water

v PLBottom .

W ) " •in

^APDry = 0.025 m

hcrest := 0.092-F wear

C r A "13 Ijrr gal

V min J w Bottom

•in

Merest = 0.022 m

Assume negligible gradient pressure drop,

hgrad : = 0

hdu := 0.56-

gal V min

449 Adu

ft2

•in

- 3 hdu = 7.641 x 10 m

67

hweir = 0.051 m

hdc : = hAP Dry + hweir + hcrest + Hju

hdc = 0.106m

hdc •\ic_aerated

<Pdc

hdc_aerated = 0.211m

This value is less than 0.61 m (24 inch) tray spacing. There should be no problem for

vapor to flow to the above tray.

Where:

Lg = Liquid Flow Rate (L/min)

Fweir = Factor accounting for a curvature of the column wall in the down comer

hgrad= The liquid gradient across the tray (m)

hdc= Total Head of clear liquid in the down comer (m)

hdc_areated = Head of clear liquid in the down comer due to aeration (m)

hdu = Frictional loss due to flow in the down comer and under the down comer onto the

tray (m)

hweir — static head by the weir (m)

hcrest= static head by the liquid crest (m)

(3 = Area of hole to active area (unit less)

Adu = Flow area under the down comer (m2)

68

4.6.3.4 Check for Down Comer Residence Time: According to the rule of thumbs, a

minimum down comer residence time should be at least 3 seconds. Based on foaming

tendency of the liquid, this residence time ranges from 3 seconds (low foaming) to 7

seconds (very high foaming) liquids (Kister, 1992).

(Ad Bottom "PL Bottom •hdc)

^Bottom + e e tdc : =

t<ic = 3.695 s

Where

ee = entrainment

tdc= Down comer residence time

This value is greater than the minimum residence time of 3 seconds, there should be

no problem to maintain the proper liquid flow.

4.6.3.5 Weeping Check: Weeping relates to the direct flow of the liquid through sieves

without maintaining a proper vapor liquid contact. Excessive weeping is estimated using

Fair's (Wankat, 2007) correlation for a surface tension head(ha). f

ho := 0.04in-

dyne

V cm J

P L B o t t o m

lb

\ dhole

V in ;

- 3 ha = 1.299 X 10 m

69

Excessive weeping is defined by equality that the term (h^p £)ry+h a) should be

greater than the value of the term (.10392+0.25119x-0.021675x 2)-

h A P Qry+h a > 0.10392+0.25119x-0.021675x 2

where x is deined as: h Weir+hcrest+hgrad t* ien

h A P _ D r y + h a = i - 0 4 7

0.10392+0.25119x-0.021675x 2 = 0.645

These calculations show that the inequality is satisfied. Therefore, Weeping should

not be a problem. All of the design calculations are presented in Appendix K.

70

CHAPTER 5: CONCLUSIONS AND RECOMMENDATIONS FOR FUTURE RESEARCH

5.1 Conclusions

In this study, reactions of canola oil with methanol to produce biodiesel were

examined in a PFR as well as in a CSTR reactors. Based on kinetic rate expression

available from the literature, a Plug Flow Reactor (PFR) was designed to determine the

reactor length as a function of conversion of canola oil. Two different reactor tube

diameters were investigated in the plug flow reactor system.

For a Plug Flow Reactor with 0.038m (1.5 inch) OD tube, nearly 91% conversion was

achieved in a 2580m long reactor at methanol to oil molar feed ratio of 12 ( as compared

to the stoichiometric ratio of 3) with 9.3 minutes of reactor residence. The total pressure

drop for this reactor was found to be around 35.463 x 105 Pa(35atms). The requirement

of a long tubular reactor is mainly due to the desired conversion of canola oil beyond

90%.Gear pumps must be used to provide the pressure of the feed system to overcome

the prevailing pressure drop.

According to the information on gear pumps from reference (Seider et al., 2004),

"Although gear pumps can be designed to operate over a wide range of flow rates and

discharge pressure, typical ranges are 10 - 1500 gpm, and up to 200 psi for high viscosity

fluids.", for a 0.038 m (1.5 inch) OD diameter PFR, the pressure drop requirements can

71

not be met with a single off-the-shelf gear pump. If two plug flow reactors of 0.038 m

(1.5 inch) OD tube are used, the developed pressure drop of 17.97 x 105 Pa (17.732atms)

per reactor would still be higher and exceeds the pressure developed by the standard size

gear pump, hence more than one gear pumps will be needed. The total estimated cost for

this PFR (diameter of 0.038 m (1.5 inch) OD tube and 2580 m length) using two gear

pumps will be $166,700 (2008). For a PFR with 0.051m (2.0 inch) OD tube, the reactor

tube length decreases to 1140 meters under the same conditions as in for PFR of diameter

0.038 m (1.5 inch) OD tube. The resulting pressure of 13.17 x 105Pa (13atms) can be

attained with a single gear pump unit. The total estimated cost of 1140m long and 2.0

inch OD PFR will be $250,000 (2008). Therefore it is concluded to use a 0.051 m (2

inch) OD tube.

In the case of a CSTR, reactor volume was determined as a function of conversion.

Under each type of reactor design study, methanol to canola oil molar feed ratio was used

as an independent parameter. The reactant flow rates were based on producing 10

million gallons of biodiesel per year. For a 7m3 volume CSTR and methanol to oil molar

feed ratio of 12 yields 90.3% conversion with 41 minutes of reactor residence time.

The estimated cost of a 316-s.s CSTR with 50% capacity over design is found to be

$77,880. The capacity over design provides 30 minutes of hold up volume to

accommodate startup and process upsets.

The CSTR reactor design also included the evaluation of three CSTR reactors in series

with a pump around loop to provide vigorous mixing. The reactor effluents were

separated by distilling the un-reacted methanol for recycle. When the value of 'R'

(defined as mole feed ratio of methanol to canola oil) is between 3 (stoichiometric feed

72

ratio) and 15, an additional distillation column will be required for product purification in

order to meet ASTM- D6751 (biodiesel product standard) which requires > 99 wt% of

biodiesel product purity. The boiling point of biodiesel is approximately 598K (325 °C).

Biodiesel is thermally decomposed when temperature exceeds 523K (250 °C). The

thermal decomposition temperature for glycerol is about 423K (150 °C). In order to keep

low distillation temperature, purification column must be operated under vacuum. Based

on Aspen simulation, 10 milli-bar of vacuum would be required to operate the distillation

column below the decomposition temperatures of these products. If the value of 'R' is

>15, product will meet the biodiesel quality as set forth under the ASTM-D6751 standard

without additional purification. The purification cost via vacuum distillation would be

$40.40/hr. To increase product purity from 87 wt% to 99 wt%, the methanol recovery

and recycle cost increases by $40/hr which comes out to be nearly the same when

compared with vacuum distillation of biodiesel product. However, the vacuum

distillation process step will be eliminated.

Because of a very large difference in the boiling points of methanol (33 8K or 65°C)

and biodiesel (598K or 325°C), separation of excess methanol can be easily accomplished

by simple distillation with nearly 100% methanol recovery for recycle.

A detailed design of a sieve tray distillation column using rigorous correlations

suggests a single column of 1.6 meters diameter and a total height of 10 m (33 ft) would

be required. This column will have 12 stages with 24 inches of tray spacing. The feed

will enter at stage 6. The column is designed to operate at 75% flooding. The column

operation for weeping and entrainment was checked, using tray hydraulic correlations.

73

5.2 Recommendations for Future Research

The results of this thesis provide the necessary guidelines for the production of 10

million gallons per year of biodiesel from canola oil. The following recommendations

for future research are made:

A reactor consisting of a reactive distillation column should be investigated as it will

serve as a reactor as well as a separator for the recycle of methanol.

It is recommended that total cost of a 'Reactive Distillation' column should be compared

with the CSTR reactor system.

If commercial scale biodiesel plants are built, there will a glut of glycerol co-product

(1 kg of glycerol is produced for every 10 kg of biodiesel manufacture). It is

recommended that some new process chemistry to convert glycerol into more useful and

economically attractive products be explored. For example, pyrolysis of glycerol may

yield synthesis gas as envisioned by the following reaction (C3H8O3 ->3CO + 4H2).

74

REFERENCES

REFERENCES

Apostolakou, A.A., Kookos, I.K., Marazoiti, C., and Angelopoulos, K.C. (2009). Techno-economic analysis of a biodiesel production process from vegetable oils. Fuel Processing Technology, 90, 1023-1031.

Byrd, A.J., Pant, K.K., and Gupta, R.B. (2008). Hydrogen production from glycerol by reforming in supercritical water over Ru/A1203 catalyst. American Chemical Society, Fuel 87, 2956-2960.

BP Statistical Review of world energy 2008. Available from BP statistical review 2008 internet pages <http://www.investis.com/bp_acc_ia/stat_review_2008/htdocs /reports/reportl 9.html>.

Chavanne, C.G. (1938). Belgian Patent 422, 8 77, Aug. 31, 1937; Chem. Abs.52:4313.

Chemical Engineering Magazine. (2009). M & S index, Pg 64, September edition.

Fogler, H.S. (2006). Elements of Chemical Reaction Engineering. 4th ed. Prentice Hall, NJ.

Gabbard, A. (1993). Coal Combustion Nuclear Resource or Danger. Oakridge National Laboratory Review (ORNL), 26, 25-33.

Haas, M.J., McAloon, A.J., Yee, W.C. and Foglia, T.A. (2006). A process model to estimate biodiesel production costs. Biosource Technology, 97, 671-678.

Kister, H.Z. (1992). Distillation Design. McGraw Hill, NY.

Klepacova, K., Mravec, D., Kaszonyi, A. and Bajus, M. (2007). Etherification of glycerol and ethylene glycol by isobutylene. Applied catalysis, 328, 1-13.

Knothe, G. (2001). Historical Perspectives on Vegetable Oil- Based Diesel Fuel. Inform 12(11), 1103-1107, Retrieved 2009-06-24.

Komers, K., Skopal, F., Stloukal, R., and Machek, J. (2002). Kinetics and mechanism of the KOH - catalyzed methanolysis of rapeseed oil for biodiesel production. European Journal of Lipid Science and Technology, 104(11), 728-737

75

http://www.investis.com/bp_acc_ia/stat_review_2008/htdocs%e2%80%a8/reports/reportl%209.html

http://www.investis.com/bp_acc_ia/stat_review_2008/htdocs%e2%80%a8/reports/reportl%209.html

Leevijit, T., Wistmethangoon, W., Prateepchaikul, G., Tongurai, G., and Allen, M. (2004). A second order kinetics of palm oil transesterification. Joint International Conference on Sustainable Energy and Environment (SEE), 3, 277-281.

Marchetti, J.M., Miguel, V.U., and Errazu, A.F. (2007). Possible methods for biodiesel Production. Renewable and Sustainable Energy Reviews, 11, 1300-1311.

McCance, R. A., Widdowson, E.M., Holland, B., and Paul, A. A. (1991). The Composition Of Food. 5th ed. Ministry of Agriculture, Fisheries and Food., Royal Society of Chemistry, Cambridge.

National Biodiesel Board, cited in Pearl, G.G. (2001, August). Biodiesel Production in the U.S. Render Magazine.

Noureddini, H., and Zhu, D. (1997). Kinetics of Transesterification of Soybean oil. JAOCS, 74, 1457-1463.

Peters, M.S., and Timmerhaus, K.D. (2003). Plant Design and Economics for Chemical Engineers. 5th ed. McGraw-Hill Companies, Inc., NY.

Schuchard, U., Sercheli, R., and Vargas, R.M. (1998). Transesterification of Vegetable Oils: A Review. J. Braz. Chem. Soc., 9, 199-210.

Seider, W.D., Sieder, J.D., and Lewin, D.R. (2004). Product and Process Design Principles:Synthesis, Analysis and Evaluation. 2nd Ed. John Wiley and Sons, Inc.

Sharma, Y.C., Singh, B., and Upadhyay, S.N. (2008). Advancements in development and characterization of biodiesel: A review, fuel, 87, 2355-2373.

Wankat, P.C. (2007). Separation Process Engineerin. 2nd Ed. Prentice Hall, NJ.

General References

Krik - Othmer. (1991). Encyclopedia of Chemical Technology. 4th ed. Wiley Inter Science, Inc.

Turner, T.L. (2005). MS thesis. Modeling and Simulation of Reaction Kinetics for Biodiesel Production. North Carolina State University

76

Appendix A MathCAD Program for Solution of PFR for Diameter 0.038m (1.5

inch OD) and 0.051m (2.0 inch OD) Tubes

MathCAD program for solution of PFR for diameter 1.5 inch OD tube:

kl := 0.05 k2 := 0.11

k3 := 0.215 k4 := 1.228

k5 := 0.242 k6 := 0.007

A := 0.0009 v0:= 0.1582

kl,k2,k3,k4,k5 and k6 are rate constants in Lit/mol*min Units

—TG(t) = — (-kl TG(t) ME(t) + k2 DG(t) BD(t)) dL v0

— DG(t) = — (kl TG(t)ME(t) - k2 DG(t) BD(t) - k3 DG(t) ME(t) + k4 MG(t) BD(t)) dL v0

— MG(t) = — (k3DG(t)ME(t) - k4MG(t)BD(t) - k5MG(t)ME(t) + k6GY(t)BD(t)) dL v0

— ME(t) = — / - k l TG(t) ME(t) + k2DG(t)BD(t) - k3 DG(t) ME(t) ... dL v0 ^+k4-MG(t)-BD(t) - k5 MG(t)-ME(t) + k6 GY(t) BD(t)

— BD(t) = — / k l TG(t) ME(t) - k2 DG(t) BD(t) + k3 DG(t) ME(t) - k4 MG(t) BD(t) dL vo ^+k5 MG(t)-ME(t)-k6 GY(t)-BD(t) j

— GY(t) = — (k5 MG(t) ME(t) - k6 GY(t) BD(t)) dL vc

77

Here:

A- Flow Area of tube 1.5 inch OD tube (m ) v0- volumetric flow rate of the feed (m /s)

TG=Q0, DG=Q1, MG=Q2, ME=Q3, BD=Q4 and GY=Q5

D(L,Q) :=

- • ( - k l Q 0 - Q 3 + k2-Q1-Q4) v0

— <kl Q0 Q3 - k2-Qi -Q4 - Id-Q, Q3 + k4 Q2-Q4) v0

— (k3Q 1 Q3 - k4 Q2 Q4 - k 5 Q 2 Q 3 + k 6 Q 5 Q 4 )

- • ( - k l -Qo-Qs + k2-Qi Q4 - k3-Q1 Q3 v, Uk4-Q 2 Q 4 -k5-Q 2 -Q 3 + k6-Q5 Q4 J

A kl Q0 Q3 - k2 Q1 Q4 + k3 Q1 Q3 - k4-Q2 Q4 ... + k5 Q2 Q3 - k6 -Q5 Q4

- • ( k 5 - Q 2 Q 3 - k 6 Q 5 Q 4 )

Note: Subscripts are vectors. Must start from 0 to n.

ic :=

f ! \

0 0 12

0 V 0 J

Npts := 3000 L0 := 0 LI := 3000

i := 1.. 2600 Wj := 1 - TGi_ i~1

1 100 Xj :=

1 - TGj-! •100

Yj := 1 - TGj_-i

1 100 Zj := 1 - TGm

1 •100

S := rkfixed(IC ,L0 ,L1 ,Npts ,D)

L := TG := S^ DG := S<2> MG := S<3> ME := BD := GY := S^

i := 0.. Npts

78

of canola oil to biodiesel versus length 0.038m (1.5 inch OD) tube.

Figure 27: Conversion

Where:

W400= 50.75

W800 = 71.03

W1200 = 80.51

W16oo= 85.78

W20oo= 89.07

W2400 = 91.27

W258o= 92.02

X400= 50.75

X8oo= 71.03

X1200= 80.51

x1600= 85.78

X 2 0 0 0= 89.07

X2400 = 91.27

X2580 — 92.02

Y 4 0 0 = 50.75

Y 8 0 0 = 7 1 . 0 3

Y 1 2 0 0 = 80.51

Y16OO= 85.78

Y2OOO= 89.07

Y 2 4 0 o= 91.27

Y2580 = 92.02

of reactor for diameter

Z400 = 6 0 . 8

Zgoo= 80.72

Zl200 = 89.06

Z 1 6 0 0= 93.24

Z2ooo= 95.54

Z24oo= 96.87

Z2580 = 97.27

79

MathCAD program for solution of PFR for diameter 2.0 inch OD tube:

A := 0.002 v0 := 0.16

Here

A- Flow Area of tube 2.0 inch tube (m2)

Using same MathCAD program used for diameter 0.038m OD, then the results are:

1C :=

' 1 0 0 12 0

o ; V

Npts := 3000 L0 := 0 LI := 2000

1 - TGn 1 - TGj_i 1 - TGj_j Wj := : 100 Xj := r-1-!--100 Y; := —-100

1

1 - TGj_i Zj := 100

1 1

S := rkfixec IC ,L0,L1 ,Npts,D)

L:=S(0> TG:=S0 > DG := S <2) MG := S ^ ME := S ^ BD := GY := S <6>

i 0.. Npts

80

L,

Figure 28: Conversion of canola oil to biodiesel versus length of reactor for diameter 0.051m (2.0 inch OD) tube.

Where:

W200 = 49.29 X200 = 49.29 Y200 = 49.29 Z200 = 59.27

W 400 = 69.78 ^400 = 69.78 Y400 = 69.78 2400 = 79.58

W 600 = 79.55 ^600 = 79.55 Y600 = 79.55 Z600 = 88.25

Wgoo= 85.01 ^800 = 85.01 Y800 = 85.01 Z800 = 92.66

W1000= = 88.44 X1000= = 88.44 Y1000= = 88.44

Z1000= =95.13

W1140= = 90.15 X1140= = 90.15 Y1140= = 90.15 Zll40= = 96.22

W1200= : 90.75 X1200= = 90.75 Y1200= = 90.75 Z1200= = 96.58

81

Appendix B PFR Sensitivity Results from Aspen Plus Simulation for Diameter

0.038m (1.5 inch OD) and 0.051m (2.0 inch OD) Tubes

Aspen input data for PFR simulation:

7 DYNAMICS

8 DYNAMICS RESULTS=ON

10 TITLE 'SIMULATRION OF PFR REACTOR FOR BIODIESEL FROM

CANOLA OIL'

12 IN-UNITS SI ENTHALPY-J/kg' FLOW='kg/hr' MASS-FLOW='kg/hr* &

13 MOLE-FLO W='kmol/hr' VOLUME-FLO W='cum/hr' PRESSURE=psia

&

14 TEMPERATURES DELTA-T=C PDROP='N/sqm' HEAT-

FLUX='kJ/hr-m'

16 DEF-STREAMS CONVEN ALL

18 SIM-OPTIONS

19 IN-UNITS ENG

20 SIM-OPTIONS OLD-DATABANK=NO

22 DESCRIPTION"

23 Chemical Simulation with English Units :

24 F, psi, lb/hr, lbmol/hr, Btu/hr, cuft/hr.

26 Property Method: NRTL

28 Flow basis for input: Mole

30 Stream report composition: Mole flow

31

33 DATABANKS 'APV71 PURE20' / 'APV71 SOLIDS' / 'APV71 INORGANIC'

&

34 / NOASPENPCD

36 PROP-SOURCES 'APV71 PURE20' / 'APV71 SOLIDS' / 'APV71

INORGANIC'

38 COMPONENTS

39 TGCANOLA C57H10406 /

40 BDIESEL C19H3602 /

41 METHANOL CH40 /

42 GLYCEROL C3H803 /

43 DG C39H7205 /

44

45

47

48

49

51

52

54

55

56

57

58

59

60

62

63

64

MG C21H4004 /

WATER H20

FLOWSHEET

BLOCK R-PLUG IN=PFR-IN OUT=PFR-OUT

BLOCK B3 IN=CANOLA MEOH OUT=PFR-IN

PROPERTIES UNIFAC TRUE-COMPS=YES

PROPERTIES ELECNRTL / IDEAL / NRTL / NRTL-RK

STRUCTURES

STRUCTURES BDIESEL CI C2 S / C2 04 D / C2 03 S / 03 &

C5 S / C5 C6 S / C6 C7 S / C7 C8 S / C8 C9 &

S / C 9 C 1 0 S / C 1 0 C 1 1 S / C l l C12S/C12C13 &

D / C 1 3 C 1 4 S / C 1 4 C 1 5 S / C 1 5 C 1 6 S / C 1 6 &

C 1 7 S / C 1 7 C 1 8 S / C 1 8 C 1 9 S / C 1 9 C20 S / &

C20 C21 S

PROP-DATA

PROP-LIST ATOMNO / NO ATOM

PVAL BDIESEL 6 1 8 / 1 9 . 36. 2.

84

66

67

68

69

70

71

72

73

74

75

76

77

79

80

81

83

84

STRUCTURES

STRUCTURES DG CI C2 S / C2 03 S / 03 C4 S / C4 05 &

D / C4 C6 S / C6 C7 S / C7 C8 S / C8 C9 S / &

C9C10S/C10C11 S / C l l C12 S /C12C13 S / &

C 1 3 C 1 4 S / C 1 4 C 1 5 D / C 1 5 C 1 6 S / C 1 6 C 1 7 S / &

C17 C18 S / C18 C19 S / C19 C20 S / C20 C21 S / &

C21 C22 S / C22 C23 S / CI 024 S / 024 C25 S / &

C25 026 D / C25 C27 S / C27 C28 S / C28 C29 S / &

C29C30S/C30C31 S /C31 C32 S / C32 C33 S / &

C33 C34 S / C34 C35 D / C35 C36 S / C36 C37 S / &

C37 C38 S / C38 C39 S / C39 C40 S / C40 C41 S / &

C41 C42 S / C42 C43 S / CI 044 S

PROP-DATA


PVAL DG 6 1 8 / 39. 72. 5.

STRUCTURES

STRUCTURES GLYCEROL CI C2 S / C2 C3 S / CI 04 S / &

85

85

87

88

89

91

92

94

95

96

98

99

100

101

102

103

104

105

C2 05 S / C3 06 S

PROP-DATA


PVAL GLYCEROL 6 1 8 / 3. 8. 3.

STRUCTURES

STRUCTURES METHANOL CI 0 2 S

PROP-DATA


PVAL METHANOL 6 8 1 / 1 . 1 . 4 .

STRUCTURES

STRUCTURES MG CI 02 S / CI C3 S / C3 04 S / C3 C5 &

S / C5 06 S / 06 C7 S / C7 08 D / C7 C9 S / &

C9C10S/C10C11 S / C l l C12S/C12C13 S / &

C13 C14S/C14C15 D / C 1 5 C 1 6 S / C 1 6 C 1 7 S / &

C 1 7 C 1 8 S / C 1 8 C 1 9 S / C 1 9 C20 S / C20 C21 D / &

C21 C22 S / C22 C23 S / C23 C24 S / C24 C25 S / &

C25 C26S

86

107

108

109

111

112

113

114

115

116

117

118

119

120

121

122

123

124

PROP-DATA


PVALMG6 1 8 /21 .40 . 4.

STRUCTURES

STRUCTURES TGCANOLA CI C2 S / CI C3 S / CI 04 S / &

04 C5 S / C5 06 D / C5 C7 S / C7 C8 S / C8 &

C9 S / C9 CIO S /C10C11 S / C l l C12S/C12 &

C13 S / C 1 3 C 1 4 S / C 1 4 C 1 5 D / C 1 5 C 1 6 S / &

C16C17S/C17C18 S / C 1 8 C 1 9 S / C 1 9 C 2 0 S / &

C20 C21 D / C21 C22 S / C22 C23 S / C2 024 S / &

024 C25 S / C25 026 D / C25 C27 S / C27 C28 S / &

C28 C29 S / C29 C30 S / C30 C31 S / C31 C32 S / &

C32 C33 S / C33 C34 S / C34 C35 D / C35 C36 S / &

C36 C37 S / C37 C38 D / C38 C39 S / C39 C40 S / &

C40 C41 S / C41 C42 S / C42 C43 S / C3 044 S / &

044 C45 S / C45 046 D / C45 C47 S / C47 C48 S / &

C48 C49 S / C49 C50 S / C50 C51 S / C51 C52 S / &

87

125 C52 C53 S / C53 C54 S / C54 C55 D / C55 C56 S / &

126 C56 C57 S / C57 C58 S / C58 C59 S / C59 C60 S / &

127 C60C61 S /C61 C62S/C62C63 S

129 PROP-DATA

130 PROP-LIST ATOMNO / NO ATOM

131 PVAL TGCANOLA 6 1 8 / 57. 104. 6.

133 ESTIMATE ALL

134 IN-UNITS ENG

225 STREAM CANOLA

226 SUBSTREAM MIXED TEMP=25. PRES=4. <atm>

227 MOLE-FLOW TGCANOLA 6.575

229 STREAM MEOH

230 SUBSTREAM MIXED TEMP=25. PRES=4. <atm>

231 MOLE-FLOW METHANOL 78.9

233 STREAM PFR-IN

234 SUBSTREAM MIXED TEMP=15. PRES=15. FREE-WATER=NO

NPHASE=1 &

88

235 PHASE=L

236 MOLE-FLOW TGCANOLA 6.575 / METHANOL 45.4

238 BLOCK B3 MIXER

239 PARAM T-EST=15

241 BLOCK R-PLUG RPLUG

242 PARAM TYPE-T-SPEC LENGTH= 1140. DIAM= 1.76 <in> NPHASE= 1

&

243 PHASE=L PRES=20. PDROP=5. <psi> FLASH=NO

244 T-SPEC 0.0 50.

245 COOLANT TOLO.OOl

246 BLOCK-OPTION RESTART=YES FREE-WATER=NO ENERGY-

BAL=YES &

247 FLASH-METHOD=INSIDE-OUT

248 REACTIONS RXN-IDS=R-1

250 EO-CONV-OPTI

301 STREAM-REPOR MOLEFLOW MASSFLOW

Table 20: Aspen stream results for various methanol to canola oil feed ratios as a the function of reactor length for diameter 0.038m (1.5 inch OD).

MEOH Reactoi Mole Ratio Length RATIO CONV

X BD

DG OUT

MG OUT

GY OUT

ME OUT

TG OUT

kmol/hr meters % mins kmol/

hr kmol/

hr kmol/

hr kmol/

hr kmol/

hr kmol/

hr 1 19.725 2200 3 49.03 9.2 6.32 1.47 0.41 13.41 1.34 3.35 2 19.725 2300 3 49.94 9.6 6.49 1.48 0.40 13.23 1.40 3.29 3 19.725 2400 3 50.78 10.0 6.66 1.48 0.39 13.07 1.47 3.24 4 19.725 2500 3 51.55 10.5 6.82 1.48 0.38 12.91 1.52 3.19 5 19.725 2580 3 52.12 10.9 6.94 1.48 0.37 12.79 1.57 3.15

6 39.45 2200 6 69.25 8.7 11.00 1.14 0.37 28.45 3.04 2.02 7 39.45 2300 6 70.32 9.1 11.28 1.12 0.36 28.17 3.15 1.95 8 39.45 2400 6 71.29 9.5 11.53 1.09 0.34 27.92 3.25 1.89 9 39.45 2500 6 72.20 10.0 11.78 1.07 0.33 27.67 3.35 1.83

10 39.45 2580 6 72.87 10.3 11.96 1.05 0.32 27.49 3.43 1.78

11 59.175 2200 9 80.92 8.3 14.13 0.77 0.29 45.05 4.26 1.25 12 59.175 2300 9 82.00 8.7 14.42 0.74 0.28 44.75 4.38 1.18 13 59.175 2400 9 82.99 9.1 14.70 0.70 0.26 44.48 4.49 1.12 14 59.175 2500 9 83.91 9.5 14.96 0.67 0.25 44.22 4.60 1.06 15 59.175 2580 9 84.59 9.9 15.15 0.65 0.24 44.03 4.67 1.01

16 78.9 2200 12 87.53 7.8 16.00 0.52 0.23 62.90 5.01 0.82 17 78.9 2300 12 88.52 8.2 16.28 0.49 0.21 62.62 5.12 0.75 18 78.9 2400 12 89.42 8.6 16.53 0.45 0.20 62.37 5.23 0.70 19 78.9 2500 12 90.23 9.0 16.76 0.43 0.19 62.14 5.32 0.64 20 78.9 2580 12 90.83 9.3 16.93 0.40 0.18 61.97 5.39 0.60

90

Table 20: Aspen stream results for various methanol to canola oil feed ratios as a the function of reactor length for diameter 0.038m (1.5 inch OD).

MEOH Reactoi Mole Ratio Length RATIO CONV

X BD

DG OUT

MG OUT

GY OUT

ME OUT

TG OUT

kmol/hr m % mins kmol/

hr kmol/

hr kmol/

hr kmol/

hr kmol/

hr kmol/

hr 1 19.73 800 3 45.50 7.5 5.67 1.43 0.44 1.12 14.06 3.58 2 19.73 900 3 48.12 8.5 6.14 1.47 0.42 1.28 13.58 3.41 3 19.73 1000 3 50.30 9.5 6.56 1.48 0.40 1.43 13.16 3.27 4 19.73 1100 3 52.09 10.5 6.93 1.48 0.37 1.57 12.79 3.15 5 19.73 1140 3 52.121 10.9 6.87 1.48 0.37 1.57 12.79 3.15 6 19.73 1200 3 53.58 11.5 7.26 1.48 0.35 1.69 12.46 3.05

7 39.45 800 6 65.09 7.0 9.98 1.22 0.43 2.63 29.47 2.30 8 39.45 900 6 68.19 8.0 10.73 1.17 0.39 2.93 28.72 2.09 9 39.45 1000 6 70.73 9.0 11.38 1.11 0.35 3.19 28.07 1.92 10 39.45 1100 6 72.84 10.0 11.95 1.05 0.32 3.42 27.50 1.79 11 39.45 1140 6 72.871 10.3 11.84 1.05 0.32 3.43 27.47 1.78 12 39.45 1200 6 74.61 11.0 12.44 0.99 0.29 3.62 27.01 1.67

13 59.18 800 9 76.64 6.7 12.98 0.89 0.36 3.79 46.20 1.54 14 59.18 900 9 79.83 7.6 13.83 0.80 0.31 4.14 45.35 1.33 15 59.18 1000 9 82.42 8.5 14.54 0.72 0.27 4.43 44.63 1.16 16 59.18 1100 9 84.55 9.5 15.14 0.65 0.24 4.67 44.04 1.02 17 59.18 1140 9 84.584 9.85 14.99 0.65 0.24 4.67 44.00 1.01 18 59.18 1200 9 86.32 10.5 15.64 0.59 0.21 4.88 43.54 0.90

19 78.90 800 12 83.49 6.3 14.88 0.65 0.29 4.55 64.02 1.09 20 78.90 900 12 86.52 7.2 15.72 0.55 0.25 4.89 63.18 0.89 21 78.90 1000 12 88.90 8.1 16.39 0.47 0.21 5.17 62.51 0.73 22 78.90 1100 12 90.80 9.0 16.92 0.41 0.18 5.39 61.98 0.60 23 78.90 1140 12 90.829 9.34 16.77 0.40 0.18 5.39 61.97 0.59 24 78.90 1200 12 92.32 9.9 17.36 0.35 0.15 5.57 61.54 0.51

91

Appendix C Shell/ Tube Configuration of a PFR

Shell/ Tube configuration for a PFR

Total_Length:= 2580m

LReactor : = 4 m

TotalLength N r u b e s := —

LReactor

N l u b e s = 6 4 5

of diameter 0.038m (1.5 inch OD):

TubeoD := 0.038m Tubqo := 0.03m

LReactor = 13.123 f t

Shell Dia= 3.228m

Shell/Tube configuration for a PFR of diameter 0.051m (2.0 inch OD):

Total_Length:= 1140m

LReactor 4 m

N Tubes : = TotalLength

LReactor

TubeoD := 0.051m TubqD := 0.045m

LReactor = 13.123ft

N-rubes = 2 8 5

Tube Pitch.- 0.10169m

Shell_Dia:= 1.25-Tube_Pitch^NTubes

Shell Dia= 7.04 ft

Shell Dia= 2.146 m

92

Tube_Pitch:= 0.10169m

Shell_Dia:= 1.25 •Tube_PitchA/ N T u b e s

Shell Dia= 10.591ft

Appendix D PFR Pressure Drop in 0.038m (1.5 inch OD) and 0.051m

(2.0 inch OD) Tubes

Pressure Drop caluclations for PFR of diameter 0.038m (1.5 inch OD) Tu

m vel:= 1— L := 4m

s

p := 800 k g f : = 0 .01

Did := 0.03m

K f := 1.5

m

AP := 4.f. + K f (vel)2-p

AP = 0.027 atm

ThisAP is for one 4m long one bend tube

Nbends := 6 4 6

Total_L:= 4m-Nbends

Total_L= 2.5 84 x 103m

APtotal : = AP-Nbends

APtotal = 17.426atm

^reactors 2

Pall Reactor : = NreaC[0rs • APtota|

Pall_Reactor = 34.853 atm

93

Pressure Drop caluclations for PFR of diameter 0.051m (2.0 inch) OD Tube

m vel := 1 .515- L := 4m

s

p := 800

D id := 0.045m

kg f := 0.01 K f := 1.5

m

A P := 4-f-f—— . V ° i d J

+ K f

(vel)2-p

AP = 0.046 atm

ThisDP is for one 4m long one bend tube

N b e n d s := 2 8 6

Total_L := 4m-Nbends

TotalL = 1.144 x 103m

APtotal := A P - N b e n d s

APtota] = 13.101 atm

N r e a c t o r s 1

Pal l Reactor reactors •APtotal

Pal l Reactor = 13.101 atm

94

Appendix E Plug Flow Reactor Cost for Diameter 0.038m (1.5 inch OD) and 0.051m (2.0 inch

OD) Tubes

Reactor Cost for diameter 0.038m (1.5 inch OD) Tube

D = 0.03 m

dollars := 1

12

L := 8465ft LL := 2580m

A := 7tD-L

A = 240.886 m

A = 2.593 x 103ft2

Area := 2593

Fm := 1.75 + 0.13

F m = 5.121

^ Area

V 100 j

Cost := _(exp( 11.0545 - 0.9228 -ln(Area) + 0.0979 -ln(Area)2)) -FM dollars

Cost = 9.707 x 10 dollars

Gear Pump Cost

gal QQ := 65 no_of_gear_pumps := 2

min

- 3 m QQ = 4.101 x 10

s

Q := 65

Fm := 2 For 316 Stainless Steel

CostQearpump :— 2[exp[ 7.2744 + 0.19868 ln(Q) + 0.029-(ln(Q)2) ]-Fm] dollars

CostcearPump = 2.193 x 10 dollars

95

MS Factor := 1538

1097.7

Total_Reactor_Plus_Gear_Pump_Cost_2008= MSFactor

Total_Reactor_Plus_Gear_Pump_Cost_2008= 166727 dollars v

Cost... CostQearpump

Approx Total Reactor Plus Gear Pump Cost 2008= 166700dollars

96

Reactor Cost for diameter 0.051m (2.0 inch OD) Tube

D := — f t D = 0.045 m dollars := 1 12

L := 8465ft LL := 2580m

A := 7iD-L

A = 362.358m2

A = 3.9 x 103ft2

Area := 3900

F m := 1.75 + 0.13

F m = 6.82

^ Area^

V 100 j

Cost := _(exp( 11.0545 - 0.9228-ln( Area) + 0.0979-ln( Area)2) )-FM dollars

Cost = 1.69 x 105 dollars

Gear Pump Cost

QQ := 4 6 - ^ min

3 QQ = 2.902 x 10"3 —

s

Q := 46

Fm := 2 For 316 Stainless Steel

CostQearpump:— [exp[ 7.2744 + 0.19868 ln(Q) + 0.029-(ln(Q)2)]-Fm]dollars

3 CostQearPump = 9.446 x 10 dollars

1538 MS Factor :=

1097.7

97

Total_Reactor_Plus_Gear_Pump_Cost_2008= MS Factor- Cost... + CostGearpump

Total_Reactor_Plus_Gear_Pump_Cost_2008= 250006 dollars

Approx_Total_Reactor_Plus_Gear_Pump_Cost_2008= 250000dollars

98

Appendix F MathCAD Program for a Solution to CSTR

v := 0.175-m min

kmol := lOOOmol

Vr := 7-m 6 := — 9 = 40 min

kl := 0.05-mol-min

k2 := 0.11 mol-min

k3 := 0.215 mol-min

k4 := 1.228-mol-min

k5 := 0.242 • mol-min

k6 := 0.007 • mol-min

C T G O : = 0 . 6 2 6 kmol

m CMEO : = 7.512

kmol

m Cdgo 0

kmol

m

C B D O : = 0 kmol

m CMGO : = 0

kmol

m CGYO : = 0

kmol

m

Guess Values

Ctg := 2 kmol

m CME : = 1

kmol

m CBD : = 1

kmol

m

Cdg := 1 kmol

m Cmg := 1

kmol

m Cqy := 2

kmol

m

Given

CMEO - CME + -k l -Ctg-Cme+ k2-CDG-CBD- k3-CDG'C]viE •• + K 4 -CMG'CBD - K 5 -CMG-CME + K 6 - C Q Y CBD

•0 = 0

CTGO - CTG+ ( - k l -CTG-CME + k2-CDG-CBD)-9 = 0

99

CDGO-CDG+ (kl-CTG'CME-k2-CDG-CBD-k3-CDG-CME+ ^-CMG'CBD)'0 = 0

CMGO~ CMG + k3-CDG'CME- k4-CjviG-CBD- k5-CMG'CME ••• + k6CGY-CBD

CGYO~CGY+ (k5-CMG-CME~ k6-CGY'CBD) -0 = 0

•0 = 0

Cbdo-Cbd- kl -Ctg-Cme- k2-CDG"CBD+ k3-CDG'CME- k4-CMG'cBD -+ k5 -CmG'Cme - k6-CoY-CBD

•0 = 0

CME^ 'CmE

Ctg Ctg

Cdg Cdg := Find

Cmg Cmg

CGy Cgy

Cbd j vCbd

TG = Triglyceride MW=884 ME= Methanol MW=32 DG= Diglyceride MW=620 MG= Monoglyceride MW=356 GY= Glycerol MW=92 BD= BioDiesel MW=296

TG + ME = DG + BD DG + ME = MG + BD MG + ME = GY + BD

Cme^ 5.906^ Ctg 0.066

Cdg 0.031

Cmg 0.013

Cgy 0.516

Cbd, v 1.606 ;

kmol

m

CTG0 = 0.626-kmol

m

CtG0-CTG „ Conv := 100

CtGO

Conv = 89.485

M a s s IN := CMEO-32-— + CTGO-884-— mol moly

100

Mass OUT := 3 2 - & C M E + 8 8 4 - & C T G + 6 2 0 - ^ C d g + 3 5 6 - ^ - C m g mol mol mol mol

+ 9 2 - ^ ] C g y + 2 9 6 - ^ i - C b d mol mol

MassIN = 2.315 — s

BDProduct := C B D v 2 9 6

kg MassOUT = 2.315 — s

kg kmol

BD Product = 4991.2 — hr

101

Appendix G CSTR Sensitivity Results from Aspen Plus Simulation

Aspen input data for CSTR simulation:

7 DYNAMICS

8 DYNAMICS RESULTS=ON

12 IN-UNITS SI ENTHALPY-J/kg' FLOW='kg/hr' MASS-FLOW='kg/hr' &

13 MOLE-FLOW='kmol/hr' VOLUME-FLOW='cum/hr' PRESSURE=psia

&

14 TEMPERATURE=C DELTA-T=C PDROP='N/sqm' HEAT-

FLUX-kJ/hr-m'

15


18 SIM-OPTIONS

19 IN-UNITS ENG


22 DESCRIPTION"



102



30 Stream report composition: Mole flow

31

33 DATABANKS 'APV71 PURE20' / 'APV71 SOLIDS' / 'APV71 INORGANIC

&

34 / NOASPENPCD


INORGANIC'

38 COMPONENTS

39 TGCANOLA C57H10406/

40 BDIESEL C19H3602/

41 METHANOL CH40 /


43 DG C39H7205 /

44 MG C21H4004 /

45 WATER H20

47

48

49

51

52

54

55

56

57

58

59

60

62

63

64

66

67

FLOWSHEET

BLOCK MIX-1 IN=CANOLA MEOH OUT=CSTRIN

BLOCK CSTR-1 IN=CSTRIN OUT=CSTROUT

PROPERTIES UNIFAC TRUE-COMPS=YES

PROPERTIES ELECNRTL / IDEAL / NRTL / NRTL-RK

STRUCTURES

STRUCTURES BDIESEL CI C2 S / C2 04 D / C2 03 S / 03 &

C5 S / C5 C6 S / C6 C7 S / C7 C8 S / C8 C9 &

S / C9C10S / C10C11 S / C l l C12S/C12C13 &

D / C 1 3 C 1 4 S / C 1 4 C 1 5 S / C 1 5 C 1 6 S / C 1 6 &

C 1 7 S / C 1 7 C 1 8 S / C 1 8 C 1 9 S / C 1 9 C20 S / &

C20 C21 S

PROP-DATA


PVAL BDIESEL 6 1 8 /19. 36. 2.

STRUCTURES

STRUCTURES DG CI C2 S / C2 0 3 S / 03 C4 S / C4 05 &

104

68

69

70

71

72

73

74

75

76

77

79

80

81

83

84

85

87

D / C4 C6 S / C6 C7 S / C7 C8 S / C8 C9 S / &

C9C10S/C10C11 S / C l l C12S/C12C13 S / &

C13 C14S/C14C15 D / C 1 5 C 1 6 S / C 1 6 C 1 7 S / &

C17 C18 S / C18 C19 S / C19 C20 S / C20 C21 S / &

C21 C22 S / C22 C23 S / CI 024 S / 024 C25 S / &

C25 026 D / C25 C27 S / C27 C28 S / C28 C29 S / &

C29 C30 S / C30 C31 S / C31 C32 S / C32 C33 S / &

C33 C34 S / C34 C35 D / C35 C36 S / C36 C37 S / &

C37 C38 S / C38 C39 S / C39 C40 S / C40 C41 S / &

C41 C42 S / C42 C43 S / CI 044 S

PROP-DATA


PVAL DG 6 1 8 /39 . 72. 5.

STRUCTURES


C2 05 S / C3 06 S

PROP-DATA

105

88

89

91

92

94

95

96

98

99

100

101

102

103

104

105

107

108



STRUCTURES


PROP-DATA


PVAL METHANOL 6 8 1 / 1 . 1 . 4 .

STRUCTURES


S / C5 06 S / 06 C7 S / C7 08 D / C7 C9 S / &

C9C10S/C10C11 S / C l l C12S/C12C13 S / &

C13 C14 S / C14 C15 D / C15 C16 S / C16 C17 S / &

C17 C18 S / C18 C19 S / C19 C20 S / C20 C21 D / &

C21 C22 S / C22 C23 S / C23 C24 S / C24 C25 S / &

C25 C26 S

PROP-DATA


106

109

111

112

113

114

115

116

117

118

119

120

121

122

123

124

125

126

PVALMG6 1 8 /21 .40 . 4

STRUCTURES


04 C5 S / C5 06 D / C5 C7 S / C7 C8 S / C8 &

C 9 S / C 9 C 1 0 S / C 1 0 C 1 1 S / C l l C12S/C12 &

C 1 3 S / C 1 3 C 1 4 S / C 1 4 C 1 5 D / C 1 5 C 1 6 S / &

C16C17S/C17C18 S / C 1 8 C 1 9 S / C 1 9 C 2 0 S / &

C20 C21 D / C21 C22 S / C22 C23 S / C2 024 S / &

024 C25 S / C25 026 D / C25 C27 S / C27 C28 S / &

C28 C29 S / C29 C30 S / C30 C31 S / C31 C32 S / &

C32 C33 S / C33 C34 S / C34 C35 D / C35 C36 S / &

C36 C37 S / C37 C38 D / C38 C39 S / C39 C40 S / &

C40 C41 S / C41 C42 S / C42 C43 S / C3 044 S / &

044 C45 S / C45 046 D / C45 C47 S / C47 C48 S / &

C48 C49 S / C49 C50 S / C50 C51 S / C51 C52 S / &

C52 C53 S / C53 C54 S / C54 C55 D / C55 C56 S / &

C56 C57 S / C57 C58 S / C58 C59 S / C59 C60 S / &

107

127

129

130

131

133

134

225

226

227

229

230

231

233

234

236

237

238

C60 C61 S / C61 C62 S / C62 C63 S

PROP-DATA


PVAL TGCANOLA 6 1 8 / 57. 104. 6.

ESTIMATE ALL

IN-UNITS ENG

STREAM CANOLA

SUBSTREAM MIXED TEMP=15. PRES=15.

MOLE-FLOW TGCANOLA 6.57

STREAM MEOH

SUBSTREAM MIXED TEMP=15. PRES=15.

MOLE-FLOW METHANOL 78.

BLOCK MIX-1 MIXER

PARAM T-EST=15

BLOCK CSTR-1 RCSTR

PARAM VOL=7. TEMP=50. PRES=15. NPHASE=1 PHASE=L

BLOCK-OPTION FREE-WATER=NO

108

239 REACTIONS RXN-IDS=R-1

241 EO-CONV-OPT

285 PROPERTY-REP PCES NOPARAM-PLUS

Table 22: Aspen steam results for various methanol to canola oil molar feed ratio as function reactor volume of CSTR.

Mole Ratio Volume RATIO CONV

T BD

TG OUT

DG OUT

MG OUT

ME OUT

GY OUT

kmol/hr 3 m % mins kmol/

hr kmol/

hr kmol/

hr kmol/

hr kmol/

hr kmol/

hr 1 19.725 3 3 46.94 18.9 6.72 3.49 1.13 0.27 13.00 1.68 2 19.725 4 3 50.00 25.5 7.36 3.29 1.13 0.25 12.37 1.91 3 19.725 5 3 52.07 32.1 7.81 3.15 1.11 0.23 11.92 2.08 4 19.725 6 3 53.58 38.7 8.15 3.05 1.10 0.22 11.57 2.20 5 19.725 7 3 54.74 45.3 8.42 2.98 1.08 0.21 11.30 2.31

6 39.45 3 6 66.04 18.4 11.03 2.23 0.86 0.26 28.42 3.21 7 39.45 4 6 70.11 25.0 11.97 1.97 0.81 0.23 27.48 3.56 8 39.45 5 6 72.89 31.7 12.63 1.78 0.77 0.21 26.82 3.81 9 39.45 6 6 74.92 38.3 13.11 1.65 0.73 0.20 26.34 3.99 10 39.45 7 6 76.49 45.1 13.49 1.55 0.71 0.19 25.96 4.14

11 59.175 3 9 76.06 17.7 13.51 1.57 0.63 0.23 45.67 4.14 12 59.175 4 9 80.14 24.0 14.48 1.31 0.56 0.20 44.69 4.51 13 59.175 5 9 82.86 30.4 15.14 1.13 0.51 0.18 44.04 4.76 14 59.175 6 9 84.80 36.8 15.61 1.00 0.47 0.17 43.56 4.94 15 59.175 7 9 86.26 43.3 15.97 0.90 0.44 0.15 43.20 5.07

16 78.9 3 12 81.14 16.7 14.82 1.24 0.49 0.21 64.08 4.64 17 78.9 4 12 84.91 22.6 15.74 0.99 0.42 0.18 63.16 4.99 18 78.9 5 12 87.34 28.6 16.34 0.83 0.37 0.16 62.56 5.22 19 78.9 6 12 89.05 34.6 16.76 0.72 0.33 0.14 62.14 5.38 20 78.9 7 12 90.32 40.6 17.07 0.64 0.31 0.13 61.83 5.50

109

Appendix H Sizing of Single CSTR

Reactor Sizing of Single CSTR

kg kg Feed := 8342 — p := 771.9 —

hr - 3 m 3 Feed m

V Flow Rate := V Flow Rate = 0.18 — p min

Reactorresidencetime := 40.6 min

Assume design Residence time=1.5*Reactor residence time

DesignResidenceTime := 1.5 -Reactorresidencetime

DesignResidenceTime = 60.9 min

Volume r := V F l o w R a t e Design Residence Time

Volume r = 10.969 m3

Assume Diameter

(Dia) 2 0 Dia := 5.5 ft Area := n Area = 2.207 mz

4

Volume r H r := H R = 4.97 m

Area

H R = 2.965

Dia

ActualFeedVolume := V F l o w R a t e -Reactorresidencetime

ActualFeedVolume = 7.313 m3

/ T~V • \ 2 Reactor Design Volume := n -HR —

Reactor Design Volume = 10.969 m3

ReactorDesignVolume PercentOverDesign

Actual_Feed_Volume

PercentOverDesign = 1.5

110

Appendix I A Graphical Representation of Purchased Cost of Jacketed and Stirred Reactors

Capacity, geu

Capacity, m 3

Figure 29: Purchased cost of jacketed and stirred reactors (Peters et al., 2003).

I l l

Appendix J

AspenPlus Input File for Three CSTR in Series with Pump Around System

10 TITLE '3 Reactors in Series with Pumparound'

12 IN-UNITS SI MASS-FLOW='kg/hr' MOLE-FLO W='kmol/hr' ENTHALPY-

FLO=kW &

13 PRESSURE=bar TEMPERATURE=C HEAT=kW-hr PDROP='N/sqm'


17 SIM-OPTIONS

18 IN-UNITS ENG


21 DESCRIPTION"





29 Stream report composition: Mole flow 30

32 DATABANKS 'APV71 PURE20' / 'APV71 SOLIDS' / 'APV71 INORGANIC

&

33 / NOASPENPCD


INORGANIC'

37 COMPONENTS

38 TGCANOLA C57H10406/

39 BDIESEL C19H3602/

40 METHANOL CH40 /


42 DG C39H7205 /

43 MG C21H4004 /

44 WATER H20

46 FLOWSHEET

47 BLOCK MIX-1 INCANOLA MEOH OUT=R-IN

48 BLOCK STR-1 IN=R1-RCY Rl-IN OUT=Rl-OUT

49 BLOCK PI IN=Rl-OUT OUT=Pl-OUT

50 BLOCK SPT1 IN=P1-0UT 0UT=R1-RCY R2-IN

51 BLOCK STR-2 IN=R2-IN R2-RCY OUT=R2-OUT

52 BLOCK STR-3 IN=R3-IN R3-RCY OUT=R3-OUT

53 BLOCK P2 IN=R2-OUT OUT=P2-OUT

54 BLOCK P3 IN=R3-OUT OUT=P3-OUT

55 BLOCK SPT2 IN=P2-OUT OUT=R2-RCY R3-IN

56 BLOCK SPT3 IN=P3-OUT OUT=R3-RCY RAW1

57 BLOCK HX1 IN=R-IN OUT=Rl-IN

58 BLOCK COL1 IN=RAW1 OUT=M-RCY RAW2

59 BLOCK SEP1 IN=5 OUT=RAW3 GLY

60 BLOCK HX2 IN=RAW2 OUT=5

62 PROPERTIES UNIFAC TRUE-COMPS=YES

63 PROPERTIES ELECNRTL / IDEAL / NRTL / NRTL-RK

65 STRUCTURES

66 STRUCTURES BDIESEL CI C2 S / C2 04 D / C2 03 S / 03 &

67 C5 S / C5 C6 S / C6 C7 S / C7 C8 S / C8 C9 &

68 S / C 9 C 1 0 S / C 1 0 C 1 1 S / C l l C12S/C12C13 &

114

69

70

71

73

74

75

77

78

79

80

81

82

83

84

85

86

87

D/C13 C14 S /C14 C15 S /C15 C16 S /C16 &

C 1 7 S / C 1 7 C 1 8 S / C 1 8 C 1 9 S / C 1 9 C20 S / &

C20 C21 S

PROP-DATA


PVAL BDIESEL 6 1 8 / 19. 36. 2.

STRUCTURES

STRUCTURES DG CI C2 S / C2 0 3 S / 0 3 C4 S / C4 0 5 &

D / C4 C6 S / C6 C7 S / C7 C8 S / C8 C9 S / &

C9C10S/C10C11 S / C l l C 1 2 S / C 1 2 C 1 3 S / &

C13 C14 S / C14 C15 D / C15 C16 S / C16 C17 S / &

C17C18S /C18C19S / C19 C20 S / C20 C21 S / &

C21 C22 S / C22 C23 S / CI 024 S / 024 C25 S / &

C25 026 D / C25 C27 S / C27 C28 S / C28 C29 S / &

C29 C30 S / C30 C31 S / C31 C32 S / C32 C33 S / &

C33 C34 S / C34 C35 D / C35 C36 S / C36 C37 S / &

C37 C38 S / C38 C39 S / C39 C40 S / C40 C41 S / &

115

88

90

91

92

94

95

96

98

99

100

102

103

105

106

107

109

110

C41 C42 S / C42 C43 S / CI 044 S

PROP-DATA


PYAL DG 6 1 8 / 39. 72. 5.

STRUCTURES


C2 05 S / C3 06 S

PROP-DATA

PROP-LIST ATOMNO / NOATOM


STRUCTURES


PROP-DATA


PVAL METHANOL 6 8 1 / 1 . 1 . 4 .

STRUCTURES


116

I l l

112

113

114

115

116

118

119

120

122

123

124

125

126

127

128

129

S / C5 06 S / 0 6 C7 S / C7 0 8 D / C7 C9 S / &

C9C10S/C10C11 S / C l l C12S/C12C13 S / &

C 1 3 C 1 4 S / C 1 4 C 1 5 D / C 1 5 C 1 6 S / C 1 6 C 1 7 S / &

C17 C18 S / C18 C19 S / C19 C20 S / C20 C21 D / &

C21 C22 S / C22 C23 S / C23 C24 S / C24 C25 S / &

C25 C26S

PROP-DATA


PVAL MG 6 1 8 /21 .40 . 4.

STRUCTURES


04 C5 S / C5 06 D / C5 C7 S / C7 C8 S / C8 &

C 9 S / C 9 C 1 0 S / C 1 0 C 1 1 S / C l l C12S/C12 &

C13 S / C13 C14 S / C14 C15 D / C15 C16 S / &

C 1 6 C 1 7 S / C 1 7 C 1 8 S / C 1 8 C 1 9 S / C 1 9 C20 S / &

C20 C21 D / C21 C22 S / C22 C23 S / C2 024 S / &

024 C25 S / C25 026 D / C25 C27 S / C27 C28 S / &

117

130

131

132

133

134

135

136

137

138

140

141

142

144

145

232

233

234

C28 C29 S / C29 C30 S / C30 C31 S / C31 C32 S / &

C32 C33 S / C33 C34 S / C34 C35 D / C35 C36 S / &

C36 C37 S / C37 C38 D / C38 C39 S / C39 C40 S / &

C40 C41 S / C41 C42 S / C42 C43 S / C3 044 S / &

044 C45 S / C45 046 D / C45 C47 S / C47 C48 S / &

C48 C49 S / C49 C50 S / C50 C51 S / C51 C52 S / &

C52 C53 S / C53 C54 S / C54 C55 D / C55 C56 S / &

C56 C57 S / C57 C58 S / C58 C59 S / C59 C60 S / &

C60 C61 S / C61 C62 S / C62 C63 S

PROP-DATA


PVAL TGCANOLA 6 1 8 / 57. 104. 6.

ESTIMATE ALL

IN-UNITS ENG

STREAM CANOLA

IN-UNITS ENG

SUBSTREAM MIXED TEMP=59.00000 PRES=15.

118

235

237

238

239

240

242

243

244

245

247

248

249

251

252

254

255

257

MOLE-FLOW TGCANOLA 6.95 <kmol/hr>

STREAM MEOH

IN-UNITS ENG

SUBSTREAM MIXED TEMP=59.00000 PRES=15.

MOLE-FLOW METHANOL 104.25 <kmol/hr

BLOCK MIX-1 MIXER

IN-UNITS ENG

PARAM PRES=1. <bar> NPHASE=1 PHASE=L T-EST=59.


BLOCK SPT1 FSPLIT

IN-UNITS ENG

FRAC Rl-RCY 0.9

BLOCK SPT2 FSPLIT

FRAC R2-RCY 0.9

BLOCK SPT3 FSPLIT

FRAC R3-RCY 0.9

BLOCK HX1 HEATER

119

258

260

261

263

264

266

267

268

269

270

271

272

273

274

276

277

278

PARAM TEMP=50. PRES=4.

BLOCK HX2 HEATER

PARAM TEMP=25. PRES=4.

BLOCK SEP1 DECANTER

PARAM TEMP=25. PRES=2. L2-COMPS=GLYCEROL

BLOCK COL1 RADFRAC

PARAM NSTAGE=6

COL-CONFIG CONDENSER=TOTAL

FEEDS RAW1 4

PRODUCTS M-RCY 1 L / RAW2 6 L

P-SPEC 1 1.1

COL-SPECS DP-COL=0. MOLE-D=83.8 MOLE-RR=2.

TRAY-SIZE 1 2 5 SIEVE

TRAY-RATE 1 2 5 SIEVE DIAM= 1.40946 P-UPDATE=NO

BLOCK STR-1 RCSTR

IN-UNITS ENG

PARAM VOL=3. <cum> TEMP=50. <C> PRES=1. <bar> NPHASE=1 &

120

279

280

281

283

284

285

286

288

289

290

291

293

294

295

297

298

300

PHASE=L


REACTIONS RXN-IDS=R-1

BLOCK STR-2 RCSTR




BLOCK STR-3 RCSTR




BLOCK PI PUMP

IN-UNITS ENG

PARAM PRES=5. <bar>

BLOCK P2 PUMP

PARAM PRES=5.

BLOCK P3 PUMP

121

301 PARAM PRES=5

303 EO-CONV-OPTI

305 SENSITIVITY S-l

306 IN-UNITS ENG

307 DEFINE OILOUT MOLE-FLOW STREAM=RAW 1

SUBSTREAM=MIXED &

308 COMPONENT=TGCANOLA

309 DEFINE OILIN MOLE-FLOW STREAM=R-IN SUBSTREAM=MIXED

&

310 COMPONENT=TGCANOLA

311 DEFINE TIME 1 BLOCK-VAR BLOCK=STR-1 VARIABLE=RT-C ALC

&

312 SENTENCE=RESULTS

313 DEFINE TIME2 BLOCK-VAR BLOCK=STR-2 VARIABLE=RT-CALC

&


315 DEFINE TIME3 BLOCK-VAR BLOCK=STR-3 VARIABLE=RT-C ALC

&

122


317 F CONV=(OILIN-OILOUT)* lOO/OILIN

318 F MIN=TIME*60

319 TABULATE 1 "CONV"

320 TABULATE 2 "TIME1"



323 VARY MOLE-FLOW STREAM=MEOH SUBSTREAM=MIXED &

324 COMPONENT=METHANOL

325 RANGE LOWER="20.85" UPPER="41.7" INCR="20.85"

327 BLOCK-REPORT INCL-BLOCKS=COLl HX1 HX2 MIX-1 PI P2 P3 SEP1

&

328 SPT1 SPT2 SPT3 STR-1 STR-2 STR-3

330 STREAM-REPOR MOLEFLOW MASSFLOW MASSFRAC INCL-

STREAMS=5 CANOLA &

331 GLY M-RCY MEOH PI-OUT P2-OUT P3-OUT R-IN Rl-IN Rl-OUT

&

332 Rl-RCY R2-IN R2-OUT R2-RCY R3-IN R3-OUT R3-RCY RAW1 &

123

333 RAW2 RAW3

335 PROPERTY-REP NOCOMPS NOPARAM-PLUS

337 REACTIONS R-l POWERLAW

338 IN-UNITS ENG

339 REAC-DATA1

340 REAC-DATA 2

341 REAC-DATA 3

342 REAC-DATA 4

343 REAC-DATA 5

344 REAC-DATA 6

345 RATE-CON 1 PRE-EXP=0.0008333 ACT-ENERGY=0.0 TEMP-

EXPONENT. &

346 T-REF=59.00000


EXPONEN=l. &

348 T-REF=59.00000

f


EXPONENT. &

124

350 T-REF=59.00000


EXPONEN=l. &

352 T-REF=59.00000


EXPONEN=l. &

354 T-REF=59.00000


EXPONEN=l. &

356 T-REF=59.00000

357 STOIC 1 MIXED TGCANOLA -1. / METHANOL -1. / BDIESEL 1. / &

358 DG 1.

359 STOIC 2 MIXED DG -1. / METHANOL -1. / BDIESEL 1. / MG &

360 1.

361 STOIC 3 MIXED MG -1. / METHANOL -1. / BDIESEL 1. / &

362 GLYCEROL 1.

363 STOIC 4 MIXED DG -1. / BDIESEL -1. / TGCANOLA 1. / &

364 METHANOL 1.

125

365 STOIC 5 MIXED MG -1. / BDIESEL -1. / DG 1. / METHANOL &

366 1.

367 STOIC 6 MIXED GLYCEROL -1. / BDIESEL -1. / MG 1. / &

368 METHANOL 1.

369 POWLAW-EXP 1 MIXED TGCANOLA 1. / MIXED METHANOL 1. / &

370 MIXED BDIESEL 0. / MIXED DG 0.

371 POWLAW-EXP 2 MIXED DG 1. / MIXED METHANOL 1. / MIXED &

372 BDIESEL 0. / MIXED MG 0.

373 POWLAW-EXP 3 MIXED MG 1. / MIXED METHANOL 1. / MIXED &

374 BDIESEL 0. / MIXED GLYCEROL 0.

375 POWLAW-EXP 4 MIXED DG 1. / MIXED BDIESEL 1. / MIXED &

376 TGCANOLA 0. / MIXED METHANOL 0.

377 POWLAW-EXP 5 MIXED MG 1. / MIXED BDIESEL 1. / MIXED DG &

378 0. / MIXED METHANOL 0.

379 POWLAW-EXP 6 MIXED GLYCEROL 1. / MIXED BDIESEL 1. / &

380 MIXED MG 0. / MIXED METHANOL 0.

Appendix K Design of Methanol Recovery Column Calculations

Top and Bottom Property Data from AspenPlus Simulation

kg P L_Top := 7 3 9 . 8 2 ^ pL_Bottom := 8 8 4 . 1 3

kg 3

m kg

m p is density o f the material

N PV_Top := 1-2639 p v _ B o t t o m : = 2 .6121 ^ °Bottom == 0 - 0 1 8 6 9 9

m

N

m m

a T o p := 0 . 0 3 7 7 9 3 — L T o p := 5 3 6 4 . 2 ^ ~ hr m

V T o p := 8045.8 kg

hr Vsottom := 14649 ^

hr

a is surface tension

b o t t o m := 2 1 4 4 5 ^ hr

m C sb Top := 0 . 0 5 9 — c s b Bottom := 0 . 0 5 6 — C s b = Fair's F lood Factor

Column Diameter

f T FLV_Top :

^Top

Lt, op

V VTop J

= 0 . 6 6 7

P V Top

P L T o p

0.5

vTop

FLV_Top = 0 .028

FLV_Bottom •'=

^Bottom

LBottom

VBottom

P VBottom

V P L Bottom

v 0 .5

^Bottom = 1.464

F L V Bottom = 0 .08

127

^nf Top := Csb_Top f \0.2 aTop

20 dyne

,PL_Top ~ P V Top j

PVTop

0.5

V cm y ft

VnfTop = 5.314 —

vnf_Top = 1.62 — s

V„f Bottom Csb Bottom r \0.2 CTBottom

20 dyne

, P L Bottom ~ PVBottom)

PV Bottom

0.5

V cm y

Vnf Bottom = 3.33 — s

VnfBottom = 1.015 — S

Assuming Flooding = 75%

Assume downcomer occupies 15% of cross-sectional area

vn_Top 0.75-Vnf Top

Vn_Top = 3.986-s

Vn Top = 1.215-s

Vn Bottom : = 0.75-Vnf Bottom

VnBottom = 0.761 — S

V_dotTop := VT, op

PVTop

ft V_dotTop = 62.447 —

V_dotTop = 1.768 — s

128

^Bottom V_dotBottom :=

P V Bottom

ft3

V_dotBottom = 55.014 — s

m3

V_dotBottom = 1-558 — s

V_dotTop AreaTop := —

V n T o p

AreaTop = 15.667 ft2

2 Areaxop = 1.456 m

Areaxop D T o p : = / 4 -

71

DTop = 1.361 m

V_dotBottom AreaBottom : -

'n Bottom

2 Area B o t t 0 m = 22.027 ft

2 Area B o t t 0 m = 2.046 m

AreaB o ttom D B o t t o m := / 4

71

^Bot tom = 5.296 ft

DBottom = 1 . 6 1 4 m

Tray Hydraulics

Rules of thumb for Tray specidicatic

Tray th ickness = t t r a y = 14 g a u g e ( 0 . 0 7 8 i n c h / 0 . 0 0 1 9 m )

H o l e diameter = 0 . 0 0 4 7 m ( 3 / 1 6 inch)

C o l u m n Area = A x

D o w n c o m e r Area = A d = 1 2 % o f A x

N e t Area = A n = 8 8 % o f A T ( V a p o u r F l o w Area)

A c t i v e A r e a = A a = ( A T - 2 A d )

H o l e Area = A h = 10% A T

Weir Length = 1 w = 7 7 % o f Diameter ( T o p or B o t t o m )

2

ATBottom := 2 2 . 1 1 5 f t

2

AT_Bottom = 2 . 0 5 5 m

An Bottom : = 0 . 8 8 - A x Bottom

A n B o t t o m = 19 .461 ft2

2 An Bottom = 1-808 m

A d B o t t o m := 0 . 1 2 - A x Bottom

2 Ad Bottom = 2 . 6 5 4 ft Ad Bottom = 0 - 2 4 7 m 2

Aa_Bottom : = ( A T_Bot tom " 2-A(j_Bottom )

2

Aa_Bottom = 16 .807 ft

2 Aa_Bottom = 1.561 m

Ah Bottom •'= 0-1 Aa_Bottom

2 AhBottom = L681 ft 2

AhBottom = 0 . 1 5 6 m

130

IwBottom 0.77-Deottom

IwBot tom = 4.078 ft

lw Bottom = 1 . 2 4 3 m

^Bot tom Vr, :=

P v Bottom' A h Bottom

v0 = 32.732-ft s

v 0 = 9 . 9 7 7 -m s

Dry Pressure Drop

dhole := — in 16

- 3 dhole = 4 . 7 6 2 x 10 m

ttray := 0.078in

t t r a y = 1 .981 x 10 m

Ahole := ™ (dhole2)

Ahole = 1-781 X 10 5 m 2

C0 := 0.85032-0.04231-f dhole

V W ) + 0 . 0 0 1 7 9 5 4 -

r dhole V

^ ttray j

131

C0 = 0.759

Pwate r := 6 1 . 0 3 -lb ft?

Pwate r = 9 7 7 . 6 0 7 kg

m

Ah Bottom Aa Bottom

|3 = 0.1

v0 = 32.732

v0 = 9 .977-

lb PL_Bottom= 55.194 — ft

PL_Bottom= 8 8 4 . 1 3 ^ m

lb Pv_Bottom =0.163 — ft

hAP_Dty := 0.003-' f tV

PV_Bottom lb

f P water 1 in

— ^PL_Bottom J

hiP_Dry = 0.996 in

132

hAP Dry = 0.025 m

For 75% flooding

Flv = 0.05, y = 0.045(relative entrainment)

v|/ := 0.045

L Bottom

LBottom

L Bottom ee := u/

3 lb ee = 2.228 x 10 —

hr

ee = 0.281 — s

Entrainment := LBottom + e e

4 lb Entrainment = 4.951 x 10 —

hr

Entrainment = 6.238 — s

Fw=Weir Correction Factor due to wall curvature. For large Diameter Column Fw approaches 1.0 otherwise use figure by Bolles

Lg = Liquid Flow rate including entrainmen (L+ee) in the units of gallons/min

(ee + LBottom) Lg :=

P l Bottom

4 l b = 4.728 X 10 —

hr

= 5.957 kg

133

L„ = 111.826 gal min

LP = 7.055 x 10 3 —

Bottom = 4.078 ft

l w Bottom = 1 . 2 4 3 m

Abscissa := \2-5

V Bottom

ft

Abscissa = 3.33

Parameter w Bottom

^ B o t t o m

Parameter = 0.77

F\vear 1 - 0 3 5 (fromgraph)

hcrest := 0.092-Fv

f La ^ gal

V min y

'w_Bottom

ft

•in

hcrest = 0.866 in

hcrest = 0.022 m

134

Assume negligible h gradient. See A P Figure

hgrad := Oin

Assume Downcomer gap= 1 inch

Gap := lin

Gap = 0.025 m

A d u := Iw Bottom - G a p

Ad u = 0.34 ft2

Ad u = 0.032 m2

hdu := 0.56 •

hdu = 0.301 in

hdu = 7.641 x 10" 3 m

hdc = total Head of clear liquid in the Downcomer

Assume Weir Height = 2.0 inches

hweir := 2in

hweir = 0 . 0 5 1 m

hdc : = h A P_Dry + h w e i r + hcrest + hdu

hdc = 4.163 in

gal

V min y

Ad u 449

ft2

•in

135

hdc = 0.106m

Note : In the operating column, the liquid in the Downcomer is Aerated. The densi aerate liquid will be less than that of clear liquid. Thus the height of liquid in Downcomer will be greater thar\j|a

For normal operatiorfdc=0.5

(j)dc := 0.5

hdc hdcae ra t ed

<Pdc

hdcae ra ted = 8.325 in

hdcae ra t ed = 0 . 2 1 1 m

This value is less than 24 inch tray spacing.

There should be no problem for vapour flow to the tray above.

Check for Downcomer time tdc

hdc = 4.163 in

hdc

tdc :

tdc = 3.695 s

The value is greater the minimum residence time of 3 seconds.

There should be no problem.

- 0.106m

( A d . Bottom "PL Bottom "hdc

^Bot tom + e e

136

Weeping Check

Caluclate surface tension head h a

Excessive weeping based on Fair correlation

if X = IVEIR+HCREST+HGRAD then

hAP_Dry+h CT >0.10392+0.25119x-0.021675x2

a := 13.2 dyne cm

dhole := 0.187in

- 3 dhole = 4 . 7 5 x 10 m

ho := 0.04in-

< a ^ dyne

V cm J

P L B o t t o m

j b

ft3

\ f dhole ^

V in J

ho = 0.051 in

- 3 dhole = 4 . 7 5 x 10 m

ho := 0.04in-

' a ^ dyne

V cm J

P L B o t t o m

lb

ftJ

\ ^ dhole

V in J

ho = 0.051 in

137

h a = 1 . 2 9 9 X 1 0 3 m

hwei r = 2 i n

hwei r = 0 . 0 5 1 m

hgrad = 0 i n

hcrest = 0.866 in

hcrest = 0 . 0 2 2 m

x := (2 + 0.865 + 0)

x = 2.865

H A P _ D R Y + h a >0.10392+0.25119x-0.021675x 2

RHS := 0.10392 + 0.2511 - x - 0.021675 -x2

RHS = 0.645

hAP Dry - 0.996 in

hAP Dry = 0.025 m

ha = 0.051 in

ha = 1.299 x 10~3m

hAP Dry + ha = 0.027 m

hAP Dry + h a = 1.047 in

LHS := 1.047

LHS (1.047) > RHS (0.645)

The inequality is satisfied. Therefore, Weeping should not be a problem.

138

BIOGRAPHICAL SKETCH

BIOGRAPHICAL SKETCH

Name of the Author: Shali Vemparala

Place of Birth: Hyderabad, India

Date of Birth: July 14, 1985

Graduate and Undergraduate Schools Attended:

University of South Alabama, Mobile, Alabama Bhoj Reddy Engineering College for Women, Hyderabad, India

Degrees Awarded:

Masters of Science in Chemical Engineering, 2010, Mobile, Alabama Bachelors of Science in Chemical Engineering, 2007, Hyderabad, India

139

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