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Ministry of Higher Education & scientific research University of Technology Chemical Engineering Department
Experimental and Dynamic Simulation of
catalyzed Esterification of n-Butanol and
Acetic acid in Reactive Distillation column
A Thesis
Submitted To The
Department of Chemical Engineering at the University of
Technology in a Partial Fulfillment of the Requirements for
the Degree of Master of Science in Chemical Engineering
By Mohammed S. Baqer
B.Sc. in Chemical Engineering, 2008
Supervised by
Ass. Prof. Dr. Khalid A. Sukkar Dr. Zaidoon M. Shakoor
June / 2011
مهورية العراقج
وزارة التعليم العالي والبحث العلمي الجامعة التكنولوجية
لألسترة المحفزة الداينميكية العملية و المحاكاتبرج التقطير التفاعليحامض الخليك في و للبيوتانول
رسالة مقدمة إلىقسم الهندسة الكيماوية في الجامعة التكنولوجية وهي جزء
متطلبات نيل درجة الماجستير في علوم الهندسة من الكيماوية
من قبل محمد سعد باقر
8002بكالوريوس في الهندسة الكمياوية
بأشراف خالد عجمي سكر. د.م .أ
زيدون محسن شكور.د 8022حزيران
Certification
I certify that this thesis entitled ( Experimental and Dynamic
Simulation of catalyzed Esterification of n-Butanol and
Acetic acid in Reactive Distillation column) was prepared
under my linguistic supervision. It was amended to meet the style of
English Language.
Signature: Name: Prof. Dr. Mumtas A. Zablouk
Date: / / 2011
Certificate of Supervisor
I certify that this thesis has been concluded under my supervision
in a partial fulfillment of the requirements for the Degree of Master
of Science in Chemical Engineering at the Chemical Engineering
Department, University of Technology.
Signature:
Name: Ass. Prof. Dr. Khalid A. Sukkar Date: / / 2011 (Supervisor)
Signature:
Name: Dr. Zaidoon M. Shakoor Date: / / 2011 (Supervisor)
In view of the available recommendations, I forward this thesis
for debate by the Examining Committee.
Signature: Name: Asst. Prof. Dr. Mohamed I. Mohamed Head of Post Graduate Committee Department of Chemical Engineering
Date: / / 2011
Certificate of Examiners
We certify that we have read this thesis and as examining committee
examined the student (Mohammed Saad Baqer) in its contents and
that in our opinion it meets the standard of a thesis for the degree of
Master of Science in chemical engineering.
Signature: Signature:
Name: Ass. Prof. Dr. Khalid A. Sukkar Name: Dr. Zaidoon M. Shakoor Data: / / 2011 Data: / / 2011 (Supervisor) (Supervisor)
Signature: Signature: Name: Ass. Prof. Dr. Malik M. Mohammed Name: Dr. Khalid F. Chasib Data: / / 2011 Data: / / 2011 (Member) (Member)
Signature: Name: Prof. Dr. Safa Aldin Abdullah Data: / / 2011 (Chairman)
Approved by the Head of the Chemical Engineering Department
Signature: Name: Prof. Dr. Mumtas A. Zablouk Head of Chemical Engineering Department Data: / / 2011
Acknowledgment Thanks be to Allah Who gave me ability to achieve this research.
I would like to thank all the people who supported my research and me during the
past year. Foremost have been my advisors Dr. Khalid A. Sukkar and Dr. Zaidoon
M. Shakoor. I appreciated very much their guidance, suggestions, their manner of
carrying out research and their intriguing way of working and thinking.
I would also like to express my acknowledgment to the Head and to the staff of
Chemical Engineering Department of the University of Technology.
Special thanks belong to Engineer Abeer S. Mahmood for her great help in analyses
using gas chromatograph.
And finally my special thanks to my family for their support and
encouragement.
Mohammed
Abstract
In this work, the production of n-butyl acetate was carried out
successfully using solid catalysts in a continuous reactive
distillation column. The present study includes three parts; kintects
study, reactive distillation unit, and mathematical modeling.
In the first part the kinetic of the esterification reaction to produce
n-butyl acetate in a heterogeneous catalyzed batch reactor was
studied, using n-butanol and acetic acid as reactants. Two types of
catalysts Dowex-50 and Amberlite CG-50 were investigated. The
catalysts were modified with 0.1 N HCl. The results show that the
modified catalysts with HCl gave higher activity than parent
catalysts types.
The results obtained from kinetics study show that the modified
Dowex-50 catalyst gives the highest conversion of n-butanol and
acetic acid to produce n-butyl acetate. The n-butanol conversion is
67% in batch reactor with activation energy equal to39.975 kJ/mol.
On the other hand, a Pseudo-Homogeneous Model was developed
to describe the reaction kinetics. The comparison between
calculated results and experimental results shows a very good
agreement between them.
In the second part, a continuous reactive distillation column made of
QVF glass was constructed. The effect of reflux ratio, acetic acid
flow rate and heat duty on the performance of reactive distillation
with the best catalyst (modified Dowex-50) was studied. It is
concluded that, when the reflux ratio increases the temperature level
along the column decreases. On the other hand, the increase of
acetic acid flow rate or heat duty lead to a slightly affect on
temperature level. Also, it was noted that when the acetic acid flow
rate increases, the reaction zone temperature increases too, while,
an opposite results was noted, when heat duty is increased the
reaction zone temperature decreased.
In the present work, unsteady state mathematical analysis was
derived using MATLAB program. The set of algebraic equations
governing composition profile in a reactive distillation column are
solved by using Gausses elimination method. The model was used
effectively to describe: compositions, flowrates and temperatures
inside the column. The results indicated that, the conversion in
reactive distillation increases directly with increasing the reflux ratio.
On the other hand, it was noted that the decreasing of the ratio
between stages holdup and reboiler holdup will increase the speed of
column response and decrease the time to reach steady state value in
dynamic distillation columns.
Contents
I
Contents
Subject Page no.
Contents …………………….………………………………………………..…..I
Nomenclature………………………………………………..……...………..…..V
CHAPTER ONE: INTRODUCTION
1.1 Introduction……………………………………………………………….....1
1.2 The Advantage and Disadvantage of Reactive Distillation………….……...4
1.3 Industrial Applications of Reactive Distillation…………….....………..…..5
1.4 Aims of the Present Work…………………………………………………...6
CHAPTER TWO: LITERATURE SURVEY
2.1 SCOPE………………………………………………………………………..7
2.2 Fundamental of Reactive Distillation………………………………..………..7
2.5 Esterification Catalysts and Processes…………………………………...….11
2.5.1 Homogeneous Acid Catalysts………………………………………….11
2.5.2 Heterogeneous acid catalysis…………………………………………..12
1) Metal ion complexes as catalysts…………………………………….12
2) Zeolites as catalysts………………………………………….……….12
3) Ion exchange resins as catalysts…………………………….………..13
2.5.3 Effect of Catalysts Type on Reactive Distillation……………….……..13
2.6 Thermodynamics of Reactive Separations…………………………………..14
2.6.1 Ideal Solution…………………………………………………………..15
2.6.2 Non Ideal Solutions…………………………………………………….15
2.6.3 Heat of Reaction………………………………………………………..15
2.6.4 Pressure Drop…………………………………………………………..16
Contents
II
2.7 Reactive Distillation Internal……………………………………….….…….16
2.7.1 Homogeneous Reactive Distillation………………………….………...16
2.7.2 Heterogeneous Reactive Distillation…………………………….……..18
2.7.2.1 Packed Reactive Distillation…………………………….………18
2.7.2.2. Trayed Reactive Distillation…………………………….……...20
2.8 Production of Butyl acetate as Case Study…………………………….…….24
2.8.1 Reaction Kinetic………………………………………………….…….24
2.8.2 Production of n-Butyl acetate (Previous work)…………………….…..25
2.9 Modeling of Reactive Distillation……………………………………….…..31
2.9.1 Equilibrium Model……………………………………………….…….31
2.9.2 Rate Based Model………………………………………………….…..35
CHAPTER THREE: EXPERIMENTAL WORK
3.1 Introduction……………………………………………………………….…37
3.2 Materials and Analysis………………………………………………….…...37
3.3 Catalysts Modification………………………………………………….……38
3.4 Experiments of Kinetic Study and Procedure………………………………..38
3.4.1 Experimental Procedure in Kinetic unit………………………………..40
3.5 Production of Butyl Acetate Using Continuous Pilot Plant Reactive
Distillation and Procedure…………………………………………………...41
3.5.1 Experimental Procedure in RD Unit……………………………………45
3.6 Sample Analysis……………………………………………………………..45
3.7 Thermocouple Calibration…………………………………………………...46
CHAPTER FOUR: Mathematical Representation, Modeling and Simulation
4.1 Introduction………………………………………………………………...48
4.2 Parameters Estimation for the Reaction Kinetics……………………………49
4.3 Mathematical Model…………………………………………………………51
4.3.1 Model Assumptions…………………………………………………...51
Contents
III
4.3.2 Model Equations………………………………………………………..54
4.3.3 Estimation of Model Parameters……………………………………….56
4.3.3.1 Antoine Model…………………………………………………..56
4.3.3.2 Equilibrium relations……………………………………………57
4.3.3.3 Bubble Point Calculation………………………………………..57
4.3.3.4 Enthalpy calculation…………………………………………….58
4.3.3.5 Tray molar holdup………………………………………………58
4.4 Numerical Computation……………………………………………………..59
CHAPTER FIVE: RESULTS AND DISCUSSION
5.1 Introduction…………………………………………………………….……61
5.2 Reaction Kinetic……………………………………………………….…….61
5.2.1 Effect of Temperature…………………………………………….……61
5.2.2 Effect of catalyst type and modification with HCl……………….……62
5.2.3 Effect of Catalyst Loading………………………………………….….63
5.2.4 Effect Feed Ratio………………………………………………….……63
5.2.5 Parameters Estimation of Reaction Kinetic…………………….………66
5.2.6 Validity of Reaction Kinetic……………………………………….…...66
5.3 Experimental Unit Results………………………………………………...…69
5.3.1 Temperature Distribution along the Reactive Distillation……………..69
5.3.2 Effect of Reflux Ratio………………………………………………….72
5.3.3 Effect of Heat Duty…………………………………………………….76
5.3.4 Effect of Feed Ratio……………………………………………………76
5.3.5 Condenser and Reboiler Temperature………………………………….76
5.4 Mathematical Model Results……………………………………………...…81
5.4.1 Validity of Mathematical Model…………………………………….....84
5.4.1.1 Comparison of Compositions Profile…………………………...85
Contents
IV
5.4.1.2 Comparison of temperature profile………………………..……86
5.5 Process Dynamics…………………………………………………………...88
5.5.1 Response to Reflux Ratio……………………………………………...88
5.5.2 Response to Feed Flow Rate………………………………………..…89
5.6 Determination of Reflux Ratio……………………………………………...92
CHAPTER SIX: CONCLUSIONS AND RECOMMENDATIONS
6.1 Conclusions…………………………………………………………………93
6.2 Recommendations for the Future Work…………………………………….94
Appendix
Appendix A: Technical Data…………………………………………………….95
Appendix B: UNIQUAC Model………………………………………………...97
Appendix C: Calibration Data…………………………………………………...99
Appendix D: Simulation and Experimental Data………………………………107
References………………………………………………………………………….114
Nomenclature
V
Nomenclature
Symbols Definitions Units A,B,C and
D
Reaction and product components (....)
ai Activity of ith
component in the bulk
of liquid phase
(….)
Ci Concentration of component i mol/m3
CP Specific heat J/mol .K
D Distillate flow rate [ mol/hr]
EA Activation energy J/mole
F Feed flow rate ml/min
Fobj. Objective function (….)
∆hr Heat of reaction [J/mol]
hi Liquid phase enthalpy of component i [J/mol]
Hi Vapor phase enthalpy of component i [J/mol]
Ka Equilibrium constant (….)
Kf Forward reaction rate constant mole/gcata hr
L Liquid molar flow rate [mol/hr]
M Molar holdup mol
Mcat Catalyst mass g
N Total no. of trays (….)
n No. of component (….)
P Total pressure atm
Psat
Saturated pressure atm
QR Reboiler Heat duty Watt
QC Condenser Heat duty Watt
R Gas constant 8.314 J/mol K
Rref. Reflux ratio (….)
rj Rate of reaction mole/ gcata.hr
Sj Side stream flow rate [mol/hr]
T Temperature K
t time hr
V Vapor molar flow rate [mol/hr]
xi liquid phase mole fraction of the
component i
(….)
Nomenclature
VI
yi vapor phase mole fraction of the
component i
(….)
Greek Latter Symbols Definitions
γi Activity coefficient of the ith
component
∆ difference
νi Stoichiometric coefficient of component i
Fugacity coefficient
Subscript Symbols Definitions
j Stage no.
i Component no.
Cata, Catalyst
Abbreviation Symbols Definitions BuOH n-Butanol
BuAc n-Butyl acetate
CD Catalytic distillation
DBE di-butyl ether
ETBE Ethyl-tert-butyl ether
EtOH Ethanol
EtAc Ethyl acetate
EQ Equilibrium
FID Flame ionization detector
FR Acetic acid flow rate (mol/min)
Nomenclature
VII
GC Gas chromatograph
HAc Acetic acid
HCP High capacity packing
HETP Height Equivalent to Theoretical Plate
H2O Water
I.D Internal diameter (cm)
MCSP Model for catalytic structured packing
MESH Material balance, phase Equilibrium, Summation, Enthalpy balance
MTBE Methyl-tert-butyl ether
NEQ Non equilibrium
NRTL Non-random to liquid activity coefficient model
O.D Outer diameter (cm)
RD Reactive distillation
TAC Total annual cost
TAME Tert-amyl- methyl ether
Unifac UniQUAC functional group activity coefficient model
UniQUAC Universal Quasi – chemical activity coefficient model
VLE Vapor-Liquid Equilibrium
Dedicated
To
The best supervisors for all times
Dr. Khalid Dr. Zaidoon
Chapter One Introduction
1
CHAPTER ONE INTRODUCTION
1.1 Introduction
Chemical manufacturing companies produce materials based on chemical reactions
between selected feed stocks. In many cases the completion of the chemical reactions
is limited by the equilibrium between feed and product. The process must then
include the separation of this equilibrium mixture and recycling of the reactants.
Usually, reaction and separation stages are carried out in discrete equipment units,
and thus equipment and energy costs are added up from these major steps [Doherty
and Buzad 1992, Agar 1999, Kelkar and Ng 1999].
In recent decades, a combination of separation and reaction inside a single unit has
become more and more popular. This combination has been recognized by the
chemical process industries for having favorable economics of carrying out reaction
simultaneously with separation. This type of new technology processes is called
reactive distillation (RD). In RD, reaction and distillation take place within the same
zone of a distillation column. Reactants are converted to products with simultaneous
separation of the products and recycle of unused reactants [Doherty and Buzad 1992,
Agar 1999, Kelkar and Ng 1999].
Reactive or catalytic distillation has captured the imagination of many recently
because of the demonstrated potential for capital productivity improvements (from
enhanced overall rates, by overcoming very low reaction equilibrium constants and
by avoiding or eliminating difficult separations), selectivity improvements (which
reduce excess raw materials use and byproduct formation), reduced energy use, and
the reduction or elimination of solvents. Some of these advantages are realized by
Chapter One Introduction
2
using reaction to improve separation, e.g., overcoming azeotropes or reacting away
contaminants; others are realized by using separation to improve reactions, e.g.,
overcoming reaction equilibrium limitations, improving selectivity, or removing
catalyst poisons [Taylor and Krishna 2000, Kai and Achim 2002].
On the other hand, RD can be also efficient both in size and cost of capital equipment
as well as in energy used to achieve complete conversion of reactants. This advantage
is clearly presented in Eastman’s methyl acetate reactive distillation process (as an
example) as shown in Figure (1.1). In this process, one hybrid reactive distillation
device replaced an entire flowsheet consisting of 11 major units plus all of their heat
exchangers, control systems, pumps, intermediate storage tanks [Kai and Achim
2002].
It is important to mention here that the development, design and operation of RD
processes are highly complex tasks. The potential benefits of this intensified process
come with significant complexity in process development and design. The nonlinear
coupling of reactions, transport phenomena and phase equilibria can give rise to
highly system-dependent features, possibly leading to the presence of reactive
azeotropes and/or the occurrence of steady-state multiplicities [Taylor and Krishna
2000]. Furthermore, the number of design decision variables for such an integrated
unit is much higher than the overall design degrees of freedom of separate reaction
and separation units.
Many investigators studied the production of methyl acetate, ethyl acetate, methyl
tertiary butyl ether (MTBE), ethyl tertiary butyl ether (ETBE), and tert-amyl- methyl
ether (TAME) [Yeong et al. 2003, Anil et al. 2007, Calvar et al. 2007, Brehelin et al.
2007, Firas 2008]. On the other hand, few works are focused on the study of the
esterification of butanol using reactive distillation technology to produce butyl acetate
[Gangadwala et al. 2003, Wang et al. 2003, and Serge et al. 2006].
Chapter One Introduction
3
Fig (1.1) The reduction of full plant to one unit for methyl acetate production using reactive
distillation technology (Eastman Kodak process) [Kai and Achim 2002].
n-Butyl acetate is an important solvent in the chemical industry. Primarily used in
coating and painting processes, it has also been applied in pharmaceutical industries
and cosmetic formulations as an artificial flavor in recent years. In addition, due to its
pleasant fruity (rather like pears) odor, it is also used as a component in synthetic
flavors of fruits. Because of its low toxicity and environmental impact, n-butyl acetate
has become an important replacement to such toxic and teratogenic solvents as ethoxy
ethyl acetate. It is also used as a reaction medium for adhesives [Steinigeweg and
Gmehling 2003, Charubala et al. 2004].
Chapter One Introduction
4
1.2 The Advantage and Disadvantage of Reactive Distillation
The benefits of RD can be summarized as follows [Kai and Achim 2002, Taylor and
Krishna 2000]:
(a) The combination of reaction and separation in one unit shows large savings in
capital cost and energy with less maintenance.
(b) Improved conversion of reactant approaching ~100%: Higher conversions are
obtained for equilibrium-limited reactions due to shifting of the equilibrium
to the right.
(c) Improved selectivity: Removing one of the products from the reaction
mixture or maintaining a low concentration of one of the reagents preventing
them from undergoing further reaction to produce by-products and hence
improved selectivity for the desired products.
(d) Significantly reduced catalyst requirement for the same degree of conversion.
(e) Avoidance of azeotropes: RD is particularly advantageous when the reactor
product is a mixture of species that can form several azeotropes with each
other.
(f) Reduced by-product formation.
(g) Heat integration benefits: If the reaction is exothermic, the heat of reaction
can be used to provide the heat of vaporisation and reduce the reboiler duty.
(h) Avoidance of hot spots and runaways using liquid vaporisation as thermal fly
wheel.
On the other hand, there are several constraints and foreseen difficulties [Kai and
Achim 2002, Taylor and Krishna 2000]:
(a) Volatility constraints: The reagents and products must have suitable volatility
to maintain high concentrations of reactants and low concentrations of products
in the reaction zone.
Chapter One Introduction
5
(b) Residence time requirement: If the residence time for the reaction is long, a
large column size and large tray hold-ups will be needed and it may be more
economic to use a reactor-separator arrangement.
(c) Scale up to large flows: It is difficult to design RD processes for very large
flow rates because of liquid distribution problems in packed RD columns.
1.3 Industrial Applications of Reactive Distillation:
The last decades have seen a significant increase in the number of experimentally
research studies dealing with RD applications [Rameshwar et al. 2004, Kai and
Achim 2002]. The following section shows the main industrial application of the
reactive distillation.
1) Etherification: MTBE, ETBE and TAME Etherification [Taylor and Krishna, 2000
and Sharma and Mahajani, 2003] and Synthesis of alcoxyalkanol[Zhang and Wan,
1991].
2) Hydrolysis: Hydrolysis of methyl acetate [Hoyme and Holcombe, 2002] and
Manufacture of glycine from glycinonitrile [Aoki and Otsubo, 2001].
3) Hydrogenation: Production of cyclopentane or cyclopentene [Silverberg et al.,
2000], Hydro desulfurization [Groten and Loescher, 2002 ,Podrebarac et al. 2001]
and Hydroconversion [Mukherjee and Louie, 2003].
5) Alkylation: Alkylation of Benzene with Ethylene [Netzer, 2001] and Synthesis of
linear alkyl benzene [Knifton et al. 2003].
6) Esterification: Methyl Acetate/Ethyl Acetate [Okur and Bayramoglu (2001), Kenig
et al. (2001), Sharma and Mahajani (2003), Shakor and Sukkar (2008)], Amyl Acetate
[Chiang et al. 2002], Methyl Isopropyl Acetate [Smejkal et al. 2001 and Hanika et al.
2001] and Butyl Acetate [Hanika et al.1999 , Lederer et al. 2002 Gangadwala 2002;
Hiwale et al. 2002; Hiwale, 2003 and Gangadwala et al., 2003].
Chapter One Introduction
6
1.4 Aims of the Present Work
Studying the production of butyl acetate using reactive distillation technology.
Studying the effect of catalyst type, reaction temperature, feed ratio and the
catalyst weight on the reaction conversion and formulate a Pseudo-
Homogenous equation to represent the reaction kinetic.
Developing an unsteady state mathematical model to describe continuous
reactive distillation column.
Evaluating the model by comparing the experimental results with the model
results at the same selected conditions.
Chapter Two Literature Survey
7
CHAPTER TWO
LITERATURE SURVEY
2.1 SCOPE
The aim of this chapter is to give a comprehensive review of literature deal with the
reactive distillation to gain a fundamental understanding of the possible reaction,
internals of reactive distillation and the mathematical models used to analyze the
dynamics or design of reactive distillation. In the present study, the production of
butyl acetate was selected as a case study.
2.2 Fundamental of Reactive Distillation
Reactive distillation is attractive in those systems where certain chemical and phase
equilibrium conditions exist. Because there are many types of reactions, there are
many types of reactive distillation designs depending on reaction system. In this
section the ideal classical situation was described, which will serve to outline the
basics of reactive distillation.
Consider the system in which the chemical reaction involves two reactants (such as A
and B) producing two products (C and D). The reaction takes place in the liquid phase
and is reversible.
(2.1)
For reactive distillation to work, it must be able to remove the products from the
reactants by distillation. This implies that the products should be lighter and / or
heavier than the reactants. In terms of the relative volatilities of the four components,
Chapter Two Literature Survey
8
an ideal case is when one product is the lightest and the other product is the heaviest,
with the reactants being the intermediate boiling components.
Figure (2.1) rpresents the flowsheet of this ideal reactive distillation column. In this
situation the lighter reactant A is fed into the lower section of the reaction section in
the column. The heavier reactant B is fed into the upper section of the reaction section
in the column. The middle of the column is the reactive section and contains NRX trays
on which the net reaction rate of the reversible reaction depends on the forward and
backward specific reaction rates (kf And kB) and the liquid holdup (or amount of
catalyst) on the tray. The vapor flowrates through the reaction section change from
tray to tray because of the heat of the reaction.
Fig.(2.1) Ideal reactive distillation column [Keil 2007].
Chapter Two Literature Survey
9
As component A flows up the column, it reacts with descending B. The very light
product C is quickly removed in the vapor phase from the reaction zone and flows up
the column. Likewise, very heavy product D is quickly removed in the liquid phase
and flows down the column.
The section of the column above where the fresh feed of B is introduced (the
rectifying section with NR trays) separates light product C from all of the heavier
components, so a distillate is produced that is fairly pure product C. The section of the
column below where the fresh feed of A is introduced (the stripping section with NS
trays) separates heavy product D from all of the lighter components, so a bottom is
produced that is fairly pure product D. The reflux flowrate and the reboiler heat duty
can be manipulated to maintain these product purities.
It is important to mention hear that, the column pressure is one of the most important
design parameters for reactive distillation [Kai and Achim 2002].
Pressure effects are much more pronounced in reactive distillation than in
conventional distillation. In normal distillation, the column pressure is selected so that
the separation is made easier (higher relative volatilities). In most systems this
corresponds to low pressure. However, low pressure implies a low reflux-drum
temperature and low-temperature coolant. The typical column pressure is set to give a
reflux-drum temperature high enough to be able to use in expensive cooling water in
the condenser and not require the use of much more expensive refrigeration [Cristhian
et al. 2005].
In reactive distillation, the temperatures in the column affect both the phase
equilibrium and chemical kinetics. A low temperature gives high relative volatilities
Chapter Two Literature Survey
10
may give small specific reaction rates that would require very large liquid holdups (or
amounts of catalyst) to achieve the required conversion. In contrast, a high
temperature may give a very small chemical equilibrium constant (for exothermic
reversible reactions), which makes it more difficult to drive the reaction to produce
products. On the other hand high temperatures may also promote undesirable side
reactions. Thus, selecting the optimum pressure in the reactive distillation column is
very important [Cristhian et al. 2005]
Reactive distillation is also different from conventional distillation in that there are
both product compositions and reaction conversion specifications. The design degrees
of freedom in a reactive distillation column must be adjusted to achieve these
specifications while optimizing some objective function such as total annual cost
(TAC). These design degrees of freedom include pressure, reactive tray holdup,
number of reactive trays , location of reactant feed streams, number of stripping trays,
number of rectifying trays , reflux ratio, and reboiler heat input [Cristhian et al. 2005].
Another design aspect of reactive distillation that is different from conventional is tray
holdup. Holdup has no effect on the steady-state design of a conventional column. It
certainly affects the dynamics but not the steady-state design. Column diameter is
determined from maximum vapor-loading correlations after vapor rates have been
determined that achieve the desired separation. Typical design specifications are the
concentration of the heavy key component in the distillate and the concentration of the
light key component in the bottoms. However, holdup is very important in reactive
distillation because reaction rates directly depend on holdup (or the amount of
catalyst) on each tray. This means that the holdup must be known before the column
can be designed and before the column diameter is known. As a result, the design
procedure for reactive distillation is iterative.
Chapter Two Literature Survey
11
2.3 Esterification Catalysts and Processes:
Esterification reactions proceed with or without a catalyst. In the absence of a catalyst,
the reaction is, however, extremely slow, since its rate depends on the autoprotolysis
of the carboxylic acid. Therefore, esterification is carried out in the presence of an
acid catalyst, which acts as a proton donor to the carboxylic acid. [Ulmann 2001].
There are two types of catalyst used in esterification reaction: Homogenous and
heterogeneous catalyst.
2.3.1 Homogeneous Acid Catalysts:
Catalysis by mineral acids has emerged as a field of growing interest and importance
in the last three decades with respect to new applications and detailed investigations
on catalysis and reaction engineering. Generally homogeneous acid catalysts consist
of inorganic mineral acids and heteropoly acids. Typical examples include sulfuric
acid, hydrochloric acid, arylsulfonic acids such as p-toluenesulfonic acid and
chlorosulfuric acid. Phosphoric acid, polyphosphoric acids and mixtures of acids are
also recommended [Charubala 2004].
Leyes and Othmer (1945) used sulfuric acid catalyst for esterification of acetic acid
and butanol.
Ronnback et al. (1997) studied the esterification kinetics of acetic acid with methanol
in presence of hydroiodic acid.
The disadvantage of mineral acids is their miscibility with the reaction medium
leading to corrosion hazards and separation problems. Hence heterogeneous or
heterogenized acid catalysts provide an attractive alternative to homogeneous catalyst.
Chapter Two Literature Survey
12
2.3.2 Heterogeneous acid catalysis:
1) Metal ion complexes as catalysts:
Metallic oxides and hydroxides of magnesium, zinc, titanium, zirconium and metal
oxide complexes have been used as catalysts for esterification reactions.
Okuhara et al. (1998) observed that solid acid heteropoly acids are water-tolerant
catalysts for various reactions such as hydrolysis of esters, hydration of alkenes and
esterification. Supporting metal on oxides improved stability of these catalysts in
water.
Timofeeva (2003) presented a review on the achievements in the field of acid catalysis
by heteropoly acids. Due to their unique physicochemical properties, heteropoly acids
can be profitably used in homogeneous, biphasic and heterogeneous systems. The
catalytic effect of heteropoly acids in acidic-type reactions depends mainly on three
factors, namely, the acidity, heteropolyanion structure and type of reaction. The
catalytic activity is more dependent on the heteropoly acid structure rather than its
composition [Timofeeva 2003].
2) Zeolites as catalysts:
Zeolites are also used as esterification catalysts. The rare earth exchanged RE H-Y
zeolite is the best of the various zeolites catalysts [Charubala 2004]. The Nb2O5.nH2O
catalyst is claimed to be more active than cation exchange resin, SiO2, Al2O3 and solid
super acids [Charubala 2004]. Nagaraju and Mehboob (1996) compared the catalytic
activity of zeolites of the types of NaX, NaY and NaZSM-5 and their protonated
forms with some conventional Lewis acids such as anhydrous ZnCl2, AlCl3, and
H2SO4 in the esterification reaction between isoamyl alcohol and acetic acid. They
observed that zeolites were more active than the conventional Lewis acid catalyst.
Chapter Two Literature Survey
13
Majid et al. (2008) investigated the effect of the catalyst type on the production of
ethyl acetate via reactive distillation using three type (Zeolit 225, Zeolit 226 and
Ambylite 400) and found that Ambylite 400 give the highest conversion
3) Ion exchange resins as catalysts
Ion exchange resin catalysts have been used for several years in esterification
reactions. Ion exchange materials may be broadly defined as an insoluble matrix
containing labile ions capable of exchanging with ions in the surrounding medium
without major physical change in its structure [Streat 1988]. Typical resin catalysts are
sulphonic acids fixed to a polymer carrier, such as polystyrene crosslinked with di-
vinyl benzene (DVB). Several types of catalysts are commercially available like
Amberlyst resins (e.g. Amberlyst –15, Amberlite IR-120, Dowex-50 WX8, and
Amberlite CG 50etc.).
2.3.3 Effect of Catalysts Type on Reactive Distillation:
In reactive distillation reaction can be autocatalytic, homogeneous or heterogeneous.
In the case of autocatalytic reactions the reaction velocity to be influenced by the
reaction temperature, in other words for reactive distillation by the pressure of the
equipment [Cristhian 2005].
Homogeneous catalysis allows the reaction velocity to be influenced by changing the
catalyst concentration. Thus the reaction velocity can be adapted over a wide range to
the needs of the distillation equipment.
Heterogeneous catalysis requires a construction to fix the catalytic particles in the
reaction zone. This may cause construction and operation problems and is in addition
a limiting factor to the catalyst concentration that can be achieved. The reaction
Chapter Two Literature Survey
14
velocity can be enhanced only to the limit set by the attainable concentration range.
Furthermore the possibility of enhancing the reaction velocity by a higher temperature
or pressure of the equipment is limited, because in general the catalyst consists of ion
exchanger particles, whose temperature range is limited [Cristhian 2005].
So homogeneous catalysis is much more flexible but has its price in an additional
separation step necessary for the catalyst recycle and by demands for expensive
materials in the case of mineral acids. Heterogeneous catalysis was simpler in
principle, but technical problems have to be solved. In general the equipment will
need more volume, for example the columns must have a bigger diameter. It should be
clear from these considerations that a single case decision is needed for every
individual design [Cristhian 2005].
In addition, the type of the catalysis is important. Homogeneous catalysis are possible
in most cases but need a separation step to purify and recycle the catalyst. This can be
avoided in heterogeneous catalysis, but here special constructions are necessary to fix
the catalyst in the reaction zone.
2.4 Thermodynamics of Reactive Separations:
Thermodynamics plays a key role in understanding of reactive separation process. The
fact that reaction and separation occur simultaneously gives rise to special challenges
both in experimental investigation and modeling the processes. There are several
contributions of thermodynamics to the field of reactive separations. Thermodynamics
provides the basic relations, such as energy balances of equilibrium condition, used in
the process models, and (models and experimental methods) for the investigation of
properties of the reacting fluid that have to be known [Kai and Achim 2002,Amado et
al. 2008].
Chapter Two Literature Survey
15
2.4.1 Ideal Solution:
In ideal solutions all molecules are the same size and all forces between molecules
(like and unlike) are equal. The ideal gas, consisting of molecules with zero volume
that do not interact, fulfills the condition of solution ideality as a special case. When
ideal gases are mixed, there is no volume change of mixing, because the molar volume
of mixture igV and molar volume of the pure species ig
iV are all equal to P
RT [Smith
2001].
2.4.2 Non Ideal Solutions:
When a liquid contains dissimilar polar species, particularly those that can form or
break hydrogen bonds, the ideal liquid solution assumption is almost always invalid
and the regular solution theory is not applicable. Non ideal solution effects can be
incorporated into K-value formulations, therefore VLE calculations are carried out by
using the activity coefficients for the liquid which are calculated to correct the
equilibrium constant [Seader and Ernest 2006]. At present there are at least four
different types of correlation for the predication of activity coefficients in chemical
systems that are normally used: Wilson, NRTL, UNIQUAC and UNIFAC
[Mandagaran et al. 2006].
2.4.3 Heat of Reaction:
The heat liberated or absorbed during reaction depends on the nature of the reacting
system, the amount of material reacting, and the temperature and pressure of the
reacting system, and is calculated from the heat of reaction Hr . When this is not
known, it is in most cases calculated from known and tabulated thermo chemical data
on heat of formation Hf or heat of combustion Hc of the reacting materials [Majid
et al.2008].
Chapter Two Literature Survey
16
2.4.4 Pressure Drop:
Structured packings have been established in the field of distillation. They have
advantages compared to other distillation column internals, such as high separation
performance or low pressure drop [Miller and Kaibel 2004].
Kreul et al. (1998) studied pressure drop in the packed reactive distillation column
majority pressure drop in distillation models it is neglected and only an overall
column pressure is studied. In dynamic systems the consideration of pressure drop in
the form of a correlation of the vapor (and, if necessary liquid) load is taken into their
studies.
Peter and Hans (1999) reported pressure drop on Katapak-S is a structured catalytic
packing for reactive distillation. They concluded that the results of experimental
studies are in good agreement with the theoretical studies.
Behrens et al. (2006) predicted and developed model for the pressure drop on catalytic
structured packings (MCSP). The open channels in the MCSP exhibits the normal
structure as encountered in high capacity packing (HCP); therefore these channels are
treated similarity to the HCP. In these channels the pressure drop is determined by
three contributions. Gas–gas interaction in the crossing flow channels, gas–liquid
interaction at the interface along the channel, and direction change related losses, the
experimental results up to flooding show good agreement with the model predictions.
2.5 Reactive Distillation Internal
2.5.1 Homogeneous Reactive Distillation:
For homogeneous RD processes, counter-current vapor-liquid contacting, with
sufficient degree of staging in the vapor and liquid-phases, can be achieved in a multi-
Chapter Two Literature Survey
17
tray column as shown in Figure (2.2) or a column with random or structured packings
as shown in Figure (2.3). The froth regime is usually to be preferred on the trays as
shown in Figure (2.4) [Kai and Achim 2002, Taylor and Krishna 2000] because of the
desire to maintain high liquid hold-up on the trays. High liquid hold-ups could be
realized by use of bubble caps, reverse flow trays with additional sumps to provide
ample tray residence time. In the Eastman process for methyl acetate manufacture
specially designed high liquid hold-up trays are used [Agreda et al. 1990].
Fig. (2.2) Counter-current vapor- liquid contacting in homogenous trayed reactive distillation
columns[Kai and Achim 2002, Taylor and Krishna 2000 ].
Chapter Two Literature Survey
18
Fig. (2.3) Counter-current vapor-liquid contacting in homogenous packed reactive distillation
columns [Kai and Achim 2002, Taylor and Krishna 2000].
Fig. (2.4) Flow regimes on trays[Kai and Achim 2002, Taylor and Krishna 2000].
2.5.2 Heterogeneous Reactive Distillation:
2.5.2.1 Packed Reactive Distillation:
For heterogeneously packed RD, hardwired-sign poses considerable challenges. The
catalyst particle sizes used in such operations are usually in the 1-3 mm range. Larger
particle sizes lead to intra-particle diffusion limitations. To overcome the limitations
of flooding the catalyst particles have to be enveloped within wire gauze envelopes.
Chapter Two Literature Survey
19
Most commonly the catalyst envelopes are packed inside the column. Almost every
conceivable shape of these catalyst envelopes has been patented; some basic shapes
are shown in Figurs (2.5-2.9).These structures are:
1. Porous spheres filled with catalyst inside them as shown in Figure (2.5a).
2. Cylindrical shaped envelopes with catalyst inside them as shown in Figure (2.5b).
3. Wire gauze envelopes with various shapes: spheres, tablets, doughnuts, etc. as
shown in Figure (2.5c).
4. Horizontally disposed wire-mesh “gutters”, filled with catalyst as shown in Figure
(2.6a).
5. Horizontally disposed wire-mesh tubes containing catalyst as shown in Figure
(2.6b).
6. Catalyst particles enclosed in cloth wrapped in the form of bales this is the
configuration used by Chemical Research and licensing in their RD technology for
etherification, hydrogenation and alkylation of aromatic compounds [Shoemaker &
Jones, 1987]. The catalyst is held together by fiberglass cloth. Pockets are sewn into
a folded cloth and then solid catalyst is loaded into the pockets. The pockets are
sewn shut after loading the catalyst and the resulting belt or “catalyst quilt” is rolled
with alternating layers of steel mesh to form a cylinder of “catalyst bales as shown
in Figure (2.7). The steel mesh creates void volume to allow for vapor traffic and
vapor/liquid contacting. Scores of these bales are installed in the reactive zone of a
typical commercial RD column. Bales are piled on top of each other to give the
required height necessary to achieve the desired extent of reaction. When the
catalyst is spent the column is shut down and the bales are manually removed and
replaced with bales containing fresh catalyst. Improvements to the catalyst bale
concept have been made over the years [Kai and Achim 2002, Taylor and Krishna
2000].
Chapter Two Literature Survey
20
7. Catalyst particles sandwiched between corrugated sheets of wire gauze as shown in
Figure (2.8). Such structures are being licensed by Sulzer (called KATAPAK-S)
and Koch-Glitsch (called KATAMAX). They consist of two pieces of rectangular
crimped wire gauze sealed around the edge, thereby forming a pocket of the order
of 1-5cm wide between the two screens.
8. Another alternative is to make the packing itself catalytically active. Where in the
raschig ring-shaped packing are made catalytically active as shown in Figure
(2.9a). Their activity is quite high; however, osmotic swelling processes can cause
breakage by producing large mechanical stresses inside the resin. An alternative
configuration is the glass-supported precipitated polymer prepared by precipitation
of styrene-divinylbenzene copolymer, which is subsequently activated by
chlorsulphonic acid. Another possibility is to coat structured packing with zeolite
catalysts [Oudshoorn, 1999]as shown in Figure (2.9b).
The catalyst can also be “cast” into a monolith form and used for counter-current
vapor-liquid contacting; Lebens (1999) has developed a monolith construction
consisting of fluted tubes as shown in Figure (2.9c).
2.5.2.2. Trayed Reactive Distillation:
The catalyst envelopes can be placed in a trayed RD column and many configurations
have been proposed.
1. Vertically disposed catalyst containing envelopes can be placed along the direction
of the liquid flow path across a tray as shown in Figure (2.10).
2. Catalyst envelopes can be placed within the down comers as shown in Figure
(2.11a).
3. Catalyst envelopes can be placed near the exit of the downcomer as shown in
Figure (2.11b).
Chapter Two Literature Survey
21
4. Trays and packed catalyst sections can also be used on alternate stages as shown in
Figure (2.11c).
5. Other designs have been proposed for tray columns with catalyst containing pockets
or regions that are fluidized by the up flowing liquid.
Fig. (2.5) Various “tea-bag” configurations. Catalyst particles need to be enveloped in wire
gauze packings and placed inside RD columns [Kai and Achim 2002, Taylor and Krishna
2000].
Fig. (2.6) Horizontally disposed (a) wire gauze gutters and (b) wire gauze tubes containing
catalyst[Kai and Achim 2002, Taylor and Krishna 2000].
Chapter Two Literature Survey
22
Fig (2.7) Catalyst bales licensed by Chemical Research and Licensing[Kai and Achim 2002,
Taylor and Krishna 2000].
Fig (2.8) Structured catalyst-sandwiches. (a) Catalyst sandwiched between two corrugated
wire gauze sheets. (b) The wire gauze sheets are joined together and sewn on all four sides. (c)
The sandwich elements arranged into a cubical collection. (d) The sandwich elements arranged
in a round collection [Kai and Achim 2002, Taylor and Krishna 2000].
Chapter Two Literature Survey
23
Fig (2.9) (a) Catalytically active Raschig ring. Adapted from Sundmacher (1995). (b)
Structured packings coated with catalyst. (c) Fluted catalyst monolith tubes.
Fig.(2.10) Catalyst envelopes placed along the liquid flow path[Kai and Achim 2002, Taylor
and Krishna 2000].
Chapter Two Literature Survey
24
Fig.(2.11) Counter-current vapor-liquid-catalyst contacting in trayed columns. (a) catalyst in
envelopes inside downcomers (b) tray contacting with catalyst placed in wire gauze
envelopes near the liquid exit from the downcomers. (C) Alternating packed layers of
catalyst and trays [Kai and Achim 2002, Taylor and Krishna 2000].
2.6 Production of Butyl acetate as Case Study
2.6.1 Reaction Kinetic:
Synthesis of butyl acetate is commonly carried out by esterification of n-butanol
(BuOH) with acetic acid (HAc) in the presence of a suitable acid catalyst. The
reaction is reversible in nature making it difficult to produce required product purity,
despite the equilibrium constant favoring the production of butyl acetate.
Simultaneous removal of product(s) during the course of the reaction is beneficial to
obtain enhanced conversion. For this purpose various methods have been adopted but
the most important one is reactive distillation
Equation (2.3) show the chemical reaction of n-butyl acetate production. This type of
reaction is kinetically controlled which is catalyzed by free protons [Ulmann, 2001].
Chapter Two Literature Survey
25
(2.3)
In the reaction mixture there are four binary azeotropes and two ternary azeotropes.
The binary azeotropes are water–butanol, water–butyl acetate, butyl acetate –butanol
and acetic acid –butanol and the ternary azeotrope are butanol–butyl acetate–water
and butanol–butyl acetate– acetic acid [Silke et al. 2000].
The azeotropes boiling points and compositions are shown in Table (2.1).
Table (2.1) Singular points in the butyl acetate system [Silke et al. 2000].
Name Composition T(Co) xBuOH xHAc xBuAc
Azeotrope 1 BuOH/BuAc/H2O 90.66 0.0895 0 0.2041
Azeotrope 2 BuAc/H2O 91.19 0 0 0.2766
Azeotrope 3 BuOH/H2O 92.96 0.2334 0 0
Water H2O 100 0 0 0
Azeotrope 4 BuOH/BuAc 116.15 0.7004 0 0.2996
n- butanol HAc 117.77 0 1 0
Acetic acid BuOH 117.97 1 0 0
Azeotrope 5 BuOH/BuAc/ HAc 121.3 0.2 0.46 0.34
Azeotrope 6 BuOH / HAc 122.69 0.5161 0.4839 0
n-butyl acetate BuAc 126.17 0 0 1
2.6.2 Production of n-Butyl acetate (Previous work):
Leyes and Othmer (1945) investigated the sulfuric acid catalyzed esterification of
acetic acid with butanol. They found that the reaction to be second order with respect
to acetic acid concentration up to 75 to 85 %, in a temperature range of 373 K-393 K.
The rate was a linear function of the catalyst concentration and the molar ratio of
Chapter Two Literature Survey
26
butanol to acetic acid. They also found that the logarithm of the rate constant was
proportional to the reciprocal of absolute temperature as per Arrhenius law of
temperature dependence.
It is important to mention that the solid heterogeneous catalysts are receiving attention
because of their obvious engineering benefits such as ease of separation and fewer
disposal and corrosion problems. Cation-exchange resin is considered to be an
important catalyst for liquid-phase reactions like esterification, etherification, etc.
[Chakrabarty et al. 1993].
Li et al. (1996) studied various zeolites catalysts, such as HX, HY, HM, and HZSM5
for esterification of butanol with acetic acid. They found that HZSM5 acts as the best
catalyst from all zeolites studied.
Liao and zhange (1997) have studied the kinetics of liquid-phase esterification of
acetic acid with butanol by using ion-exchange resin as catalysts. The experimental
results showed that the reaction is an apparent first-order reaction, with the apparent
rate constant 3.5× 10-2
min-1
.
Janowsky et al. (1997) studied the kinetics in the presence of Lewatit SPC 108 and
118 catalysts and proposed a pseudo-homogeneous (PH) kinetic model for the
esterification reaction. They proposed the expression to represent the following
equilibrium constant.
(2.4)
Chapter Two Literature Survey
27
Altiokka and Citak (2003) studied the kinetics of homogeneous and ion exchange
resin (IR-120) catalyzed for esterification of acetic acid with butanol. They reported
that presence of resin catalyst reduces the activation energy from 59.3 to 49 kJ/mol.
Blagov et al. (2006) studied the synthesis of n-butyl acetate by comparing three ion-
exchange resin catalysts (Purolite CT 269, Amberlyst 46, and Amberlyst 48). They
found that the three catalysts show only minor differences in their activity.
Bozek and Gmehling (2006) investigated the reaction kinetics and chemical
equilibrium of transesterification of methyl acetate and n-butanol to n-butyl acetate
and methanol in the temperature range of (40 to 57 oC) using acidic ion-exchange
resin, Amberlyst 15. They concluded that the chemical equilibrium constant obtained
from kinetic experiments was in qualitative agreement with the calculated standard
thermodynamic properties. Also they studied the influence of the catalyst loading,
initial reactant molar ratio, and temperature on the kinetics.
Izci et al (2009) studied the kinetics of esterification of acetic acid with isobutanol
using Amberlite IR-122. Experiments were carried out in a stirred batch reactor at
different temperatures (50 to 75 oC) under atmospheric pressure. They found that the
equilibrium constant is equal to 4 in the temperature range. The possible mechanism
of reaction is mathematically treated using the theories of the Eley-Rideal model. The
reaction rate constants and the adsorption coefficients for isobutanol and water were
determined from the experimental data at the same temperature intervals.
On the other hand, very few papers studied the experimental production of n-butyl
acetate in a continuous mode via Reactive Distillation.
Chapter Two Literature Survey
28
Hanika et al. (1999) used a column (inner diameter (ID) of 81 mm) that consisted of a
catalytic zone packed with Katapak-S packing and two separation zones that were
equivalent to 20 theoretical stages. Two different configurations of catalytic
distillation have been considered. In the first set, only a catalytic distillation column
was used; in the second set, a primary fixed-bed reactor that was packed with ion-
exchange resin, followed by a catalytic distillation column, was studied. A mixture of
acetic acid and butanol (in excess) was preheated and either fed into the catalytic zone
of the column or into the pre-reactor. The output from the prereactor, containing an
almost-equilibrium mixture of acetic acid, butanol, butyl acetate, and water, was
preheated almost to its boiling point and fed into the catalytic zone of the RD column.
Janowsky et al. (1997) performed butyl acetate synthesis experiments to study steady-
state column performance at three different pressures over a range of 0.65-1.105 bar.
They used a packed column in their experiment, and each section of the column was
equivalent to 15 theoretical stages. The stripping section, which was filled with
catalyst, acts as a reactive section, and feed that comes from a pre-reactor, containing
a slight excess of butyl acetate, was introduced at the top of the reactive section. At
higher pressure, they observed a significant amount of 1-butene at the top of the
column. Also, the unwanted byproduct di-butyl ether (DBE) was observed, with the
main product (butyl acetate) in the bottoms. They were able to eliminate the formation
of 1-butene by decreasing the column pressure up to 0.65 bar; however, they were
unable to eliminate DBE from the bottoms.
Excellent work on the synthesis of butyl acetate has been reported by Steinigeweg and
Gmehling. (2003) They studied the thermodynamic properties, reaction kinetics, and
RD system through experiments and simulation. Katapak-S (Sulzer ChemTech) filled
with strongly acidic-ion-exchange resin (Amberlyst-15) was used as a catalyst. The
Chapter Two Literature Survey
29
activity coefficients for the liquid phase were calculated by the UNIQUAC equation.
A pseudo-homogeneous activity-based model was used to describe the rate equation.
The experiments were performed in a continuous RD column by introducing a fresh
mixture of acetic acid and butanol as a feed. The effect of different parameters, such
as reboiler duty, feed location, composition, molar ratio, and pressure, was studied.
Experiments have shown that the most suitable feed location was the top of the
catalyst bed and, by increasing number of reactive stages, conversion increases.
A maximum conversion of 98% was realized. All the simulations were performed
using a steady-state simulator (Aspen Plus). Comparison of the experimental data with
simulation results indicated that an equilibrium stage model is capable of describing
the column profiles quantitatively. The same model and simulator were used further to
predict the performance of the column with feed that contains butyl acetate and water.
They suggested that a pre-reactor, followed by an RD column, is the best process
alternative. The experiments with a four-component feed to the RD column were not
performed.
It is important to mention here that the most important side reactions in the production
of esters are caused by dehydration of alcohols leading to formation of alkenes and by
self-condensation of alcohols leading to ethers [Sergej et al. 2006]. Whereas in
conventional esterification processes, side reactions usually do not play a major role,
they are important in heterogeneously catalyzed reactive distillation.
The alkene is always the lightest boiling component and is, hence, quickly removed
from the reacting liquid, so that its formation is enhanced by distillation. This may
become crucial not only with respect to the selectivity of the process, but also for the
process stability as accumulation of alkenes in the column may lead to qualitative
changes of process behavior .The ether , as well as the product ester , are usually
Chapter Two Literature Survey
30
heavy boiling components and, therefore, found in the bottom of the reactive
distillation column. Hence, the ether is an impurity in the final product and can only
be tolerated with in specified limits. Therefore, the mitigation of the side reactions is a
key issue in the design of improved esterification processes by reactive distillation.
The side reactions of the n-BuAc formation can generally not be studied together with
the main reaction , as they proceed at much lower speed , so that only very little
amounts of side products are formed in typical reaction kinetic studies of the main
reaction. This explains why almost no quantitative data on the side reactions of
interest are found in the literature.
Gangadwala et al. (2003) studied the kinetics of the etherification of nBuOH in the
presence of several acidic ion-exchange resins including Amberlyst 15. They found
that the etherification practically does not occur at low temperatures (60–90◦C) in
conventional batch reaction experiments with a charge of HAc and n-BuOH. They
also performed separate kinetic experiments with pure n-BuOH as charge at higher
temperatures (113–126◦C) and high catalyst loading in the range 7-17 g of
catalyst/gmol of n-BuOH which allowed to develop a kinetic model of the
etherification based on the Langmuir–Hinshelwood–Hougen–Watsonapproach.
Sergej et al. (2006) studied the side reaction kinetics of the heterogeneously catalyzed
esterification of n-BuOH with acetic acid in an isothermal fixed bed loop reactor at
temperatures between 100 and 120◦C using three different ion-exchange catalysts.
They observed di-n-butylether, sec.-butyl-n-butyl ether, sec.-butanol and sec.-butyl
acetate in these experiments. On the other hand, they concluded that surface-
sulfonatedion-exchange catalysts are extraordinarily attractive for the production of n-
BuAc by reactive distillation.
Chapter Two Literature Survey
31
Ajay et al. (2005) studying the production of butyl acetate in a continuous catalytic
distillation system and that it is feasible to obtain high-purity butyl acetate in a
reactive distillation (RD) column with almost-quantitative conversion. So they believe
that the formation of di-butyl ether (DBE) in RD is due to the large amount of n-
butanol present in the reactive zone, compared to acetic acid. This is unlikely to
happen in conventional fixed-bed or slurry reactors. So they recommended that one
should use butanol as a limiting reactant, to avoid the formation of DBE as a side
product. However, a large excess of acetic acid in the feed may also be undesired,
because one would either get impure butyl acetate as the bottom product or lose acetic
acid from the top of the column through an aqueous layer. The best position for the
introduction of the feed is at the top of the reactive zone.
2.7 Modeling of Reactive Distillation
Two primary approaches are available in the literature for modeling reactive
distillation columns
2.7.1 Equilibrium Model:
A schematic diagram of an equilibrium stage is shown in Figure (2.12a) .Vapor from
the stage below and liquid from the stage above are brought into contact on the stage
together with any fresh or recycle feeds. The vapor and liquid streams leaving the
stage are assumed to be in equilibrium with each other. A complete separation process
is modeled as a sequence of these equilibrium stages in Figure (2.12b).
The equations that model equilibrium stages are known as the MESH equations,
MESH being an acronym referring to the different types of equation.
Chapter Two Literature Survey
32
The M equations are the material balance equations; the total material balance takes
the form
,
(2.5) Mj is the hold-up on stage j. With very few exceptions, M j is considered to be the
hold-up only of the liquid-phase. It is more important to include the hold-up of the
vapor phase at higher pressures. The component material balance (neglecting the
vapor hold-up) is
(2.6)
In the material balance equations given above υi,m represents the
stoichiometric coefficient of component i in reaction m and represents the reaction
volume.
The E equations are the phase equilibrium relations
(2.7)
The S equations are the summation equations
(2.8)
(2.9)
The H is the enthalpy balance given by
(2.10)
Chapter Two Literature Survey
33
Fig (2.12) (a) The equilibrium stage. (b) Multi-stage distillation column [Taylor and Krishna
2000].
The enthalpy in the time derivative on the left-hand side represents the total enthalpy
of the stage but, for the reasons given above; this will normally be the liquid-phase
enthalpy.
Under steady-state conditions all of the time derivatives in the above equations are
equal to zero.
Davies et la. (1979) described a variation on the standard EQ stage model that is
depicted in Fig. (2.13). The vapor-liquid contacting section is modelled as a
conventional vapor-liquid equilibrium stage (without reaction). The outgoing liquid
stream passes to a reactor where chemical equilibrium is established. The stream
leaving this reactor passes on to the next equilibrium stage. The disadvantage of this
approach is that it fails to properly account for the influence that chemical equilibrium
Chapter Two Literature Survey
34
has on vapor-liquid equilibrium (and vice versa).The model is used to predict the
temperature and composition profiles in a 76mm diameter column in which
formaldehyde is reacting with water and methanol. Good agreement between
predicted and measured values is claimed, but the figures provided in their paper are
small and hard to read.
Barbosa and Doherty (1988) pointed out that the EQ stage model equations (including
those that account for simultaneous phase and chemical equilibrium) can be rewritten
so that they are identical in form to the EQ model equations in the absence of
chemical reactions. The advantage of this approach is that existing algorithms and
programs can be used to solve the equations. All that is required is to replace that part
of the program that carries out the phase equilibrium calculations with a new
procedure that computes the phase and chemical equilibrium computation and
evaluates the transformed variables.
Fig. (2.13). Equilibrium stage model used by Davies et al. (1979).
Chapter Two Literature Survey
35
2.7.2 Rate Based Model:
The basic idea of the NEQ cell model is shown in Figure (2.14). Each stage is divided
into a number of contacting cells; these cells describe just a small section of a single
tray. The vapor entering a stage is divided into cells m in total, in the first horizontal
row. The liquid entering the stage is, similarly, divided into cells n in total, in the first
vertical column. Also the liquid flow is divided equally into cells in a vertical column.
Any feed entering the stage is also apportioned to the entering row, or column, of cells
in the same manner. By choosing an appropriate number of cells in each flow
direction, one can model the actual flow patterns on a tray. A column of cells can
model plug flow in the vapor phase, and multiple columns of cells can model plug
flow in the liquid phase. When the number of well-mixed cells in any flow direction is
four or more, we have essentially plug flow of that phase. Various degrees of
backmixing in the vapor and liquid phases can be modeled by choosing the number of
well-mixed cells to lie between 1 and 4. Correlations are available in the literature to
estimate the number of well-mixed cells in the liquid flow direction [Bennett and
Grimm 1991]. The staging in the liquid phase is strongly dependent on the column
diameter. Liquid phase staging is in particular important for large-diameter columns.
The assumption of plug flow for the vapour phase is a good approximation and
therefore a choice of 4 cells in the vertical direction is able to deal with this situation.
Further details of the implementation of the cell model can be found in Higler, Taylor
and Krishna (1998) and Higler, Krishna and Taylor (1999) who have developed a
steady-state version for RD columns.
Chapter Two Literature Survey
36
Fig.(2.14) (a)Schematic representation of an NEQ cell model for a stage j.(b)Balance relations
for a representative cell.(c) Composition and temperature profiles within the vapor and liquid
“ films” [Taylor and Krishna 2000].
Chapter Three Experimental Work
37
CHAPTER THREE
EXPERIMENTAL WORK
3.1 Introduction
This chapter describes the experimental equipments and the procedure that used in the
studying of the reaction kinetic and column dynamic of n-butyl acetate production.
The production process was carried out continuously using packed reactive
distillation technology.
In the first part of the experimental work, the kinetic of the reaction was studied in a
batch reactor using two types of ion exchange resins (Dowex-50 WX8 and Amberlite
CG 50). The second part of experimental work was carried out in a pilot plant
(continuous packed reactive distillation) specialized to study the effect of some
parameters (such as: acetic acid flow rate, heat duty and reflux ratio) on the
performance and temperature distribution along the reactive distillation column.
3.2 Materials and Analysis
In the present investigation, many chemicals were used in the experiential work:
Acetic acid (99.8%) was supplied by Rioel-de Haën chemicals Germany, n-butanol
was manufactured by Gainland Chemical Company U.K. of analytical grade (99 %).
On the other hand, two types of solid catalysts (resins) were used Dowex-50WX8, of
size (16-40 mesh), supplied by fluka, which is a strong acid ion-exchange resin, and
Amberlite CG 50, of size (100-200 mesh) supplied by Hopkin &Williams.
To remove impurities, prior to use, the catalyst was washed several times with
distilled water until the supernatant liquid was colorless. The catalyst was then dried
at 60oC for six hours to remove the moisture. The dried catalysts were stored in
desiccator for further use.
Chapter Three Experimental Work
38
3.3 Catalysts Modification
The two types of solid catalysts (heterogeneous catalysts) were modified by treatment
with (1N) HCl to increase acidity of catalysts. The modification procedure was
carried out as following [Firas 2008]:
1-Each 100 g of catalyst is stirred for 3 hrs with 500 ml hydrochloric acid solution
(1N) at 25oC.
2-The modified catalyst is filtered and washed several times with distilled water to be
free of chloride ions.
3-Then, the modified catalyst is dried at 60oC for 6 hours.
For the each of the two types of catalysts were used in present investigation, we
investigate the original type and modified type.
3.4 Experiments of Kinetic Study
Figures (3.1) and (3.2) show the view and the schematic diagram of the experimental
set-up respectively. The reactor consisted of a two-necked Pyrex flask of 500 ml
capacity fitted with a reflux condenser and sampling device. The flask (reactor) was
placed in the oil bath on/off controller was used to maintain a constant temperature
(within ±1 oC) inside the reactor. The reflux condenser was connected with a chiller
to avoid any loss of volatile compounds. The reaction mixture was magnetically
stirred at about 1200 rpm. Liquid samples of (2cm3) were taken using a syringe
through a porous filter, to avoid catalyst lost.
Chapter Three Experimental Work
39
Fig (3.1) View of experimental set-up for measuring of the reaction rate and the reaction
equilibrium.
Fig (3.2) Schematic diagram of the kinetic study experimental apparatus.
Chapter Three Experimental Work
40
3.4.1 Experimental Procedure in Kinetic unit:
In a typical run, the acetic acid and n-butanol were charged into the reactor and heated
to desired temperature. Once the desired reaction temperature was attained, the
catalyst was charged to the reactor, and this time was considered as a zero reaction
time. The samples were taken periodically for two hours.
To prevent evaporation losses and stop any further reaction, all samples taken were
directly cooled and stored in a refrigerator at around 5 o C until they analyses with gas
chromatograph.
The operating conditions of the kinetic study are given in Table (3.1).
Table (3.1): Set of the kinetic study experiments using Dowex-50 and Amberlite CG catalysts
[Gangadwala et al. 2003].
Exp. No.
Catalyst Type
Treated
with
HCl
T(oC)
Feed
Ratio
Catalyst
weight (gm)
Initial
molar
holdup
(mol)
1 Dowex-50 yes 90 1/1 24 5
2 Dowex-50 yes 80 1/1 24 5
3 Dowex-50 yes 70 1/1 24 5
4 Dowex-50 No 90 1/1 24 5
5 Dowex-50 yes 90 1/1 12 5
6 Dowex-50 yes 90 1/1 36 5
7 AmberliteCG yes 90 1/1 24 5
8 AmberliteCG No 90 1/1 24 5
9 Dowex-50 yes 90 1/2 24 7.5
10 Dowex-50 yes 90 2/1 24 3.75
Chapter Three Experimental Work
41
3.5 Production of Butyl Acetate Using Continuous Reactive Distillation
The continuous reactive distillation column was constructed to study the effects of
reflux ratio and heat duty on the performance of heterogeneous catalyst (the best type
of catalysts which is modified Dowex-50) that was selected from experiment in the
kinetic study. Figures (3.3) and (3.4) show a view and the schematic diagram of pilot
plant of continuous packed reactive distillation unit respectively.
The reactive distillation column made of QVF- glass consists of three sections:
stripping, reactive and rectifying sections. A column of O.D=4.7cm, I.D=4.4 cm and
total height with the reboiler and the condenser of 330 cm. The bottom of the column
was connected to the reboiler which consisted of a 2000cm3 capacity round flask
heated by mantel heater connected to voltage controller to control the heat duty. The
top condenser was connected with a chiller to avoid any loss of volatile compounds.
The reactive section (45 cm height) was filled with the selected solid catalysts
(Dowex-50) in the form of bales contain the catalyst particles. The bales were
prepared in the lab as shown in Figure (3.5) to reduce pressure drop in reactive
distillation, the preparing method described in section 2.7.2.1. On the other hand, the
stripping section of 65 cm height and the rectifying section of 90 cm heights packed
with glass rushing ring (I.D=6 mm, O.D=8 mm,10 mm height ). It is important to
mention here that the whole column is thermally insulated with a thick layer of glass
wool to prevent any heat loss from the unit.
Chapter Three Experimental Work
42
Fig (3.3) View of experimental continuous reactive distillation unit.
Chapter Three Experimental Work
43
Fig (3.4) Schematic diagram of a packed continuous reactive distillation unit.
Chapter Three Experimental Work
44
Fig (3.5) Balles sheet contain the catalyst particles prepared in the laboratory.
The operation of pilot plant unit was carried out continuously through pumping the
acetic acid and n-butyl alcohol to the reactive section using two dosing pumps. The
dosing pumps (ALLDOS Germany) were calibrated and adjusted to the desirable flow
rate before the start of each experiment.
The interface system (Computerized Temperature Measurement System) has eight
thermocouples sensors type-T (Copper-Constantan) was used to measure the
temperature distributed along the reactive distillation unit as shown in Figure (3.4).
Chapter Three Experimental Work
45
3.5.1 Experimental Procedure in RD Unit:
In a typical run, the reboiler is charged with 10 mole n-butanol, 10mole acetic acid
before the start of the operation. The acetic acid is the highest boiling point
component, so it is fed above the catalytic packing of the column continuously and
the n-butanol which is of a lower boiling point fed in the lower part of the reactive
section. The column is operated under total reflux conditions for 30 min. Then the two
feeds pumps are turned on and the valve that controls the reflux is opened. The
bottom product pumped periodically with a constant amount using a syringe. Samples
from the top and bottom product are taken every 30 min for the GC analysis. The
operating condition of each experiment is given in Table (3.2).
Table (3.2): Reactive distillation experiment set.
Exp. No. Catalyst
weight(g) Feed ratio Reflux ratio
Heat duty
(watt)
1 100 1:1 1 440
2 100 1:1 2 440
3 100 1:1 4 440
4 100 1:1 2 380
5 100 2:1 2 440
3.6 Sample Analysis
Analysis of the samples were taken throughout the experimental runs were carried by
the gas chromatograph type Shimadzu GC 2014 as shown in Figure (3.6). This unit
is equipped with a Porapack-Q capillary column, with a total length of 30 m and an
inside diameter of 0.35 mm. On the detection side, the GC has a Flame Ionization
Chapter Three Experimental Work
46
Detector (FID), which is suitable for detection of organic compounds and usually
gives better results (using N2 as carrier gas at 50.0KPa; split ratio 50; temperature
program 110 oC hold for 1min, heat at 10
oC/min to 140
oC hold for 1 min). The gas
chromatograph was calibrated using different concentration of the injection material;
Appendix C shows calibration curves and equation.
Figure (3.6): Gas chromatographic analysis.
3.7 Thermocouple Calibration
The thermocouples type T (Copper-Constantan) was calibrated before being used to
measure the temperature of the experiments. Therefore, each thermocouple was
Chapter Three Experimental Work
47
immersed in the constant temperature bath (consisted of a glass beaker filled with oil
and stirred with a magnetic stirrer). The temperature in the bath was measured using
thermometer. At each temperature measured by the thermometer, the interface system
was used to measure the temperature three times to increase the accuracy and using
the average of the readings was taken.
Figure (3.7) describes the calibration procedure and Appendix C show calibration
curves and equation.
Fig (3.7) Schematic diagram of thermocouple calibration process.
Chapter Four Modeling and Simulation
48
CHAPTER FOUR
Mathematical Representation, Modeling and Simulation
4.1 Introduction
Mathematical modeling of any chemical process is a set of equations including the
necessary input data to solve the equations, whose solution gives a specified data
representative of the process to a corresponding set input that allows us to predict the
behavior of chemical process system [Cristhian 2005].
In recent years, process simulation has been used extensively to analyze the dynamic
chemical process or design controllers and study their effectiveness in controlling the
process. The simulation operations make it possible to evaluate the influence of
variables on any process theoretically. Dynamic simulation involves the solution of
ordinary differential equations. Also by comparing the experimental results with
simulation results, one can decide if it is necessary to develop a more detailed model
or it is possible to introduce simplifying assumptions to the model. The simulation is
also used to fix the experimental conditions needed for design, optimization, and
control [Kai and Achim 2002, Taylor and Krishna 2000].
The operation of reactive distillation leads to complex interactions between vapor–
liquid equilibrium, mass rates and chemical kinetics. Therefore, most of the chemical
reactions carried out in RD column are kinetically controlled, for this reaction kinetics
have a significant influence on RD process performance. Therefore, in the first part of
this chapter, the reaction kinetics of n-butyl acetate formation from n-butanol and
acetic acid reaction on acidic solid catalyst named Dowex 50 was studied.
In the second part, unsteady state model was developed to simulate continuous
reactive distillation columns for n-butyl acetate production.
Chapter Four Modeling and Simulation
49
The set of ordinary differential equations representing components mass balance
integrated numerically to evaluate the compositions within the column as a function of
time, while the set of total mass and heat balance equations converted to a linear
equations with the aid of finite difference approximation and then solved to evaluate
the vapor and liquid flow rates by using Gausian Elimination Method. The trays
temperatures were determined by using Newton Raphsin's Method.
A computer program written by sofMATLAB environment (version 7) is used to
perform all the calculation within this chapter.
4.2 Parameters Estimation for the Reaction Kinetics
The esterification of the acetic acid (HAc) with n-butanol (n-BuOH) to produce n-
butyl acetate (BuAc) and water (H2O) was taken into consideration as case study. This
reversible reaction is represented by the equation:
(2.3)
Reaction kinetics of the n-BuAc synthesis is usually described by second order
Pseudo-Homogeneous Models [Steinigeweg and Gmehling, 2002; Gangadwala et al.,
2004]. Then, the reaction kinetic model has the following form:
(4.1)
where, (4.2)
Chapter Four Modeling and Simulation
50
- Mcat is the catalyst mass.
- Kf and Ka are the forward reaction rate constant and equilibrium constant,
respectively.
- ai is the activity of the ith
component in the bulk liquid phase, xi is mole fraction of
the ith component.
- γi is the activity coefficient of the ith component.
The initial reaction rates in the carried experiments are calculated by the following
expression:
(4.3)
where Ci is n-butanol concentration and xi is the conversion of the n-butanol.
In the present work, the kinetic model for the esterification reaction was developed
using the experimental data obtained over a temperature range of 70-90 Co, a catalyst
loading 12, 24 and 36 g and for feed mole ratio (acetic acid/butanol) 2/1, 1/1 and 1/2.
Experimental reaction rate data was fitted using a nonlinear regression method. The
objective of the fitting is to minimize the sum of the squares of the difference between
the calculated and experimental values. For each reaction temperature, 8 data points
were used to calculate the parameters.
The following steps were used to calculate the parameters of reaction kinetic:
1- The first step is determining the rate constants of the reaction by getting the
concentration difference for each time interval using equation (4.1).
2- Depending on assumed forward reaction rate constant and equilibrium constant
using equation (2.4), simulation of the kinetic reactions is performed.
Chapter Four Modeling and Simulation
51
3- The simulation results are compared with experimental results at each measured
point. All deviations between experimental and calculated values are squared
and summed up to form an objective function Fobj. :
Fobj. = Σ (exp. conc. – calc. conc.)2
(4.4)
4- For each experimental data, the concentration-time data and new value of Fobj is
calculated. The rate constants corresponding to the minimum Fobj are stored and
considered improved rate constants for final or next iteration.
5- The optimizations proceed until the absolute difference between two successive
objective functions is less than a predefined tolerance.
The final obtained values of Kf and Ka will be the best rate constants. The effect of
temperature on reaction rate constants K is modeled using Arrhenius expression. The
activation energy (EA) of the reaction is calculated from two measuring points at
different temperatures:
(4.5)
4.3 Mathematical Model
4.3.1 Model Assumptions:
In the present work, the reactive distillation column is modeled as a tray column,
using reactive and non reactive stages where appropriate. Packed distillation column
model can be achieved by considering a specific height of packing as theoretical plate
depending on the idea of Height Equivalent to a Theoretical Plate (HETP). In the
Chapter Four Modeling and Simulation
52
present study each 10 cm of column assumed as single theoretical stage[Cristhian
2005].
Figure (4.1) represents a scheme of the reactive tray in continuous reactive distillation
column. By starting from the lower point, the reboiler is numbered as stage number
one and the first stage (section 1) of packing column is numbered as stage (2), then we
count from the bottom to the top. The last tray of the column is thus stage number
(19), and the condenser is considered as stage number (20).
yV kk xL kk 11
yv kk 11 xL kk
Fig (4.1) Scheme of the reactive tray in continuous reactive distillation column.
Therefore, the proposed model includes the following assumption:
1. The vapor and liquid stream leaving a stage are in thermodynamic equilibrium with
one another.
2. Column pressure is constant at atmospheric pressure with a neglected vapor holdup.
3. Perfect mixing on all stages and in all vessels (condenser and reboiler).
4. Ideal vapor phase for all components in the mixture.
5. The condenser and the reboiler are treated as equilibrium stages and all stages
efficiency is assumed 100%.
6. Total condensation of the top condenser.
Figure (4.2) represents continuous reactive distillation column, there is vapor liquid
equilibrium in the reboiler and condenser which can be assumed as a theoretical stage.
A+B C+ D
Chapter Four Modeling and Simulation
53
Making the total material, component, and energy balances on the various section of
continuous reactive distillation column, and further simplification of the equation lead
to drive the present model.
Fig (4.2) Continuous reactive distillation column.
L N-1
Lj-1
Vk+1
L1
Feed 1 (HAc)
Feed 2 (BuOH)
Stage 1
Stage w+1
Stage w
Stage w+1
Stage N-1
Stage 2
Stage j-1
Stage j
Stage j+1
Stage k-1
Stage k
Stage k+1
Stage 20
LN D
VN-1
L2
V1
Vw
-1
Vk
Vj+1
Vj
Lj+1
Vw+1
Lj
Lk
Lk+1
Lw
Lw+1
L1
V2
LN-1
Chapter Four Modeling and Simulation
54
4.3.2 Model Equations:
I. Reboiler: (stage 1)
a) Total material balance:
(4.6)
b) Component Material Balance:
(4.7)
c) Energy Balance:
(4.8)
d) Summation:
(4.9)
II. Non Reactive Trays: (stage j)
a) Total material balance:
(4.10)
b) Component material balance:
(4.11)
c) Energy balance:
(4.12)
Chapter Four Modeling and Simulation
55
d) Summation:
(4.13)
III. Reactive tray: (stage k)
a) Total material balance:
(4.14)
b) Component material balance:
(4.15)
c) Energy balance:
(4.16)
d) Summation:
(4.17)
IV. Condenser: (stageN)
a) Total material balance:
(4.18)
Chapter Four Modeling and Simulation
56
b) Component material balance:
(4.19)
c) Energy balance:
(4.20)
d) Summation:
(4.21)
e) Reflux Ratio:
Rref. = LN/D (4.22)
4.3.3 Estimation of Model Parameters:
4.3.3.1 Antoine Model:
The vapor pressure of each component in this study is obtained by using Antoine
equation [Sinnott and Tower 2009]:
(4.23)
Appendix (A), Table (A-4) shows the parameters of Antoine equation for all
components used in this study.
Chapter Four Modeling and Simulation
57
4.3.3.2 Equilibrium relations:
For non-ideal mixture additional variable γi appears to represent the degree of
deviation from ideality.
(4.24)
The UNIQUAC model was used to determine the activity coefficients in the liquid
phase while the vapor phase was assumed ideal. Appendix B shows the UNIQUAC
model equation and parameter.
4.3.3.3 Bubble Point Calculation:
The most widely employed numerical method for estimating bubble point of a mixture
is the Newton Raphson's technique. For distillation process the liquid of each tray is at
its bubble point and the vapor above the plate is at its dew point. The bubble point of
multi component mixture can be calculated by trial and error on the equilibrium
relationships.
(4.25)
When liquid at its bubble point then,
(4.26)
Moreover, when the vapor is at its dew point then,
(4.27)
To estimate the bubble point the Newton Raphson's iterative method equation is
written in the form
Chapter Four Modeling and Simulation
58
(4.28)
where
(4.29)
(4.30)
4.3.3.4 Enthalpy calculation:
The enthalpy of vapor and liquid phases mixtures were calculated by using the
following two equations:
(4.31)
(4.32)
(4.33)
(4.34)
4.3.3.5 Tray molar holdup:
The model is based on constant volume holdup on trays; therefore the molar holdup
on all trays and in the reflux drum is calculated using the following equation.
(4.35)
Chapter Four Modeling and Simulation
59
4.4 Numerical Computation
To calculate the composition of each component on each stage (condenser, trays, and
reboiler), the component mass balance equation are linearized and then solved using
fourth order Runge-kutta integration method. At each time step the vapor and liquid
flow rates were calculated depending on total mass and heat balance equations which
solved using Gauss Elimination Method as shown in matrix below:
The flow chart of simulation program which simulates the continuous reactive
distillation is shown in Figure (4.3). Since the number of equations is equal to the
number of variables, then the model can be solved to evaluate the following
parameters:
Liquid flow rate in the column.
Liquid composition profiles.
Vapor composition profiles.
Amount of distillate and bottom product.
Temperature profiles in the column.
Reaction rate profiles.
Chapter Four Modeling and Simulation
60
Fig (4.3) Flow chart of simulation program for continuous reactive distillation column.
Chapter Five Results and Discussion
61
CHAPTER FIVE RESULTS AND DISCUSSION
5.1 Introduction
This chapter includes the results of experimental work and the theoretical results
calculated from dynamic simulation model. The experimental results are divided into
two parts, the first one consisted of studying the reaction kinetic of acetic acid reacting
with n-butanol to produce n-butyl acetate and water, while the second part consisted
of studying the effect of various parameters on the composition and temperature of
reboiler and condenser for continuous reactive distillation column used to produce n-
butyl acetate and water from the reaction of acetic acid with n-butanol.
Theoretical results included the results of reaction kinetic parameters using dynamic
optimization method, simulation of composition and temperature profiles for all stages
by applying equilibrium model, comparisons between experimental and simulation
results, studying dynamic response of calculated model to various magnitudes of step
change in reflux ratio and feed flow rate and finally using dynamic simulation model
for selecting the optimum reflux ratio to operate the reactive distillation to maximize
the production of n-butyl acetate.
5.2 Reaction Kinetic
5.2.1 Effect of Temperature:
Figure (5.1) shows the experimental results of n-butanol conversion at three different
temperatures (70, 80 and 90 oC). According to this figure, the reaction rate is directly
proportional to the reaction temperature.
Chapter Five Results and Discussion
62
From these three experiments the forward reaction rate constants Kf and equilibrium
constant Ka were calculated at different temperatures using Matlab programs. The
Arrhenius equation was used to determine the apparent activation energy for the
forward reaction rate constant and equilibrium constant.
0
10
20
30
40
50
60
70
0 20 40 60 80 100 120 140
% n
-Bu
tan
ol
Co
nvers
ion
Time [min]
T=90 C
T=80 C
T=70 C
Fig (5.1) Effect of temperature on n-butanol conversion (catalyst Dowex 50 prepared with
HCl, wt: 24gm, ACH:BuOH=1:1).
5.2.2 Effect of catalyst type and modification with HCl:
Figure (5.2) shows n-butanol conversion using two catalysts, the first one is Dowex-
50 catalyst modified with HCl, while, the other is Dowex-50 without modification. It
is clear from this figure that the catalyst modification with HCl has a small effect on
increasing the conversion. This result is due to the fact that, the catalyst used is
already strong acidic resin and the modification with HCl has a weak effect on
increasing the acidity of the catalyst.
Chapter Five Results and Discussion
63
In Figure (5.3) the n-butanol conversion is plotted verses time for the two catalysts,
the first one is Amberlite CG catalyst modified with HCl while the other is Amberlite
CG without modification. In this figure, the catalyst modification with HCl decreases
the reaction rate.
By comparing the results of Figures (5.2) and (5.3), it is clear that the modefied
Dowex-50 gives a better total conversion of n-butanol around 67 %, while, using
Amberlite CG without modification gives lower total conversion of 27.5 %.
5.2.3 Effect of Catalyst Loading:
It is important to mention here, that the catalyst load regards a very important factor
in heterogeneous reaction.
Figure (5.4) represents the n-butanol conversion at various catalysts loading (12, 24
and 36 g) respectively. It is clear from this figure that increasing the amount of
catalyst loading will increase the conversion but does not have a significant effect on
final n-butanol conversion.
5.2.4 Effect Feed Ratio:
Figure (5.5) shows the n-butanol conversion at various acetic acid to n-butanol feed
ratios (1/2, 1/1 and 2/1) respectively. It is clear from this figure that using 100%
excess amount of acetic acid will increase the n-butanol from 67 % to 86 %, while
using 100% excess amount of n-butanol leads to a decrease in the acetic acid from
67 % to 40 %.
Chapter Five Results and Discussion
64
0
10
20
30
40
50
60
70
0 20 40 60 80 100 120 140
% n
-Bu
tan
ol
Co
nvers
ion
Time [min]
Modified catalyst with HCL
Not Modified with HCL
Fig (5.2) Effect of Catalyst Modification on n-butanol conversion (catalyst Dowex 50, wt:
24gm, T=90 C ,HAc:BuOH=1:1).
0
5
10
15
20
25
30
0 20 40 60 80 100 120 140
% n
-Bu
tan
ol
Co
nvers
ion
Time [min]
Modified catalyst with HCL
Not Modified with HCL
Fig (5.3) Effect of catalyst Modification on n-butanol conversion (catalyst AmberliteCG ,
wt:24gm,T=90 C ,HAc:BuOH=1:1).
Chapter Five Results and Discussion
65
0
10
20
30
40
50
60
70
0 20 40 60 80 100 120 140
% n
-Bu
tan
ol
Co
nvers
ion
Time [min]
36 gm
24 gm
12 gm
Fig. (5.4) Effect of catalyst weight on n-butanol conversion (catalyst Dowex 50 prepared with
HCl, T=90 C ,HAc:BuOH=1:1).
0
10
20
30
40
50
60
70
80
90
0 20 40 60 80 100 120 140
% n
-Bu
tan
ol
Co
nvers
ion
Time [min]
HAc:BuOH=1:2
HAc:BuOH=1:1
HAc:BuOH=2:1
Fig (5.5) Effect of feed mole ratio on n-butanol conversion (catalyst Dowex 50 prepared with
HCL, wt: 24gm).
Chapter Five Results and Discussion
66
5.2.5 Parameters Estimation of Reaction Kinetic:
The experimental results of reacting mixture compositions for the experiments in
Table (3.1) were fitted using a nonlinear regression method. Different values of the
forward reaction rate constants Kf and equilibrium constant Ka were calculated at
different temperatures using Matlab software as shown in the Table below.
Table (5.1) Forward and equilibrium constants at different temperature.
Temperature oC 70 80 90
Kf (mol/g.hr) 0.6912 1.5931 1.9101
Ka (-) 45.3068 12.7934 21.2913
The Arrhenius equation was used to determine the apparent activation energy for
forward reaction rate constant and equilibrium constant, the temperature dependency
of the constants was found to be:
(5.1)
(5.2)
The activation energy of the n-butanol esterification reaction was found to be
39.975kJ/mol.
5.2.6 Validity of Reaction Kinetic:
The model calculations for different operating conditions are compared with
experimental data by applying the predicted values in reaction rate equation (2.4).
Fourth order Runga-Kutta integration method was used, the four components
compositions were calculated for a given time range. The experimental and predicted
Chapter Five Results and Discussion
67
results are plotted in Figures (5.6), (5.7) and (5.8) for temperatures 70, 80 and 90 oC
respectively. The solid lines in these figures represent the predicted mole fractions
while the points represent the experimental mole fractions. Due to the presence of the
reacting mixture and reaction in equimolar proportion therefore the two reactants and
two products have the same profile. It was concluded from these three figures that the
predicted kinetic model has a great representation of n-butanol esterification with
acetic acid on Dowex-50 catalyst.
0
0.1
0.2
0.3
0.4
0.5
0.6
0 20 40 60 80 100 120 140
Mo
le fra
cti
on
Time [min]
HAc
BuOH
BuAc
H2O
Fig (5.6) Comparison between experimental data and rate equation results (catalyst Dowex 50
prepared with HCl, wt:24 gm ,ACH:BuOH=1:1, T=90 C), solid line represents rate equation
results.
Chapter Five Results and Discussion
68
0
0.1
0.2
0.3
0.4
0.5
0.6
0 20 40 60 80 100 120 140
Mol
e fr
actio
n
Time [min]
HAc
BuOH
BuAc
H2O
Fig (5.7) Comparison between experimental data and rate equation results (catalyst Dowex 50
prepared with HCl, wt:24 gm, ACH:BuOH=1:1, T=80 C), Solid line represents rate equation
results.
0
0.1
0.2
0.3
0.4
0.5
0.6
0 20 40 60 80 100 120 140
Mo
le fr
acti
on
Time [min]
HAc
BuOH
BuAc
H2O
Fig (5.8) Comparison between experimental with rate equation results (catalyst Dowex 50
prepared with HCL, wt:24 gm, ACH:BuOH=1:1, T=70 C), solid line represents rate equation
results.
Chapter Five Results and Discussion
69
5.3 Experimental Unit Results
In this part, the reactive distillation performance was tested through studying several
variables and parameters such as temperature distribution, reflux ratio, heat duty and
feed ratio. These parameters are very important in reactive distillation design and
operation.
5.3.1 Temperature Distribution along the Reactive Distillation:
In order to study the effect of reflux ratio on the temperature distribution along the RD
column, Figure (5.9) shows the effect of three levels of reflux ratio (1, 2 and 4) on the
temperature distribution along the reactive distillation under steady state operation
period.
It is important to mention here that the esterification processes are considered as an
exothermic reaction; therefore, benefits of heat integration are obtained because the
heat generation in the chemical reactions is used for vaporization. On the other hand,
the reactants and the product must have suitable volatility to maintain high
concentrations of reactants and low concentration of the products in the reaction zone.
Therefore, in esterification processes the temperature distribution and liquid-vapor
interaction must be known.
Figure (5.9) shows that there is an inverse relationship between reflux ratio and the
temperature distribution. Thus, as the reflux ratio increases the temperature
distribution level along the column axial decreases. The explanation of such behavior
is based on phenomenon of interaction between liquid-vapor equilibrium in reactive
distillation.
This result has a very good agreement in comparison with the results of Majid et al.
2008 where they worked on a smaller but similar column for the production of ethyl
Chapter Five Results and Discussion
70
acetate. They also found that the reflux ratio has a little effect on the level of
temperature distribution which can be seen obviousl at Reflux (2 and 4).
At reflux ratio equal to 1, a high increase in the temperature under the reaction zone in
the column was noticed although no floating or block in the column were noticed and
the pressure was kept at the atmospheric.
Fig (5.9) Temperature distribution along the reactive distillation column at different reflux
ratio.
Figure (5.10) shows the temperature distribution along the column at two amounts of
heat duty 380 and 440 watts. From this figure it can be concluded that by lowering the
heat duty, the rate of reaction increased since the temperature at the reaction zone
increased although most of the column was at lower temperature but the temperature
rose at the upper part of the column due to the heat of reaction.
The study of the effect of the feed flow rate on the temperature, lead to conclusion that
increasing of flowrate of acetic acid has little effect on the tempurature level along the
85
90
95
100
105
110
115
120
0 66 107 139 154 174 194 249
Tem
pra
ture
( c
)
column length (cm)
Ref 1
Ref 2
Ref 3
Chapter Five Results and Discussion
71
column.But it increases the tempurature in the reaction zone since the rate of reaction
increased as shown in Figure (5.11).
Fig (5.10) Temperature distribution along the reactive distillation column at different heat
duty.
Fig (5.11) Temperature distribution along the reactive distillation column at different mole
ratio.
85
90
95
100
105
110
115
0 66 107 139 154 174 194 249
Tem
pra
ture
( c
)
Column Length(cm)
440 watt
380 watt
85
90
95
100
105
110
115
0 66 107 139 154 174 194 249
Tem
pra
ture
( c
)
Column Length(cm)
1:1 mole ratio
2:1 mole ratio
Chapter Five Results and Discussion
72
5.3.2 Effect of Reflux Ratio
In this section the effect of the reflux ratio on the distillate and still composition is
studied, different reflux ratios were used under the same experimental condition. The
ranges of reflux ratio values reported by various investigators [Janowsky et al. 1997,
Hanika et al. 1999, Steinigeweg and Gmehling 2003] vary between 1and 6.
The present investigation was conducted at value of reflux ratio 1, 2 and 4.
In all experimental runs, the feed mole ratio of HAc and BuOH was kept at ratio of
1:1 and all the other operating conditions were kept constant.
Figure (5.12), (5.13) and (5.14) show the distillate and reboiler composition at
deferent reflux ratio. From the comparison between the three figures, it was noticed
that there is a slight effect of the reflux ratio on the mole fraction of the desired
product and this is attributed to the steady state of the reaction zone appearance at a
time more than 4 hours (the time of experiment). It can be noticed from Figure (5.12-
b) that the concentration of butyl acetate in the reboiler increased slowly from 0.214 to
0.365 and continued to increase with time.
By increasing the reflux ratio, Figures (5.13) and (5.14) show that the time needed to
reach the steady state became longer so the effect doesn’t appeared and the
concentration seemed almost constant throughout the time of the experiment.
Another explanation can be considered to explain the low concentration of the desired
product (n-butyl acetate) is
1) The highly non ideal nature of the quaternary mixture accompanied by the large
number of azeotropes.
2) The tendency of the reverse reaction to occur on certain stage.
The same explanation was given by Seader et al. (2006)
Chapter Five Results and Discussion
73
(a) Distillate composition
(b) Reboiler composition
Fig (5.12) Distillate and reboiler composition for (reflux ratio =1, ACH:BuOH=1:1,catalyst wt
100gm and (440 W) heat duty).
Chapter Five Results and Discussion
74
(a) Distillate composition
(b) Reboiler composition
Fig (5.13) Distillate and reboiler composition for (reflux ratio =2, ACH:BuOH=1:1,catalyst wt
100gm and (440 W) heat duty).
Chapter Five Results and Discussion
75
(a) Distillate composition
(b)Reboiler composition
Fig (5.14) Distillate and reboiler composition for (reflux ratio =4, ACH:BuOH=1:1,catalyst wt
100gm and (440 W) heat duty.
Chapter Five Results and Discussion
76
5.3.3 Effect of Heat Duty
Figure (5.15) shows the composition of the condenser and reboiler with time after
decreasing the heat duty to 380 watt. It’s clear from the trend of this figure, that the
system required more time to reach a steady state.
5.3.4 Effect of Feed Ratio
Figure (5.16) shows the composition of the condenser and reboiler with time after
increasing the flow rate of acetic acid by 100%. It’s clear from the trend of this figure,
that the system required more time to reach a steady state.
5.3.5 Condenser and Reboiler Temperature
Figures (5.17) to (5.21) represent the temperature behavior in the condenser and the
reboiler along the time of the experiments after 30 min of total reflux. These figures
show that a stable temperature exists in the condenser and reboiler along the time for
all experiments.
Chapter Five Results and Discussion
77
(a) Distillate composition
(b)Reboiler composition
Fig (5.15) Distillate and reboiler composition for (reflux ratio =2, ACH:BuOH=1:1,catalyst wt
100gm and (380 W) heat duty).
Chapter Five Results and Discussion
78
(a) Distillate composition
(b)Reboiler composition
Fig (5.16) Distillate and Reboiler composition for (reflux ratio =2, ACH:BuOH=2:1,catalyst wt
100gm and (440 W) heat duty).
Chapter Five Results and Discussion
79
Fig (5.17) Temperature profile of the reboiler and the condenser with time for (reflux ratio =1,
ACH:BuOH=1:1,catalyst wt 100gm and (440 W) heat duty).
Fig (5.18) Temperature profile of the reboiler and the condenser with time for (reflux ratio =2,
ACH:BuOH=1:1,catalyst wt 100gm and (440 W) heat duty).
Chapter Five Results and Discussion
80
Fig (5.19) Temperature profile of the reboiler and the condenser with time for (reflux ratio =4,
ACH:BuOH=1:1,catalyst wt 100gm and (440 W) heat duty).
Fig (5.20) Temperature profile of the reboiler and the condenser with time for (reflux ratio =2,
ACH:BuOH=1:1,catalyst wt 100gm and (380 W) heat duty).
Chapter Five Results and Discussion
81
Fig (5.21) Temperature profile of the reboiler and the condenser with time for (reflux ratio =2,
ACH:BuOH=2:1,catalyst wt 100gm and (440 W) heat duty).
5.4 Mathematical Model Analysis Results
Dynamic simulation has been carried out on the continuous reactive distillation
column using MATLAB software. Dynamic simulation model consists of a large
number of nonlinear ordinary differential equations and demands much information
about the system (feed flowrates, initial compositions, vapor and liquid flowrates,
liquid hold up in all stages at every instant, energy balance data, and vapor-liquid
equilibrium data). Table (5.1) contains the column specification values.
In simulation program, the stages were numbered from the bottom of the column
(reboiler) to the top of the column (condenser). The column was divided into 18
theoretical stages and two stages for reboiler and condenser respectively. The same
condition of exp. No. (2) in Table (3.2) was used in the simulation program. At the
start-up period of the continuous reactive distillation column operation, reboiler, trays,
and reflux drum are assumed to have the same composition of feeds. As in
Chapter Five Results and Discussion
82
experimental work, at startup period the column may be brought to the steady state by
following the column at total reflux procedure for 30 min. Then the production phase
began by switching on the distillate production according to a specific reflux ratio.
Table (5.1) Reactive distillation column specification
Acetic acid/n-butanol/n-butyl acetate/water) Continuous
Distillation
Total no. of trays including the reboiler and condenser 20
No. of trays in rectifying section 6
No. of trays in reactive section 4
No. of trays in stripping section 8
Reflux Ratio 2
Acetic acid feed tray 7
N-butanol feed tray 10
Feed flow rate of acetic acid (kmol/h) 4
Feed flow rate of n-butanol (kmol/h) 4
Top product flow rate (kmol/hr) 2.6
Bottom product flow rate (kmol/hr) 0.42
Total catalyst weight 100 gm
Condenser drum holdup 62 cm3
Reactive section stages holdup 47 cm3
Stripping and rectifying stages holdup 62 cm3
Rebolier Holdup 1500 cm3
Reboiler duty 440 Watt
Column diameter 2.54 cm
Figures (5.22) and (5.23) represent the simulation results of the distillate and bottom
product composition profiles respectively for the continuous reactive distillation
column model. It is evident from these two figures that the continuous reactive
distillation column takes nearly 6 hrs to reach a steady state.
Due to the water having higher volatility than the other three components in the
column, it does not appear in the bottom product, while, it is gradually concentrated in
the distillate, the same results were achieved in the present experimental work. The
Chapter Five Results and Discussion
83
steady state composition of the lightest component (water) in the distillate product is
44.5 %. The reactants (acetic acid and n-butanol) appear in equal proportions in the
distillate as a result of that at steady state the bottom product does not contain any
amount of reactants and the reactants consumed according to their stoichiometric
proportion. The appearing of acetic acid and n-butanol in the distillate indicates that
the conversion in the column does not reach a completion.
Figure (5.23) shows that, the n-butyl acetate is concentrated gradually in the bottom
of the catalytic distillation column until it reaches maximum value about (100 %) after
6 hours of startup. The column reaches a steady state values after 6 hrs due to huge
difference in hold-ups between the reboiler and stages within the column. The time
necessary to reach a steady state value increases as the difference in hold-ups
increases.
Fig (5.22) Mole fraction in condenser with respect to time (model results)
Chapter Five Results and Discussion
84
Fig (5.23) Mole fraction in reboiler with respect to time (model results).
5.4.1 Validity of Mathematical Model:
The dynamic simulation results of reactive distillation column show that the column
required at least five hrs to reach a steady state values and due to that all of five
experiments did not work more than four hrs. Therefore, the model cannot be
validated by comparing the model results with the experimental work results. To
overcome this problem, the model results were compared with the experimental and
theoretical results of other researchers. Our model was compared with the results of
Hanika et al. (1999), where they studied the production of n-butyl acetate in
continuous reactive distillation experimentally and theoretically, as following:
Chapter Five Results and Discussion
85
5.4.1.1 Comparison of Compositions Profile:
Figure (5.24) represents the model results for the steady state composition profiles of
four components within the column, while Figure (5.25) represents the Hanika model
results for compositions profile within the column.
In spite of that Hanika took 50 stages within the column, but both figures have the
same trend. In Figure (5.24) a drop in n-butyl acetate composition appears in the
middle of the column as a result of increasing the n-butanol composition in the feed
stage.
Fig (5.24) Mole fraction with respect to tray (model results).
Chapter Five Results and Discussion
86
Fig (5.25) Mole fraction with respect to tray (Hanika model results).
5.4.1.2 Comparison of temperature profile:
Figure (5.26) represents the model results for the temperature profile within the
column at steady state, while, Figure (5.27) represents the experimental and
theoretical temperature profile of Hanika’s study.
In both figures the temperature in the rebolier is around 126 Co. In spite of the
difference in column specifications between the two studies, the agreement between
the temperature distributions is very clear. Temperature profile is smooth in the
rectifying and reaction zones, while, the temperature profile in the stripping zone is
very steep, corresponding to separation of very different boiling point components (
n-butylacetate, acetic acid and n-butanol).
Figure (5.28) shows the steady state vapor and liquid flowrates within the column.
The temperature deviation between the two figures is zero while there is a significant
HAc
Chapter Five Results and Discussion
87
deviation in composition as a result of the deference in assumptions, specifications
and VLE data of the two models.
Fig (5.26) Tray temperature with respect to tray number (model results).
Fig (5.27) Tray temperature with respect to tray number (Hanika model results).
Chapter Five Results and Discussion
88
Fig (5.28) Liquid and vapor flow rates with respect to tray No. (model results)
5.5 Process Dynamics
5.5.1 Response to Reflux Ratio:
Dynamic simulations enable the study of the influence of disturbances in the product
specifications on the process output. Table (5.3) shows the steady state distillate and
bottom products compositions.
Figures (5.29 a, b, c, d and e) show the dynamic response of distillate molar fractions
and temperature versus time for a ±50% change in the reflux ratio. In Figure (5.29)
increasing reflux ratio from 2 to 3 makes the reactant composition (acetic acid and n-
butanol) in the distillate decrease from 13.53% to 13.1 % and water composition
increases from 45% to 46.1 % as an indication of increasing the conversion of the
reactants to produce water and n-butyl acetate. Such behavior is attributed to
Chapter Five Results and Discussion
89
increasing the reaction rate in reacting zone due to the increase of the residence time
in reaction zone. The reboiler composition is not affected by increasing reflux ratio
but its flowrate will increase. Decreasing reflux ratio from 2 to 1 makes the reactant
composition (acetic acid and n-butanol) in the distillate increase and water
composition decrease.
Table (5.3) Steady state condition within the column
Top product flow rate (kmol/hr) 2.6
Bottom product flow rate (kmol/hr) 0.42
Distillate composition
BuOH 0.1354610
HAc 0.1351752
BuAc 0.2799193
H2O 0.4494445
Bottom product composition
BuOH 0.0001191
HAc 0.0000117
BuAc 0.9998693
H2O 0.0000000
Figures (5.29 e) shows the distillate temperature response versus time for a +50% and
-50% step change in the reflux ratio for the distillate and bottom.
5.5.2 Response to Feed Flow Rate:
Figures (5.30 a, b, c, d and e) represent the dynamic response of the distillate molar
fractions respond to various step change (±50%) in acetic acid feed flow rate. In this
figure, increasing acetic acid feed flow rate lead to decrease the compositions of the
other three components in the distillate while decreasing the acetic acid flow rate will
increase the n-butanol composition in the distillate.It is evident from these figures
that, the new steady state, corresponding to the new values set for the reflux ratio or
acetic acid feed flow rate is reached after a transient period of approximately 2 hrs.
Chapter Five Results and Discussion
90
a-N
-but
anol
com
posi
tion
b-A
ceti
c ac
id c
ompo
sitio
n
c-B
utyl
ace
tate
com
posi
tion
d-W
ater
Com
posi
tion
e-C
onde
nser
Tem
pera
ture
Fig (5.29) Condenser mole fraction and temperature with respect to time for a ±50 % step change in reflux ratio .
0.132
0.134
0.136
0.138
0.14
0.142
0.144
10 11 12 13 14
N-b
uta
no
l
Time )hr)
Reflux Ratio +50%
Reflux Ratio -50%
0.132
0.134
0.136
0.138
0.14
0.142
0.144
10 11 12 13 14
Ace
tic
Aci
d
Time )hr)
Reflux Ratio +50% Reflux Ratio -50%
0.27
0.275
0.28
0.285
0.29
0.295
0.3
10 11 12 13 14
Bu
tyl A
ceta
te
Time )hr)
Reflux Ratio +50% Reflux Ratio -50%
0.41
0.42
0.43
0.44
0.45
0.46
0.47
10 11 12 13 14
Wat
er
Time )hr)
Reflux Ratio +50% Reflux Ratio -50%
362.6
362.8
363
363.2
363.4
10 11 12 13 14
Te
mp
era
ture
( K
)
Time )hr)
Reflux Ratio +50% Reflux Ratio -50%
Chapter Five Results and Discussion
91
a-N
-but
anol
com
posi
tion
b-A
ceti
c ac
id c
ompo
sitio
n
c-B
utyl
ace
tate
com
posi
tion
d-W
ater
Com
posi
tion
e-C
onde
nser
Tem
pera
ture
Fig (5.30) Condenser mole fraction and temperature with respect to time for a ±50 % step change Acetic acid feed flow rate.
0
0.1
0.2
0.3
0.4
10 11 12 13 14
N-b
uta
no
l
Time )hr)
Acetic acid feed flowrate (+50%) Acetic acid feed flowrate (-50%)
0
0.1
0.2
0.3
0.4
10 11 12 13 14
Ace
tic
Aci
d
Time )hr)
Acetic acid feed flowrate (+50%) Acetic acid feed flowrate (-50%)
0.25
0.26
0.27
0.28
0.29
0.3
10 11 12 13 14
Bu
tyl A
ceta
te
Time )hr)
Acetic acid feed flowrate (+50%) Acetic acid feed flowrate (-50%)
0.39
0.4
0.41
0.42
0.43
0.44
0.45
0.46
10 11 12 13 14
Wat
er
Time )hr)
Acetic acid feed flowrate (+50%) Acetic acid feed flowrate (-50%)
362
363
364
365
366
367
10 11 12 13 14
Tem
pe
ratu
re (
K )
Time )hr)
Acetic acid feed flowrate (+50%) Acetic acid feed flowrate (-50%)
Chapter Five Results and Discussion
92
5.6 Determination of Reflux Ratio
The purity of n-butyl acetate in the bottom product depends on the catalytic
distillation column regime. Optimum operating policies, i.e., optimum reflux ratio
were estimated by simulating the continuous reactive distillation column for different
but constant reflux ratios thereby maximizing the production rate of n-butyl acetate.
Searching optimum reflux ratio carried out for several reflux ratios ranging from 1 to
6. For each of the different but constant reflux ratio, optimal operating conditions
were derived. Figure (5.31) shows the effect of increasing the reflux ration on n-
butanol conversion within the column. In this figure increasing the reflux ratio from 2
to 6 will increase the n-butanol conversion from 49.56 % to 92%.
Fig (5.31) Effect of reflux ratio on n-butanol conversion in reactive distillation column (model
result)
The esterification of acetic acid and n-butanol is a reversible and kinetically controlled
reaction. To increse the conversion of reactant a further option arise which is using a
prereactor before continous reactive diatillation processes. Chemical equilibrium can
be obtained in the prereactor and the reactive distillation column should enhance the
conversion to nearly 100%.
40
50
60
70
80
90
100
1 2 3 4 5 6 % n
-Bu
tan
ol C
on
vers
ion
Reflux Ratio
Chapter Six Conclusions and Recommendations
93
CHAPTERSIX CONCLUSIONS AND RECOMMENDATIONS
6.1 Conclusions
The following major conclusions can be drawn from the results obtained:
1- The results show that the modified catalysts with hydrochloric acid gave higher
activity than parent catalysts types due to acidity modification of active sites of
the catalysts.
2- The results obtained from kinetics study shows that the modified Dowex-50
catalyst gives the highest conversion rate of n-butanol and acetic acid to produce
n-butyl acetate. The n-butanol conversion is 67% in batch reactor with activation
energy equal to39.975 kJ/mol.
3- In the present work, a Pseudo-Homogeneous Model was developed to describe the
reaction kinetic. The comparison between model results and experimental
results shows very good agreement between them.
4-It is concluded that, when the reflux ratio increases the temperature level along the
column decreases. On the other hand, the increase of acetic acid flow rate or heat
duty lead to slightly affecting the temperature level.
5-It was noticed that when the acetic acid flow rate increases, the reaction zone
temperature increases too. On the other hand, an opposite conclusion was noted
that, when heat duty is increased the reaction zone temperature decreased.
6- The results indicated that, the conversion in reactive distillation increases directly
with increasing the reflux ratio. This result could be regarded as to increase the
resident time in the reaction zone when the reflux ratio increased.
7- It was observed that, the use of equimolar reactants (equimolar feed moles) is
necessary to produce the best purity and productivity of desired product in the
continuous reactive distillation column.
Chapter Six Conclusions and Recommendations
94
8-Decreasing the ratio between stages holdup and reboiler holdup will increase the
speed of column response and decreases the time to reach steady state value in
dynamic distillation columns.
6.2 Recommendations for the Future Work
For the future work, the following suggestions can be considered:
1-Experimental work could be extended through investigation of other reactions
such as: etherification reaction to produce oxygenates like ETBE, MTBE, TAME
and TAEE.
2-Driving Non-equilibrium mathematical models for the same system and comparing
the result with that of the present model.
3- Studying the effect of various types of feedback controllers on the performance of
continuous reactive distillation columns.
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113
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Appendix
95
Appendix A
Technical Data
Table (A-1) Physical Properties [R.Sinnott and G.Towler 2009]
Component
Boiling
Point (oC)
Density
(kg/m3)
Molecular weight
(gm/gmol)
H at
(298K)
(KJ/mol)
Latent
heat
(J/mol)
Butanol 117.7 810 74.123 -274.86 43124
Acetic acid 117.9 1094 60.052 -435.13 23697
Butyl acetate 126 898 116.16 -486.76 36006
Water 100 998 18.015 -242 40683
Table (A-2) Specific heat of Vapor [R.Sinnott and G.Towler 2009].
Component
CPV=a+b*T+c*T^2+d*T^3 (J/mol.K)(T in K)
a b c d
Butanol 3.266 41.801E-2 -2.328 E-4 46.85 E-9
Acetic acid 4.84 25.485E-2 -1.753 E-4 49.488 E-9
Butyl
acetate
13.62 54.8889 E-2 -2.278 E-4 -7.913E-10
Water 32.243 19.238 E-4 10.555 E-6 -3.596 E-9
Appendix
96
Table (A-3) specific heat of Liquid from [R.Sinnott and G.Towler 2009]
Component
CPL=a+b*T+c*T^2+d*T^3*e*T^4 (J/kmol.K)(T in K)
a b c d e
Butanol 191200 -730.4 2.2998 0 0
Acetic acid 139640 -320.8 0.8985 0 0
Butyl
acetate
111850 384.52 0 0 0
Water 276370 -2090.1 8.125 -0.014116 9.3701E-06
Table (A-4) Constants of Antoine Equation [R.Sinnott and G.Towler 2009].
Component
ln(P)=a-(b/(T+c))
(P in mmHg) (T in K)
a b c
Butanol 17.216 3137 -69.15
Acetic acid 16.808 3405.6 -56.34
Butyl acetate 16.171 3151.1 -69.15
Water 18.304 3816.4 -46.13
Appendix
97
Appendix B
UNIQUAC Model
The following equations represent UNIQUAC model. This model distinguishes
two contributions termed combinatorial Co and residual Rs.
(B-1)
Where,
(B-2)
(B-3)
Where,
(B-4)
(B-5)
(B-6)
Where z=10
Appendix
98
Table (B-1) Binary interaction parameters for the UNIQUAC model (Kathel and Amiya
(2010)). (cal/mol)
Component r q
Acetic acid 2.2024 2.072
n-Butanol 4.4544 4.052
n-Butyl acetate 4.8274 4.196
Water 0.92 1.4
Table (B-2) Binary interaction parameters for the UNIQUAC model (Kathel and Amiya
(2010)). (cal/mol)
A11=0 A12=-141.7686 A14=-298.4444 A14=-444.594
A21=148.2844 A22=0 A24=82.5446 A24=68.0084
A41=712.2449 A42=24.6486 A44=0 A44=685.71
A41=527.9296 A42=581.1471 A44=461.4747 A44=0
Appendix
99
Appendix C Calibration Data
Thermocouple Calibration:
Table (C-1) Thermocouple Calibration Data.
Measured
Temp.
-------------
Real
Temp.
No.1 No.1
av. No.2
No.2
av. No.3
No.3
av. No.4
No.4
av. No.5
No.5
av. No.6
No.6
av. No.7
No.7
av. No.8
No.8
av.
16
20.5
20.27
20.2
20.27
20.6
20.33
21.1
20.8
21.3
21.03
20.5
20.47
20.7
20.7
22.2
21.73 20.1 20.2 20.2 20.7 20.7 20.4 20.6 21.4
20.2 20.4 20.2 20.6 21.1 20.5 20.8 21.6
20
25.9
26.47
26.3
26.6
26.2
26.63
27.1
27.13
26.7
27.57
27
27.23
26.7
27.17
28.8
28.9 26.2 26.8 26.6 27.2 27.9 27.1 27.4 28.2
27.3 26.7 27.1 27.1 28.1 27.6 27.4 29.7
30
36.4
38.53
37.2
37.57
37.6
38
37
37.7
37.7
38.6
37.6
38.23
37.6
38.47
39.2
40.17 37.6 37.7 37.5 37.6 38.7 38 38.4 40.6
41.6 37.8 38.9 38.5 39.4 39.1 39.4 40.7
40
47.9
48.17
48.1
48.33
48.4
48.57
48.3
48.77
49.2
49.23
48.5
48.67
48.5
48.87
51.1
51.03 48.3 48.4 48.5 49.3 48.8 48.7 48.6 50.9
48.3 48.5 48.8 48.7 49.7 48.8 49.5 51.1
50
59.4
59.7
59.4
59.37
59.8
59.8
59.2
59.47
60
60.77
59.6
60.1
60.2
60.57
62.6
62.57 59.8 59.2 59.5 59.5 61.2 60 60.6 62.7
59.9 59.5 60.1 59.7 61.1 60.7 60.9 62.4
60
70
70.4
70.1
70.47
70.4
70.13
70.1
70.07
70.4
70.53
70.4
70.1
70.4
70.87
71
71.33 70.6 70.5 70.2 70.1 70.5 70.1 70.8 70.8
70.6 70.8 69.8 70 70.7 69.8 71.4 72.2
70
82.6
83.4
80.8
81.73
81
81.73
79.9
80.5
81.8
82.7
80.5
81.07
82.7
82.7
83
83.37 82.3 81.5 82.1 81.2 83.2 81.2 82.2 83
85.3 82.9 82.1 80.4 83.1 81.5 83.2 84.1
80
93.1
92.87
92.4
92.37
92.8
92.57
92.4
92.17
92.8
92.57
92.2
92.03
92.6
92.23
92.6
92.63 93 92.4 92.5 92 92.6 92.1 92.2 92.7
92.5 92.3 92.4 92.1 92.3 91.8 91.9 92.6
90
103.6
104.5
103.7
104.6
104.1
104.7
103.8
104.5
104.2
105.1
103.8
104.6
104.2
105
105.1
105.7 104.6 104.7 104.7 104.7 105.2 104.6 105.2 105.6
105.2 105.4 105.4 105.1 105.9 105.3 105.7 106.4
100 116.4
117.3 116.8
117.7 116.8
117.5 116.8
117.7 117.4
118.2 117.3
118 117.4
118.2 118.1
118.7
117.4 117.6 117.6 117.7 118.1 117.9 118.1 118.6
Appendix
100
118.2 118.6 118 118.5 119 118.8 119.1 119.3
110
128.7
129.1
129.1
129.9
129.5
129.5
129.3
129.3
129.4
129.5
129.4
129.2
129.9
130
130.2
130.1 129.1 129.8 129.6 129.2 129.5 129.1 130.1 129.9
129.5 130.9 129.5 129.3 129.7 129.1 130.1 130.2
120
139.9
141.3
140.3
141.6
139.9
140.4
136.5
138.7
140.1
140.8
138.9
139.8
138.6
140.6
141.2
141.4 141.6 141.9 140.4 139.4 141.1 139.9 141.5 141.4
142.5 142.5 141 140.2 141.2 140.6 141.6 141.5
130
153.2
153.7
153.3
153.7
153.4
153.6
153.4
153.1
153.6
154.2
153.2
153.2
153.4
153.5
154.6
154.2 153.9 154.2 153.8 153.2 154.6 153.5 153.6 154.1
154 153.7 153.7 152.7 154.5 153 153.4 154
Table (C-2) Thermocouple Calibration equations.
Thermocouple calibration equations
No.1 Real Temp. = -1.9569+0.8671*x
No.2 Real Temp. = -1.5281+0.8632*x
No.3 Real Temp. = -1.8987+0.8681*x
No.4 Real Temp. = -2.2323+0.8751*x
No.5 Real Temp. = -2.6359+0.8711*x
No.6 Real Temp. = -2.2427+0.873*x
No.7 Real Temp. = -2.4588+0.8706*x
No.8 Real Temp. = -4.0567+0.8792*x
Appendix
101
Fig.(C-1) Calibration curve of thermocouple No.1.
Fig.(C-2) Calibration curve of thermocouple No.2.
0
20
40
60
80
100
120
140
160
20.27 26.47 38.53 48.17 59.70 70.40 83.40 92.87 104.47 117.33 129.10 141.33 153.70
Re
al T
em
p. (
c)
Measurd Temp. (c)
0
20
40
60
80
100
120
140
160
20.3 26.6 37.6 48.3 59.4 70.5 81.7 92.4 104.6 117.7 129.9 141.6 153.7
Re
al T
em
p. (
c)
Measured Temp.(c)
Appendix
102
Fig.(C-3) Calibration curve of thermocouple No.3.
Fig.(C-4) Calibration curve of thermocouple No.4.
0
20
40
60
80
100
120
140
160
20.3 26.6 38.0 48.6 59.8 70.1 81.7 92.6 104.7 117.5 129.5 140.4 153.6
Re
al T
em
p.(
c)
Measurd Temp. (c)
0
20
40
60
80
100
120
140
160
180
20.8 27.1 37.7 48.8 59.5 70.1 80.5 92.2 104.5 117.7 129.3 138.7 153.1
Re
al T
em
p. (
c)
Measurerd Temp. (c)
Appendix
103
Fig.(C-5) Calibration curve of thermocouple No.5.
Fig.(C-6) Calibration curve of thermocouple No.6.
0
20
40
60
80
100
120
140
160
180
21.0 27.6 38.6 49.2 60.8 70.5 82.7 92.6 105.1 118.2 129.5 140.8 154.2
Re
al T
em
p.(
c)
Measured Temp. (c)
0
20
40
60
80
100
120
140
160
180
Re
al T
em
p. (
c)
Measured Temp. (c)
Appendix
104
Fig.(C-7) Calibration curve of thermocouple No.7.
Fig.(C-8) Calibration curve of thermocouple No.8.
0
20
40
60
80
100
120
140
160
180
Re
al T
em
p. (
c)
Measured Temp. (c)
0
20
40
60
80
100
120
140
160
180
Re
al T
em
p. (
c)
Measured Temp. (c)
Appendix
105
GC Calibration:
Table (C-3) GC calibration data.
mole fraction(x)
Area
Acetic
acid
Butanol butyl acetate
0 0 0 0
0.1 6734430 20743469 26716094
0.2 11500877 28522304 33112659.5
0.25 14560330 39705340 39509225
0.3 16165876 51008442 50844979
0.4 19509518 61870670 88102623
0.5 27602374 88942713 116977315
0.75 40117594
Table (C-4) GC calibration equations.
material calibration equations
Acetic acid mole fraction(x) = -0.0109+1.8994E-8*Area
Butanol mole fraction(x) = 0.0086+5.8118E-9*Area
butyl acetate mole fraction(x) = 0.037+4.1975E-9*Area
Fig.(C-9) GC Calibration Curves
0.00
2.00
4.00
6.00
8.00
10.00
12.00
14.00
16.00
18.00
0 0.2 0.4 0.6 0.8
Are
a x
10
00
00
00
Mole fraction
acetic acid
Butanol
butyl acetate
Appendix
106
Appendix D
Simulation and Experimental Data
D-1 Kinetic Data
Table (D-1.1) Run No.1 experimental data.
t (min) HAc(%mole) buoH(%mole) buAc(%mole) H2o(%mole) conversion
0 0.5 0.5 0 0 0
10 0.364813798 0.364813798 0.135186202 0.135186202 27.03724042
20 0.290428901 0.290428901 0.209571099 0.209571099 41.91421981
30 0.241252248 0.241252248 0.258747752 0.258747752 51.74955047
45 0.205943276 0.205943276 0.294056724 0.294056724 58.81134487
60 0.190037934 0.190037934 0.309962066 0.309962066 61.99241324
90 0.174193675 0.174193675 0.325806325 0.325806325 65.1612651
120 0.165081972 0.165081972 0.334918028 0.334918028 66.98360557
Table (D-1.2) Run No.1 simulated data
t (min) HAc(%mole)2 buoH(%mole)3 buAc(%mole)4 H2o(%mole)5
0 0.5 0.5 0 0
10 0.35621491 0.35621491 0.143785 0.143785
20 0.28301576 0.28301576 0.216984 0.216984
30 0.24199484 0.24199484 0.258005 0.258005
45 0.20963365 0.20963365 0.290366 0.290366
60 0.19403007 0.19403007 0.30597 0.30597
90 0.18242595 0.18242595 0.317574 0.317574
120 0.17942964 0.17942964 0.32057 0.32057
Appendix
107
Table (D-1.3) Run No.2 experimental data.
t (min) HAc(%mole) buoH(%mole) buAc(%mole) H2o(%mole) conversion
0 0.5 0.5 0 0 0
10 0.365632921 0.365632921 0.134367079 0.134367079 26.87341577
20 0.328198482 0.328198482 0.171801518 0.171801518 34.36030369
30 0.283304677 0.283304677 0.216695323 0.216695323 43.33906464
45 0.240945608 0.240945608 0.259054392 0.259054392 51.81087833
60 0.221540266 0.221540266 0.278459734 0.278459734 55.69194681
90 0.193475168 0.193475168 0.306524832 0.306524832 61.30496649
120 0.180449288 0.180449288 0.319550712 0.319550712 63.91014239
Table (D-1.4) Run No.2 simulated data
t (min)
HAc(%mole) buoH(%mole) buAc(%mole) H2o(%mole)
0 0.5 0.5 0 0
10 0.40038804 0.40038804 0.099612 0.099612
20 0.33563495 0.33563495 0.164365 0.164365
30 0.29126159 0.29126159 0.208738 0.208738
45 0.247222 0.247222 0.252778 0.252778
60 0.21929349 0.21929349 0.280707 0.280707
90 0.18852362 0.18852362 0.311476 0.311476
120 0.17415705 0.17415705 0.325843 0.325843
Table (D-1.5) Run No.3 experimental data .
t (min) HAc(%mole) buoH(%mole) buAc(%mole) H2o(%mole) conversion
0 0.5 0.5 0 0 0
10 0.426429341 0.426429341 0.073570659 0.073570659 14.71413179
20 0.386452021 0.386452021 0.113547979 0.113547979 22.70959584
30 0.381345493 0.381345493 0.118654507 0.118654507 23.73090145
45 0.314415305 0.314415305 0.185584695 0.185584695 37.11693908
60 0.273244735 0.273244735 0.226755265 0.226755265 45.35105299
90 0.231449198 0.231449198 0.268550802 0.268550802 53.71016048
120 0.204792375 0.204792375 0.295207625 0.295207625 59.04152491
Appendix
108
Table (D-1.6) Run No.3 simulated data
t (min) HAc(%mole) buoH(%mole) buAc(%mole) H2o(%mole)
0 0.5 0.5 0 0
10 0.43567597 0.43567597 0.064324 0.064324
20 0.38649445 0.38649445 0.113506 0.113506
30 0.34788707 0.34788707 0.152113 0.152113
45 0.30380599 0.30380599 0.196194 0.196194
60 0.27117207 0.27117207 0.228828 0.228828
90 0.22697782 0.22697782 0.273022 0.273022
120 0.19955698 0.19955698 0.300443 0.300443
Table (D-1.7) Run No.4 experimental data.
t (min) HAc(%mole) buoH(%mole) buAc(%mole) H2o(%mole) conversion
0 0.5 0.5 0 0 0
10 0.403474634 0.403474634 0.096525366 0.096525366 19.30507329
20 0.318698249 0.318698249 0.181301751 0.181301751 36.26035014
30 0.267856183 0.267856183 0.232143817 0.232143817 46.42876347
45 0.230134571 0.230134571 0.269865429 0.269865429 53.97308575
60 0.20796129 0.20796129 0.29203871 0.29203871 58.40774194
90 0.206903037 0.206903037 0.293096963 0.293096963 58.61939263
120 0.171284153 0.171284153 0.328715847 0.328715847 65.74316934
Table (D-1.8) Run No.5 experimental data.
t (min) HAc(%mole) buoH(%mole) buAc(%mole) H2o(%mole) conversion
0 0.5 0.5 0 0 0
10 0.401620568 0.401620568 0.098379432 0.098379432 19.67588642
20 0.299159312 0.299159312 0.200840688 0.200840688 40.16813769
30 0.262405748 0.262405748 0.237594252 0.237594252 47.51885049
45 0.228340159 0.228340159 0.271659841 0.271659841 54.33196813
60 0.20809164 0.20809164 0.29190836 0.29190836 58.38167203
90 0.188910044 0.188910044 0.311089956 0.311089956 62.2179912
120 0.175577296 0.175577296 0.324422704 0.324422704 64.88454071
Appendix
109
Table (D-1.9) Run No.6 experimental data .
t (min) HAc(%mole) buoH(%mole) buAc(%mole) H2o(%mole) conversion
0 0.5 0.5 0 0 0
10 0.257178674 0.257178674 0.242821326 0.242821326 48.56426527
20 0.207716344 0.207716344 0.292283656 0.292283656 58.45673114
30 0.188628262 0.188628262 0.311371738 0.311371738 62.27434765
45 0.175005004 0.175005004 0.324994996 0.324994996 64.99899925
60 0.167027526 0.167027526 0.332972474 0.332972474 66.59449472
90 0.160675354 0.160675354 0.339324646 0.339324646 67.86492927
120 0.154741972 0.154741972 0.345258028 0.345258028 69.0516057
Table (D-1.10) Run No.7 experimental data.
t (min) HAc(%mole) buoH(%mole) buAc(%mole) H2o(%mole) conversion
0 0.5 0.5 0 0 0
10 0.433513095 0.433513095 0.066486905 0.066486905 13.29738102
20 0.426419086 0.426419086 0.073580914 0.073580914 14.71618276
30 0.411008721 0.411008721 0.088991279 0.088991279 17.79825578
45 0.408771584 0.408771584 0.091228416 0.091228416 18.24568311
60 0.403182561 0.403182561 0.096817439 0.096817439 19.36348778
90 0.385107867 0.385107867 0.114892133 0.114892133 22.97842663
120 0.36865134 0.36865134 0.13134866 0.13134866 26.26973207
Table (D-1.11) Run No.8 experimental data.
t (min) HAc(%mole) buoH(%mole) buAc(%mole) H2o(%mole) conversion
0 0.5 0.5 0 0 0
10 0.376644063 0.376644063 0.123355937 0.123355937 24.67118749
20 0.376046786 0.376046786 0.123953214 0.123953214 24.79064272
30 0.37368003 0.37368003 0.12631997 0.12631997 25.26399394
45 0.371723104 0.371723104 0.128276896 0.128276896 25.65537926
60 0.369333012 0.369333012 0.130666988 0.130666988 26.13339757
90 0.364505548 0.364505548 0.135494452 0.135494452 27.09889041
120 0.357511702 0.357511702 0.142488298 0.142488298 28.4976596
Appendix
110
Table (D-1.12) Run No.9 experimental data .
t (min) HAc(%mole) buoH(%mole) buAc(%mole) H2o(%mole) conversion
0 0.333333333 0.666666667 0 0 0
10 0.182603715 0.515937049 0.150729618 0.150729618 22.60944269
20 0.174249171 0.507582505 0.159084162 0.159084162 23.86262428
30 0.147675416 0.481008749 0.185657917 0.185657917 27.84868762
45 0.112491606 0.445824939 0.220841728 0.220841728 33.12625917
60 0.095315083 0.428648417 0.23801825 0.23801825 35.7027375
90 0.081254022 0.414587356 0.252079311 0.252079311 37.81189665
120 0.068685613 0.402018946 0.26464772 0.26464772 39.69715803
Table (D-1.13) Run No.10 experimental data .
t (min) HAc(%mole) buoH(%mole) buAc(%mole) H2o(%mole) conversion
0 0.666666667 0.333333333 0 0 0
10 0.492091978 0.158758644 0.174574689 0.174574689 52.37240671
20 0.436986111 0.103652777 0.229680556 0.229680556 68.90416685
30 0.41099629 0.077662957 0.255670376 0.255670376 76.70111287
45 0.396797855 0.063464522 0.269868811 0.269868811 80.96064342
60 0.38882831 0.055494977 0.277838357 0.277838357 83.35150699
90 0.383356967 0.050023633 0.2833097 0.2833097 84.99291
120 0.382041228 0.048707895 0.284625439 0.284625439 85.38763162
D-2 Pilot Plant Experimental Data
Table (D-2.1) experiment No. 1:
Distillate Composition Reboiler Composition
Time (min)
Acetic acid
Butanol Butyl Acetate
Water Acetic acid
Butanol Butyl Acetate
Water
0 0 0.1556 0.3378 0.5066 0.3793 0.4067 0.214 0
30 0 0.1653 0.3226 0.5121 0.3635 0.402 0.2346 0
60 0 0.1427 0.3405 0.5168 0.3516 0.3967 0.2517 0
90 0 0.1413 0.3229 0.5359 0.337 0.391 0.272 0
120 0 0.1423 0.3213 0.5364 0.3218 0.3845 0.2937 0
150 0 0.1471 0.3159 0.537 0.304 0.3837 0.3123 0
180 0 0.1472 0.3091 0.5437 0.2833 0.3809 0.3357 0
210 0 0.1459 0.3049 0.5492 0.2641 0.3802 0.3558 0
240 0 0.1425 0.3064 0.5511 0.2485 0.386 0.3655 0
Appendix
111
Table (D-2.2) experiment No. 2:
Distillate Composition Reboiler Composition
Time (min)
Acetic acid
Butanol Butyl Acetate
Water Acetic acid
Butanol Butyl Acetate
Water
0 0 0.1589 0.3356 0.5055 0.3973 0.3845 0.2181 0
30 0 0.1267 0.3506 0.5228 0.3826 0.3812 0.2362 0
60 0 0.1309 0.3508 0.5183 0.3983 0.3717 0.23 0
90 0 0.1328 0.3439 0.5234 0.3962 0.3745 0.2293 0
120 0 0.1291 0.3474 0.5235 0.3944 0.3811 0.2244 0
150 0 0.1323 0.3439 0.5237 0.3892 0.389 0.2218 0
180 0 0.1424 0.3411 0.5165 0.3845 0.3968 0.2187 0
210 0 0.1408 0.3375 0.5217 0.3807 0.4036 0.2157 0
240 0 0.1406 0.3439 0.5155 0.3718 0.4147 0.2135 0
270 0 0.1371 0.3502 0.5127 0.3695 0.421 0.2095 0
Table (D-2.3) experiment No. 3:
Distillate Composition Reboiler Composition
Time (min)
Acetic acid
Butanol Butyl Acetate
Water Acetic acid
Butanol Butyl Acetate
Water
0 0 0.1537 0.3357 0.5107 0.3879 0.3811 0.231 0
60 0 0.1245 0.3486 0.527 0.3947 0.3677 0.2376 0
120 0 0.1287 0.3473 0.524 0.3929 0.3761 0.231 0
180 0 0.127 0.3474 0.5256 0.3833 0.3841 0.2326 0
240 0 0.1299 0.349 0.521 0.3762 0.3933 0.2305 0
Table (D-2.4) experiment No. 4:
Distillate Composition Reboiler Composition
Time (min)
Acetic acid
Butanol Butyl Acetate
Water Acetic acid
Butanol Butyl Acetate
Water
0 0 0.1527 0.3372 0.51 0.3902 0.383 0.2268 0
60 0 0.127 0.3431 0.5299 0.3923 0.3799 0.2277 0
120 0 0.1142 0.3551 0.5307 0.3896 0.3791 0.2313 0
180 0 0.1205 0.3537 0.5258 0.3779 0.3906 0.2315 0
240 0 0.1157 0.3567 0.5276 0.3654 0.4028 0.2318 0
Appendix
112
Table (D-2.5) experiment No. 5:
Distillate Composition Reboiler Composition
Time (min)
Acetic acid
Butanol Butyl Acetate
Water Acetic acid
Butanol Butyl Acetate
Water
0 0 0.131 0.3155 0.5536 0.3598 0.3559 0.2843 0
60 0 0.1104 0.3466 0.5429 0.3797 0.3487 0.2715 0
120 0 0.1145 0.3473 0.5382 0.3942 0.3405 0.2654 0
180 0 0.1084 0.3481 0.5435 0.3993 0.3216 0.279 0
Table (D-2.6) Temperature profile along the colome.
Column
Length(cm)
Temprature(co)
experiment
No. 1
experiment
No. 2
experiment
No. 3
experiment
No. 4
experiment
No. 5
0 119.7229291 112.2118467 109.24215 108.9360356 110.5117019
66 117.8982565 101.1251452 99.03267913 97.99722166 100.285308
107 115.8897038 94.88448185 93.11073193 93.47603845 93.68228449
139 111.8240169 94.53571167 92.7772192 94.44797982 94.06722409
154 105.9440544 92.89541312 91.72064203 93.50388333 92.95942262
174 94.62899446 94.86662776 94.11907893 97.43287758 97.82925316
194 89.71390683 89.90875418 88.86393712 90.00284441 89.78475975
249 88.69163964 88.25763373 87.27189077 88.49239751 87.94620093
الخالصة
في هذا البحث تم انتاج خالت البيوتيل بنجاح باستخدام برج التقطير التفاعلي المستمر مع .استخدام عامل مساعد صلب
:وقد أنجزت هذه الدراسة من خالل عدة مراحل هي كما يلي .مرحله دراسه ميكانيكيه التفاعل :أوال . مختبرية مرحله الدراسه العمليه في منظومه:ثانيا
.مرحله الموديل الرياضي :ثالثا تم دراسة ميكانيكيه تفاعل أألسترة إلنتاج خالت البيوتيل في مفاعل :المرحله االولى
ول الوجبة و باستخدام عامل مساعد غير متجانس مستعملين حامض الخليك و كحعمل نوعين من العوامل المساعدة ََ ََ ََ Dowex) البيوتانول كمواد متفاعلة وحيث اسَت
( Amberlite CG 50)و( 50, ( Hcl 0.1 N) وقد تم تطوير العوامل المساعدة باستعمال حامض الهايدروكلوريك
غير النتائج اظهرت بأن العامل المساعد المعدل أظهر فعالية أكبر من العامل المساعدية التفاعل بأن كما وأظهرت النتائج التي تم الحصول عليها من دراسة ميكانيك. المعدل
قد أعطى أعلى نسبة تحول للبيوتانول في مفاعل ( Dowex 50 )العامل المساعد )مستعملين( مول \كيلو جول 63...5) وبطاقة تنشيط قدرها % 76الوجبه و البالغة
Pseudo-Homogeneous Model ) حيث كانت المقارنه ,لوصف ميكانيكية التفاعل .بين النتائج النظريه و العمليه ذات توافق عالي جدا
ريادية لبرج تقطير تفاعلي مستمر مصنوع من زجاج مختبرية منظومة :المرحلة الثانية QVF نسبة الراجع وسرعة جريان حامض الخليك وكمية ) لدراسة تاثير التغير في
Dowex 50 (على أداء برج التقطير التفاعلي مستخدمين افضل عامل مساعد (الحرارة ( Modified. جزء جزء االنتزاع و لكل من( 9., 53, 73)تبلغ ارتفاعات البرج
و ( O.D=4.7 cm) يبلغ القطرالخارجي للبرج و,جزء التجزئه على التوالي التفاعل و cm (330)ويبلغ ارتفاع البرج الكلي ( I.D=4.4 cm) القطر الداخلي
لقد تم استنتاج انه عندما تزداد نسبة الراجع فإن مستوى درجات الحرارة على طول البرج الحرارة ذو و من ناحية اخرى فإن زيادة معدل جريان حامض الخليك أو زيادة كمية, يقل
.تأثير قليل على مستوى الحرارة في البرجلوحظ بأنه عند زيادة معدل جريان حامض الخليك فإن درجة حرارة منطقة التفاعل تزداد
. بينما عند زيادة كمية الحرارة فإن درجة حرارة منطقة التفاعل تقل , ايضا مستقرة باستعمال برنامج إشتقاق موديل رياضي للحالة الغير : المرحلة الثالثة
(MATLAB ) ,ستعمال طريقة حيث تم حل مجموعة المعادالت الجبرية با(Gausses
Elimination Method .) وقد تم استخدام الموديل بكفاءة لوصف مكونات وجرياناظهرت النتائج بأن نسبة التحول تزداد بصورة مباشرة مع , ودرجة الحرارة داخل البرج
ة الراجع و هذه النتيجة يمكن تفسيرها بزيادة زمن البقاء في منطقة التفاعل عند زيادة نسبحجم انه عند تقليل النسبة بين: من ناحية اخرى تم مالحظة .زيادة نسبة الراجع
الصواني و حجم المرجل فإن هذا يزيد من سرعة استجابة البرج و يقلل من الزمن الالزم . ة لكي يصل الى الحالة المستقر
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