Experimental and Dynamic Simulation of catalyzed ... · Certification I certify that this thesis...

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


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].


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].


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.


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].


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,


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].


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


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.


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.


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.


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


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].


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].


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-


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 ].


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.


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].


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).


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].


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].


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].


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].


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


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)


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.


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


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


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.


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.


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)


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


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).


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.


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.


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.


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.


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


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.


42

Fig (3.3) View of experimental continuous reactive distillation unit.


43

Fig (3.4) Schematic diagram of a packed continuous reactive distillation unit.


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).


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


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


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.


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)


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.


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


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


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


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)


(4.10)

b) Component material balance:

(4.11)

c) Energy balance:

(4.12)


55

d) Summation:

(4.13)

III. Reactive tray: (stage k)


(4.14)


(4.15)

c) Energy balance:

(4.16)

d) Summation:

(4.17)

IV. Condenser: (stageN)


(4.18)


56


(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.


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


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)


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.


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.


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.


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 %.


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).


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).


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


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.


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.


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


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


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


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)


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).


74


(b) Reboiler composition




75


(b)Reboiler composition


100gm and (440 W) heat duty.


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.


77






78



Fig (5.16) Distillate and Reboiler composition for (reflux ratio =2, ACH:BuOH=2:1,catalyst wt



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).




80






81



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


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


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)


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:


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).


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


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).


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


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.


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)


0.41

0.42

0.43

0.44

0.45

0.46

0.47

10 11 12 13 14

Wat

er

Time )hr)


362.6

362.8

363

363.2

363.4

10 11 12 13 14

Te

mp

era

ture

( K

)

Time )hr)



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)


0.25

0.26

0.27

0.28

0.29

0.3

10 11 12 13 14

Bu

tyl A

ceta

te

Time )hr)


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)


362

363

364

365

366

367

10 11 12 13 14

Tem

pe

ratu

re (

K )

Time )hr)



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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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.


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



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



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



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



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


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 .


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


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



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



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



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



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



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



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



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


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



Time (min)

Acetic acid


Water Acetic acid


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



Time (min)

Acetic acid


Water Acetic acid


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



Time (min)

Acetic acid


Water Acetic acid


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



Time (min)

Acetic acid


Water Acetic acid


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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