KINGDOM OF SAUDI ARABIA KING SAUD UNIVERSITY

COLLEGE OF ENGINEERING DEPARTMENT OF CHEMICAL ENGINEERING

MODELING AND SIMULATION OF A REACTIVE DISTILLATION UNIT FOR PRODUCTION OF MTBE

ميثل ثالثي بيوتيل اإليثر نمذجة ومحاآاة وحدة تقطير متفاعل إلنتاج مادة

SUBMITTED BY

FAHAD S. AL-HARTHI 424121504

SUPERVISED BY

PROF. DR. AHMED E. ABASAEED PROF. DR. IBRAHIM S. AL-MUTAZ

A Thesis submitted in partial fulfillment of the requirements for the degree of Master of Science in Chemical Engineering

2nd Semester 1428-1429 AH

JUNE 2008

KINGDOM OF SAUDI ARABIA KING SAUD UNIVERSITY COLLEGE OF ENGINEERING DEPARTMENT OF CHEMICAL ENGINEERING

MODELING AND SIMULATION OF A REACTIVE DISTILLATION UNIT FOR PRODUCTION OF MTBE

Submitted by:

Eng. Fahad Al-Harthi 424121504

Supervised by:

Prof. Dr. Ahmed E. Abasaeed Prof. Dr. Ibrahim S. Al-Mutaz

Examination Committee Members

Prof. Ahmed E. Abasaeed

Prof. Ibrahim S. Al-Mutaz

Prof. Kamil M. Wagialla

Prof. Mohammed Asif

Prof. Khalid I. Al-Humaizi

A thesis submitted in partial fulfillment of the requirements for the degree of Master of Science in Chemical Engineering

2nd Semester 1428-1429 AH

JUNE 2008

السعودية العربية اململكة سعود امللك جامعة

الكيميائية اهلندسة قسم – اهلندسة كلية

ميثل ثالثي  إلنتاج مادة نمذجة ومحاآاة وحدة تقطير متفاعل )MTBE( بيوتيل اإليثر

: إعداد

الحارثي سفر بن فهد / المهندس : من آل إشراف

الحاج سعيد أبا أحمد : الدآتور األستاذ و

المعتاز صالح بن إبراهيم : الدآتور األستاذ

رسالة مقدمة الستكمال متطلبات درجة الماجستير

في الهندسة الكيميائية هـ ١٤٢٩جمادى الثانية

June 2008

i

TABLE OF CONTENTS Table of contents …………………………………………………………………………………….. i

List of Figures…………………………………………………………………………………………. iii

List of Tables……………………………………………………………………………………….…. v

Nomenclature ………………………………………………………………………………………… vi

Dedication …………………………………………………………………………………………… viii

Acknowledgments……………………………………………………………………………………. ix

Abstract …….…………………………………………………………………………………………. xi

CHAPTER 1 : INTRODUCTION ………………………………………………………………….. 1

1.1 Properties of MTBE……………………………………………………………………… 2

1.2 Chemistry of MTBE……………………………………………………………………… 3

1.3 Demand of MTBE……………………………………………………………………….. 3

1.4 Commercial Technology of MTBE…………………………………………………….. 5

(a) lsobutylene from Steam Crackers…………………………………………………. 6

(b) Isobutylene from Refineries………………………………………………………… 6

(c) Isobutylene by dehydrogenation lsobutane……………………………………….. 7

(d) Isobutylene from TBA……………………………………………………………….. 7

1.5 Introduction to Conventional Process…………….………………………………..…. 8

1.6 Concept of Reactive Distillation Process…………………………………………….. 9

1.7 Production of MTBE by Reactive Distillation……….………………………………… 10

1.7.1 ABB LUMMUS Process...……………………………..………………. 11

1.7.2 UOP Process ……………………………………………..………….... 12

1.7.3 Huntsman TBA/PO Process……………………………..……………. 14

1.8 Thesis outline and Work Methodology……………………………………………….. 15

1.8.1 Research Outline…………………………………………………………………15

1.8.2 Research Methodology……………..………...…………………………………15

CHAPTER 2 : REACTIVE DISTILLATION

2.1 Introduction to Reactive Distillation……………………………………………………. 17

2.2 Basics and Advantages of Reactive Distillation…………………………………..…. 17

2.2.1 Importance of RD……………………………………………………………… 18

2.2.2 Constraints and Difficulties in RD implementation……………………….... 23

2.2.3 The complexity of RD…………………………………………………………. 24

2.2.4 Model Available for design of RD columns……………………………….... 25

ii

CHAPTER 3 : LITERATURE SURVEY

3.1 Overall Review……………………………..………………………………….……….... 28

3.2 Computational Methods…………………………….….……………………….……... 30

CHAPTER 4 : MODELING & SIMULATION OF MTBE PROCESS

4.1 MTBE Production by PRO/II …………………………………………………………… 40

4.1.1 Introduction ……………………………………………………………………… 40

4.1.2 Process description …………………………………………………………….. 40

• Catalyst used……………………………………………………………………43

• MTBE Recovery section……………………………………………………….43

• Methanol Recovery section……………………………………………………44

4.2 Modeling and Simulation of MTBE RD unit………….……………………………...... 45

4.3 Thermodynamics Data….………………………..…….…………………………........ 46

4.4 MTBE Reaction Kinetics……………..……………………………………………........ 48

4.5 Modeling Procedure………..………………………………………………………........ 51

4.5.1 Model Equations for a single Stage…..……..…………………………......... 53

4.6 Modeling and simulation framework………………………….………………..…........ 54

CHAPTER 5 : RESULTS AND DISCUSSION

5.1 Results of the simulation work…………….…………...……………..…….................. 57

5.2 Results discussion………………………….…………...……………..……................... 62

5.3 Cases studies (Optimization)…..…..….…………...……………..…………….…........ 62

CHAPTER 6 : CONCLUSIONS AND RECOMMENDATIONS

6.1 Conclusions…………………………………………………………………………..… 69

6.2 Recommendations …..…..…………….……………………………………………… 71

REFERENCES……………………………………..…………………………………………..….. 73 APPENDIXES APP (A) : Reaction model code. …………………………………………………………..…… 76

APP (B) : MTBE Process Simulation Procedure using PRO/II software. ………………….. 80

APP (C) : Process Input-File used in PRO/II software. ………………………………..…….. 98

iii

LIST OF FIGURES

Figure No Page

Figure 1.1 : Publications and Patents on Reactive Distillation (1973-2003)……………………….….... 1

Figure 1.2 : MTBE Global Demand Distribution……………………………………………………………. 5

Figure 1.3 : MTBE Production Processes……………...…………………………………….…………….. 6

Figure 1.4 : MTBE Production via conventional Processes …………………………………….............. 9

Figure 1.5 : CDTECH Process Flow Diagram……………………………………………….……............. 12

Figure 1.6 : ETHERMAX MTBE Process Flow Diagram…………………………………........…………. 13

Figure: 2.1 Simple Reactive Distillation sketch…………………………….........................……............ 18

Figure 2.2: Processing schemes for a reaction sequence A +B <==> C+ D …….………....…………. 19

Figure 2.3: Processing schemes for the esterification reaction MeOH + AcOH<==>MeOAc + H2O… 20

Figure 2.4: (a) Reactive distillation concept for synthesis of MTBE from the acid catalysed reaction between MeOH and iso-butene. The butene feed is a mixture of reactive iso- butene and non-reactive n-butene. (b) Reactive distillation concept for the hydration of ethylene oxide to ethylene glycol. (c) Reactive distillation concept for reaction between benzene and propene to form cumene. (d) Reactive distillation concept for reaction production of propylene oxide from propylene chlorohydrin and lime. The reactive sections are indicated by grid line...……………………………………………… 21

Figure 2.5: Hydrodesulphurisation of gas oil carried out in (a) co-current trickle bed reactor and (b) counter-current RD unit…………………………..…………………………………………...………………. 22

Figure 2.6: Transport processes in RD. (a) homogeneous liquid phase reaction, and (b) heterogeneous catalyzed reactions………………………………………………………...…..………..…………….……..… 25

Figure 2.7 : Length and time scales in RD……………..……………...………..……………………..……. 25

Figure 3.1: Ternary LLE diagram for Acetone-water-MIBK system at 5atm. …………………………..…33

Figure 3.2: RD Column Configuration……………..…..………..………………………………………...... 34

Figure 3.3: Liquid Composition profile in RD Column..………..……………………...……...…………….. 34

Figure 3.4 Nonreactive residue curve map for the system IBTE+ MeOH ↔ MTBE at 101.32 kPa…… 36

Figure 3.5 Reactive residue curve map for the system IBTE+ MeOH ↔ MTBE at 101.32 kPa………. 37

Figure 3.6 Reactive phase diagram for the system IBTE/ MeOH /MTBE at 101.32 kPa………………. 37

Figure 3.7 Reactive phase diagram for the system IBTE/ MeOH /MTBE at 810.56 kPa………………. 38

Figure 4.1: MTBE Plant Flowsheet.………………………………..……………...…..……......……………..41

Figure 4.2: Thermodynamic Data of PRO/II software…………..………………………..…...……………..47

Figure 4.3: A schematic representation of a catalytic distillation column………..………....……………..52

iv

Figure 4.4: Representation of a generic equilibrium stage………...……………..……….....……………..52

Figure 4.5: MTBE Process Flowsheet PRO/II………………………………..…..………….....…………….55

Figure 5.1: Feed and products concentration profile among column trays...……......……………57

Figure 5.2: Temp Profile of column T-1………………………………..…………..…....…......…………….58

Figure 5.3: Rates Profile of column T-1……………………………………………..………………………. 58

Figure 5.4: Density Profile of column T-1…………………………………………..………….…………….59

Figure 5.5: Viscosity Profile of column T-1………………………………………..………..………………..59

Figure 5.6: Temp Profile of column T-2…………………………………………..………..…..…………….59

Figure 5.7: Rates Profile of column T-2……………………………………..……………………………….59

Figure 5.8: Density Profile of column T-2…………………………………..………………….……………..60

Figure 5.9: Viscosity Profile of column T-2………………………………..………………..………………. 60

Figure 5.10: Temp Profile of column T-3…………………………………..……………..………………….60

Figure 5.11: Rates Profile of column T-3…………………………………..……………….………………..60

Figure 5.12: Density Profile of column T-3………………………………..…………………………..……..61

Figure 5.13: Viscosity Profile of column T-3………………………………..…………...…..……………….61

Figure 5.14: Results of changing the MeOH/IBTE ratio……………..……...…..…………………………63

Figure 5.15: Results of changing the number of trays of RD column………...…..………………………64

Figure 5.16: Results of changing the location of feed stream into RD column …………………………65

Figure 5.17: Results of changing the Reflux Ratio…………..…………...…..…………………………….66

Figure 5.18 Results of changing the location of reaction zone .…………...…..…………………………..67

v

LIST OF TABLES Table page

Table 1.1: Typical Properties of Oxygenates……………………………………………………............... 2

Table 1.2 : MTBE competitive strengths and weaknesses……………………………………................ 3

Table 1.3 : Global MTBE demand (thousand tons per year) …………………………………................ 5

Table 1.4: Typical isomer distribution in C4 product stream (volume %)……………………................. 7

Table 4.1: Reactor Feed Data………..……………….…………………………………………................. 42

Table 4.2: Reaction Stoichiometry………………………………………………………..…….................. 43

Table 4.3: Approximate Catalyst Properties of Amberlyst 15……………………….….……................ 43

Table 4.4: Binary Interaction Parameters for SRKM_VLE…………………………………..…............... 46

Table 5.1: Summary Table…………………………………………………………………..……............... 57

Table 5.2: Summary of MTBE column T-1…………………………………….………….….…................. 58

Table 5.3: Summary of water wash column T-2………………………………………….….…................. 59

Table 5.4: Summary of Methanol Recovery column T-3…………………………..….……….................. 60

Table 5.5: Overall Report…………………………………..…………………………………..….................. 61

vi

NOMENCLATURE Notations

a,b,c Order of reaction of species A,B and C respectively

∆HA Heat of adsorption of methanol, J/mol

∆HC Heat of adsorption of MTBE, J/mol

aij The mixing rule

Xi Mole fraction of component i

Xj Mole fraction of component j

Kij Binary interaction parameter

Ks Surface reaction rate constant, (gmole/g catalyst)

KA Equilibrium adsorption constant for A, g cat/gmole

KC Equilibrium adsorption constant for C, g cat/gmole

K Equilibrium constant for the overall reaction

KAO , KCO, KSO Preexponential factor

rs Rate of surface reaction, (mol/gcat)/h

CA IBTE concentration, mole/l

CB MEOH concentration, mole/l

CC MTBE concentration, mole/l

R Gas constant, 8.314 J/mol.K

T Temperature, K

E1 Activation Energy of the forward reaction, J/mol

KX Mole fraction Equilibrium ratio

Kγ Ratio of activity coeffiecients at equilibrium

Xk,i The mole fraction of liquid phase for species j in stage k.

Yk,i The mole fraction of vapor phase for species i leaving stage k

Lk The molar liquid flow rate leaving stage k.

Vk The molar vapor flow rate leaving stage k.

Fk The feed flow rate to stage k.

hk The molar enthalpy of liquid phase in stage k.

HK The molar enthalpy of vapor phase in stage k.

Hjk The heat of reaction in stage k.

Qk The heat duty in stage k.

γk,i The activity coefficient of component i in stage k.

Kk,i The vapor-liquid equilibrium constant.

ρk The molar density of the liquid on stage k.

vk The volumetric hold-up of liquid on stage k

P Pressure, bar

NC Number of components

NT Number of trays

vii

u,w Constants, typically integers

Subscripts:

i Tray index

j Component index

Superscripts: F refers to a feed

D refers to a draw

L refers to a liquid property

V refers to a vapor property

viii

DEDICATION To my mother, who will be always my inspiration. This work is also dedicated to the ones that encourage me, my wife , my daughter Lujain, and my brothers.

ix

ACKNOWLEDGMENT

Firstly, I thank Allah almighty for His blessings and providing me with the

strength and power to succeed.

This research compiles the upmost results of a 2-year period of intense and

challenging work. Independently on the nature of their contribution (intellectual or/and

emotional), many people have pushed toward the completion of this work. To all of

them I am deeply and sincerely indebted.

My first words of thanks go, sincerely, to my promoter, Prof. Ibrahim Al-Mutaz.

I thank him for his continuous guidance and scientific insight during these 2+ years

and, on top of this, for his admirable kindness and politeness. I fully enjoyed the long

hours we spent together, discussing a wide range of topics. Also. I would like to thank

my advisor Prof. Ahmed Abasaeed for his support and I am very much thankful for

allocating time in your tight agenda to discuss and repeatedly correct all the thesis

chapters. Really, Without all your comments and suggestions this thesis probably

would have remained as a bunch of unconnected manuscripts.

Additionally, I would like to thank all the people who supported my research

and me during the past two years. Foremost have been my wife who has always

encouraged me to work hard and to continue this work.

Moreover, my thanks also extended to my family members, relatives and

colleagues for supporting me throughout my research period.

x

א

:مقدمةأبرز المواد من (MTBE) المعروفة عالميًا باالسم المختصر) مثيل ثالثي بيوتيل اإليثر(تعتبر مادة

آتيني والحد من األآسجينية التي تتمتع بخصائص فيزيائية ممـتازة عند إضافتها إلى الوقـود لزيادة العدد األو

آبير جدًا في الحد من االنبعاثات الضارة الملوثة للبيئة من غازات ودورها )مقارنة بالرصاص(التلوث البيئي

وتؤدي في نفس الوقت إلى احتراق الوقود بشكل أفضل وأوفر من مادة الرصاص وإعطاء طاقة عادم السيارات

باستخدام ثالث MTBEمادة حويل لقيم البيوتان والميثانول إلىوتعد أفضل طريقة إلنتاجه باستخدام تقنية ت. أعلى

الجزئية للبيوتان وإزالة الهيدروجين من االيزوبيوتان وتفاعل االيزوبيوتان مع عمليات وهي تغيير البنية

.الميثانول

يًا عبر السنوات األخيرة للتوافق مع تصاعد استخدامها عالم MTBEوقد تنامت الطاقه اإلنتاجية لمادة

سالمة وصحة البيئة، وقد أثبتت المادة نجاحها في هذا المجال، حيث استجابة لألنظمة المنادية بالمحافظة على

.يجري استخدامها في أوروبا والعديد من األسواق العالمية األخرى بصورٍة طبيعية وآمنة

االيزو بيوتين آمواد متفاعلة في الطريقة التقليدية عبر إضافة مادة الميثانول إلى MTBEيتم إنتاج مادة

ثم يتم إمـرار ناتج التفاعـل إلى . درجة مئوية 60-45في مفاعل بوجود مادة حفازة عند درجة حرارة تتراوح بين

واستعادة المواد التي لم تتفاعل ليتم استخدامها مره (MTBE)عـدد من أبراج التقطيـر ليتم فصل المادة المنتجـة

.أخرى

م 1980وهي المتبعة حديثًا ويعود تاريخ هذه الطريقة إلى عام MTBEرى إلنتاج الـ هناك طريقة أخ

عبر هذه الطريقة وتتلخص في إنتاج وفصل الـ MTBEحيث سجل العالم سميث أول براءة اختراع إلنتاج الـ

MTBE آنياً ، وتسمى هذه الطريقة بـعملية التقطير المتفاعلReactive Distillation Process وتمتاز ،

خفض تكاليف إنشاء المصنع وذلك بتقليل عدد المعدات واألجهزة، : هذه الطريقة بعدة نواحي تقنية واقتصادية مثل

المرغوبة، زيادة آفاءه التفاعل والفصل مما يؤدي إلى زيادة معدل اإلنتاج تقليل نسبة إنتاج المواد الجانبية غير

.واالستفادة من الحرارة الناتجة من التفاعل في برج التقطير حيث تساعد في عملية الفصل

سوف يتم في هذا البحث عمل مسح مرجعي مكثف لألبحاث السابقة في هذا المجال وذلك بهدف دراسة

بطريقة التقطير المتفاعل وسيستخدم هذا MTBEبناء نموذج رياضي يحاآي إنتاج مادة هذه الطريقة ومن ثم

.النموذج في دراسة تأثير العوامل التشغيلية والتصميمية المختلفة

أوضحت بان استخدام عملية PRO/IIنتائج عملية المحاآاة التي تم الحصول عليها باستخدام برنامج

وبنقاوة % 99.2بنسبة عالية جداً تصل الى MTBEالى االيزو بيوتينة تحويل التقطير المتفاعل يزيد من نسب

%.99.7عالية ايضاً تصل الى

xi

ABSTRACT

Methyl tertiary butyl ether (MTBE) is primarily used in gasoline blending as an octane

enhancer to improve hydrocarbon combustion efficiency. Of all the oxygenates, MTBE is

attractive for a variety of technical reasons. It has a low vapor pressure. It can be blended with

other fuels without phase separation. It has the desirable octane characteristics.

MTBE is produced via direct addition of methanol to isobutylene using sulphonated ion

exchange resin as catalysts. There are two ways to produce MTBE, one is the conventional

process which is mainly a reactor and separate distillation column with conversion range 87-

92%. Another method for the production of MTBE is newly established and date back to the

way in 1980 as the scientist Smith recorded the first patent for the production of MTBE

through this method, this method called Reactive Distillation Process, and there are a lot of

features that makes this process attractive and practical with a conversion reached 99.2%.

The main objectives from this research are developing a mathematical model for the

MTBE production via reactive distillation column at steady state, coding and embedding the

model into simulation software (PRO/II) and using the developed model to assess the effect of

some critical design and operating parameters on column performance and doing some case

studies.

PRO/II software has the capabilities of solving Reactive Distillation Processes

utilizing Chemdist Algorithm provided by the software, but it requires knowledge of Material

and Energy balances equations as well as Thermodynamics and kinetics of reaction.

In this work, we have made an extensive literature survey for the previous research

done in this area to study this technology and then we have developed a mathematical model

that simulates the production of MTBE in a reactive distillation unit and that model was used

to study the impact of operational factors and different design parameters by doing several case

studies "Optimization".

The main conclusion of this work is that the higher conversion of IBTE to MTBE can

be obtained by applying Reactive Distillation approach, we have obtained 99.2 % IBTE

conversion and high selectivity for MTBE with 99.7%. Moreover, the optimum

Methanol/Isobutylene ratio was found equal to 1.0 and the optimum number of trays for

Reactive Distillation column is 30 trays.

CHAPTER -1-

INTRODUCTION

- 1 -

CHAPTER -1-

INTRODUCTION

Methyl tertiary butyl ether (MTBE) is primarily used in gasoline blending as an

octane enhancer to improve hydrocarbon combustion efficiency. Of all the

oxygenates, MTBE is attractive for a variety of technical reasons. It has a low vapor

pressure. It can be blended with other fuels without phase separation. It has the

desirable octane characteristics and is becoming increasingly important as stricter air

pollution control measures are implemented.

MTBE can be produced by addition of methanol to isobutylene in the liquid

phase over an acidic catalyst consisting of sulfonated macroporous ion exchange

resins.

In 1980, Smith registered a patent to process MTBE for Chemical Research

and Licensing Company. He used a reactive distillation system, containing catalytic

packing. The pilot plant was 3 inches in diameter and it was used to predict the

operation variables of a large commercial plant [1].

Due to its potential for improved process design and the success of its

commercial applications, reactive distillation gained the interest of both academics

and industry toward its use and application in commercial processes. Figures 1.1

show the growing interest in reactive distillation in recent years. These data were

compiled from the ACS databases CAPLUS, CHEMCATS and CHEMLIST. Some 400

publications were taken into consideration for the period of 1973 to May 2003 [2].

Figure 1.1 Publications and Patents on RD -Historical Trends (1973-2003).

- 2 -

1.1 PROPERTIES OF MTBE

Oxygenates are hydrocarbons that contain one or more oxygen atoms.

The primary oxygenates are alcohols and ethers, including: fuel ethanol, methyl

tertiary butyl ether (MTBE), ethyl tertiary butyl ether (ETBE), and tertiary amyl

methyl ether (TAME). The physical properties for these components are shown

in Table 1.1.

The 1977 Clean Air Act amendments set requirements for "substantially

similar gasoline," which requires that oxygenates be approved by the U.S. EPA

before they are allowed to be used in gasoline. In 1981 the EPA allowed the

blending of MTBE up to 11 volume %, and extended the limit to 15 volume % in

1988.

Oxygenates are added to motor vehicle fuels to make them burn more

cleanly, thereby reducing toxic tailpipe pollution, particularly carbon monoxide.

Oxygenates are favored not only for their vehicle emission benefits but also

their blending properties in motor gasoline (e.g., octane).

Table 1.1: Typical Properties of Oxygenates.

Ethanol MTBE ETBE TAME

Chemical formula CH3CH2OH CH3OC(CH3)3 CH3CH2OC(CH3)3 (CH)3CCH2OCH3 Oxygen content, % by weight 34.73 18.15 15.66 15.66

Octane, (R+M)/2 115 110 111 105 Blending vapor pressure, RVP 18 8 4 1.5

Source: National Petroleum Council, U.S. Petroleum Refining: Meeting Requirements for Cleaner Fuels and Refineries (Washington, DC, August 1993) Appendix L.

There is an increasing demand for MTBE as a gasoline additive.

Currently, the worldwide consumption of MTBE reached 6.6 billion gallons of

which 65% is consumed in the United States [3].

Methyl tertiary-butyl ether is an excellent gasoline blending component

because it is a high octane enhancer and has low volatility. It has blending

properties similar to gasoline’s, but do not exhibit undesirable properties such

as azeoptrope formation, water pick-up, or phase separation. Gasoline

- 3 -

containing MTBE is accepted for transport on common carrier pipelines in the

United States and by major and independent refiners.

Table 1.2 summarizes the relative strengths and weaknesses of MTBE

compared to other octane enhancers. MTBE is a widely accepted octane

enhancer and enjoys several significant strengths compared to competing

octane enhancers. Its major weakness is that one of its key feedstocks,

isobutylene, is limited by petrochemical and refinery operations.

Table 1.2 : MTBE competitive strengths and weaknesses

Although MTBE has a density less than that of water, when dissolved in

water, MTBE responds to localized groundwater gradients and aquifer

recharge. A growing number of studies have detected MTBE in ground water in

United State; in some instances these contaminated waters are sources of

drinking water. Low levels of MTBE can make drinking water supplies

undrinkable due to its offensive taste and odor. Therefore, the jurisdiction of the

U.S. Department of Transportation (DOT) has set some restricted Regulation to

minimize releases from gasoline pipelines or another potential source of leaks.

MTBE is believed to have entered the water supply from leaking

underground storage tank systems including underground lines that contained

gasoline and/or from surface spills at gas stations. Once in the ground, MTBE

behaves differently from other gasoline constituents such as benzene . Unlike

petroleum hydrocarbons, it is highly water soluble, not easily absorbed into soil,

and is more resistant to biodegradation. Thus, with widespread use, MTBE

has the potential to occur in high concentrations in groundwater, travel far from

leak sources, and accumulate to become a concern for the entire region.

Strengths Weaknesses - High octane - Low volatility - Blending characteristics similar to gasoline - Widely accepted in marketplace by

consumers and refiners - Reduces carbon monoxide and exhaust - hydrocarbon emissions

- Availability of economical isobutylene feedstocks is limited

- Possible methanol supply constraints

- Health hazard

- 4 -

1.2 CHEMISTRY OF MTBE

Methyl tertiary-butyl ether (MTBE) can be formed by the addition of

methyl alcohol to the highly reactive double bond in isobutylene, as shown in

the following equation:

Methanol Isobutylene MTBE

MTBE synthesis occurs in the liquid phase at 40°C-100°C and 100-150

psig as an exothermic reaction (∆HR= -16,060 Btu/lb-mole). In the presence of

a small amount of acidic cation exchange resin catalyst the reaction proceeds

quantitatively. Indeed, there are few reactions in industrial chemistry that

demonstrate such high selectivity. Several processes, as will be discussed

later, have been devised for carrying out this reaction efficiently on the

isobutylene contained in the so-called raffinate-I that results when butadiene is

removed from the C4 fraction obtained either from steam or catalytic cracking.

The C4 fraction from the former source is preferred because of its higher

isobutylene concentration. This fraction contains, in addition to isobutylene,

butene-I, and butane-2. The removal of isobutylene from raffinate-I provides a

mixture of the normal butenes known as raffinate-Il [2]. 1.3 DEMAND OF MTBE

The importance of MTBE is increasing from year to year since the global

demand for MTBE is expected to grow at 4.0% annually from 1994 to 2010, as

shown in Table I.3. In the 2000-2010 period, however, growth will slow to 1.7 %

annually from 8.1 % from 1994-2000 as shown in Fig 1.3 [4].

( ) ( ) 3332233 CHOCCHCHCCHOHCH −−⇔=+−

- 5 -

Table 1.3 : Global MTBE demand (thousand tons per year). Adapted from CHEM SYSTEM.

Figure 1.2 : MTBE Global Demand Distribution

1-4 COMMERCIAL TECHNOLOGIES OF MTBE

Methyl tertiary-butyl ether (MTBE) is produced through the reaction of

isobutylene with methanol. The MTBE reaction is equilibrium limited. Higher

temperatures increase the reaction rate, but the conversion level is lower.

Lower temperatures shift the equilibrium toward ether production, but more

catalyst inventory is required. Therefore, conventional MTBE units are

designed with two reactors in series. Most of the etherification reaction is

achieved at an elevated temperature in the first reactor and then finished at a

thermodynamically favorable lower temperature in the second reactor [2].

Figure 1.3 shows the different routes to obtain IC4 for MTBE production, which

vary largely due to the different isobutylene feedstock sources. As shown in the

figure, isobutylene can be produced as a:

• Byproduct of ethylene manufacture in steam cracker units

- 6 -

• Byproduct of fluid catalytic cracking operations.

• Main product of an isobutane dehydrogenation unit

• Coproduct (via tertiary-butyl alcohol-TBA) of propylene oxide manufacture,

as operated by ARCO Chemical and Texaco Chemical Company

Figure 1.3 : MTBE Production Processes

These various supply sources for isobutylene are briefly reviewed below.

(a) lsobutylene from Steam Crackers The mixed butylene stream produced from ethylene plants has a

relatively high isobutylene content. Depending on the feedstock, operating

conditions, plant configuration, and other factors, the isobutylene content can

vary from 35 to 50 vol%, with about 44 % considered typical. One benefit of

MTBE manufacture from steam cracker sourced mixed butylenes is that it

provides an easy way to separate butylene isomers [2].

(b) Isobutylene from Refineries Isobutylene is also produced as a byproduct of fluid catalytic cracking

(FCC) operations in refineries. The isobutylene concentration is much lower in

this mixed butylene stream compared to the steam cracker mixed C4s stream.

Table 1.4 shows the relative butane/butylene isomer distribution for typical FCC

and steam cracker C4 streams [2].

The higher isobutylene concentration in steam cracker mixed butylene

streams reduces the capital and operating costs compared to feeding FCC

mixed butylenes.

- 7 -

Table 1.4: Typical isomer distribution in C4 product stream (volume %)

(c) Isobutylene by dehydrogenation lsobutane One possibility in alleviating the potential problem of insufficient

isobutylene is to dehydrogenate isobutane. lsobutane may, in turn, be made by

the isomerization of n-butane, a reaction well established, for isobutane is

prepared in large volumes for alkylation reactions in the refinery. Another

approach involves the skeletal isomerization of butene-1.

These reactions are shown in the following equations:

(d) Isobutylene from TBA High capital costs are involved in the isomerization and dehydrogenation

reactions. tertiary-Butyl alcohol (TBA), made by ARCO and Texaco as a

coproduct of propylene oxide production, is another source of isobutylene. This

involves dehydration to isobutylene, as shown in the following equation [2].

- 8 -

A number of companies have technology for MTBE manufacturing: all

the routes are based on liquid phase etherification of isobutylene with

methanol. Selected technology holders and licensers are: ARCO, CDTECH

(ABB Lummus Crest and Chemical Research & Licensing), Snamprogetti,

Phillips Petroleum, HuelslUOP, Sumitomo Chemical, Erdoelchemie, IFP.

Description of some commercial conventional Processes are given below:

1.5 A TYPICAL CONVENTIONAL PROCESS

As mentioned in the previous section, MTBE is produced via direct

addition of methanol to isobutylene using sulphonated ion exchange resin as

catalysts. The technology features a two-stage reactor system of which the first

reactor is operated in a recycle mode. With this method, a slight expansion of

the catalyst bed is achieved which ensures very uniform concentration profiles

within the reactor and, most importantly, avoids hot spot formation. Undesired

side reactions, such as the formation of dimethyl ether (DME), are minimized.

The reactor inlet temperature ranges from 45°C at start-of-run to about 60°C at

end-of-run conditions. The catalyst used in this process is a cation-exchange

resin. Isobutylene conversions of 97% are typical for FCC feedstocks. Higher

conversions are attainable when processing steam-cracker C4 cuts that

contain isobutylene concentrations of 25% as illustrated in figure 1.4 [2].

MTBE is recovered as the bottoms product of a distillation unit. The

methanol-rich C4 distillate is sent to the methanol-recovery section. Water is

used to extract excess methanol and recycle it back to process. The isobutylene-

depleted C4 stream may be sent to a raffinate stripper or to a molsieve-based

unit to remove other oxygenates such as DME, MTBE, methanol and tert-

butanol.

- 9 -

Figure 1.4 : MTBE Production via Conventional Processes

An alternative to the conventional process is Reactive Distillation process.

1.6 CONCEPT OF REACTIVE DISTILLATION PROCESS

Reactive distillation is a process where simultaneous chemical reaction

and vapor-liquid phase separation take place in the presence of a

heterogeneous catalyst. This represents an exciting alternative to traditional

liquid phase chemical reaction processing. In reactive distillation separation of

product(s) from unconverted reactants allows for greater conversion, because

product removal displaces equilibrium and forces the reaction to completion [4].

Both reaction and distillation acting simultaneously offers certain

advantages that cannot be matched by conventional processing. The

advantages of reactive distillation are specific to each system. However, the

followings are some of the significant advantages of reactive distillation:[4]

- 10 -

• Effecting distillation and reaction simultaneously reduces the capital

costs and includes benefits such as reduction of recycle, optimization of

separation, lower requirements of pumps, instrumentation and piping.

• An equilibrium reaction can be driven to completion by separation of the

products from the reacting mixture.

• Elimination of possible side reactions by removal of the products from

the reaction zone. This can serve to increase selectivity.

• Savings associated with energy costs, through use of the energy

released by exothermic reactions for vaporization. This reduces the

reboiler heat duty for boil-up that is supplied normally by steam.

• Non-reactive azeotropes may disappear under reactive distillation

conditions.

• Improved materials use. With reactive distillation, material proper usage

is attained through removal of the reaction product(s) from the reaction

phase. Also, elimination of by-products formation may allow use of

lesser quantities of reactants.

• Reduction of hot spots, because the liquid vaporization provides a sink

for thermal energy.

Synthesis of chemicals through reactive distillation has been mainly

applied to processes such as esterifications, hydrolysis reactions, trans-

esterifications, and etherifications. A detailed list of reactions where reactive

distillation is advantageous is cited by Doherty and Malone in 2001 [5]. While

the use of reactive distillation is well established in the chemical process

industry, very little is known regarding the importance of mass transfer

resistances in limiting catalyst effectiveness in MTBE synthesis [6].

1.7 PRODUCTION OF MTBE BY REACTIVE DISTILLATION UNIT

Most of the new commercial MTBE plants utilize reactive distillation

technology in which chemical reaction and fractionation of products are

combined into a single unit operation.

- 11 -

In the production of MTBE in reactive distillation process, a pre-reactor

and a reactive distillation column replaced the conventional process. The

reactive column completes the exothermic reaction and simultaneously

conducts the separation. It is also possible to incorporate the pre-reactor into

the reactive column thereby lowering capital costs and improving energy

integration [7]. The reaction between methanol and isobutylene to form MTBE

is equilibrium limited and the feed to a conventional reactor requires a relatively

low methanol/isobutylene feed ratio in the feed.

Maintaining the methanol/isobutylene ratio below 1.05 allows recovery of

the product MTBE as the bottoms product from a distillation tower with the

overhead product being at the azeotropic composition of methanol and

isobutylene. Feeding a higher methanol/isobutylene ratio results in recovery of

the MTBE from the bottoms at the methanol/MTBE azeotropic composition

[8,9]. Use of reactive distillation allows more economical recovery of the MTBE

product. Typical Reactive Distillation processes are:

1.7.1 ABB LUMMUS Process

The process of production of MTBE using ABB LUMMUS process

consists of three basic steps as listed below [2]:

• Isomerization of n-butanes to isobutane.

• Dehydrogenation of isobutane to isobutylene (CATOFIN Process).

• Etherification of isobutylene with methanol to MTBE (CDTech

Process). MTBE is formed by the catalytic etherification of isobutylene with

methanol, as shown in Figure 1.5 Isobutylene-rich feedstock from the

dehydrogenation unit is combined with fresh methanol from offsite storage and

recycle methanol from the methanol recovery of the MTBE unit and fed to the

primary fixed bed reactor. The CDTech process is based on a two step reactor

design, consisting of a fixed bed reaction followed by final conversion in a

catalytic distillation column [2].

After reaction in the primary reactor, the partly reacted mixture is cooled

to the bubble point and pumped to the catalytic distillation reaction column that

- 12 -

combines reaction and fractionation sections. The reaction is continued in the

column and the MTBE product is separated from the un-reacted C4s. The

column allows a high conversion (99 %) of isobutylene.

Methanol and C4s form a minimum boiling azeotrope, so that the

methanol is carried into the catalyst packing where the reaction proceeds. As

MTBE is formed, it is removed from the reaction zone by distillation, which

allows the reaction to proceed well beyond the limit set by chemical equilibrium

for a conventional reactor system. MTBE leaves the reaction column as the

bottoms product together with a small amount of reaction byproducts. All

byproducts formed in the MTBE reaction are gasoline compatible components.

They include TBA (formed by the reaction of isobutylene and water) and

diisobutylene (formed by the dimerization of isobutylene). Any residual

methanol distills overhead with the unreacted C4s and is routed to the

methanol extraction column. The recovered methanol (methanol column

overhead) is recycled to the MTBE reactor feed. The methanol column bottoms

product (water) is recycled to the methanol extraction column.

Figure 1.5 : CDTECH Process Flow Diagram

1.7.2 UOP Process The process of production of MTBE using UOP Process consists of three

basic steps as listed below [2]:

• Butane isomerization (Butamer Process).

• Isobutane dehydrogenation (Oleflex Process).

- 13 -

• Isobutylene Etherification (Ethermax Process).

In the Ethermax process isobutylene reacts selectively with methanol to

yield MTBE. The reaction proceeds in the liquid phase at mild conditions in the

presence of a sulfonic acid ion exchange resin. A simplified flow diagram of the

process is shown in Figure 1.6. The reaction of isobutylene with methanol is

conducted in the presence of a small excess of methanol relative to that

required for the stoichiometric reaction of the isobutylene contained in the C4

feed.

Figure 1.6 : ETHERMAX MTBE Process Flow Diagram

The reaction is virtually 100 % selective except for minor side reactions

owing to the presence of certain impurities in the feed. Thus, feed water will

give rise to equivalent amounts of tertiary-butyl alcohol (TBA) in the product,

whereas feed isoamylene will yield tertiary-amyl methyl ether (TAME). These

byproducts are usually unimportant, since they can be used equally as gasoline

blending agents and need not be separated from the MTBE product.

The Ethermax process combines the fixed bed reactor technology

(originally developed by Huels) with Koch Engineering’s RWD reaction with

distillation technology to overcome equilibrium limitations found with

conventional fixed bed technologies. The methanol and C4 streams are

combined, heated, and charged to a fixed bed reactor. In the reactor,

- 14 -

condensation of isobutylene and methanol takes place under relatively mild

conditions. Although a tubular reactor can be used, UOP has opted for an

adiabatic packed bed reactor for the large flow rates associated with Oleflex-

based units.

The catalytic section of the RWD column uses a structured packing that

contains catalyst to overcome reaction equilibrium constraints by continuously

fractionating the product from unreacted components. The key to success with

this type of technology is proper distribution of liquid and vapor in the reaction

zone; efficient contact of reactants with catalyst; and instantaneous

fractionation of products. Thus, the KataMax structured packing, developed by

Koch Engineering provides an efficient mass and heat transfer between vapor,

liquid, and solid catalyst, and exhibits a high hydraulic capacity.

MTBE is taken off as RWD bottoms and C4 raffinate is recovered in the

overhead fraction. Residual methanol in the raffinate stream is recovered and

recycled with a water wash system to reduce methanol consumption to about

stoichiometric levels.

1.7.3 Huntsman TBA/PO Process Huntsman has a 181,000 metric ton per year propylene oxide plant in

Port Neches, Texas, which started up in 1994. The plant, originally built by

Texaco was designed with Texaco technology. The byproduct TBA is

converted to 14,000 barrels per day of MTBE (560,000 metric tons). The

Texaco/Huntsman process is believed to be essentially similar to Lyondell's

propylene oxide production process. However, MTBE is produced directly from

TBA with no intermediate production of isobutylene as in the Lyondell process,

according to patent information.

The Texaco/Huntsman process route consists of the well-known steps of

peroxidation of isobutane to TBHP and TBA; epoxidation of propylene with

TBHP to propylene oxide and TBA; and the reaction of TBA with methanol to

form MTBE, although this has not been previously practiced as a one step

reaction.

- 15 -

1.8 THESIS OUTLINE AND WORK METHODOLOGY 1.8-1 Research Objectives

The main aim of this work is to develop a mathematical model for the

production of MTBE in a reactive distillation unit at steady state operation. The

model will incorporate rate data, thermodynamics properties and vapor-liquid

non-idealities. A suitable coding language (FORTRAN) will be used to code the

relevant model equations into used simulation software. The coded equations

will be embedded into PROII software using UAS feature "User Added

Subroutine". The simulation package will be used to assess the effects of some

important design and operating variables on RD column performance.

1.8-2 Research Methodology

The methodology that will be approached to achieve the objectives of this

research topic can be summarized as follows:

Collect and choose the required data for the RD unit (thermodynamics,

kinetics, operating parameters…etc).

Formulate a mathematical model for the RD unit.

Code the relevant model equations using a suitable coding language.

Embed the coded equations using a suitable embedding technique such as

User Added Subroutines (UAS), Excel Visual Basic for Applications (VBA)

or CapeOpen interface into a simulation package whenever it is possible.

Use a suitable simulation software (PROII, Aspen Plus) to solve the model

equations.

Perform sensitivity analysis on the effect of some important design,

operating and kinetics parameters on column performance.

- 16 -

CHAPTER -2-

REACTIVE DISTILLATION

- 17 -

CHAPTER -2-

2.1 INTRODUCTION TO REACTIVE DISTILLATION

Currently there is considerable academic and industrial interest in multi-

functional reactors, involving in-situ separation of products from the reactants

[10]. Reactive distillation is one of the most common means of in-situ product

removal and has been receiving increasing attention in recent years as an

alternative to the conventional reaction-followed-by-distillation processes [11].

Doherty and Buzad in 1992 have placed this subject in historical perspective

and list references to show that the advantages of reactive distillation were

recognized as early as in 1921. [12]

Reactive distillation is potentially attractive whenever a liquid phase

reaction must be carried out with a large excess of one reactant. Under such

circumstances, conventional processes incur large recycle costs for excess

reactant. Reactive distillation, on the other hand can be carried out closer to

stoichiometric feed conditions, thereby eliminating recycle costs. Both

homogeneous and heterogeneous catalysed chemical reactions can be carried

out in a reactive distillation column.

Process 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 equilibrium can give

rise to highly system-dependent features, possibly leading to the presence of

reactive azeotropes and/or the occurrence of steady-state multiplicities

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 [13].

2.2 BASICS AND ADVANTAGES OF REACTIVE DISTILLATION

Reactive distillation is an old concept that combines chemical reaction

and physical separation in the same unit. It is a unit operation that combines a

reactor as an integral part of the distillation column as depicted in Figure 2.1. It

- 18 -

can be utilized for either equilibrium reactions or non-equilibrium (irreversible)

reactions. In the first case, the withdrawal of products as they are formed

results in an increase in the conversion that can be achieved. This increase is

achieved through a shift in the equilibrium, based on Le Chatelier’s principle. In

the second case, it is generally applied to systems where products may react

with reactants, causing a decrease in product yield in conventional reactors.

The term catalytic distillation is also used for such systems where a catalyst (homogeneous or heterogeneous) is used to accelerate the reaction. In this thesis we use the generic name reactive distillation, with the acronym RD, to cover both catalyzed or unanalyzed reactions systems.

Figure 2.1 Simple Reactive Distillation sketch.

2.2-1 Importance of RD

Consider the reversible reaction scheme A+ B ⇔C+ D where the boiling

points of the components follow the sequence A, C, D and B. The traditional

flow-sheet for this process consists of a reactor followed by a sequence of

distillation columns; see Figure 2.2 (a). The mixture of A and B is fed to the

reactor, where the reaction takes place in the presence of a catalyst and

reaches equilibrium. A distillation train is required to produce pure products C

and D. The unreacted components, A and B, are recycled back to the reactor.

- 19 -

In practice the distillation train could be much more complex than the one

portrayed in Figure 2.2 (a) if one or more azeotropes are formed in the

mixture. The alternative RD configuration is shown in Figure 2.2 (b). The RD

column consists of a reactive section in the middle with non-reactive rectifying

and stripping sections at the top and bottom. The task of the rectifying section

is to recover reactant B from the product stream C. In the stripping section, the

reactant A is stripped from the product stream D. In the reactive section the

products are separated in-situ, driving the equilibrium to the right and

preventing any undesired side reactions between the reactants A (or B) with

the product C (or D). For a properly designed RD column, virtually 100%

conversion can be achieved.[14]

Figure 2.2 Processing schemes for a reaction sequence A+ B ⇔ C+ D where C and D are both desired products. (a) Typical configuration of a conventional process consisting of a reactor followed by a distillation train. (b) The reactive distillation configuration. The components A, C, D and B have increasing boiling points. The reactive sections are indicated by grid lines.

The most spectacular example of the benefits of RD is in the production

of methyl acetate. The acid catalyzed reaction MeOH + AcOH ⇔ MeOAc +

H2O was traditionally carried out using the processing scheme shown in Figure

2.3 (a), which consists of one reactor and a train of nine distillation columns. In

the RD implementation (see Figure 2.3 (b)) only one column is required and

nearly 100 % conversion of the reactant is achieved. The capital and operating

costs are significantly reduced.

- 20 -

Figure 2.3 Processing schemes for the esterification reaction MeOH + AcOH ⇔ MeOAc + H2O (a) Conventional processing scheme consisting of one reactor followed by nine distillation columns. (b) The reactive distillation configuration. The reactive sections are indicated by grid lines.

For the acid catalyzed reaction between iso-butene and methanol to

form methyl tert-butyl ether, isobutylene + MeOH ⇔ MTBE , the traditional

reactor-followed-by-distillation concept is particularly complex because the

reaction mixture leaving the reactor forms three minimum boiling azeotropes.

The RD implementation requires only one column to which the butenes feed

(consisting of a mixture of n-butene, which is non-reactive, and iso- butene

which is reactive) and methanol are fed near the bottom of the reactive

section.[13]

The RD concept shown in Figure 2.4 (a) is capable of achieving close to

100% conversion of iso-butene and methanol, along with suppression of the

formation of the unwanted dimethyl ether. Also, some of the azeotropes in the

mixture are “reacted away”.[11]

For the hydration of ethylene oxide to mono-ethylene glycol:

EO + H2O EG, the RD concept, shown in Figure 2.4 (b) is advantageous for

two reasons [15].

Firstly, the side reaction EO+EG DEG is suppressed because the

concentration of EO in the liquid phase is kept low because of its high volatility.

- 21 -

Secondly, the high heat of reaction is utilized to vaporize the liquid

phase mixtures on the trays. To achieve the same selectivity to EG in a

conventional liquid phase plug flow reactor would require the use of 60%

excess water Similar benefits are also realized for the hydration of iso-butene

to tert-butanol [16] and hydration of 2-methyl-2-butene to tert-amyl alcohol

[17].

Figure 2.4 (a) Reactive distillation concept for synthesis of MTBE from the acid catalysed reaction between MeOH and iso-butene. The butene feed is a mixture of reactive iso- butene and non-reactive n-butene. (b) Reactive distillation concept for the hydration of ethylene oxide to ethylene glycol. (c) Reactive distillation concept for reaction between benzene and propene to form cumene. (d) Reactive distillation concept for reaction production of propylene oxide from propylene chlorohydrin and lime. The reactive sections are indicated by grid lines.

Several alkylation reactions, Aromatic + Olefin ⇔ Alkylaromatic, are best carried out using the RD concept not only because of the shift in the reaction equilibrium due to in-situ Separation but also due to the fact that the undesirable side reaction, Alkyl Aromatic+Olefin ⇔ dialkyl aromatic , is suppressed. The reaction of propene with benzene to form cumene, Benzene+Propene ⇔ Cumene ([18]; see Figure 2.4 (c)), is advantageously carried out in a RD column because not only is the formation of the undesirable di-isopropylbenzene suppressed, but also the problems posed by high exothermicity of the reaction for operation in a conventional packed bed reactor are avoided. Hot spots and runaway problems are alleviated in the RD concept where liquid vaporisation acts as a thermal flywheel. The alkylation of isobutane to isooctane, isobutene+ n-butane ⇔ isooctane , is another reaction that benefits from a RD implementation because in-situ separation of the product prevents further alkylation: isooctane + n-butene ⇔ C12H24 [18].

- 22 -

The reaction between propylene chlorohydrin (PCH) and Ca(OH)2 to produce propylene oxide (PO) is best implemented in an RD column, see Figure 2.4 (d). Here the desired product PO is stripped from the liquid phase by use of live steam, suppressing hydrolysis to propylene glycol [19].

Co-current gas-liquid down flow trickle bed reactors are widely applied for hydroprocessing of heavy oils. This co-current mode of operation is disadvantageous in most hydroprocesses [11], and counter-current flow of gas and liquid would be much more desirable as shown in Figure 2.5. The counter-current reactor shown in Figure 2.5 (b) is essentially a RD column wherein the H2S is stripped from the liquid phase at the bottom and carried to the top. The quantitative advantages of the RD implementation for hydroprocessing are brought out in a design study carried out by Van Hasselt in 1999. For a 20,000 bbl/day hydrodesulphurisation unit with a target conversion of 98% conversion of sulphur compounds, the catalyst volume required for a conventional trickle bed reactor is about 600 m3. For counter-current RD implementation the catalyst volume is reduced to about 450 m3.

Figure 2.5 Hydrodesulphurisation of gas oil carried out in (a) co-current trickle bed reactor and (b) counter-current RD unit

From the foregoing examples, the benefits of RD can be summarized as follows:

a) Simplification or elimination of the separation system can lead to significant capital savings.

b) Improved conversion of reactant approaching 100 %. This increase in

conversion gives a benefit in reduced recycle costs.

- 23 -

c) Improved selectivity. Removing one of the products from the reaction mixture or maintaining a low concentration of one of the reagents can lead to reduction of the rates of side reactions 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. RD conditions can allow the azeotropes to be “reacted away” in a single vessel.

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 vaporization and reduce the reboiler duty.

h) Avoidance of hot spots and runaways using liquid vaporization as

thermal fly wheel.

2.2-2 Constraints and Difficulties in RD implementation

Against the above mentioned advantages of RD, there are several

constraints and foreseen difficulties [20].

1. 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.

2. 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.

3. 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.

4. Process conditions mismatch. In some processes the optimum

conditions of temperature and pressure for distillation may be far from

optimal for reaction and vice versa.

- 24 -

2.2-3 The complexity of RD

The design and operation issues for RD systems are considerably more

complex than those involved for either conventional reactors or conventional

distillation columns. The introduction of an in-situ separation function within the

reaction zone leads to complex interactions between vapour-liquid equilibrium,

vapour-liquid mass transfer, intra-catalyst diffusion (for heterogeneously

catalysed processes) and chemical kinetics. Figure 2.6 shows the various

transfer processes in homogeneous and heterogeneous RD.

In heterogeneous RD the problem is exacerbated by the fact that these

transfer processes occur at length scales varying from 1 nm (pore diameter in

gels, say) to say a few meters (column dimensions); see Figure 2.7. The time

scales vary from 1 ms (diffusion within gels) to say a few hours (column

dynamics). The phenomena at different scales interact with each other. Such

interactions, along with the strong non-linearities introduced by the coupling

between diffusion and chemical kinetics in counter-current contacting, have

been shown to lead to the phenomenon of multiple steady states and complex

dynamics, which have been verified in experimental laboratory and pilot plant

units [36]. Successful commercialization of RD technology requires careful

attention to the modelling aspects, including column dynamics, even at the

conceptual design stage [12].

In some cases the reactor and distillation paradigms do not translate

easily to RD. The potential advantages of RD could be nullified by improper

choice of feed stage, reflux, amount of catalyst, boilup rate, etc. Thus, it is

possible to decrease conversion by increasing the amount of catalyst under

certain circumstances. Increased separation capability could decrease process

performance [21].

- 25 -

Figure 2.6 Transport processes in RD. (a) homogeneous liquid phase reaction,

and (b) heterogeneous catalyzed reactions.

Figure 2.7 Length and time scales in RD.

2.2-4 Models available for design of RD columns

A variety of models exist in the literature for design of RD columns. They

can be classified in the following manner.

1. Steady-state equilibrium (EQ) stage model

2. Dynamic EQ stage model

3. Steady-state EQ stage model with stage efficiencies

- 26 -

4. Dynamic EQ stage model with stage efficiencies

5. Steady-state nonequilibrium (NEQ) stage model, where the interphase

mass transfer is described by rigorous Maxwell-Stefan diffusion

equations

6. Dynamic NEQ stage model

7. Steady-state NEQ cell model, developed by Higler in 1999, in order to

account for staging of the vapour and liquid phases during cross-current

contacting on a distillation tray [22].

The equilibrium stage model assumes that the vapor and liquid streams

leaving a given stage are in thermodynamic equilibrium with one another.

These models can be coupled with the assumption of chemical equilibrium at

each stage or the kinetics can be described using an nth order kinetic reaction

model.

The column is described by a group of equations that model the

equilibrium stages in a column configuration. These are known as the MESH equations. MESH stands for:

• M: Material balance equations for each component and total mass.

• E: Equilibrium equations.

• S: Summation equations or composition constraints.

• H: Heat or energy balance equations.

- 27 -

CHAPTER -3-

LITERATURE SURVEY

- 28 -

CHAPTER -3- 3.1 OVERALL REVIEW

The combination of chemical reaction with distillation in only one unit is

called reactive distillation. The performance of reaction with separation in one

piece of equipment offers distinct advantages over the conventional, sequential

approach [7]. In reactive distillation (RD) chemical reactions occur within the

distillation column to achieve specific goals, such as to obtain high conversions

and high purity products as well to minimize side reactions. Currently RD has

various applications, such as the production of methyl tertiary -butyl ether

(MTBE) and ethylene oxide (EO) and others.

Reactive distillation is being used in industrial applications with more

frequency because of increasing research and development of this technology,

a result of commercial and academic experience and success. Examples of

commercialized technologies are the Ethermax process from UOP-Huls and

Koch Engineering that uses the KataMax structured packing from Koch and

Catacol, a low cost reactive distillation technology for etherification from IFP’s

Industrial Division [22].

So far, there is no generally accepted method for the design of

distillation with reaction. Most of the systematic methods available possess

limitations because of their simplified assumptions. Moreover, these methods

have rarely been proven with a variety of reactive distillation processes and

they do not consider the design in detail [12].

In spite of the advances in separation with reaction processes, reactive

distillation still relies on intuition and expertise. A reactive distillation problem

can be studied using different approaches including: feasibility, simulation,

modeling, design and experimental studies in the laboratory and the pilot plant.

A combination of all of these methods gives rise to the most accurate

solution to the problem. One very important aspect of predicting the behavior in

these systems is the model used to design and simulate the reactive distillation

process. In the literature, the most common models that have been developed

and proven are the equilibrium stage model and the non-equilibrium stage

- 29 -

model. The equilibrium stage model is based on the conventional equilibrium-

stage model of a distillation column with the addition of the reaction terms in

the mass and energy balances. The non-equilibrium stage model for reactive

distillation, also known as the rate-based model, is an extension of the

conventional rate based model for distillation. In this chapter, a discussion of

the important aspects of modeling, simulation, design and analysis of reactive

distillation is provided.

Matthias et al, [24] presented a conceptual design methodology for the

reactive distillation columns, the method assesses feasibility of a proposed

reactive distillation, designs the column and allows evaluation of the design for

both fully reactive and "hybrid" column configurations. Stage composition lines

are used to represent all possible liquid compositions in a column section for

specified product compositions and for all reflux or reboil ratios. Reaction

equilibrium is assumed on each reactive stage, and vapor-liquid equilibrium is

assumed on all stages. The methodology is illustrated by application to an ideal

reactive system and for MTBE production. They developed a new graphical

design methodology to assess feasibility and design columns for proposed

reactive distillation processes. Also, they concluded that the methodology is

restricted to systems with a single-feed two-product columns. It is assumed that

equilibrium reaction take place in the liquid phase only.

A simulation study of the impact of various process configuration

catalyst types and catalyst loading on the production rate and separation of the

MTBE product was carried by Shah et al, [25]. They mentioned in their paper

the advantages and disadvantages of each configurations. For reactive

distillation, the catalyst distribution on a range of trays and the feed location is

considered for the optimal production/separation scenario. They also studied

an approach, based purely on steady state analyses, for synthesizing effective

control structures for reactive distillation (RD) columns [26]. The main idea was

to analyze the steady-state relationships between the manipulated (input)

variables and potential controlled (output) variables to identify input-output (IO)

pairings that are sensitive and avoid steady state multiplicities providing a large

range of nearly linear operating region around the base case design. The

MTBE case study shows that input and output multiplicity in the IO relation

- 30 -

occurs in the steady-state relationships. It also shows that the occurrence of

multiplicity depends on the control strategy implemented.

Rivera and Johan [27] considered the residue curve mapping technique

(RCM) a powerful tool for the flow-sheet development and preliminary design of

conventional multi-component separation processes. An RCM-based feasibility

analysis has been applied to the homogeneous RD synthesis of MTBE at 11

atm from methanol and isobutylene and in the presence of n-butane. The

reaction space, defined in terms of transformed composition variables, has

been divided into sub-regions characterized by separation boundaries. A

feasibility analysis of the RD process has been performed based upon the

location of the reacting mixture, defined initial separation sequences have been

generated according to the feed transformed-composition. In the generated

sequence, high purity MTBE has been obtained as a product, due to the

appearance of a pseudo reactive azeotrope, which imposes limitation to the

separation task.

The progress in chemical engineering unit operations which laid out the

catalytic distillation column is a significant advancement in chemical

engineering technology. Some of the equilibrium-limited reactions show an

improving yield due to the continuous removal of the products from the reaction

zone. A substantial research work, mainly experimental has been done for this

combined unit operation, most of which have appeared in patents [28 & 29].

However, essential research aspects of the reactive distillation process

mathematical modeling, design, optimization and control have been covered

yet. Also, little attention has been paid to plate efficiency when reaction occurs.

Research and development work on reactive distillation has focused on:

• Computational methods to solve the simulations chemical reaction and

vapor-liquid equilibrium equations.

• Chemical equilibrium aspects in reactive systems (e.g. thermodynamic

models).

3.2 COMPUTATIONAL METHODS The Computational methods used for reactive distillation are extensions

to the algorithms developed for the solution of the equations for conventional

- 31 -

distillation. The first attempts to model reactive distillation were by using the

simplified plate-to-plate calculations. Rigorous mathematical models for

computer simulation were not developed until the 1970's. Since that time,

various techniques have been developed that allow the rigorous solution of the

equations. These techniques include equation partitioning methods, and

Newton-Raphson based methods.

Nelson in 1971 modified the Tierny-Bruno algorithm based on tray-by-

tray calculation, by taking into consideration the non-ideal vapor liquid

equilibrium. Which represented in material balance equations no longer linear

in composition [30].

Suzuki et al. in 1971 applied the successive iteration method to reactive

distillation problems, and concluded that it converges rapidly and it is stable.

However, difficulties with convergence arise for systems with non-ideal

solutions, because of non linearity of the equations. In addition, successive

iteration methods have the disadvantage that as the solution is approached,

the progress of iteration calculations decelerates [31].

Murthy in 1984 considered an extension of the Newton-Raphson

algorithm to columns in which chemical reactions occur, while Venkataraman et

al. in 1990 were developed an extension of the " inside-out" approach

proposed by Boston that combines the advantages of Newton's method with

those of the inside-out strategy very effectively [32].

Jelinek and Hlavacek in 1976 applied the relaxation method to solve

steady state countercurrent equilibrium stage separation with chemical reaction

problems. They confirmed the suitability of this method where an azeotrpoe

exists [33].

Teirney and Riquelme in 1982 proposed a correction algorithm which

gives quadratic convergence near the solution. Its use was demonstrated on a

sample problem of the separation of meta-and para-xylene using experimental

results from an earlier study done by Satio et al. in 1971. They found that their

solution agree well with the numerical solution obtained by Satio et al, but not

with Saito's experimental data. Teirney and Riquelme in 1982 claimed the

equilibrium in the reaction is not satisfied in each stage. This contradicts the

experimental results of Satio et al. in 1971, showing that the equilibrium

- 32 -

constant is almost constant in the range of operating temperature between (50-

100 ˚C) [34].

The reactive distillation column has no reboiler, although it has to be

used in conjunction with the gas/liquid polymer reactor. The model was based

on the material balance equations describing the system, and the simulation

was carried out by using a fourth-order Runge-Kutta method. The usefulness of

the model was limited because of the number of assumptions and

simplifications that were made in development of the model, such as:

neglecting enthalpy balances, neglecting tray hydraulics [34].

Reactive Distillation (RD) is state-of-the-art multifunctional reactor

concept that integrates reaction and distillation in a single process unit. Some

recent reactions that have been proposed to utilize RD technology involve non-

condensable species like hydrogen which proposed by Kamath et al in 2005.

Reactions involving liquid phase splitting have also been examined in RD .

Very little research has been done for such complex cases and the potential of

RD for such reactive systems has not been thoroughly investigated [36].

This work focuses on the presence of both non-condensables and liquid

phase splitting in RD using the one-step synthesis of methyl isobutyl ketone

(MIBK) from acetone as a case study. Traditionally, MIBK is produced in three

steps by the condensation of acetone to diacetone alcohol followed by its

dehydration to mesityl oxide (MO) and then its hydrogenation to MIBK. The first

two reactions are reversible and limited by thermodynamic equilibrium. But the

third reaction (MO to MIBK) is fast and irreversible and hence the overall

reaction (acetone to MIBK) is expected to be an irreversible reaction. However,

the presence of water as a product inhibits the reaction rate considerably

leading to pseudo-equilibrium [36].

The work employs the RD technology to minimize the amount of water in

the reactive zone by simultaneous distillation and thereby increase the overall

conversion of acetone. The hydrogenation reaction results in the appearance of

a non-condensable species in the RD column while the ternary acetone-water-

MIBK system results in liquid-liquid phase splitting (Figure 3.1).

- 33 -

Figure 3.1 Ternary LLE diagram for Acetone-water-MIBK system at 5atm.

A potential RD configuration for the synthesis of MIBK was developed as

shown in Figure 3.2. The column configuration is capable of achieving a very

high conversion of around 95% by employing a very high reflux ratio (defined

as ratio of the liquid refluxed to the liquid distillate). The liquid phase

composition profile in the column is shown in Figure 3.3. The simultaneous

distillation in the reactive zone keeps the water concentration below 35% which

would not have been possible in a conventional reactor. Because of the

presence of a binary acetone-water azeotrope at higher pressures, a large

number of stages may be required to reduce the water concentration in the top.

As expected, phase-splitting is observed in the stripping section. However, the

presence of sufficiently large amount of acetone in the reactive zone ensures a

homogenous liquid phase [35].

- 34 -

Figure 3.2 RD Column Configuration.

Figure 3.3 Liquid Composition profile in RD Column.

It was concluded from the above work that there is a further scope for

improvement of the proposed RD configuration. The column performance is

very sensitive to the feed locations and their flow rates. The presence of non-

condensable hydrogen poses additional complexities in the VLLE

computations. A successful RD process is not only capable of achieving a

much higher conversion as compared to the conventional route but also leads

to process intensification and compactness.

- 35 -

An approach, based purely on steady-state analyses, for synthesizing

effective control structures for reactive distillation (RD) columns . The main idea

is to analyze the steady-state relationships between the manipulated (input)

variables and the potential controlled (output) variables to identify input–output

(IO) pairings that are sensitive and avoid steady-state multiplicities providing a

large range of nearly linear operating region around the base case design.

Traditional SISO control loops are then implemented using these IO pairings to

obtain control structures that maintain the column near the design product

purity and conversion for the anticipated primary disturbances. The Niederlinski

Index is used to eliminate dynamically unstable pairings in control structures

with multiple loops. The approach is demonstrated on an example MTBE RD

column. The impact of steady-state multiplicities on control structure design is

highlighted. [25]

This work illustrates the systematic analysis of the steady-state IO

relationships for synthesizing effective control structures for RD columns. Such

control structures are necessary to provide robust, stable, safe and economical

column operation to tide over disturbances entering the column. Effective

control structures can be obtained by identifying IO pairings that avoid

multiplicities and allow a sizeable nearly linear operating region around the

base case. The MTBE case study shows that input and output multiplicity in the

IO relations occurs in the steady-state relationships. It also shows that the

occurrence of multiplicity depends on the control strategy implemented.

For the MTBE example, a control structure that uses the reboiler duty to

control Tray 11 temperature in the stripping section and the butene feed to

control Tray 10 isobutylene composition in the reactive section is found

suitable. Proper choice of the manipulated variables and the tray locations for

the measurements in the two loops is especially critical for effective column

regulation. At a more general level, the results show that maintaining the

stoichiometric balance of the fresh feeds is needed for the effective control of

double feed RD columns.

M. H. M. Reis et al in 2006, presented a novel approach for establishing

the route for process intensification through the application of two developed

- 36 -

softwares to characterize reactive mixtures is presented. A robust algorithm

was developed to build up reactive phase diagrams and to predict the

existence and the location of reactive azeotropes. The proposed algorithm

does not depend on initial estimates and is able to compute all reactive

azeotropes present in the mixture. It also allows verifying if there are no

azeotropes, which are the major troubles in this kind of programming. An

additional software was developed in order to calculate reactive residue curve

maps. Results obtained with the developed program were compared with the

published in the literature for several mixtures, showing the efficiency and

robustness of the developed software's. [40]

Figure 3.4 shows the nonreactive residue curve map for the system

under consideration. The binary azeotrope between methanol and tert-butyl

methyl ether (MTBE) is a saddle point and between isobutylene and methanol

an unstable node appears, leading to the occurrence of a distillation boundary.

This distillation boundary splits the diagram in two different regions of

distillation, making the separation of the ternary mixture impossible. [40]

Figure 3.4 Nonreactive residue curve map for the system IBTE+ MeOH ↔ MTBE at 101.32 kPa.

Figure 3.5 shows the reactive residue curve map for this system, where

it can be seen that, under atmospheric pressure, this reactive mixture does not

present any azeotrope, showing how the chemical equilibrium influences the

phase equilibrium. This conclusion is in agreement with the results of Barbosa

- 37 -

and Doherty (1988b). Figure 3.6 shows the reactive phase diagram for

isobutylene/methanol/MTBE at 101.3 kPa, confirming that this reactive system

is zeotropic.

Figure 3.5 Reactive residue curve map for the system IBTE+ MeOH ↔ MTBE at 101.32 kPa.

Figure 3.6 Reactive phase diagram for the system IBTE/ MeOH /MTBE at 101.32 kPa.

Maier et al. (2000) studied this same system at 810.56 kPa (8 atm) and

with different constant values for the equilibrium constant (Kr). Similar results

were obtained in this work. When Kr is equal to 49.0, the mixture forms two

azeotropes at different compositions. Figure 3.7 shows the reactive phase

equilibrium diagram for this case. [40]

- 38 -

Figure 3.7 Reactive phase diagram for the system IBTE/ MeOH /MTBE at 810.56 kPa.

In actual processes to produce MTBE through reactive distillation, the

source of iC4 consists of C4 cuts available from steam or catalytic crackers. In

order to analyze this multicomponent system, it is considered that C4 stream

consists, basically, of isobutylene and n-butenes, being n-butenes inert

components in the reaction.

For this quaternary system (ethanol/isobutylene/MTBE/n-butene) the

search of azeotropes were carried out at pressures of 1013.2, 2026.4 and

4052.8 kPa (10, 20, and 40 atm). Although in this high pressure it is not correct

to consider the vapor as an ideal mixture, this assumption will be still

considered, in order to compare the obtained results with those published by

Maier et al. (2000). [40]

- 39 -

CHAPTER -4-

MODELING AND SIMULATION OF MTBE PROCESS

- 40 -

CHAPTER -4- 4.1 MTBE PRODUCTION BY PROII

4.1-1 Introduction

This section will demonstrate the use of PRO/II in the simulation of the

synthesis of methyl tert-butyl ether (MTBE), also the simulated process for

MTBE production will be depicted and each unit in the process will be

elaborated. A PRO/II simulation model of a typical MTBE plant is presented

here. The process plant includes a reactor along with an azeotropic distillation

column for separation of the MTBE product.

A reactive distillation section is added to the MTBE azeotropic column in

order to increase the overall conversion to MTBE. This is followed by the

methanol recovery section which includes a liquid-liquid extractor. The

thermodynamics successfully predicts the azeotropic removal of methanol from

the MTBE product stream.

4.1-2 Process Description

There are several variations on MTBE plant designs. In general, an

MTBE plant is comprised of the three sections, a reactor section, a MTBE

recovery section, and a methanol recovery section. For the Ethermax process,

the MTBE recovery section includes a second reaction zone in the distillation

column. The complete process flow diagram for the MTBE reactive distillation

plant model used in this simulation is given in Figure 4.1.

- 41 -

Figure 4.1 MTBE Plant Flowsheet.

MTBE is manufactured by catalytically reacting isobutylene and

methanol in a fixed-bed reactor at a moderate temperature and pressure. The

reaction is exothermic and reversible, and is carried out in the liquid phase over

a fixed bed of ion-exchange resin-type catalyst. It is highly selective since

methanol reacts preferentially with the tertiary olefin.

In the Ethermax MTBE process, modeled here, an isobutylene-rich

mixed C4 stream is mixed with fresh methanol along with a small amount of

recycle methanol and fed to the reactor section. The reactors are cooled to

under 200 °F to prolong catalyst life and to minimize the undesirable side

reactions such as dimerization of isobutylene.

The methanol-to-isobutylene ratio in the reactor feed is kept low to

minimize the costs of recovering unreacted methanol, and to facilitate the

operation of the MTBE column which will be discussed later. Generally, this

ratio is maintained close to 1:1. Table 4.1 contains the reactor feed composition

used in this model.

- 42 -

Table 4.1 Reactor Feed data Stream Name Stream Number

C4 Feed 2

Methanol Feed 1

Methanol Recycle 20

Flowrate (kg-mol/hr) 850 277.5 4.3 Temperature (°C ) 16 16 44 Pressure (kPa) 1620 1620 1724

Component Mole % N-butane 9.0 0.0 0.0 Isobutane 41.0 0.0 0.0 1-butene 7.0 0.0 0.0

Cis 2-butene 4.0 0.0 0.0 Trans 2-butene 6.0 0.0 0.0

Isobutylene (IBTE) 33.0 0.0 0.0 MTBE 0.0 0.0 0.0

Tert-butanol (TBA) 0.0 0.0 0.0 Water 0.19 0.0 6.98

Di-isobutylene (DIB) 0.0 0.0 0.0 Methanol (MEOH) 0.0 100 93.02

An isobutylene conversion to MTBE of 90 to 93% is easily achieved in

the reactor. Overall isobutylene conversions higher than those obtained in the

standard process can be achieved by either recycling a portion of the MTBE

column overhead product, or by providing a second reactor unit and MTBE

column downstream of the first MTBE column. The cost-effectiveness of these

options vary from plant to plant, but both require greater capital expenditure. In

the reactive distillation process, no major increase in capital expenditure is

required and overall isobutylene conversions of over 99% are easily obtained.

Any water in the reactor feed (from recycle methanol) is instantly

converted to t-butanol. Another impurity, di-isobutylene, is formed by the

dimerization of isobutylene. While the formation of di-isobutylene and t-butanol

should be minimized, their presence in small concentrations in the MTBE

product is acceptable since these byproducts also have very high octane

numbers. Table 4.2 shows the three main reactions used in the stoichiometric

- 43 -

reactor model. The base component and the fraction converted are also

shown.

Table 4.2 Reaction Stoichiometry Reaction Base Component Conversion %

2 (IBTE) = DIB IBTE 0.25 H

2O + IBTE = TBA H

20 100.00

IBTE + MEOH = MTBE MEOH 93.0

Catalyst Used A common catalyst for the MTBE synthesis process is the Amberlyst 15

polymeric catalyst developed by Rohm and Haas. Approximate properties of

the commercial form of this catalyst, along with suggested operating conditions

are provided below in Table 4.3.

Table 4.3 Approximate Catalyst Properties of Amberlyst 15

Physical form Spherical beads Ionic form Hydrogen Acid site concentration 1.8 meg/ml (4.9 meg/g) Moisture content 53% Apparent density 770 g/l Particle size 0.35-1.2 mm Shrinkage: Wet to methanol Wet to MTBE

4% 12%

Porosity 0.30 cc/g Average pore diameter 250 A Surface area 45 m

2/g

Bulk density 48 lb/ft3

Maximum operating temperature 120 °C Minimum bed depth 0.61 m Flowrate, LHSV 1-5 hr

-1

MTBE Recovery Section

In the Hüls process, the reactor products are processed in the MTBE column

where MTBE, along with t-butanol, dimerized butylene, and a trace amount of

methanol, are removed as the bottoms product. In the Ethermax process,

further reaction of the isobutylene to MTBE takes place in a section of the

distillation column containing the catalyst resin in tower packing. The MTBE is

- 44 -

removed as the bottoms product in a manner similar to the standard process.

The MTBE product is greater than 99.5% pure and requires no further

purification.

The key to operating the MTBE column is to have a sufficient amount of

C4s in the column feed to form azeotropes with the methanol in the feed.

Conversely, if a proportionately large amount of methanol is present in the

column feed, it may result in breakthrough of methanol with the MTBE bottoms

product. Therefore, suitable azeotrope formation is possible only when a limited

excess of methanol is used in the reactor feed. In this manner, unreacted

methanol, which has a higher boiling point than MTBE, is fractionated away

from the MTBE bottoms. The overhead product containing non-reactive linear

butenes, iso and normal butanes, and unreacted methanol and isobutylene, is

sent to the methanol recovery section.

Methanol Recovery Section

In the methanol recovery section, the MTBE column overhead product is

water washed to extract methanol. This unit is simulated as a liquid-liquid

extraction column. The raffinate, which contains less than 10 ppm methanol, is

suitable for recovering high purity C4 isomers, or as a feed to an alkylation unit.

The extract phase which contains water, methanol and small amounts of

dissolved hydrocarbons is warmed and flashed to remove the hydrocarbons.

The resultant methanol-water mixture is fractionated to recover methanol as

the overhead product. The methanol (with a trace of water) is recycled to the

MTBE reactor. The wash water stream from the bottoms, along with a small

amount of makeup water, is returned to the water wash column.

- 45 -

4.2 MODELING AND SIMULATION OF MTBE PROCESS USING PRO/II

PRO/II software has the capabilities of solving Reactive Distillation Processes

utilizing Chemdist Algorithm provided by the software, but it requires knowledge of

Material and Energy balances equations as well as Thermodynamics and kinetics of

reaction.

The Chemdist algorithm in PRO/II is a Newton based method which is suited to

solving non-ideal distillation problems involving a smaller number (10 vs. 100) of

chemical species. These conditions are generally encountered in chemical distillations

as opposed to crude fractionation where the I/O algorithm would be a better choice.

Chemdist is designed to handle both vapor-liquid and vapor-liquid-liquid equilibrium

problems as well as chemical reactions.

All derivatives for the Jacobian matrix are calculated analytically. User-added

thermodynamic options that are used by Chemdist must provide partial derivatives

with respect to component mole fractions and temperature. Chemdist uses the chain

rule to convert these to the needed form.

PRO/II automatically generates numerical estimates for the reaction rate

derivatives. In many cases, this is sufficient. However, certain reactions require the

use of more accurate analytical derivatives that provide better solutions. Reactions

requiring these more accurate analytical derivatives include: reversible reactions,

exothermic reactions, and/or reactions where the equilibrium is sensitive to

temperature. There are some assumption will be considered to simplify the simulation

work which can be listed as :

1. The reactor is operated under steady state conditions.

2. The reactor is isothermal operating condition (55 C), it is maintained by

circulating a coolant.

3. The pressure drop across reactor is specified with 69 kpa.

4. One dimensional homogeneous model is considered.

5. Dispersion of mass and energy terms are negligible and a plug flow

reactor model is assumed.

6. Complete mixing of gases entering the collection chamber before the

catalyst beds. No reaction takes place inside the chamber, the reaction

- 46 -

is only in the liquid phase.

7. Physical properties are considered as a function of temperature and

pressure, and SRKM thermodynamic method is used.

4.3 THERMODYNAMIC DATA

The VLE fractionators are simulated well with PRO/II's modified Soave-

Redlich-Kwong (SRKM) equation of state method. For this method, PRO/II

contains extensive, built-in databanks that encompass binary interaction

parameter data for the majority of component pairs present in this simulation.

However, binary interaction data (Kijs) are directly supplied for 8 component

pairs to improve the accuracy of the separations in the columns as

demonstrated in Figure 4.2. The binary interaction parameters are listed below

in Table 4.4. The thermodynamic set used for VLE thermodynamics is referred

to as SRKM_VLE. The used components and their abbreviation are as follows:

Table 4.4 Binary Interaction Parameters for SRKM_VLE Comp I NC4

n-butane

1BUT 1-butene

BTC2Cis2-butene

BTT2 Trans2-butene

IBTEIsobutylene

MEOH Methanol

TBA Tert-Butanol

MTBE

Comp j TBA MEOH MEOH MEOH MEOH TBA H20 DIB

kija 0.0469 0.136 0.136 0.136 0.13553 -

0.07397

-0.145 0.05785

kjia 0.1260 -0.0323 -0.0323 -0.0323 -0.0322 -

0.05522

-0.253 -0.0093

kijb 0.0 0.0 0.0 0.0 0.0 0.0 0.0 -10.144

kjib 0.0 0.0 0.0 0.0 0.0 0.0 0.0 6.17

UOM K K K K K K K K

- 47 -

Figure 4.2 Thermodynamic Data of PRO/II software.

Transport properties are needed in order to use the rigorous heat

exchanger model in the MEOH recovery section. Transport property

calculations are set to pure-component averages by default and can be

modified by modifying the thermodynamic method of choice.

The liquid extraction unit is simulated using the SRKM method for VLLE

thermodynamics with binary interaction data again supplied as part of the input.

The thermodynamic set used for VLLE thermodynamics is referred to as

SRKM_VLLE. The L1KEY component (i.e., the predominant component in the

L1 liquid phase) is specified as n-butane. The L2KEY component is specified to

be water. Explicitly specifying the key components eliminates the need for

PRO/II to find an appropriate immiscible pair, reducing the computation time.

All the azeotropes are properly predicted. The form of SRKM is presented as:

………..………(4.1)

Where :

Xi= Mole fraction of component i

Xj= Mole fraction of component j

Kij= Kji = Binary interaction parameter

aij= The mixing rule, Cij= constant

]))/()(()1[()( 21

ijcjiijiijijjiij XXXkkkaaa +−+−=

- 48 -

The accuracy of correlating vapor-liquid equilibrium data using a cubic equation

of state can be improved further by choosing an appropriate mixing rule. The

original mixing rule was derived from the van der Waals one-fluid

approximation:

)/()()/( 22 wbubvvTabvRTP ++−−= …………………………………....……..(4.2)

where: P = the pressure T = the absolute temperature v = the molar volume u,w = constants, typically integers

∑∑=i j

ijji axxa …………………………………………………………....……..(4.3)

∑=

iiibxb …………………………………...............................................…….(4.4)

where:

Xi = mole fraction of component i.

The binary interaction parameter, kij, is introduced into the mixing rule to

correct the geometric mean rule of parameter a in the general cubic equation of

state (4.2):

The original mixing rule is capable of representing vapor-liquid equilibria

for nonpolar and/or slightly polar systems using only one (possibly

temperature-dependent) binary interaction parameter.

4.4 MTBE REACTION KINETICS

The algorithm used for the reactive distillation column model is a

Newtonian-based algorithm. Therefore, in order to accurately model the MTBE

reaction in the distillation column, we need to determine not only the reaction

rate of the reaction, but also the temperature and composition derivatives of the

rate. These derivatives may be generated numerically by an estimation

method, or analytically by an expression based on the reaction rate equation.

The MTBE reactions fit all three of the reaction types, requiring the use

of analytical derivatives. The reaction rate expression and its analytical

derivatives can be easily and readily entered by the user in the Procedure Data

- 49 -

category of input. The MTBE reaction rate expression used in this simulation

model is based on the rate expression described in a paper by Al-Jarallah [38].

First, in the Reaction Data category of input, the stoichiometry of the

forward reaction is given (IBTE + MEOH = MTBE). The kinetic data will be

provided later on in the Procedure Data category of input using FORTRAN-like

language as the procedure named ALJD. The procedure data used in the

reactive distillation column model is entered in the Kinetic Procedure data entry

window shown in App (A).

The MTBE synthesis reaction can be represented by:

A + B ↔ C ………………………………………………..…..………………….(4.5)

Where, A, B and C denote methanol, isobutylene and MTBE

respectively, In general, the forward reaction is order a in A and order b in B,

and the reverse reaction is order c in C.

The rate of surface reaction, rs, is assumed to be the rate controlling

step, as there were no mass transfer limitations. There are two possible

mechanisms by which this surface reaction takes place:

1. Reaction between adsorbed molecules of both A and B on adjacent

active centers, and

2. Reaction between one adsorbed reactant and the other reactant in

solution [37].

The first mechanism is the Langmuir-Hinshelwood mechanism and the

second one is the Rideal-Eley mechanism as discussed by Smith in 1981 and

Satterfield in 1980. In these references the reaction is assumed to be a simple

reaction, that is, the reaction is first order in all species. The following rate

equations were derived for general orders of reaction a, b and c. For a

Langmuir-Hinshelwood model, the rate of reaction can be represented by the

following equation [37]:

( ) ⎥⎥⎦

⎤

⎢⎢⎣

⎡

+++

−×=

+baCCBBAA

cC

bB

aAb

BaAss CKCKCK

kCCCkkkr

1/

…………..……….…………….(4.6)

For the case of the Rideal-Eley mechanism, there are two possibilities in which

either one of the two reactants is adsorbed on the catalyst and then reacts with

- 50 -

the other reactant in solution. For the case when the methanol (A) is adsorbed

and reacted with the isobutylene (B) in solution, the final rate equation is:

( ) ⎥⎥⎦

⎤

⎢⎢⎣

⎡

++

−×= a

CCAA

cC

bB

aAa

Ass CKCKkCCC

kkr1

/…………………………..………………..….(4.7)

For the case when isobutylene is absorbed and reacted with methanol in

solution, the final rate equation is:

( ) ⎥⎥⎦

⎤

⎢⎢⎣

⎡

++

−×= b

CCBB

cC

bB

aAb

Bss CKCKkCCC

kkr1

/…………………………..…………………...(4.8)

For a given set of a, b and c the unknown parameters in Equation (4.6), (4.7)

and (4.8) are the surface reaction rate constant, Ks the equilibrium adsorption

constants KA, KB and KC and the thermodynamic equilibrium constant, K. This

equilibrium constant can be calculated from experimental concentration data in

which concentration equilibrium has been reached [37].

Since K = Kx Kγ …………………………………………………………………………..(4.9)

The dependence of the rate constant, Ks, on temperature was determined from

the Arrhenius equation,

Ks = Kso exp (-E1 / RT)………………………………………………..………....(4.10)

The values of Kso and E1 were found from the least squares fit of Equation

(4.10) Thus:

Ks = 1.2 x 1013 exp (-87,900/RT)……………………………………………..(4.11)

The dependence of the adsorption constants, KA and KC, was

determined from the van’t Hoff equation,

KA = KAo exp (∆HA / RT)………………………………………………..…........(4.12)

- 51 -

KC = KCo exp (∆HC / RT)………………………………………………..…...…..(4.13)

The values of KAO, KCO, ∆HA and ∆HC where obtained from the least

squares fit of the above two equations. Thus:

KA = 5.1 x 10-13 exp (97,500/RT)……………………………………………….(4.14)

and

KC = 1.6 x10-16 exp (119,000/RT)…………………………………….………..(4.15)

The reaction rate equation described by Al-Jarallah takes into account

the forward and the reverse reaction. We have modified Al-Jarallah's rate

equation for MTBE process to simulate the effect of catalyst loading on the

reaction rate. This was achieved by removing the catalyst terms from the

concentration terms. The modified reaction rate is given by:

( ) ⎥⎥⎦

⎤

⎢⎢⎣

⎡

++

−×=

+baBBAA

cC

bB

aA

Ass CKCKkCCC

kkr1

/…………………………..………………(4.16)

where: • Ks: surface reaction rate constant= 1.2x1013 exp(-87900/RT) in (gmole/g catalyst)

• KA: equilibrium adsorption constant = 5.1x10-13 exp(97500/RT) in g catalyst/gmole

• KC : equilibrium adsorption constant = 1.6x10-16 exp(119000/RT) in g catalyst/gmole

• K : equilibrium constant

• CA : IBTE concentration in mole/l

• CB : MEOH concentration in mole/l

• CC : MTBE concentration in mole/l

4.5 MODELING PROCEDURE In modeling this process, the Steady-state equilibrium stage model, will all its

underlined assumptions has been used. The main feature of the model is the

assumption that the vapor and liquid streams leaving a given stage are in a

thermodynamic equilibrium.

The reaction kinetics in the reaction zone was assumed to follow Al-Jarallah

kinetics [37]. A typical schematic diagram of a catalytic distillation column is shown in

Figure 4.3.

- 52 -

Figure 4.3 A schematic representation of a catalytic distillation column.

A schematic diagram of an equilibrium tray is shown in Figure 4.4. It is

assumed that the tray has one feed stream Fk one vapor side streams SV one liquid

side stream SL and coolings Q occurs.

Lk+1 V

k Feed SVk

Fk Qk

HEAT DUTY

SLk

Lk

Vk-1

Figure 4.4 Representation of a generic equilibrium stage.

Stage K

Tk, Pk ∆Rk, Mk

- 53 -

4.5-1 Model equations for a single stage

The model equations for a generic stage k and component i are presented

based on the commonly used distillation column, with incorporation of the

reaction terms, they may be expressed as follows:

1. Component material balance on stage k:

0)()(1

,,,,,11,11 =±++−+−+ ∑=

−−++

r

jijkikkikkkikkkikkikk RZFYSVVXSLLYVXL …………..…(4.17)

Where; Xk,j is the mole fraction of liquid phase for species i in stage k. Yk,i is the mole fraction of vapor phase for species i leaving stage k (Lk+SLk) is the molar liquid flow rate leaving stage k. (Vk+SVk) is the molar vapor flow rate leaving stage k. Fk is the feed flow rate to stage k. ∑R jk,i is the rate of disappearance of component i due to reaction j on stage k

(i.e. when the sign is negative).

2. Total material balance on stage k:

0)()(11 =+++−+−+ −+ kkkkkkkk RFSVVSLLVL …………….…………….(4.18)

3. Energy balance on stage k:

0)()(1

1111 =+−++−+−+ ∑=

−−++ jk

r

jjkkfkkkkkkkkkkkk HRQhFHSVVhSLLHVhL ………..(4.19)

Where;

hk denotes the molar enthalpy of liquid phase in stage k. HK denoted the molar enthalpy of vapor phase in stage k. Hj,k is the heat of reaction in stage k. Qk is the heat duty in stage k.

For stages without feed or side streams or no reaction, these terms are assigned a value of zero in Equations (4.17-4.19).

4. Component equilibrium relationship on stage k:

Yk,i = γk,i Kk,i X k,i ……………………………….....…….………..(4.20)

Where;

γk,i denotes the activity coefficient of component i in stage k. Kk,i denotes the vapor-liquid equilibrium constant.

- 54 -

5. Constraint equations:

∑ =−i

ikY 01, …………………………………………………….………..(4.21)

∑ =−i

ikX 01, …………………………………………………….…..…..(4.22)

The unknowns, alternatively referred to as iteration or primitive variables:

(X, Y, L, V)i , where i= 1, NT are solved for directly using an augmented Newton-Raphson method. Specification equations involving the iteration variables are added directly to the above equations and solved simultaneously.

4.6 MODELING AND SIMULATION FRAMEWORK

It has been assumed in this work that the production of MTBE using

PRO/II is a steady-state process. Therefore, the necessary process flowsheet

for production of MTBE by Reactive Distillation has been built and developed

using PRO/II as shown in Figure 4.5 , the detailed procedure is described step

by step in App. (B). All required files (Model file and process input file) are

given in App. (A) and App. (C), respectively.

55

Figure 4.5 MTBE Process Flowsheet PRO/II .

56

CHAPTER -5-

RESULTS AND DISCUSSION

57

CHAPTER -5- 5.1 RESULTS OF THE SIMULATION WORK

In this section, the outcome of simulation work will be presented for each

equipment with some demonstrated graphs. The used software which is PRO/II

can extract the results on excel file after running the model as shown below.

Table 5.1 Summary Table Stream Name 1 2 3 5 6 7Stream Description MEOH FEED OLEFINS T-1 OVHD

Phase Liquid Liquid Liquid Liquid Liquid Liquid

Temperature (C?) 16.0000000 16.0000000 43.5000000 72.0000000 43.5000000 129.0753174Pressure, bar 16.2000008 16.2000008 15.8549995 14.8199997 6.2100000 6.9749999

Flowrate, Kgmol/hr 277.5000000 850.0000000 1131.8957520 869.3738403 575.9934692 278.0000000

Composition NC4 0.0000000 0.0900000 0.0675875 0.0879966 0.1328171 0.0000005 IC4 0.0000000 0.4100000 0.3078933 0.4008669 0.6050468 0.0000001 1BUTENE 0.0000000 0.0700000 0.0525667 0.0684401 0.1032997 0.0000002 BTC2 0.0000000 0.0400000 0.0300382 0.0391087 0.0590279 0.0000014 BTT2 0.0000000 0.0600000 0.0450572 0.0586630 0.0885424 0.0000007 IBTE 0.0000000 0.3300000 0.2478143 0.0202760 0.0039012 0.0000000 MTBE 0.0000000 0.0000000 0.0000000 0.3013479 0.0000071 0.9976990 MEOH 1.0000000 0.0000000 0.2488773 0.0226821 0.0073578 0.0003628 TBA 0.0000000 0.0000000 0.0000000 0.0002155 0.0000000 0.0006739 H2O 0.0000000 0.0000000 0.0001655 0.0000000 0.0000000 0.0000000 DIB 0.0000000 0.0000000 0.0000000 0.0004033 0.0000000 0.0012612

Tray Number0 6.0 12.0 18.0 24.0 30.0

Frac

tion

0

0.20

0.40

0.60

0.80

1.00

COLUMN T-1

Liquid Fraction of IBTELiquid Fraction of MEOHLiquid Fraction of MTBELiquid Fraction of NC4

Figure 5.1 Feed and products concentration profile among column trays.

58

Table 5.2 Summary of MTBE column T-1 Column T-1 Summary

Net Flow RatesTray Temp. Pressure Liquid Vapor Feed Product Duties

C BAR KG-MOL/HR M*KJ/HR1 43.5 6.21 633.6 576.0 -23.66552 50.6 6.21 661.6 1209.53 51.3 6.24 659.7 1237.64 51.7 6.26 658.5 1235.75 52.1 6.29 657.3 1234.46 52.4 6.32 655.6 1233.37 52.8 6.35 651.6 1231.58 53.5 6.37 636.0 1227.69 54.5 6.40 614.4 1214.5

10 56.2 6.43 582.6 1195.511 59.1 6.46 539.7 1166.412 63.5 6.48 494.3 1126.813 68.7 6.51 462.8 1084.614 73.4 6.54 445.4 1054.215 76.9 6.57 1206.9 1036.7 869.416 91.0 6.59 1180.5 928.917 105.5 6.62 1202.3 902.518 115.4 6.65 1237.5 924.319 121.0 6.67 1264.7 959.520 123.9 6.70 1281.9 986.721 125.5 6.73 1292.5 1003.922 126.5 6.76 1299.2 1014.523 127.1 6.78 1303.8 1021.224 127.5 6.81 1307.2 1025.825 127.9 6.84 1309.8 1029.226 128.2 6.87 1311.9 1031.827 128.4 6.89 1313.8 1033.928 128.6 6.92 1315.4 1035.829 128.8 6.95 1316.9 1037.430 129.1 6.97 1038.9 278.0 23.7889

13579

11131517192123252729

0.0 20.0 40.0 60.0 80.0 100.0 120.0 140.0Temperature, C

Tray

13579

11131517192123252729

0.0 200.0 400.0 600.0 800.0 1000.0 1200.0 1400.0

Vapor Rate, KG-MOL/HR

Tray

0.0 200.0 400.0 600.0 800.0 1000.0 1200.0 1400.0

Liquid Rate, KG-MOL/HRVapor

Liquid Figure 5.2 Temp Profile of column T-1 Figure 5.3 Rates Profile of column T-1

59

13579

11131517192123252729

0.000 5.000 10.000 15.000 20.000 25.000

Vapor Density, KG/M3

Tray

520.000 540.000 560.000 580.000 600.000 620.000

Liquid Density, KG/M3Vapor

Liquid

13579

11131517192123252729

0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000

Vapor Viscosity, PAS

Tray

0.0001 0.0001 0.0001 0.0001 0.0002 0.0002 0.0002

Liquid Viscosity, PASVapor

Liquid Figure 5.4 Density Profile of column T-1 Figure 5.5 Viscosity Profile of column T-1 Table 5.3 Summary of water wash column T-2

Column T-2 Summary Net Flow Rates Tray Temp. Pressure Liquid Vapor Feed Product Duties C BAR KG-MOL/HR M*KJ/HR

1 38.2 7.92 947.2 375.0 572.1 2 38.2 7.92 947.2 0.0 3 38.2 7.92 947.3 0.0 4 38.2 7.92 948.1 0.0 5 38.2 7.92 0.0 576.0 378.8

1

2

3

4

5

38.2 38.2 38.2 38.2 38.2 38.3Tem perature, C

Tray

1

2

3

4

5

0.0 0.2 0.4 0.6 0.8 1.0

Vapor Rate, KG-MOL/HR

Tray

947.0 947.2 947.4 947.6 947.8 948.0 948.2

Liquid Rate, KG-MOL/HRVapor

Liquid Figure 5.6 Temp Profile of column T-2 Figure 5.7 Rates Profile of column T-2

60

1

2

3

4

5

0.000 0.200 0.400 0.600 0.800 1.000

Vapor Density, KG/M3

Tray

603.700 603.800 603.900 604.000 604.100 604.200 604.300 604.400

Liquid Density, KG/M3Vapor

Liquid

1

2

3

4

5

0.0000 0.2000 0.4000 0.6000 0.8000 1.0000

Vapor Viscosity, PAS

Tray

0.0003 0.0003 0.0003 0.0003 0.0003 0.0003

Liquid Viscosity, PASVapor

Liquid Figure 5.8 Density Profile of column T-2 Figure 5.9 Viscosity Profile of column T-2 Table 5.4 Summary of Methanol Recovery column T-3

Column T-3 SummaryNet Flow Rates

Tray Temp. Pressure Liquid Vapor Feed Product DutiesC BAR KG-MOL/HR M*KJ/HR

1 30.0 1.03 96.9 4.4 -4.07302 74.1 1.38 105.9 101.33 76.5 1.40 103.7 110.34 81.7 1.42 99.4 108.15 92.8 1.44 95.8 103.86 104.6 1.46 95.9 100.27 108.8 1.48 96.2 100.38 109.9 1.50 96.3 100.69 110.4 1.51 96.4 100.7

10 110.8 1.53 483.1 100.7 378.811 111.6 1.55 483.6 108.712 112.2 1.57 484.0 109.213 112.8 1.59 484.4 109.614 113.3 1.61 484.8 110.015 113.7 1.63 485.1 110.416 114.2 1.65 485.5 110.717 114.6 1.67 485.8 111.018 114.9 1.69 486.1 111.319 115.3 1.71 486.4 111.620 115.7 1.73 111.9 374.4 4.5077

1

3

5

7

9

11

13

15

17

19

0.0 20.0 40.0 60.0 80.0 100.0 120.0 140.0Temperature, C

Tray

13579

1113151719

98.0 100.0 102.0 104.0 106.0 108.0 110.0 112.0 114.0

Vapor Rate, KG-MOL/HR

Tray

0.0 100.0 200.0 300.0 400.0 500.0 600.0

Liquid Rate, KG-MOL/HRVapor

Liquid Figure 5.10 Temp Profile of column T-3 Figure 5.11 Rates Profile of column T-3

61

13579

1113151719

0.000 0.500 1.000 1.500 2.000

Vapor Density, KG/M3

Tray

0.000 200.000 400.000 600.000 800.000 1000.000

Liquid Density, KG/M3Vapor

Liquid

13579

1113151719

0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000

Vapor Viscosity, PAS

Tray

0.0000 0.0001 0.0002 0.0003 0.0004 0.0005 0.0006

Liquid Viscosity, PASVapor

Liquid Figure 5.12 Density Profile of column T-3 Figure 5.13 Viscosity Profile of column T-3 Table 5.5 Overall Report Stream (Summary) UOM 1 10 11 12 13 14Name 1 10 11 12 13 14Description MEOH FEED 10 C4'S 12 13 14Phase Liquid Liquid Liquid Liquid Mixed MixedThermodynamic System S1 S1 S1 S1 S1 S1Total Molar Rate kg-mol / hr 277.5 375.0209471 572.136588 378.87777 378.87777 378.87777Total Mass Rate kg / hr 8891.699411 6756.40571 32942.70972 6886.158242 6886.158242 6886.158242Temperature C 15.9999939 37.9999939 38.1693034 38.1787967 98.9999939 99.0980516Pressure bar 16.2 7.93 7.92 7.92 7.575 2.41Total Molecular Weight 32.04216003 18.01607554 57.57840072 18.17514456 18.17514456 18.17514456Total Specific Enthalpy kJ / kg 23.45157451 169.3191479 92.15106329 166.9490502 422.8609698 422.8609697Total Cp J/kg-K 2454.101878 4366.688375 2503.383034 4341.893717 0 0Total Molar Component Rates kg-mol / hr NC4 0 0 76.4962078 0.005569965 0.005569965 0.005569965 IC4 0 0 348.4874116 0.015603494 0.015603494 0.015603494 1BUTENE 0 0 59.49979477 0.000171456 0.000171456 0.000171456 BTC2 0 0 33.99945592 0.000241614 0.000241614 0.000241614 BTT2 0 0 50.99964646 0.000192594 0.000192594 0.000192594 IBTE 0 0 2.247034765 5.202E-06 5.202E-06 5.202E-06 MTBE 0 1.68576E-18 0.004041775 2.16251E-05 2.16251E-05 2.16251E-05 MEOH 277.5 0.021275489 0.003515716 4.255771777 4.255771777 4.255771777 TBA 0 3.14002E-36 2.8833E-10 2.73157E-13 2.73157E-13 2.73157E-13 H2O 0 374.9996714 0.399479196 374.6001924 374.6001924 374.6001924 DIB 0 0 1.03273E-11 1.28821E-23 1.28821E-23 1.28821E-23Total Molar Component Fractions fraction NC4 0 0 0.133702702 1.47012E-05 1.47012E-05 1.47012E-05 IC4 0 0 0.60909828 4.11835E-05 4.11835E-05 4.11835E-05 1BUTENE 0 0 0.103995787 4.52538E-07 4.52538E-07 4.52538E-07 BTC2 0 0 0.059425418 6.3771E-07 6.3771E-07 6.3771E-07 BTT2 0 0 0.089138936 5.08327E-07 5.08327E-07 5.08327E-07 IBTE 0 0 0.003927445 1.373E-08 1.373E-08 1.373E-08 MTBE 0 4.49511E-21 7.06435E-06 5.70768E-08 5.70768E-08 5.70768E-08 MEOH 1 5.67315E-05 6.14489E-06 0.011232572 0.011232572 0.011232572 TBA 0 8.37292E-39 5.03954E-13 7.20964E-16 7.20964E-16 7.20964E-16 H2O 0 0.999943269 0.000698223 0.988709874 0.988709874 0.988709874 DIB 0 0 1.80503E-14 3.40006E-26 3.40006E-26 3.40006E-26Vapor Molar Rate kg-mol / hr N/A N/A N/A N/A 0.000577937 0.029560113Vapor Mass Rate kg / hr N/A N/A N/A N/A 0.030431601 1.20838919Vapor Molecular Weight N/A N/A N/A N/A 52.65556269 40.87904483Vapor Mole Fraction fraction N/A N/A N/A N/A 1.52539E-06 7.80202E-05Vapor Specific Enthalpy kJ / kg N/A N/A N/A N/A 629.8510681 950.3195163Vapor CP J/kg-K N/A N/A N/A N/A 2096.037884 2012.784279Vapor Composition fraction NC4 N/A N/A N/A N/A 0.13081499 0.122759833 IC4 N/A N/A N/A N/A 0.711658135 0.422089572 1BUTENE N/A N/A N/A N/A 0.007379891 0.004578206 BTC2 N/A N/A N/A N/A 0.004576501 0.005051828 BTT2 N/A N/A N/A N/A 0.006164664 0.004781536 IBTE N/A N/A N/A N/A 0.000268759 0.000144091 MTBE N/A N/A N/A N/A 8.803E-07 2.36786E-06 MEOH N/A N/A N/A N/A 0.010672162 0.032527184 TBA N/A N/A N/A N/A 7.31467E-14 1.93873E-13 H2O N/A N/A N/A N/A 0.128464018 0.408065382 DIB N/A N/A N/A N/A 2.18059E-20 4.35717E-22Vapor Component Rate kg-mol / hr NC4 N/A N/A N/A N/A 7.56028E-05 0.003628795 IC4 N/A N/A N/A N/A 0.000411294 0.012477015 1BUTENE N/A N/A N/A N/A 4.26511E-06 0.000135332 BTC2 N/A N/A N/A N/A 2.64493E-06 0.000149333 BTT2 N/A N/A N/A N/A 3.56279E-06 0.000141343 IBTE N/A N/A N/A N/A 1.55326E-07 4.25934E-06 MTBE N/A N/A N/A N/A 5.08758E-10 6.99943E-08 MEOH N/A N/A N/A N/A 6.16784E-06 0.000961507 TBA N/A N/A N/A N/A 4.22742E-17 5.7309E-15 H2O N/A N/A N/A N/A 7.42441E-05 0.012062459 DIB N/A N/A N/A N/A 1.26024E-23 1.28798E-23Liquid Molar Rate kg-mol / hr 277.5 375.0209471 572.136588 378.87777 378.8771922 378.84821Liquid Mass Rate kg / hr 8891.699411 6756.40571 32942.70972 6886.158242 6886.127812 6884.949852Liquid Molecular Weight 32.04216003 18.01607554 57.57840072 18.17514456 18.17509196 18.17337306Liquid Mole Fraction fraction 1 1 1 1 0.999998475 0.99992198Liquid Specific Enthalpy kJ / kg 23.45157451 169.3191479 92.15106329 166.9490502 422.860055 422.7683945Liquid CP J/kg-K 2454.101878 4366.688375 2503.383034 4341.893717 4133.847462 4134.720639Liquid Composition fraction NC4 0 0 0.133702702 1.47012E-05 1.45017E-05 5.12388E-06

62

5.2 RESULTS DISCUSSION

The results of this simulation shown in previous section indicate that the

overall conversion of IBTE is 99.2%, the conversion can be calculated as

following:

% 9.29 (%) Conv.

100280.5

2.32 - 280.5 (%) Conv.

100(IBTE)

(IBTE) - (IBTE) (%) Conv.

inlet

outletinlet

=

×=

×=

Also with a selectivity to MTBE of 99.7%. In the reactive distillation

column itself, 87.2% of the IBTE fed to the column is converted to MTBE. The

MTBE product is 99.77% pure and needs no further purification.

There are a number of factors that affect the overall conversion rate of

IBTE. Some of these are:

• MeOH to IBTE ratio.

• Number of reaction trays.

• Type of catalyst used.

• Reflux ratio

Note, however, that while the IBTE conversion in the conversion reactors

increases as the MEOH:IBTE ratio is increased, the overall IBTE conversion

reaches a maximum, then decreases as the MEOH:IBTE ratio is increased.

This is due to the fact that more MTBE product is carried upward through the

column stripping section into the reaction trays. This promotes the reverse

reaction of MTBE to methanol and IBTE, thus reducing the overall conversion

of IBTE.

5.3 CASE STUDIES

The validated MTBE primary reactor and reactive distillation column models

were used to develop new process flow diagrams for the MTBE unit.

Concerning the evaluation of plant performance, there are some parameters

were taken into account for the plant optimization as follows:

63

o MeOH to IBTE ratio.

o Number of trays.

o Location of feeds into RD column.

o Reflux ratio of RD column.

o Number of reaction trays.

Case Title: Case-1: Changing the Methanol/ Isobutylene Ratio Case Clarifications:

• R1 = 0.90 ==> (MeOH)Feed= 252.5 kgmol/hr • R2 = 0.95 ==> (MeOH)Feed= 266.5 kgmol/hr • R3 = 0.98 ==> (MeOH)Feed= 277.5 kgmol/hr • R4 = 1.0 ==> (MeOH)Feed= 280.5 kgmol/hr (Base case) • R5 = 1.05 ==> (MeOH)Feed= 294.5 kgmol/hr

MeOH/IBTE Ratio 0.900 0.950 0.980 1.000 1.050 Conversion(%) 87.500 92.300 96.970 99.200 96.970 MTBE Purity(wt%) 87.900 92.700 99.600 99.700 99.600

Case Observations and Comments: - The overall conversion of IBTE into MTBE is increased when the Methanol/Isobutylene ratio increased but it dropped when the ratio exceed 1.0 as shown in the figure 5.14. - The purity of MTBE is increased when the Methanol/Isobutylene ratio increased as shown in the figure 5.14. The last two runs are showing almost stable purity. - The optimum Methanol/Isobutylene ratio is 1.0 as demonstrated in figure 5.14.

Figure 5.14 Results of changing the MeOH/IBTE ratio

64

Case Title: Case-2: Changing the Number of trays for RD column Case Clarifications:

• R1 = The RD column consists of 20 trays • R2 = The RD column consists of 25 trays • R3 = The RD column consists of 30 trays (Base case) • R4 = The RD column consists of 32 trays • R5 = The RD column consists of 35 trays • R6 = The RD column consists of 40 trays

RD Trays 20 25 30 32 35 40 Conversion(%) 97.00 98.00 99.20 99.20 99.20 99.20 MTBE Purity(wt%) 97.20 99.20 99.70 99.70 99.80 99.80

Case Observations and Comments: - The overall conversion of IBTE to MTBE is increased when the number of trays increased as shown in the figure 5.15. - The purity of MTBE is increased when the number of trays increased as shown in the figure 5.15. The last three runs are showing almost stable purity. - The last two runs are almost comparable which indicates that the both parameters will get stable. - The optimum number of trays are 30 (The max limit to have the desired results).

Figure 5.15 Results of changing the number of trays of RD column

65

Case Title: Case-3: Changing the Location of feed stream into RD column Case Clarifications:

• R1 = The feed stream entered the RD column at tray # 8 • R2 = The feed stream entered the RD column at tray # 10 • R3 = The feed stream entered the RD column at tray #12 • R4 = The feed stream entered the RD column at tray #15 (Base case) • R5 = The feed stream entered the RD column at tray #17

Note:- The Reaction zone applied in this case is from 11-13 trays of RD column Feed stream to tray 8 10 12 15 17 Conversion(%) 94.50 95.30 97.91 99.20 99.20 MTBE Purity(wt%) 94.50 95.80 98.40 99.70 99.70

Case Observations and Comments: - The overall conversion of IBTE to MTBE is increased when the location of feed stream changed to the next tray as shown in the figure 5.16. - The purity of MTBE is increased when the location of feed stream changed to the next tray as shown in the figure 5.16. - The last two runs are almost comparable which indicates that the both parameters will get stable. - The optimum location for the feed stream is tray 15 (The middle of the column).

Figure 5.16 Results of changing the location of feed stream into RD column

66

Case Title: Case-4: Changing the Reflux Ratio for RD column Case Clarifications:

• R1 = The Reflux Ratio is equal to 0.8 • R2 = The Reflux Ratio is equal to 0.9 • R3 = The Reflux Ratio is equal to 1.0 • R4 = The Reflux Ratio is equal to 1.1 (Base case) • R5 = The Reflux Ratio is equal to 1.2

Reflux Ratio 0.8 0.9 1.0 1.1 1.2 Conversion(%) 98.33 98.75 99.16 99.20 99.20 MTBE Purity(wt%) 98.80 99.30 99.60 99.70 99.80

Case Observations and Comments: - The overall conversion of IBTE to MTBE is increased when the Reflux Ratio increased as shown in the figure 5.17. - The purity of MTBE is increased when the Reflux Ratio increased as shown in the figure 5.17. - The overall conversion of IBTE into MTBE is getting stable for the last three runs at the value 99.2 %.

Case-4: Changing the Reflux Ratio for RD column

98.20

98.40

98.60

98.80

99.00

99.20

99.40

99.60

99.80

100.00

0.6 0.7 0.7 0.8 0.8 0.9 0.9 1.0 1.0 1.1 1.1 1.2 1.2 1.3

Reflux Ratio

Conversion(%) MTBE Puirty(wt%)

Figure 5.17 Results of changing the Reflux Ratio

67

Case Title: Case-5: Changing the Location of reaction zone in RD column Case Clarifications:

• R1 = The Reaction zone from 9-13 trays • R2 = The Reaction zone from 10-13 trays • R3 = The Reaction zone from 11-13 trays (Base case) • R4= The Reaction zone from 12-13 trays • R5 = The Reaction zone only tray # 13

Reaction zone 9-13 10-13 11-13 12-13 13 Conversion(%) 99.20 99.20 99.20 99.17 98.50 MTBE Purity(wt%) 99.80 99.80 99.70 99.70 99.20

Case Observations and Comments: - The overall conversion of IBTE to MTBE is decreased when the location of reaction zone get changed as shown in the figure 5.18. - The purity of MTBE is stable then decreased when the reaction zone get small (only one tray) as shown in the figure 5.18. - The optimum location for the reaction zone is tray 11-13 (The middle of the packing bed).

Case-5: Changing the Location of reaction zone in RD column

98.40

98.60

98.80

99.00

99.20

99.40

99.60

99.80

100.00

0 1 2 3 4 5 6

Reaction Zone

Conversion(%) MTBE Puirty(wt%)

Figure 5.18 Results of changing the location of reaction zone

68

CHAPTER -6-

CONCLUSIONS AND RECOMMENDATIONS

69

6.1 CONCLUSIONS

• The combination of reaction and distillation helps in achieving products of

higher purity and higher conversion of reactants as compared to old

conventional processes.

• As shown in this study it is possible to obtain an overall conversion of IBTE

around 99.2% with purity of MTBE almost 99.7%.

• The mathematical model developed has shown satisfactory results in

simulating a reactive distillation column for the etherification of methanol and

isobutylene to form MTBE.

• The model and computational technique has also been successful in

determining the effect of various design variables namely reflux ratio, feed plate

location, reaction zone, Methanol/Isobutylene molar ratio.

• The selectivity of IBTE to MTBE is 99.7%. In the reactive distillation column

itself, 87.2% of the IBTE fed to the column is converted to MTBE. The MTBE

product is 99.77% pure and needs no further purification.

• The overall conversion of IBTE to MTBE is increased when the

Methanol/Isobutylene ratio increased but it dropped when the ratio exceed 1.0

as shown in the figure 5.14.

• The purity of MTBE is increased when the Methanol/Isobutylene ratio

increased as shown in the figure 5.14. The last two runs are showing almost

stable purity.

• The overall conversion of IBTE to MTBE is increased when the number of trays

increased as shown in the figure 5.15.

• The purity of MTBE is increased when the number of trays increased as shown

in the figure 5.15. The last three runs are showing almost stable purity.

• The optimum number of trays are 30 (The max limit to have the desired

results).

• The overall conversion of IBTE to MTBE is increased when the location of feed

stream changed to the next tray as shown in the figure 5.16.

• The purity of MTBE is increased when the location of feed stream changed to

the next tray as shown in the figure 5.16. but the last two runs are almost

comparable which indicates that the both parameters will get stable.

70

• The optimum location for the feed stream is tray 15 (The middle of the

column).

• The overall conversion of IBTE to MTBE is increased when the Reflux Ratio

increased as shown in the figure 5.17.

• The purity of MTBE is increased when the Reflux Ratio increased as shown in

the figure 5.17.

• The overall conversion of IBTE to MTBE is decreased when the location of

reaction zone get changed as shown in the figure 5.18.

• The purity of MTBE is stable then decreased when the reaction zone get small

(only one tray) as shown in the figure 5.18.

• The optimum location for the reaction zone is tray 11-13 (The middle of the

packing bed).

• It is not recommended to exceed the MEOH/IBTE ratio more than 1.0 to

achieve the desired conversion as well as the MTBE purity, this due to

limitation observed from the simulation results.

• It is not recommended to increase the number of trays more than 30 trays

since the conversion of IBTE and purity of MTBE were not increased much

when the trays increased up to 40. this will also maintain the capital cost.

• It is recommended to keep the feed stream into RD column at stage no. 15

which represents the middle of the column.

• It is recommended to keep the reflux ration equal to 1.1 since the conversion

and purity of MTBE were no effected when it increased more than 1.1.

• It is recommended to keep the location of the reaction zone in the space of tray

11-13 (The middle of the packing bed).

• It is recommended to keep close monitoring system on the recycled methanol

to the reactor to monitor the moisture content in the stream to avoid forming

impurities such as t-butanol.

• It is recommended to maintain a sufficient amount of C4s in the column feed to

have smooth operation for the MTBE RD column. Also to form azeotropes with

the methanol in the feed.

71

6.2 RECOMMENDATIONS

• It is recommended to perform an Economic Evaluation study for RD process

against the conventional process.

• It is recommended also, to perform a dynamic study.

• It is recommended to validate the results by using another data or another

simulation software.

• It is recommended to perform an effective control study on control structures

for RD column to have effective system to control the critical parameters.

• It is recommended to study the possibility of having quality performance system

for the feed and recycled streams to avoid forming undesired products.

• It is recommended to use this model as a tool for design, optimization and

control of Reactive Distillation processes.

72

BIBLIOGRAPHY

73

REFERENCES

1- J. H. Gregor, S.T. Bakas, and M. A. Ulowetz, "Converting field butanes into MTBE",

Seventy-First Annual Convention Gas Processors Association, Anaheim, California,

1992, sku P1992.36.

2- Sandra Viviana, " Computer Simulation Studies for the production of 7-Tertadecene by

Reactive Distillation", Mississippi State University , USA, 2 August 2003.

3- D. N. Nakamura, “HP in Processing,” Hydrocarbon Processing, Vol. 77(1), 15, 1998.

4- CHEM System Inc, "Methyl tertiary-Butyl Ether (MTBE)", PERP 94/95-4, 1996.

5- M. F. Doherty, and M. F. Malone, "Conceptual Design of Distillation Systems",

McGraw Hill, New York, 2001

6- E. J. Chang and S. M. Leiby, “Ethers Help Gasoline Quality,” Hydrocarbon Processing,

Vol. 71(2), 41–44, 1992.

7- J. D. Seader, "Separation Process Principles", John Wiley and Sons, Inc., New York,

1998.

8- J. L. DeGmo, V. N. Pdekar, and V. Pinjala, "Consider Reactive Distillation," Chem

Eng. Prog. 88, 43-50, 1992.

9- R. K. Toghiani, H. Toghiani, and G. Venkateswarlu, "Vapor-liquid Equilibria for methyl

ter-butyl ether+methanol and ter-amyl methyl ether+methanol", Fluid Phase Equilibria

122, 157-168, 1996.

10- A. Tuchlenski, A. Beckmann, D. Reusch, R. Dussel, U. Weidlich, and R. Janowsky,

"Reactive Distillation-Industrial Applications, Process Design and Scale-up", Chemical

Engineering Science 56, 387-394, 2001.

11- Krishna, R. and Sie, S.T. "Strategies for Multiphase Reactor Selection", Chem.Engng

Sci., 49, 4029-4065, 1994.

12- Agreda, V.H., Partin, L.R. and Heise, W.H. "High-purity methyl acetate via reactive

distillation", Chem. Eng. Progress, 2, 40 – 46, 1990.

13- Doherty, M.F. and G. Buzad, "Reactive Distillation by Design", Chem. Eng. Res .&

Design., Trans I.Chem.E., Part A, 70, 448-458, 1992.

14- Siirola, J.J., "An Industrial Perspective on Process Synthesis", A.I.Ch.E. Symposium

Series No. 304, 91,222 – 233, 1995.

15- Stichlmair, J.G. and Frey, T., "Reactive distillation processes", Chemical and

Engineering Technology, 22, 95-103, 1999.

16- Ciric, A.R., and D. Gu., "Synthesis of nonequilibrium reactive distillation by MINLP

optimization", A.I.Ch.E.J., 40, 1479 – 1487,1994

74

17- Gonzalez, J. C., & Fair, J. R., "Preparation of tertiary amyl alcohol in a reactive

distillation column.1 Reaction kinetics, chemical equilibrium, and mass transfer

issues". Industrial and Engineering chemistry Research, 36, 3833-3844,1997.

18- Velo, E., Puigianer, L. & Recasens, F., "Inhibition by product in the liquid-phase

hydration of iso-butane to tert-butyl alcohol – kinetics & equilibrium studies", Industrial

and Engineering chemistry Research, 27, 2224-2231, 1988.

19- Shoemaker, J.D. and E.M. Jones, "Cumene by catalytic distillation", Hydrocarbon

Processing, June 1987, 57 – 58, 1987.

20- Subawalla, H., Gonzalez, J.C., Seibert, A.F., and Fair, J.R., "Capacity and Efficiency

of Reactive Distillation Bale Packing: Modeling and Experimental Validation", Ind. Eng.

Chem. Res., 36, 3821-3832, 1997

21- Van Gulijk, C. , "Using computational fluid dynamics to calculate transversal dispersion

in a structured packed bed", Comput. Chem. Eng., 22, S767-S770, 1998.

22- Higler, A., Krishna, R., Ellenberger, J. and Taylor, R., "Counter-current operation of a

structured Catalytically Packed Bed Reactor: Liquid phase mixing and mass transfer",

Chem. Eng. Sci., in press, 1999c.

23- Frey, S. J., Ozmen, S. M., Hamm, D. A., Pinjala, V., & DeGarmo, J. L. , "Advanced

ether technology commercialized". Paper presented at the world conference on clean

fuels and air quality control, Washington, DC, 1993.

24- G. Matthias, M. Ramona, Dragomir and J. Megan, "Conceptual design of reactive

distillation columns using stage composition lines", Chemical Engineering Science 60,

5049-5068, 2005.

25- V. B. Shah, D. Bluck, J. W. Kovach III, "The sensitivity of the design and operability of

the MTBE processes with respect to changes in reaction parameters and process

configurations", Refining LNG and Petrochmasia 1994.

26- P. S. Bhanu, S. Ram, M.V Pavan, and K. Nitin, "Steady state analysis for reactive

distillation control: An MTBE case study", Journal of Loss Prevention in the process

Industries 18, 283-292, 2005.

27- C. A. Rivera and G. Johan, "Feasibility of Equilibrium controlled reactive distillation

process: application of residue curve mapping", Computers and Chemical Engineering

28, 17-25, 2004.

28- Smith Jr., L.A. " U.S. Patent 4443559", Catalytic distillation structure, 1984.

29- Johnson, K.H. , "U.S. Patent 5189001", Catalytic distillation structure,1993.

30- Nelson, P. A., "Countercurrent equilibrium stage separation with reaction". American

Institute of chemical Engineers Jornal, 17, 1043-1049, 1971.

75

31- Suzuki, I., Yagi, H., Komatsu, H., and Hirata, M., " Calculation of multicomponent

distillation accompanied by chemical reaction", J. Chem. Eng. Jap., 4, 26-33, 1971.

32- Venkataraman, S., Chan, W. K., & Boston, J. F., "Reactive distillation using ASPEN

PLUS". Chemical Engineering Process, 86(8), 45-54, 1990.

33- Jelink, J., & Halvacek, V., "Steady state Countercurrent equilibrium stage separation

with chemical reaction by relaxation method". Chemical Engineering

communications,2, 79-85, 1976.

34- Saito, S., Michishita, T., & Maeda, S. , " Separation of meta- and para-xylene mixture

by distillation accompanied by chemical reaction". Journal of Chemical Engineering

Japan, 4, 37-43, 1971.

35- Grosser, J.H., Doherty, M.F. and Malone, M.F., "Modeling of reactive distillation

systems", Ind. Chem. Eng. Res., 26, 983-989, 1987.

36- Kamath R., Qi Z., Sundmacher K., Aghalayam P., Mahajani S., "Modelling and

Optimization for Energy Saving and Pollution Reduction", 8th conference on Process

Integration, The Italian Association of Chemical Engineering (AIDIC) 309, 2005.

37- Al-Jarallah, A. M., M. A. B. Siddiqui and A. K. K. Lee, “Kinetics of Methyl Tertiary Butyl

Ether Synthesis Catalyzed by Ion Exchange Resin”, Canadian J. Chem. Eng. , 66,

802-7, 1988.

38- A. Abufares, and P.L. Douglas, "Mathematical Modeling and Simulation of an MTBE

Catalytic Distillation Process using Speedup and Aspenplus", Chemical Engineering

Research and Design 79 (A1), 3-12, 1995.

39- J. M. Bader, S. Guesneux, Axens, "Increase MTBE plant productivity", Hydrocarbon

Processing, October 2005.

40- M. H. M. Reis; L. F. S. Mascolo; M. R. Wolf-Maciel, " Development of a robust

algorithm to compute reactive azeotropes ", Brazilian Journal of Chemical

Engineering, vol.23 no.3 , July/Sept. 2006.

76

APPENDIX (A)

REACTION MODEL CODE

77

APPENDIX (A) : MTBE Reaction Model Code

In this section the entire code of reaction kinetics will be shown below along with a

description of purpose. The first portion of the defines real and integer variables REAL KS , KA , KC , KALJ , KREH1 , KREH2 , KIZQ , KEQREF INTEGER IBTE , MEOH , MTBE

Next, the indices for the components are set, a value is given for the gas constant in

J/gm-mole K, and the basis for the temperature values in the procedure is set to an

absolute basis. In addition, the temperature and composition rate derivatives are

initially set equal to zero. $ $ INITIALIZE DATA: $ SET INDEXES FOR COMPONENTS $ DEFINE GAS CONSTANT IN JOULES/GM-MOLE K $ NOTE: R COULD HAVE BEEN RETRIEVED IN INPUT UNITS BY R=RGAS. $ HOWEVER, SINCE THE REACTION BASIS WON'T CHANGE, AND $ RGAS WILL CHANGE WITH THE DEFAULT UNITS, THIS $ ELIMINATES ONE POSSIBLE SOURCE OF ERROR. $ INITIALIZE THE LOCAL VARIABLE TK TO THE ABSOLUTE TEMPERATURE. $ NOTE: THE TEMPERATURE BASIS FOR THE FLOWSHET MUST BE C OR K. $ SET TEMPERATURE AND COMPOSITION DERIVATIVES TO ZERO. $ IBTE = 6 MTBE = 7 MEOH = 8 $ R = 8.314 $ TK = RTABS $ DO 1000 I1 = 1,NOR DRDT(I1) = 0.0 DO 1000 I2 = 1,NOC 1000 DRDX(I2,I1) = 0.0

The surface reaction rate constant, ks, and the equilibrium adsorption constants, KA,

and KB, are calculated using the expressions given previously as (1a), (1b), and (1c).

$ CALCULATE THE SURFACE REACTION RATE CONSTANT, KS, AND THE $ EQUILIBRIUM ADSORPTION CONSTANTS KA AND KB. THE ACTIVATION $ ENERGY IS IN J/GM-MOLE. $ UNITS: KS - (GM-MOLE / GM CATALYST)**1.5 /HOUR $ KA - GM-CATALYST / MOLE $ KC - GM-CATALYST / MOLE $ KS = 1.2E+13*EXP(-87900.0/(R*TK)) KA = 5.1E-13*EXP( 97500.0/(R*TK)) KC = 1.6E-16*EXP(119000.0/(R*TK))

78

Next, the derivatives of these constants are computed and are used later on in

calculating the rate derivatives. $ DKSDT = KS * 87900.0 / R / (TK*TK) DKADT = KA * (-1.0) * 97500.0 / R / (TK*TK) DKCDT = KC * (-1.0) * 119000.0 / R / (TK*TK)

Then the bulk concentration of components A, B, and °C per gram of catalyst

(RHOA, RHOB, and RHOC) are determined from the liquid mole fractions of the

components (XLIQ), the density of the liquid, and the catalyst loading (GCAT) in g/l.

Note that the liquid density, DENS, obtained directly from PRO/II using the predefined

variables, RLMRAT and RLVRAT, is in the user-specified units of kg-moles/m3 (SI

units). Our basis for calculations is gm-moles/l and the conversion factor between

these kg-moles/m3 and gm-moles/l is 1.0. Also, note that the value of GCAT used here

is 12.4 g/l. This value is used because it is the catalyst loading at which data for the Al-

Jarallah rate equation was collected. $ ---- CALCULATE THE EQUILIBRIUM CONSTANT. $ $ UNITS - (GM-MOLES/GM-CATALYST)/HOUR $ PHASE - LIQUID PHASE REACTION $ CATALYST - ION EXCHANGE RESIN AMBERLYST 15, $ THE EQUILIBRIUM SHOULD BE INDEPENDANT OF THE CATALYST

Expressions for the equilibrium constant and its derivative as functions of temperature

are provided based on equilibrium data published by Al-Jarallah et al. $ KALJ = EXP(-17.31715+(7196.776/TK)) $ DKALJDT = - KALJ * 7196.776 / (TK*TK)

Then the reaction rate and rate derivatives with respect to temperature and

composition are determined. $ BULK CONCENRATIONS OF COMPONENTS PER GRAM OF CATALYST, XLCONC IS $ IN MOLES/FLOW VOLUME. XLCONC WILL BE PASSED TO THE PROCEDURE $ IN USER INPUT UNITS. INTERNALLY TO PRO/II, IT IS IN SI UNITS $ (KG-MOLE/CUBIC METER). THE BASIS FOR THESE REACTION EQUATIONS $ IS GM-MOLES/LITER. THE CONVERSION FACTOR FROM INPUT UNITS OF $ KG-MOLES/CUBIC METER TO THE REACTION BASIS OF GM-MOLES/LITER $ IS ONE. THEREFORE, XLCONC CAN BE USED WITH NO CONVERSION. $ RHOA=(XLCONC(MEOH)/GCAT) |-THIS SHOULD BE EQUIVALENT TO BELOW. $ RHOB=(XLCONC(IBTE)/GCAT) | IT HAS BEEN WRITTEN EXPLICITLY BELOW $ RHOC=(XLCONC(MTBE)/GCAT) | TO MAKE IT OBVIOUS HOW TO DO THE $ ANALYTICAL DERIVATIVES.

79

$ CALCULATE DENSITY IN MOLES / VOLUME $ GCATX = 12.4 DENS=RLMRAT/RLVRAT RHOA=(XLIQ(MEOH)*DENS/GCATX) RHOB=(XLIQ(IBTE)*DENS/GCATX) RHOC=(XLIQ(MTBE)*DENS/GCATX) $ ---- CALCULATE REACTION RATE AND DERIVATIVES BY TERMS $ ---- UNITS - RATE - GRAM-MOLE / GRAM CATALYST / HR. $ $ DENOMINATOR & DERIVATIVES. $ RDEN = 1.0 + ( KA*RHOA ) + 0.0 + ( KC*RHOC ) DRDDT = RHOA*DKADT + 0.0 + RHOC*DKCDT DRDDME = KA/GCATX*DENS DRDDIB = 0.0 DRDDMT = KC/GCATX*DENS $ $ FIRST FACTOR IN RATE EQUATION. FACT1 = KS *KA/RDEN DFAC1DT = DKSDT*KA/RDEN + KS*DKADT/RDEN - KS*KA/RDEN**2 * DRDDT $ $ SECOND FACTOR IN RATE EQUATION. FACT2 = RHOA*RHOB**0.5 - RHOC**1.5/KALJ DFAC2DT = 0.0 + RHOC**1.5/KALJ**2 * DKALJDT $ COMBINE TERMS TO CALCULATE RATE AND DERIVATIVES. $ -- RATE EQUATION (RATE PER ONE GRAM OF CATALYSIS). RATE = FACT1 * FACT2 $ -- RATE TEMPERATURE DERIVATIVE. DRDT(1) = DFAC1DT * FACT2 & + FACT1 * DFAC2DT $ -- RATE COMPOSITION DERIVATIVES. DRDX(MEOH,1) = -KS*KA/RDEN**2 * DRDDME * FACT2 & + FACT1 * (RHOB**0.5/GCATX*DENS) DRDX(IBTE,1) = -KS*KA/RDEN**2 * DRDDIB * FACT2 & + FACT1 * (RHOA/2.0/RHOB**0.5/GCATX*DENS) DRDX(MTBE,1) = -KS*KA/RDEN**2 * DRDDMT * FACT2 & - FACT1 * (1.5* RHOC**0.5/GCATX/KALJ*DENS)

It is important to note, however, that the rate and rate derivatives calculated above are

computed on a basis of 1 gram of catalyst. The reactive distillation algorithm requires

that these values (RRATES, DRDT, and DRDX) be supplied on a unit reaction

volume basis. Therefore, the rate and rate derivatives are multiplied by the grams of

catalyst per unit volume, GCAT. $ ---- CONVERT RATE EQUATION AND DERIVATIVES TO A STRAIGHT VOLUME BASIS $ ---- BY MULTIPLING THE BASE RATE BY THE GRAMS OF CATALYST/UNIT VOLUME. $ ---- THE RATE IS RETURNED IN INPUT UNITS, KG-MOLES/CUBIC METER/HOUR. $ RRATES(1) = GCAT * RXFACT * RATE DRDT(1) = GCAT * RXFACT * DRDT(1) DRDX(MEOH,1) = GCAT * RXFACT * DRDX(MEOH,1) DRDX(IBTE,1) = GCAT * RXFACT * DRDX(IBTE,1) DRDX(MTBE,1) = GCAT * RXFACT * DRDX(MTBE,1) RETURN

80

APPENDIX (B)

MTBE PROCESS SIMULATION PROCEDURE USING PRO/II

81

APPENDIX (B) : MTBE Process Simulation Procedure using PRO/II Process Simulation General Data

SI units are used in this simulation. The total calculation sequence is specified. The

calculator CAL0 is processed before the MTBE column in order to set the reaction factor

equal to 1.0 on the first pass through the flowsheet.

In order to check the overall material balance, PRO/II is instructed to print out an overall

flowsheet mass balance in the Miscellaneous Report Options menu.

82

Component Data All the components in the simulation are available in the PRO/II databank.

Thermodynamic Data The VLE fractionators are simulated well with PRO/II's modified Soave-Redlich-Kwong

(SRKM) equation of state method. For this method, PRO/II contains extensive, built-in

databanks that encompass binary interaction parameter data for the majority of component

pairs present in this simulation. In this case, however, binary interaction data (kijs) are directly

supplied for 8 component pairs to improve the accuracy of the separations in the columns.

The binary interaction parameters are listed below in Table 1. The thermodynamic set used

for VLE thermodynamics is referred to as SRKM_VLE.

Table 1. Binary Interaction Parameters for SRKM_VLE Comp I NC4 1BUTENE BTC2 BTT2 IBTE MEOH TBA MTBE

Comp j TBA MEOH MEOH MEOH MEOH TBA H20 DIB

kija 0.0469 0.136 0.136 0.136 0.13553 -0.07397 -0.145 0.05785

kjia 0.1260 -0.0323 -0.0323 -0.0323 -0.0322 -0.05522 -0.253 -0.0093

kijb 0.0 0.0 0.0 0.0 0.0 0.0 0.0 -10.144

kjib 0.0 0.0 0.0 0.0 0.0 0.0 0.0 6.17

UOM K K K K K K K K

Transport properties are needed in order to use the rigorous heat exchanger model in

the MEOH recovery section. Transport property calculations are set to pure-component

83

averages by default and can be modified by modifying the thermodynamic method of choice.

Modification is not required in this simulation, as pure-component averages should be used.

The liquid extraction unit is simulated using the SRKM method for VLLE thermodynamics with

binary interaction data again supplied as part of the input. The thermodynamic set used for

VLE thermodynamics is referred to as SRKM_VLLE. Note that the L1KEY component (i.e.,

the predominant component in the L1 liquid phase) is specified as n-butane. The L2KEY

component is specified to be water. Explicitly specifying the key components eliminates the

need for PRO/II to find an appropriate immiscible pair, reducing the computation time. All the

azeotropes are properly predicted.

Stream Data Feed Stream

The mixed C4 feed stream, and the methanol feed stream are specified in the normal

manner, using the compositions and stream conditions given in Table 2.

Recycle Stream The composition of the recycle methanol-water stream from the MEOH recovery

section is estimated initially for the first run through the flowsheet (see Table 2).

84

The amount of wash water in stream 10 (the feed to column T-2) is provided. The

temperature and pressure of the cooling water stream (CW) for the condenser for column T-3

is provided, along with an estimate of the flowrate. An estimated value is given for the flowrate

of the make-up water stream, MKUP.

Unit Operations MTBE Reaction The MTBE reaction section of the plant is shown in Figure 2 below.

85

Figure 2. MTBE Reaction Section

With reference to the previous figure, mixed C4s (stream 2) are combined with fresh

methanol (stream 1) and recycle methanol (stream 20) and pre-heated in a heat exchanger

(HX-1) to 43.5 C. The heated feed (stream 3) is then sent to a conversion reactor (RX-1)

which is maintained at 55 °C by circulating a coolant. A pressure drop of 69 KPa through the

reactor is also specified.

Since this is a conversion reactor. The three reactions defined in Table 3 take place in this

reactor at the specified conversion levels. The extent of reaction are defined based on the

values in Table 3.

86

The stoichiometries of the major and minor reactions in the MTBE process are provided in the

reaction definitions data entry window.

MTBE Distillation and Recovery The MTBE distillation and recovery section of the plant is shown in Figure 3 below.

FIGURE 3. MTBE Distillation and Recovery Section

87

The reactor product (stream 4) exchanges heat with the MTBE column bottoms

product in exchanger HX-2. Normally, this would create a thermal calculation loop. However,

since the temperature of stream 5 is known, this process is modeled by two separate heat

exchangers, HX-2A and HX-2B as shown in Figure 3. Stream 4 from the reactors is heated to

72 °C in HX-2A to produce stream 5. The product of column T-1, stream 7, is cooled in

exchanger HX2-B to produce the MTBE product stream 8. The duty of exchanger HX-2B is

defined to be equal to the duty in HX-2A. This approach avoids an unnecessary calculation

loop since the temperature of stream 5 is fixed at 72 °C.

The heated stream 5 is fed to tray 15 of the 30 tray MTBE column (T-1). The MTBE

column is simulated with the CHEMDIST algorithm using the SIMPLE initial estimate

generator (IEG). A top pressure of 621 KPa and a column pressure drop of 76.5 KPa are

given. The condenser is operated at a fixed temperature of 43.5 °C and pressure of 621 KPa.

The control specifications are a bottoms flowrate of 278 kgmoles/hr and a reflux ratio of 1.1.

The condenser and reboiler duties are varied to achieve these specifications.

88

The next step is to provide all the information required for specifying the reaction trays in the

distillation column.

Implementing Reactive Distillation You can visualize the reaction zone of a distillation column as a series of boiling pot

reactors. On each reaction tray sits a bed of solid catalyst. Each tray is connected to the next

in the forward direction (down the column) by the flow of liquid from one tray to the next, and

in the reverse direction by the vapor flow moving up from one tray to the previous tray.

For the reactive distillation process, the reaction zone (trays 8 through 13) is specified

in the column data entry window.

89

Note that the liquid volume of each of the reaction trays is also specified using the Reaction

Volumes button in the Options section of the data entry window, and that the concentration of

the dry catalyst (GCAT, in g/l) is specified using a DEFINE statement specified using the

Subroutine/Procedure Data button in the Options section of the data entry window.

90

A value of 360 g/l is given for GCAT to represent commercial catalyst loadings (corresponding

to a wet catalyst density of 770 g/l at 53% moisture content — see Table 1). The reaction

factor, RXFACT, is used to demonstrate how the reaction rate in the simulation model can be

varied to match data from an actual plant. For this case, RXFACT is set equal to 1.0,

indicating that the reaction rate has not been adjusted.

MTBE Kinetic Model The algorithm used for the reactive distillation column model is a Newtonian-based

algorithm. Therefore, in order to accurately model the MTBE reaction in the distillation column,

we need to determine not only the reaction rate of the reaction, but also the temperature and

composition derivatives of the rate. These derivatives may be generated numerically by an

estimation method, or analytically by an expression based on the reaction rate equation.

PRO/II automatically generates numerical estimates for the reaction rate derivatives.

In many cases, this is sufficient. However, certain reactions require the use of more accurate

analytical derivatives that provide better solutions. Reactions requiring these more accutate

analytical derivatives include: reversible reactions, exothermic reactions, and/or reactions

where the equilibrium is sensitive to temperature .

The MTBE reactions fit all three of the reaction types, requiring the use of analytical

derivatives. The reaction rate expression and its analytical derivatives can be easily and

readily entered by the user in the Procedure Data category of input. The MTBE reaction rate

expression used in this simulation model is based on the rate expression described in a paper

by Al-Jarallah. In this case, we will detail how to enter the reaction rate and the associated

91

analytical derivatives for the MTBE reaction.

First, in the Reaction Data category of input, the stoichiometry of the forward reaction

is given (IBTE + MEOH = MTBE). The kinetic data will be provided later on in the Procedure

Data category of input using FORTRAN-like language as the procedure named ALJD.

The reaction rate equation described by Al-Jarallah takes into account the forward and

the reverse reaction. We have modified Al-Jarallah's rate equation for this case to simulate

the effect of catalyst loading on the reaction rate. This was achieved by removing the catalyst

terms from the concentration terms. The modified reaction rate is given by:

( ) ⎥⎥⎦

⎤

⎢⎢⎣

⎡

++

−×=

+baBBAA

cC

bB

aA

Ass CKCKkCCC

kkr1

/

where:

ks = surface reaction rate constant = 1.2x10

13 exp(- 87900/RT) in (gmole/g catalyst)

(1a) Ka = equilibrium adsorption constant = 5.1x10

-13 exp(97500/RT) in g catalyst/gmole

(1b) KC = equilibrium adsorption constant = 1.6x10

-16 exp(119000/RT) in g catalyst/gmole

(1c) Keq

= equilibrium constant C

A = IBTE concentration in mole/l

CB = MEOH concentration in mole/l and C

C = MTBE concentration in mole/l

Kinetic Data The procedure data used in the reactive distillation column model is entered in the Kinetic

Procedure data entry window shown below.

92

Once the column is converged, the top and bottom product compositions are known.

Exchanger (HX-2B) is now simulated for heat exchange between the column feed (see HX-

2A) and the bottom product (stream 7). The duty in this exchanger is set equal to the duty in

exchanger HX-2A. The cooled hot side fluid is the MTBE product (stream8).

Pump P-1 pumps the liquid distillate (stream 6) at a pressure of 827 KPa to the

methanol recovery section.

A calculator (CONVERSION) is set up to compute the conversions of IBTE and MEOH

to MTBE in the reactive distillation column itself.

93

Methanol Recovery The methanol recovery section of the process is shown in Figure 4.

FIGURE 4. MTBE Distillation and Recovery Section

The methanol-C4s azeotrope (stream 6P) is delivered by pump P-1 to heat exchanger

HX-3 where it is cooled to 38 °C against cooling water (CW). The exchanger also calculates

the utility (CW) requirement given a CW delivery temperature of 16 °C and a return

temperature of 32 °C. The cooled process stream is fed to the bottom of the water wash

column (T-2).

Column T-2 is simulated as a liquid-liquid extractor with 5 theoretical trays.

Recirculating wash water is fed to the top of the column. A top pressure specification of 792

KPa is given. This column uses the VLLE SRK thermodynamic set (SRKM_VLLE) defined

previously in the Thermodynamic Data Category of the input file.

The raffinate leaves the top of the column (stream 11) and contains the unreacted and

non- reactive C4s. The extract phase (stream 12) exits at the bottom. It enters the cold side

(HX4A) of the feed-bottoms heat exchanger where it is warmed to 99 °C against the recycle

94

wash water (stream 21) which in turn is cooled (in unit HX-4B described later on).

Valve V-1 drops the pressure of the heated methanol-water stream (13) to 241 KPa

generating a mixed phase stream (14) which is adiabatically flashed in unit D-1. The vapor

phase (stream 15) containing the dissolved hydrocarbons which have been released is vented

as a flare gas; the liquid phase (stream 16) is pumped (P-2) to the methanol column to

recover methanol.

The methanol column (T-3) is simulated with 20 theoretical trays. The feed (stream 17)

enters on tray 10. The column top pressure is 138 KPa; the pressure drop through the column

is 34.5 KPa. A Subcooled, Fixed Temperature condenser type operating at 30 °C and 103.5

KPa is specified. The separation of methanol from water is readily solved using the I/O

algorithm and conventional IEG. The performance specifications are 99.5% recovery of

methanol in the overhead product and 99.95% recovery of water in the bottoms product. Tray

rating calculations are done for this column for 610 mm diameter sieve trays throughout the

column.

95

A calculator (CAL1) computes the total loss of water as a result of carry over with the

C4s (stream 11), the vent gas (stream 15) and by consumption in the reactor. This total

quantity is the amount of make-up water required. The flowrate of the make-up water stream

(MKUP) is established through a procedure call to the PRO/II stream function SRXSTR.

96

Pump P-4 pumps the recovered wash water from the methanol column bottoms

combined with make-up water as stream 21 to heat exchanger HX4B. This unit represents the

hot side of the exchanger HX-4 (see HX-4A described previously) and calculates the exit

temperature for stream 22.

Trim cooler (HX-5) further cools the wash water (stream 22) to the desired

temperature of 38 °C before it (stream 10) goes back to the water wash column. At this stage,

the first recycle loop between unit T-2 (water wash column) and HX-5 (trim cooler) is closed.

The second recycle loop between unit HX-1 (feed heater) and P-3 (recycle pump) is closed

when the pump P-3 recycles the overhead (stream 19) from the top of the methanol column

(T-3) to the reactor section.

Then, as an illustrative example, a rigorous heat exchanger (RC-1) is used to

rigorously rate the methanol column condenser. This rigorous heat exchanger is modeled as

an attached heat exchanger to column T-3. This unit takes as its input the exchanger's

mechanical data such as shell and tube dimensions, tube layout pattern, the baffle cut and

shell and tube side nozzle sizes. A fouling factor of 0.00035 m2-K/kW is used for the

condenser cooling water side. The ZONES option is selected to determine where phase

changes occur in the exchanger. An extended data sheet is printed in the output.

97

98

APPENDIX (C)

PROCESS INPUT FILE USED IN PRO/II

99

Appendix (C): Process Input-File used in PRO/II $ Generated by PRO/II Keyword Generation System <version 8.0> $ Generated on: Thu Jan 24 01:49:16 2008 TITLE PROJECT=MasterThises, PROBLEM=MTBE PLANT, USER=Fahad Harthi, & DATE=Dec 2007 PRINT INPUT=ALL, STREAM=ALL, RATE=M, MBALANCE DIMENSION SI, TEMP=C, PRES=BAR SEQUENCE DEFINED=CONVERSION,HX-1,RX-1,HX-2A,CAL0,T-1,HX-2B,P-1, & HX-3,T-2,HX4A,V-1,D-1,P-2,T-3,CAL1,P-4,HX4B,HX-5,P-3,RC-1 COMPONENT DATA LIBID 1,NC4/2,IC4/3,1BUTENE/4,BTC2/5,BTT2/6,IBTE/7,MTBE/8,MEOH/ & 9,TBA/10,H2O/11,244TM1P,,DIB THERMODYNAMIC DATA METHOD SYSTEM=SRKM, SET=S1, DEFAULT KVAL(VLE) SRKM 1,9,0.046973,0.126027,0,0,0,0,1,1 SRKM 3,8,0.136,-0.0323,0,0,0,0,1,1 SRKM 4,8,0.136,-0.0323,0,0,0,0,1,1 SRKM 5,8,0.136,-0.0323,0,0,0,0,1,1 SRKM 6,8,0.135525,-0.032271,0,0,0,0,1,1 SRKM 8,9,-0.073971,-0.055222,0,0,0,0,1,1 SRKM 9,10,-0.145,-0.253,0,0,0,0,1,1 SRKM 7,11,0.05785,-0.0093,-10.144,6.17,0,0,1,1 METHOD SYSTEM(VLLE)=SRKM, L1KEY=1, L2KEY=10, SET=S2 KVAL(VLE) SRKM 1,9,0.046973,0.126027,0,0,0,0,1,1 SRKM 3,8,0.136,-0.0323,0,0,0,0,1,1 SRKM 4,8,0.136,-0.0323,0,0,0,0,1,1 SRKM 5,8,0.136,-0.0323,0,0,0,0,1,1 SRKM 6,8,0.135525,-0.032271,0,0,0,0,1,1 SRKM 8,9,-0.073971,-0.055222,0,0,0,0,1,1 SRKM 9,10,-0.145,-0.253,0,0,0,0,1,1 SRKM 7,11,0.05785,-0.0093,-10.144,6.17,0,0,1,1 STREAM DATA PROPERTY STREAM=1, TEMPERATURE=16, PRESSURE=16.2, PHASE=M, & COMPOSITION(M,KGM/H)=8,277.5 PROPERTY STREAM=2, TEMPERATURE=16, PRESSURE=16.2, PHASE=M, & RATE(M)=850, COMPOSITION(M)=1,9/2,41/3,7/4,4/5,6/6,33, & NORMALIZE PROPERTY STREAM=10, TEMPERATURE=38, PRESSURE=7.93, PHASE=M, & COMPOSITION(M,KGM/H)=10,375 PROPERTY STREAM=20, TEMPERATURE=44, PRESSURE=17.24, PHASE=M, & COMPOSITION(M,KGM/H)=8,4/10,0.3 PROPERTY STREAM=CW, TEMPERATURE=21, PRESSURE=6.9, PHASE=M, & RATE(LV)=175, COMPOSITION(M)=10,100 PROPERTY STREAM=MKUP, TEMPERATURE=38, PRESSURE=3.5, PHASE=M, & COMPOSITION(M,KGM/H)=10,500 NAME 1,MEOH FEED/2,OLEFINS/20,MEOH RECYC/MKUP,MKUP WATER/ & 6,T-1 OVHD/8,MTBE/11,C4'S/15,FLARE GAS RXDATA RXSET ID=ST1

100

REACTION ID=1 STOICHIOMETRY 6,-2/11,1 REACTION ID=2 STOICHIOMETRY 6,-1/9,1/10,-1 REACTION ID=3 STOICHIOMETRY 6,-1/7,1/8,-1 RXSET ID=ALJX REACTION ID=ALJ0 STOICHIOMETRY 6,-1/7,1/8,-1 PROCEDURE DATA PROCEDURE(KINETIC) ID=ALJD, NAME=MTBE Process PDATA GCAT,RXFACT CODE REAL KS , KA , KC , KALJ , KREH1 , KREH2 , KIZQ , KEQREF INTEGER IBTE , MEOH , MTBE $ $ INITIALIZE DATA: $ SET INDEXES FOR COMPONENTS $ DEFINE GAS CONSTANT IN JOULES/GM-MOLE K $ NOTE: R COULD HAVE BEEN RETRIEVED IN INPUT UNITS BY R=RGAS. $ HOWEVER, SINCE THE REACTION BASIS WON'T CHANGE, AND $ RGAS WILL CHANGE WITH THE DEFAULT UNITS, THIS $ ELIMINATES ONE POSSIBLE SOURCE OF ERROR.. $ INITIALIZE THE LOCAL VARIABLE TK TO THE ABSOLUTE TEMPERATURE. $ NOTE: THE TEMPERATURE BASIS FOR THE FLOWSHET MUST BE CENTIGRADE $ OR KELVIN. $ SET TEMPERATURE AND COMPOSITION DERIVATIVES TO ZERO. $ IBTE = 6 MTBE = 7 MEOH = 8 $ R = 8.314 $ TK = RTABS $ DO 1000 I1 = 1,NOR DRDT(I1) = 0.0 DO 1000 I2 = 1,NOC 1000 DRDX(I2,I1) = 0.0 $ $ CALCULATE THE SURFACE REACTION RATE CONSTANT, KS, AND THE $ EQUILIBRIUM ADSORPTION CONSTANTS KA AND KB. THE ACTIVATION $ ENERGY IS IN J/GM-MOLE. $ UNITS: KS - (GM-MOLE / GM CATALYST)**1.5 /HOUR $ KA - GM-CATALYST / MOLE $ KC - GM-CATALYST / MOLE $ KS = 1.2E+13*EXP(-87900.0/(R*TK)) KA = 5.1E-13*EXP( 97500.0/(R*TK)) KC = 1.6E-16*EXP(119000.0/(R*TK)) $

101

DKSDT = KS * 87900.0 / R / (TK*TK) DKADT = KA * (-1.0) * 97500.0 / R / (TK*TK) DKCDT = KC * (-1.0) * 119000.0 / R / (TK*TK) $ $ ---- CALCULATE THE EQUILIBRIUM CONSTANT. $ $ UNITS - (GM-MOLES/GM-CATALYST)/HOUR $ PHASE - LIQUID PHASE REACTION $ CATALYST - ION EXCHANGE RESIN AMBERLYST 15, $ THE EQUILIBRIUM SHOULD BE INDEPENDANT OF THE CATALYST $ $ KALJ = EXP(-17.31715+(7196.776/TK)) $ DKALJDT = - KALJ * 7196.776 / (TK*TK) $ $ BULK CONCENRATIONS OF COMPONENTS PER GRAM OF CATALYST, XLCONC IS $ IN MOLES/FLOW VOLUME. XLCONC WILL BE PASSED TO THE PROCEDURE $ IN USER INPUT UNITS. INTERNALLY TO PRO/II, IT IS IN SI UNITS $ (KG-MOLE/CUBIC METER). THE BASIS FOR THESE REACTION EQUATIONS $ IS GM-MOLES/LITER. THE CONVERSION FACTOR FROM INPUT UNITS OF $ KG-MOLES/CUBIC METER TO THE REACTION BASIS OF GM-MOLES/LITER $ IS ONE. THEREFORE, XLCONC CAN BE USED WITH NO CONVERSION. $ $ $ RHOA=(XLCONC(MEOH)/GCAT) |-THIS SHOULD BE EQUIVALENT TO BELOW. $ RHOB=(XLCONC(IBTE)/GCAT) | IT HAS BEEN WRITTEN EXPLICITLY BELOW $ RHOC=(XLCONC(MTBE)/GCAT) | TO MAKE IT OBVIOUS HOW TO DO THE $ ANALYTICAL DERIVATIVES. $ $ CALCULATE DENSITY IN MOLES / VOLUME $ GCATX = 12.4 DENS=RLMRAT/RLVRAT RHOA=(XLIQ(MEOH)*DENS/GCATX) RHOB=(XLIQ(IBTE)*DENS/GCATX) RHOC=(XLIQ(MTBE)*DENS/GCATX) $ $ ---- CALCULATE REACTION RATE AND DERIVATIVES BY TERMS $ ---- UNITS - RATE - GRAM-MOLE / GRAM CATALYST / HR. $ $ DENOMINATOR & DERIVATIVES. $ RDEN = 1.0 + ( KA*RHOA ) + 0.0 + ( KC*RHOC ) DRDDT = RHOA*DKADT + 0.0 + RHOC*DKCDT DRDDME = KA/GCATX*DENS DRDDIB = 0.0 DRDDMT = KC/GCATX*DENS $ $ FIRST FACTOR IN RATE EQUATION.

102

FACT1 = KS *KA/RDEN DFAC1DT = DKSDT*KA/RDEN + KS*DKADT/RDEN - KS*KA/RDEN**2 * DRDDT $ $ SECOND FACTOR IN RATE EQUATION. FACT2 = RHOA*RHOB**0.5 - RHOC**1.5/KALJ DFAC2DT = 0.0 + RHOC**1.5/KALJ**2 * DKALJDT $ $ COMBINE TERMS TO CALCULATE RATE AND DERIVATIVES. $ -- RATE EQUATION (RATE PER ONE GRAM OF CATALYSIS). RATE = FACT1 * FACT2 $ $ -- RATE TEMPERATURE DERIVATIVE. DRDT(1) = DFAC1DT * FACT2 & + FACT1 * DFAC2DT $ -- RATE COMPOSITION DERIVATIVES. DRDX(MEOH,1) = -KS*KA/RDEN**2 * DRDDME * FACT2 & + FACT1 * (RHOB**0.5/GCATX*DENS) DRDX(IBTE,1) = -KS*KA/RDEN**2 * DRDDIB * FACT2 & + FACT1 * (RHOA/2.0/RHOB**0.5/GCATX*DENS) DRDX(MTBE,1) = -KS*KA/RDEN**2 * DRDDMT * FACT2 & - FACT1 * (1.5* RHOC**0.5/GCATX/KALJ*DENS) $ $ ---- CONVERT RATE EQUATION AND DERIVATIVES TO A STRAIGHT VOLUME BASIS $ ---- BY MULTIPLING THE BASE RATE BY THE GRAMS OF CATALYST/UNIT VOLUME. $ ---- THE RATE IS RETURNED IN INPUT UNITS, KG-MOLES/CUBIC METER/HOUR. $ RRATES(1) = GCAT * RXFACT * RATE $ DRDT(1) = GCAT * RXFACT * DRDT(1) $ DRDX(MEOH,1) = GCAT * RXFACT * DRDX(MEOH,1) DRDX(IBTE,1) = GCAT * RXFACT * DRDX(IBTE,1) DRDX(MTBE,1) = GCAT * RXFACT * DRDX(MTBE,1) RETURN UNIT OPERATIONS CALCULATOR UID=CONVERSION, NAME=CONVERSION OF IBTE-MEOH TO MTBE RESULT 1,IN - MEOH/2,IN - IBTE/3,IN - MTBE/4,OUT - MEOH/ & 5,OUT - IBTE/6,OUT - MTBE/20,IBTE CONV/21,MEOH CONV DEFINE P(1) AS STREAM=4, RATE(KGM/H), COMP=8,WET DEFINE P(2) AS STREAM=4, RATE(KGM/H), COMP=6,WET DEFINE P(3) AS STREAM=4, RATE(KGM/H), COMP=7,WET DEFINE P(4) AS STREAM=6, RATE(KGM/H), COMP=8,WET DEFINE P(5) AS STREAM=6, RATE(KGM/H), COMP=6,WET DEFINE P(6) AS STREAM=6, RATE(KGM/H), COMP=7,WET DEFINE P(7) AS STREAM=7, RATE(KGM/H), COMP=8,WET DEFINE P(8) AS STREAM=7, RATE(KGM/H), COMP=6,WET DEFINE P(9) AS STREAM=7, RATE(KGM/H), COMP=7,WET PROCEDURE $ --- LOAD RATES R( 1) = P( 1)

103

R( 2) = P( 2) R( 3) = P( 3) R( 4) = P( 4) + P( 7) R( 5) = P( 5) + P( 8) R( 6) = P( 6) + P( 9) $ --- CALCULATE CONVERSION R(20) = ( R(2) - R(5) ) / R(2) R(21) = ( R(1) - R(4) ) / R(1) $ --- DISPLAY RESULTS DISPLAY R( 1: 9 ) DISPLAY R( 20:21 ) RETURN HX UID=HX-1, NAME=FEED HEAT COLD FEED=1,2,20, L=3, DP=0.345 OPER CTEMP=43.5 CONREACTOR UID=RX-1, NAME=REACTORS FEED 3 PRODUCT L=4 OPERATION ISOTHERMAL, TEMPERATURE=55, DP=0.69 RXCALCULATION MODEL=STOIC RXSTOIC RXSET=ST1 REACTION 1 BASE COMPONENT=6 CONVERSION 0.0025 REACTION 2 BASE COMPONENT=10 CONVERSION 1 REACTION 3 BASE COMPONENT=8 CONVERSION 0.93 HX UID=HX-2A, NAME=FEED-BTMS-A COLD FEED=4, L=5, DP=0.345 OPER CTEMP=72 CALCULATOR UID=CAL0, NAME=COPY RXFACT PROCEDURE IF (R(2) .NE. 1.0) R(1) = 1.0 $Set RXFACT TO 1 on first call. R(2) = 1.0 RETURN COLUMN UID=T-1, NAME=MTBE COLUMN PARAMETER TRAY=30,CHEMDIST=35 FEED 5,15, SEPARATE PRODUCT OVHD(M)=6, BTMS(M)=7,280, SUPERSEDE=ON CONDENSER TYPE=TFIX, PRESSURE=6.21, TEMPERATURE=43.5 DUTY 1,1,,CONDENSER DUTY 2,30,,SIDEHC2 PSPEC PTOP=6.21, DPCOLUMN=0.765 PRINT COMPOSITION=M, PROPTABLE=ALL ESTIMATE MODEL=SIMPLE, RRATIO(L)=1.1 SPEC ID=COL1SPEC1, STREAM=7, RATE(KGM/H),TOTAL,WET, VALUE=278 SPEC ID=COL1SPEC2, RRATIO, PHASE=L, VALUE=1.1 DEFINE GCAT AS 360 DEFINE RXFACT AS 1

104

VARY DNAME=CONDENSER,SIDEHC2 PLOT LOG, XCOMPONENT=6,6/8,8/7,7/1,1 TSIZE SECTION(1)=2,29,VALVE, DMIN=381, FF=80, DPCALC=0 VLLECHECK CHECK=OFF LVOL 8,5/9,5/10,5/11,5/12,5/13,5 RXTRAY REFERENCE=ALJX, LOCAL=L_ALJX, KPROCEDURE=ALJD, TRAY=8, & 13 RXSET LOCAL=L_ALJX REACTION ID=ALJ0, COPTION=KINETICS KINETICS PEXP(C,KG,M3,KPA,HR) HX UID=HX-2B, NAME=FEED-BTMS-B HOT FEED=7, L=8, DP=0.345 DEFINE DUTY(KJ/HR) AS HX=HX-2A, DUTY(KJ/HR) PUMP UID=P-1, NAME=T-1 OVHD FEED 6 PRODUCT L=6P OPERATION EFF=65, PRESSURE=8.27 HX UID=HX-3, NAME=COOLER HOT FEED=6P, L=9, DP=0.345 UTILITY WATER, TIN=16, TEMPERATURE=32 CONFIGURE COUNTER OPER HTEMP=38 COLUMN UID=T-2, NAME=WATER WASH PARAMETER TRAY=5,LLEX=25 FEED 9,5/10,1 PRODUCT OVHD(L1,M)=11, BTMS(M)=12,185, SUPERSEDE=ON PSPEC PTOP=7.92 PRINT PROPTABLE=PART ESTIMATE MODEL=SIMPLE METHOD SET=S2 HX UID=HX4A, NAME=FEED-BTMS COLD FEED=12, L=13, DP=0.345 OPER CTEMP=99 VALVE UID=V-1, NAME=VALVE FEED 13 PRODUCT M=14 OPERATION PRESSURE=2.41 FLASH UID=D-1, NAME=SEPARATOR FEED 14 PRODUCT V=15, L=16 ADIABATIC PUMP UID=P-2, NAME=FEED PUMP FEED 16 PRODUCT L=17 OPERATION EFF=65, PRESSURE=6.9 COLUMN UID=T-3, NAME=MEOH COLUMN PARAMETER TRAY=20,IO=10 FEED 17,10 PRODUCT OVHD(M)=19, BTMS(M)=18,182, SUPERSEDE=ON CONDENSER TYPE=TFIX, PRESSURE=1.035, TEMPERATURE=30 DUTY 1,1,,CONDENSER DUTY 2,20,,SIDEHC2

105

PSPEC PTOP=1.38, DPCOLUMN=0.345 PRINT PROPTABLE=PART ESTIMATE MODEL=CONVENTIONAL, RRATIO=10 SPEC ID=COL3SPEC1, STREAM=19, RATE(KGM/H), COMP=8,WET, DIVIDE, & STREAM=17, RATE(KGM/H), COMP=8,WET, VALUE=0.995 SPEC ID=COL3SPEC2, STREAM=18, RATE(KGM/H), COMP=10,WET, DIVIDE, & STREAM=17, RATE(KGM/H), COMP=10,WET, VALUE=0.9995 VARY DNAME=CONDENSER,SIDEHC2 TRATE SECTION(1)=2,19,SIEVE, PASSES=1, DIAMETER(TRAY)=610, & DIAMETER(SIEVEHOLE,IN)=0.5, WEIR=50.8, DCC=38.1 CALCULATOR UID=CAL1, NAME=MAKEUP SEQUENCE STREAM=MKUP DEFINE P(1) AS STREAM=11, RATE(KGM/H), COMP=10,WET DEFINE P(2) AS STREAM=15, RATE(KGM/H), COMP=10,WET DEFINE P(3) AS STREAM=19, RATE(KGM/H), COMP=10,WET PROCEDURE R(1) = P(1) + P(2) + P(3) CALL SRXSTR(SMR,R(1),MKUP) RETURN PUMP UID=P-4, NAME=WATER PUMP FEED 18,MKUP PRODUCT L=21 OPERATION EFF=65, PRESSURE=8.62 HX UID=HX4B, NAME=FEED-BOTS HOT FEED=21, L=22, DP=0.345 DEFINE DUTY(KJ/HR) AS HX=HX4A, DUTY(KJ/HR) HX UID=HX-5, NAME=COOLER HOT FEED=22, L=10, DP=0.345 OPER HTEMP=38 PUMP UID=P-3, NAME=RECYCLE PUMP FEED 19 PRODUCT L=20 OPERATION EFF=65, PRESSURE=17.24 HXRIG UID=RC-1, NAME=T-3 COND TYPE TEMA=AES TUBES FEED=CW, L=WOUT, LENGTH=5.75, OD=19, BWG=14, PASS=2, & PATTERN=90, PITCH=25.4, FOUL=0.00035, METHOD=S1 SHELL DPUNIT=0.49244, METHOD=S1, ID=381 BAFFLE CUT=0.18 SNOZZLE TYPE=CONV, ID=152,102 TNOZZLE ID=102,102 PRINT EXTENDED, ZONE ATTACH COLUMN=T-3, TYPE=CONDENSER RECYCLE DATA ACCELERATION TYPE=WEGSTEIN LOOP NUMBER=1, START=T-2, END=HX-5,WEGSTEIN LOOP NUMBER=2, START=HX-1, END=P-3,WEGSTEIN CASESTUDY OLDCASE=BASECASE, NEWCASE=CS1 PARAMETER ID=MTBE, COLUMN=T-1, SPEC(1),STEP(ACT)=0, CYCLES =1,1 RESULT ID=RESULT1, STREAM=15,FRACTION, COMP=1,WET END