modeling and simulation of fcc risers · figure 1.1 fluid catalytic cracking unit 2 figure 3.1...

Republic of Iraq Ministry of Higher Education and Scientific Research University of Technology Chemical Engineering Department

MODELING AND SIMULATION OF FCC RISERS

A RESEARCH

SUBMITTED TO THE CHEMICAL ENGINEERING DEPARTMENT OF THE UNIVERSITY OF TECHNOLOGY IN THE PARTIAL FULFILLMENT OF THE REQUIREMENTS

FOR THE DEGREE OF

HIGHER DIPLOMA IN CHEMICAL ENGINEERING

(PETROLEUM REFINING AND GAS TECHNOLOGY)

By WALEED KHALID FADHIL

B.Sc. in Chem. Eng. 1998

March 2012

بسم اهللا الرحمن الرحيم

منه نتم ا أ ذنارا فإ ر خض ر األ الشج كم من ل ل ع ي ج ذال ۞ 80 یس ۞ وند وق ت

صدق اهللا العظيم

Dedicated to

My Family With My Deep Love

بالعامل المساعد ة الریاضیة لوحدة التكسیر المحفزالنمذج

FCC riserبالعامل المساعد یاضي لمفاعل وحدة التكسیر المحفزیشتمل البحث على عمل نموذج ر

التفاعل یكون بین اربع مكونات ان و Plug flow reactorاعتبار المفاعل مثالي على یعتمد البحث

والفحم Gasolineوالبنزین Light gasesوالغازات الخفیفة VGO feedیة ذافتراضیة هي التغ

Coke او ما یسمى بـFour lumps model .تم لوصف حركیة تفاعالت التكسیر خالل المفاعل

العتماد على عالقة خطیة بین المحتوى احساب دالة فقدان الفعالیة للعامل المساعد خالل المفاعل ب

تفعیل النموذج الریاضي كان ما یعادله من فعالیة متبقیة للعامل المساعد.الكاربوني للعامل المساعد و

الحد المصافي العراقیة الجدیدة, UOPباستخدام الوحدة التجاریة للتكسیر بالعامل المساعد من تصمیم

الریاضي قادر على ایجاد درجات حرارة الخلط عند كل من وعاء الخلط ذج.النمو واعطى نتائج مقاربة

MxR chamber الضافة الى امكانیة عرض االداء الفیزیائي ومقدار ادخول المفاعل, بومنطقة بدایة

في برنامج اكسل لحل معادالت worksheetانشاء تم االنتاجیة للمكونات على طول المفاعل.

الریاضي والستخدامه كاداة جیدة لدراسة اداء الوحدة جراء اي تغییر یحصل في الضروف ذجالنمو

شغیلیة للوحدة.الت

III

ACKNOWLEDGMENTS

First of all, thanks to Allah, who enabled me to achieve this research.

I wish to express my sincere gratitude and thankfulness to my

supervisor Dr. Shakir M. Ahmed for his kind supervision and continuous

advices during the research.

My grateful thanks to Prof. Dr. Mumtaz A. Zablouk, the Chairman of

the Department of Chemical Engineering at the University of Technology

for the provision of research facilities.

Special thanks to Assist. Prof. Dr. Mohammed F. Abid for his help

and support.

I would like to express my sincere appreciation to Assis. Lec. Farooq

A. Mehdi for his help. Also, my respectful regards to all the staff of

Chemical Engineering Department of University of Technology.

And finally my special thanks to my family for their support and

encouragement.

IV

ABSTRACT

In the present work a mathematical model for the riser reactor of Fluid

Catalytic Cracking has been developed. The riser is considered as a plug

flow reactor incorporating the four lumps model for kinetics of cracking

reactions. Catalyst deactivation function is calculated based on linear

relationship between the catalyst coke content and its retention activity.

The model has been validated using the plant data of a commercial FCC

unit with RxCat technology developed by UOP. The model can predict the

mixing temperatures, at MxR chamber and riser inlet; also shows the

physical performance and productivity all over the riser height. An

interactive excel worksheet is constructed and used as a powerful tool for

solving the model equations and studying the effect of any change in

operating variableson the unit performance.

V

CONTENTS

Certification I

Acknowledgements III

Abstract IV

Contents V

List of Figures VII

List of Tables IX

Nomenclature X

CHAPTER – 1Introduction 1

1.1. Introduction 1

1.2. Aim and scope of Work 4

CHAPTER – 2Literature Survey 5

CHAPTER – 3Mathematical Modeling 10

3.1. Introduction 10

3.2. Reactor / Regenerator Material & Energy Balances 10

3.2.1. Material Balance 11

3.2.1.1. Reactor Material Balance 11

3.2.1.2. Regenerator Materials Balance 12

3.2.2. Energy Balance 13

3.2.2.1. Reactor Energy Balance 13

3.2.2.2. Regenerator Energy balance 13

3.3. Riser model 14

3.3.1. Model Assumptions 14

3.3.2. Cracking Reaction Kinetics 15

3.3.3. Concentration, Temperature, Pressure and Coking

time Profiles in the Riser 17

3.3.4. Catalyst Deactivation 18

3.3.5. Riser Hydrodynamics 19

VI

3.3.6. Mixing Temperatures 21

3.4. Heat of Combustion at the Regenerator 23

3.5. Model Solution 23

CHAPTER – 4 Results and Discussion 26

4.1. Introduction 26

4.2. Case Study 26

4.3. Model Results 28

CHAPTER – 5 Conclusions and Recommendations 38

5.1. Conclusion 38

5.2. Recommendations for FutureWork 39

Appendix A -Fluidized Catalytic Cracking Technologies A-1

Appendix B - Variables of FCCunits B-1

Appendix C - Computer Programs C-1

Appendix D -Glossary of Terms Used In This Work E-1

References R-1

VII

LIST OF FIGURES

Figure No. Title Page No.

Figure 1.1 Fluid Catalytic Cracking Unit 2

Figure 3.1 schematic of FCCU reactor/regenerator system used in present model

11

Figure 3.2 Input and output streams for reactor and regenerator in FCCU

12

Figure 3.3 A volume element in the riser reactor 14

Figure 3.4 Schematic of four lumped reactions 16

Figure 3.5 Mathematical representation of reactor riser used in the model

21

Figure 3.6 computational flow diagrams for riser reactor model

25

Figure 4.1 Four lump concentration profile vs. riser height 28

Figure 4.2 Riser temperature profile 28

Figure 4.3 Riser pressure profile 30

Figure 4.4 Gas phase molecular weight vs. riser length 30

Figure 4.5 Gas phase density vs. riser height 30

Figure 4.6 Gas phase mass rate vs. riser height 31

Figure 4.7 Slip factor vs. riser height 32

Figure 4.8 Gas phase and catalyst velocities vs. riser height

32

Figure 4.9 Gas phase and catalyst residence times vs. riser height

33

Figure 4.10 Gas phase void fraction vs. riser height 34

VIII

Figure No. Title Page No.

Figure 4.11 Predicted catalyst activity along the riser height

34

Figure 4.12 Gasoline yield vs. feed conversion 35

Figure 4.13 Feed conversion vs. riser height 35

Figure A-1(a) Stacked Reactor regenerator configuration A-2

Figure A-1(b) Side by side Reactor regenerator configuration A-2

Figure A-2 KBR’s counter-current regeneration design A-5

Figure A-3 Lummus FCCU Process Flow Diagram A-7

Figure A-4 S&W / IFP FCCU design A-10

Figure A-5 Mix zone temperature control & Feed injection nozzle

A-11

Figure A-6 Shell’s FCC & MILOS-FCC designs A-13

Figure A-7 PentaFlow Packing & Feed nozzels configurations

A-14

Figure A-8 UOP’s reactor/regenerator FCCU design A-16

Figure B-1 FCC reaction network B-4

Figure B-2 Principal Reactions in Fluid Catalytic Cracking B-5

Figure B-3 Evolution in structure of FCC catalysts before 1990

B-6

Figure B-4 Catalyst activity retention vs. Carbon on regenerated catalyst

B-8

Figure C-1 Constructed Excel worksheet For FCC unit C-8

IX

LIST OF TABLES

Table No. Title Page No.

Table 1.1 Gasoline Pool Example 2

Table 4.1 Kinetic parameters with Modified frequency factors used in present model

27

Table 4.2 Mixing temperatures at MxR chamber and Riser inlet temperature

27

Table 4.3 Model predicted and plant values comparison 36

Table 4.4 Case study results 36

Table A-1 KBR & ExxonMobil FCCU Technologies A-6

Table A-2 LUMMUS FCCU Technologies A-8

Table A-3 S&W / IFP FCCU Technologies A-11

Table A-4 Shell’s FCCU Technologies A-14

Table B-1 Feedstock Crackability B-2

Table B-2 Typical FCC unit products B-3

Table B-3 Effect of operating Temperature of the reactor on the performance of a fluidized bed cracking

B-11

Table C-1 Variables used in Polymath program C-1

X

NOMENCLATURES

A Riser cross section area (m2)

Ar Archimedes number (-)

Cp Heat capacity (kJ/kg.K)

D Riser diameter (m)

d Particle diameter (m)

Ej Activation energy (kJ/kmole)

Fr Fround number

H Heat enthalpy (kJ/s)

Hj Heat enthalpy of jth reaction (kJ/kg)

Kj Kinetic reaction rate constant of jth reaction

Koj frequency factor or preexponential factor for jth reaction

Kuop UOP characterization factor

L Riser Height (m)

MW Molecular weight (kg/kg mole)

m Mass rate (kg/s)

P Riser pressure (Pascal)

R Universal ideal gas constant (atm ∙ m3/kmole ∙ K)

Re Reynold number

S Sulfur (kg/s)

Sph Sphericity

SG Feed specific gravity

T Temperature (K)

t Gas phase residence time (second)

tc Catalyst residence time (second)

u Velocity (m/s)

WHSV Weighted hourly space velocity (1/hr)

X Conversion (wt %)

yi Weight fraction of ith lump

XI

z Axial position of riser height (m)

Greek letters

ε Voidage

ϕ Catalyst deactivation function

ρ Density (kg/m3)

ψ Slip factor Catalytic cracker efficiency

∆ Difference

μ Viscosity (Pa.s)

σj Ratio of frequency factor of jth lump reaction per frequency

factor of VGO to gasoline reaction

Subscripts

air Air for regeneration

cok Coke

cat Catalyst

ds Dispersion or Atomizing steam

f Feed

fg Flue gas

fl Feed in the liquid phase

fv Feed in the vapor phase

g Gas phase

in Flowing in

j 1,2,3,4 and 5 for the reactions VGO to GLN, VGO to LGS,

VGO to COK, GLN to LGS, and GLN to COK respectively

ls Lifting steam

mix1 Mixing temperature at MxR chamber

mix2 Mixing temperature at riser inlet

XII

o Superficial

out Flowing out

p Particle

pr Products

rcat Regenerated catalyst

rcoke Coke on the regenerated catalyst

s Steam

si Riser and reactor inlet steam

so Riser and reactor outlet steam

scat Spent catalyst

scok Coke on the spent catalyst

t Terminal velocity

xcat Carbonized catalyst

xcok Coke of carbonized catalyst

Abbreviations

AF Advanced Fluidization

BPSD Barrel Per Stream Day

CB & I Chicago Bridge & Iron

CCR Conradson Carbon Residue or Catalyst Circulation Rate

CFD Computational Fluid Dynamics

COK Coke

CRC Coke on the Regenerated Catalyst

CSC Coke on the Spent Catalyst

C/O Catalyst to Oil ratio

E-cat Equilibrium Catalyst

EMRE ExxonMobil Research and Engineering

FCC Fluidized Catalytic Cracking

FCCU Fluidized Catalytic Cracking

FEED Front End Engineering Design

XIII

FF Fresh Feed

GLN Gasoline

HCO Heavy Cycle Oil

IFP Institute France Petrol

KBR Kellogg Brown & Root

LCO Light Cycle Oil

LGS Light Gases

LPG Liquefied Petroleum Gas

MTC Mix Temperature Control

RON Research Octane Number

RSS Riser Separator Stripper

ODE Ordinary Differential Equation

TSS Third Stage Separator

UOP Universal Oil Products

VDS Vortex Disengaging System

VGO Vacuum Gasoil

VSS Vortex Separation System

CHAPTER ONE

INTRODUCTION

Chapter One Introduction

1

1.1. INTRODUCTION Fluid catalytic cracking (FCC) technology is a technology with more than

60 years of commercial operating experience. The process is used to

convert higher-molecular-weight hydrocarbons to lighter, more valuable

products through contact with a powdered catalyst at appropriate

conditions. The primary purpose of the FCC process has been to produce

gasoline, distillate, and C3/C4 olefins from low-value excess refinery gas

oils and heavier refinery streams. FCC is often the heart of a modern

refinery because of its adaptability to changing feedstocks and product

demands and because of high margins that exist between the FCC

feedstocks and converted FCC products. As oil refining has evolved over

the last 60 years, the FCC process has evolved with it, meeting the

challenges of cracking heavier, more contaminated feedstocks, increasing

operating flexibility, accommodating environmental legislation, and

maximizing reliability [1]. In the environmental protection field, FCC unit

play a significant role by producing the gasoline with lower benzene

content as clarified in the gasoline pool example (Table 1.1)

Refineries use fluid catalytic cracking to correct the imbalance between the

market demand for gasoline and the excess of heavy high boiling range

products resulting from the distillation of crude oil. [2]

The fluid catalytic cracking (FCC) unit consists of a reaction section and a

fractionating section that operate together as an integrated processing unit.

The reaction section includes two reactors, the riser reactor, where almost

all the endothermic cracking reactions and coke deposition on the catalyst

occur, and the regenerator reactor, where air is used to burn off the

accumulated coke. The regeneration process provides, in addition to

reactivating the catalyst powders, the heat required by the endothermic

cracking reactions, (Figure 1.1). [3]


2

Figure 1.1: Fluid Catalytic Cracking Unit [4]

Table 1.1

Gasoline Pool Example [5, 6]

Gasoline source

% vol. of Pool

% vol. Bz. RON

% vol Pool RON

FCC 35 0.8 88 33.8 Reformate 30 4.5 94 31 Alkylate 20 0 94 20.6 Isomerate 15 0.6 89 14.6

Isomerization Alkylation

FCCU Reformer


3

A modern FCC unit comprises different sections such as a riser reactor, a

stripper, a disengager, a regenerator, a main fractionator, catalyst transport

lines (spent catalyst standpipe and regenerated catalyst standpipe) and

several other auxiliary units such as: cyclones, air blower, expander, wet

gas compressor, feed pre-heater, air heater, catalyst cooler, etc [7]. The

proprietary new designs and technologies that have been developed by the

major FCC designers and licensors are briefly described in the

Appendix A.

Because of the importance of FCC unit in refining, a construction of

mathematical model that can describe the dynamic behavior of FCC unit

equipments in steady state is very important. Accurate model can be used

as a powerful tool to study the effect of process variables on the

performance and productivity of the system [7].

Simulation studies also provide guidance in the development of new

processes and can reduce both time and investment [8]. The effective

simulation of the fluid catalytic cracking operation requires knowledge of

reaction kinetics, fluid dynamics, feed and catalyst effects [9].

The riser reactor is probably the most important equipment in a FCC unit.

The modeling of a riser reactor is very complex due to complex

hydrodynamics and unknown multiple reactions, coupled with mass

transfer resistance, heat transfer resistance and deactivation kinetics. A

complete model of the riser reactor should include all the important

physical phenomena and detailed reaction kinetics [10].


4

1.2. AIMS AND SCOPE OF WORK

The main objectives of the present work are:

1. A short literature review of previous FCC riser modeling and

simulation studies.

2. Formulation of a mathematical model that can describe the reaction

kinetics and physical performance in the riser section of FCC unit by

using four lump model for kinetics description with linearly scaled

up frequency factors of Arrhenius equations.

3. The quantities of lifting and dispersion steam in all calculation steps

will be considered.

4. Model validation against a commercial scale FCC unit designed by

UOP is to be investigated.

5. An interactive excel worksheet for solving model equations and

studying the unit performance is to be constructed.

CHAPTER TWO

LITERATURE SURVEY

Chapter Two Literature Survey

5

Modeling of riser reactor is very complex due to complex

hydrodynamics, unknown multiple reactions coupled with mass transfer and

heat transfer resistances. Also, the conditions keep changing all along the riser

height due to cracking which causes molar expansion in the gas phase and

influences the axial and radial catalyst density in the riser. In the literature,

numerous models of FCC riser are available with varying degrees of

simplifications and assumptions.

Ali et al. (1997) [3]; Arbel et al. (1995) [11]; Han et al. (2001) [12],

developed a mathematical model of an industrial FCC unit, includes one

dimensional mass, energy, and species balance; their models was based on the

assumption of instantaneous and complete vaporization of the feed when

contacted with hot regenerated catalyst assuming modern high efficiency feed

injection systems. These types of modeling are normally simple to formulate

and to solve. They are more suitable when the interest is to explore the

influence of operating conditions, test a kinetic model or when the simulation

includes not only the riser, but also other equipments like the regenerator and

the stripper. The simplest kind of these models is the homogeneous version,

where both the vapor phase (hydrocarbon feed & products vapors) and the

solid phase (catalyst & coke) are moving at the same velocity. The

heterogeneous version considers different velocities for the two phases,

resulting in different residence times for each phase inside the riser.

The simplest hydrodynamic models assume steady state ideal plug flow

reactor.


6

Ali et al. [3] and Han et al. [12] used the four-lump kinetic models to

describe the behavior of cracking reactions, while Arbel et al. [11] used more

complex ten-lump model.

Theologos and Markatos (1993) [13] proposed a three dimensional

mathematical model considering two phase flow, heat transfer, and three lump

reaction scheme in the riser reactor. The authors developed the full set of

partial differential equations that describes the conservation of mass,

momentum, energy and chemical species for both phases, coupled with

empirical correlations concerning interphase friction, interphase heat transfer,

and fluid to wall frictional forces. The model can predict pressure drop,

catalyst holdup, interphase slip velocity, temperature distribution in both

phases, and yield distribution all over the riser. Theologos et al. (1997) [14]

coupled the model of Theologos and Markatos (1993) [13] with a ten lump

reaction scheme to predict the yield pattern of the FCC riser reactor.

An integrated dynamic model for the complete description of the fluid

catalytic cracking unit (FCC unit) was developed by Bollas et al. (2002) [15];

the model simulates successfully the riser and the regenerator of FCC and

incorporates operating conditions, feed properties and catalyst effects.

Erthal et al. (2003) [16], developed a one dimensional, mathematical

model, they considered in their model gas-solid flow that occurs in FCC risers,

two equations of momentum conservation applied to the compressible gas

flow and solid flow respectively, the model considers also the drag force and

heat transfer coefficient between two phases; four lump model used for

cracking reactions description.


7

Souzaa et al. (2003) [17], combined a 2-D fluid flow field with a 6- lumps

kinetic model and used two energy equations (catalyst and gas oil) to simulate

the gas oil cracking process inside the riser reactor.

Das et al. (2003) [18], performed the three-dimensional simulation of an

industrial-scale fluid catalytic cracking riser reactor using a novel density-

based solution algorithm. The particle-level fluctuations are modeled in the

framework of the kinetic theory of granular flow. The reactor model includes

separate continuity equations for the components in the bulk gas and inside the

solid phase.

Berry et al. (2004) [19], modified the two-dimensional hydrodynamic

model to make it predictive by incorporating the slip factor for the calculation

of the cross-sectionally averaged voidage. The model has been coupled with

the four-lump kinetic model to predict the effect of operating conditions on

profiles of conversion, yield, temperature and pressure in the riser.

Hassan (2005) [20], developed Material and energy balance calculations to

design Fluidized catalytic cracking (FCC) unit from Iraqi crude oil. She used

the visual basic program in her work.

With regard to reaction and kinetics, Xu et al. (2006) [21] proposed a

seven lump kinetic model to describe residual oil catalytic cracking, in which

products especially coke were lumped separately for accurate prediction.

Because in recent studies, kinetics was developed accounting for coke

formation leading to catalyst deactivation. The reactor block is modeled as a

combination of an ideal Plug Flow Reactor (PFR) and a Continuously Stirred

Tank Reactor (CSTR).


8

On the other hand, Krishnaiah et al. (2007) [22], a steady state simulation

for the fluid cracking was investigated, the riser reactor was modeled as a plug

flow reactor incorporating four lump model for cracking reactions; they

studied the effect of the operating variables on FCC unit performance, a

catalyst to oil ratio, air rate and gasoil inlet temperature have been chosen as

operating variables.

Souza et al. (2007) [23], bi – dimensional fluid flow combined with six

lumps kinetic model and two energy equations are used to model the gasoil

mixture flow and the cracking process inside the riser reactor.

Gupta et al. (2007) [24] proposed a new kinetic scheme based on

pseudocomponents cracking and developed a semi-empirical model for the

estimation of the rate constants of the resulting reaction network. Fifty

pseudocomponents (lumps) are considered in this scheme resulting in more

than 10,000 reaction possibilities. The model can be easily used to incorporate

other aspects of the riser modeling.

Ahari et al. (2008) [25], a one dimensional adiabatic model for riser reactor

of FCC unit was developed, the chemical reaction were characterized by a

four lump kinetic model, in their study, four cases of industrial riser operating

conditions have been adopted and the modified kinetic parameters are used to

eliminate the deviations between calculated and real values, also simulation

studies are performed to investigate the effect of changing process variables.

Based on Ahari et al. (2008) [25] study, Heydari et al. (2010) [26] performed

an excessive analysis to gasoline yield throughout the riser with respect to

different inlet mixing temperatures, different feed rates and different catalyst

to oil ratios.


9

Shakoor (2010) [27] developed a computer program using MATLAB 7

software to determine the rate constants of FCC unit cracking reactions

represented by six lump model and at any certain temperature.

Baurdez et al. (2010) [28] proposed a method for steady-state/transient,

two phase gas–solid simulation of a FCC riser reactor. Authors used a simple

four lump kinetic model to demonstrate the feasibility of the method

Osman et al. (2010) [29] developed a kinetic model to simulate the riser of

a residue fluid catalytic cracking unit (RFCC) at steady state. The model based

on combination the material and energy balance equations with seven lump

model and a modified two dimensional hydrodynamic model. Simulation has

been performed based on the data from an operating unit at Khartoum

Refinery Company (KRC). MATLAB environment has been used to solve and

analyze the kinetic model and process variables.

A control system of a fluidized-bed catalytic cracking unit has been

developed by AL-Niami (2010) [30]. In this work the dynamic and control

system based on basic energy balance in the reactor and regenerator systems

have been carried out. For the control system, the important input variables

were chosen to be the reactor temperature and the regenerator temperature.

CHAPTER THREE

MATHEMATICAL

MODELING

Chapter Three Mathematical Modeling

10

3.1. INTRODUCTION

In this chapter, a mathematical model for the riser of an industrial FCC is

developed, based on the reactor/regenerator configuration presented in the

Figure 3.1. The preheated raw oil and steam are introduced into the reactor

riser at a point near the base of the riser above the MxR Chamber. Here, the

feed is contacted with a controlled amount of regenerated catalyst and lift

media from the MxR Chamber. The regenerated catalyst flow is controlled

to maintain a desired reactor temperature, and the spent catalyst

recirculation to the MxR Chamber is controlled manually or by ratio to the

regenerated catalyst flow. Feed and steam are mixed and injected through

the feed nozzles distributors. At the distributors the riser diameter increases

to allow for the expansion of hydrocarbon vapors as the oil is vaporized

when it meets the catalyst. As a result of feed vaporization, the cracking

reactions start and the density of the oil decreases causing an increase in the

velocity of the vapor/gas phase. The increasing gas phase velocity

accelerates the velocity of the catalyst and the riser behaves as a transport

bed reactor. The Gasoil is converted to gasoline range hydrocarbons, light

gases and coke. The cracking reactions’ by product (coke) gets deposited

on the catalyst surface and decreases its activity as the catalyst moves

toward the exit of the riser. Because the riser volume is small, it limits the

contact time between the catalyst and hydrocarbon to 5 seconds or less, and

prevents over cracking of the feed[10].

3.2. REACTOR/REGENERATOR MATERIAL & ENERGY

BALANCES

The material and energy balances around the reactor and the

regenerator can be calculated by defining the input and output streams

(Figure 3.2).


11

Figure 3.1 schematic of FCC unit reactor/regenerator system used in

present model [31]

3.2.1. MATERIAL BALANCE

The material balance for any system at steady state is defined as:

Mass in = Mass out

3.2.1.1. REACTOR MATERIAL BALANCE

Mass in = Mass of (feed + steam + regenerated catalyst)

Mass out = Mass of (reactor vapor + spent catalyst + steam)

Where, the oil feed contains small quantity of sulfur, portion of the sulfur

goes with spent catalyst and burned to SO2 in the regenerator, the

remainder exists with products; and steam inlet is equal to summation of

lifting steam, injection steam and stripping steam.


12

Assuming steam inlet does not condense and is present in the exit vapor

products at the same rate, therefore reactor material balance can be

expressed as[33]:

mf + msi + mrcat = mpr + mso + mscat (3.1)

Since, mscat = mrcat + mcokandmsi = mso

Then, eq. (3.1) eliminated to:

mf = mpr + mcok (3.2)

3.2.1.2. REGENERATOR MATERIAL BALANCE

Mass in = Mass of (spent catalyst + air for coke burning)

Mass out = Mass of (flue gases + regenerated catalyst)

OR

mscat + mair = mfg + mrcat (3.3)

Figure 3.2 Input and output streams for reactor and regenerator in FCCU

∆H Reaction

∆H Combustion

of Coke

FEED (mf)

Steam (msi)

Product vapors

+ steam (mpr+mso) Flue gases

(mfg)

Air (mair)

Radiation losses

Heat removal

Radiation losses

mrcat

mscat


13

3.2.2. ENERGY BALANCE

The hot regenerated catalyst supplies the bulk of the heat required to

vaporize the liquid feed to provide the overall endothermic heat of

cracking, and to raise the temperature of dispersion steam and inert gases to

the reactor temperature. [40] The energy balance equation at steady state may

be written as:

Energy in + Energy produced = Energy out + Energy consumed

3.2.2.1. REACTOR ENERGY BALANCE

Energy in = Energy of (feed + regenerated catalyst + steam)

Energy produced = 0

Energy out = Energy of (reactor vapors + spent catalyst + radiation losses)

Energy consumed = Heat of reaction

If the Reactor temperature is the reference base temperature, then

-ΔHfeed - ΔHsteam + ΔHregenerated catalyst = ΔHradiation losses + ΔHReaction

Or

ΔHregenerated catalyst = ΔHfeed + ΔHsteam + ΔHradiation losses +ΔHReaction (3.4)

3.2.2.2. REGENERATOR ENERGY BALANCE

Energy in = Energy (air with moisture + spent catalyst + coke)

Energy produced = Combustion heat of coke

Energy out = Energy (flue gas with moisture + regenerated catalyst

+removed by catalyst cooler + radiation losses)

Energy consumed = 0

If the Regenerator temperature is the reference temperature then,

ΔHcombustion of coke = ΔHcatalyst + ΔHair + ΔHsteam+ ΔHcoke+ ΔHremoved +

ΔHradiation losses (3.5)


14

The enthalpy change for the spent and regenerated catalyst is given by

ΔHspent catalyst = mcatCpcat(Regen. Temp.- Reactor Temp.)

ΔHregenerated catalyst = mcatCpcat(Reactor Temp. –Regen. Temp.)

At steady conditions,

ΔHspent catalyst + ΔHregenerated catalyst = 0

3.3. RISER MODEL

For numerical computation, riser is divided into equal sized segments

of thickness (dz), forming sequential equal sized volume elements (see

Figure 3.3).

Figure 3.3 A volume element in the riser reactor

3.3.1. MODEL ASSUMPTIONS

In order to develop a mathematical model for the riser reactor, the

following assumptions are introduced:

z

dz

Z = 0

Z = L


15

a. One dimensional transported plug flow reactor prevails in the riser

without radial and axial dispersion

b. Steady state operation

c. The riser wall is adiabatic.

d. Viscosities and heat capacities for all components in vapor phase are

constant along the riser.

e. The pressure change through the riser length is due to the static head of

catalyst in the riser.

f. The coke deposited on the catalyst does not affect the fluid flow

g. Instantaneous vaporization occurred in entrance of riser.

h. Each volume element is assumed to contain two phases (i) solid phase

(catalyst and coke) and (ii) gas phase (vapors of feed and product

hydrocarbon, and steam).

i. Each volume element, solid and gas phases are assumed to be well

mixed so that heat and mass transfer resistances can be ignored, and the

two phases have the same temperature.

j. The gas-solid flow is fully developed along all the riser height.

3.3.2. CRACKING REACTIONS KINETICS

The FCC process involves a network of reactions producing a large

number of compounds. Therefore, lumping models can be used to describe

the reaction system in terms of the feed and a defined number of products,

the agglomeration of many chemical compounds into a single compound

(called a lump),should exhibit some or several common properties(i.e.

boiling point, molecular weight, reactivity).In this work four lump model

scheme has been selected (Figure 3.4). This scheme consists of (VGO feed,

Light gases, Gasoline, and Coke), it is more realistic and simple to solve,

with more lumps, the mathematic becomes more complicated.


16

Figure 3.4 Schematic of four lumped reactions

According to this scheme, a part of gasoline is also converted to light gases

and coke. It is assumed that cracking reaction rate is second order with

respect to Gasoil, and first order with respect to Gasoline, and the reactions

take place only in the vapor phase[3]. Rate constants (Kj) for cracking

reactions follow the Arrhenius dependence on temperature (equation 3.6). = (3.6)

Where, the kinetic parameters (Koj and Ej) for cracking reactions are

selected from the literatures[23, 25]. In order to fit the predicted gasoline yield

with industrial gasoline yield, the selected frequency factors can be scaled

linearly by dividing each one by the modified frequency factor (Ko1)ofthe

reaction feedstock → gasoline: = (3.7)

While, the selected activation energies are used directly in the industrial

scale unit model. This approach has been adopted by Ancheyta (2011) [34].

VGO Gasoline

Light gases

Coke

K2

K1

K3

K4

K5


17

3.3.3. CONCENTRATION, TEMPERATURE, PRESSURE AND

COKING TIME PROFILES IN THE RISER

In order to calculate the concentration profile for each lump

throughout the riser height, a differential material balance can be applied

along the riser, the following set of equations is obtained[3, 22, 33]:

For VGO lump: = − ∙ ∙ ∙ [ + + ] ∙ (3.8)

For gasoline lump: = ∙ ∙ ∙ [ ∙ − ( + ) ∙ ] (3.9)

For light gases lump: = ∙ ∙ ∙ [ ∙ + ∙ ] (3.10)

For coke lump: = ∙ ∙ ∙ [ ∙ + ∙ ] (3.11)

The temperature profile along the riser can be calculated using following

energy balance equation [3, 22, 33]: = − ∙ ∙ ∙ + ( ∆ + ∆ + ∆ ) ∙ +( ∆ + ∆ ) ∙ (3.12)

Where, the term {– ϕ [(K1∆H1+K2∆H2+K3∆H3) y12 + (K4∆H4+K5∆H5) y2]}

represents the energy absorbed by endothermic reaction of vapor phase in

the riser [16].

The pressure change throughout the riser can be predicted by[34]: = − ∙ ∙ (1 − ) (3.13)


18

The catalyst residence time can be calculated using following equation [33]: = ∙ ∙ + [ ] ∙ ∙ (1 − ) ∙ ∙ ∙ (3.14)

The variation of the vapor phase mass flow rate (mg) throughout the riser

can be predicted by the following equation [33]: = ( + + ) + + (3.15)

Where, the quantities of dispersion steam (mds) and lifting steam (mls) that

inlets the riser are controlled as 1wt% and 3.5wt% of the feed rate,

respectively. The vapor phase density considered ideal gas and calculated

by: = ∙ 101325 ∙ (3.16)

The average vapor phase molecular weight expressed as [16]: = 1 + + + ( )/ (3.17)

3.3.4. CATALYST DEACTIVATION

The catalyst deactivation function ( )due to coke deposition,

generally, there are two ways of its representation, one that depends on the

catalyst contact time and the other one depends on the catalyst coke

content. Since the functions that depend on the contact time do not account

for the efficiency of the regenerator, i.e. the catalyst activity at the riser

inlet. Therefore, in this work a single deactivation function depending on

the catalyst coke content has been formulated from (Figure B-4) in

appendix B using curve fit technique:


19

= 1 − 45 × ( × ) + (3.18)

3.3.5. RISER HYDRODYNAMICS

After complete vaporization of feed, only solid phase (catalyst &

coke) and vapor phase (steam, hydrocarbon feed and product vapor) are

left. Based on the model assumption, the two phases in fully developed

flow, the empirical correlation (equation 3.19) developed by Patience et

al. (1992) for calculating slip factor can be applied. Slip factor, defined as

the ratio of interstitial gas velocity to average solids velocity.

= = ∙ = 1 + 5.6 + 0.47 ∙ . (3.19)

Where = ∙ (3.20)

And = ∙ (3.21)

The superficial gas velocity is calculated by: = ∙ (3.22)

And the average particle velocity calculated by [25]: = ∙ ∙ (1 − ) (3.23)

By combination of equations (3.19),(3.22) and (3.23), the average void

fraction of the vapor phase can be evaluated by:


20

= ∙ ∙ ∙ + ∙ (3.24)

Therefore, the gas and particle velocities can be evaluated by: = (3.25)

= (3.26)

The residence time of the gas phase can be represented as the ratio of

distance and velocity as: = (3.27)

In order to calculate the particle terminal velocity, many correlations used

and can be found in the literatures. In general, the terminal velocity is

usually calculated for three zones: Stokes, intermediate and Newton zone,

and it classified according to Archimedes numbers which defines the

border between different zones. The Stokes regime holds for Ar< 32.9,

Intermediate regime is valid for 32.9 <Ar< 106.5, and the Newton regime is

defined for Ar> 106.5. In this work the simple correlation (equation3.29) for intermediate regime has been employed for calculating Reynolds

number based on particle terminal velocity[35, 36, 37]:

= ∙ − ∙ ∙ (3.28)

= 18 + (2.3348 − 1.7439 ∙ ℎ) ∙ . (3.29)

= ∙ ∙ (3.30)


21

3.3.6. MIXING TEMPERATURES

In order to calculate the mixing temperature (Tmix1) at the MxR

chamber, where the hot regenerated catalyst is mixed with the carbonized

catalyst which has the same riser outlet temperature, and the mixed

catalysts are lifted by steam (Figure3.5). The energy balance around MxR

chamber can be applied to obtain the following equation:

= ( + ) + ( + ) − ( ) ( + ) + ( + ) +

(3.31)

Figure 3.5 Mathematical representation of reactor riser used in the model

MxR chamber mxcat

mxcok Tout

mrcat mrcok Trcat

mls Tls

mcat mcok mls Tmix1

mf Tfl API Cpfv

mds Tds

mcat mcok mg Tmix2

mscat mscok Tout

mg Tout

Rege

nera

tor

Carbonized catalyst

Regenerated catalyst

(Lifting steam)


22

For calculating the riser inlet mixing temperature (Tmix2), where the

lifting (regenerated/carbonized catalysts + steam) are mixed with injected

(feed + dispersion steam), the energy balance (equation3.32) around the

riser inlet can be formulated.

( − ) + ( − ) + ( − )= ( − ) + 2.32 × × [ 1(1.8 − 459.67) − 2(1.8 − 459.67) + 3 − ] (3.32)

Solving equation (3.32) for Tmix2, and taking the positive root for the quadratic equation gives: = − + √ − 4 2 (3.33) Where = 2.32 × 3.24 × × 1 (3.34) = + + + − 2.32 × 1654.812× × 1 − 2.32 × 1.8 × × 2 (3.35) = 2.32 × (459.67) × × 1 + 2.32 × 459.67 × × 2 + 2.32× × 3 − 2.32 × × − − × ( + + ) (3.36)

Note that equation (3.33) represents the initial boundary condition to the

differential equation (3.12)


23

3.4. HEAT OF COMBUSTION AT THE REGENERATOR

At the regenerator where coke on the catalyst is burning off, the heat

consumed (kJ/s) for catalyst heating up inside the regenerator can be

calculated: ∆ = ( − ) (3.37)

And the heat consumed (kJ/s) for heating up the coke, air and moisture plus

heat losses and removed are assumed 38% of total heat of combustion of

the coke. Therefore, equation 3.5 becomes: ∆ = ∆ 1 − 0.38 (3.38)

Therefore, the heat of combustion at the regenerator per 1Kg of coke burnt,

can be calculated as follows:

∆ = ∆ . × (3.39)

3.5. MODEL SOLUTION

A computer program presented in Appendix C for the model

simulation was developed using polymath version 5.1 and Microsoft Excel

worksheet 2007, based on the 4th order Runge – Kutta method numerical

technique; and a sequential approach has been chosen in this solution. In

the present work the height of each volume element was kept 5cm. Further

decrease in the height of the volume element had no appreciable effect on

the results.

The sequence of calculation steps is listed below and the flow diagram for

the same sequence is given in Fig. 3.6. The model results and discussions

are presented in chapter four.


24

Sequence of calculation steps:

1. Read the input data required for calculation of MxR chamber

temperature.

2. Calculate Tmix1 (equation 3.31)

3. Read the input data required for calculation of riser inlet temperature.

4. Calculate Tmix2 (equation 3.33)

5. Read the input initial values of ODEs.

6. Calculate the variable parameters (ϕ, K1, K2, K3, K4, K5, MWg, ρg,

mg, Uo, Ar, Ret, Ut, ψ, εg, Ug, Up, X) using the appropriate

correlations. In this step all the calculated variable parameters are

represent the conditions of current volume element. As the conditions

at the exit of current volume element are the same as the conditions at

the inlet of the next volume element, therefore, use these calculated

values as an initial values for calculation the next volume element

conditions, equations (3.6) and (3.15) to (3.30).

7. Go to the next volume element with (z = z + 0.05).

8. Calculate the new values of ODEs depending on the calculated values

in step 6, equations (3.8) to (3.14).

9. Repeat steps 6 - 8 until the sum of increment height equals the height

of the riser.

10. Calculate the yields and conversion at the exit of the riser.

11. Calculate the cracking efficiency, selectivity, WHSV, delta coke,

∆Hcombustion of coke.


25

Figure 3.6 computational flow diagrams for riser reactor model

START

From E. B. Calculate Tmix1, eq. (3.31)

From E. B. Calculate Tmix2, eq. (3.33)

INPUT mf , Tf , Kuop , API , SG , mds , Tds , Cpg

Calculated and updated values ϕ , K1, K2 , K3 , K4 , K5, MWg , ρg

, mg , uo , Ar , Ret , ut , ψ , εg , ug , up , X, t

Equations (3.6) and (3.15) to (3.30)

Z < L

END

NO

YES

Calculate the yields and conversion

Z = 0

INPUT initial values of ODEs from (3.8) to (3.14)

y1=1 , y2= y3= y4= 0 , tc= 0, T = Tmix2, P = Pin

INPUT mrcat , mrcok,Trcat , mxcat , mxcok , Tout , mls , Tls , Cpcat , Cpcok , Cps

Solve above seven ODEs by Runge – Kutta method

Z = Z + 0.05

CHAPTER FOUR

RESULTS AND

DISCUSSIONS

Chapter Four Results and discussion

26

4.1. INTRODUCTION

As discussed in previous chapter, the material and energy balance

equations were combined with reaction kinetics and the hydrodynamic

model equations to obtain a model capable to predicting the yield pattern

along the riser height. In this chapter the proposed model is validated

directly by comparing the model results with plant data. Model results are

plotted in the following figures with a brief discussion. An Excel worksheet

was developed for modeling of new and existing FCC unit, and to predict

the effect of any change in operating parameters on unit performance.

4.2. CASE STUDY

For model validation, commercial FCC unit (30000 BPSD) designed

to handle hydrotreated VGO feed was selected. The unit operates for

maximum gasoline mode, therefore, no recycle occurs at normal operation [1], with (0.5/1) mixing ratio of carbonized catalyst to regenerated catalyst

at MxR chamber.

Table 4.1 shows the kinetic parameters for cracking reactions with

adjustable frequency factors utilizing the productivity of studied case. This

approach has been adopted by many researchers[25, 26, 34].


27

Table 4.1 Kinetic parameters with Modified frequency factors used in present model

Frequency factor Koj*

Activation energy[16]

Ej (kJ/kmol)

Heat of reaction[25]

∆Hj (kJ/kg)

Gasoil to gasoline 12500 57359 393 Gasoil to light gases 1950 52754 795 Gasoil to coke 16 31820 1200 Gasoline to light gases 2650 65733 1150 Gasoline to coke 550 66570 151

*For feedstock cracking the Koj units are (wtfrac./s); for other lump cracking the Koj units are (s-1)

4.3. MODEL RESULTS

Table 4.2 lists the calculated and design values for mixing

temperatures at MxR chamber and riser inlet region.

Table 4.2 Mixing temperatures at MxR chamber and Riser inlet temperature Present model Licensor design data MxR temperature (K) 912.74 912 Riser inlet temperature (K) 833.99 834

Product yield profiles and actual yield values have been modeled at the

riser exit shown in Figure 4.1. It shows the chemical reactions are faster

near the riser inlet where the higher gradients of the variables take place.


28

Figure 4.1 Four lump concentration profile vs. riser height

The riser temperature profile is plotted in Figure 4.2, where the

temperature drops quickly in the first few meters of the riser bottom zone,

indicating that most of the reactions occur at the riser bottom.

Figure 4.2 Riser temperature profile

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

0 5 10 15 20 25 30 35

wt f

ract

ion

Riser Height (m)

VGO present modelGLN present modelLGS present modelCOK present modelPlant dataPlant dataPlant dataPlant data

800

805

810

815

820

825

830

835

840

0 5 10 15 20 25 30 35

Tem

pera

ture

(K)

Riser Height (m)

Present model

Plant data Tout

Plant data Tin


29

The predicted pressure profile of the riser is shown in Figure 4.3.The

pressure drops 15.3 kPa only, while the actual pressure drop that exists in

FCC risers is (35 – 62) kPa[40]. This difference is due to the simple

assumption of present model, where static head of catalyst prevails,

neglecting the frictions effect.

Figure 4.3 Riser pressure profile

The gas phase molecular weight along the riser height is shown in

Figure4.4, where the molecular weight decreases due to the increase of

light products percentage in the gas phase towards the riser exit. Although,

the steam quantity is 4.5% of the total gas phase, however, the model gives

better result of molecular weight when taking into account this small

quantity of steam. As a result of decreasing the gas phase molecular

weight, the gas density also decreases as shown in Figure 4.5.

225000

227000

229000

231000

233000

235000

237000

239000

241000

243000

245000

0 5 10 15 20 25 30 35

Pres

sure

(Pa)

Riser Height (m)

Present model


30

Figure 4.4 Gas phase molecular weight vs. riser length

Figure 4.5 Gas phase density vs. riser height

0

20

40

60

80

100

120

140

160

180

200

0 5 10 15 20 25 30 35

Kg/K

mol

e

Riser Height (m)

Present Model

Plant data

0

1

2

3

4

5

6

7

8

0 5 10 15 20 25 30 35

Den

sity

(Kg/

m3)

Riser Height (m)

Present Model

Plant data


31

The mass flow rate of the gas phase through the riser is plotted in

Figure4.6; it is decreases slightly due to the coke formation.

Figure 4.6 Gas phase mass rate vs. riser height

The slip factor is plotted in the Figure 4.7; slip factor is high at the

beginning of the riser indicating the gas velocity is much greater than the

particle velocity, while the gas phase velocity increase due to cracking

reactions; it accelerates the catalyst velocity resulting decrease in slip

factor. The slip factor values may range from 1.2 to 4, where 2 considered a

typical in a commercial FCC unit[7].

The gas phase and solid phase velocities are plotted in Figure 4.8. The

figure shows the two phases at the maximum values at the riser exit.

50.5

51

51.5

52

52.5

53

53.5

54

0 5 10 15 20 25 30 35

Gas

pha

se m

ass

rate

(Kg/

s)

Riser height (m)

Present Model

Plant data


32

Figure 4.7 Slip factor vs. riser height

Figure 4.8 Gas phase and catalyst velocities vs. riser height

0

1

2

3

4

5

6

0 5 10 15 20 25 30 35

ψ

Riser Height (m)

Slip factor

0

2

4

6

8

10

12

14

16

0 5 10 15 20 25 30 35

Velo

city

(m/s

)

Riser Height (m)

Model Gas Phase velocity

Model Solid phase velocity


33

Figure 4.9 shows the gas phase and catalyst residence time in the riser.

The gas residence time is 2.2 seconds, and this agrees with riser design

criteria (1 to 5 seconds).While the catalyst is heavier than the gas, it stays

about 7 seconds inside the riser.

Figure 4.9 Gas phase and catalyst residence times vs. riser height

Figure 4.10 indicates a molar expansion of the gas phase as a function

of cracking reaction. It shows that at 2m of the riser height is locally 92%

occupied by the gas phase, and this agrees with Figure 51 of Gauthier

(2009) [41]The catalyst deactivation function is plotted in Figure 4.11. The

catalyst activity at the beginning of the riser is around 0.8 indicates to the

effect of mixed carbonized catalyst of lower activity, while the catalyst

activity at the riser exit is around 0.44.

0

1

2

3

4

5

6

7

8

9

0 5 10 15 20 25 30 35

Resi

denc

e tim

e (s

ec.)

Riser height (m)

Vapor residence time

Catalyst residence time


34

Figure 4.10 Gas phase void fraction vs. riser height

Figure 4.11Predicted catalyst activity along the riser height

0.7

0.75

0.8

0.85

0.9

0.95

1

0 5 10 15 20 25 30 35

εg

Riser Height (m)

Present Model

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

0 5 10 15 20 25 30 35

ϕ

Riser Height (m)

Present Model


35

Figure 4.12 shows the maximum gasoline yield occurs at 82.4wt% of feed

conversion; this conversion located at 31.65 m of riser height (Figure 4.13),

it agrees well with the riser size of a commercial FCC unit.

Figure 4.12 Gasoline yield vs. feed conversion

Figure 4.13 Feed conversion vs. riser height

0

0.1

0.2

0.3

0.4

0.5

0.6

0 20 40 60 80 100

Gas

olin

e yi

eld

(wt f

ract

ion)

feed

Conversion (wt %)

Present Model

0

10

20

30

40

50

60

70

80

90

0 5 10 15 20 25 30 35

Conv

ersi

on w

t%

Riser height (m)

Present Model


36

Table 4.3 shows a good agreement between predicted and plant data for

the variable parameters with acceptable deviation percentages; while

Table4.4 shows the case study results.

Table 4.3 Model predicted and Plant data comparison

Variable parameter Model prediction value Plant data % dev.

MxR chamber temperature (K) 912.74 912 0.081 Riser inlet temperature (K) 833.99 834 -0.0012 Light gases yield (kg/s) 14.0844 14.65 - 3.86 Gasoline yield (kg/s) 25.6186 25.7 - 0.32 Unconverted VGO yield (kg/s) 8.94922 8.9 - 0.55 Coke yield (kg/s) 2.54777 2.56 - 0.48 Riser outlet temperature (K) 802.888 802 0.11 Riser outlet pressure (kPa) 228.747 228.81 0.027 Gas phase outlet MW (kg/kmole) 71.656 72.6 -1.3

Gas phase outlet density (kg/m3) 2.457 2.43 1.11

Gas phase outlet mass rate (kg/s) 50.9562 51 -0.086

Table 4.4 Case study results

Unit variable value

Efficiency 60.6 Selectivity 1.54 Delta Coke 1.2 C/O ratio 6.09

WHSV (hr-1) 77.44 ∆H combustion of coke (kJ/kg coke) 80459.8

CHAPTER FIVE

CONCLUSION

AND

RECOMMENDATION

Chapter Five Conclusions and Recommendations

38

5.1. CONCLUSIONS

In the present work a mathematical model has been developed for the

riser reactor of new reactor/regenerator configuration designed by UOP.

Based on the results of present study, the following conclusions can be

derived:

• The formulated energy balance equation (3.31) for calculation of

the mixing temperature (Tmix1) at MxR chamber considers the effect

of coke content for both regenerated and circulated carbonized

catalysts, making the model flexible and more predictive.

• The developed procedure for calculation of the riser inlet mixing

temperature (Tmix2) depends on defining specific gravity and Kuop

property of the feed into the energy balance equation (3.33). This

approach is more predictive than defining heat capacity, normal

boiling point and heat of vaporization of the feed which has adopted

by many others.

• Considering the small quantity of the steam with hydrocarbons

inside the riser converges the model results to better values,

especially when calculating the gas phase molecular weight, where

the deviation percentage was reduced from 16% to -1.3%.

Chapter Five Conclusions and Recommendations

39

5.2. RECOMMENDATIONS FOR FUTURE WORK:

The following suggestions for future work can be considered:

• Using CFD to visualize the unit performance in order to improve

design of internals.

• Developing the model to include the cyclone, VSS, and regenerator

performance as well.

• Developing the model using more lumps for kinetics, multi-

dimensions for the riser, and multi-phase system.

• Developing the model to taking account the friction effects between

phases with the wall of the riser and between phases itself.

• Study the effect of change of any operating conditions variables on

unit performance, i.e. feed temperature, feed type and C/O ratio.

Appendix A FCC technologies

A-1

APPENDIX A

A-1. Fluidized Catalytic Cracking Technologies

There are number of different proprietary designs that have been

developed for modern FCC units. Each design is available under a license that

must be purchased from the design developer by any petroleum refining

company desiring to construct and operate an FCC of a given design.

Basically, there are two different configurations for an FCC unit: the "stacked"

type, where the reactor and the catalyst regenerator are contained in a single

vessel with the reactor above the regenerator Figure A-1a, and "side-by-side"

type, where the reactor and catalyst regenerator are in two separate vessels

Figure A-1b. The major FCC designers and licensors are: [2, 42]

A- Stacked configuration:

• Kellogg Brown & Root (KBR)

B- Side-by-side configuration:

• CB&I Lummus

• Stone & Webster Engineering Corporation (The Shaw Group) /

Institut Francais Petrole (IFP)

• Shell Global Solutions International

• ExxonMobil Research and Engineering (EMRE)

• Universal Oil Products (UOP) - currently fully owned subsidiary

of Honeywell


A-2

Figure A-1a Stacked Reactor regenerator configuration [43]

Figure A-1b Side by side Reactor regenerator configuration [43]


A-3

These two different configurations can be classified into small scale and

large scale units as follows:

• A “stacked” configuration of the reactor and regenerator, generally

applicable to small scale units processing less than 20,000 BPSD.

This was used for the small scale unit, with a feed rate of 2,500

BPSD.

• A “side-by-side” configuration of the reactor and regenerator,

generally applicable to large scale units processing 20,000 or more

barrels per stream day of fresh feed. This was used for the large

scale unit, with a feed rate of 62,000 barrels per stream day (BPSD).

In general, all the major licensors agree that the “stacked” design requires

less plot space and less structural steel, thus lowering the installed cost. As

capacity increases, the vessel sizes also increase to a point where structural

cost exceeds this advantage. As a rule of thumb, a capacity of 20,000 BPSD is

the cut-off size where “side-by-side” design becomes more economical than

“stacked” design, according to UOP. There are other factors that affect the

cost/BPSD for the same capacity and thus affect the selection of the FCC

configuration. These factors include location, available land space, and local

building codes.

Over the years, with improvements in technology and increasing gasoline

demand, most of the refiners have modified their FCC units to increase

capacity. Several FCC units have been revamped to process more than

100,000 BPSD. According to the major licensors, the typical capacity of new

FCC units is 40,000 to 60,000 BPSD.


A-4

Less vendor detail and previous design information was available for

estimation of the small scale FCC. In general, small scale FCC units are no

longer economic, and very few units under 20,000 BPSD are being built

today[43].

A-2. New designs for FCC units

The theory behind the design of any FCC unit configurations is the

same, unless the variations of design by the licensors, the component that

make up the FCC unit are the same: the riser, reactor, and regenerator [43].

A brief description of new technologies and advanced features identified by

each licensor is describing below.

A-2.1. KBR and ExxonMobil

Feed enters through the proprietary ATOMAX-2™ feed injection

system, (Figure A-2). Reaction vapors pass through a right angle turn and are

quickly separated from the catalyst in a positive pressure CLOSED

CYCLONE system which minimizes dry gas make and increases gasoline

yield. Spent catalyst flows to the regenerator through a stripper equipped with

DYNAFLUX™ baffles technology, featuring proprietary FLUX Tubes™ and

Lateral Mixing Elements™ which Reduces entrained hydrocarbons by 80

percent compared to conventional stripper designs and may achieve six

percent or less hydrogen in the coke.

REGENMAX™ Provides staged regeneration in a simple, single regenerator

vessel where counter-current flow of catalyst and air contacting is carried out;

spent catalyst is evenly distributed across the top of the regenerator bed,

causing a chemical reaction 2C + 2NO → 2CO + N2 , only 5 to 7 percent of

the nitrogen in the coke escapes as NOx.


A-5

Figure A-2 KBR’s counter-current regeneration design [46]

Catalyst flow from the regenerator to the external vertical riser is controlled by

riser outlet temperature, which regulates the regenerated catalyst slide valve.

A plug valve located in the regenerator bottom head controls the level in the

stripper by regulating the catalyst flow from the spent catalyst standpipe.

Flue gas flows to an external plenum and then to the flue-gas system.

A Cyclofines Third Stage Separator (TSS) may be used to remove particulates

from the flue gas for protection of power recovery expander and/or

compliance with particulate emissions standards, where TSS is capable to

achieve less than 50 mg/Nm3 particulates in flue gas [44, 45].


A-6

Riser Quench Technology consists of a series of nozzles uniformly spaced

around the upper section of riser. A portion of the feed or a recycle stream

from the main fractionator is injected through the nozzles into the riser to

rapidly reduce the temperature of the riser contents. The heat required to

vaporize the quench is supplied by increased fresh feed preheat or by

increased catalyst circulation. This effectively increases the temperature in the

lower section of the riser above that which would be achieved in a

nonquenched operation, thereby increasing the vaporization of heavy feeds,

increasing gasoline yield, olefin production, and gasoline octane [1].

The process features developed by KBR & ExxonMobil for FCCU and its

advantages are tabulated in Table A-1.

Table A-1

KBR & ExxonMobil FCCU Technologies [45] Process features Process benefits

ATOMAX-2™ Feed Nozzles Produces 43 percent smaller droplet size than the leading industry ATOMAX-1™ nozzles

Riser Quench Technology Improves gasoline yield and octane, and prevents over-cracking that result in undesirable products

Closed Cyclone Riser Termination

Minimizes dry gas make and increases gasoline yield using a simple compact cyclone system

DynaFlux™ Stripping baffles Lower regenerator temperature, increase catalyst circulation and higher conversion.

REGENMAX™ Reduces regenerator emissions and entrainment and provides clean catalyst burning under partial CO combustion at minimum investment

MAXOFIN™ Offers the ability to adjust propylene yield, propylene/ethylene ratio and olefins/ liquid fuels ratio as market conditions demand


A-7

A-2.2. LUMMUS, a CB&I company

The Lummus process incorporates an advanced reaction system, high-

efficient catalyst stripper and a mechanically robust, single stage fast fluidized

bed regenerator. Oil is injected into the base of the riser via proprietary

MICRO-Jet injection nozzles. Catalyst and oil vapor flow upwards through a

short-contacting time, all vertical riser where raw oil feedstock is cracked

under optimum conditions (Figure A-3).

Reaction products exiting the riser are separated from the spent catalyst in a

patented, DIRECT-COUPLED cyclone system. Product vapors are routed

directly to fractionation, thereby eliminating nonselective post-riser cracking

reactions and maintaining the optimum product yield slate. Spent catalyst

containing only minute quantities of hydrocarbons is discharged from the

diplegs of the direct-coupled cyclones into the cyclone containing vessel. The

catalyst flows down into the stripper containing proprietary MODULAR

GRID (MG) baffles.

Figure A-3 Lummus FCCU Process Flow Diagram [47]


A-8

Trace hydrocarbons entrained with spent catalyst are removed in the MG

stripper using stripping steam. The MG stripper efficiently removes

hydrocarbons at low steam rate. The net stripper vapors are routed to the

fractionators via specially designed vents in the direct-coupled cyclones.

Table A-2 LUMMUS FCCU Technologies [47]

Process Features Process Benefits Micro-Jet™ feed injectors • Uniformly contact feed with catalyst

• Minimal erosion and catalyst attrition • Low pressure drop

Short contact time riser reactor • Minimal back-mixing • Efficient catalyst/oil contacting

Patented direct-coupled cyclones at the end of the riser reactor for quick and efficient recovery of product vapors

• Minimal after- cracking • Low dry gas yield and delta coke • Minimal hydrocarbon loading in the

stripper Modular Grid (MG) catalyst stripper design

• Highly efficient removal of hydrocarbon product vapors from the catalyst

• Reduced delta coke • Low stripping steam requirement

Dual diameter catalyst regenerator and turbulent bed combustion

• Low carbon on regenerated catalyst • Efficient use of combustion air • Reduced after-burning and NOx

emissions Regenerated catalyst standpipe with external hopper

• Smooth, stable catalyst flow over a wide operating range

• Insensitive to unit upsets Spent catalyst square-bend transfer line and distribution of spent catalyst into the center of the regenerator

• Improved slide valve pressure differentials

• Lower catalyst hydrothermal deactivation

• Stable spent catalyst flow • Lower capital and operating cost


A-9

Catalyst from the stripper flows down the spent catalyst standpipe and through

the slide valve. The spent catalyst is then transported in dilute phase to the

center of the regenerator through a unique square bend spent catalyst transfer

line. This arrangement provides the lowest overall unit elevation. Catalyst is

regenerated by efficient contacting with air for complete combustion of coke.

The resulting flue gas exits via cyclones to energy recovery/flue gas treating.

The hot regenerated catalyst is withdrawn via an external withdrawal well.

The well allows independent optimization of catalyst density in the

regenerated catalyst standpipe. Maximizes slide valve pressure drop and

ensures stable catalyst flow back to the riser feed injection zone [44].

The process features developed by Lummus for FCCU and its advantages are

tabulated in Table A-2.

A-2.3. Stone & Webster Engineering Corporation (The Shaw Group)/

AXENS, Institut Francais Petrole (The IFP Group)

Catalytic and selective cracking in a short contact time riser where oil

feed is effectively dispersed and vaporized through a proprietary feed

infection system. Operation is carried out at a temperature consistent with

targeted yields. The riser temperature profile can be optimized with the

proprietary Mixed Temperature Control (MTC) system, (Figure A-5).

Reaction products exit the riser reactor through a high efficiency, close-

coupled, proprietary riser termination device RSS (Riser Separator Stripper).

Spent catalyst is pre-stripped followed by an advanced high-efficiency packed

stripper prior to regeneration, (Figure A-4). The reaction product vapor may

be quenched to give the lowest possible dry gas and maximum gasoline yield.


A-10

Figure A-4 S&W / IFP FCCU design [48]

Final recovery of catalyst particles occurs in cyclones before the product

vapor is transferred to the fractionation section.

Catalyst regeneration is carried out in a single regenerator equipped with

proprietary air and catalyst distribution systems, and may be operated for

either full or partial CO combustion.

RSS

Packed stripper

Riser Separator Stripper


A-11

Figure A-5 Mix zone temperature control (left), Feed injection nozzle (right) [49]

Table A-3 S&W / IFP FCCU Technologies [1, 49]

Process Features Process Benefits Feed Injection • Primary Atomization: oil impact on target, Steam cross

shearing. • Secondary Atomization: Filamenting in barrel, spry

shaped by tip Mix zone Temperature Control (MTC)

• Provide an independent control of the mix temperature • Achieve a high mix-zone temperature and lower riser

temperature Riser Termination • Rapid separation of products and catalyst

• Compact design • Easy start-up and operation • Reduces dry gas • Reduces coke • Sealed or open operation

Packed stripper • 95% of the whole area is open to flow • Increased catalyst circulation rate and capacity • Stripping is more effective • Robust construction

Cold wall design • Allows the use of carbon steel for construction • Reduces the skin temperatures • Less thermal expansion of the components • Minimizing the need for expansion joints • Minimum capital investment


A-12

The internals of the vessels and transfer lines are covered in insulating and

abrasion resistant refractory, Stone & Webester developed this technology to

make high temperature, high Olefins FCC a reality.

Reliable operation is ensured through the use of advanced fluidization

technology combined with a proprietary reaction system. Unit design is

tailored to the refiner’s needs and can include wide turndown flexibility [44].

The process features developed by Stone & Webster / IFP for FCCU and its


A-2.4. SHELL Global Solution International B.V

Shell’s high-performance feed nozzle system feeds hydrocarbons to a short

contact-time riser. This design ensures good mixing and rapid vaporization

into the hot catalyst stream (Figure A-6).

Cracking selectivity is enhanced by the feed nozzles (Figure A-7) and

proprietary riser internals, which reduce catalyst back mixing while reducing

overall riser pressure drop.

Riser termination design incorporates reliable close-couple cyclones that

provide rapid catalyst/hydrocarbon separation. It minimizes post-riser cracking

and maximizes desired product yields, with no slurry clean up required.

Stripping begins in the staged stripper, equipped with high capacity baffle

structure, (Figure 2.7).


A-13

Figure A-6 Shell’s FCCU design [44]

A single stage partial or full burn regenerator delivers excellent performance

at low cost. Proprietary internals are used at the catalyst inlet to disperse

catalyst, and the catalyst outlet to provide significant catalyst circulation

enhancement. Catalyst coolers can be added for more feedstock flexibility.

Cyclone systems in the reactor and regenerator use a proprietary design, thus

providing reliability, efficiency and robustness. Flue gas cleanup can be

incorporated with Shell’s third-stage separator [44], (Figure A-6).

The process features developed by Shell Global Solutions for FCC unit and its



A-14

Figure A-7 PentaFlow Packing (left), Feed nozzels configurations (right)[50]

Table A-4

Shell’s FCCU Technologies [51]

Process Features Process Benefits Feed injection system • Good riser coverage and mixing with the

catalyst. Resulting higher gasoline yields, less dry gas, lower steam consumption and reduced pressure requirements

Riser internals • improves catalyst distribution and reduces spent catalyst reflux, which minimizes nonselective thermal cracking

Close-coupled reactor cyclones with coke catcher

• providing high separation efficiency, minimizes post-riser cracking and reactor vessel coking

High-efficiency stripper • PentaFlow stripper packing removes up to 95% of hydrocarbons, open design prevents plugging, enhances catalyst flux and facilitates access for maintenance.

Catalyst circulation enhancement technology

• Improves circulation rates by up to 50% and is applicable to both the stripper and the regenerator standpipes; it improves stability and pressure build-up by optimizing the catalyst condition near the inlets.

Third-stage separator (TSS) technology

• Reduces flue gas particulate emissions to less than 50 mg/Nm3. It also protects the flue gas system from erosion


A-15

A-2.5. Universal Oil Products UOP

UOP’s process (Figure A-8) uses a side by side reactor/regenerator

configuration and a patented pre-acceleration zone to condition the

regenerated catalyst. Modern OPTIMIX feed distributors inject the feed into

the riser, which terminates in a vortex separation system (VSS). A high

efficiency stripper then separates the remaining hydrocarbons from the

catalyst, which is then reactivated in a combustor style regenerator, with

RxCat technology, a portion of the catalyst that is pre-stripped by the riser

termination device can be recycled back to the riser via a standpipe and the

MxR chamber.

The reactor zone features a short contact time riser, state of the art riser

termination device for quick separation of catalyst and vapor, with high

hydrocarbon containment (VSS/VDS technology) and RxCat technology,

wherein a portion of the pre-stripped (carbonized) catalyst from the riser

termination device is blended with the hotter regenerated catalyst in a

proprietary mixing chamber (MxR) for delivery to the riser.

With this technology, the reactor temperature can be lowered to reduce

thermal cracking with no negative impact on conversion, thus improving

product selectivity. The ability to vary the carbonized/regenerated catalyst

ratio provides considerable flexibility to handle changes in feedstock quality

and shortens the time for operating adjustments by enabling rapid switches

between gasoline, olefins or distillate operating modes. Since coke yield can

be decreased at constant conversion, capacity and reaction severity can be

increased, and CO2 emissions reduced. Furthermore, because the catalyst

delivered to the regenerator has a higher coke content, it requires less excess

oxygen at a given temperature to sustain the same kinetic combustion rate.


A-16

Figure A-8 UOP’s reactor/regenerator FCCU design [31]

The combustor style regenerator burns coke in a fast-fluidized environment

completely to CO2 with very low levels of CO.

UOP also offers two process technologies for maximizing propylene from

feedstocks traditionally processed in FCC units. The PetroFCC and RxPro

processes are specifically designed to meet increased propylene production

requirements but are flexible to also operate in maximum gasoline mode, if

required.


A-17

Both processes utilize commercially proven technology and mechanical

features found in a conventional UOP FCC design, but are operated at process

conditions that promote light olefin and/or aromatics production for

petrochemical applications. The commercially-proven PetroFCC technology

provides a cost-effective means for producing moderate quantities of

propylene from moderate quality feedstocks. the newest entry to the enhanced

propylene platform at UOP, the RxPro process, employs a multi-stage reaction

system with targeted olefin recracking to achieve a highest yield of propylene

(>20 wt% FF) for a given reaction severity and feedstock quality[1, 44].

Appendix B Variables of FCC units

B-1

APPENDIX B

VARIABLES OF FCC UNITS

B-1. FCC unit Feedstocks and Products

Modern FCC units can take a wide variety of feedstocks and can adjust

operating conditions to maximize production of gasoline, middle distillate

(LCO) or light olefins to meet different market demands. [42]

The FCC unit is also an important source of butene and pentene olefins used

in refinery processes such as the alkylation unit. [44]

B-1.1. Feedstocks

The main feedstock used in FCC unit is the Gas Oil boiling between

316oC and 566 oC (600 oF and 1050 oF). FCC unit can accept a broad range of

gas oil feedstocks such as: [1]

v Atmospheric gas oils.

v Vacuum gas oils.

v Coker gas oils.

v Thermally cracked gas oils

v Solvent deasphalted oils

v Lube extracts

v Hydrocracker bottoms

These gas oils can be considered mixtures of aromatic, naphthanic and

paraffinic molecules. There are also varying amounts of contaminants such as

sulphur, nitrogen and metals, particularly in the higher boiling fractions. These

differences in feed composition and in contaminants affect the operating

conditions required to obtain desired yields. To protect the catalyst, feed pre-

treatment by hydrotreating is required in order to remove contaminants and

improve cracking characteristics and yields.


B-2

The principal limitation on charge stocks are the Conradson Carbon Residue

(CCR) and metal contaminants. The effect of Conradson carbon is to form a

deposit on the catalyst and reduce catalyst activity, promote coke and

hydrogen formation. [33]

Paraffinic atmospheric and vacuum gas oils takes as charge stocks because it's

more easily cracked in the catalytic cracker. [42]

Feedstocks can be ranked in terms of their Crackability, or easily to convert in

FCC unit. Crackability is a function of the relative proportions of paraffinic,

naphthenic, and aromatic species in the feed as shown in table B-1 [1].

Table B-1

Feedstock Crackability[1]

Range of

characteristic factor Kuop Relative Crackability Feedstock type

> 12 High Paraffinic

11.5 – 11.6 Intermediate Naphthenic

< 11.3 Refractory Aromatic

FCC gasoline yield largely depends on changes in feed quality, catalyst

properties, and operating variables, for Feedstock Quality: Paraffinic

feedstocks produce the most gasoline yield (but the lowest octane). The

common indicators of any increase in feed paraffinicity are: [40]

• Increase in the Kuop factor

• Increase in the aniline point

• Increase in the nickel-to-vanadium ratio

• Decrease in the fraction of "cracked" material (Coker Gasoil)


B-3

B-1.2. Products

Table B-2 indicates the typical yields and the essential characteristics of the

products obtained by the catalytic of vacuum distillate (VGO) and the required

treatments.

Table B-2

Typical FCC unit products [53]

Products Yield wt%

Product characteristics – Complementary Treatments - Valorizations

GAS C2 – & H2S

3 – 5 • Large amount of H2S produced with high sulfur feeds

• Requiring purification treatments (amine washing – Claus plant)

C3 CUT - Propane - Propylene

5 – 9 • 70% propylene in the C3 fraction • Propylene recovery by distillation for further

use in petrochemicals C4 CUT

- Butanes - Butenes

6 – 12 • 50 – 60% butenes in the C4 fraction • Possible valorization of butenes in

ALKYLATION plant to increase gasoline production

• Selective valorization of isobutene in MTBE plant

GASOLINE C5 – C11

45 – 55 • Most desirable product • Medium octane rating for unleaded gasoline

RON 92 – 93, MON 79 – 80 • Relatively good stability • Corrosive and bad odor mercaptans require

sweetening by MEROX type process • Possibility to improve octane numbers by

separating a gasoline heart cut and then processing it.

GAS OIL (LCO)

15 – 20 • Very aromatic • Low cetane number • S% generally high • Good diluents for heavy fuel oil

HEAVY streams (HCO + SLURRY)

6 – 10 • Very aromatic products • Diluents for heavy fuel oil

COKE 4 – 6 § Self – consumed as fuel in the regenerator


B-4

B-2. Catalytic Cracking Reactions

The main reaction in the FCC is the catalytic cracking of paraffin, olefins,

naphthenes and side chains in aromatics. A network of reactions occurring in

the FCC is illustrated in Figure B-1. The VGO undergoes the desired ‘primary

cracking’ into gasoline and LCO. A secondary reaction also occurs, which

must be limited, such as a hydrogen transfer reaction which lowers the

gasoline yield and causes the cycloaddition reaction. The latter could lead to

coke formation (needed to provide heat for catalyst regeneration) [33].

Figure B-1 FCC reaction network [54]

The main reactions occurs in the FCC riser reactor can be summarized in the Figure B-2

Heavy Gasoil

Primary Cracking

Gasoline +

LCO

Primary Cracking

Secondary Cracking

Gases H2, C1, C2,

C3, C4

Residue +

Coke

Oligomerization

Cycloaddition Dehydrogenation

Dehydrogenation


B-5

Figure B-2 Principal Reactions in Fluid Catalytic Cracking [48, 55]

Paraffins Cracking

Paraffins + Olefins

Olefins *

Cracking LPG Olefins

Cyclization

Isomerization

H Transfer

Cyclization

Condensation Dehydrogenation

Naphthenes

Branched Olefins

Paraffins

Coke

H Transfer Branched Paraffins

Naphthenes

Cracking Olefins

Dehydrogenation Cyclo-Olefins

Isomerization Naphthenes with different rings

Dehydrogenation

Aromatics

Side-chain

Cracking Unsubstitued aromatics + Olefins

Transalkylation Different alkylaromatics

Dehydrogenation

Condensation Polyaromatics

Alkylation

Dehydrogenation Condensation

Coke

Aromatics

* Mainly from cracking, very little in feed


B-6

B-3. FCC unit Operation Variables

The primary variables available to the FCC unit operator for

maximizing unit conversion or specific desired product for a given feedstock

quality can be divided into two groups, catalytic and process.

B-3.1. Catalytic activity and catalyst design

Earlier FCCU catalysts were natural occurring clays, which suffered from

low cracking activity, poor stability, and poor fluidization characteristics.

Later synthetic silica/alumina catalysts were developed, which were more

active and stable. These were replaced with present day zeolite (crystalline

silica – alumina or molecular sieves) catalysts with greatly increased activity,

stability, and improved selectivity. It has been realized for some time that

better FCC unit yield could be achieved with shorter contact time between the

feedstock and the catalyst by using zeolite catalysts. With the earlier catalysts

of low activity and long residence time, some of the gasoline formed cracked

further in the catalyst bed to LPG resulting in lower gasoline yield [56].

Fig. B-3 Evolution in structure of FCC catalysts before 1990 [5, 57]


B-7

Figure B-3 shows continuous improvements in the performance of FCC

catalysts from the beginning until 1990

Modern FCC catalysts are fine powders with a bulk density of 0.80 to

0.96g/cc, made up primarily of silica and alumina and containing acid sites

that enables the catalyst to crack reactions, and having a particle size

distribution ranging from 10 to 150 μm and an average particle size of 60 to

100 μm [2, 56].

Cracking catalysts for the fluid process are characterized by certain physical

and chemical properties. Physically, the catalyst must have: (a) satisfactory

fluidization characteristics, and (b) sufficient resistance to attrition so that

excessive loss of fine particles is not encountered. On the chemical side, the

catalyst must: (a) have adequate activity and maintenance of activity with age;

and (b) promote the formation of the desired reaction products with minimum

production of undesired by-products such as gas and coke. [58]

Amorphous catalysts have higher attrition resistance and are less costly than

zeolitic catalysts. Most commercial catalysts contain approximately 15%

zeolites and thus obtain the benefits of the higher activity and gasoline

selectivity of the zeolites and the lower costs and make-up rates of the

amorphous catalysts [42].

During the process, Catalyst deactivation occurs due to the metal content of

the feed such as Nickel and Vanadium, where vanadium can destroy the

zeolite activity, and coke deposition which temporarily block some of the

catalytic sites; The spent catalyst entering the regenerator contains between

0.4 to 2.5 wt% coke [40], this coke is burned off in the regenerator and

Regenerated catalyst contains from 0.01 to 0.4%wt residual coke (CRC) [42],


B-8

this is an important parameter for a unit operator to monitor periodically. The

CRC is an indicator of regenerator performance (Figure B-4).

Catalyst loss also occurs due to the attrition, in order to maintain the catalyst

activity and replace the lost catalyst, daily makeup rate of fresh catalyst is

added, typically 0.045 to 0.14 Kg per barrel of fresh feed [40].

Figure B-4 Catalyst activity retention vs. Coke on regenerated catalyst [40]

In general, the deactivation in a given unit is largely a function of the unit’s

mechanical configuration, its operating condition, the type of fresh catalyst

used, and the feed quality. The primary criterion for adding fresh catalyst is to

arrive at an optimum equilibrium catalyst activity level. A too-high

equilibrium catalyst (E-cat) activity will increase delta coke on the catalyst,

resulting in a higher regenerator temperature. The higher regenerator

temperature reduces the catalyst circulation rate, which tends to offset the

activity increase. The amount of fresh catalyst added is usually a balance

between catalyst cost and desired activity [40, 59].


B-9

B-3.2. Process Variables

B-3.2.1. Independent variables: where directly regulated, usually with control devices

B-3.2.1.1. Feed Rate

As the flow rate of the feed decrease, the space velocity decrease and the

contact time increase, cause to increase conversion [60].

Conversion has been observed in some units to increase 1% absolute for a

3-5% relative decrease in fresh feed rate [61].

Charge rate is set as desired by the refiner based on the existing economics [31].

B-3.2.1.2. Feed Preheat Temperature

Decreasing the feed preheat temperature at constant riser outlet

temperature will [31, 62]:

§ Increase the catalyst circulation rate

§ Slightly increase conversion

§ Increase the coke production rate 0.1 – 0.15 %wt on feed for every

10oC reduction in preheat temperature

§ Decrease the delta Coke (kg coke/kg catalyst)

§ Decrease the Regenerator temperature

B-3.2.1.3. Reactor Pressure

According to UOP Operating Manual [31], the pressure in the reactor is

normally varied very little. There is a trade off here as a higher reactor

pressure would reduce the Gas Concentration Unit gas compressor

horsepower requirements, but it would also increase the main air blower

horsepower. Higher pressure would also reduce the required size of the

vessels.


B-10

The olefin content of the products will decrease with an increase in the

hydrocarbon partial pressure. Conversion will increase somewhat. Coke

laydown will increase slightly, an effect that may be offset by adding steam or

inert gas to reduce the hydrocarbon partial pressure. This may, however,

defeat the original purpose of raising the reactor pressure.

Reactor pressure normally varies slightly with changes in feed rate and

loading of the main column. The operator has some element of control, but

pressure must be kept within narrow limits around the design value to avoid

problems with riser and cyclone velocities. Normal first stage cyclone inlet

velocities are in the (20 m/s) range. Higher velocities are better for cyclone

efficiency, but worse with respect to the greater amount of catalyst carried up

to them as the vessel superficial velocity increases. Higher velocity also

increases cyclone erosion problems.

B-3.2.1.4. Reactor Temperature

Reactor temperature is the prime control of reactor severity. Increasing the

reactor temperature set point will signal the regenerated catalyst valve to

increase the hot catalyst flow as necessary to achieve the desired riser outlet

temperature. The increased catalyst circulation rate results in increased

conversion.

Higher temperature accelerates intermolecular motions, and assists the

transformation of the rate of chemical process. The effect of temperature on

the rate constant is described by the Arrhenius equation: = exp( − ) ( . 1)

The cracking rate constant is enhanced by the temperature of the riser and the

conversion of the feed into light products increases as a result [60].


B-11

Higher cracking temperature also increase gasoline octane and LPG olefinicity

-with their potential for alkylation feed- , cracked gasoline from high

temperature operation are particularly useful blending components for lead-

free gasoline [54, 61]. The higher octane is due to the higher rate of primary

cracking reactions relative to secondary hydrogen transfer reactions which

saturate olefins in the gasoline boiling range and lowers gasoline octane [61].

An increase in the unit conversion does not necessarily mean an increase in

gasoline yield, because Beyond a certain temperature, gasoline yield will be

negatively impacted [1, 42] (Table B – 3).

Table B – 3

Effect of operating Temperature of the reactor on the performance of a

fluidized bed cracking; Feed: Mid – continent gasoil [60]

Aver. Reactor temperature, oC 510 538 558 Feed conversion, vol. % 85 87 89 yields Gasoline, vol. % 71 68 64

Propylene, vol. % 7.8 10.6 11.5 Butenes, vol. % 7.4 8.5 10.6 Coke, wt % 3.6 3.4 3.5

Research octane rating of gasoline, clear 87 90 93

However, cracking at higher temperatures also produces more gas and coke

which are less desirable products. The optimum temperature is when just

enough coke is formed on the catalyst so that the heat of combustion in the

regenerator can provide all the heat required for reaction when the catalyst

returns to the riser-reactor [43].


B-12

B-3.2.2. Dependent Variables: are responsive to changes in the

independent variables

B-3.2.2.1. Catalyst – To – Oil Ratio

Catalyst to Oil ratio (C/O) is the ratio of kg/hr of catalyst circulated to

kg/hr of fresh feed. It is dependent variable, where it increases with an

increase in reactor temperatures, and decreases with higher regenerator or feed

preheat temperatures.

With UOP RxCat Technology, the riser Catalyst/Oil ratio (riser-C/O), where it

is the ratio of kg/hr of both regenerated and recirculated spent (carbonized)

catalyst to kg/hr of fresh feed. This (riser-C/O) ratio can be directly

manipulated by the amount of spent catalyst that is recirculated to the MxR

chamber without affecting the heat balance of the reactor/regenerator system.

The increases in spent catalyst recirculation to riser at a constant reactor

temperature will [31]:

§ Increase conversion.

§ Decrease light gas yield.

§ Increase C3 and C4 yields.

§ Increase gasoline aromatic content.

§ Increase gasoline yield.

§ Decrease LCO.

§ Increase regenerator temperature.

§ Decrease C/O rate.

B-3.2.2.2. Regenerator Temperature

The coke yield of a given cat cracker is essentially constant. However, a

more important term is delta coke. Delta coke controls the regenerator

temperature. Reducing delta coke will lower the regenerator temperature.


B-13

Many benefits are associated with a lower regenerator temperature. The

resulting higher cat/oil ratio improves product selectivity and/or provides the

flexibility to process heavier feeds. Many factors influence delta coke, these

are [40]:

§ Feedstock quality. The quality of the FCC feedstock impacts the

concentration of coke on the catalyst entering the regenerator. A

"heavier" feed containing a higher concentration of coker gas oil will

directionally increase the delta coke as compared with a "lighter," resid-

free feedstock.

§ Feed/catalyst injection. A well-designed injection system provides a

rapid and uniform vaporization of the liquid feed. This will lower delta

coke by minimizing non-catalytic coke deposition as well as reducing

the deposits of heavy material on the catalyst.

§ Riser design. A properly designed riser will help reduce delta coke by

reducing the back-mixing of already "coked-up" catalyst with fresh

feed. The back-mixing causes unwanted secondary reactions.

§ Cat/oil ratio. An increase in the cat/oil ratio reduces delta coke by

spreading out some coke-producing feed components over more catalyst

particles and, thus, lowering the concentration of coke on each particle.

§ Reactor temperature. An increase in the reactor temperature will also

reduce delta coke by favoring cracking reactions over hydrogen transfer

reactions. Hydrogen transfer reactions produce more coke than cracking

reactions.

§ Catalyst activity. An increase in catalyst activity will increase delta

coke. As catalyst activity increases so does the number of adjacent sites,

which increases the tendency for hydrogen transfer reactions to occur.

Hydrogen transfer reaction.

Appendix C Computer Programs

C-1

APPENDIX C

COMPUTER PROGRAMS

A Polymath Program for solution the FCC riser model is stated below. The

program may solve for any variable and flexible to accept different

operating conditions. Table C-1 lists some variables used in this program.

Excel worksheet is used as a powerful tool to solve the same equations

used in Polymath. Polymath solves the ODEs by Runge - kutta method

built inside its solver, while ODEs need to be introduced in details to Excel

and solved by iteration on cells using Runge - Kutta method. The main

interface simulator of an Excel worksheet is shown in Figure C-1.

Table C – 1

Variables used in Polymath program

variable Unit Description

Mug Pa.s Hydrocarbon vapor viscosity Phy - Catalyst deactivation function Rhog kg/m3 Density of gas phase Rhocat kg/m3 Density of catalyst particles Sph - Sphericity v - Voidage of gas phase w - Slip factor

The programmed Equations in Polymath are:

Independent variable: z Initial value: 0 Final value: 32

[1] d(y1)/d(z) = -J*(K1+K2+K3)*y1^2 Initial value: 1

[2] d(y2)/d(z) = J*((K1*y1^2-(K4+K5)*y2)) Initial value: 0

[3] d(y3)/d(z) = J*(K2*y1^2+K4*y2) Initial value: 0


C-2

[4] d(y4)/d(z) = J*(K3*y1^2+K5*y2) Initial value: 0

[5] d(T)/d(z) = -S*((K1*H1+K2*H2+K3*H3)*y1^2+

(K4*H4+K5*H5)*y2) Initial value: 834

[6] d(P)/d(z) = -RhoCat*9.81*(1-v) Initial value: 258000

[7] d(tc)/d(z) = (A*RhoCat*w)/(mcat*w+((1/Mwtg)*mg*(1-y4)*RhoCat*

0.082*101325*T/P)) Initial value: 0

[8] Mwtg = 1/(y1/MwtVGO+y2/MwtGLN+y3/MwtLGS+

((mds+mls)/mf)/18)

[9] Rhog = (P*Mwtg)/(0.082*T*101325)

[10] mg = mf*(y1+y2+y3)+mds+mls

[11] Uo = mg/(A*Rhog)

[12] Ar = (Rhog*(RhoCat-Rhog)*9.81*dp^3)/Mug^2

[13] Ret = Ar/(18+(2.3348-1.7439*Sph)*Ar^0.5)

[14] Ut = (Ret*Mug)/(Rhog*dp)

[15] w = 1+5.6*(9.81*D)^0.5/Uo+0.47*(Ut/(9.81*D)^0.5)^0.41

[16] v = (RhoCat*mg)/(Rhog*mcat*w+RhoCat*mg)

[17] Ug = Uo/v

[18] Up = Ug/w

[19] J = A*v*phy*Rhog/mg

[20] S = A*v*Rhog*phy/(mcat*Cpcat+mg*Cpg)

[21] K1 = Ko1*exp(-E1/(8.3*T))


C-3

[22] K2 = Ko2*exp(-E2/(8.3*T))

[23] K3 = Ko3*exp(-E3/(8.3*T))

[24] K4 = Ko4*exp(-E4/(8.3*T))

[25] K5 = Ko5*exp(-E5/(8.3*T))

[26] D = 1.37

[27] mf = 51.2

[28] mcat = mrcat+mxcat

[29] Mug = 1.179e-5

[30] RhoCat = 970

[31] dp = 60e-6

[32] Sph = 1

[33] MwtVGO = 360.04

[34] MwtGLN = 99

[35] MwtLGS = 46.5

[36] CpVGOgas = 3.56

[37] Cpg = CpVGOgas

[38] Cpcok = 1.675

[39] Cpcat = 2.3

[40] Ko1 = 12600

[41] Ko2 = 2070

[42] Ko3 = 25.1


C-4

[43] Ko4 = 2735

[44] Ko5 = 0.006

[45] E1 = 57359

[46] E2 = 52754

[47] E3 = 31820

[48] E4 = 65733

[49] E5 = 66570

[50] H1 = 393

[51] H2 = 795

[52] H3 = 1200

[53] H4 = 1150

[54] H5 = 151

[55] L = 32

[56] A = 1.474

[57] phy = -45*(y4*mf+mxcok)/mcat+1

[58] x = (1-y1)*100

[59] mds = 0.01*mf

[60] mls = 0.035*mf

[61] mrcat = 208

[62] mxcat = 104

[63] mcok = mrcok+mxcok


C-5

[64] mrcok = 0.013728

[65] mxcok = 1.28

[66] Trcat = 980

[67] Tout = 805

[68] Cps = 2.4

[69] Tls = 523

[70] Tmix1 = ((mrcat*Cpcat+mrcok*Cpcok)*Trcat+(mxcat*Cpcat+

mxcok*Cpcok)*Tout-mls*Cps*Tls)/ (mrcat*Cpcat+

mrcok*Cpcok+mxcat*Cpcat+mxcok*Cpcok+mls*Cps)

[71] A1 = -a1+ (a2+a3*SG)*Kuop

[72] A2 = A1+ (a4-a5*Kuop)/SG

[73] A3 = A2/1000

[74] A4 = (1+a6*Kuop)*(a7-a8/SG)/1000000

[75] A5 = -(1+a9*Kuop)*(a10-a11/SG)/1000000000

[76] HL = A3*(1.8*Tf-259.67)+A4*((1.8*Tf)^2-259.67^2)+

A5*((1.8*Tf)^3-259.67^3)

[77] F1 = f1*SG+f2*Kuop*(Kuop-f3)

[78] F2 = f4*SG+f5*Kuop*(Kuop-f6)

[79] F3 = f7*SG-f8*Kuop*(Kuop-f9)-f10

[80] a = 2.32*3.24*mf*F1

[81] b = mcat*Cpcat+mcok*Cpcok+mds*Cps+mls*Cps-2.32*1654.812*

mf*F1-2.32*1.8*mf*F2


C-6

[82] c = 2.32*(459.67)^2*mf*F1+2.32*459.67*mf*F2+2.32*mf*F3-

2.32*HL*mf-mds*Cps*Tds-(mcat*Cpcat+mcok*Cpcok+mls*

Cps)*Tmix1

[83] Tmix2 = (-b+ (b^2-4*a*c) ^0.5)/ (2*a)

[84] SG = 0.8841

[85] Kuop = 12.19

[86] Tf = 547

[87] Tds = 523

[88] t=z/Ug

POLYMATH Results POLYMATH Report 02-10-2012 , Rev5.1.233 Calculated values of the DEQ variables Variable initial value minimal value maximal value final value z 0 0 32 32 y1 1 0.1747862 1 0.1747862 y2 0 0 0.5003684 0.5003646 y3 0 0 0.2750897 0.2750897 y4 0 0 0.0497595 0.0497595 T 834 802.89656 834 802.89656 P 2.44E+05 2.287E+05 2.44E+05 2.287E+05 tc 0 0 7.6269989 7.6269989 MwtLGS 46.5 46.5 46.5 46.5 MwtGLN 99 99 99 99 mf 51.2 51.2 51.2 51.2 A 1.474 1.474 1.474 1.474 dp 6.0E-05 6.0E-05 6.0E-05 6.0E-05 Sph 1 1 1 1 mls 1.792 1.792 1.792 1.792 MwtVGO 360.04 360.04 360.04 360.04 mds 0.512 0.512 0.512 0.512 Mwtg 189.48476 71.656029 189.48476 71.656029 mg 53.504 50.956315 53.504 50.956315 mxcat 104 104 104 104 mrcat 208 208 208 208 E1 5.736E+04 5.736E+04 5.736E+04 5.736E+04 E2 5.275E+04 5.275E+04 5.275E+04 5.275E+04 E3 3.182E+04 3.182E+04 3.182E+04 3.182E+04 E4 6.573E+04 6.573E+04 6.573E+04 6.573E+04 E5 6.657E+04 6.657E+04 6.657E+04 6.657E+04 D 1.37 1.37 1.37 1.37 Rhog 6.6721777 2.4570839 6.6721777 2.4570839 mcat 312 312 312 312


C-7

Mug 1.179E-05 1.179E-05 1.179E-05 1.179E-05 RhoCat 970 970 970 970 Ar 97.97987 36.239765 97.97987 36.239765 Ret 4.108341 1.681099 4.108341 1.681099 Ut 0.1209933 0.1209933 0.1344423 0.1344423 Uo 5.4402789 5.4402789 14.069561 14.069561 w 4.8897177 2.5803509 4.8897177 2.5803509 CpVGOgas 3.56 3.56 3.56 3.56 Cpg 3.56 3.56 3.56 3.56 Cpcok 1.675 1.675 1.675 1.675 Cpcat 2.3 2.3 2.3 2.3 Ko1 1.26E+04 1.26E+04 1.26E+04 1.26E+04 Ko2 2070 2070 2070 2070 Ko3 25.1 25.1 25.1 25.1 Ko4 2735 2735 2735 2735 Ko5 0.006 0.006 0.006 0.006 K1 3.1747015 2.3030001 3.1747015 2.3030001 K2 1.0144193 0.7550937 1.0144193 0.7550937 K3 0.2531086 0.2118213 0.2531086 0.2118213 K4 0.2055465 0.1422815 0.2055465 0.1422815 K5 3.996E-07 2.753E-07 3.996E-07 2.753E-07 H1 393 393 393 393 H2 795 795 795 795 H3 1200 1200 1200 1200 H4 1150 1150 1150 1150 H5 151 151 151 151 L 32 32 32 32 v 0.8360283 0.8360283 0.9615194 0.9615194 mxcok 1.28 1.28 1.28 1.28 x 0 0 82.521382 82.521382 Ug 6.5072905 6.5072905 14.632633 14.632633 Up 1.3308111 1.3308111 5.6707921 5.6707921 phy 0.8153846 0.4479301 0.8153846 0.4479301 J 0.1253032 0.0306117 0.1253032 0.0306117 S 0.0073829 0.0017351 0.0073829 0.0017351 mrcok 0.013728 0.013728 0.013728 0.013728 mcok 1.293728 1.293728 1.293728 1.293728 Trcat 980 980 980 980 Tout 805 805 805 805 Cps 2.4 2.4 2.4 2.4 Tls 523 523 523 523 Tmix1 912.74207 912.74207 912.74207 912.74207 Kuop 12.19 12.19 12.19 12.19 SG 0.8841 0.8841 0.8841 0.8841 A1 -613.66162 -613.66162 -613.66162 -613.66162 A2 45.266496 45.266496 45.266496 45.266496 A3 0.0452665 0.0452665 0.0452665 0.0452665 Tf 547 547 547 547 A4 4.471E-04 4.471E-04 4.471E-04 4.471E-04 A5 -7.737E-08 -7.737E-08 -7.737E-08 -7.737E-08 F1 1.856E-04 1.856E-04 1.856E-04 1.856E-04 a 0.0714201 0.0714201 0.0714201 0.0714201 F2 -0.443224 -0.443224 -0.443224 -0.443224 Tds 523 523 523 523 HL 363.65147 363.65147 363.65147 363.65147 b 783.58545 783.58545 783.58545 783.58545 F3 177.5477 177.5477 177.5477 177.5477 c -7.032E+05 -7.032E+05 -7.032E+05 -7.032E+05 Tmix2 833.99118 833.99118 833.99118 833.99118 t 0 0 2.1868928 2.1868928

FCCU SIMULATOR : Base case 30000 BPD FCCU

Notes : Product Kg/s wt% feedInput Value color LGS 14.08437 27.50854

GLN 25.61864 50.03641Riser diam. = 1.37 m TGO 8.949219 17.47894Riser Height = 32 m

WHSV 77.44749 hr-1

Tout 802.888 KPout 228747.2 Paselectivity 1.540309

Product Kg/s wt% feed Conv. 82.5 wt%COK 2.547766 4.976105 efficiency 60.63472 wt%∆ Coke 1.218287 wt%

80459.84 KJ/Kg coke

mds 0.512 Kg/sC/O 6.09375 Tds 523 KTmix2 833.9912 KPin 244000 Pa. mf 51.2 Kg/s

Tf 547 KAPI 28.55

mcat 312 Kg/s Kuop 12.19mcok 1.293728 Kg/s Mwt 360 Kg/KmoleTmix1 912.7421 K

mrcat 208 Kg/smrcok 0.013728 Kg/sTrcat 980 K

mls 1.792 Kg/sTls 523 K

mxcat 104 Kg/smxcok 1.28 Kg/sTout 805 K

Proportional value colorCalculated Value color

MxR

Rise

r

∆Hcomb =

Regenerator

Figure C-1

Appendix D Glossary

E - 1

APPENDIX D

GLOSSARY OF TERMS USED IN THIS WORK

Activity: is a measure of how fast one or more reactions proceed in

presence of the catalyst. High catalyst activity means fast reactions rates

and short contact time required (1 – 5 second) to achieve high conversion.

API Gravity: is an "artificial" scale of liquid gravity defined by: (141.5/SG)

- 131.5. The scale was developed for water = 10. The main advantage of

using °API gravity is that it magnifies small changes in liquid density.

Back-Mixing: is the phenomena by which the catalyst travels more slowly

up the riser than the hydrocarbon vapors.

Carbonium Ion: is a positively charged (CH5+) ion which is formed by

adding a hydrogen ion (H+) to paraffin.

Catalyst/oil ratio (C/O): The weight of circulating catalyst fed to the

reactor of a fluid-bed catalytic cracking unit divided by the weight of

hydrocarbons charged during the same interval.

Catalyst Cooler: is a heat exchanger that removes heat from the regenerator

through steam generation.

Catalyst Stability: resistance to time – dependent deactivation by coke.

Conradson carbon: A test used to determine the amount of carbon residue

left after the evaporation and pyrolysis of oil under specified conditions.

Expressed as weight percent; ASTM D-189.

Appendix D Glossary

E - 2

Conversion: is often defined as the percentage of fresh feed cracked to

gasoline, lighter products, and coke. Raw conversion is calculated by

subtracting the volume or weight percent of the FCC products (based on

fresh feed) heavier than gasoline from 100, or:

Conversion = 100 – (LCO + HCO + DO) vol% or wt%

Dense Phase: is the region where the bulk of the fluidized catalyst is

maintained.

Dilute Phase: is the region above the dense phase which has a substantially

lower catalyst concentration.

Decanted Oil, Slurry, Clarified Oil, or Bottoms: Are the heaviest and often

the lowest priced liquid product from a cat cracker.

Delta Coke: is the difference between the coke content of the spent catalyst

and the coke content of the regenerated catalyst.

Efficiency = (%Gasoline)/ (% Conversion) x 100

HCO: Heavy FCC cycle gas oil product.

LCO: Light FCC cycle gas oil product.

LHSV: Liquid hour space velocity; volume of feed per hour per volume of

catalyst.

Refractory: is a cement-like material used to stand abrasion and erosion.

Riser: is a vertical "pipe" where virtually all FCC reactions take place.

Appendix D Glossary

E - 3

Selectivity: the ratio of the yield of desirable product (gasoline) to the yield

of undesirable products (coke and gas).

Severity: The degree of intensity of the operating conditions of a process

unit. Severity may be indicated by percent disappearance of the feed, or

operating conditions alone.

Slip Factor: is the ratio of catalyst residence time to hydrocarbon vapors

residence time in the riser.

Space velocity: The volume (or weight) of gas and/or liquid passing

through a given catalyst or reactor space per unit time, divided by the

volume (or weight) of catalyst through which the fluid passes. High space

velocities correspond to short reaction times. See LHSV and WHSV.

WHSV: Weight hour space velocity; weight of feed per hour per weight of

catalyst.

REFERENCES

R-1

REFERENCES

[1] Meyers R. A., "Handbook of Petroleum Refining Processes", Third

edition, McGraw-Hill (2004).

[2] Wikipedia, the free encyclopedia, "Fluid Catalytic Cracking",

http://www.en.wikipedia.org, (2011)

[3] Ali H., Rohani S. and Corriou J. P., “Modeling and Control of a Riser

Type Fluid Catalytic Cracking (FCC) Unit”, IChemE, Vol. 75, Part A,

May (1997).

[4] www.UOP.com

[5] COLORADO school of mines, "Fluidized Catalytic Cracking",

http://ebookdock.com/fluidized-catalytic-cracking.html

[6] Ronald F., Colwell P.E., http://www.processengr.com/

ppt_presentations/Msat_regulation_options.pdf , “Benzene in Gasoline

Regulations & Remedies” Process Engineering Associates, LLC, (2009)

[7] Fernandes J. L., Pinheiro C. I. C., Oliveira N. and Ramoa Ribeiro F.,

“Modeling and Simulation of an Operating Industrial Fluidized

Catalytic Cracking (FCC) Riser”, ENPROMER.

[8] Madhusudana R. Rao, Rengaswamy R., Suresh A. K. and Balaraman K.

S., “Industrial Experience with Object – Oriented Modeling, FCC Case

Study”, IChemE, Part A, April (2004).

[9] Bollas G. M., Lappas A. A., Iatridis D. K. and Vasalos I. A., “Five-

Lump Kinetic model with selective catalyst deactivation for the

prediction of the product selectivity in the fluid catalytic cracking

process”, Elsevier, Catalyst Today 127(2007) 31-43.

http://www.en.wikipedia.org

http://www.UOP.com

http://ebookdock.com/fluidized-catalytic-cracking.html

http://www.processengr.com/

R-2

[10] Gupta R. K., Kumar V. and Srivastava V. K., “Modeling of Fluid

Catalytic Cracking Riser Reactor: A Review”, International Journal of

chemical engineering, Vol. 8, (2010).

[11] Arbel A., Huang Z., Rinard I. H., and Shinnar R., “Dynamic and

Control of Fluidized Catalytic Crackers. 1. Modeling of the Current

Generation of FCC’s”, Ind. Eng. Chem. Res. 34(1995) 1228-1243.

[12] Han I. S., and Chung C. B., “Dynamic Modeling and Simulation of a

Fluidized Catalytic Cracking Process. Part II: Property estimation and

simulation”, Chem. Eng. Sci., Vol. 56, (2001).

[13] Theologos K.N., and Markatos N.C., “Advanced modeling of fluid

catalytic cracking riser-type reactors”, AIChE Journal, 39 (1993) 1007-

1017.

[14] Theologos K.N., Nikou I.D., Lygeros A.I., and Markatos N.C.,

“Simulation and design of fluid catalytic cracking riser-type reactors”,

AIChE Journal, 43 (1997) 486-493.

[15] Bollas G. M., Lappas A. A., and Vasalos I. A., “An Integrated Riser-

Reactor Dynamic Model for the Simulation of Pilot and Commercial

FCC units”, (2002).

[16] Erthal R. H., Negrao C.O. R., and Rossi L. F. S., “Modeling the Riser of

a Fluid Catalytic Cracking Unit”, 17th international congress of

mechanical eng., 10-14, Nov., (2003).

[17] Souzaa J. A., Vargasa J.V.C., Von Meiena O. F. and Martignonib W.,

“Numerical simulation of FCC risers”, Engenharia Termica, 4(2003)

17-21.

R-3

[18] Das A. K., Baudrez E., Marin G. B. and Heynderickx G. J., “Three-

dimensional simulation of a fluid catalytic cracking riser reactor”, Ind.

Eng. Chem. Res., 42 (2003) 2602-2617.

[19] Berry T. A., McKeen T. R., Pugsley T. S. and Dalai A. K., “Two-

dimensional reaction engineering model of the riser section of a fluid

catalytic cracking unit”. Ind. Eng. Chem. Res., 43 (2004) 5571-5578.

[20] Hassan Sh. Khazaal, “Computer Aided Design of a Fluidized Catalytic

Cracking Unit using Iraqi feed stocks”, MSC thesis, University of

Technology, (2005).

[21] Xu O., Su H., Mu S. and Chu J., “7-lump kinetic model for residual oil

catalytic cracking”, J. Zhejiang Univ. Sci. A., 7(2006) 1932-1941.

[22] Krishnaiah D., Gopikrishna V., Bono A. and Sarbatly R., “Steady State

Simulation of a Fluid Catalytic Cracking Unit”, Journal of Applied

Science, 7 (15) (2007) 2137-2145.

[23] Souza J. A., Vargas J. V. C., Von Meien O. F., and Martignoni W. P.,

“Modeling and Simulation of Industrial FCC Risers”, Thermal

Engineering, Vol. 6, No.1, June (2007).

[24] Gupta R.K., Kumar V., and Srivastava V.K, “A new generic approach

for the modeling of fluid catalytic cracking (FCC) riser reactor”,

Chemical Engineering Science, 62 (2007) 4510-4528.

[25] Ahari J. S., Farshi A. and Forsat K., “ Mathematical Modeling of the

Riser Reactor in Industrial FCC Unit”, Petroleum & Coal, 50 (2)

(2008), 15-24,

R-4

[26] Heydari M., Ebrahim H. A., and Dabir B., “Modeling of an Industrial

Riser in the Fluid Catalytic Cracking Unit”, American Journal of

Applied Sciences 7(2) (2010) 221-226.

[27] Shakoor Z. M., “Estimation of Kinetic Parameters of Complex

Reactions by MATLAB Software”, University of Technology, Iraq,

(2010).

[28] Baudrez E., Heynderickx G. J., Marin G. B. “Steady-state simulation of

fluid catalytic cracking riser reactors using a decoupled solution

method with feedback of the cracking reactions on the flow”, Chemical

Engineering Research and Design, 88 (2010) 290-303.

[29] Osman A. M., Shah Maulud A., Ramasamy M., and Mahadzir S.,

“Steady State Modeling and Simulation of the Riser in an Industrial

RFCC Unit”, Journal of Applied Sciences, 10 (2010) 3207-3214.

[30] AL-Niami S. A., Mehdi F. A., “Decoupling control of a FCC unit using

fuzzy logic”, first scientific conference for petroleum refining and gas,

university of technology, Iraq, (2011).

[31] L. M. Wolschlag and K. A. Couch, UOP LCC, “Upgrade FCC

Performance”, hydrocarbon processing, Sep. (2010) 57-65.

[32] UOP FLUID CATALYTIC CRACKING (FCC) AND RELATED

PROCESSES,www.UOP.com.

[33] Fahim M.A., Al-Sahhaf T.A., Elkilani A.S., " Fundamentals of

Petroleum Refining" Elsevier (2010)

[34] Ancheyta J., “Modeling And Simulation Of Catalytic Reactors For

Petroleum Refining”, Wiley Publication, (2011)

http://www.UOP.com

R-5

[35] Svoboda K., Kalisz S., Miccio F., Wieczorek K., and Pohorely M.,

“Simplified Modeling of Circulating Flow of Solids between a Fluidized

bed and a vertical Pneumatic transport tube reactor connected by

orifices”, Elsevier, Powder Technology, 192(2009) 65-73.

[36] Coulson J. M., and Richardson’s J. F., “Chemical Engineering”, Vol. 2,

Fifth edition, Butterworth-Heinemann (2002).

[37] Rabinovich E., and Kalman H., “Flow regime diagram for vertical

pneumatic conveying and fluidized bed systems”, Elsevier, Powder

Technology, 207 (2011) 119-133.

[38] J. M. Llanes and M. Miranda, CEPSA, La Rábida, Spain and S.

Mullick, AspenTech, Houston, Texas, “Use modeling to fine-tune

cracking operations”, hydrocarbon processing, Sep. (2008) 123-132.

[39] Hernandez-Barajas J. R., Vazquez-Roman R. and Felix-Flores Ma. G.,

”A comprehensive estimation of kinetic parameters in lumped catalytic

cracking reaction models”, Elsevier, Fuel 88 (2009) 169-178.

[40] Sadeghbeigi R., "Fluid Catalytic Cracking Handbook", Gulf publishing

(2000).

[41] Gauthier T. A., “Current R&D Challenges for Fluidized Bed Processes

in the Refining Industry”, International Journal of chemical reactor

engineering, Vol. 7, (2009).

[42] Gary J. H., Handwerk G. E., "Petroleum Refining Technology and

Economics", Forth Edition, Marcel Dekker, Inc. (2001).

R-6

[43] Report S., "Cost Estimates of Small Modular Systems", National

Renewable Energy Laboratory (2006), Available electronically at

http://www.osti.gov/bridge.

[44] Hydrocarbon Processing Journal, “Refining Processes Handbook”,

(2011).

[45] Kellogg Brown and Root LLC (KBR) Catalog, “Fluid Catalytic

Cracking”, (2007).

[46] www.KBR.com

[47] www.Lummus.CBI.com

[48] Oil & Gas Journal, “Fluid Catalytic Cracking Poster”, 17, Mar., (2003)

[49] Lawler D., Letzsch W., Dhaidan F., “Advances in FCCU Technology

for the Production of Olefins”, Stone & Webster Engineering

Corporation.

[50] Nieskens M., Shell presentation, “The Fluidized Catalytic Cracking

Process”, 4, Oct., (2007)

[51] www.shell.com

[52] Jay Ross, Stephane Wambergue and Jean-Paul Lepage, “A New

Separator Helps FCC Adapt to a New Refinery-Petrochemical Role”,

Axens.

[53] ENSPM formation industrie – IFP training – "Refining Processes and

Petroleum Products" (1998).

[54] Venuto P. B., Habib E. T., "Fluid Catalytic Cracking with Zeolite

Catalysts", Marcel Dekker Inc., (1979).

http://www.osti.gov/bridge

http://www.KBR.com

http://www.Lummus.CBI.com

http://www.shell.com

R-7

[55] http://www1.eere.energy.gov/industry/petroleum_refining/

bandwidth.html, “Energy Bandwidth for Petroleum Refining Processes”

by Industrial Technologies Program (ITP), (2006)

[56] Parkash S., "Refining Processes Handbook", Elsevier (2003).

[57] Nieskens M. - Shell presentation (US), Houston “MILOS – Shell’s

ultimate flexible FCC technology in delivering diesel/propylene”,

March, (2008).

[58] Murphree E.V., Gohr E.J., and Brown C.L., "Fluid Catalytic Cracking",

the Science of Petroleum, Vol. V, Part II, (1953).

[59] Gupta R. K., “Modeling and Simulation of Fluid Catalytic Cracking

Unit", Ph.D Thesis, Thapar University, India, (2006).

[60] Decroocq D., “Catalytic Cracking of Heavy Petroleum Fractions”, Gulf

Publishing, (1984).

[61] http://thefccnetwork.com, technical papers/ FCC Catalysts and

Additives/ “Alternative Routs to High Conversion” (2010)

[62] Hobson G.D., “Modern Petroleum Technology”, Institute of Petroleum,

London, (1984).

[63] Sertic Bionda K., Gomzi Z., and Fabulic Ruszkowski M., “Kinetics of

Gas Catalytic Cracking”, Akademiai Kiado Budapest, Vol. 88, No. 1

(2006) 111-118.

[64] Bollas G. M., and Vasalos I.A., “Modeling Small-Diameter FCC Riser

Reactors. A Hydrodynamic and Kinetic Approach”, Ind. Eng. Chem.

Res., 41(2002) 5410-5419.

http://www1.eere.energy.gov/industry/petroleum_refining/

http://thefccnetwork.com

modeling and simulation of fcc risers · figure 1.1 fluid catalytic cracking unit 2 figure 3.1...

Documents