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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
I
II
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]
Chapter One Introduction
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
Chapter One Introduction
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].
Chapter One Introduction
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.
Chapter Two Literature Survey
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.
Chapter Two Literature Survey
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).
Chapter Two Literature Survey
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.
Chapter Two Literature Survey
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).
Chapter Three Mathematical Modeling
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.
Chapter Three Mathematical Modeling
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
Chapter Three Mathematical Modeling
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)
Chapter Three Mathematical Modeling
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
Chapter Three Mathematical Modeling
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.
Chapter Three Mathematical Modeling
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
Chapter Three Mathematical Modeling
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)
Chapter Three Mathematical Modeling
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:
Chapter Three Mathematical Modeling
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:
Chapter Three Mathematical Modeling
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 32.9 106.5. 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)
Chapter Three Mathematical Modeling
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)
Chapter Three Mathematical Modeling
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)
Chapter Three Mathematical Modeling
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.
Chapter Three Mathematical Modeling
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.
Chapter Three Mathematical Modeling
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].
Chapter Four Results and discussion
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.
Chapter Four Results and discussion
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
Chapter Four Results and discussion
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
Chapter Four Results and discussion
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
Chapter Four Results and discussion
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
Chapter Four Results and discussion
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
Chapter Four Results and discussion
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
Chapter Four Results and discussion
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
Chapter Four Results and discussion
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
Chapter Four Results and discussion
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
Appendix A FCC technologies
A-2
Figure A-1a Stacked Reactor regenerator configuration [43]
Figure A-1b Side by side Reactor regenerator configuration [43]
Appendix A FCC technologies
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.
Appendix A FCC technologies
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.
Appendix A FCC technologies
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].
Appendix A FCC technologies
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
Appendix A FCC technologies
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]
Appendix A FCC technologies
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
Appendix A FCC technologies
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.
Appendix A FCC technologies
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
Appendix A FCC technologies
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
Appendix A FCC technologies
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
advantages are tabulated in Table A-3.
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).
Appendix A FCC technologies
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
advantages are tabulated in Table A-4.
Appendix A FCC technologies
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
Appendix A FCC technologies
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.
Appendix A FCC technologies
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.
Appendix A FCC technologies
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.
Appendix B Variables of FCC units
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 proportion