UNIVERSITY OF CALGARY
Thermal and Catalytic Steam Reactivity Evaluation of Athabasca Vacuum Gasoil
By
Gustavo Luis Trujillo Ferrer
A THESIS
SUBMITTED TO THE FACULTY OF GRADUATE STUDIES
IN PARTIAL FULFILMENT OF THE REQUIREMENTS FOR THE
DEGREE OF MASTER OF SCIENCE
DEPARTMENT OF CHEMICAL AND PETROLEUM ENGINEERING
CALGARY, ALBERTA
APRIL 2008
© Copyright by Gustavo Luis Trujillo Ferrer 2008
ISBN: 978-0-494-44637-9
Abstract The use of water as source of hydrogen for heavy oil upgrading has long been
considered. Aquaconversion is a Selective Catalytic Steam Cracking (SCSC)
process developed in the 90’s and had shown promising results when applied for
the upgrading of Venezuelan heavy oils. This work is aimed at applying
aquaconversion for the upgrading of Northern Alberta heavy oils. To that end a
pilot plant was setup and tested to include all important features of said process.
Athabasca Vacuum Gasoil (AVGO) was selected as heavy feedstock for testing.
First, thermal cracking was studied and used to ensure repeatability and
reproducibility of the pilot plant performance. Secondly, steam thermal cracking
was undertaken using the same conditions of SCSC, however without
incorporating the catalyst. Finally, different catalysts were incorporated and the
results were compared with thermal and steam thermal cracking results. The
comparison was based on yield and conversion levels, gas composition and
activation energies of the cracking reactions. No significant difference in the
yield and conversion were attained when steam was introduced to the thermal
cracking process. Notably SCSC increased the conversion and the liquid yield
and reduced the tendency to form coke and gases. The activation energy of
SCSC was approximately 6 times lower than steam thermal cracking, which in
turn, was similar to that found for thermal cracking.
Acknowledgements
I would like to take this opportunity to express my special and deep thanks
to my supervisor Dr. Pedro Rafael Pereira Almao for the opportunity of being part
of his excellent team and his support and guidance. I feel very honoured of
making research with you and very privileged for all what I have learned working
with you.
I wish to thank the staff of the Chemical and Petroleum Engineering
department, and also to the institutions for their financial support: Schulich
School of Engineering at the University of Calgary, the Alberta Ingenuity Centre
for in situ Energy (AICISE), the National Science and Engineering Research of
Canada (NSERC), and the Alberta Energy Research Institute (AERI).
I also want to thanks to Cristina Wong and especially to Amely Pereira for
your dedication and support on the design, construction and planning the starting
up of the pilot plant. To Mike Grigg and Pat Walsh for their help in the
development of the automatic control system and the electrical work. To Lante
Carbogniani, Lina Marcela, Linda Catalina, Dr. Azfar Hassan and Luis Pineda for
all their support and effort displayed in the solution of the analytical issues. To
Dr. Francisco Lopez and Dr. Carlos Scott for their support in the catalyst
preparation. And my very special thanks to Redescal Gomez for all his support
and his huge contribution in the accomplishment of this work, our team work was
unique and absolutely remarkable.
Thanks to all the people from the Catalysis for Upgrading and Hydrogen
Production group: Carmen, Enzo, Maria Laura, Herbert, John, Eumir, Luis
Gerardo, Latif, Mazin, Behdad and Alejandro, for the wonderful and extraordinary
working environment. I feel really fortunate to closely work with all of you. Thanks
to Michael Wheeler for his patience on helping me to improve my English. Also
very thanks to Clementina, my sister in Canada, for to have convinced me to
come to this country looking for a new horizon and staying working together to
reach the goals that we had proposed to accomplish here, yes Clemen, we did it!
To my dear mom Cirila, to mom in law Maria Griselda, to my supporting
mom Renata and to Mr. Onofre, they helped me to do one part of the work that
was hard to cover for me.
And last but not least, my very deep thanks to my beloved family for giving
me the inspiration and motivation to take this challenge, first of all to my wife
Eleonora for believing on me unconditionally and giving me all the support to
keep going in this new objective in this new country and sharing with me the bad
and the good, to my sweet daughter Marianna, she always surprising me and
who supported me and helped me to complete this work more than I would have
ever imagined, and my little son Gustavito, my little man, the backwater of peace
at the moments of anguish, he always giving me his sweetness and the illusion
for finishing what I had started, yes my boy, now we can play. I wouldn't have
finished this walk if I would not have had them with me.
Dedication
To my wife Eleonora, my daughter Marianna and my little son Gustavito, the
most important reasons of my life.
To Cirila and Luis, my mom and my dad, to whom I owe what I am.
To Narkys, Memo, Niorko, Nano and Nuby, my brothers and sisters, my eternal
gang.
To the big one who has been always with me.
To Canada.
Table of Contents
Abstract .................................................................................................................ii Acknowledgements .............................................................................................. iii Dedication ............................................................................................................ v Table of Contents .................................................................................................vi List of Tables ........................................................................................................ix List of Figures ....................................................................................................... x List of Symbols and Abbreviations.......................................................................xii
Chapter 1: Introduction ......................................................................................... 1 1.1 Background................................................................................................. 1 1.2 Motivation ................................................................................................... 5 1.3 Objectives ................................................................................................... 6
Chapter 2: Literature Review................................................................................ 8 2.1 Thermal Cracking........................................................................................ 9
2.1.1 Chemistry of the Thermal Cracking................................................... 13 2.1.2 Severity and Conversion in Thermal Cracking Processes................. 17
2.2 Visbreaking ............................................................................................... 18 2.2.1 Definition and Fundamentals of Visbreaking ..................................... 18 2.2.2 Visbreaking: State of the Art Technology. ......................................... 19
2.2.2.1 ABB Lummus Giobal/Shell Global Solutions ............................ 20 2.2.2.2 AXENS..................................................................................... 24 2.2.2.3 Foster WheeIer/UOP................................................................ 27 2.2.2.4 Kellogg Brown & Root (KBR) ................................................... 29 2.2.2.5 Lurgi Öl Gas Chemie................................................................ 30
2.3 Catalytic Steam Cracking.......................................................................... 32 2.3.1 Thermal Catalytic Steam Reforming ................................................. 34 2.3.2 Unsupported Ultradispersed Catalysis .............................................. 35 2.3.3 Selective Catalytic Steam Cracking .................................................. 36
2.3.3.1 Reaction Mechanism................................................................ 37 2.3.3.2 UOP / PDVSA-Intevep / Foster Wheeler: Aquaconversion ...... 38
2.4 The Thermodynamic Equilibrium of Water/Heavy Hydrocarbons Reactions................................................................................................. 40
2.5 Remarks ................................................................................................... 42
Chapter 3: Experimental ..................................................................................... 43 3.1 Experimental Setup................................................................................... 43
3.1.1 Pilot Plants: Basic Considerations..................................................... 43 3.1.2 Process Overview ............................................................................. 46 3.1.3 Feed and Pre-heater Section ............................................................ 49 3.1.4 Reactor.............................................................................................. 51 3.1.5 Products Collection ........................................................................... 54 3.1.6 Process Line ..................................................................................... 56 3.1.7 Control and Data Acquisition............................................................. 57
3.2 Reactivity Tests......................................................................................... 59 3.2.1 Experimental Plan ............................................................................. 59 3.2.2 Operational Criteria ........................................................................... 60 3.2.3 Feed Characterization....................................................................... 61 3.2.4 Preparation of the Catalyst................................................................ 62 3.2.5 Products Analysis.............................................................................. 63
3.2.5.1 Gases....................................................................................... 63 3.2.5.2 Liquid Product .......................................................................... 63
Chapter 4: Results and Discussion..................................................................... 67 4.1 Bench Scale Pilot Plant: Performance Evaluation..................................... 67 4.2 Reactivity Tests......................................................................................... 73
4.2.1 Thermal Processing Evaluation......................................................... 73 4.2.2 Steam Processing Evaluation ........................................................... 81 4.2.3 Selective Catalytic Steam Processing Evaluation ............................. 87
4.2.3.1 Ultradispersed Catalyst Test .................................................... 87 4.2.3.2 Fixed Bed Tests ....................................................................... 90
4.3 Studied Processes Comparison................................................................ 93 4.3.1 Activation Energy .............................................................................. 95
4.4 Global Conversion Map for Athabasca VGO Thermal and Steam Catalytic Cracking.................................................................................... 96
Chapter 5: Conclusions and Recommendations............................................... 100 5.1 Conclusions ............................................................................................ 100 5.2 Recommendations .................................................................................. 101
References ....................................................................................................... 103
Appendix A: Reactors Sizing Estimation .......................................................... 110
Appendix B: Catalytic Steam Cracking Pilot Plant ............................................ 111
Appendix C: Tags Assigned for Points of Temperature and Pressure.............. 112
Appendix D: Algorithm for the Design of Vapors Condenser............................ 114
Appendix E: Catalysts Preparation Procedure.................................................. 119
Appendix F: Certified Gas composition and Calibration Result for Gas Chromatographer ..................................................................................... 124
Appendix G: Operational Data and Experimental Results for Thermal Cracking Runs.......................................................................................... 125
Appendix H: Operational data and Experimental Results for Steam Cracking Runs ......................................................................................................... 127
Appendix I: Operational Data and Experimental Results for Catalytic Runs..... 129
Appendix J: Physical Properties of Athabasca Vacuum Gasoil ........................ 132
List of Tables
Table 2-1 Bond dissociation energies[12,13] ....................................................... 15
Table 3-1 Comparison of commercial and typical pilot-scale operations .......... 45
Table 3-2 Design specifications for the Catalytic Steam Cracking unit............. 48
Table 3-3 Length and volume specifications of flooded process line sections.. 57
Table 3-4 Operational periods for each run performed on the pilot plant unit ... 61
Table 4-1 Operational conditions and liquid yield results for SET1 runs........... 68
Table 4-2 SimDist results for SET1-A and SET1-B .......................................... 69
Table 4-3 Operational conditions and variable results for SET2 runs............... 70
Table 4-4 SimDist results for SET2-A and SET2-B .......................................... 72
Table 4-5 Summary of experimental conditions evaluated and result obtained for Thermal Cracking evaluation. .................................................. 74
Table 4-6 Summary of experimental conditions evaluated and results obtained for Steam Cracking evaluation...................................................... 82
Table 4-7 Summary of experimental conditions evaluated and results obtained for Selective Catalytic Steam Cracking evaluation using ultradispersed catalyst................................................................................. 88
Table 4-8 Summary of experimental conditions evaluated and results obtained for Selective Catalytic Steam Cracking evaluation using a fixed bed reactor. ................................................................................................. 92
Table 4-9 Summary results of Athabasca Vacuum Gasoil for different processes evaluated.................................................................................... 94
Table 4-10 Activation energy results ................................................................ 95
List of Figures
Figure 1-1 Primary energy consumption estimation, 2005-20030 (quadrillion BTU)[1] ......................................................................................... 1
Figure 1-2 Growth in oil demand 2005-2007[2]..................................................... 2
Figure 1-3 Alberta’s conventional crude oil and crude bitumen reserves[2].......... 3
Figure 2-1 Reaction scheme for Thermal Cracking reactions[12,13] .................... 14
Figure 2-2 Shell Soaker Visbreaking Process (SSVB) flow diagram[46] ............. 20
Figure 2-3 Shell Deep Thermal Conversion process (SDTC) flow diagram[46]... 22
Figure 2-4 Shell Thermal Gasoil process (STGP) flow diagram[46] .................... 24
Figure 2-5 The Tervahl T process[49] ................................................................. 25
Figure 2-6 The Tervahl H process[49] ................................................................. 26
Figure 2-7 The Tervahl C process[47] ................................................................. 27
Figure 2-8 Foster Wheeler/UOP Visbreaking process[50] ................................... 28
Figure 2-9 Kellogg Brown & Root visbreaking technology[47] ............................. 30
Figure 2-10 Lurgi Öl Gas Chemie technology[47]................................................ 31
Figure 2-11 Aquaconversion technology[47] ....................................................... 39
Figure 2-12 Molecular structure of Morichal1 (C100H107N2S2O) ......................... 41
Figure 3-1 Bench scale plant for reactivity and catalysts evaluation.................. 47
Figure 3-2 Feed tank specifications................................................................... 49
Figure 3-3 Detail of probes installed inside the reactors.................................... 52
Figure 3-4 Detail of the reactors zones with corresponding probe inside .......... 53
Figure 3-5 Double ended cylinder used as heavy product tank......................... 55
Figure 3-6 Double pipe heat exchanger configuration used as condenser........ 55
Figure 3-7 Graphic summary of the experimental plan...................................... 60
Figure 4-1 Simulated distillation curves for virgin VGO and liquid products from SET1 runs ........................................................................................... 68
Figure 4-2 Simulated distillation curves for virgin VGO and liquid products from SET2 runs ........................................................................................... 71
Figure 4-3 Temperature effect on SimDist result for liquid product. Curves at (a) pressure=160 psig and (b) pressure= 260 psig...................................... 75
Figure 4-4 Pressure effect on SimDist results for liquid product. Curves at (a) temperature= 400ºC, (b) temperature= 410ºC and (d) temperature= 420ºC. ......................................................................................................... 76
Figure 4-5 Variation of MCR in liquid product .................................................... 77
Figure 4-6 Variation of conversion with temperature and pressure. .................. 79
Figure 4-7 Temperature effect on SimDist result for liquid product. Curves at (a) pressure=160 psig and (b) pressure= 260 psig...................................... 83
Figure 4-8 Variation of MCR in liquid product .................................................... 84
Figure 4-9 Variation of conversion with temperature, pressure and spatial velocity ........................................................................................................ 86
Figure 4-10 Athabasca VGO conversion map ................................................... 96
Figure 4-11 Products selectivity versus Athabasca VGO conversion ................ 97
Figure 4-12 Relative MCR versus Athabasca VGO conversion......................... 98
List of Symbols and Abbreviations
Symbol AVGO bbl C cc cP He HTSD HVGO k Lpm LVGO M mA MCR N PH1 PH2 PI SCSC SimDist S TBP TI TIC V VAC VGO
Definition Athabasca Vacuum Gasoil Barrel Carbon Cubic centimetre Centi Poise Helium High temperature simulated distillation Heavy vacuum gasoil Kilo Litres per minute Light vacuum gasoil Mega Mili ampere Carbon as MicroCarbon value Nitrogen Pre-heater 1 Pre-heater 2 Pressure indicator Selective Catalytic Steam Cracking Simulated distillation Sulphur True boiling point Temperature indicator Temperature indicator and controller Volt Volt alternate current Vacuum gasoil
1
Chapter 1: Introduction
1.1 Background
According to the Energy Information Administration, total primary energy
consumption in North America, including energy for electricity generation, will
grow by 1.1 percent per year from 2005 to 2030 (Figure 1-1) and fossil fuels
account for 87% of the growth [1].
Figure 1-1 Primary energy consumption estimation, 2005-20030 (quadrillion BTU)[1]
Total world oil demand reached 84.3 million bbl/d in 2006, up 0.9% or 0.8
million bbl/d from 2005. Demand is expected to grow by 1.8% or almost 1.5
million bbl/d in 2007. Figure 1-2 illustrates growth in oil demand across the globe
between 2005 and 2007 [2].
2
Figure 1-2 Growth in oil demand 2005-2007[2]
Oil reserves in the world can be classified in two types:
1. Conventional crude oils, characterized by a high API gravity
(greater than 20 ºAPI), and relatively low levels of contaminants
and residue and,
2. Unconventional crude oils i.e, heavy oils, extra heavy oils and
bitumen, with API gravity values lower than 20 API, high levels of
sulphur, nitrogen and metals and considered low quality oils, very
difficult to produce and process [3],
World oil production has been focused on the economic advantages of
extraction and refining of conventional crude oil and which continues to take
priority over the unconventional oils However, conventional crude oil reserves
continue to decline with each year as the energy demands of China and India
and other parts of the world continue to grow.
Alberta has vast reserves of bitumen and major technological advances in
the past 10-15 years have proven that extraction and refining are viable
particularly given the high price of conventional crude oil in world markets. Figure
3
1-3 compares the initial sizes of Alberta’s conventional crude oil and bitumen
reserves, and the current remaining reserves. While most of Alberta’s known
conventional crude oil reserves have been exploited, only a fraction of the crude
bitumen has been tapped [2].
Figure 1-3 Alberta’s conventional crude oil and crude bitumen reserves[2]
In this environment of increasing energy demand and economic growth and,
notwithstanding the difficulties of extraction and processing heavy residues, oil
sands and crude bitumen, they represent an immense and real alternative for the
conventional crude. The total estimated in-situ and recoverable bitumen reserves
in Canada are 27.7 billion m3 (174 billion barrels), similar to 2002. To date, only
2 per cent of the initial established crude bitumen reserve has been produced.
Total raw bitumen production, which exceeded total conventional crude oil
production for the first time in 2001, and has continued its growth, accounting for
70 per cent of Alberta’s total crude oil and raw bitumen production in 2006 [4].
Due to their low quality characteristics, unconventional oils add complexity
to the production, upgrading and refining processes. The choice for many
companies to process and carry out the upgrading of unconventional crude have
4
been mainly focused in conventional thermal processing technologies such as
delayed coking and visbreaking, because of their lower investment and operating
costs, high reliability and ability to economically produce finished products and
reduce oil production. However, the continuous oscillation in the oil prices and
the tight refinery margins have been a very important driving force to search and
introduce new technologies for the upgrading of the bottom of the barrel that will
allow both processing those cheaper feedstock and reducing residue production
[5]. In this sense, countries like Canada and Venezuela with large reserves of
bitumen and heavy crude oil are facing two important challenges. The first of
them is to find an economical way to incorporate the vast reserves of this low
cost extra heavy oil into the refinery market and second how to process these
crude oils in the refinery to maximize transportation fuels while reducing the
production of high sulphur fuel oil.
Various processes are known for converting heavy hydrocarbons into more
desirable liquid and gas products. These processes include visbreaking and
extreme thermal cracking. However these processes are characterized by low
conversion rates and sometimes a large percentage of undesirable by-products
such as coke which, among other things, can pose transportation and disposal
problems.
New technologies based in Selective Catalytic Steam Cracking (SCSC)
have been developed in order to reduce the viscosity of the heavier components
of the refinery’s fuel oil pool, reduce fuel oil productions and to increase the yield
of distillates for a minimal incremental investment [5]. This process enables the
refiner to profit from the additional yields of distillate and lighter products
produced at conversion levels significantly higher than those attainable with
conventional visbreaking technologies. The unique features of SCSC are the
inclusion of steam and ultradispersed catalysts to the thermal process in order to
cause moderate hydrogen incorporation from water to the thermal products [6].
The proposed SCSC reaction sequence proceeds by the catalytic dissociation of
water into hydroxyl and hydrogen free radicals, concurrently with the formation of
5
hydrocarbon free radicals due to the thermal cracking of the feedstock. Thus,
hydrogen free radicals saturate the resulting hydrocarbon free radicals to
produce lighter hydrocarbons, reducing aromatic condensation reactions and
formation of coke precursors [6].This novel process allows refiners to reach
higher conversion levels than those achieved in visbreaking, with the additional
advantage of still being able to produce a stable converted product. This process
has been developed by PDVSA-INTEVEP and it supports Aquaconversion
technology.
1.2 Motivation
The group “Catalysis for Bitumen Upgrading and Hydrogen Production” was
created at the Department of Chemical and Petroleum Engineering at the
University of Calgary in January 2004 with the arrival of the first contingent of
professionals ready to develop research thesis under the supervision of Dr.
Pereira Almao. The group had developed experience on processing Extra Heavy
Oils from the Orinoco Belt in Venezuela. It had also been involved in the
development of two catalytic processes to upgrade such bitumen,
Aquaconversion, a catalytic steam cracking process and HDH-Plus, a
hydroconversion process. The feedstock from Northern Alberta, although similar
in composition to the Venezuelan bitumen, needed to be tested with these
processes to assess the applicability of them to its upgrading. For the purpose a
reactivity unit needed to be setup and a representative heavy fraction had to be
selected. Developing a pilot test unit to assess the reactivity of Northern Alberta
heavy oil fractions requires identifying the smallest scale significant enough to
include all the effects that may be important for said processes, and to use an
adequate heavy fraction that is sufficiently easy and economic to handle in that
scale. Vacuum Gas Oil (VGO) was selected as feedstock for two reasons; the
first of a practical and analytical nature. VGO has a viscosity low enough to be
transported with the available pumps and it does not contain asphaltenes, which
6
helped to simplify an initial understanding of the reactivity of Athabasca.
Secondly VGO is the second largest heavy fraction in bitumen, typically around
30% wt.
Vacuum gas oil (VGO), is conventionally treated using fluid catalytic
cracking (FCC) to obtain gasoline as the product of interest. In accordance with
the latest trends of the market, one particularly desirable product to obtain is
diesel, also known as light gasoil or LGO. However, conventional FCC provides a
small conversion to LGO of about 15% of the feedstock [55] and this diesel is of
poor cetane index. The Application of the catalytic steam cracking process
(SCSC) to the Athabasca vacuum gasoil may result in good conversion to diesel
with reduced levels of undesirable by-products like gas or coke.
In this sense, to have a reliable setup where evaluations of heavy
hydrocarbon feedstock such as Athabasca vacuum gasoil (AVGO), vacuum
residue and bitumen can be carried out not only would open a potential of
investigation in heavy oils in the University of Calgary but also would offer high
possibilities for newer processing applications to the large reserves of bitumen
existing in Alberta.
1.3 Objectives
The main objective of this study is the evaluation of Athabasca vacuum
gasoil (VGO) reactivity via Thermal Cracking processing and Selective Catalytic
Steam Cracking processing in order to generate a reactivity data base and
determine increments in lighter fraction of the product. An experimental bench
scale unit is built as an important part of the scope of this work, to carry out all
the experiments. To accomplish this objective, the thermal processing of the
feedstock is studied as reference in order to evaluate and compare the effect of
water and catalyst on the conversion and product yields.
The specific objectives are:
1. Construction of a bench plant for thermal and catalytic
reactivity evaluation of heavy oils and fractions.
7
2. Test for production of converted streams of Athabasca VGO
via:
i. Thermal processing.
ii. Catalytic Steam processing (applying an ultra-
dispersed catalyst system)
3. Characterization and comparison of converted VGO via:
i. Simulated distillation
ii. On line gas chromatography
iii. Microcarbon Residue method
8
Chapter 2: Literature Review
Thermal cracking of heavy oils, vacuum residues and bitumen has been
advanced as an upgrading process so as to meet the increasing world demand
for lighter and middle distillates [9]. Technologies for upgrading heavy crude oils
can broadly be divided into carbon rejection and hydrogen addition processes.
Carbon rejection redistributes hydrogen among the various components,
resulting in fractions with increased H/C atomic ratios and fractions with lower
H/C ratios, and is commonly present in thermal cracking processes (delayed
coking, fluid thermal cracking, flexicoking, etc.). On the other hand, hydrogen
addition processes involve reaction of heavy crude oils with an external source of
hydrogen resulting in an overall increase in H/C ratio that can be found in hydro-
processing reactions (hydrotreatment, hydrocracking and hydro-visbreaking)
[8,32]. Visbreaking is the oldest of these processes and the most well known and
widely applied, contributing about 32% in terms of the total heavy oil volume
processed and followed closely by delayed coking with 30% [32].
The main target in the heavy crude upgrading is to increase the H/C ratio in
order to obtain better products, however carbon rejection technologies require
the addition of large amounts of hydrogen downstream to saturate the alkenes
and aromatics which are produced or in the reactions of hydrodesulphurization,
hydro-denitrogenation, hydro-deoxidation and demetallization, which also
improve the quality of the derived oil [10].
If a catalyst is used to improve or make the thermal cracking more selective,
the process is called Thermal Catalytic Cracking. Water normally doesn’t
dissociate at the thermal cracking conditions, however, to be able to dissociate
water at this conditions, a specific type of catalytic function is required. If this
catalytic function is added to the thermal catalytic cracking catalyst the whole
process is called selective catalytic steam cracking process. This last option can
be proved advantageous not only economically but also operationally, and could
be an interesting option in the mild conversion of heavy oils.
9
2.1 Thermal Cracking
Thermal cracking was used commercially in the production of oils from coal
and shale before the petroleum industry began. The precise origins of oil
fractions cracking are unknown. It is rumored that it was around 1861, when it
was discovered that heavier petroleum products could be decomposed to lighter
and could be used to increase the production of kerosene [8].
Generally speaking, thermal processes are those processes of primary
improvement of heavy crude oils and residues in which the decomposition,
rearrangement or combination of hydrocarbons molecules is carried out through
the application of heat and without the aid of a catalyst [11]. Particularly thermal
cracking can be defined as thermal decomposition of big molecules to small
molecules, applying high temperatures (higher than 400°C) and moderate
pressures (50 to 200 psig) [8,12]. The process is spontaneous and the non
selective hydrocarbon disintegration occurs simultaneously with condensation
and polymerization reactions, as a result molecules with a wide range of
molecular weight and boiling points are produced. This means that the products
of the reaction are not only lighter than the feed, but also a heavier residual with
less hydrogen content than the feed is produced, which depending on its size
and concentration, could produce some precipitate. Consequently, the
conversion levels in thermal cracking are usually limited by the stability of the
residue generated [8,11,12,13]. In all these processes the targets are usually as
follows: higher yields of light cracked oil, increasing total liquid yield, reduction of
utility cost and reduction of construction cost [26].
Thermal processing of hydrocarbon feedstock has been the study focus of
many research groups in the past century. Some researchers developed highly
theoretical models trying to detail the mechanism of thermal cracking reactions
such as the work of free radical mechanism as proposed by Rice [14] and
defined more accurately by Kossiakoff and Rice [15] where the theory of free
10
radical and the application of the concept of chain reactions to the thermal
decomposition of the organic compounds is extensively developed.
There are pioneer works on the thermal cracking of gasoil like the one of
Geniesse and Reuter [16] in which varying residence times from 0.75 to 4800
seconds and temperatures from 430°C to 700°C were evaluated. They were able
to prove that an increase of 17°C halves the time necessary to produce the same
results at the lower temperature.
Thermal cracking of pure hydrocarbons like n-hexadecane and others in the
range of C15-C20 has also been investigated by researchers due to their
similarities in molecular weight, average boiling point and other physical
characteristics to heavy oils, gaining considerable knowledge into the reaction
mechanisms of more complex feedstock. In their study on thermal cracking of
higher paraffins, Voge and Good [17] thermally cracked n-dodecane, iso-decane
and n-hexadecane at 500°C and the results at one atmosphere were consistent
with the predictions of the Rice radical chain theory [14], amplified by Kossiakoff
and Rice [15]. Fabuss et al [18] studied the thermal cracking of pure cetane (n-
hexadecane) in a flow reactor at 1100 °F to 1300 °F and pressures of 200 to
1000 psi. The average molecular weight of the product steadily dropped with
increased per cent cracking. The amount of coke deposits were far less than
those previously reported in the literature at that time, probably because of high
flow rates. The rate of reaction was substantially limited by the rate of heat
transfer. Wonsky [19] simulated the kinetics of thermal cracking of high
molecular weight normal paraffins (C16-C20) using an analog computer. The
results give over-all reaction orders and activation energies comparable with the
available data in the range of 475°C to 625°C and 0.2 to 70 atm. Results could
be approximately correlated over a wide range of conditions by an over-all
reaction order of 1.3 and activation energies of 221 kJ/mol. Most notable is the
effect of molecular weight on the overall reaction rate; the overall reaction rate
decreased with decreasing molecular weight. Blouri et al. [20] studied the
controlled cracking in the liquid phase of pure heavy hydrocarbons (n-
11
hexadecane, 6-methyleicosane, 1-phenyldodecane and C21-C27 paraffins in a
micro-reactor between 350°C and 440°C for residence times varying from 0.5 to
4 hours at nitrogen or hydrogen pressures of 20 bar with the intention of
simulating visbreaking of heavy oils. Kinetic data obtained, such as the order of
reaction and the activation energy were similar to those of a radical type
cracking. Savage [21] developed a reaction model for the pyrolysis of binary
mixtures of compounds using mixtures of n-pentadecylbenzene and n-
tridecyclohexane as asphaltene model compounds. The study provides an
excellent description of the use of these model compounds to simulate the
thermal cracking of more complex hydrocarbons (asphaltenes). Khorashed and
Gray [22] carried out experiments similar to Fabuss et al. [18] at lower
temperatures (380°C - 450°C) and higher pressure (13.9 MPa) developing a
simple kinetic model based on a free-radical mechanism to account for the
observed product distributions and overall n-C16 conversion and within the
relatively low conversion employed in this study, agreement between predicted
and experimental results was quite satisfactory.
Closely linked with the studies on thermal cracking of pure hydrocarbons
are the studies on thermal cracking of heavy oils. In these studies the reactor
conditions are not as applicable as pure hydrocarbons like n-hexadecane simply
due to the complexity of the feed to the reactor, the huge number of components
and reactions that characterize the system in addition to the difficulty to develop
a detailed kinetic model. However, the results presented related to thermal
cracking of heavy oils are valuable for determining chemical and physical
properties of cracked products, kinetic modeling and reactivity characterization
[30].
The first models recorded in the literature for predicting thermal cracking
product distributions were developed at the University of Calgary [24] by
Hayashitani et al. [23]. In this work, studies are carried out about thermal
cracking reaction models based on pseudo reactions mechanism proposed in
order to be incorporated into numerical simulators of thermal recovery processes
12
for the Athabasca oil sands. Hayashitani's model was enhanced by the work of
Lin et al. [25], which also includes a literature review of alternative cracking
models. Takatsuka et al. [26] developed a practical model of thermal cracking of
residual oil stating reactions of four lumped hypothetical components defined to
represent polymerization and condensation of residue. The model confirmed the
performance of various types of reactors such as batch, semi-batch and
complete stirred-tank reactor (CSTR) both in laboratory experiments and in
industrial settings. Del Bianco et al. [27, 28, 29] carried out a series of studies
about thermal cracking of petroleum residues. In the first one the thermal
cracking was studied using a batch reactor and long residence times. The
characteristics of the distillate product collected compared well with that
predicted by first-order reaction. The results of this work were used as reference
data in a second study of residue thermal cracking in the presence of a hydrogen
donor solvent [28, 29], using a batch micro-reactor first and then a continuous
stirred tank reactor at bench scale plant level, confirming that the visbreaking
performances in terms of net conversion, product quality and inhibition of coke
formation can be substantially improved adding a hydrogen donor solvent which
also allows operations at much higher severities than conventional thermal
cracking processes.
Kawai and Kumata [31] investigated the thermal cracking mechanism of
heavy oil, specially the effect of hydrogen partial pressure in the thermal cracking
reaction having as model compounds 1, 2-diphenylethane and dibenzylsulfide to
represent a part of the asphaltene structure. The authors concluded that free
radical recombination was not affected by the change of hydrogen partial
pressure within the range studied.
Kataria et al. [32] studied the thermal cracking of vacuum residues and
asphalts in a batch reactor varying temperatures in the range of 400°C to 430°C ,
residence time from 0 to 15 min and keeping pressure at a constant value of 174
psig. They concluded that the thermal cracking of the studied feedstock gives
gas and distillates as products without observing coke formation. They also found
13
that for a given feed, the composition of the gas fractions didn’t change by
changing severity. In addition, they concluded that the product yields and their
corresponding activation energies at the conditions evaluated are a function of
the severity and feed characteristic. They also noticed that the stability of liquid
products decreases with increasing conversion of vacuum residue. Kataria also
reported that the activation energy for vacuum gasoil cracking was in the range
of 265-398 KJ/mol. In addition, Jasvinder et al. [33, 34] has developed studies on
thermal cracking of residual feedstocks (short residue, visbreaker feed and
asphalt) in a batch reactor, confirming that thermal cracking of residues follows
the first order reaction kinetics. They found the activation energies to be in the
range of 102-206 kJ/mol.
2.1.1 Chemistry of the Thermal Cracking
Thermal cracking reactions play an important role in fluid flow in the
reservoir because the flowing oils thermally transformed do not have the same
fluid properties as the original oil in place. Thermal cracking reactions are also
important for the design of thermal conversion units and upgrading facilities [23].
The thermal cracking reaction chemistry is quite complex, and the degree of
complexity increases when cracking heavier hydrocarbons at higher operating
severity. The enormous multiplicity of chemical reaction coupled with non-uniform
heat transfer makes the development of a comprehensive pyrolysis theory a
major challenge for chemists and mathematicians [35]. As previously indicated
the hydrocarbon reactions at high temperatures are based on a mechanism
proposed by Rice in 1931 [14] and defined more accurately by Kossiakoff and
Rice in 1943 [15]. In these studies thermal decomposition of hydrocarbon is
defined fundamentally as a chain reaction that involves free radicals. The
concept postulates three basic types of reactions, which are represented in
Figure 2-1.
14
1. Chain Initiations, or the initial formation of radicals:
CnH2n+2 → CmH2m+1 ° + C(n-m)H2(n-m)+1 ° Eq. 1
2. Chain Propagations, through reaction of radicals with molecules:
CnH2n+2 + CmH2m+1 ° → CnH2n+1 ° + CmH2m+2 Eq. 2
CnH2n+1 ° → CmH2m + C(n-m)H2(n-m)+1 ° Eq. 3
3. Chain Terminations causing the disappearance of radicals:
CnH2n+1 ° + CmH2m+1 ° → CnH2n + CmH2m+2 Eq. 4 CnH2n+1 ° + CmH2m+1 ° → CnH2n+2 + CmH2m Eq. 5
CnH2n+1 ° + CmH2m+1 ° → C(n+m)H2(n+m)+2 Eq. 6
Figure 2-1 Reaction scheme for Thermal Cracking reactions[12,13]
In the chain initiation step two radicals are produced by the cleavage of the
C-C bonds of paraffin molecules. During the chain propagation step many
different reactions are involved including hydrogen abstraction, addition, radical
decomposition, and radical isomerization. The reaction (2) is called β-scission
because the C-C bond located two carbons away from the hydrogen deficient
carbon breaks forming an olefin and a smaller alkyl radical. Chain termination
step is the reverse of chain initiation, causing the disappearance of radicals.
Chain reaction takes place when reactions (2) and (3) are much more frequent
than initiation and termination reactions. As a result, free radicals numbers
increase until reaching a steady-state concentration that allows the thermal
cracking propagation reactions to continue. On the other hand, combinations
resulting from termination reactions may produce heavier compounds than the
ones present in the feedstock.
15
The thermal cracking of a complex hydrocarbon mixture such as petroleum
is determined by the reactivity of its constituents. Bond dissociation energy is the
energy needed for homolytic breaking of one mole of a covalent bond into two
mole of free radicals in the gas phase: A:B → A. + .B [36]. Bond breaking is
always endothermic and the sign of ΔHd is positive. The bond dissociation
energies shown in Table 2-1 give an idea of the difficulty for bond breaking in
different types of bonds found among the hydrocarbon compounds that constitute
the crude oil.
Table 2-1 Bond dissociation energies [12,13]
In general, for compounds of the same molecular weight their relative
thermal reactivity depends on the chemical nature of the feedstock. Therefore,
the reactivity of the chemical families has the following descendent order
[11,2,3]:
Paraffins > Naphthenes (Cycloparaffins) > Olefins > Aromatics
Type of Bond Bond Dissociation Energy (Kcla/mol)
C – C 82.6 C = C 145.8 C ≡ C 199.9
C – H (n-alkane) 98.7 C – H (aromatic) 110.5
H – H 104.2 C – O 85.5 C – S 65 S – S 84 S – H 83
C – N (amine) 72.8 C = N 147 C ≡ N 212.6
(Kcal/mol)Type of Bond Bond Dissociation Energy (Kcla/mol)
C – C 82.6 C = C 145.8 C ≡ C 199.9
C – H (n-alkane) 98.7 C – H (aromatic) 110.5
H – H 104.2 C – O 85.5 C – S 65 S – S 84 S – H 83
C – N (amine) 72.8 C = N 147 C ≡ N 212.6
(Kcal/mol)
16
On the other hand, the thermal disintegration tends to increase with the
molecular weight or the boiling point range. This observation may be
contradictory but the big hydrocarbon molecules, which were not under severe
thermal cracking, have more bonds available to be broken, in consequence,
there is a higher probability for new scissions to take place. As a result, the
higher the cracking time, the higher the yield of light products. The value of the
light product yield reaches a maximum which then decreases gradually at the
expense of formation of gases and coke. The cracked fractions with lower
molecular weight are thermally more stable than the others with higher molecular
weight. For this reason, naphtha yields are the result of heavier cracked fractions
such as gas oil, produced in the first stage of the thermal reaction.
Heavier liquid feeds are composed principally of paraffins, cycloparaffins,
aromatic ring structures, and occasionally, olefins. The paraffinic hydrocarbons
have a high hydrogen-to-carbon ratio and decompose to produce high yields of
gaseous products and low yields of heavy aromatic oils. The cycloparaffins
decompose by ring rupture accompanied by the formation of alkyl radicals.
These compounds require higher activation energy levels compared with
paraffins. The hydrogen-to-carbon ratio is lower than paraffins but higher than
aromatics, and therefore, the yield of gaseous products and heavy liquids is
intermediate (i.e., between paraffins and aromatics). Aromatic ring structures are
highly heat-stable and are refractory compounds. The hydrogen-to-carbon ratios
are low, and olefin yields from aromatic feeds are negligible. These compounds
are tar precursors and cause the build-up of coke in the cracking coils through
condensation reactions. Even though the alkanes are not normally the main
constituents of most of the residues, the model of the thermal chemistry of the
paraffins offers an easy platform to understand the thermal reaction of any other
compounds family.
17
2.1.2 Severity and Conversion in Thermal Cracking Processes
The terms severity and conversion are used to measure the extent of
cracking. The severity in the thermal cracking processes is given by a
combination of the reaction time and temperature, which determines also the
conversion and characteristics of the product. Therefore, large reaction times at
lower temperatures may be equivalent to short reaction times at higher
temperatures. The treatment of heavy oils goes from moderate treatments, in
order to reduce the viscosity of the feeds, to the ultra-pyrolysis in order to reach
the complete conversion of olefins. Thermal processes at moderate and high
severity are commonly used for the heavy oils processing. Meanwhile, processes
at similar severity to ultra-pyrolysis are used commercially to reach the thermal
cracking of ethane, propane, butane and light feedstocks to produce ethylene
and other olefins of petrochemical interests.
Conversion of a single component is measured by the rate of its
disappearance relative to its concentration in the feed. However, when cracked
as mixed feed the single compound is also formed as a product of cracking from
other larger compounds. The measured conversion of a single compound in a
mixed feed provides only an approximation of the true conversion. For liquid
feeds, such as naphtha, it is impractical to calculate true conversion. Instead,
several other indicators are measured/calculated to define severity of operation.
This includes propylene-to-methane ratio, molecular collision parameter, average
molecular weight of the complete product, hydrogen content in the C5+ product,
etc.
18
2.2 Visbreaking
2.2.1 Definition and Fundamentals of Visbreaking
Applied since the beginning of last century, visbreaking is one of the
simplest and least expensive processes that have been used to reduce the
production of residual and to increase the distilled yields. As its name indicates,
this process reduces the viscosity of residuals (atmospheric and vacuum)
through mild liquid phase pyrolysis of large molecules to smaller ones to form
lighter and gaseous products. In principle the technology of this process is very
simple and the thermal reactions (being essentially first order reactions) take
place at low pressure and without catalyst. The severity of this process is
controlled by operational variables such as temperature and residence time.
Most of the commercial units have been modified in the past decades from coil-
type (high temperature cracking occurs in a dedicated coil in the heater with low
residence time) to coil soaker (cracking begins in the heater at lower temperature
than in the coil-type but occurs primarily in a downstream soaker drum for a
lower controlled amount of time) which has allowed reduction of the process
temperature by about 10ºC and as a result, this modification has enhanced the
yields of gas and light liquid products. In addition, the concurrent visbroken
products obtained achieve relatively superior quality of the fuel oil.
For a given feed the extent of conversion, selectivity to gas and light liquid
products, coking behavior and stability of the visbroken product are directly
related to the feed characteristics, such as paraffin, asphaltene content,
aromaticity, hydrogen content (hydrogen-to-carbon ratio), heteroatom content,
etc., varying to a large extent from feed to feed. In research related to the effect
of feedstock in visbreaking, Yan [38] studied the visbreaking of crude and
residues reporting that the tendency to form coke is directly proportional to the
asphaltene content, supporting the theory that coke is produced from
condensation reaction of asphaltenes present in the feed. On the other hand,
Brauch et al. [39] found that using feeds with less than 1% of sulfur content, the
19
relative increase of asphaltenes was more drastic that in those feeds with sulfur
content higher than 2%.
The conversion extent of a particular feed within the stability range is called
the crackability of the feed [37]. To investigate the role of feed properties on
crackability, several research groups [40-44] have studied the kinetics of
visbreaking of heavy oils and residues from a large number of sources in a batch
or continuous reactor. Some work have already been commented in the previous
section, noting that these studies were primarily focused on thermal cracking
behaviour of feeds such as model compounds having similar characteristics to
heavy oil (viscosity, average boiling point, etc.) and also in actual heavy oils that
include atmospheric residue, vacuum residue, and their blends which act as
major feeds for the visbroken units.
It is well known that operating parameters of the visbreaking units are
adjusted to produce the maximum reduction of feedstock viscosity, without
making the product unstable. Hence, conversion of the process is limited by the
stability of the visbroken residue. Increasing severity produces a change in the
chemical composition of product that progressively modifies the peptization of
asphaltenes, which tend to precipitate and cause troubles like clogging pipes
(fouling) and increase the instability of the residue during storage. As a result,
miscibility or solubility characteristics of residue components are closely related
to their stability. From the experimental perspective, visbreaking conversion limit
is set forth through empirical stability tests that may be carried out using different
methods. These methods are selected by refineries based on the experience
acquired from their own products.
2.2.2 Visbreaking: State of the Art Technology.
Although new thermal cracking units are now under development,
visbreaking still remains as a technology broadly applied in refineries. Major
Licensers of visbreaking technology include ABB Lummus Global/Shell
Solutions, Axens, Foster Wheeler/UOP, Kellogg Brown & Root, Lurgi Öl Gas
20
Chemi and UOP/PDVSA-Intevep/Foster Wheeler. A review of these major
commercial licensers is briefly described next.
2.2.2.1 ABB Lummus Giobal/Shell Global Solutions
More than two decades ago, ABS Lummus Global became the licensing
agent for Shell Thermal Conversion technologies, which includes Shell Soaker
Visbreaking (SSVB), Shell Deep Thermal Conversion (SDTC), the Shell Thermal
Gasoil (STGP), the Shell Deep Thermal Gasoil (SDTG) process, and Shell
Vacuum Flashing (SVF).
2.2.2.1.1 Shell Soaker Visbreaking (SSVB)
In this process, Figure 2-2, preheated feed is cracked in a heater and
soaker drum before being flashed and quenched in an atmospheric fractionator
[46]. Liquid residue in the bottom of the fractionator and a side-stream of gasoil
are steam-stripped and cooled before exiting the battery limits. Alternatively, the
gasoil side draw can be combined with the liquid residue. The HVGO can be also
recovered from the cracked residue if a vacuum flasher is included.
Figure 2-2 Shell Soaker Visbreaking Process (SSVB) flow diagram[46]
21
This visbreaking process offers several advantages over coil-type
processes according to Shell and ABB. Due to lower cracking temperatures and
longer residence time and the lower pressure drop in the heater, it provides
higher tolerance for feedstock changes and operating upsets, enhanced process
control, extended heater run lengths, faster turnarounds, and lower power
requirements. The investment cost can be lowered by 15% due to the smaller
heater size, reducing the size requirement of the waste heat recovery system
and decreasing fuel consumption. This process also addresses the problem of
back mixing that occurs in the soaker through the use of special internals. This
boosts the conversion by 0.4-0.7% without reducing fuel oil stability and also
increases distillate yields by 1-2% and lowers cutter stock usage.
As of 2002, more than 80 Shell Soaker Visbreaking units have been built,
were converted from coil visbreakers and crude units or are in the construction
phase. Over 70% of the total visbreaking capacity built during the last 10 years
was based on this Shell technology. The 1998 US investment cost of this type of
visbreaker, excluding treating facilities, is US$1.0k/bbl to US$1.4k/bbl. Typically,
for each barrel of 180°C feed sent to this unit, 16,000 kcal (63.5 kBtu) of fuel, 0,5
kWh of electricity, and 0.1 m3 of cooling water are consumed while 18 kg of
steam is produced [50].
With typically 20% of the vacuum residue feed converted to distillate and
lighter products, Shell Soaker Visbreaking is one of the lowest cost conversion
process options.
2.2.2.1.2 Shell Deep Thermal Conversion (SDTC)
In this process, Figure 2-3, preheated vacuum residue is charged to the
visbreaker heater and from there to the soaker, where the deep conversion takes
place [46, 50]. The soaker effluent is flashed in a cyclone and the cyclone
overheads are charged to the flash zone of the atmospheric fractionator to
22
produce the desired products like gas, LPG, naphtha, kerosene and gasoil. The
bottoms from this unit and the liquid phase from the cyclone are combined and
sent to a vacuum flasher to recover additional gas oils. The vacuum flasher’s
bottom, which is also called liquid coke stream, can be used for gasification, in
bitumen production, or as a binder used to make materials such as pipes and
bricks. Using this product as gasifier feed is a good economic option since its
caloric value is six times its blending value.
Figure 2-3 Shell Deep Thermal Conversion process (SDTC) flow diagram[46]
This type of vacuum flasher, which is based on Shell Vacuum Flashing
technology, can run for more than 18 months without requiring cleaning.
The SDTC process offers important advantages over other residue
upgrading options: a coke handling system as in delayed coking is not required
since the bottom stream from the flasher is a liquid, providing an advantage in
terms of investment cost. It also utilizes the hydrogen in the feed to produce
higher-quality distillates than obtained through delayed coking. Finally, the
distillate yield from this process (45-60 wt%) is higher than that from traditional
23
soaker visbreaking due to the conversion achieved and the use of the vacuum
flasher [50].
The 1998 US$ investment cost of a SDTC unit, excluding treating facilities,
ranges from US$1.3k/bbl to US$1.6k/bbl. The low energy usage and
maintenance needs of an SDTC unit enable this capital expenditure to be
recovered in less than one year. Typically, for each barrel of feed (at 180°C) that
is sent to this unit, 26 Mcal (103 kBtu) of fuel, 0.5 kWh of electricity, and 0.15 m3
of cooling water are consumed while 20 kg of steam is produced [50].
As of 2002, four licensed SDTC units were operating including two that
were revamped from Shell Soaker Visbreakers The first non-Shell refinery to
license the STDC technology was the Litvinov refinery of Ĉeska Rafinerska in the
Czech Republic [47].
2.2.2.1.3 Shell Thermal Gasoil Process (STGP)
In this process, Figure 2-4, atmospheric or vacuum residue feed is cracked
in a heater and soaker before being routed to a cyclone. The overhead stream
from the cyclone is sent to an atmospheric fractionator, which separates out gas,
naphtha, gasoil, and HGO. The bottom streams from the cyclone and fractionator
are vacuum flashed to recover VGO from the cracked residue. The HGO from
the atmospheric fractionator and the VGO from the vacuum flasher are cracked
in a distillate thermal cracking heater using coil cracking technology, and the
cracked distillates are routed to the fractionator. The cracked residue stream that
exits the bottom of the vacuum flasher is commonly used to produce commercial
fuel oil. The heater and soaker can also be adjusted to SDTC conditions so that
liquid coke can be produced [46, 50].
As of 2002, a total of 13 STGP units had been built or were currently in the
construction phase The 1998 US$ investment cost of this type of unit excluding
treating facilities, ranges from US$1.4k/bbl to US$1.6k/bbl. Typically, for each
barrel of feed (at 180°C) that is sent to this unit, 34kcal (135 Btu) of fuel, 0.8 kWh
24
of electricity, and 0.17 m3 of cooling water are consumed while 29 kg of steam is
produced [50].
Figure 2-4 Shell Thermal Gasoil process (STGP) flow diagram[46]
2.2.2.2 AXENS
Axens offers three variations of its TervahI process: Tervahl T, Tervahl H
and Tervahl C.
2.2.2.2.1 Thervahl T Process
In this process, Figure 2-5, the hot vacuum residue conversions achieve
around 30 wt%. The temperature of the feedstock is raised to the desired
temperature using heat recovered in the stabilization section and a coil heater
before the stream is sent to a soaker where it is held for a specified residence
time. The output from the soaker drum is then quenched and sent to a stabilizer,
which yields a product stream and gas. The gas produced from the process is
used for fuel [48,49]. With this process conversions around 33% can be reached.
25
Figure 2-5 The Tervahl T process[49]
2.2.2.2.2 The Tervahl H Process
In the Tervahl H hydrovisbreaking process, Figure 2-6, the feedstock
passes through the heat recovery section and is combined with an hydrogen rich
stream that has been compressed to 100 bar (10 MPa) and then sent to a fired
heater to raise the temperature of the combined stream before it enters a soaker
where it is held as in the Tervahl T process. The effluent from this drum is
quenched with recycle hydrogen and sent to a high- pressure hot separator. The
gas from this unit is sent to a high-pressure cold-separator for the recovery of the
hydrogen stream to be recycled to the heater and soaking drum effluent. The
liquid streams from both separators are sent to the stabilization section where
purge gas and liquid product is separated. This process increases conversion by
about 8% over Tervahl T and achieves a greater reduction in product viscosity
[49].
26
Figure 2-6 The Tervahl H process[49]
2.2.2.2.3 The Tervahl C Process
The Tervahl C process, Figure 2-7, is configured in the same way as
Tervahl H, but a stream of 50 ppm of an ultra-dispersed molybdenum additive is
combined to the feed stream after the fire heater to increase hydrogen efficiency
and produce a stable product. Other benefits claimed include a 40% conversion
and a further drop in residue viscosity [47].
The 1994 US Gulf Coast investment costs of 15kbsd Tervahl T and
Tervahl H visbreaking units were US$1.5k/bpsd and US$2.05k/bpsd respectively.
These amounts include the cost of desalting, topping, and engineering. For each
barrel of crude fed to a Tervahl T unit, typically 2.4 kg of fuel oil and 0.3 kWh of
electricity are consumed while 2.4 kg of steam is produced. For a Tervahl H unit,
the utility requirements for each barrel of feed are 2.4 kg of fuel oil, 1.9 kWh of
electricity, 4.8 kg of steam, and 4.8 m3 of hydrogen [47].
27
Figure 2-7 The Tervahl C process[47]
2.2.2.3 Foster WheeIer/UOP
Foster Wheeler and UOP combined their visbreaking technologies in 1993
offering a coil-type process that utilizes high furnace outlet temperatures and
short reaction times [51].
The unit used in this process, Figure 2-8, can be configured in three ways
depending on the objective of the refiner. In the first option, the feed is charged to
the visbreaker heater reaching temperatures of about 427°C, causing partial
evaporation and mild cracking [50]. This stream is then sent to the second zone,
the soaking coil, where it remains for a specific time depending on the desired
conversion; in addition, steam is injected into the heater coils to help control
residence time and limit coke formation. The heater outlet stream is quenched
with gasoil or fractionator bottoms or a combination of both to stop the cracking
reaction and prevent the fuel oil product from becoming unstable. The vapour-
liquid mixture enters the fractionator to be separated into gas, naphtha, gas oil,
and visbreaker fuel oil (visbroken residue). A portion of the gasoil product is
steam-stripped and mixed with the visbreaker fuel oil to reduce its viscosity [50].
28
Figure 2-8 Foster Wheeler/UOP Visbreaking process[50]
The second option has the same configuration except that the bottom
stream is sent to a vacuum tower from the fractionator. The liquid that exits from
the bottom of this tower is steam-stripped and the vapour overhead is partially
condensed into distillate products that can be processed to transportation fuels.
Sidestreams of heavy vacuum gasoil (HVGO) and light vacuum gasoil (LVGO)
are removed from the column and combined, being used part of this product for
fuel oil viscosity reduction while the remaining portion is utilized as cracker
feedstock.
The third option is similar to the second except that the combined gasoil
stream is sent to a thermal cracker before being routed back to the fractionator,
increasing the yield of light distillates and reducing the yield and pour point of the
fuel oil stream, which is produced by blending vacuum tar and atmospheric gas
oil.
Innovations in Foster Wheeler coil visbreaker design have made it more
beneficial for refiners to use this unit rather than soaker visbreakers. Proprietary
valves and piping arrangements have been included in the coil heater allowing
29
individual heater passes to be decoked without shutting down the entire unit. This
improvement extends the heater run lengths, which are traditionally longer in
soaker units due to their lower furnace outlet temperatures. In addition, in order
to eliminate redundant equipment and reduce installation costs, coil visbreakers
have been integrated with vacuum units.
Over 50 Foster Wheeler/UOP visbreaking units have been installed
worldwide. The investment cost of a 10k-40k bpsd unit ranges from
US$785/bpsd to US$1,650/bpsd. Typically, for each barrel of feed charged to this
unit, 119.5kBtu, 0.0358 kW/bpsd of power, 6.4 lb of medium-pressure steam, and
71.0 gal of cooling water are consumed [50].
2.2.2.4 Kellogg Brown & Root (KBR)
The technology of visbreaking offered by KBR, Figure 2-9, claims to
increase distillate production and to reduce heavy fuel oil output by 25% [52]. In
this process, a mild cracking of the feedstock takes place in the furnace which is
heated at 427-510°C while controlling the pressure at 100-300 psig. The
temperature in the furnace effluent is decreased with a specially designed
quench valve, which cools the stream before it is flashed to a fractionation tower
where the gas, gasoline, gas oil, and residue product streams are separated.
The company also offers a soaker visbreaking process, which allows longer
residence times and lower reaction temperatures. The recovery of heavy gas oil
(HGO) product is avoided with this process and the fuel requirements in the
heater are also lowered.
KBR has developed some techniques for recycle cracking of visbroken
gasoil as well as special design methods which allow determination to be made
of furnace and quench parameters such as soaking coil volume and cracking
severity.
30
Figure 2-9 Kellogg Brown & Root visbreaking technology[47]
KBR has licensed its visbreaking process for more than 50 plants around
the world. Each commercial unit typically requires an investment of
US$0.9k/bpsd to US$1.1k/bpsd. For each barrel of feed, approximately 80 kBtu
of fuel and 0.5 kWh of electricity are consumed while 50 lb of steam is produced
[47].
2.2.2.5 Lurgi Öl Gas Chemie
In this process, Figure 2-10, the feedstock is preheated to more than 300°C
and sent through a heater where temperature reaches a value no higher than
450°C, leading the feed to partial vaporization and starting the cracking
reactions. The effluent from the heater is sent to a soaking drum where it remains
until the desired value of conversion is reached. The cracked product coming
from the soaker at temperature no lower than 400°C is sent to a fractionator
where separation of gas, naphtha, light gas oil, heavy gasoil and residue is
carried out.
31
Figure 2-10 Lurgi Öl Gas Chemie technology[47]
Lurgi claims that its visbreaking process provides significant advantages
such as a low investment cost, feedstock flexibility, and a heat exchange system
that reduces fuel consumption [47].
32
2.3 Catalytic Steam Cracking
Before reviewing the application of catalysts in steam cracking processes, a
brief overview is provided about how steam has been used in combination with
thermal processing.
It is important to mention that the hydrocarbon reactions in aqueous (liquid
and gaseous) media have been studied in diverse areas. In the area of
production, gases generated during the extraction of crude by alternate injection
of steam (H2, CO2, CH4, etc.) motivated Clark et al [69, 70] to develop studies to
understand the chemical transformations that organic compounds undergo in
aqueous media. These chemical reactions that happen in the reservoir in
presence of water are called aquathermolysis [9].
The aquathermolysis is associated directly to activities of oil production, but
also it has been studied to try to understand the petroleum formation. Between
1990 and 1997 Siskin, Katritzky et al [71, 72] published a series of articles in
which they demonstrate the reactivity in aqueous media (at normal conditions
and supercritical conditions) of aromatic and aliphatic compounds with and
without hetero-atoms (sulphur, oxygen, nitrogen), which were chosen as
representative molecules of kerogen composition. The conditions of the reservoir
under the oil recovery scheme of steam injection were simulated, with
temperature in the range of 200 -320ºC. Some of the major conclusions of these
works were:
1) Most of the evaluated compounds reacted in the presence of water, and
they generated compounds of smaller molecular weight and/or deheterogenized;
2) The formation of coke precursors was not observed.
On the other hand, in the processing area, the steam cracking of light oil
fraction is the most important process for producing a wide range of unsaturated
hydrocarbons for petrochemical use. Steam cracking is the thermal cracking and
reforming of hydrocarbons with steam at low pressure, high temperature and
very short residence times (generally less than 1 second) [73]. In commercial
applications of thermal cracking, such as visbreaking and delayed cocking,
33
steam is also injected into the heater coils and mixed with the feed to help control
residence time and limit coke formation.
Previous applications use the thermal fundaments with steam injection;
however, the addition of catalyst to the thermal-steam process in order to
produce an upgraded oil fraction, and even more, the use of ultradispersed
catalysts has been notably investigated in the last 20 years. The Catalytic Steam
Cracking can be defined as a process of moderate conversion of oil residues and
heavy crude oils, in which the hydrogen generation is made at low pressures
through the catalytic dissociation of the water. As its name indicates, the process
is based on reactions of thermal cracking that are carried out in the presence of
steam and catalysts. The use of steam as a source of hydrogen and the use of
bi-functional ultradispersed catalysts allows increasing the conversion of thermal
process like Visbreaking, and to maintain or to surpass the quality of thermally
cracked products.
Therefore, in the Steam Catalytic Cracking process are implicit reactions
and principles of:
1) Steam Catalytic Reforming.
2) Thermal Cracking.
3) Unsupported Ultradispersed Catalysis.
In order to have a better understanding on how the combination of these
three areas allowed the development of the Selective Catalytic Steam Cracking
(SCSC), as a new process which supports Aquaconversion technology,
developed by PDVSA-Intevep for the upgrading of heavy oil and heavy oil
fractions, it is necessary to explain briefly the principles on which each one of
them are based. Thermal Cracking already has been extensively commented on
previous section, a review of steam catalytic reforming and unsupported
ultradispersed catalyst is presented in the following sections.
34
2.3.1 Thermal Catalytic Steam Reforming
The reaction of light hydrocarbons with steam in the presence of catalysts is
known as Catalytic Steam Reforming [74]. This process has been used since
1930 for obtaining hydrogen, and in the synthesis gas elaboration (CO + H2O).
Depending on the transformation in the hydrocarbon molecule, the process can
be classified in two categories namely total and selective [75].
The feed in total catalytic steam reforming (generally natural gas and/or
naphtha), is totally gasified to hydrogen and carbon monoxide according to the
following reaction:
CnHm + 2nH2O nCO2 + (2n+m/2)H2
In selective catalytic steam reforming only part of the hydrocarbon
(generally an alkylaromatic) is transformed by steam to H2, CO and also to
aromatic compounds with smaller number of carbon atoms than those of the
feed, according to the following reaction:
CnHm + H2O CxHy + gas(H2, CO2, …) x < n
Besides the previous reactions, reactions like steam displacement (steam
shift reaction) also take place, favorably in a temperature range of between
300°C and 600°C, and methane formation through CO Hydrogenation:
CO + H2O CO2 + H2 Water Gas Shift Reaction
CO + 3 H2 CH4 + H2O CO Hydrogenation
The selective rupture of C-C bonds by the steam requires that well
differentiated C-C bonds exist in the molecule. In this sense, the alkyl groups in
aromatic hydrocarbons are primarily the ones that are transformed by the steam.
It is for this reason that the selective catalytic reforming is also well known like
steam dealkylation.
35
2.3.2 Unsupported Ultradispersed Catalysis
Unsupported catalysts offer an alternative technology to conventional
supported catalysts, with benefits of reducing pore plugging problems, increasing
the accessibility of active sites to the reactants with bigger molecular size, and
allowing the catalyst to be recovered, recycled or discarded. The aims of this
catalysis are:
a) To generate in situ the active form of the catalyst, through the thermal
decomposition of the substances that contain the catalytically active
metal, also well-known as catalytic precursors, and to produce fine solid
particles suspended in the oil media augmenting its accessibility and its
reactivity,
b) To obtain high levels of metallic particles dispersion in the residue, and
c) To maintain a high catalytic activity. Furthermore, the reduction of the
catalyst size to the micro or/and nano scale, will enhance the metal
dispersion in the media, producing an ultradispersed catalyst.
Through the unsupported dispersed catalysis it is claimed that instead of
the hydrocarbon molecule approaching the active site of a typical supported
catalyst, it is more efficient for the dispersed catalyst to approach the molecule.
This better contact between reactants and active phase increases the possibility
of catalytic reaction resulting in a higher conversion of residuals to light products
and reduction of the amount of carbonaceous products.
Many methods are known to produce ultradispersed catalysts such as:
solutions spray drying, electrical arc of metallic wires containing the active
metals, ultra fine grinding of preformed catalysts, preparation of supported
catalysts in which the active phases are impregnated and, dispersed as nano
particles and ulterior destruction of the support, micro emulsion in which
precursors of the active phases are dissolved in the water phase. Previous
studies have found the micro emulsions method offers significant technological
36
advantages over the other methods [62, 76], in versatility of incorporation of the
catalyst to the process, preparation control, particle size determination, multi-
component catalyst production, etc. In the micro-emulsion method, the catalyst
chemical formulation starts with the definition of the precursor salts of the active
components of the formulation. The stabilization in solution of all the chemicals
required in the aqueous solution used to prepare the micro emulsion is essential.
In the unsupported dispersed catalysis for the selective catalytic steam
cracking, transition metals like Mo, V, Ni, Co, etc. are used for hydrogenation,
and alkaline metals like K or Na are used with the hydrolizants proposes.
The technical literature related to unsupported dispersed catalysis, although
not very abundant, has been published in the last 25 years. At the beginning of
the nineties several highly active dispersed catalysts were developed for heavy
crude oil hydroconversion [77,78,79], most prepared starting from naphthenates,
acetylacetonates and carboxylic acids as catalytic precursors.
2.3.3 Selective Catalytic Steam Cracking
Selective Catalytic Steam Cracking (SCSC) represents a new way to
reduce viscosity of heavy crude oils and vacuum residue and to increase the
yield of distillates. The unique features of SCSC are the inclusion of steam and
ultradispersed catalysts to the thermal process (visbreaking) in order to cause
moderate hydrogen incorporation from water to the thermal products. This novel
process allows refiners to reach higher conversion levels than those achieved in
visbreaking, with the additional advantage of still producing a stable converted
product. This process was developed by PDVSA-INTEVEP and it supports
Aquaconversion technology [5,6,80].
During visbreaking a great number of reactions take place: dealkylation,
aromatic condensation, cracking, dehydrogenation, among others. Although
these reactions occur also in SCSC, some of them, such as dehydrogenation of
naphtenic rings, are likely to be reduced, while dealkylation increases in
37
comparison to visbreaking. The global effect in visbreaking is that feedstock
undergoes chemical changes in such extension, that they could cause alteration
of the equilibrium conditions between asphaltenes and resins, inducing aromatic
condensation with asphaltene formation, flocculation and finally precipitation.
Although saturation reactions are easily promoted by catalysts and elevated
hydrogen pressures, neither of these conditions exists in a conventional
visbreaking unit. The aromatic condensation dilemma was solved by using a
novel oil-soluble, dual-catalyst system, which, at conventional visbreaking
processing conditions, converts water into hydrogen and then inserts the
hydrogen at the critical point in the thermal cracking reaction sequence where the
asphaltene forming condensation reactions occur.
This hydrogen-transfer mechanism inhibits aromatic condensation and
produces a more stable visbroken product, which has a higher hydrogen content
and lower asphaltene and Conradson carbon contents than the product from a
conventional visbreaking unit.
2.3.3.1 Reaction Mechanism
SCSC reaction mechanism proceed by the unique interaction between the
two non-noble metal catalysts. The first catalyst enhances the dissociation of
water into hydrogen and oxygen free radicals. The highly reactive hydrogen free
radicals that are formed accelerate the thermal cracking rates of the paraffinic
components of the feedstock and stabilize the resulting thermal products by
saturating olefinic free radicals.
The second catalyst minimizes the condensation reactions by promoting the
addition of hydrogen to the aromatic free radical. The result is the formation of a
smaller aromatic component as well as additional hydrogen free radicals and
carbon dioxide.
This entire reaction sequence effectively reduces formation of coke
precursors and the undesirable aromatic-condensation reactions so the refiner
38
can benefit from the viscosity reduction associated with higher visbreaking
conversion and still produce a stable visbroken product. The sequence of the
reaction mechanism is presented next:
(1) R-Rn’ R + Rn’ Thermal Cracking
(2) H2O CAT H• + OH• Catalytic Dissociation of Water
(3) R• + Rn’• + 2H• CAT R-H + Rn’-H Saturation of organic free
radicals by hydrogen free radicals.
(4) Rn’• + 2OH• CAT Rn-1’ + CO2 + H2 Oxidation / Reforming
(5) Rn’• + R• Rn’-Rn’ + R-R Condensation
The highest activation energy value in the previous reactions has been
found in the range of 40-60 kcal/mol and corresponds to thermal cracking
reaction [81]. For this reason, equation (1) is considered the rate-limiting step in
this mechanism.
2.3.3.2 UOP / PDVSA-Intevep / Foster Wheeler: Aquaconversion
In 1996 UOP and Foster Wheeler formed an alliance with Intevep, the
research and technology support center of Petroleos de Venezuela S.A.
(PDVSA), to commercialize and further develop its Aquaconversion catalytic
hydrovisbreaking process.
In the Aquaconversion process, which can be used to replace or modify
both coil-type and soaker visbreakers to achieve higher conversion levels, the
feed, steam and an oil-soluble dual catalyst system are sent to a heater and
reactor. In these units, thermal cracking takes place resulting in the formation of
olefinic free radicals. These radicals are saturated by highly reactive hydrogen
free radicals that are formed when the catalyst system promotes water
dissociation. The catalyst system also promotes hydrogen addition to aromatic
free radicals limiting the occurrence of the reactions that form asphaltenes and
39
cause stability problems. After cracking and hydrogenation, the reactor effluent is
quenched before entering a fractionator. A recovery system separates and
recycles the catalyst from the fractionator bottoms stream.
Figure 2-11 Aquaconversion technology[47]
According to the developers of Aquaconversion, the use of the catalyst
system allows refiners to achieve greater conversion and viscosity reduction than
in conventional visbreaking processes. This includes up to 40 wt% conversion of
500°C + residues without the occurrence of product stability problems [82]
Consequently, refiners can benefit from lower fuel oil yields as well as reduced
cutter stock usage and higher distillate production.
Compared to a conventional visbreaker, a new Aquaconversion unit will
cost about 30% more because of the use of the catalyst addition and recovery
sections. The cost to revamp a 20k/bpsd visbreaker to Aquaconversion
technology is estimated at US$15-$20MM. The payback period for a revamp of
this kind would range from one to one and a half years [28].
40
2.4 The Thermodynamic Equilibrium of Water/Heavy Hydrocarbons Reactions
The thermodynamic application in the organic reactions has been
particularly useful in predicting the equilibrium reaction and evaluating the
thermodynamic feasibility of a given process. The thermodynamic approach for
the equilibrium reaction and for the spontaneity of a process can be summarized
in terms of Gibbs energy change,
ΔG = -R x T x Ln(K)
ΔG= Change of Gibbs free energy R= Universal gas constant T= Temperature K= Equilibrium constant
If ΔG is negative, the reaction can be carried out spontaneously. If ΔG is
bigger than zero but smaller than 10 kcal the process is not possible although it
could be investigated. If ΔG is bigger than 10 kcal, the reaction is not possible
thermodynamically unless it is carried out at extreme conditions.
Vacuum residue can be assumed to be a colloidal suspension of
asphaltenes in an oily medium formed by aliphatic and aromatic hydrocarbons
and resins. Asphaltene is known as the non-soluble fraction of the crude oil in
organic solvents (n-pentane or n-heptane). These compounds differ in weight
and composition depending on the origin of the crude oil, the type of residue from
which it has been precipitated, the solvent used and the precipitation conditions.
Due to the large molecular complexity of this compound, it has not been possible
to formulate an exact molecular structure in spite of the different methods that
have been attempted, resulting in complex molecular matrixes consisting of poly-
condensed aromatics in which aliphatic chains are bonded with different
alkylaromatics and aromatic structures fundamentally with sulfur and nitrogen.
For this reason, the thermodynamic evaluation of asphaltene molecular rupture is
41
carried out considering average molecular structures that better represent the
chemical composition of the asphaltene.
In her study about the progressive rupture of an asphaltene molecule and
their thermodynamic feasibility, Salerno, G [83] evaluated the possible different
ruptures in an asphaltene model molecule (MORICHAL1) in the presence and in
the absence of steam, and in a range of temperature between 600 °K and 850
°K. The molecular model structure is presented in Figure 2-12.
Figure 2-12 Molecular structure of Morichal1 (C100H107N2S2O)
The first rupture reaction for this molecule can be written as:
C100H107N2S2O + 65H2O C20H15SN + C28H14NS + C20H20 + 32CO2 + SO2 + 94H2
According to her results, Salerno concluded that the first rupture reactions
are thermodynamically feasible in the whole evaluated range of temperature and
besides, a great amount of energy is necessary for breaking the multiple C-C
bonds in order to produce H2 and CO2, for that reason this reaction is considered
highly endothermic.
Salerno also evaluated final ruptures reactions the same as the first
ruptures, the results indicate that this reactions are also feasible
thermodynamically.
42
2.5 Remarks
The literature review provides the background to put into context the
research demarche of this thesis.
Thermal cracking has been introduced both in the mechanistic aspects
needed to interpret the results of our tests, and the characteristics of the
industrial setups that provide support to the design of our pilot plant as they will
be introduced in the experimental section.
In particular this work focused on the use of conditions such as temperature
and residence time similar to the ones that soaker type of visbreaking would use.
In this work is presented the relevant aspects found within the limited open
information available for steam catalytic processing of heavy oil fractions. This
type of processing was introduced as a novelty at the end of the last decade and
claims the use of setups similar to visbreaking thus the suitability of our pilot plant
to incorporate a comparative study of said process with thermal cracking. The
results and discussion section will make use of some of the concepts and papers
here disclosed.
43
Chapter 3: Experimental
3.1 Experimental Setup
3.1.1 Pilot Plants: Basic Considerations.
There is not a commonly accepted definition of what constitutes a pilot
plant. Different organizations tend to define what a pilot plant is based on size,
purpose, operating group and so on. For the intentions of this research, a pilot
plant is defined as specialized equipment intended primarily for research and
development purposes.
A more general definition for a pilot plant could be given as a tool intended
to allow investigation of a process or process problem on a manageable scale in
a realistic manner in a timely fashion [58]. The word tool implies that all pilot plant
operations are a means to an end and not an end in themselves. Manageable
scale implies a limited commitment of time, money and resources to achieve an
end. This requires some experience for the expenditure of time and effort to
define objectives and the desired end before the program begins. Realistic
manner implies a scale large enough to include all effects that may be important.
To some degree this conflicts with the desire to maintain a manageable scale.
The idea is to try to identify the smallest scale in which the data generated is
significant enough to the area being investigated. Timely implies some
contemplated endpoint for the program. While a basic research program may be
relatively open-ended in that it strives to achieve a better understanding of the
basic theory underlining a process, most pilot plant operations are more focused
and goal-oriented.
Pilot plant work is undertaken for a variety of different goals. Some of the
most common are:
• Performing process development
• Demonstrating technical feasibility
• Providing technical services
• Supporting market development.
44
In our particular case and because the built pilot plant is involved in an
investigation about heavy oil upgrading, our main goals are related with
performing process development and to demonstrate technical feasibility.
Process development can include process optimization, patent development,
patent defence, research and development. In general the pilot plant needs to be
capable of evaluating competing processes and determining which is the most
efficient or effective.
On the basis of size, the generally accepted classification of pilot plants is
as follow:
1. Laboratory-scale, benchtop test plants or micro units: These are pilot plants
that generally fit on a benchtop or inside a small laboratory hood. In general
they are in the range of 0.5 to 1.0 m2 and use 1/16” to ¼” in tubing for piping.
Traditionally totally manual and continuously attended, many new versions
are automated and designed to run continuously and unattended. Two of the
plants currently existing in the Catalysis group where this thesis was
performed.
2. Integrated pilot plants or research-scale pilot plants: This remains the
workhorse of many chemical processes industries with R&D organizations.
They may vary in size from several frames or pallets to a unit occupying a
small building. In general they are in the range of 2 to 14 m2 and use ¼” to 1”
in tubing. They are usually automated and may frequently be designed for
unattended operations.
3. Demonstration units, semiworks units or prototype units: These units are
designed to operate at the lower end of plant scale. They are very large, in
the order of 900 m2 or more and are built with commercial pipe sizes typically
in the range of 1” to 8” in. They resemble the plant in automation and
operation.
It may be tempting to design and built a pilot plant using the same
specifications as for a full-scale commercial facility; however, doing so can add
to the cost and schedule of the project. It can also have a severe impact on the
45
quality and operability of the pilot plant, and may in the extreme case, cause the
project to be abandoned [59]. Often, the uniqueness of pilot plants and factors
such as cost, space minimization, time restrictions and pilot plant life are not
carefully considered. The principle of keeping the pilot plant as simple as
possible is important without falling in the trap of blinding pass on basic
specifications and common safety rules.
Factor Commercial Scale Pilot Scale
Target Continuous generation of on specification products
Process knowledge and understanding. Operational
understanding. Scale-up data. Product sampling.
Scale Ton/h kg/h
Operation Continuous, maximizing up-time Operations in campaigns depending
on operational stability and/or sampling requirements.
Design life Tens of years 1-10 yrs
Maintenance As much as possible during operation Between campaigns
Operational Mode Steady-state Chasing steady-state
Data Acquisition and Control
As required to obtain steady-state To obtain steady-state and to collect
the necessary process data for scaling-up
Operating Temperature and Pressure
Commercially optimal conditions Beyond commercially optimal conditions to establish optimum
Source of Design Data Pilot plant Laboratory data, simulations and
experience.
Capital Project Timescale Several years 1 yr
Need for Operational Flexibility
Modest. Not mandatory Considerable
Table 3-1 Comparison of commercial and typical pilot-scale operations
Table 3-1 shows a comparison of commercial and typical pilot scale
operations and requirements to have an idea of the differences between both
46
scales [60]. These issues have to be taking into consideration at the moment to
plan the construction of a pilot plant, to minimize the risk of failure of the pilot
plant performance and to take maximum advantage of the unit in the
accomplishment of its objectives. In our specific case, the pilot plant design and
construction must attend to additional requirements of operability within a
university learning environment, which pushes into a size reduction and extra
risk minimization considerations.
3.1.2 Process Overview
In order to reach the objectives of this thesis it was required to build a setup
to simulate a commercial visbreaking unit at research scale to evaluate the
reactivity of Athabasca VGO under thermal, thermal-steam and thermal-steam-
catalytic conditions.
The research scale pilot plant was built with the purpose of carrying out
experimental evaluations focusing on heavy hydrocarbon feedstock. It is a plant
conceived to carry out studies under a wide range of variables and conditions, It
also has an adequate size and accessibility that allows quick implementation of
eventual modifications required to improve and expedite the results of the studies
or to switch the unit to operate under different modes as it will be described soon.
The thermal and catalytic steam cracking bench plant unit is shown in
Figure 3-1. In the feed section the feedstock without or with ultra-dispersed
catalyst is placed in the feed tank and is fed into the unit where a positive
displacement metering pump can deliver up to 850 cc/h as maximum flow at 450
psig. Nitrogen or helium are used before each experiment for initial flushing of
the unit and to create an inert atmosphere to ensure no oxygen is present inside
and to adjust the pressure of the system, setting the back pressure valve to the
value required for the test. Nitrogen is also used after each run as a purge gas to
flush and drain the liquid remaining in the pipe lines. The feed is passed through
a preheat section where the temperature can be raised to the range of 200 to
350°C , close to the reactor temperature reducing in that way the heat load
47
required in the reactor. A water pump with a flow range of 0.001 – 12.0 ml/min
and a heater were installed just before the reactor inlet to generate and provide
steam in the case that it is required.
Figure 3-1 Bench scale plant for reactivity and catalysts evaluation
Two tubular up flow reactors were installed in the reaction zone and their
specifications can vary according to the requirements of the research, for this
case both with 100 cc of capacity. Just before and after the reaction zone two
manually actuated 3-way valves allow aligning the reactor required for the
experiment. A fixed bed can be also placed in both reactors if that is required by
the test conditions. Once at the inlet to the reactor, temperature of the stream is
increased to that of the test (350 – 500°C), assuming an isothermal operation
throughout the length of the reactor.
The effluent coming out of the reactor goes to the collection zone, reaching
first a hot separator, where the temperature of the heavy product can be
controlled in the range of room temperature to 350°C. The non-condensed light
products and steam coming from the reactor and hot separator are sent through
a water-cooled single tube heat exchanger where the condensable light fraction
Heat Exchanger 225 ºC
FeedPump
FeedTank
PreHeater200-350 ºC
WaterPump
WaterTank
ReactionSystem
350-500 ºC
Hot Separator50 - 350 ºC
To GC and flow meter
Back PressureValve
WaterHeater
300 C - 450 ºC
ColdSeparator
25 ºC
In water Out water
.
.
Heat Exchanger 225 ºC
FeedPump
FeedTank
PreHeater200-350 ºC
WaterPump
WaterTank
ReactionSystem
350-500 ºC
Hot Separator50 - 350 ºC
To GC and flow meter
Back PressureValve
WaterHeater
300 C - 450 ºC
ColdSeparator
25 ºC
In water Out water
.
.
48
can be obtained. Non-condensable vapors (mainly C1 – C5 hydrocarbons, H2,
CO, CO2 and traces of H2S) and the condensed fraction pass through a back
pressure valve, which maintain a constant pressure in the unit, and then are
directed to the cold separator where the condensed fraction is collected. Non-
condensable vapors leaving the cold separator are passed through the gas flow
meter (wet test meter) and sent to the gas chromatograph for compositional
analysis or introduced into the KOH trap to remove H2S traces in the gas stream.
The heavy products and light products condensed at their respective tanks are
drained to collecting vessels after mass balance periods without affecting the
pressure of the system. A detailed flow process scheme can be seen in Appendix
B. Table 3-2 shows the range applied for the main operational variables
considered in the unit design.
Feed flow (mL/h)
H2O flow (mL/h)
Pressure (psig)
Preheating temperature (ºC)
Reaction temperature (ºC)
VHSV (h-1)
30 – 850
0 – 720
0 – 450
200 – 350
350 – 500
0.3 – 8.5
Table 3-2 Design specifications for the Catalytic Steam Cracking unit
The unit was built on a metallic structure in the form of a table 90x100 cm with
two levels. In the lower level were placed the feed and water pumps and the
control and electricity boxes. The tank bases were welded to two of the sides of
the structure, so the feed zone tanks were on one side and the collection zone
tanks were on the other, the rest of the equipment was placed on the top of the
table. In addition and although this research thesis was focused on evaluations of
vacuum gasoil reactivity, the plant was built providing heating capabilities along
the tube lines and tanks where liquid feed could be flowing or accumulated thus
49
allowing processing of a wide range of heavy oils (from light gas oils to vacuum
residues and bitumen).
This unit was also conceived for processing versatility, allowing testing
reactivity of heavy hydrocarbons under Thermal, Thermal-Steam and Thermal-
Steam-Catalytic (dispersed and fixed bed) cracking conditions. Later on
hydroprocessing capabilities were also incorporated, their description being out
of the scope of this thesis.
3.1.3 Feed and Pre-heater Section
This zone consists of a feed tank, a wash oil tank, feed pump, pre-heater
system and steam injection pump. The process lines were built with 3/8” tubing
316 SS. The feed tank has a capacity of 10 l that allows a continuous operation
in the unit between 9 hours and 10 days depending on the feed flow rate. The
detailed scheme of the feed tank is shown in Figure 3-2. The feed tank was
manufactured at the department workshop using 316 stainless steel having
dimensions of 6.63” diameter and 15” length.
Figure 3-2 Feed tank specifications
0,134” min
6,625”
3/8NPTx3/8Tub
1/4” NPTx1/4” Tub8,25” min
0,8”0,39”
15”’
V=10 l
Cover
N° of bolts = 6Bolt diameter = 0,32” min
Cover
N° of bolts = 6Bolt diameter = 0,32” min
50
A concave shell of the same material was welded to the bottom with a
connection for draining. The tank was wrapped with a heating tape and isolated
with ceramic fibre supplied by Thermal Ceramics, to keep them warm.
Temperature is set and controlled according to the feed used, in a range of 25 to
120ºC, in order to maintain the viscosity of the feedstock in a value under 400 cP,
for which the feed remains fluid enough and problems associated with difficulty of
pumping are eliminated. Nitrogen can be injected on the top of the tank once it is
filled with the feedstock and pressure can be kept between 6 and 15 psig to
maintain an inert atmosphere in contact with the feed.
Feed flow is measured using a burette system, where the volume displaced
is read in an interval of time given. This method should be improved especially if
a heavy feedstock is used since it can be too viscous and sticky and can plug the
burette. For pumping, a gear pump was used in order to reduce to a minimum
any pulse or instability of the flow which could cause small intervals of stopping
flow representing a high risk in the hot areas because of coke formation and
plugging. The feed pump installed is a Precision Metering Pump Zenith H-9000,
0.3 cc/rev, maximum operation temperature: 120ºC, maximum discharge
pressure: 1500psig.
After the feed pump, the system is protected using spring-type pressure
relief valve set at 750 psig. If for any reason the process line is plugged
downstream, this valve will open and will discharge the flow to the feed tank,
avoiding rupturing the line by over pressure or causing a bigger damage to the
unit.
A feed preheat section with two heating units was installed to reduce the
heat load to the reactor. The pre-heaters used were recovered from a pre-
existing unit already dissembled, and tested with vacuum gasoil to ensure the
pumped feed was able to reach the temperature of 380ºC. Testing demonstrated
that temperature was reached without any problem requiring a low power duty
(only 20% of the maximum capacity).
51
All the process lines were wrapped with heating tapes and temperature
controlled, in order to maintain temperature profile along the line. Heating tapes
were covered with ceramic fibre to avoid heating loses and for safety issues.
3.1.4 Reactor
The cracking process takes place in a tubular up flow plug reactor. The
reaction system consists of an arrangement of two vessels in which either one of
them can be aligned to the process. The vessel selected as the process reactor
is aligned using a 3-way valve installed at the inlet and outlet of the reaction
system. The fact of having two reactors available in the unit allows more
versatility in the operation, one of them can be used as spare in the case that the
other one is coke plugged, or can also be filled with a catalytic fixed bed which
can be switched from the empty tube reactor during the operation.
In this work three vessels were constructed to be used as reactors, two
were identical in dimensions, one to be installed in the pilot plant and the other
kept as a spare. A third reactor, longer and thinner than the two previous was
also installed in the unit, running parallel to the first described reactor. Even
though it was not under the scope of this research, the intention of having two
different reactor diameters was to evaluate in the future the L/D effect, which is
essential to determine which flow regime, plug flow or mixing can be more
effective to the process(es) under study. When sizing the vessels two factors
were taken into consideration:
a) Spatial velocity: Defined as the ratio of the volumetric feed rate to the
volume of the reactor, the volume of the vessel had to match with the feed pump
specifications in order to be able to run the experiment at a spatial velocity in the
range of 0.3 to 8 hr-1 required for this and future research.
b) L/D ratio bigger than 10, in order to minimize fluid dynamic problems and
to ensure that any differential of feed volume entering the reactor will run along
the reactor volume with minimal back mixing (plug flow). One reactor with a lower
52
L/D (wider diameter) thus ensures progression toward back mixing to check the
effect of flow regime on the catalytic process. These features of the pilot plant
weren’t tested during the part of research planned for this MSc Thesis.
c) Tubing diameter available in the market: This was related to commercial
availability of 316 stainless steel tubing to be used for constructing the vessel,
which also must be able to handle a maximum pressure of 850 psi and a
maximum temperature of 500ºC.
Appendix A shows results of some calculations developed in a spread sheet
with commercial 316 stainless steel piping diameters and the range of spatial
velocities requested. According to these results, the effective volume for both
sizes of vessels was estimated to be 100 cc. For the shorter reactors was chosen
¾” outer diameter and calculated 50 cm of length, whereas for the larger reactor
was chosen to have a 5/8” outer diameter with 86.4 cm of length.
Figure 3-3 Detail of probes installed inside the reactors.
Calculations were made considering a thermocouple probe of 1/8” diameter
inside each reactor in order to monitor the internal temperature (custom made
and provide by OMEGA). These customized probes were made of stainless steel
316SS sheath with 1/8” diameter and with a grounded junction. The
thermocouples installed inside the probe were type K, which has a maximum
29
cm
b b b
29
cm
29
cm
b=9.7 cm
29
cm
b b b
29
cm
29
cm
29
cm
b b b
29
cm
29
cm
b=9.7 cm
25 cm
a a a
25 cm
a=8.3 cm
25 cm
a a a
25 cm
25 cm
a a a
25 cm
a=8.3 cm
53
service temperature of 870ºC. Figure 3-3 shows detail of the probes placed
inside the reactors.
Figure 3-4 Detail of the reactors zones with corresponding probe inside
Each reactor was divided depending on their length, into two or three equal
sections. The shorter one was divided in two sections of 25 cm, and the larger
one is divided in three sections of about 29 cm (see Figure 3-4 for details). Each
reactor section was provided with a set of two thermocouples installed in the
probe for monitoring the internal temperature of each section, and was also
wrapped externally with their own heating tape in order to individually control the
heat supplied to each section.
25 cm
25cm
a
a
a a=8.3 cm
25 cm
25cm
a
a
a
25 cm
25cm
a
a
a a=8.3 cm29 cm
29 cm
29 cm
b
b
bb=9.7 cm
29 cm
29 cm
29 cm
b
b
b
29 cm
29 cm
29 cm
b
b
bb=9.7 cm
54
To determine the best heating tape length to be installed on each reactor
section, the following equation is applied [62]:
Where:
C = length of the tubing D = diameter of the tubing W = wide of the tape plus the space between coils T = thickness of the tape.
3.1.5 Products Collection
Liquid product, and gas product and unreacted feedstock exit the reactor
and this stream is directed to the products collection zone, which consists of two
hot separator tanks, one back pressure regulator valve, one condenser and a two
cold separator vessels.
The hot separator tanks consist of a double ended sample cylinder of 2250
cc capacity and body material of 304L stainless steel, supplied by Swagelok®,
product code 304L-HDF8-2250. These tanks can handle a maximum pressure of
950 psig at an operation temperature of the 400°C.
Figure 3-5 shows an example of double ended cylinder similar to the two
installed in the unit.
This set of two hot separator tanks were conceived to have one aligned and
receiving products during any running condition (starting up, stabilization period,
mass balance) while the other tank is being drained for weighting and analysis
(for mass balance purposes) of the product of the previous condition; once
drained it is re-pressurized at the current operation condition in order to get it
back aligned to the process and ready to receive the product of the next mass
balance condition. The alignment of the tank to the process line is carried out
[ ]22 )(14.3 TDWWCL +×+×=
55
with a manually actuated 3-way valve. The top end of the tank has an
arrangement of fittings and concentric tubing which allows through the internal
tubing the entrance of the mixture of gas and liquid coming from the reactor, and
by the external tubing the gas exit from the tank to the vapors condenser.
Figure 3-5 Double ended cylinder used as heavy product tank
From the hot separator tanks, vapors are directed to the condenser, which
consists of a “double pipe” counter flow heat exchanger where cooling water
flows through the external pipe and vapors coming from the heavy product tank
run through the internal tube in which any condensable light hydrocarbon and
steam are condensed at 25-30 °C. Figure 3-6 shows the heat exchanger
configuration used as condenser.
Figure 3-6 Double pipe heat exchanger configuration used as condenser
In order to keep the same dimension for the process line, 3/8” OD was used
for internal tubing. After appropriate calculations for double pipe condensers it
Hot product
Water inWater out
Hot product
Water inWater out
56
was determined for the external pipe of the condenser a diameter of 2.5 cm and
24 cm of length, estimating to reduce the temperature of vapors from a maximum
of 170 °C to 26 °C. Appendix D shows a calculus algorithm developed for
condenser estimations as well as assumed values of variables assumed and
properties used.
The operating pressure was controlled by a back pressure valve model
JBP-2-S-M-G-1-6-F-5, body of 316 stainless steel, spring range 5 - 500 psig,
supplied by Tundra Boiler & Instrumentation. This valve is manually set to
regulate the pressure required within the range it allows. Vapors from the heavy
product tank pass through the back pressure to drop to atmospheric pressure.
The mixture of liquids and non-condensable vapors from the back pressure
valve pass to the cold separators where liquid product is retained. The cold
separators consist in two plastic containers of 250 cc of capacity. The alignment
of each container to the process line is done with a manually actuated 3-way
valve, applying the same operational criteria described for the hot separator
tanks. Gases not condensed coming out of the plastic container were then split
into two streams, one stream was passed through the gas chromatographer and
then rejoined with the second stream to be blown through a sodium hydroxide
solution to absorb traces of H2S produced in the process. Finally the rejoined gas
stream was measured using a Shinagawa Precision Wet Gas Meter model W-
NK-2.5B; max flow 25 lpm, five litres per revolution.
3.1.6 Process Line
The process line interconnects all equipment from the feed tank to heavy
product tanks and light product vessels and was built using 3/8” seamless tubing
supplied by Swagelok®, body material of 316L Stainless Steel and 0.035” of
nominal wall thickness and Swagelok® fitting connections of the same material.
Table 3-3 shows length and volumes estimate for each process line section from
the feed tank to a 3-way valve located just before the hot separator tanks,
including pre-heaters and reactor pipes. Total liquid dead volume was estimated
57
to be approx. 440 cc and later measured flooded with liquid, these experimental
measures confirmed this value.
Section Length (cm) Diameter (cm) Volume (cm3) From feed tank to pump 75.0 0.77 35.4 From pump to pre-heater 1 143.0 0.77 67.4 Pre-heater 1 31.0 1.26 38.5 From Pre-heater 1 to pre-heater 2 76.0 0.77 35.8 Pre-heater 2 62.0 1.26 77.0 From Pre-heater 2 to Reactor 116.0 0.77 54.7 Reactor 49.0 -- 100.0 From Reactor to three way valve 75.0 0.77 35.4
Table 3-3 Length and volume specifications of flooded process line sections
3.1.7 Control and Data Acquisition
A computer based data acquisition and control system was set up to collect
the temperature and pressure data as well as to control the heating tapes and
pre-heaters for the Unit.
Three points of pressure in the unit are measured by electronic MPI
pressure transmitter model code MT104P1MXS-3/8”, 3/8” NPT connection, 1000
psi / 4-20 mA supplied by MorHEAT Inc., and a total of thirty eight points of
temperature measurement were inserted using standard type K thermocouples
(supplied by Cole Palmer). Twenty thermocouples are placed along the process
line to measure internal fluid temperature (TI), including the four thermocouples
placed in the probe to measure the temperature in the core of the reactor. There
are two 1 kWatt pre-heaters, and sixteen heating tapes (of various wattages) all
of them wired to 120 VAC via solid state relays (one for each heater). Eighteen
thermocouples are placed to measure the temperature of the heaters, fourteen
for the heating tapes and two for the pre-heaters, which provide the feedback for
temperature control loops (TIC). The pre-heaters and heating tapes are computer
controlled, and each one has its own PID control loop (with unique tuning
parameters). The tuning parameters are kept in a spread sheet compatible file
58
and can be modified during run time via pull down menu selection. If necessary,
the PID tuning parameters can be modified adjusting the proportional, integral
and derivative setting. Appendix C describes the tags assigned and the location
of the pressure indicators (PI), temperature indicators (TI) and temperature
indicator and controller (TIC) installed in the unit.
The computer measures all thermocouple data approximately every 5
second and shows it on the computer display. The computer also logs the data
into a spread sheet compatible file at one minute intervals.
The data acquisition system is comprised of National Instruments
Labview™ 7.1 software and National Instruments hardware. The following list is
a brief description of the hardware used and its purpose in the pilot plant:
• FP-1601 Ethernet Module – provides communication between the data
acquisition hardware and the controlling computer.
• FP-TC-120 Thermocouple Input Module – 8 Channel – measurement
of the thermocouples.
• FP-AI-100 Analog Input Module – 8 Channel – general analog signal
measurement (pressure transducers, etc).
• FP-RLY-420 / PCI-6602 8 Channel SPST Relay Output Module / eight
32-bit counter/timers. Control to turn the various heaters on or off
through solid state relays.
• FP-TB-1 Universal Terminal Base, Screw Terminals – termination for
the above modules to their respective signals.
• FP-PS-4 Power Supply, 24VDC, Universal Power Input Din Rail Mount
– power supply to drive the data acquisition modules.
As the PCI-6602 can not directly drive the heaters, Solid State Relays
(Crydom 2425) with 5VDC control input and 240VAC @ 25Amps switching
capabilities are used. Preventively, the solid state relays are mounted on heat
sinks (Crydom HS2) to dissipate excess heat. Each circuit is fuse protected and
has an indicator lamp to show circuit state.
59
3.2 Reactivity Tests
3.2.1 Experimental Plan
The reactivity evaluation of the Athabasca VGO was divided into three steps.
The operating conditions were established according to evaluations developed
by Pereira et al for vacuum gasoil and established in the US patent Nº 6,030,522
[55].
First set of experiments:
a) Thermal Cracking evaluation: running thermal-cracking in the pilot plant.
The process conditions selected to test include two pressure levels, 160
and 260 psig; a feedstock flow of 200 cc/h; temperatures of 400°C, 410°C,
420°C and 430°C and spatial velocity of 2 h-1.
Second set:
b) Steam-Thermal Cracking evaluation: running at similar process conditions
as thermal cracking evaluation but now studying the effect of injecting
steam just upstream of the reactor. The process conditions include the
same pressures as thermal cracking, steam flow of 10 cc/h and
temperatures of 410°C , 420°C and 430°C . Spatial velocity of 2 h-1.
Third set:
c) Evaluating Catalytic Steam Reactions of Athabasca VGO in the presence
of both ultradispersed catalysts and fixed bed catalysts, studying the effect
of variables and changes in the catalytic matrix. The process conditions
include a pressure of 260 psig, water to hydrocarbon ratio of 0.05,
temperatures of 430°C and 440°C. The first catalytic matrix was prepared
based on nickel as the transition metal and potassium as the alkali metal
(Matrix A). The second catalytic matrix, Matrix B, is similar to Matrix A but
now calcium as alkali-earth metal is also added. Spatial velocities of 1 and
2 h-1 for ultradispersed and 2 h-1 for fixed bed catalysts were used.
60
Figure 3-7depicts a summary of different evaluations to be carried out and
the parameters and variables considered in order to accomplish this study.
Figure 3-7 Graphic summary of the experimental plan
3.2.2 Operational Criteria
Getting a stable operation of the pilot plant to test a set of experimental
conditions requires executing a sequence of required events in a minimum length
of time so that this event can be completed without any problem.
In this sense, the starting and adjustment of flows, pressure and
temperature conditions take a minimum time of three hours. In order to reach
stationary conditions and to ensure that the fluid coming out of the reactor and
the collected products are stable and they do not change their characteristics
with time, which means steady state has been reached. A rule of thumb
indicating that at least 3 times the volume of the reactor must be flowed through it
before starting any mass balance period was used. The required time to
complete this step depends on the feed flow or spatial velocity set in the
experiment. For the evaluations to be developed in this research, a 3 hours
Pressure, psiTemperature, °CH2O/HC, wt/wtCatalytic System
Thermal Cracking Steam Cracking
160 / 260400/410/420
--
160 / 260410/420/430
0.05-
260430/440
0.05Matrix A/Matrix BTransition/Alkali
2604300.05
Matrix BTransition/Alkali
(metal)(metal)
Non-CatalystTest
CatalystTest
UDCatalytic Steam
Cracking
Fixed BedCatalytic Steam
Cracking
Pressure, psiTemperature, °CH2O/HC, wt/wtCatalytic System
Thermal Cracking Steam Cracking
160 / 260400/410/420
--
160 / 260410/420/430
0.05-
260430/440
0.05Matrix A/Matrix BTransition/Alkali
2604300.05
Matrix BTransition/Alkali
(metal)(metal)
Non-CatalystTest
CatalystTest
UDCatalytic Steam
Cracking
Fixed BedCatalytic Steam
Cracking
61
stabilization period has been considered which implies from three to 6 times the
reactor volume was flowed through before mass balances were started.
The duration of the mass balance will depend on the minimum volume
required for the laboratory analyses to be carried out and to keep some quantity
as witness for repetition or verification purposes. Mass balance duration between
1,5 to 2 hours to collect an amount of 300 cc to 400 cc of product can be
considered sufficient. More than one mass balance was produced per condition
evaluated, typically two mass balances per condition.
The shutdown step of the unit, that includes cooling down the plant, is
satisfactorily carried out in about three hours.
Table 3-4 summarizes the events to be carried out for each experiment
performed to test one single set of conditions and the duration of each event
required. The maximum total hours required per run is also presented.
Operational events (hours per run)
Adjusting operational conditions (hours)
Reaching steady state conditions (hours)
Mass balance period (1 to 2 mass balance 2 hours each one)
Shut down (hours)
Total hours per run (maximum)
Hours
3
3
4
4
14 hours
Table 3-4 Operational periods for each run performed on the pilot plant unit
3.2.3 Feed Characterization
The heavy hydrocarbon feed used in this study was Athabasca Vacuum
Gasoil (AVGO), also called Athabasca virgin VGO, supplied by Suncor Energy
and the physical properties (carbon MCR, viscosity, sulphur and nitrogen
content) and simulated distillation data are provided in Appendix J.
62
3.2.4 Preparation of the Catalyst
The evaluation of physical characteristics and properties of ultradispersed
and fixed bed catalysts are not the focus of this research. The method of
preparation of ultradispersed catalyst via emulsions which was used is based on
the steam conversion process and catalyst studies performed by Pereira et all
[62], and on results obtained by Vazquez Alejandro in his MSc research thesis
[63]. Fixed bed catalyst preparation is based on development established by
Pereira et al [55].
1. Preparation of Ultra Dispersed Catalyst: According to the Steam
Conversion Process and Catalyst patent N° 5,885,441 [62], the non-
supported dispersed catalyst system is provided in the form of a water-in-oil
emulsion, having an average droplet size of maximum 10 microns and having
a ratio of water to oil by volume of between about 0,05 to about 0,4. The
catalytic emulsion is made so as to include a first alkali metal, for this
research potassium was selected but also could be sodium or mixtures
thereof, and a second metal which may preferably be a Group VIII non-noble
metal. For this case nickel was chosen but also could be applied cobalt, or an
alkaline earth metal such as calcium or magnesium or mixtures thereof.
Detailed preparation procedure of ultradispersed catalyst emulsion is
described in Appendix E.
2. Preparation of fixed bed catalyst: catalyst on fixed bed may suitable
be prepared through either co-impregnation or consecutive impregnation
methods by adding aqueous solutions of at least one transition metal selected
from group VIII of the periodic table of elements and/or alkali metal solutions
over the support, following by drying and calcination (The catalytic matrix
prepared for this research was made having nickel oxide and potassium oxide
supported on γ-alumina. Such a catalyst may suitably be prepared by
impregnating γ-alumina with an aqueous solution of potassium nitrate, drying
the impregnated support and calcining at a temperature of about 450ºC. The
resulting solid is then impregnated with a second solution of nickel nitrate;
63
Ni(NO3)2.6H2O, dried and also calcined. The resulting NiO-KxOy/γ-alumina
catalyst provides a stable fixed bed supported catalyst. A detailed preparation
procedure is shown in Appendix E.
3.2.5 Products Analysis
Gas produced was on line repeatedly sampled and analyzed with a gas
chromatograph during the mass balance periods. Liquid products of each mass
balance were also sampled for later analyses. Following is a description of the
analysis carried out.
3.2.5.1 Gases
The gas analyses for each run were performed on line using an SRI
multiple gas analyzer model 8610#2, 120 V TCD detector with a 10 port gas
sampling valve, and an assemble of 3’ molecular sieve / 6’ Hayesep-D columns.
The GC was previously calibrated with a hydrocarbon mixture gas and it took 40
minutes for each analysis. Appendix F provides the certificated composition
received for the calibration gas used and some examples of the gas
chromatographic results obtained.
3.2.5.2 Liquid Product
To evaluate and compare changes in the quality of the liquid products two
methods were applied: High temperature simulated distillation (SimDist) and
Carbon as % of carbon as MicroCarbon (MCR). A brief description of both
methods is developed next.
3.2.5.2.1 High temperature simulated distillation (HTSD) – ASTM D-7169-
2005
Simulated distillation (SimDist) is a gas chromatography (GC) technique
which separates individual hydrocarbon components in the order of their boiling
64
points, and is used to simulate the time-consuming laboratory-scale physical
distillation procedure known as true boiling point (TBP) distillation [64]. High
temperature simulated distillation is an alternative for determining true boiling
point distributions of heavy petroleum samples. The technique is calibrated by
correlating n-paraffins’ elution times with their known (or accepted) atmospheric
equivalent boiling points (AEBP). HTSD determines the true boiling point (TBP)
distribution of petroleum products up to a final boiling point (FBP) of 720°C. In its
present form, HTSD is the most convenient GC alternative for analysis of
samples containing distillation residua.
Reportedly, the estimated accuracy of the correlation between crude assay
distillation and HTSD yield at each cut point results in standard deviations of <
2% weight. The precision of HTSD cut points up to 1000°F is reportedly better
than 0.5% weight [65 ].
Simulated distillations were performed in the analytical area of our group in
an Agilent Gas Chromatograph Model 6890N used for HTSD. Chromatographic
analysis was performed with SimDist Expert 8 software provided by Separation
Systems [66]. Capillary columns P/N SS-112-102-01 from Separation Systems (5
m x 0.53 mm, 0.1 μm film megabore column) were used for the analysis. The
chromatographic events were controlled with the GC ChemStation software
provided by Agilent Technologies (Rev.A.10.02 [1757]). Sample solutions were
prepared in CS2 (about 0.2g sample/ 12.5 g solvent) and 0.2 µL injected into a
special cold on-column injector designed by Separation Systems. Experimental
conditions were set up following the standard ASTM-D7169-2005 procedure [67].
Results are presented as a TBP curve of temperature versus weight
percentage of mass recovered.
65
3.2.5.2.2 Microcarbon Residue method
Carbon residue is the residue that remains as a solid residue after
evaporation and pyrolysis of crude oil/ bitumen under given conditions. It is
indicative of the coke forming tendency of the oil under thermal degradation
conditions, like in refinery coking operations. In literature [84] there are three
methods that have been generally used and accepted as standards for
determining carbon residue:
1. Conradson method (ASTM D-189),
2. Ramsbottom method (ASTM D-524) and
3. Microcarbon Residue method (ASTM D-4530).
These three methods differ from each other with regards to the
experimental setup; however the principle remains the same. A known mass of a
sample is heated under controlled conditions. The mass is recorded after
heating. The difference in the mass reflects the amount of residue present.
Microcarbon residue (MCR) method has attracted more attention because of the
small amounts of sample used and simpler experimental set up.
Microcarbon analyses of liquid samples were performed in the analytical
area of our group. The methodology used is presented as following: Samples with known MCR (0.35-24.5 wt %) were obtained from PCA,
Texas, USA. These samples were used to get the calibration curve for our set up
for MCR determination. MCR determination was also carried out using thermo
gravimetric analysis equipment (TGA), from TA Instruments, coupled with a
Quadrupole mass spectrometer (MS), from Pfeiffer Instruments. The sample
(~10 mg) was placed in the ceramic sample holder. The reference holder was
kept empty. The system was purged with He at a flow of 300 cc/min for 20 min,
or until the mass spectrometer showed no trace of oxygen. He flow was then
lowered to 100 cc/min. Samples were heated up to 500°C for 15 min at a heating
rate of 10ºC/min. The mass loss was recorded at 500°C and after the system
reached ambient temperatures. This was done to compare the results with the
66
MCR values obtained by using a muffle furnace and weighing the sample before
and after heating at ambient temperatures.
Calibration curves obtained from TGA and Furnace experiments showed
that both TGA and furnace measurements could be used to determine MCR.
Since this set up allowed analyzing multitude of samples, its use was preferred
over TG for routine analyses.
MCR determination was carried out using a custom made apparatus to
analyze a maximum of 26 samples. The apparatus was placed in a Barnstead
Muffle furnace, equipped with programmable temperature controller.
An analytical Balance from Mettler with ± 0.01 mg sensitivity was used for
weighing the samples before and after heating.
The MCR apparatus was custom made of Aluminium and stainless steel in
order to make it light but robust. A known mass (10-40 mg) of sample was placed
in a 2 cc glass sample holder. The sample holders were then placed on the
platform adjacent to the purge tubing. Twenty six N2 purge tubes, 3/4” long and
1/8” diameter were installed for purging each sample. A glass cover 4” wide and
2” high with an orifice of 1/8” is then placed to shield the samples from air. The
system was purged for 45 min. The system was designed to ensure similar N2
flow through all the tubes. Then after 45 min the samples were heated 15 min at
500°C using a ramp of 10°C/min to reach that temperature. After heating, the
sample was allowed to cool under N2 until the temperature dropped to 200°C .
The samples were then placed in a desiccator before measuring the final mass of
the sample.
At this stage reference samples were also placed with the test samples
during each analysis to ensure correct determination of residual microcarbon.
Each sample was analyzed at least three times to ensure reproducibility of the
measurement and to check for low values dispersion.
67
Chapter 4: Results and Discussion
The results obtained to fulfill the objectives of this thesis are presented
and explained in this chapter and are divided into two major sections.
The first section comprises the results obtained to evaluate the reliability of
the pilot plant performance based on the repeatability of selected parameters of
each test. The second section contains reactivity evaluation of Athabasca
vacuum gasoil via Thermal, Steam and Selective Catalytic-Steam processing. A
comparison of performance for these processes is also developed.
4.1 Bench Scale Pilot Plant: Performance Evaluation.
One of the objectives of this research was the construction of a versatile
pilot plant unit for testing upgrading technologies including thermal and catalytic
processing with the addition of steam or hydrogen. As a first step toward this goal
it was required to demonstrate the reliability of the unit with respect to its steady
operation and the repeatability of results.
Results of two series of tests, SET1 and SET2, are presented and
compared in order to evaluate the ability of the unit to consistently perform the
same run and to provide the same data values under identical operational
conditions when it is conducted by the same operator, a characteristic better
known as repeatability [53]. For each series, two runs were carried out. The first
set (SET1) under thermal processing conditions at a pressure of 160 psig and a
reactor temperature of 400ºC, was intended to evaluate the repeatability of feed
pumping, liquid recovering as percentage of liquid yield and the liquid product
quality given by simulated distillation (SimDist). The same operator carried out
the second SET1 run 3 days after the first one.
Table 4-1 shows the operational conditions selected and results obtained
for the two SET1 runs. It can be seen that values of pressure, reactor
temperature and feed flow pumped were the same or almost the same in both
68
runs. Also, the weight of liquid product collected was very close and as result,
liquid yield values were practically the same, with a very low standard deviation.
Test SET1 Run Set1-A Set1-B Date 28-Jun-05 30-Jun-05 Pressure, psig 151 150 Temperature, C 400 400 Feed Flow, cc/h 203.40 202.84 Mass of Feed Pumped, g 384.63 383.57 Liquid Product Collected, g 381.60 378.2 Liquid Yield, % 99.21 98.6 Liquid yield, Stand Dev 0.43
Table 4-1 Operational conditions and liquid yield results for SET1 runs
50
150
250
350
450
550
650
0 10 20 30 40 50 60 70 80 90 100
Destilled Off, %
Boi
ling
Poin
t, C
SET1-A SET1-B Virgin VGO
Figure 4-1 Simulated distillation curves for virgin VGO and liquid products from
SET1 runs
The liquid product from runs performed for SET1 was recovered, sampled
and analyzed by simulated distillation (SimDist) and the resulting curves are
shown in Figure 4-1. The SimDist curve for virgin vacuum gasoil used as
69
feedstock is also presented. Comparing the SimDist of products they are
practically overlapping, and remarkably different from the SimDist feedstock
curve. This difference is due to the cracking reactions taking place during thermal
processing, which produced changes in the liquid product distribution. The liquid
product quality of the two different thermal cracking tests is almost the same,
thus demonstrating repeatability.
Test SET1-A Set1-B
Boiling Point (ºC ) Weight Percent Off, wt%
Weight Percent Off, wt%
Standard Deviation
68.7 0.60 0.61 0.01 125.7 2.53 2.65 0.09 174.1 5.26 5.57 0.22 216.3 8.05 8.56 0.36 253.5 11.80 12.51 0.51 286.8 16.85 17.76 0.64 316.3 25.12 26.17 0.74 343.2 37.75 38.85 0.78 368.6 52.03 53.14 0.78 391.2 63.27 64.35 0.76 412.0 72.29 73.31 0.72 431.3 79.99 80.93 0.66 449.2 85.45 86.32 0.61 465.7 89.48 90.30 0.58 481.0 92.27 93.05 0.55 496.0 94.17 94.93 0.53 509.0 95.52 96.27 0.53 522.0 96.46 97.21 0.53 534.0 97.20 97.95 0.53 545.0 97.75 98.50 0.53 556.0 98.20 98.95 0.53 566.1 98.57 99.32 0.53 575.0 98.87 99.63 0.54 583.9 99.12 99.89 0.54
Table 4-2 SimDist results for SET1-A and SET1-B
The repeatability in liquid product distribution is detailed in the SimDist
results shown in Table 4-2, for both SET1 runs. From this table it can be seen
that the maximum standard deviation is less than 0.8, a result that is well the
acceptable deviation range.
The second set of experiments (SET2) was conducted to evaluate the
repeatability of the unit under steam processing conditions by injecting 5wt% of
70
water in the line immediately before the reaction zone. In this set the same
operator performed the second run 30 days after the first run.
This time the focus of the evaluation was not only on liquid recovery and
liquid SimDist analysis, but also on the variables required to carry out the
reactivity study of the Athabasca vacuum gasoil developed subsequently: mass
of gas recovering and percentage of gas yield, conversion, key gas compounds
ratios (CH4/CO2, H2/CO2), carbon as Microcarbon residue (MCR, ASTM D-4530)
and mass balance. In this set the second run was performed by same operator
30 days after the first run.
Test SET2 Run Set-2A Set2-B Pressure, psig 261 260 Temperature,°C 420 420
Standard Deviation
Total mass pumped, g 397.72 398.82 0.78 Liquid Product Collected, g 371.73 368.67 2.16 Liquid Yield, wt% 93.47 92.44 0.72 Gas product collected, g 25.12 27.09 1.39 Gas yield, wt% 6.32 6.79 0.33 Product weight percentage-off at 350ºC, wt% 45.00 45.15 0.11 Conversion 350°C +, wt% 27.68 28.67 0.70 Mass balance, wt% 99.76 99.23 0.37 CH4/CO2 20.36 19.91 0.32 H2/CO2 2.29 2.82 0.37
Table 4-3 Operational conditions and variable results for SET2 runs
The operational conditions and results obtained for runs Set2-A and Set-
2B corresponding to SET2, are presented in Table 4-3. The pressure and
temperature values, 260 psig and 420°C respectively, used for this test differed
from those used in SET1 in order to evaluate the repeatability of the unit under
these new conditions and to verify if changes in operational conditions would
affect the repeatability. This new set of pressure and temperature would be
applied later in the reactivity evaluation.
From Table 4-3 it can be observed that the standard deviation for liquid and
gas products collected is lower than 2.5, which represents a percentage of
relative standard deviation of less than 0.6%, and is a very acceptable value in
71
terms of repeatability. Percentage of liquid and gas yield with standard deviations
less than 0.75 are also considered to be acceptable.
0
100
200
300
400
500
600
700
0 10 20 30 40 50 60 70 80 90 100Weight Percent Off, wt%
Boilin
g Po
int,
C
SET2-A SET2-B Virgin VGO .
Figure 4-2 Simulated distillation curves for virgin VGO and liquid products from SET2 runs
The percentage of weight recovery at 350°C resulting from the SimDist
analysis is similar for both runs. Both SimDist curves overlap each other as can
be seen in Figure 4-2 where only the end of the curve shows a separation, which
can also be seen in Table 4-4 with a standard deviation higher than 4. This last
finding is attributable to water traces found in the liquid sample remaining from
the steam injection which to some extent is due to the limited capabilities of the
hot separation tanks to completely separate water, thus compromising the
accuracy of SimDist results at the end of the curve. This situation was not
observed in the thermal runs due to absence of water in the process. Since this
deviation occurs beyond the 500ºC, which is considered unconverted material,
no impact or modification to the conversion and yield features of the processes
under study in this unit might be expected.
72
Test SET2-A Set2-B
Boiling Point (°C) Weight Percent Off, wt%
Weight Percent Off, wt%
Standard Deviation
0 163.4 163.5 0.07 5 170.6 171.2 0.42 10 219 220.1 0.78 15 257.6 258.5 0.64 20 285.7 286.2 0.35 25 304.3 304.4 0.07 30 317.5 317.6 0.07 35 329 329.1 0.07 40 339.5 339.7 0.14 45 349.4 349.7 0.21 50 359.1 359.7 0.42 55 369.2 370 0.57 60 380.1 381.2 0.78 65 391.8 393.4 1.13 70 404.6 406.7 1.48 75 418.3 420.9 1.84 80 433.9 437.8 2.76 85 455.1 461.6 4.60 90 489.4 505.6 11.46 95 614 -- --
100 -- --
Table 4-4 SimDist results for SET2-A and SET2-B
One of the most important calculated parameters to be considered in the
reactivity test was the percentage of conversion of the 350°C+ cut, used
afterward to estimate the conversion and activation energies. This parameter
was also considered in the repeatability evaluation. Table 4-3 shows the
conversion values for SET2 runs, the standard deviation between both results
was 0.76, a very low value that serves to confirm the repeatability of the
conversion in both runs. Table 4-3 also shows general mass balance results, as
the ratio of inlet mass to outlet mass in percentage terms, and ratio of hydrogen/
methane to CO2. All these results show a standard deviation for both runs lower
than 0.4, which is considered sufficiently low to conclude that the repeatability is
also confirmed
73
4.2 Reactivity Tests
A thermal preliminary testing of Athabasca vacuum gasoil (AVGO)
represents an essential reference for the catalytic steam processing of bitumen
heavy fractions. The study undertaken focuses on the evaluation of the feedstock
in a range of temperatures from 400ºC to 440ºC and at pressure conditions of
160 psig and 260 psig, both typical of thermal cracking and visbreaking.
For the purposes of this research the thermal cracking reactivity would
allow to confirm or discard whether the upgraded product obtainable with SCSC
evaluation is the result of the heating level, the steam addition or the contribution
of steam and catalyst in a joint chemical action.
4.2.1 Thermal Processing Evaluation
The experimental comparisons of thermal cracking runs were based on a
condition of equal spatial velocity in the reactor. Runs were carried out at two
pressure conditions, 160 and 260 psi and at three different temperatures namely,
400, 410 and 420ºC, to evaluate how the pressure and temperature variations
affected the quality of cracked products. Considerable difficulty was encountered
operating at temperatures higher than 420ºC under both pressures, because of
accumulation of coke produced in the reactor pipe as well as in the inlet and
outlet lines of the reactor. This problem forced the shutdown of the plant before
completing the runs on several occasions. This fact was confirmation of a
severity limit being reached at this temperature for processing of VGO under
thermal cracking. Thus, Athabasca VGO can only be processed under thermal
cracking in a stable and reliable fashion, with minimal reactor maintenance
(usually required to clean off coke deposits) if the temperature is maintained
below 420ºC in a pressure range between 160 and 260 psi and with a spatial
velocity of 2h-1 . A lower spatial velocity (longer residence time) would reduce the
temperature level in order to maintain a stable, solid deposits free, operation.
74
Table 4-5 provides a summary of the experimental conditions evaluated and
the results obtained. Detailed data on operational conditions and flows of feed
and products as well as gas chromatography analysis of the gas products for
each run is summarized in Appendix G. The experimental tag of the run follows
the format with the word “Thermal”, as the name of the process involved. The
first digit corresponds to the pressure condition applied, 1 for 160 psig, 2 for 260
psig, a dash and the second digit for the temperature condition applied, 1 for
400ºC, 2 for 410ºC and 3 for 420ºC.
Table 4-5 Summary of experimental conditions evaluated and result obtained for
Thermal Cracking evaluation.
Table 4-5 shows that at both pressures the liquid yield tended to decrease
with the increase of reaction temperature, however it was also observed that this
tendency to decrease was slightly higher at 260 psig than at 160 psig. Thermal
cracking of liquid hydrocarbons produce gases, liquid and coke as products.
Considering that coke production was negligible at the conditions shown (no
coke deposits were observed neither inside the reactor nor inside any pipe line
for the tests reported), the tendency for the gas yield to increase with reaction
temperature is only compensated by the reduction of liquid yield. Simply stated,
Test Thermal
1-1 Thermal
1-2 Thermal
1-3 Thermal
2-1 Thermal
2-2 Thermal
2-3
Pressure, psig 160 159 162 262 261 261
Temperature,°C 400 410 420 400 410 420
VHSV, hr-1 2.00 2.00 2.00 2.00 2.00 2.00
Liquid yield, wt% 98.58 96.38 95.90 98.16 94.09 93.25
Gas yield, wt% 2.75 3.78 6.71 4.61 5.64 6.92 Product weight percentage-off at 350ºC, wt% 36.91 41.68 46.11 38.92 41.73 46.60
Conversion 350ºC+ , wt% 12.50 20.92 27.29 15.65 22.87 29.94
Carbon MCR, wt% 0.30±0.02 0.52±0.22 0.86±0.41 0.36±0.24 0.55±0.05 0.66±0.21
%H2 0.00 0.00 0.00 0.00 0.00 0.00
%CH4 56.37 52.82 46.67 47.28 47.28 48.07
%CO2 3.14 2.13 1.88 2.63 2.33 1.86
CH4/CO2 17.97 24.83 24.81 17.95 20.28 25.82
Mass balance, wt% 101.33 100.16 102.61 102.77 99.73 100.17
75
the lower the liquid yield, the higher the gas yield, and the total of yield
percentages of liquid and gas should amount to 100%. The results of general
mass balance demonstrates this approach, and even though mass balance
values are in the range of 97% to 103%, this span is considered acceptable and
the small margin of error can be attributed to errors in feed flow measurements,
liquid product collection and gas flow produced determination. The maximum
decrease in liquid yield percentage or increase in gas yield percentage was
reached in the Thermal 2-3 run, which was the most severe run in terms of
reaction temperature.
In this research the evaluation was focused on the improvement of the
350ºC+ cut of the Athabasca VGO. This improvement was accomplished when
an increase in the light fraction (350ºC-) of the thermal cracked liquid product was
obtained. Table 4-5 also shows the values of weight percentage at 350ºC of
liquid product obtained by simulated distillation (SimDist). It can be seen that this
value increases as long as the reaction temperature increases, and that this
tendency was observed at both operating pressures.
Figure 4-3 Temperature effect on SimDist result for liquid product. Curves at (a) pressure=160 psig and (b) pressure= 260 psig
0
100
200
300
400
500
600
700
0 10 20 30 40 50 60 70 80 90 100Weight Percent Off, wt%
Boi
ling
Poi
nt, C
T= 400 C T= 410 C T= 420 C VirginVGO
0
100
200
300
400
500
600
700
0 10 20 30 40 50 60 70 80 90 100Weight Percent Off, wt%
Boi
ling
Poi
nt, C
T= 400 C T= 410 C T= 420 C VirginVGO
(a) P=160 psig (b) P=260 psig
76
SimDist curves of virgin VGO and liquid products for the three reaction
temperatures at each pressure of operation evaluated are shown in Figure 4-3. It
should be noticed that the higher the reaction temperature the larger the shift of
the SimDist curves to the right, which means a larger increment of lights product
350ºC- in the liquids. As anticipated, the quality of liquid product was strongly
affected by temperature. Detailed SimDist results for each run are presented in
Table G-2 in Appendix G. From Table 4-5 it can also be seen that at the same
reaction temperature the weight percentage-off at 350ºC of liquid product given
by SimDist is very close for both operating pressures.
Figure 4-4 indicates that at the same reaction temperature the SimDist
curves are almost overlapping and the shift to the right from the feed curve
occurred when the reaction temperature was increased for both pressures.
According to this result, the quality of liquid product in terms of SimDist seems to
be minimally effected by pressure within the range evaluated.
Figure 4-4 Pressure effect on SimDist results for liquid product. Curves at (a) temperature= 400ºC, (b) temperature= 410ºC and (d) temperature= 420ºC.
0
100
200
300
400
500
600
700
0 10 20 30 40 50 60 70 80 90 10Weight Percent Off, wt%
Boi
ling
Poi
nt, C
P= 160 psig P= 260 psig VirginVGO
0
100
200
300
400
500
600
700
0 10 20 30 40 50 60 70 80 90 100Weight Percent Off, wt%
Boi
ling
Poi
nt, C
P= 160 psig P= 260 psig VirginVGO
0
100
200
300
400
500
600
700
0 10 20 30 40 50 60 70 80 90 10Weight Percent Off, wt%
Boi
ling
Poi
nt, C
P= 160 psig P= 260 psig VirginVGO
(a) T= 400 C (b) T= 410 C
(c) T= 420 C
0
100
200
300
400
500
600
700
0 10 20 30 40 50 60 70 80 90 10Weight Percent Off, wt%
Boi
ling
Poi
nt, C
P= 160 psig P= 260 psig VirginVGO
0
100
200
300
400
500
600
700
0 10 20 30 40 50 60 70 80 90 100Weight Percent Off, wt%
Boi
ling
Poi
nt, C
P= 160 psig P= 260 psig VirginVGO
0
100
200
300
400
500
600
700
0 10 20 30 40 50 60 70 80 90 10Weight Percent Off, wt%
Boi
ling
Poi
nt, C
P= 160 psig P= 260 psig VirginVGO
(a) T= 400 C (b) T= 410 C
(c) T= 420 C
77
In Table 4-5 it can also be observed that the increase of percentage of
carbon MCR (Micro carbon) with the increase in temperature. Figure 4-5 clearly
demonstrates this trend. Even though it is clear that there is a general trend in
the increase of MCR with the increase of the temperature at the two pressure
values tested, it is unclear, due to the impossibility of testing beyond 420ºC (coke
deposits would plug the reactor), if a crossing of the trends with a reduction in the
positive slope for higher pressure conditions effectively takes place. Another fact
to be taken into consideration is the lack of accuracy of this analysis indicated by
the significant deviation in the measurements of MCR in this range of relatively
low MCR values. Table 4-5 illustrates that the standard deviation in most of
values range from 10% to 50%.
0
0.2
0.4
0.6
0.8
1
1.2
1.4
390 400 410 420 430
Reaction Temperature, C
Car
bon
MC
R, %
P=160 psig P=260 psig
Figure 4-5 Variation of MCR in liquid product
78
The increase of carbon MCR values with the increase of processing
temperature is attributed to the increase of hydrocarbon macromolecules in the
liquid product due to polycondensation which is an inevitable part of the reaction
mechanism of thermal cracking. The percentage of carbon MCR is usually
proportional to the quantity of coke formed from the feed under more severe
cracking conditions such as when heavy oil fractions are processed by Delayed
Coking or Resid Fluid Catalytic Cracking and FCC. In short, the higher the value
of carbon MCR the higher is the tendency of the product to form coke relative to
the original feed. In this evaluation all MCR values for the products collected
were higher than the feedstock as it was expected; also the highest values were
obtained at the highest reactor temperatures.
The conversion of the Athabasca VGO was one of the most interesting
parameters evaluated in this research. The conversion of the heavy fraction of
the feedstock having a boiling point greater than 350ºC is determined as follows
by equation 4-1 [54]:
Conversion (%) = 100))(
xR
CRR
i
fi +− Equation 4-1
Wherein:
Ri is the amount of hydrocarbon in the Athabasca vacuum gasoil having a
boiling point greater than 350°C.
Rf is the amount of hydrocarbon in the liquid product having a boiling point
greater than 350ºC.
C is the amount of coke produced in the process
Based on the results of percentage of liquid yield and gas yield together and
considering accurate the mass balance results, the coke produced during the
process is considered to have had a negligible effect on the calculations.
Table 4-5 shows conversion values for each run. It was found that the
conversion increased with the increase in reaction temperature and that the
incremental increase in the conversion was almost 8% per each 10ºC of reaction
temperature increase. Pressure was also found to increase the conversion;
79
however this increment was small, being about 2.5% from 160 psig to 260 psig,
compared to the increase due to the reaction temperature. Hydrogenation (H
transfer) reactions, typically favoured by pressure increases, may account for the
slight difference.
Figure 4-6 shows the trend of the conversion and indicates how these
trends result in two almost parallel lines representing the operation pressures
evaluated as the reactor temperature increases.
Partial gas chromatography results are shown in Table 4-5. Complete
chromatography of the gaseous products for each run is given in Table G-3 in
Appendix G. The analysis was done for hydrogen, CO2 and HC gaseous
components. However H2S and CO which are also found to be present in the
product gas after cracking reactions were not analyzed due to lack of sensitivity
of the chromatograph.
0
5
10
15
20
25
30
35
390 400 410 420 430
Reaction Temperature, C
Con
vers
ion,
%
P=160 psig P=260 psig
Figure 4-6 Variation of conversion with temperature and pressure.
In gases from thermal cracking, methane is obtained as a major fraction
[32]. The gases obtained in this thermal cracking evaluation demonstrate this
similarly. At high severity, there was a slight increase in C3 and C4 fractions. At
low severity, first the lower alkyl chain components cracked, possibly from bigger
molecules, however, with an increase in the severity, the higher component
80
fractions began to crack. CO2 was produced at an average of 2.35% in the gas
stream throughout the evaluation. No oxidizing agent was present in the media,
so this CO2 must be produced by reaction of the oxygen of the carboxylic and
other oxygenated species indigenously present in the vacuum gasoil. Because of
the high production of methane and the relatively low production of CO2, the
CH4/CO2 ratio shows an averaged value of 22 that is consistent with normal
expectations in thermal cracking.
81
4.2.2 Steam Processing Evaluation
The Steam processing evaluation, as an intermediate process between
thermal cracking and selective catalytic steam cracking (SCSC), was conducted
in this research to:
1. Evaluate the reactivity of Athabasca vacuum gasoil with the addition of
steam, under similar conditions of thermal cracking evaluation. The goal
being to depict any possible reactivity of steam with VGO in the absence
of catalyst and,
2. Apply these results at exactly the same level of steam in the catalytic
steam cracking of VGO that will be used as the base line in order to
determine the upgrading accomplished by the catalyst action.
The steam process evaluation was carried out at the same pressure
conditions used in thermal cracking, 160 and 260 psig. However temperature
conditions were 410ºC, 420ºC and 430ºC. Temperatures above 430ºC could not
be evaluated because of coke formation problems. Even though experimental
comparisons of steam cracking runs were based on equal spatial velocity in the
reactor, one change in spatial velocity, 1 hr-1, was also included to observe how
the product was affected. Results of the steam processing evaluation followed
similar trends as the thermal evaluation already presented. For that reason some
comments and considerations already expressed in the thermal evaluation are
omitted in what follows.
Table 4-6 provides the summary of experimental conditions carried out for
this evaluation and their corresponding results. Detailed data of operational
conditions, flows and products as well as gas chromatography are shown in
Appendix H. A similar structure to the one used for Thermal evaluation was used
for the naming of each run. An additional run, Steam2-3a, is the one
corresponding to the spatial velocity variation run previously mentioned. As can
be seen in this table, the percentage of liquid yield had a tendency to decrease
with the increase of reaction temperature, and it is also seen that the decrease is
slightly higher at 260 psig than at 160 psig. When comparing runs Steam2-3 and
82
Steam2-3a it can be seen that a 50% reduction of spatial velocity produced a
slight decrease in % of liquid yield. It also produces an increase in conversion, a
sensible increase in MCR and also in the variability of it as the standard deviation
of MCR for Steam2-3a suggests.
Test Steam1-1 Steam1-2 Steam1-3 Steam2-1 Steam2-2 Steam2-3 Steam2-3a
Pressure, psig 159 161 161 261 260 259 260
Temp.,°C 410 420 430 410 420 430 430
VHSV, hr-1 2 2 2 2 2 2 1
Water injection, wt% 5.01 4.93 4.97 5.03 5.01 4.99 5.03
Liquid yield, wt% 94.67 93.86 90.78 93.19 92.44 89.91 89.49
Gas yield, wt% 3.72 4.77 6.79 5.55 6.79 7.48 9.81 Product weight percentage-off at 350ºC, wt% 40.16 45.26 53.47 40.37 45.15 52.50 57.28
Conversion 350+C, wt% 20.30 27.71 40.58 21.82 28.67 39.92 45.83
Carbon MCR, wt% 0.40±0.09 0.82±0.59 2.18±0.05 0.52±0.14 1.15±0.21 2.75±0.11 5.12±1.34
%H2 8.92 8.91 11.31 7.44 5.53 6.83 8.80
%CH4 42.93 46.47 46.11 42.79 38.98 43.13 45.70
%CO2 2.56 1.80 1.76 3.04 1.96 1.75 3.31
CH4/CO2 16.74 25.84 26.17 14.07 19.91 24.68 13.81
H2/CO2 3.48 4.96 6.42 2.45 2.82 3.91 2.66
Mas Balance, wt% 98.40 98.63 97.56 98.74 99.23 97.39 99.30
Table 4-6 Summary of experimental conditions evaluated and results obtained for Steam Cracking evaluation.
The results of mass balance ranges between 97.3% and 99,3% indicate a
differential of less than 3% with respect to the target value of 100%, which can be
considered acceptable The small margin of error can be attributed to
measurements of feed and water flow pumped, liquid product collected or gas
flow produced. Coke deposits that were formed at the highest severities were
minor and do not seem to have had a noticeable impact on the global mass
balance.
Regarding weight percentage-off at 350ºC of liquid product (see Table 4-6),
it can be seen that this value increased with the increase of reaction temperature,
and this tendency was also evident at both pressure. Reduction of spatial velocity
83
produced an increase in the weight percentage-off 350ºC, with the highest value
obtained at a spatial velocity equal to 1hr-1, meaning more time for cracking
reaction to occur leading to higher production of light ends and gases.
Figure 4-7 shows the SimDist curves in the range 150 to 400ºC the boiling
point of virgin VGO and liquid products for the three reaction temperatures at
each operating pressure. It can be seen that at the boiling point of 350ºC and the
same reaction temperature but different pressure conditions, the curves almost
overlap. By maintaining a constant pressure but varying the reaction
temperature, the curves tend to move to the right of the feed curve (Virgin VGO)
resulting in the increase of atmospheric distillate yields in the same manner as in
the thermal cracking process. Detailed SimDist results for each run are
presented in Table H-2 in Appendix H.
Figure 4-7 Temperature effect on SimDist result for liquid product. Curves at (a) pressure=160 psig and (b) pressure= 260 psig
Regarding the percentage of carbon MCR results, Table 4-6 shows how this
parameter increased strongly with the increase of reaction temperature. Although
at 410ºC the pressure does not show much impact on MCR, the impact was
150
200
250
300
350
400
0 5 10 15 20 25 30 35 40 45 50 55 60Percentage Distilled Off, %
Boi
ling
Poi
nt, C
Virgin VGO P=160 T=410P=160 T=420P=160 T=430P=260 T=410P=260 T=420P=260 T=430
150
200
250
300
350
400
0 5 10 15 20 25 30 35 40 45 50 55 60Percentage Distilled Off, %
Boi
ling
Poi
nt, C
Virgin VGO P=160 T=410P=160 T=420P=160 T=430P=260 T=410P=260 T=420P=260 T=430
84
more significant when the reaction temperature was higher than 420ºC. The
effect of spatial velocity was even more important since reducing the spatial
velocity by half almost doubled the MCR value as shown in Figure 4-8. The
standard deviations for MCR in this evaluation were lower than the ones obtained
in thermal processing, however for the Steam 2-3a test the error was observed to
be high, being almost 30% of the %MCR value. This MCR variability observed for
the highest severity condition of the run 2-3a doesn’t mean inaccuracy of the
MCR measurement since it would show up at all other lower severity evaluations.
It is very possibly due to the incipient formation of coke particles, which were also
confirmed by observation of deposits occurring at the hotter zones in the plant
(reactor and near by zones). This close to instability condition produces variability
in the sampling. This had also been observed during similar thermal cracking
studies performed on heavy oils at PDVSA-Intevep [84].
0
1
2
3
4
5
6
7
400 410 420 430 440
Reaction Temperature, C
Car
bon
MC
R, %
P=160 psig P=260 psig VHSV= 1hr-1
Figure 4-8 Variation of MCR in liquid product
Regarding the global conversion of Athabasca VGO under steam cracking
processing conditions, it was again found that this parameter significantly
increases with reaction temperature, and that pressure has almost no effect. It is
85
important to underline that at the same temperature and pressure, thermal
cracking and steam cracking reach about the same VGO conversion levels,
28.6% ± 1.4%. This would suggest that no major effect of the water presence
was experienced at least in the range of 160 to 260 psig. Nevertheless, the
steam presence allowed a more stable operation at slightly higher conversions
than thermal cracking by flushing out of the reactor the micro carbon particles
being formed at that severity level. This is the reason for the use of low steam
proportions during industrial thermal cracking processes. We were able to
operate the pilot plant in a stable condition at higher thermal conversion levels in
the presence of steam. The fact that this fluid dynamic effect was evidenced in
the pilot plant indicates its suitability for examination and evaluation of many
features of scaled up industrial processes. An increase in the proportion of CO2
released with respect to CH4 at longer residence times was also noticed, which
may be indicative of an initiation of steam reforming reactions even in the
absence of catalyst.
Figure 4-9 shows the trend of the conversion with reaction temperature at
both pressure conditions, and also the point of conversion at a spatial velocity of
1hr-1. As expected an increase in the residence time had a positive effect on the
conversion, with an increment of more than 5%, when reducing the spatial
velocity by one half.
The last apparent advantageous result of increased conversion with the
reduction of spatial velocity had a major drawback. Although the Steam2-3a run
corresponding to VHSV= 1hr-1 was completed without any operating problems,
an increment of the pressure differential in the reaction zone was observed
during the cooling down of the unit. Subsequent inspection showed significant
coke deposition in the reactor inlet pipeline and also a small coke accumulation
in a strainer located in the reactor outlet pipeline. It all indicated a latent risk of
coke plugging which would likely have resulted in the eventual shut-down of the
pilot plant if the unit had worked at this condition for much longer. This incipient
86
coke deposition explains the variability of MCR for this particular condition tested
under steam thermal cracking, see Table 4-6.
15
20
25
30
35
40
45
50
400 410 420 430 440
Reaction Temperature, C
Con
vers
ion,
%
P=160 psig P=260 psig VHSV= 1hr-1
Figure 4-9 Variation of conversion with temperature, pressure and spatial velocity
Partial results of gas chromatography are shown in Table 4-6. In addition,
complete chromatography analysis of the gaseous products for each run are also
presented in Table H-3 in Appendix H. A similar quality of gas analysis described
for thermal cracking was performed for steam cracking studies.
As in the thermal evaluations, the collected methane was the most
important fraction in the gas product [32]. CO2 was also observed as a product
resulting from reactions of the carboxylic compounds resident in vacuum gasoil.
Some hydrogen was also detected by chromatography, most probably arising
from some water steam reactions incipiently occurring due to high temperatures
and long residence time in the reactor zone.
.
87
4.2.3 Selective Catalytic Steam Processing Evaluation
To carry out the UD catalytic evaluation of the Athabasca vacuum gasoil, an
emulsion containing 5% water in gasoil was prepared, according to the
procedure previously outlined in Chapter 3. Active metals forming the catalytic
matrix were present initially in the aqueous phase, and after vaporization of the
water by heating these metals remained ultradispersed in the oil media ready to
carry out the catalytic action at reactor conditions.
4.2.3.1 Ultradispersed Catalyst Test
In this evaluation the effect of two different catalytic matrixes were
investigated. The first one was produced by a combination of 300 ppm of nickel
and 416 ppm of potassium embedded in the media. The second one was
prepared using a combination of 300 ppm of nickel, 415 ppm of potassium and
415 ppm of calcium. All the runs were carried out at 260 psig and the
denomination of each run uses the prefix “UDCat”, to refer to the use of ultra
dispersed catalyst. The first digit after the prefix corresponds to the catalytic
matrix used, 1 for K-Ni and 2 for K-Ni-Ca. After the hyphen a second digit is used
to refer the spatial velocity condition evaluated 1 for WHSV=1 and 2 for
WHSV=2. Finally, the last digit represents the reaction temperature applied in
matrix 2, 1 for 430ºC and 2 for 440ºC.
Table 4-7 provides a summary of the experimental conditions investigated
and the results obtained. Detailed data of operational conditions, feed and
product flows as well as gas chromatography analysis for each run is shown in
Appendix I. In Table 4-7 can be seen that the conversion reached a value of
36% for the first matrix at a spatial velocity of 2 hr-1, however the conversion
increased to 51% when the spatial velocity was reduced 50%. A 50% conversion
was also found in matrix 2 at a spatial velocity of 1 hr-1. The fact that both of
these matrixes evaluated at the same spatial velocity reach such high conversion
values are justifiable grounds to pursue further evaluation of different metals and
matrix combinations that have not previously been used in ultradispersed
88
catalysts for heavy oils upgrading to try to achieve even higher levels of
conversion.
Matrix 2 was run at 440ºC and its effect on conversion increase was also
observed. A 56% conversion was obtained which is 5% more than the value
obtained at 430ºC. The value reported is the result of two consecutive mass
balances; the difference between balances was 1.17 standard deviation which
falls within the range of repeatability tested for thermal cracking (section 4.1). It is
important to mention that for this run, no signs of coke formation or deposition
were found in the internals after a thorough inspection, including the reactor,
which was found to be clean and free of coke. This finding suggests that runs at
reaction temperatures higher than 440ºC using the ultradispersed catalyst
formulation are still possible without the risk of coke plugging and with the benefit
of increasing the conversion even further. The exhaustion of feedstock prevented
us from further confirming this hypothesis..
Test UDCat 1-1 UDCat 1-2 UDCat 2-1 UDCat 2-2
Pressure, psig 260 260 260 260
Temperature,°C 430 430 430 440
VHSV, hr-1 2 1 1 1
Catalytic matrix K-Ni K-Ni K-Ni-Ca K-Ni-Ca
Liquid yield, wt% 90.44 86.69 87.55 91.49
Gas yield, wt% 9.22 12.95 12.03 9.38 Product weight percentage-off at 350ºC, wt% 49.70 59.76 59.62 65.71
Conversion 350+C, wt% 36.00 50.92 50.26 55.87
Carbon MCR, wt% 2.04±0.44 4.24±0.32 4.93±0.67 5.25±0.67
%H2 6.86 9.31 3.49 4.81
%CH4 38.27 37.54 45.41 45.23
%CO2 6.16 4.89 4.37 4.07
CH4/CO2 6.21 7.68 10.40 11.15
H2/CO2 1.12 1.90 0.80 1.21
Mass Balance, wt% 99.66 99.64 99.59 100.86
Table 4-7 Summary of experimental conditions evaluated and results obtained for Selective Catalytic Steam Cracking evaluation using ultradispersed catalyst.
89
In Table 4-7 it can be seen that the liquid yield values ranged from 87% to
92%. The decrease in the liquid yield corresponds with the increase of
conversion obtained. This increment of conversion also leads to a noticeable
increase of light ends in the cut 350ºC-, with values of 59.7, 59.6 and 65.7% as
shown in Table 4-7, but also to an increase in the gas production and as a
consequence in the gas yield. General mass balances closed very close to
100%, representing very acceptable numbers and showing good operational
performance.
Values of carbon MCR shown in Table 4-7 reveal increases with the greater
levels of severity even though the standard deviation was acceptable, which is
indicative of low solid coke formation, rather than from steam cracking at similar
conditions. Results for MCR at 430ºC and a spatial velocity of 1 hr-1 doubled in
both matrixes compared with the values obtained at the same reaction
temperature and spatial velocity of 2 hr-1. The reaction temperature increase as
observed in run UDCat2-2, also produced an increase in carbon MCR which,
combined with the lower spatial velocity, resulted in the higher carbon MCR
value.
It is important to note that the MCR analysis included ashes remaining after
the heating period established by the method. Furthermore, metals added to the
process as ultradispersed catalysts ended up in the final residual sample. Under
the pyrolysis conditions of the ASTM method these metals may act as enhancers
of MCR [81]. The real MCR number would need to be measured after separation
of the metals from the sample, which could not be performed in a practical way
without extracting carbonaceous solids that are a component part of the real
microcarbon intended to be measured. A reliable verification of a reduced
tendency to form solid deposits in the catalytic runs is the inspection of the
reactor zone after reaction, which in all catalytic tests was found to be perfectly
clean and free of carbonaceous deposits despite being performed at the highest
severity conditions.
90
Partial results of gas chromatography are shown in Table 4-7. Complete
chromatography of the gaseous products for each run is also presented in Table
I- 1 in Appendix I. Gas analysis was done for hydrogen, CO2 and HC gaseous
components of product gas stream. However H2S and CO that were also found
to be present in the product gas after cracking reactions were not analyzed due
to lack of sensitivity of the chromatographer used.
In this evaluation methane was also found to be the highest fraction in the
gas product stream, which was evidence that thermal effect still has a strong
impact on cracking reactions. CO2 was also produced; however the fraction of
this component is clearly higher than results previously shown in Thermal and
Steam evaluation. The increase of CO2 fraction indicates that the carboxylic acid
present in vacuum gasoil had decomposed to produce CO2. In addition, and as
mentioned in Chapter 3 steam reforming reactions in the presence of a catalyst
designed for that purpose, produced additional CO2 and hydrogen. A lower ratio
of CH4/CO2 observed in matrix 1 compared with matrix 2 indicates the probability
that steam reforming reactions are bigger in matrix 1 than in matrix 2. Ca could
also be converted into Calcium Carbonate during the reaction, which acts as a
CO2 scavenger reducing the CO2 proportion in the gases.
Hydrogen was detected by chromatography, which could be produced by
reforming reactions or by catalytic dissociation of water, also mentioned in
Chapter 2. In this case part of the hydrogen produced was reacting to stabilize
the free radicals formed thus contributing to higher liquid yields and to a reduced
tendency to form deposits inside the reactor.
4.2.3.2 Fixed Bed Tests
After finishing the UD catalyst evaluation, a fixed bed reactor was prepared
according to the patent already described in Chapter 3 [62]. The metals
impregnated to the support are the same used in the ultradispersed catalyst
matrix 2.
91
Table 4-8 shows experimental conditions investigated and the results
obtained. Detailed data of operational conditions and flows of feed and products
as well as chromatographic analysis of the gases produced for each run is shown
in Appendix I.
For this evaluation with a fixed bed packed reactor, the term spatial velocity
loses much of its usefulness as a basis of comparison with the results of
ultradispersed catalysts obtained in an unpacked reactor. In the latter case the
reaction volume available for reaction was 100 cc, whereas in the packed reactor
with catalyst occupying much of the volume the available porous volume for the
feed to flow was only 36 cc. For the packed bed case the term mass hourly
spatial velocity (MHSV) is the one usually applied and as defined below as:
MHSV = mass feed flow, g/hr Equation 4-2
mass of catalytic fixed bed, g The evaluation was carried out at MHSV = 1, 0.75 and 0.5 hr-1, however,
since the free volume of the reactor was 36 cc as previously mentioned, the
corresponding value of VHSV based on that volume was also calculated and
shown in Table 4-8 just for purpose of comparison.
The 5 runs performed in the pilot plant for this evaluation were carried out
continuously; conditions changing after the mass balance from each test were
completed. As can be seen in the table, FixBed1, FixBed2 and FixBed3 were the
experimental runs keeping a pressure of 260 psig and a temperature of 430ºC,
and varying the MHSV. Even though runs were increasing in severity, values of
conversion unexpectedly decrease. The FixBed4 run was a return point of the
initial condition of FixBed1 however the conversion diminished considerably. The
next condition FixBed5, was evaluated after increasing the temperature to 440ºC.
Although a higher conversion was obtained, the result was still low compared to
92
the conversion obtained in FixBed1. These results clearly pointed to a
deactivation of the catalyst properties.
Test FixBed1 FixBed2 FixBed3 FixBed4 FixBed5
Pressure, psig 260.00 260.00 260.00 260.00 260.00
Temp.,°C 430.00 430.00 430.00 430.00 440.00
MHSV, g/gh 1.00 0.75 0.50 1.00 1.00
VHSV, hr-1 2.30 1.72 1.15 2.30 2.30
Catalytic matrix K-Ni-Ca K-Ni-Ca K-Ni-Ca K-Ni-Ca K-Ni-Ca
Liquid yield, wt% 91.32 88.69 94.23 93.58 90.04
Gas yield, wt% 11.11 10.62 10.27 11.11 10.55 Product weight percentage-off at 350ºC, wt% 52.54 49.09 51.46 42.56 47.03
Conversion 350ºC +, wt % 38.95 36.43 35.64 24.38 32.91
Carbon MCR, wt% 2.69 2.11 2.36 2.27 2.07
%H2 55.02 54.47 55.62 34.06 42.08
%CH4 17.51 18.68 20.54 30.97 23.80
%CO2 3.12 3.62 3.53 6.17 5.05
CH4/CO2 5.61 5.16 5.83 5.02 4.71
H2/CO2 17.70 15.07 15.78 5.52 8.33
Mass Balance, wt% 102.43 99.30 104.50 104.70 100.59
Table 4-8 Summary of experimental conditions evaluated and results obtained for Selective Catalytic Steam Cracking evaluation using a fixed bed reactor.
Partial results of gas chromatography are shown in Table 4-8. In this
evaluation hydrogen, with more than 50%, was the compound with the highest
percentage in the gas product followed by methane. This served to confirm the
effectiveness of this catalyst for the generation of hydrogen through partial steam
reforming at the conditions explored. Nevertheless, deactivation, perhaps due to
an excessive accumulation of condensed molecules, requires design
modifications in the plant to accommodate for the excess of molecular
condensation. Perhaps swing guard reactors or swing reactors with more
frequent regeneration cycles. These aspects are recommended for future work.
Results of carbon MCR were lower than obtained in UD evaluation. It is
possible that coke started to accumulate in the fixed bed and if so, that would be
the reason for the reduction. It is important to note that the duration of this
93
evaluation was less than 30 hours, which resulted in a quick deactivation of the
fixed bed. This represents a clear advantage for to the UD evaluation, since the
metals dispersed in the media did not require any support, thus avoiding pore
plugging and deactivation problems encountered with the fixed bed evaluation.
Nevertheless, development of a fixed bed steam catalytic cracking process could
be of interest for other applications involving lighter fractions such as the
alkylaromatic streams in the range of atmospheric gasoil. This research is also
being undertaken in this research group [85].
4.3 Studied Processes Comparison
Table 4-9 shows the most significant facts obtained in the reactivity
evaluation of Athabasca vacuum gasoil. Results from the Thermal and Steam
evaluations were based on a pressure of 260 psig, the only value used in the
catalytic tests.
It is an established fact that non-coking thermal cracking severity limits are
determined by the onset of solid formation. In that sense thermal cracking
reached that critical condition at 420ºC, which is about 10ºC lower than steam
cracking. Steam cracking didn’t show a significant difference in conversion when
both were tested at the same temperature and spatial velocity which implies no
effect of water other than physical stripping.
The conversion in the steam cracking evaluation was improved when the
spatial velocity was lowered by half or stated another way, when the residence
time was doubled. However this conversion was lower than the one obtained with
UD catalysts 1 and 2 at the same spatial velocity. At this point it is important to
mention that this condition of Steam evaluation (T=430ºC and VHSV=1) is
presented as a limit condition for steam cracking even though coke deposits
were found inside the reactor at the end of the run. On the other hand UD
catalyst evaluations showed clean internals (no coke formation) at the same
condition, which indicates a further increase in the conversion at the same
94
reaction temperature The conversion of the fixed bed runs did not reach the
values of the UD catalyst evaluation and presented severe problems of catalyst
deactivation due to coke formation and pore plugging on the bed.
Fix Bed Test Matrix 1 Matrix 1 Matrix 2 Matrix 2 Matrix 2
Temperature, C 420 430 430 430 430 430 440 430VHSV, hr-1 2 2 1 2 1 1 1 2.3Conv 350+C 29.9 39.9 45.8 36.0 51.0 50.3 55.9 39.0Relative µCarbon 1.0 5.3 9.8 1.0 2.1 2.4 2.6 1.4CH4/CO2, (%v/%v) 25.8 24.7 13.8 6.2 7.7 10.4 11.2 5.6
Conditions and Results Thermal Tests Steam TestsUD Tests
Table 4-9 Summary results of Athabasca Vacuum Gasoil for different processes evaluated
Carbon MCR is presented as a relative value with respect to the value
obtained at the lower severity condition in their respective process evaluations.
The comparison is made in this manner so as to eliminate the effect of the
catalyst on the MCR evolution. Since the ratio of UD catalyst/feed was about the
same in all cases the trend (as opposed to the absolute number) of MCR with
severity is for practical purposes, independent of the catalyst presence. The
relative MCR evolves faster with severity (either via residence time or heat level)
for Thermal Cracking and Steam Thermal Cracking than it does in the Catalytic
process. The relative MCR observed in the fixed bed is low; however a significant
amount of coke was accumulated in the fixed bed during the process, which is
responsible for the catalyst deactivation thus yielding the comparison worthless.
As anticipated, the CH4/CO2 ratio in the thermal evaluations reached the
highest values. In the Steam evaluation this ratio was reduced only when the
lower spatial velocity (longer residence time) was used. However in the UD
catalyst runs the CH4/CO2 ratio was the lowest. The reduction is primarily due to
a CO2 increase in the gas product whose volume percentage in the gas evolved
from around 1.5 in Thermal Cracking to less than 3 in Steam Cracking to about
4-6 in Catalytic Steam Cracking. This increase is due to steam reforming
95
reactions attributed to the selective action of the catalyst and according to
reaction mechanisms already explained in chapter 2.
4.3.1 Activation Energy
The apparent activation energy was calculated for each one of the
processes evaluated, the results are shown in Table 4-10. For thermal cracking
and steam cracking a first order kinetic model was assumed whereas for the UD
catalyst evaluation a kinetic model between zero order and first order was
assumed, which was considered to be reasonable based on the literature [57].
Fixed bed catalyst data were not used for this comparison due to unreliability of
the data caused by the fast deactivation of the catalyst.
Activation Energy, kJ/mol
Thermal Cracking 305
Thermal Steam Cracking 301
Selective Catalytic Steam Cracking 43 / 66
Table 4-10 Activation energy results
It can be seen that thermal and steam thermal processing show the same
apparent activation energy, which confirms a very similar reaction mechanism.
The activation energy values obtained are well within the range of those reported
in the literature. Kataria et al [32] reported the activation energy for VGO thermal
cracking, at conditions similar to the ones evaluated in this research, to be in the
range of 264 – 398 kJ/mol.
For the selective catalytic steam cracking evaluation the activation energy
obtained was between 43 and 66 kJ/mol, depending on the assumption of a first
order or a zero order kinetic.
The difference in activation energies between the thermal processes, with
or without steam, and the catalytic process is considerable, from which it is
96
reasonable to conclude that there is a clear existence of two different
mechanisms, a catalytic one predominant for the tests using catalysts and a well
known thermal cracking pattern reported widely in the literature
4.4 Global Conversion Map for Athabasca VGO Thermal and Steam Catalytic Cracking
Further comparative features are shown below that illustrate the differences
between the Aquaconversion type of catalytic processing and the thermal
cracking process as applied to Athabasca VGO.
Figure 4-10 Athabasca VGO conversion map
Firstly, we can observe in Figure 4-10, the conversion map so far obtained
for Athabasca VGO. In it are illustrated the characteristic features of heavy
hydrocarbons conversion vs. increased severity via temperature increase.
Thermal cracking can only be sustained at a moderate conversion level if
coke plugging of the reactor is to be avoided. Steam cracking, which improves
conversion by stripping and solids flushing out of the reaction zone, enables
5
15
25
35
45
55
390 400 410 420 430 440 450Reaction Temperature, C
Con
vers
ion,
%
Steam Steam VHSV= 1hr-1 Thermal UDCat1-1UDCat1-2 UDCat2-1 UDCat2-2 FixBed1
Thermal Cracking
Steam ThermalCracking
CatalyticSteam ThermalCracking
Con
vers
ion,
wt%
°
5
15
25
35
45
55
390 400 410 420 430 440 450Reaction Temperature, C
Con
vers
ion,
%
Steam Steam VHSV= 1hr-1 Thermal UDCat1-1UDCat1-2 UDCat2-1 UDCat2-2 FixBed1
Thermal Cracking
Steam ThermalCracking
CatalyticSteam ThermalCracking
Con
vers
ion,
wt%
5
15
25
35
45
55
390 400 410 420 430 440 450Reaction Temperature, C
Con
vers
ion,
%
Steam Steam VHSV= 1hr-1 Thermal UDCat1-1UDCat1-2 UDCat2-1 UDCat2-2 FixBed1
Thermal Cracking
Steam ThermalCracking
CatalyticSteam ThermalCracking
Con
vers
ion,
wt%
°
97
conversion to be increased without modifying the thermal reaction mechanism
(trend with temperature is the same).
Secondly, the introduction of an ultradispersed catalyst with steam to
generate hydrogen and oxygen free radicals by water splitting prevents the
formation of large condensate molecules. This last step incorporates mechanistic
modifications that are reflected in lower activation energy as previously estimated
and increased liquid production as visible in the Figure 4-11
Figure 4-11 Products selectivity versus Athabasca VGO conversion
Figure 4-11 shows the evolution of atmospheric distillates (350ºC-) and
gas selectivites (green for steam processing and blue for the selective catalytic
steam processing). Figure 4-12 depicts the relative MCR evolution with
conversion; From Figure 4-12 we can compare the rapid growth of relative MCR
for steam cracking while the catalytic steam cracking remains stagnant with
severity. MCR evolution means an increased tendency to form coke. From the
information in both graphs it can be said that the catalytic process allows
increasing conversion without producing a parallel increase in gas or coke, which
0
5
10
15
20
25
30
35
40
45
50
0 10 20 30 40 50 60AVGO Conversion, %wt
Pro
duct
s S
elec
tivity
, %w
t
Gas Selectivity %wt Distillate Selectivity %wt, ,
0
5
10
15
20
25
30
35
40
45
50
0 10 20 30 40 50 60AVGO Conversion, %wt
Pro
duct
s S
elec
tivity
, %w
t
Gas Selectivity %wt Distillate Selectivity %wt, ,
98
typically happen with thermal processes, in consequence lighter liquids are more
selectively produced as clearly shown in Figure 4-11.
Figure
Figure 4-12 Relative MCR versus Athabasca VGO conversion
In summary, Thermal cracking of AVGO can be run stably up to a
conversion level of about 30%, which in the pilot plant setup was attained at
420ºC. The atmospheric distillates (350-ºC) reached at that conversion level a
selectivity value of about 18% while the gases, consisting mainly of
hydrocarbons, reach a selectivity of around 10%, Steam cracking does a bit
better while keeping the same activation energy, which implies a non chemical
effect of steam. The conversion in the presence of steam can reach about 40%
with 25% selectivity to atmospheric distillates. The catalyst designed to
dissociate water and to early stabilize free radicals that conduce to coke
deposits, allowed a remarkable increase in atmospheric distillates selectivity
reaching an elevated 44%, while keeping the gas selectivity at around 13% this
by increasing the residence time and temperature simultaneously. Conditions as
severe as the ones reached for the catalytic process can not be applied to
0
2
4
6
8
10
12
0 10 20 30 40 50 60
Conversion, wt%
Rel
ativ
e M
CR
Steam cracking Selective catalytic steam cracking
0
2
4
6
8
10
12
0 10 20 30 40 50 60
Conversion, wt%
Rel
ativ
e M
CR
Steam cracking Selective catalytic steam cracking
99
thermal or steam thermal cracking without a dramatic deposition of coke and
consequent fast plugging of the reactor and process lines.
100
Chapter 5: Conclusions and Recommendations
5.1 Conclusions
The following conclusions and recommendations stem from the results
and discussions just presented:
A bench scale plant for evaluation of reactivity and catalytic steam
cracking (UD and Fixed Bed) of heavy oil feedstock was built and its
reliability was proven by testing the reactivity of Athabasca VGO under
thermal and steam processing conditions and achieving repeatability of
results (liquid yields, quality of products, conversion).
The effect of variables such as temperature and pressure and spatial
velocity were evaluated for thermal and steam processing. Their
comparison was performed based on Simulated Distillation analysis as
well as micro-carbon (MCR) to obtain conversions and selectivities. The
results show only slight sensitivity of conversion, yields and MCR with
pressure, but they were strongly affected by temperature and residence
time (inverse of spatial velocity) within the operating range scanned.
Activation energies found for Thermal and Steam cracking were very
similar in the typical range of thermal processes (~300 kJ/mole), the
Catalytic Steam Cracking activation energy is about 6 times lower.
By applying selective catalytic steam cracking the feedstock can be
processed at higher temperature (440ºC) to attain conversion levels not
possible to reach in Thermal and Steam Tests due to accelerated coke
formation and potential plugging at that severity.
101
The Selective Catalytic Steam Cracking process allows within the range of
the experiments performed increasing conversion without producing a
parallel increase in gas or coke, typically observed with thermal
processes. In consequence lighter liquids are more selectively produced.
The relative increase of MCR is higher in the case of Steam Tests
compared to catalytic tests. This trend revealed substantial differences in
condensation and coke formation reactions that presumably allow the
catalytic process to evolve towards liquid yields beyond the limits of
thermal steam cracking.
The ratio of CH4/CO2 in Steam Test almost doubled from VHSV=2hr-1 to
VHSV=1hr-1 representing a large amount of gas produced, particularly
methane, however in catalytic tests this increase is less with the CH4/CO2
ratio almost reduced by half, basically due to the increase in CO2. These
results are also explained by the occurrence of the steam reforming
reactions targeted with the catalyst.
5.2 Recommendations
• Some modifications and improvements to the pilot plant would facilitate
the operation of the plant and would increase the reliability of the results of
the experiments. The feed measurement would probably be improved by
incorporating a digitalized weight tank able to register changes of mass
during run times. The current system using burettes is not absolutely
accurate since the readings depend on the visual ability of the operator.
Substituting some manual valves with electronically controlled devices
actuated from the main panel would allow a semi-automation of the unit.
A technical assessment of these modifications is recommended.
102
• Having tested the thermal and steam catalytic processing of AVGO, it is
relevant to test more complex feedstock as vacuum residue or bitumen
with this pilot plant unit. The use of heavier feedstock will offer a more
comprehensive approach to the upgrading investigation being carried out
in our research group.
• Based on the results obtained using ultradispersed catalysts for steam
processing of AVGO, improvements in the catalyst performance might be
enhanced by combining other transition and alkali metals to create new
catalytic matrixes.
• Analysis of ultradispersed catalysts remaining in the liquid product, were
outside the scope of this thesis. However, further research is pending in
this area. Evaluation of catalytic activity of metals that remains in the liquid
product as well as the possibility of recycling to the reactor is still in the
study stage at this time.
• Design modifications in the plant to accommodate for the excess of
molecular condensation deactivation observed on the fixed bed catalytic
processing can be considered. Options like swing guard reactors or swing
reactors with catalyst regeneration in parallel are recommended as well to
be evaluated for future work
103
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65. Villalanti, D., Raia, J., Maynard, J., Arias. “Yield Correlations between Two Evaluation Techniques: Crude Oil Assay Distillation versus HTSD,” Triton Analytics Corporation; Shell Development Company. home.earthlink.net/~villalanti/HTSD.pdf
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110
Appendix A: Reactors Sizing Estimation
Spat Vel (h-1) 4.00 3.33 3.00 2.00 0.60 0.250 0.125OD in R=L/D ID (cm) V (cc) feed flow(cc/h) feed flow(cc/h) feed flow(cc/h) feed flow(cc/h) feed flow(cc/h) feed flow(cc/h) feed flow(cc/h) L (cm) V probe (cc) VReact-Vprob Ladd (cm) 1/2 50 1.09 51.16 187.36 156.14 140.66 93.68 28.05 11.71 5.86 54.61 4.32 46.84 4.61 1/2 60 1.09 61.40 224.84 187.36 168.80 112.42 33.66 14.05 7.03 65.53 5.19 56.21 5.54 5/8 50 1.34 94.19 355.55 296.30 266.93 177.78 53.23 22.22 11.11 66.93 5.30 88.89 3.77 5/8 45 1.34 84.77 320.00 266.67 240.24 160.00 47.90 20.00 10.00 60.24 4.77 80.00 3.39 5/8 50 1.26 78.05 292.29 243.58 219.44 146.15 43.76 18.27 9.13 62.87 4.98 73.07 4.01 5/8 68.73 1.26 107.29 401.79 334.82 301.64 200.89 60.15 25.11 12.56 86.41 6.84 100.45 5.51 5/8 50 1.17 62.23 230.46 192.05 173.02 115.23 34.50 14.40 7.20 58.29 4.62 57.61 4.32 5/8 60 1.17 74.68 276.55 230.46 207.62 138.28 41.40 17.28 8.64 69.95 5.54 69.14 5.19 3/4 30.11 1.66 107.41 413.85 344.87 310.70 206.92 61.95 25.87 12.93 49.86 3.95 103.46 1.83 3/4 26.25 1.66 93.64 360.80 300.66 270.87 180.40 54.01 22.55 11.27 43.47 3.44 90.20 1.60 3/4 30.00 1.66 107.02 412.34 343.61 309.56 206.17 61.73 25.77 12.89 49.68 3.93 103.08 1.83 3/4 30.00 1.57 92.02 353.12 294.27 265.11 176.56 52.86 22.07 11.04 47.24 3.74 88.28 1.92 3/4 30.00 1.48 76.90 293.52 244.60 220.36 146.76 43.94 18.35 9.17 44.50 3.52 73.38 2.04 3/4 33.00 1.48 84.59 322.88 269.06 242.40 161.44 48.33 20.18 10.09 48.95 3.88 80.72 2.24 3/4 35.00 1.48 89.72 342.45 285.37 257.09 171.22 51.26 21.40 10.70 51.92 4.11 85.61 2.38 3/4 36.00 1.48 92.29 352.23 293.52 52.73 176.11 52.73 22.01 11.01 53.40 4.23 88.06 2.451 50.00 2.12 373.11 1458.91 1215.76 218.40 729.45 218.40 91.18 45.59 105.90 8.38 364.73 2.381 41.10 2.12 306.70 1199.22 999.35 179.52 599.61 179.52 74.95 37.48 87.05 6.89 299.81 1.961 40.00 2.12 298.49 1167.13 972.61 174.72 583.56 174.72 72.95 36.47 84.72 6.71 291.78 1.901 30.00 2.12 223.87 875.35 729.45 131.04 437.67 131.04 54.71 27.35 63.54 5.03 218.84 1.431 20.00 2.12 149.24 583.56 486.30 87.36 291.78 87.36 36.47 18.24 42.36 3.35 145.89 0.95
Light blue: Reactors sizing selected
111
Appendix B: Catalytic Steam Cracking Pilot Plant
112
Appendix C: Tags Assigned for Points of Temperature and Pressure Temperature Control and Indicators (TIC):
TIC 100: Feed tank (50 -120 °C)
TIC 101: Pump outlet – PH1 inlet (50 -120 °C)
TIC 102: Steam generation (100 - 400 °C)
TIC 103: Feed tank outlet - pump inlet (50 -120 °C)
TIC 200: PH1 (100 - 200 °C)
TIC 201: PH1 outlet - PH2 inlet (100 - 200 °C)
TIC 202: PH2 (200 - 350 °C)
TIC 203: PH2 outlet - PI1 (300-350 °C)
TIC 204: PI1 – Steam injection (300 - 400 °C)
TIC 205: Steam injection – reactor inlet (400 °C)
TIC 300: Reactor A, zone 1 (400- 460 °C)
TIC 301: Reactor A, zone 2 (400- 460 °C)
TIC 302: Reactor A, zone 3 (400- 460 °C)
TIC 400: Reactor B, zone 1 (400- 460 °C)
TIC 401: Reactor B, zone 2 (400- 460 °C)
TIC 500: Outlet reactor – 3-way valve (100 - 230 °C)
TIC 501: 3-way valve – Heavy product tank 1 (80 - 350 °C)
TIC 502: 3-way valve - Heavy product tank 2 (80 - 350 °C)
Total: 18
Temperature Indicators (TI):
TI 100: Feed tank
TI 101: Feed tank outlet
TI 102: Feed pump inlet
113
TI 103: Feed pump outlet
TI 201: Outlet PH1
TI 203: Outlet PH2
TI 205: Inlet reactor
TI 300: Reactor zone 1, point 1
TI 301: Reactor zone 1, point 2
TI 302: Reactor zone 2, point 3
TI 303: Reactor zone 2, point 4
TI 400: Reactor outlet
TI 401: Before 3-way valve
TI 402: Condenser Inlet
TI 403: Condenser outlet
TI 404: Spare
TI 500: Heavy product tank 1 inlet
TI 501: Heavy product tank 2 inlet
TI 502: Heavy product tank 1
TI 503: Heavy product tank 2
Total: 20
Pressure Indicator (PI)
PI 001: Before reaction system
PI 002: After reaction System
PI 003: Before Back Pressure Regulator
Total: 3
114
Appendix D: Algorithm for the Design of Vapors Condenser
Characteristics of the vapour product from the heat exchanger 2:
Tsat ( °C ) ρl (kg/m3) ρg(kg/m3) Cp (kj/kg C) k (w/m C) μ (cp) Hfg (Kj/Kg)
120 888.4 2.7727 3.8571 0.13859 0.91746 1789.4
Tii = 175 C, Tiq = 26 C
mi = 1.92 x 10-5 m3/h
Twi = 25 C
Di = 3/8” , Do = 0.75”, 1”, 1.5”, 1.25”, 2”, 2.5”
Lmax = 50 cm
Vapors saturation enthalpy = 1789.4 kJ/kg
The gas enters the condenser with a flux of approximately 1.92 x 10-5 m3/h
and about 160 °C.
1) Get Density of the liquid phase (ρl), density of the gaseous phase (ρg), calorific
capacity of liquid phase (Cp), Thermal conductivity of liquid phase (k) and
viscosity of the liquid phase (μ) of the inlet vapor (GP) at the temperature of
saturation of the gaseous products.
2) Get Density (ρ), Calorific capacity (Cp), Thermal conductivity (k) and viscosity
(μ) of water at 25 °C.
3) Calculate Prank’s number of water and gaseous product.
k
Cp μ*Pr =
4) Calculate the mass flow of GP (mi)
ρ*vmm =
115
Where mv= volumetric flow
5) Assume a water flow (mw) ≈ 10 kg/hr and determine the temperature of water
in the outlet (Two)
ww
wewo mCpqTT*
−=
Where: Twe= Temperature of water in the inlet
q = Heat flux.
Cpw= Calorific capacity of water.
mw= mass flow of water.
6) Calculate LMTD
Where: Tie= Temperature of GP in the inlet.
Tio = Temperature of GP in the outlet
7) Choose one of the external diameters given (Do).
8) Assume a longitude (L).
10) Determine Reynold’s number for the water and the gaseous product.
( ) μ***4Re
iw
ww DD
m+Π
=
11) (Because the mass flows are very small, the Reynolds number is going to be
very small, showing laminar flows; therefore the equation that is going to be used
next is for this type of flow.) Determine the Nussel’s number for the annular pipe.
⎟⎟⎠
⎞⎜⎜⎝
⎛−−
−−−=
weio
woie
weiowoie
TTTTLn
TTTTLMTD )()(
116
32
PrRe***04.01
PrRe***668.066.3
⎟⎠⎞⎜
⎝⎛+
⎟⎠⎞⎜
⎝⎛
+=
LD
LD
Nuh
h
Dh=Do-Di
12) Calculate the area of the internal pipe (Ai) and the annular area (Ao)
LDALDA
oo
ii
****
Π=Π=
13) Determine the convection coefficient of water (hw)
ioh
h
DDDDkNuh
−=
= *
14) Get the fouling factor of water (Rfw=0.0003) , the fouling factor of oil
(Rfo=0.0009) and the thermal conductivity of the steel (ks=18.3 w/m°C).
15) Determine the resistance (Ro) of the annular part and the resistance (Rs) of
the stainless steel.
owo
wo AhA
RfR*1
+=
LkD
DLnR
s
iis ***2
1
Π
⎟⎠⎞⎜
⎝⎛
=
16) Calculate the convection coefficient for the tube part (hi), assuming that Tsat=
120 °C
117
( )( )
41
3
******
*555.0 ⎟⎟⎠
⎞⎜⎜⎝
⎛
−
′−=
issatL
fgLgLLi DTT
hkgh
μρρρ
Where:
g=Gravity
( )ssatLfgfg TTCphh −+=′ **83
hfg= Enthalpy of saturation.
Ts= Temperature of the steel wall.
Because Ts is unknown, hi is left in terms of it
17) To determine Ri the convection coefficient found in the last step is replaced in
the next equation as well as all the other known terms.
i
o
iii A
RfAh
R +=*1
After the replacement of terms;
0066.01*
)120(*1384505171300594)120(*03146.007519.0
41
⎟⎟⎠
⎞⎜⎜⎝
⎛−+
−+=
s
si T
TR Equation 1
18) By doing energy balance in the condenser the result is:
o
weo
s
os
i
ssat
RTT
RTT
RTT −
=−
=− Equation 2
With E1 and E2 and a method for root finding (e.g. secant method) Ts is
determined, and, therefore; Ri.
118
19) Determine the heat transfer coefficient
)(*1
osio RRRAU
++=
20) Determine a Length (Lcalc)
LMTDDUqLcalci *** Π
=
21) Is Lcalc≠L?
Yes
Go to step 8. End if after 20 tries
No
End if
. - Repeat from step 16 with Tsat= 140, 150, 1601.
. - Repeat from step 7 with every annular diameter size.
. - Choose The Best Annular Diameter With Its Resulting Length
1 The properties of the GP at the different temperatures of saturation are very similar to the ones at 120°C, therefore; to make the calculations easier, the same properties were used.
119
Appendix E: Catalysts Preparation Procedure
A) Ultradispresed catalyst: Emulsion preparation.
It is required to prepare an emulsion with 400 ppm of Potassium (K) and
(400/3) 133 ppm of Nickel (Ni). To do this we need to prepare a surfactant that
will approximately be 1 % of the complete solution.
In different experiments done by Dr. Carlos Scott from Bitumen Upgrading
and Hydrogen Production group at University of Calgary, the most stable
emulsion with the Athabasca VGO was obtained with a combination of
surfactants, SPAN 80 (S80) and TWEEN 80 (T80), prepared in the proportion of
65.1 % of S80 and 34.9 % of T80.
It was determined that for each run in the pilot plant approximately 2200
ml of emulsion was needed. Except for the first run that will need approximately
an extra litre to fill the dead volume of the plant and push any impurities that
could be in the lines.
In order to have enough surfactant for a few days, approximately 100 g of
surfactant were prepared. This required 65.1 g of S80 and 34.9 g of T80 for 100
g of solution.
It was measured in a balance 65.209 g of S80 and 34.980 g of T80 which
is a total of 100.189 g of surfactant.
The emulsion will need to have approximately 2.5 % of the Nickel solution
and approximately 2.5% of the Potassium solution. The calculation to determine
the amount of K and Ni to be added to water in order to obtain a solution of each
compound was done as followed.
a) Production of the Potassium solution:
Approximately 400 ppm of potassium (K) have to be present
in the emulsion, so is necessary to determine the mass of K in order to prepare
10 L of emulsion to have enough for at least two runs.
120
Vol. of K solution= 10000 ml*0.025=250 ml
[K]solution= (400_mg/Lemul*10_L)/250 ml=16 mg/ml sol=0.016 g/ml sol.
Mass of K = 0.016_g/ml sol*250ml= 4 g.
Solids of KOH with a purity of 85% were available and, therefore, were
used. In order to determine the amount of solids to be used to obtain 4 g of K the
next calculations were done.
Knowing that in 56 g of KOH the amount of pure potassium is 39 gr.
Mass of KOH= 4 g* 56 gr/39 g = 5.74 g
Mass of solids= 5.74/0.85 = 6.76 g
The amount of KOH solids that was weight in the laboratory was 6.754 g.
This amount was added to 100 ml of water, then the solution was agitated and
water was added until 250 ml were reached. The amount of KOH solids weighted
has 3.998 gs of K and the concentration of this component in the solution is
0.01599 g/ml. The concentration of KOH in the emulsion is 399.75 ppm. These
values are within experimental error of the calculated values, therefore are
acceptable.
b) Production of the Nickel solution:
In the emulsion is needed 133 ppm of nickel (Ni), so is necessary to
determine the mass of Ni to prepare 10 l. of emulsion to have at least enough
amounts for two runs.
Vol of Ni solution=10000 ml* 0.025= 250 ml
[Ni]solution= (133 mg/l*10 Lt)/250 ml=5.32 mg/ml sol =0.00532 gr/ml sol
Mass of Ni= 0.00532 g/ml sol*250 ml=1.33 g
Solids of Nickel Acetate (NiAc) with a purity of 98% were used to prepare
the solution. In order to determine the amount of solids to be used in order to
obtain 1.33 gr of Ni the next calculations were done.
Knowing that in 249 g of Nickel acetate the amount of pure Nickel present
is 58.7 g.
121
Mass of NiAc=1.33 g*249 gr/58.7 g =5.64 g
Mass of Ni solids= 5.64 g/0.98 g = 5.757 g
The amount weighted in the laboratory is exactly as calculated.
c) Preparing the emulsion:
Approximately 3500 ml of emulsion were necessary to prepare for
the first run of the plant, but it was not possible to prepare all at once because of
unavailability of a container big enough for that amount. Only 2500 ml
approximately were prepared on this first day.
Because the mass percentage of the components in the emulsion
is known, is possible to determine an approximation of the density of it.
ρemul= (X*ρ)VGO+(X*ρ)Nisol+(X*ρ)KOHsol+(X*ρ)surf
ρVGO= 0.9455 g/ml
ρNisol=1 g/ml
ρKOHsol=1 g/ml
ρsurf= 0.989 g/ml
ρemul=(0.94*0.9455)+(0.025*1)+ (0.025*1)+(0.01*0.989) = 0.948 g/ml.
With the density is now possible to know the mass of emulsion in
2500 ml.
Mass of emulsion = 2500 ml*0.948 g/ml= 2370 g
Mass of VGO =Mass of emulsion*0.94
Mass of VGO = 2370 g*0.94=2227.8 g
Mass of Ni and K= 62.5 g each.
Mass of surfactant =25 g
In this first preparation a mistake occurred while weighting and instead of
2227.8 g of VGO, 2245 g were weighted. It was also weighted 25 g of Surfactant
and 62.5 g of each catalyst. A correction had to be carried out, this mean that
instead of 2370 g of Emulsion, 2395 g were prepared.
122
To prepare the emulsion the VGO and the surfactant are added in a
beaker. Then the beaker has to be heated between 45°C and 50°C (for mixtures
Residue+VGO use 70°C), and agitated at 900 rpm. When the mixture reaches
the desired temperature, the KOH solution is added slowly. After the KOH
solution is added is necessary to wait 15 minutes before adding the NiAc
solution. The NiAc solution has to be added slowly too, and then wait for 15
minutes before stopping the agitation. If the emulsion is not needed right away,
can stay agitated at 500 rpm until needed.
The density of the emulsion was measured and the result was 0.953 g/ml.
Final Volume of emulsion =2395 g/0.953 g/ml = 2513.1 ml
Final % of VGO = (2245 g/2395 g)*100 =93.74%
Final % of surfactant = (25 g/2395 g)*100 =1.04%
Final % of solutions of Ni and K = (125 g/2395 g)*100 =5.22%
B) Fixed Bed catalyst: impregnated support preparation
i) Support pre-treatment
There is a previous know-how that a mild acid pretreatment on the alumina
can eliminate support contaminations and enhance the dispersion of metal such
as rhodium [85], therefore this concept was used in this present research work.
The support, γ-Al2O3 was placed in a 250 ml beaker and a 0.1 M HCl solution
was then added to it. The resulted mixture was kept at room temperature for 12
hrs in the hood. The beaker was then transferred to a muffle furnace and dried at
120°C for 12 hrs. The resulting support was stored in a desiccator.
123
ii) Incipient wet impregnation for Ni catalyst
This method was used to deposit the active metal (Ni) either on the blank
support or on pre-coated support (by Ca and K layer) and it is described as
follows:
First a known amount of Ni(NO3)2.6H2O was dissolved in distilled water.
Then the activated γ-alumina was placed in a glass dish and the aqueous
solution of nickel nitrate was added to it. The initial green solution was gradually
turned into light green, while the solid became green. This change in colour was
a visual indication that impregnation is completed. Then the catalyst was dried at
room temperature overnight, transferred to the muffle furnace and dried at 120°C
for 2 hrs. The catalyst was then calcined in air at 500°C for 8 hrs. Thus the
resulting dark green solid was transferred to a bottle, sealed and kept for later
use.
iii) Sequential wet impregnation (coating with K and Ca oxides)
This method was used to deposit a layer of the desired potassium and
calcium oxides on the support prior to the impregnation with nickel. The
procedure is described as follows:
A known amount of the activated γ-alumina (100 g) was placed in a glass
dish. An aqueous solution of the desired concentration of Ca(NO3)2 (6.71 g) and
KNO3 (2.71 g) in 66 ml of distilled water was then added slowly to the dish
containing the activated γ-Al2O3. After 12 hrs of standing in the hood, the coated
γ-Al2O3 was transferred to the muffle furnace for drying at 120°C for 2 hrs. After
drying, the catalyst was calcined at 500°C for 8 h under air flow. The resulting
white solid was transferred to a glass dish and impregnated with an aqueous
solution of Ni(NO3)2.6H2O (29.65 g) in 66 ml of distilled water, following the
same procedure described in the previous section. This procedure was used to
prepare the catalyst.
124
Appendix F: Certified Gas composition and Calibration Result for Gas Chromatographer
Component Certified concentration %
Tested concentration %
Standard deviation
Methane 51.4 49.41 1.41Hydrogen 13.6 14.05 0.32Ethylene 1.5 2.93 1.01
Carbon Dioxide 0.2 0.19 0.01Ethane 15.9 15.76 0.10
Propylene 3.1 3.21 0.08Propane 8.2 8.03 0.12
i-Butane/1-Butene 3.4 3.53 0.09n-Butane 2.6 2.88 0.20i-Pentane 0.1 0.01 0.06
Certified gas mixture provided by Praxair Distribution Inc.
125
Appendix G: Operational Data and Experimental Results for Thermal Cracking Runs
Test Thermal 1-1 Thermal 1-2 Thermal 1-3 Thermal 2-1 Thermal 2-2 Thermal 2-3 Date 22-Mar-06 10-Apr-06 17-Apr-06 24-May-06 29-May-06 May 31,2006
Length of mass balance, h 2.00 2.00 2.00 2.00 2.00 1.50
Pressure, psig 160.00 159.00 162.00 262.00 261.00 261.00
Temp.,°C 400.00 410.00 420.00 400.00 410.00 420.00
VHSV, hr-1 2.00 2.00 2.00 2.00 2.00 2.00
Feed flow pumped, cc/h 199.40 200.99 199.89 200.90 200.30 201.10
Total mass pumped, g 377.06 380.07 377.99 379.90 378.80 380.28
Total liquid product, g 371.72 366.32 362.52 372.92 356.41 354.62
Product gas flow, cc/min 73.89 99.77 176.23 114.13 137.79 174.30
Total gas product, g 10.36 14.36 25.36 18.49 21.36 26.31
Liquid yield, wt% 98.58 96.38 95.90 98.16 94.09 93.25
Gas yield, wt% 2.75 3.78 6.71 4.87 5.64 6.92
VGO SimDis at 350C, wt% 28.92 28.92 28.92 28.92 28.92 28.92
Product SimDis at 350C, wt% 36.91 41.68 46.11 38.92 41.73 46.60
Conversion 350+C, wt% 12.50 20.92 27.29 15.65 22.87 29.94
Distillate Selectivity, wt% 10.50 15.83 21.52 13.06 14.55 20.45
Gas Selectivity, wt% 3.86 5.31 9.44 6.85 7.93 9.73
Carbon MCR, wt% 0.30±0.02 0.52±0.22 0.86±0.41 0.36±0.24 0.55±0.05 0.66±0.21
%H2 0.00 0.00 0.00 0.00 0.00 0.00
%CH4 56.37 52.82 46.67 43.38 47.28 48.07
%CO2 3.14 2.13 1.88 2.58 2.33 1.86
CH4/CO2 17.97 24.83 24.81 16.83 20.28 25.82
H2/CO2 0.00 0.00 0.00 0.00 0.00 0.00
Mass balance, wt% 101.33 100.16 102.61 103.03 99.73 100.17
Table G- 1 Operational conditions and experimental results for different thermal cracking runs
126
Run % Off Thermal 1-1 Thermal 1-2 Thermal 1-3 Thermal 2-1 Thermal 2-2 Thermal 2-3
0 163.8 163.1 162.7 163.1 163 162.6 5 225.6 195 172.3 214.4 185.6 169.7
10 277 256 222.1 267.3 246.3 216.5 15 301.7 287.4 259.6 295.7 281 253.9 20 315.9 305.5 286.7 311.6 302 281.6 25 327.2 318 304.3 323.4 315.7 301.1 30 337.2 328.5 317 333.4 327 314.7 35 346.6 338 328.1 343.1 337.1 326.3 40 355.5 347.1 338.2 351.9 346.8 336.8 45 364.6 355.7 347.9 360.8 356 346.9 50 373.8 364.6 357.4 369.8 365.4 356.6 55 383.7 373.7 367.2 379.4 375.3 366.6 60 394.2 383.5 377.8 389.6 385.8 377.4 65 405.2 393.8 389.1 400.3 397.2 389.1 70 416.6 404.8 401.4 411.5 409.2 401.8 75 428.9 416.3 414.6 423.3 421.9 415.5 80 443.7 428.7 429.4 436.8 436.8 431.1 85 462.9 443.8 448.6 454.1 456.4 451.8 90 494.5 463.6 476.8 478.6 486.5 483.9 95 -- 496.8 547.8 537.6 590.2 584.4 100 -- -- -- -- -- --
Table G- 2 SimDist results for thermal cracking runs
Compound \ Test Thermal 1-1 Thermal 1-2 Thermal 1-3 Thermal 2-1 Thermal 2-2 Thermal 2-3 Hydrogen ND ND ND ND ND ND
Methane HAYD 56.37 52.82 46.67 43.38 47.28 48.07
CO2 3.14 2.13 1.88 2.58 2.33 1.86
Ethylene 0.13 0.00 2.32 0.25 0.19 1.40
Ethane 19.91 21.48 22.94 22.20 22.14 23.30
Propylene 3.22 3.92 4.07 3.50 2.88 2.63
Propane 11.32 13.02 13.35 16.36 15.22 14.09
i-Butane/ 1 Butene 0.40 1.03 3.40 4.28 3.48 3.25
n-Butane 5.52 5.60 5.36 7.46 6.49 5.39
i-Pentane 0.00 0.00 0.00 0.00 0.00 0.00
TOTAL 100.00 100.00 100.00 100.00 100.00 100.00
Table G- 3 Gas chromatography for thermal cracking runs
127
Appendix H: Operational data and Experimental Results for Steam Cracking Runs
Test Steam 1-1 Steam 1-2 Steam 1-3 Steam 2-1 Steam 2-2 Steam 2-3 Steam 2-3a
Date 03-Aug-06 21-Jul-06 27-Jul-06 10-Aug-06 23-Aug-06 31-Jul-06 23-Feb-07
Length of mass balance, h 2.00 2.00 2.00 2.00 2.00 2.00 1.50
Pressure, psig 159 161 161 261 260 259 260
Temp.,°C 410 420 430 410 420 430 430
VHSV, hr-1 2 2 2 2 2 2 1
Feed flow pumped, cc/h 200.58 203.90 202.24 199.70 200.36 201.36 100.00
Water flow pumped, cc/h 10.00 10.00 10.00 10.00 10.00 10.00 5.00
Water injection, wt% 5.01 4.93 4.97 5.03 5.01 4.99 5.03
Total mass pumped, g 399.23 405.51 402.37 397.57 398.82 400.71 149.02
Total liquid product, g 377.96 380.62 365.26 370.50 368.67 360.28 133.35
Product gas flow, cc/min 105.70 138.49 195.53 151.42 174.86 206.96 141.39
Total gas product, g 14.87 19.34 27.30 22.06 27.08 29.96 14.62
Liquid yield, wt% 94.67 93.86 90.78 93.19 92.44 89.91 89.49
Gas yield, wt% 3.72 4.77 6.79 5.55 6.79 7.48 9.81
VGO SimDis at 350C, wt% 28.92 28.92 28.92 28.92 28.92 28.92 29.42
Product SimDis at 350C, wt% 40.16 45.26 53.47 40.37 45.15 52.50 57.28
Conversion 350+C 20.30 27.71 40.58 21.82 28.67 39.92 45.83
Distillate Selectivity, wt % 12.80 19.08 27.60 12.24 18.03 25.72 30.94
Gas Selectivity, wt % 5.24 6.71 9.55 7.81 9.55 10.52 13.90
Carbon MCR, wt% 0.40±0.09 0.82±0.59 2.18±0.05 0.52±0.14 1.15±0.21 2.75±0.11 5.12±1.34
%H2 8.92 8.91 11.31 7.44 5.53 6.83 8.80
%CH4 42.93 46.47 46.11 42.79 38.98 43.13 45.70
%CO2 2.56 1.80 1.76 3.04 1.96 1.75 3.31
CH4/CO2 16.74 25.84 26.17 14.07 19.91 24.68 13.81
H2/CO2 3.48 4.96 6.42 2.45 2.82 3.91 2.66
Mas Balance, wt% 98.40 98.63 97.56 98.74 99.23 97.39 99.30
Table H- 1 Operational conditions and experimental results for different steam cracking runs
128
Run %Off Steam 1-1 Steam 1-2 Steam 1-3 Steam 2-1 Steam 2-2 Steam 2-3 Steam 2-3a
0 163.8 163.3 163.1 163.8 163.5 163.3 163.4 5 193.9 171.8 166.4 193.9 171.2 166.5 167.8 10 257.0 223.3 180.5 257.0 220.1 180.9 184.6 15 288.9 261.9 220.7 288.9 258.5 221.3 220.4 20 307.1 289.0 248.6 307.1 286.2 250.1 243.6 25 319.8 306.2 272.6 319.8 304.4 274.5 266.0 30 330.5 318.8 292.8 330.5 317.6 294.8 285.8 35 340.4 329.8 307.9 340.4 329.1 309.9 300.6 40 349.7 339.9 320.1 349.7 339.7 322.3 313.0 45 358.8 349.5 331.4 358.8 349.7 333.4 324.2 50 368.1 359.0 342.5 368.1 359.7 344.7 334.4 55 378.0 368.9 353.3 378.0 370.0 355.3 345.3 60 388.6 379.5 364.5 388.6 381.2 366.6 355.7 65 400.0 390.9 377.0 400.0 393.4 378.9 367.3 70 412.3 403.6 391.4 412.3 406.7 392.8 380.4 75 424.9 417.2 408.6 424.9 420.9 408.8 396.0 80 440.3 432.4 429.1 440.3 437.8 427.0 415.1 85 461.1 453.1 462.6 461.1 461.6 453.4 441.4 90 497.4 486.4 563.6 497.4 505.6 505.2 502.0 95 609.9 672.9
100
Table H- 2 SimDist results for steam cracking runs
Compound \ Test Steam 1-1 Steam 1-2 Steam 1-3 Steam 2-1 Steam 2-2 Steam 2-3 Steam 2-3a
Hydrogen MS 8.92 8.91 11.31 7.44 5.53 6.83 8.80
Methane 42.93 46.47 46.11 42.79 38.98 43.13 45.70
CO2 2.56 1.80 1.76 3.04 1.96 1.75 3.31
Ethylene 2.68 2.64 2.49 2.02 2.33 2.07 1.47
Ethane 18.65 20.56 19.31 19.53 21.06 20.68 19.50
Propylene 4.87 4.06 3.64 4.22 4.85 3.78 1.94
Propane 10.89 10.58 9.42 12.11 13.85 12.44 10.39
i-Butane/ 1 Butene 3.53 2.99 2.46 2.42 4.60 3.51 3.03
n-Butane 4.98 4.33 3.51 4.77 6.82 5.78 5.73
i-Pentane 0.00 0.25 0.00 1.66 0.03 0.03 0.13
TOTAL 100.00 102.59 100.00 100.00 100.00 100.00 100.00
Table H- 3. Gas chromatography for thermal cracking runs
129
Appendix I: Operational Data and Experimental Results for Catalytic Runs
Test UDCat 1-1 UDCat 1-2 UDCat 2-1 UDCat 2-2
Date 17-Feb-07 21-Feb-07 22-Mar-07 22-Mar-07 Length of the balance, h 1.50 2.00 2.00 1.50
Pressure, psig 260 260 260 260
Temp.,°C 430 430 430 440
VHSV, hr-1 2 1 1 1
Water content in feed, % 5.00 5.00 5.00 5.00
Catalytic matrix K-Ni K-Ni K-Ni-Ca K-Ni-Ca
Feed density, g/cc 0.952 0.952 0.953 0.95
Feed flow pumped, g 301.29 203.44 152.24 151.85
Total liquid product, g 272.49 176.36 133.29 138.92
Product gas flowt, cc/min 236.52 172.49 121.50 130.20
Total gas product, g 27.77 26.35 18.32 14.24
Liquid yield, wt% 90.44 86.69 87.55 91.49
Gas yield, wt% 9.22 12.95 12.03 9.38
VGO SimDist at 350C, wt% 28.92 28.92 28.92 28.92
Product SimDist at 350C, wt% 49.70 59.76 59.62 65.71
Conversion 350+C, wt% 36.00 50.92 50.26 55.87
Distillate Selectivity, wt% 22.55 32.20 32.75 43.89
Gas Selectivity, wt % 13.0 18.2 16.9 13.2
Carbon MCR, wt% 2.04±0.44 4.24±0.32 4.93±0.67 5.25±0.67
%H2 6.86 9.31 3.49 4.81
%CH4 38.27 37.54 45.41 45.23
%CO2 6.16 4.89 4.37 4.07
CH4/CO2 6.21 7.68 10.40 11.15
H2/CO2 1.12 1.90 0.80 1.21
Mas Balance, wt% 99.66 99.64 99.59 100.86
Table I- 1 Operational conditions and experimental results for different ultradispersed catalyst runs
130
Weight Percent Off, wt% / Run UDCat 1-1 UDCat 1-2 UDCat 1-3 UDCat 2-1 UDCat 2-2
0 163.3 163.3 163.4 164.2 164.0 5 172.4 167.4 166.6 167.5 167.0
10 216.2 183.3 177.0 179.4 174.2 15 249.2 216.2 206.7 211.9 194.4 20 275.0 241.6 231.9 235.7 222.7 25 295.1 263.6 254.5 258.9 243.3 30 309.3 282.1 274.2 278.1 263.8 35 320.8 297.4 292.2 294.9 282.3 40 331.1 309.7 306.4 308.1 297.7 45 341.1 320.4 318.5 319.3 310.6 50 350.6 330.4 329.9 330.0 322.1 55 360.2 340.4 341.3 340.3 333.3 60 370.2 350.4 352.8 350.8 344.7 65 381.3 360.6 365.2 361.6 356.7 70 393.6 371.8 379.4 373.6 370.0 75 407.3 384.8 396.5 387.7 386.2 80 422.4 400.1 417.6 404.6 406.8 85 442.0 418.2 448.8 425.2 439.3 90 473.5 443.6 532.6 458.2 522.4 95 569.1 496.3 -- 550.2 619.0 100 656.6 674.2 -- -- --
Table I-2 SimDist results for steam ultradispersed catalyst runs
Compound \ Test UDCat 1-1 UDCat 1-2 UDCat 1-3 UDCat 2-1 UDCat 2-2
Hydrogen MS 6.86 9.31 7.34 3.49 4.81
Methane 38.27 37.54 40.62 45.41 45.23
CO2 6.16 4.89 4.37 4.37 4.07
Ethylene 2.09 1.27 1.16 1.31 1.23
Ethane 16.92 17.71 18.46 20.69 19.30
Propyllene 4.94 4.27 3.76 1.40 1.88
Propane 12.66 14.17 13.95 12.99 12.67
i-Butane/1-Butene 4.13 3.02 2.80 0.46 1.28
n-Butane 7.73 6.30 6.11 9.26 9.27
I-Pentane 0.05 1.52 1.41 0.48 0.19
Table I- 3 Gas chromatography for ultradispersed catalysts runs
131
Test FixBed1 FixBed2 FixBed3 FixBed4 FixBed5
Date 01-Mar-07 01-Mar-07 01-Mar-07 01-Mar-07 01-Mar-07 Length of the balance, h 1.50 2.00 2.00 1.50 1.50
Pressure, psig 260.00 260.00 260.00 260.00 260.00
Temp.,°C 430.00 430.00 430.00 430.00 440.00
MHSV, g/gh 1.00 0.75 0.50 1.00 1.00
VHSV, hr-1 2.30 1.72 1.15 2.30 2.30
Water injection, % 5.9 7.47 11.03 5.87 5.82
Catalytic matrix K-Ni-Ca K-Ni-Ca K-Ni-Ca K-Ni-Ca K-Ni-Ca
Feed flow pumped, gr 118.92 123.56 80.50 120.04 121.21
Water pumped, g 7.73 9.98 9.98 7.49 7.49
Total liquid product, gr 115.66 118.41 85.25 119.34 115.88
Product gas flow, cc/min 208.12 158.67 103.64 163.59 167.27
Total gas product, g 13.22 13.12 8.27 13.34 12.79
Liquid yield, wt% 91.32 88.69 94.23 93.58 90.04
Gas yield, wt% 11.11 10.62 10.27 11.11 10.55
VGO SimDist at 350C, wt% 28.92 28.92 28.92 28.92 28.92
Product SimDist at 350C, wt% 52.54 49.09 51.46 42.56 47.03
Conversion 350+C, wt% 38.95 36.43 35.64 24.38 32.91
Carbon MCR, wt% 2.69 2.11 2.36 2.27 2.07
%H2 55.02 54.47 55.62 34.06 42.08
%CH4 17.51 18.68 20.54 30.97 23.80
%CO2 3.12 3.62 3.53 6.17 5.05
CH4/CO2 5.61 5.16 5.83 5.02 4.71
H2/CO2 17.70 15.07 15.78 5.52 8.33
Mass Balance, wt% 102.43 99.30 104.50 104.70 100.59
Table I- 4 Operational conditions and experimental results for different catalytic fixed bed runs
Compound \ Test FixBCat 1 FixBCat 2 FixBCat 3 FixBCat 4 FixBCat 5 Hydrogen MS 55.02 54.47 55.62 34.06 42.08 Methane 17.51 18.68 20.54 30.97 23.80 CO2 3.12 3.62 3.53 6.17 5.05 Ethylene 0.43 0.62 0.42 0.08 1.26 Ethane 8.36 8.47 8.05 11.03 10.19 Propyllene 1.17 1.28 0.79 2.33 2.73 Propane 6.69 6.20 3.14 6.69 6.90 i-Butane/1-Butene 2.36 2.12 3.50 2.88 2.90 n-Butane 5.21 4.41 2.51 4.52 4.90
I-Pentane 0.08 0.08 1.75 0.00 0.01
Table I-1Gas chromatography for fixed bed catalyst runs
132
Appendix J: Physical Properties of Athabasca Vacuum Gasoil
Weight Percent Off, wt% / Run
Athabasca VGO
0 163.3 5 172.4
10 216.2 15 249.2 20 275.0 25 295.1 30 309.3 35 320.8 40 331.1 45 341.1 50 350.6 55 360.2 60 370.2 65 381.3 70 393.6 75 407.3 80 422.4 85 442.0 90 473.5 95 569.1 100 656.6
Table J-1 Simulated distillation of Athabasca vacuum gasoil
Physical propierty Value Density 225ºC, g/cc 0.946 Viscosity @ 20ºC, cP 371.1 Viscosity @ 21ºC, cP 340.6 Viscosity @ 22ºC, cP 311.2 MCR, wt% 0.25 Sulfur content, wt% 3.3 Nitrogen content, wt% 0.11
TableJ-2 Physical properties of Athabasca VGO
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