Ref. AXENS A0066-2011
FIXED BED
CATALYST HANDBOOK
MB4-A.DOT
3
2
May 2013
April 2011
1
April 2004
Rev.
Date
M. THERY
M. THERY
PY. LEGOFF
F. CHOPINET
J. COOK
J. DE BONNEVILLE
X. DECOODT
JM. DEVES
H. DEVILLE
P. DUHAUT
PY. LEGOFF
F. LEPELTIER
R. ODELLO
M. THERY
Redacted by
X. DECOODT
P. MEGE
PY. LEGOFF
J. DE BONNEVILLE
H. DEVILLE
A. LECORRE
Checked by
Approved by
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TABLE OF CONTENT
1. PREFACE ....................................................................................................... 4
2. CHEMICAL REACTIONS AND CATALYST ......................................................... 6
2.1 Introduction ................................................................................................ 7
2.1.1 Objective of chapter 2 ................................................................................ 7
2.1.2 General considerations - Thermodynamics and kinetics ........................... 7
2.2 Chemical reactions ...................................................................................... 8
2.2.1 Fundamental reactions............................................................................... 8
2.2.2 Kinetic analysis of the chemical reactions................................................ 16
2.3 Catalyst ...................................................................................................... 22
2.3.1 Activity, selectivity, stability ..................................................................... 22
2.3.2 Reforming catalyst characteristics ........................................................... 22
2.3.3 Catalysis mechanism ................................................................................ 23
2.3.4 Catalyst contaminants .............................................................................. 26
2.3.5 Catalyst distribution in reactors ............................................................... 37
2.4 Process variables ....................................................................................... 39
2.4.1 Independent variables .............................................................................. 39
2.4.2 Pressure .................................................................................................... 39
2.4.3 Temperature ............................................................................................. 40
2.4.4 Space velocity ........................................................................................... 40
2.4.5 Hydrogen to hydrocarbon ratio and hydrogen partial pressure .............. 41
2.4.6 Feed quality .............................................................................................. 41
3
START-UP PROCEDURE FOR FRESH CATALYST ............................................. 44
3.1 Preparation of the unit - General considerations – Unit Dry-out............. 45
3.2 Catalyst loading ......................................................................................... 46
3.3 Catalyst drying........................................................................................... 46
3.4 Catalyst reduction ..................................................................................... 48
3.5 Catalyst sulfiding ....................................................................................... 49
3.6 Oil-in .......................................................................................................... 50
3.7 Startup phase ............................................................................................ 51
3.7.1 First temperature level: 460°C ................................................................. 52
3.7.2 Second temperature level: 470°C ............................................................ 52
3.7.3 Third temperature level: 480°C ................................................................ 52
3.7.4 High severity operation:  480°C ............................................................. 52
3.7.5 WAIT increase summary ........................................................................... 53
4
NORMAL OPERATION ................................................................................. 58
4.1 Change of feed rate .................................................................................. 59
4.1.1 Increase of feed rate ................................................................................ 59
4.1.2 Decrease of feed rate ............................................................................... 59
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4.2
4.3
4.4
4.5
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Change of reformate octane number ....................................................... 60
Recycle gas water content. Water and chloride injection ........................ 60
Sulfur content – hydrogen sulfide concentration ..................................... 62
Operating parameters............................................................................... 62
4.5.1 Pressure .................................................................................................... 62
4.5.2 Temperature ............................................................................................. 62
4.5.3 Space velocity ........................................................................................... 64
4.5.4 Hydrogen to hydrocarbon ratio ............................................................... 65
4.5.5 Feed quality .............................................................................................. 66
4.5.6 Butane content of the reformate ............................................................. 68
4.5.7 Start of run WAIT calculation ................................................................... 68
4.5.8 Cycle length .............................................................................................. 69
4.5.9 Delta C5+ yield (wt %) for different feedstocks/versus RONC .................. 70
4.6 Troubleshooting ........................................................................................ 85
4.6.1 General ..................................................................................................... 85
4.6.2 Unexpected decrease of the octane number .......................................... 85
4.6.3 Loss of product yield ................................................................................ 87
4.6.4 Unexpected T reduction ........................................................................ 88
4.6.5 High hydrocracking rate and risk of temperature runaway ..................... 89
5
SHUTDOWN / RE-STARTUP PROCEDURES ................................................... 92
5.1 Normal shutdown ..................................................................................... 93
5.2 Re-startup.................................................................................................. 93
5.3 Shutdown for regeneration ...................................................................... 94
5.4 Emergency shutdown ............................................................................... 94
5.4.1 Recycle compressor failure ...................................................................... 95
5.4.2 Loss of feed ............................................................................................... 95
5.4.3 Other pumps failure ................................................................................. 95
5.4.4 Utilities failure: fuel gas ............................................................................ 96
5.4.5 Utilities failure: cooling water .................................................................. 96
5.4.6 Utilities failure: power supply .................................................................. 96
5.4.7 Utilities failure: HP steam ......................................................................... 96
5.4.8 Utilities failure: instrument/power or air ................................................. 96
5.4.9 Major leak - fire ........................................................................................ 97
6
REGENERATION .......................................................................................... 98
6.1 General ...................................................................................................... 99
6.2 Preparation of the unit ........................................................................... 101
6.3 Coke combustion - oxychlorination - calcination ................................... 104
6.3.1 Coke combustion .................................................................................... 105
6.3.2 Catalyst sieving and reloading ................................................................ 106
6.3.3 Catalyst reactivation ............................................................................... 107
6.4 Emergency handling procedure for regeneration .................................. 110
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6.5 Sulfur and Sulfate removal procedures .................................................. 111
6.5.1 Sulfur removal procedure ...................................................................... 111
6.5.2 Sulfate removal procedure ..................................................................... 112
6.6 Miscellaneous ......................................................................................... 114
7
SAFETY AND HEALTH RECOMMENDATIONS .............................................. 122
7.1 General .................................................................................................... 123
7.2 List of health and safety data sheets ...................................................... 123
7.3 Catalyst safety data sheet ....................................................................... 123
8
ANALYTICAL CONTROL.............................................................................. 124
8.1 Recommended methods and frequency ................................................ 125
8.1.1 Feed ........................................................................................................ 125
8.1.2 Products .................................................................................................. 126
8.1.3 Catalyst ................................................................................................... 128
8.2 IFP Analytical methods ............................................................................ 128
9
MISCELLANEOUS ...................................................................................... 130
9.1 Chemicals specifications ......................................................................... 131
9.2 TBP - ASTM Boiling range transformation .............................................. 135
9.3 Reformate RVP versus butane content .................................................. 138
10 TECHNICAL ASSISTANCE SERVICES FOR FIXED BED UNIT............................ 140
10.1 Catalyst performance estimation on site ............................................... 141
10.2 Unit follow up at Axens offices ............................................................... 141
10.3 Catalyst analysis ...................................................................................... 141
10.4 Follow up of the regeneration and the start up ..................................... 142
10.5 Catalyst optimisation performance ........................................................ 142
10.6 Operator training .................................................................................... 142
10.7 Training simulators ................................................................................. 142
10.8 Catapac – Texicap – Catalyst sampler ..................................................... 143
11 TYPE OF REACTORS – CATALYST LOADING ................................................ 151
11.1 Type of reactors ...................................................................................... 152
11.2 Catalyst loading procedure .................................................................... 158
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1. PREFACE
REV. 4
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This handbook is not an operating manual.
Its purpose is to provide USER staff with the necessary background and information to
understand how the process works with RG / PR series Catalysts.
It also gives brief instructions on how to prepare the unit for start-up, how to start-up,
how to operate, how to shut it down, how to regenerate the catalyst and how to prevent
or correct operational upsets.
The information supplied in this document is valid for any unit using RG / PR series
catalysts.
The purpose of Reforming process using these catalysts is to produce a high octane
number reformate, which is a main component to gasoline pool, and a hydrogen rich gas.
Reformer feed is either straight run naphthas or cracked naphthas generally mixed with
straight run. Due to the presence of contaminants in all cases and to the specific
characteristics of cracked naphthas, more or less elaborate naphtha pretreating is always
necessary.
A high temperature (in the range of 500°C) is required to promote the chemical reactions
which improve octane number. Hence the need for a preheating of the feed. More over,
some of the desirable reactions are highly endothermic. This leads to split the inventory
of the catalyst into several reactors with intermediate heaters.
RG / PR series catalysts are operated with average reaction pressures between 12 and 30
barg. Octane Number (RON clear) of the reformate as high as 102 can be obtained.
Before implementing any special procedure not included in this handbook contact first
one AXENS Technical Assistance Adviser or our web site www.axens.net.
THIS DOCUMENT CONTAINS AXENS’
CONFIDENTIAL INFORMATION.
IT SHALL NOT BE REPRODUCED IN WHOLE OR IN PART.
IT SHALL BE USED ONLY BY STAFF WITHIN YOUR COMPANY
REQUIRING THE INFORMATION.
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2. CHEMICAL REACTIONS AND CATALYST
REV. 4
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2.1
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Introduction
2.1.1 Objective of chapter 2
The aim of the information given in this chapter is to provide enough theoretical
background, in the simplest possible way, to supplement the instructions given in the
chapters that follow, i.e. Start-up of unit, Operation of the unit and Shutdown of the unit.
It is expected that this theoretical support helps the operators to better understand the
reasons of the operating instructions and enables them to make considered decisions,
should the circumstances deviate from what is covered in the Operating Instructions.
2.1.2 General considerations - Thermodynamics and kinetics
For any chemical reaction the thermodynamics dictates the possibility of its occurrence
and the amount of products and unconverted reactants. In fact, some reactions are
100% completed i.e. all the reactants are converted into products. Others are in
equilibrium i.e. part of the reactants only are converted. The amount of products and
reactants at equilibrium depends upon the operating conditions and is dictated by the
thermodynamics. Note that thermodynamics do not mention the time required to reach
the equilibrium or the full completion of a reaction.
Kinetics dictates the rate of a chemical reaction. Kinetics is dependent upon operating
conditions but can also be widely modified through the use of properly selected
catalysts. One reaction (or a family of reactions) is generally enhanced by a specific
catalyst.
In other words, thermodynamics dictates the ultimate equilibrium composition assuming
the time is infinite. Kinetics enables to forecast the composition after a finite time. Since
time is always limited, when several reactions proceed simultaneously, kinetics is
generally predominant.
A heterogeneous catalyst generally consists of a support (alumina, silica, magnesia...) on
which (a) finely divided metal(s) is (are) dispersed.
The metal is always responsible for the catalytic action. Very often, the support has also
a catalytic action linked to its chemical nature.
A catalyst is not consumed but can be deactivated either by impurities in the feed or by
some of the products of the chemical reactions involved, resulting in coke deposit on the
catalyst.
The different chapters of the section describe:
• The various chemical reactions involved in the process as well as the effect of
the operating conditions.
• The catalyst characteristics.
• The catalytic mechanism or catalysis.
• The catalyst contaminants.
• The process variables.
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2.2
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Chemical reactions
2.2.1 Fundamental reactions
The chemical reactions involved in reforming processes are of two types:
• Desirable reactions, i.e. reactions which lead to an increased octane number
and to high purity hydrogen production. These are the reactions to promote.
• Adverse reactions, i.e. reactions which lead to a decreased octane number, a
decrease in hydrogen purity or a loss in products yield. These are the reactions
to minimize.
The heats of the reactions mentioned hereafter as well as their relative rate are
necessary to understand the process. They are listed for the ease of reference in Table 1,
below. A catalyst is being used to promote the desirable reactions at the expense of the
adverse ones through its action on reaction kinetics.
TABLE 1
REFORMING REACTIONS
HEAT OF REACTION - RELATIVE RATE OF REACTION
REACTIONS
Naphthenes dehydrogenation
Paraffin dehydrocyclization
Isomerization:
Paraffins
Naphthenes
Cracking
(1)
(2)
REV. 4
HEAT OF REACTION
(1) KCAL/MOLE
RELATIVE RATE
(2) APPROX.
- 50
- 60
+2
+4
+ 10
30
1 (base)
3
0.5
Heat of reaction < 0 = endothermic reaction.
For pressure below 15 bar.
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A
9/ 172
Desirable reactions with hydrogen production
a) Naphthenes dehydrogenation
Naphthenic compounds i.e. cyclohexane, methylcyclohexane, dimethylcyclohexane up to
C10 naphthenes are dehydrogenated respectively into benzene, toluene, xylenes, C 9 and
C10 aromatics with the production of 3 moles of hydrogen per mole of naphthene.
The cyclohexane reaction, for instance, proceeds as follows:
CH
CH
2
CH
HC
2
2
(m)
HC
CH
+ 3H
HC
2
CH
CH
2
2
Cyclohexane
HC
2
CH
CH
Benzene
(m) Catalyst Metallic function
(a) Catalyst Acidic function
Note: Cyclohexane and benzene are generally schematically represented as follows:
Cyclohexane
Benzene
Thermodynamically the reaction is highly endothermic and is favored by high
temperature and low pressure. In addition the higher the number of carbon atoms, the
higher the aromatics production at equilibrium.
From a kinetic view point, the rate of reaction increases with temperature (Refer to
figure 2-2) and is not affected by the hydrogen partial pressure (Refer to figure 2-1). The
rate of reaction is high compared to other reactions (Table 1). It also increases with the
number of carbon atoms.
At the selected operating conditions the reaction almost total. It is promoted by the
metallic function of the catalyst. Since it yields a high octane product, promoting this
reaction is most desirable: refer to octane number below:
REV. 4
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Note that cyclopentane, present in feeds having a low ASTM D86 IBP, is an undesirable
component as either it does not react or its cycle is broken and results in low carbon
parafins or olefinsand hydrogen consumption.
RON
MON
Cyclohexane
83
77.2
Methylcyclohexane
74.8
71.1
1.3 dimethylcyclohexane
71.7
71.0
Benzene
114.8
> 100
Toluene
120
103.5
m-Xylene
117.5
115.0
RON:
MON:
Research Octane Number
Motor Octane Number
Throughout this document, “octane” is generally used for “octane number”.
b) Effect of parameters on naphthene dehydrogenation
The tables below summarize the effect of the main parameters governing the
dehydrogenation and dehydrocyclization reactions.
Thermodynamics dictates the equilibrium which could be theoretically reached (i.e. if the
time was infinite). Kinetics dictates the rate of reaction, i.e. the possibilities to reach a
state close to equilibrium in a finite time.
Increase of
Pressure
Temperature
H2/HC ratio (1)
Effect on dehydrogenation due
to thermodynamics
to kinetics
decreases
unaffected
increases
increases
slightly decreases
slightly decreases
(1) Ratio of pure hydrogen (mole) to hydrocarbon feed (mole).
c)
Paraffins dehydrocyclization
This is a multiple step process which applies either to the normal paraffins (linear) or isoparaffins (branched). It involves a dehydrogenation with a release of one hydrogen mole
followed by a molecular rearrangement to form a naphthene and the subsequent
dehydrogenation of the naphthene. The molecular rearrangement to build a naphthene
is the most difficult reaction to promote but the subsequent aromatization of the
naphthene yields a noticeable octane increase.
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The reaction can be summarized as follows:
CH
2
CH
CH
2
3
CH
2
CH
2
CH
2
CH
3
CH
2
CH
(m)
CH
3
CH
2
CH
CH
CH
2
CH
2
CH
CH
2
CH
3
2
CH
3
2
C H
7 14
C H
7 16
CH
+H
CH
2
CH
(a)
H C
2
CH
3
CH
2
CH
CH
2
CH
3
2
Methylcyclohexane
CH
2
CH
CH
2
CH
H C
2
CH
2
CH
2
CH
C
CH
3
(m)
HC
CH
CH
3
+ 3H
2
CH
Toluene
The paraffin dehydrocyclization step becomes easier as the molecular weight of the
paraffin increases, however the tendency of paraffins to hydrocrack increases
concurrently (Refer to figure 2-3).
Kinetically, the rate of dehydrocyclization increases with low pressure and high
temperature (figures 2-1 and 2-2), but altogether, at the selected operating conditions,
this rate is much lower than that of naphthene dehydrogenation (30/1). The reaction is
promoted by both catalytic metallic and acidic functions.
d) Effect of parameters on paraffin dehydrocyclization
Increase of
Pressure
Temperature
H2/HC ratio
REV. 4
Effect on dehydrocyclization due
to thermodynamics
to kinetics
decreases
decreases
increases
increases
slightly decreases
slightly decreases
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B
12/ 172
Desirable reactions without hydrogen production
a) Linear paraffins isomerization (a)
Reaction is as follows:
CH
7 16
CH
7 16
These reactions are fast, slightly exothermic and do not affect the number of carbon
atoms. The thermodynamic equilibrium of isoparaffins to paraffins depends mainly on
the temperature. The pressure has no effect.
Iso-N paraffin equilibria
Carbon atom
C4
C5
C6
% Isoparaffin at 500°C
44
58
72
C7
80
C8
88
The paraffins isomerization results in a slight increase of the octane number. From a
kinetic view point (figures 2-1 and 2-2), high temperature favors isomerization but
hydrogen partial pressure has no effect. These reactions are promoted by the acidic
function of the catalyst support.
b) Napththenes isomerization
The isomerization of an alkylcyclopentane into an alkylcyclohexane involves a ring
rearrangement and is desirable because of the subsequent dehydrogenation of the
alkylcyclohexane into an aromatic. Owing to the difficulty of the ring rearrangement, the
risk of ring opening resulting in a paraffin is high.
The reaction is slightly exothermic. The reaction can be summarized as follows:
(a)
CH3
CH3
Alkylcyclopentane
(Ethylcyclopentane)
Alkylcyclohexane
(Methylcyclohexane)
Theoretically, at the selected operating temperature (about 500°C) the thermodynamics
(
limits the alkylcyclohexane formation. But the( subsequent dehydrogenation of the
alkylcyclohexane into an aromatic shifts the reaction towards the desired direction. This
type of reaction is also easier for higher carbon number.
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The octane number increase is significant when considering the end product (aromatics)
as shown:
RON
MON
• Ethylcyclopentane
=
67.2
61.2
• Methylcyclohexane
=
74.8
71.1
• Toluene
=
120
103.5
C
Adverse reactions
a) Cracking
Cracking reactions include hydrocracking and hydrogenolysis reactions.
Hydrocracking affects either paraffins (normal or iso) or olefins. It involves both the acid
and metallic function of the catalyst. It is, to some extent, a parallel reaction to paraffin
dehydrocyclization.
It can be represented schematically by a first step of dehydrogenation which involves the
metallic function of the catalyst, followed by a cleavage of the resulting olefin and the
hydrogenation of the subsequent short chain olefin. The second reaction is promoted by
the acidic function of the catalyst.
(m)
+H
2
CH
7 14
CH
7 16
(a)
+
+H
2
C H
3 8
C H
4 8
C H
7 14
(m)
+H
2
CH
4 8
CH
4 10
The first reaction involves the same reactants as the dehydrocyclizationand is likewise
catalysed by the metallic function.
At the selected operating conditions, hydrocracking reaction could be almost complete.
Fortunately it is somewhat limited by its kinetics. Compared to its desirable concurrent
reaction (dehydrocyclization), hydrocracking becomes significant as the temperature
increases. It is also favored by high pressure.
The main effects of hydrocracking are:
• a decrease of paraffins in the reformate which results in an increase of the
aromatics percentage (i.e. an increase in octane) and a loss of reformate.
• a decrease in hydrogen production.
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• an increase of LPG production.
b) Hydrogenolysis
This undesirable reaction has some similarity with hydrocracking since it involves
hydrogen consumption and cleavage of bonds. But it is promoted by the metallic
function of the catalyst and leads to lighter hydrocarbon C1 + C2 - even less valuable than
LPG (C3 + C4).
It can be represented schematically as follows:
(m)
+H
CH
2
4
+
CH
6 14
CH
7 16
or
(m)
+H
CH
7 16
2
CH +
2 6
CH
5 12
Like hydrocracking it is exothermic and favored by high pressure and high temperature.
Naphten’s rings opening is also under (m)
control of metal function.
or
CH3 + H2
C6 - H14
CH3 - C5H9
(m)
or
+ H2
CH3 - C6H11
REV. 4
or
C7 - H16
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c)
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Hydrodealkylation
Hydrodealkylation is the breakage (or cleavage) of the branched radical (-CH3 or -C2H5) of
an aromatic ring.
Xylene (two radical groups) can be dealkylated into toluene (one radical group) which in
turn can be dealkylated to benzene.
The standard representation is:
(m)
+H
2
+CH4
H3C
Xylene
+H
2
Toluene
(m)
Toluene
+ CH
4
Benzene
Hydrodealkylation consumes hydrogen and produces methane. It is favored by high
temperature and high pressure and promoted by the metallic function of the catalyst.
d) Alkylation
Alkylation is a condensation reaction which adds an olefin molecule on an aromatic ring.
It results in an aromatic with an increased molecular weight. The reaction proceeds as
follows:
CH
(m)
+ CH = CH - CH
3
2
HC
CH
Benzene
Propylene
3
3
Isopropylbenzene
This reaction, promoted by the catalyst metallic function, is not hydrogen consuming.
But it leads to heavier molecules which may increase the end point of the product. In
addition the high molecular weight hydrocarbons also have a high tendency to form
coke. This reaction must be avoided.
e) Transalkylation (Alkyl disproportionation)
Two toluene rings (one branched CH3 radical) can disproportionate to produce one
benzene ring (no branched radical) and one xylene ring (two branched radicals), as
(m)
shown:
+
+
Toluene
REV. 4
Toluene
Benzene
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This reaction, promoted by the catalyst metallic function, occurs mainly in very severe
conditions of temperature and pressure.
f)
Coking
Coke formation on the catalyst results from a very complex group of chemical reactions,
the detailed mechanism of which is not fully known yet.
Coke formation is linked to heavy unsaturated products such as polynuclear aromatics
(or polycyclics which can be dehydrogenated) resulting either from the feed or from the
polymerization of aromatics involved in some of the reforming reactions
(dehydrocyclization, disproportionation...). Traces of heavy olefins or diolefins may also
result from the reforming reactions (dehydrocyclization, alkylation, for instance) and
promote coke formation.
A high end boiling point of the feed means greater amount of polyaromatics and then a
higher coking tendency. Since condensation is promoted by high temperature, poor
distribution in a reactor favors local high temperatures and coke build up.
Coke deposit on the catalyst reduces the active surface area and greatly reduces catalyst
activity.
2.2.2 Kinetic analysis of the chemical reactions
The effect of the main operating conditions on the rate of the reactions involved in the
reforming process using the selected catalyst is summarized below.
A
Effect of hydrogen partial pressure
Figure 2-1 shows, with a logarithmic scale, the relative rate of the various reactions as a
function of hydrogen partial pressure. The dehydrogenation rate is used as reference and
taken at 100 (Log 100 = 2). Other reaction rates are measured against this reference.
At 10 barg hydrogen partial pressure, the dehydrogenation of naphthene is about 10
times, faster than isomerization, 30 times faster than dehydrocyclization and 50-60 times
faster than cracking (hydrocracking and hydrogenolysis).
At relatively high pressure (above 20 barg) the rate of coking is low compared to the
other reactions but it increases noticeably at lower pressure.
To sum up, figure 2-1 shows that there is an incentive to operate at low pressure:
cracking rate will be reduced and dehydrocyclization rate increased as well as the
coking rate.
On another hand thermodynamics also favors low pressure for dehydrogenation and
dehydrocyclization. The only drawback of low pressure is the high coking rate.
B
Effect of temperature
Temperature influences the rate of the various reactions as shown in Figure 2-2. Energy
of activation is calculated from the slope of the curves. Dehydrogenation has a moderate
energy of activation (~ 20 Kcal. mole -1) as does isomerization (~ 25 Kcal. mole -1) and
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consequently temperature only slightly increases the rate of these reactions.
Dehydrocyclisation has a higher energy of activation (35 Kcal. mole -1) and consequently
temperature increases the rate of reaction.
Cracking and coking have higher energy of activation (45 and 35 Kcal. mole-1
respectively). The rate of these undesirable reactions is more significantly increased by
temperature.
To sum up, a higher temperature clearly favors the undesirable reactions more than
the desirable one. However a moderate temperature rise is required during the
catalyst life to maintain catalyst activity and therefore product octane.
C
Effect of carbon number
The kinetic study of the chemical reactions becomes even more complicated owing to
the presence of molecules with different numbers of carbon atoms.
As is the case for thermodynamic equilibria, it appears that the rates of the reactions are
affected by the length of the chain of the reactant. Figure 2-3, presents the rates of
dehydrocyclization and cracking of C6 to C10 paraffins related to that of n-heptane, as a
function of the number of carbon atoms of reactant.
Figure 2-3 shows that the cracking reaction rate, (the curve represents in fact the sum of
hydrocracking and hydrogenolysis), increases regularly with the number of carbon atoms,
whereas dehydrocyclization rate exhibits a sudden increase between hexane and
heptane as well as between heptane and octane, while the variation between the higher
homologues remains relatively slight.
To sum up, the dehydrocyclization of C6 paraffins to benzene is more difficult than that
of C7 paraffin to toluene, which itself is more difficult than that of C 8 paraffin to
xylenes. Accordingly the most suitable fraction to feed a reforming process is the C7- C10
fraction.
CONCLUSIONS:
From the above analysis it can be concluded:
a) Dehydrogenation reactions are very fast, about one order of magnitude faster
than the other reactions.
b) Low pressure favors all desirable reactions and reduces cracking. To
compensate the detrimental effect of low pressure on coking, low pressure
reformer requires continuous catalyst regeneration. For semi regenerative
reformer the recommended lowest operating pressure to have acceptable
cycle lengh is ~ 12 kg/cm2g.
(c) An increase in temperature favors the kinetics of dehydrogenation,
isomerization, dehydrocyclization, but accelerates the degradation reactions
(cracking, coking) even more. Consequently an increase in temperature leads
to an increased octane associated with a decrease in reformate yield.
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(d) The reaction rates of such important reactions as paraffins dehydrocyclization
increase noticeably with the number of carbon atoms. Cyclization is faster for
C8 paraffin than for C7, and for C7 than for C6. Consequently the C7 - C10 fraction
is the most suitable feed.
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RELATIVE RATE OF REACTION
VERSUS HYDROGEN PARTIAL PRESSURE
FIGURE 2-1
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RELATIVE RATE OF REACTION
VERSUS TEMPERATURE
FIGURE 2-2
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RELATIVE RATE OF REACTION
VERSUS NUMBER OF CARBON ATOMS
FIGURE 2-3
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Catalyst
2.3.1 Activity, selectivity, stability
The main characteristics of a catalyst other than its physical and mechanical properties
are:
• The activity which expresses the catalyst ability to increase the rate of the
reactions sought after. It is measured by the temperature at which the catalyst
must be operated to produce a reformate of a given octane number, for a given
feed and given operating conditions.
• The selectivity expresses the catalyst ability to favor desirable reactions rather
than others. It is practically measured by the C5+ reformate and hydrogen
yields, for a given feed and octane number, and given operating conditions.
• The stability characterizes the change with time of the catalyst performance (i.e.
activity, selectivity) when operating conditions and feed are stable. It is chiefly
the coke deposit which affects stability, through its inhibition of the catalyst
acidity and decrease of metal contact area. Traces of metal in the feed also
affect stability adversely. Stability is generally measured by the amount of feed
treated per unit weight of catalyst (i.e. m3 of feed per kg of catalyst). C5+ wt
reformate yield, at steady conditions, is also an indirect measure of the stability.
2.3.2 Reforming catalyst characteristics
The fixed bed Reforming catalysts are bimetallic catalysts consisting of platinum plus
rhenium promotor on an alumina support. Some of them can have an additional
promotor, as a consequence they are trimetallic catalysts.
The main features of these catalysts are:
• High purity alumina support having a strong mechanical resistance.
• High stability and selectivity due to the platinum associated with rhenium.
• High regenerability.
The combination of these qualities gives the following advantages:
• High reformate yield.
• High hydrogen yield.
• High on-stream factor (long cycle duration).
• Low catalyst inventory.
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2.3.3 Catalysis mechanism
A
Bifunctional catalyst
The catalyst affects reaction rates through its two different functions: metallic and
acidic, which promote different type of reactions.
Dehydrogenation and hydrogenation reactions are enhanced by the metal activity.
Structural rearrangements of the molecules (from linear to cyclic for instance) which
involve a reorganization of the carbon bonds are primarily catalyzed by the acidic
function of the support.
Because of its high activity in dehydrogenation and dehydrocyclization, platinum has
been selected for the base catalytic metal. Promotor has been added to improve catalyst
selectivity and stability.
The support is a high purity alumina (acidic function) which is chiefly active for the
cyclization of paraffins to aromatics and for isomerization reactions.
In short, the main reactions involved in reforming processes are catalyzed essentially
either by the acid support or the metal functions, as indicated below:
Dehydrogenation
Metallic function
Dehydrocyclization
Metallic + acidic functions
Isomerization
Metallic + acidic functions
Hydrogenolysis
Metallic function
Hydrocracking
Metallic + acidic function
The key point for good catalyst activity, selectivity and stability is the proper
balance between the two functions.
A1
Metallic function
For a maximum catalyst activity the metal must be highly dispersed on the alumina
support and under the minimum possible particle size (actual figure is in the range of
1.10-6 mm). This high dispersion and micrometric particle size which result from the
special manufacturing process must be maintained during the catalyst life by the use of
proper operating conditions and restored after regeneration.
In fact catalyst ageing and coke deposit require a temperature increase which favors
metal agglomeration and particle growth. The coke burning during the regeneration
results in further chlorine elution owing to combustion water. Therefore an
oxychlorination step is required to restore the chlorine level but also to favor the redispersion of the metal. A wet reduction leads to sintering of the metallic phase.
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The poisoning of the metallic function is covered in detail in chapter 2.3.4.
A2
Acidic function
A certain level of acidity of the catalyst support is required to promote some of the
desired reactions (isomerization, dehydrocyclization). The optimum level of acidity
changes somewhat with the desired performance (maximization of gasoline and
hydrogen yields or gasoline and LPG production).
The acidity of the catalyst is dependant on the amount of chlorine which is fixed and
presence of hydroxyl groups on alumina surface. This equilibrium is under controlled of
water / chloride balance.
During operation chlorine elution (leaching) from the catalyst is a function of the recycle
gas moisture. The chlorine content of the catalyst is kept constant by injecting a
chlorinated agent into the reformer feed. To be noted, chlorine favours the bonding of
metals/support. Consequently low chlorine levels are promoting metals sintering or
agglomeration.
The chlorine content of the catalyst must be in the range 0.9% to 1.1% wt. Moreover it is
known that alumina based catalysts require some moisture to activate the acidic
function.
A simplified representation of the catalyst support chemical structure, after chlorination
is as follows:
OH
Cl
OH
Al
Al
Al
O
O
The simplified theory, generally accepted today, suggests that the optimum acidity level
(which varies with the requested performance) is a function of concentration of the -OH
groups and the -Cl groups.
The relative concentrations, in turn are a function of the water and HCl content in the
recycle gas, because of the equilibrium which exists, between water and HCl in the
recycle and -OH and -Cl groups on the catalyst.
For the purpose of Reforming, this optimum acidity level is obtained for water content in
the recycle in the range 15 - 25 ppm vol. The associated HCl content should be
approximately 0.2 to 0.5 vol. ppm.
The basic recommendation is then to operate with 0.9 to 1.1 wt% of chloride on the
catalyst.
Starting from the optimum catalyst structure (the right balance of -OH vs. -Cl), an excess
of water in the recycle gas will shift the balance towards excess of -OH and reduces the
activity.
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This is referred to as chloride elution or leaching by water and can be represented
schematically as follows:
OH
Cl
Al
Al
+H O
2
OH
OH
Al
Al
O
+ HCl
O
An over-chloride catalyst under normal moisture level will tend to enhance
hydrocracking reactions with major production of C3 and C4.
A catalyst with normal chloride level and under too dry operating conditions (< 5-7 ppm.
Vol) will tend to enhance cracking reactions with major production of C1 and C2.
Of course, over-chlorination of the catalyst may result merely from the accidental
presence of chlorine, or uncontrolled addition, in the feed.
There are a couple of other occurrences worth mentioning:
• If a catalyst is excessively dry (i.e. it has been operated with a deficiency of
water for some time) it will exhibit a very high acidic function indicated by an
increased hydrocracking activity.
• If a catalyst excessively dry undergoes a water upset (amount of water in the
recycle over 50 ppm vol.) a situation may occur where the water displaces the
chlorine from the first reactors towards the last reactor with a subsequent
temporary increase of the acidic function and hydrocracking activity in the last
reactor.
To conclude, for an optimum operation of reformer unit:
• Highly dispersed metal clusters in the alumina support volume obtained by
optimal regeneration conditions
• The chloride content of the catalyst must be maintained between 0.9 to
1.1 wt%. Operators can adjust the chloride injection rate based on catalyst
analysis.
For units without catalyst sampling devices, control of the correct chloride
content on catalyst is more difficult. Optimum control is maintained by careful
monitoring of water and chloride injection rates, by monitoring of recycle gas
HCl content and ultimately by following the evolution of unit performance.
• The water content in the recycle must be maintained between 10 to 20 volume
ppm.
• HCl traces (0.2 to 0.5 volume ppm), measured by dedicated Dräeger tubes,
scale 0 to 10 vol. ppm, have to be detected in the recycle gas.
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Alteration of catalyst activity
The causes and consequences of catalyst activity loss (due to an unbalance of either the
acidic or metallic function) are listed in Table 2.
TABLE 2
Initial consequences
Causes
Decrease of acidic function
Increase of acidic function
Decrease of metallic function
• Elution of chloride due to high • Over chlorinated catalyst due to: • Temporary reversible poisoning
water content in the recycle
by sulfur
– chlorine in the feed
gas.
•
Permanent poisoning by metals.
– or too low water in recycle gas.
• Nitrogen compounds (loss of
• High water content in the recycle
Cl through NH4Cl) in the feed.
gas (upset) on a very dry catalyst
(the acidic function increase is
temporary).
• Decreased octane.
• Slight increase in octane.
• Large decrease in octane.
• Decreased LPG production.
• Decrease in liquid product and H2 • Decreased delta T in first reactor.
yields.
• Increased
Cl
production
• Decreased C1 + C2 production.
related to C1-C4 cut.
• Increased LPG production.
• Increased liquid production.
• Increased recycle gas H2 • Decreased C1 production related • Large decrease of H production.
2
purity.
to C1-C4 cut.
• Decreased recycle gas H2 purity.
• Increased liquid product yield. • Decreased recycle gas purity.
• Decreased T in last reactor.
2.3.4 Catalyst contaminants
Catalyst contaminants are classified in two categories. Temporary poisons (sometimes
called inhibitors) and permanent poisons.
Temporary poisons are those which can be removed from the catalyst without a
shutdown and for which the catalyst proper activity and selectivity is restored once the
contaminant disappears.
The effect of temporary poisons, if the operator maintains the operating conditions
prevailing before the poisoning, is a temporary decrease of performance.
The most common temporary poisons of reforming catalysts are sulfur, organic
nitrogen, water, oxygenated organics and halogens.
Permanents poisons are those which induce a loss of activity which cannot be recovered,
even with a regeneration and which is so severe that the catalyst must be replaced.
For conventional fixed bed catalysts as well as for continuously regenerated catalysts, the
main permanent poisons are arsenic, lead, copper, iron, nickel, chromium, sodium,
potassium.
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In order to ensure the optimum use of the catalyst, a proper design shall include:
• The removal of poisons from the feed prior to its introduction to the unit.
• The necessary procedures to remove, as far as possible, the temporary poisons
from the contaminated catalyst.
Impurities from the feed are removed by pretreating feed in adequatly designed units.
However their efficiency is never complete and generally limited depending upon the
type of impurities to be removed. In addition a poor adjustment of the operating
conditions of the pretreating unit results in a decreased efficiency.
A smooth and successful operation of the reforming unit requires the proper
adjustment and control of the operating conditions of the pretreating unit.
A
Temporary poisons
• Sulfur
Sulfur is the most common impurity found in the feed of any reforming unit. The
maximum allowable concentration is 0.5 ppm wt expressed as S. Whenever possible,
operation at lower sulfur content will provide additional catalyst stability and selectivity.
– Mechanism:
Poisoning is caused by H2S, either contained in the unit feed, or resulting from the
decomposition, on the catalyst, of sulfur compounds contained in the feed.
H2S reacts with platinum according to the equilibrium reaction:
Pt + H2S
PtS + H2
and with rhenium according to the equilibrium reaction:
Re + H2S
ReS + H2
and consequently it reduces the activity of the catalyst while decreasing the metallic
contact area.
The same type of reaction occurs with H2S towards the rhenium, and further reduces the
catalyst activity.
– Effect of sulfur contamination:
Sulfur contamination inhibits the metal function of the catalyst. This is indicated by:
• A decrease in hydrogen yield.
• A decrease in recycle gas purity.
• An increase in hydrocracking (LPG yield increase).
• An increased coking rate.
• A reduced temperature drop in the reactors, especially first one and sometimes
an increase T across the second reactor.
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Prevention and causes of contamination:
Sulfur removal is achieved by pretreating the naphtha feed, which results in H2S
production. Poor operation of the pretreater is generally the cause of sulfur poisoning of
the reforming catalyst:

Either low activity of the hydrotreater catalyst.

Sudden change of feed characteristics (EBP, total sulfur).

Too low hydrotreater reactor temperature or hydrogen partial pressure.

Or too high temperature at SOR conditions leading to possible
recombination of olefins with H2S.
which leads to an insufficient sulfur removal. Unsatisfactory operation of the
hydrotreater stripper can also result in dissolved H2S being fed to the reforming unit. In
such a case water content of the reforming feed also increases.
–
Detection:
Analytical methods are available to detect sulfur in the unit feed. A very easy way,
however, is to check sulfur content in the recycle gas using Draëger tubes. The H2S
content in the recycle which corresponds to the 0.5 ppm wt in the feed, is approximately
1 ppm volume. H2S detection can also be performed on the stabilizer column off-gas
(about 5 ppm volume in stabilizer column off-gas corresponding to 1 ppm volume in the
recycle gas).
–
Remedies:
When the sulfur content in the recycle gas increases, the reactor inlet temperature must
be preferably reduced but in all case not increase to compensate the loss of activity.
Typically for an H2S level of 5 vol. ppm in the recycle the reactor inlet temperature must
be lowered to 480°C; the reformer feed must be reduced accordingly to maintain
product quality (octane number). These conditions must be maintained until the cause
of the upset has been found and corrected. Contemporarily chloride injection shall be
increased by about 1 wt ppm.
The high severity operating conditions can only be resumed when the H 2S content in the
recycle gas is lower than 1 ppm vol.
In no instance shall the lost activity due to sulfur poisoning be compensated with
temperature.
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• Nitrogen
Nitrogen is less frequently present in the reforming feed than sulphur. Scarcely present in
straight run naphtha, nitrogen is a usual impurity of cracked naphtha and may also result
from injection of amine based corrosion inhibitors.
The maximum allowable concentration in the feed is 0.5 wt ppm expressed as organic
nitrogen.
Organic compounds containing nitrogen are responsible for inhibition but nitrogen gas
itself (N2) has no detrimental effect.
–
Mechanism:
Contamination is due to NH3 formed by decomposition of compounds containing organic
nitrogen, on the catalyst. Then NH3 which is alkaline reacts with chlorine decreasing the
acidic function of the catalyst and producing ammonium chloride NH 4Cl. This compound
is volatile in the conditions of the reactors and is eliminated inducing a loss of chlorine.
The reaction can be represented schematically as follows:
NH + H O +
3
2
Cl
OH
Cl
Al
Al
Al
O
and HCl + NH
–
O
3
Cl
OH
OH
Al
Al
Al
O
NH
4
+ HCl + NH
3
O
Cl
Effect of nitrogen contamination:
Nitrogen contamination reduces the acidic function and is indicated by:
 A decrease in octane.
 A slightly increased hydrogen production.
 A reduced reactor temperature drop.
On top of this, ammonium chloride in the recycle gas can deposit in coolers, separators,
stabilizer cold trays, creating mechanical problems, as it becomes solid under 80°C
It is worth remembering that 0.5 ppm wt of organic nitrogen in the feed leads to
approximately 2 T/year of NH4 Cl for a 1 106 T/y unit.
–
Prevention and causes of contamination:
Organic nitrogen removal is also achieved by naphtha pretreating of the feed. But it shall
be emphasized that nitrogen removal is more difficult than sulfur efficient nitrogen
removal often requires the use of a specific catalyst, also active for desulfurization, but
generally operating at higher hydrogen partial pressure. Naphtha, with high nitrogen
content must not be fed to a pretreater not designed for it.
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Cracked naphthas are generally characterized by high organic nitrogen content,
consequently cracked naphthas shall never be introduced to a pretreater designed to
process straight-run feeds without getting technical advice from the licensor and/or
the catalyst manufacturer.
In the pretreating unit, decomposition of nitrogen compounds gives NH3. However the
amount is generally limited and easy to remove by stripping. The corrosion inhibitor
(amine based), usually injected in the pretreatment stripper shall be selected to be
decomposed at the condition of the stripper overhead line to avoid contamination of the
stripper bottom product (Reforming feed). In fact, the presence of nitrogen compounds
in the feed is typically due to a low activity of the pretreatment catalyst towards
denitrification.
–
Detection:
There is no available method for ammonia detection in the recycle gas. Thus laboratory
analyses need to be performed on the feed to detect nitrogen compounds.
–
Remedies:
When nitrogen contamination is detected operators must:
 Increase the chlorination agent injection.
 Not try to make up for the drop in octane number of the reformate by an
increase of the reactor inlet temperature. This will only increase the loss of
chlorine
 Take the necessary actions to lower the nitrogen content down to the
acceptable figure of 0.5 ppm wt.
• Water and oxygenated organic compounds:
Oxygenated organic compounds (methanol, MTBE, TAME, phenol...) are converted into
water at reactor conditions.
Water is not exactly a poison since some water is necessary to activate the acidic
function of the catalyst. However, in usual practice, elimination of water from reforming
feed is a major concern of operators, because an excess of water leads to a decrease in
catalyst activity.
Water is often present in naphtha feeds, moreover water is frequently injected in
reforming feed hydrotreaters to remove formed salts in the cold part of the reaction
section. The maximum allowable content in feed is set to achieve approximately 20 ppm
volume in the recycle gas for typical conditions (P = 14 barg; T = 40°C at the separator).
When feeding a reformer unit from storage, route the sweet naphtha (ex-storage)
through the HDT stripper (see paragraph 4-3).
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Mechanism:
Water affects the acidic function of the catalyst and decreases the dehydrocyclization of
paraffins.
–
Prevention and causes of contamination:
Water removal is usually achieved in the stripper of the feed hydrotreater. Generally
contamination by water results from a poor operation of this equipment (insufficient
bottom temperature, water not drained in the reflux drum…..)
–
Detection:
Since water contamination is a major concern for the operator, on line analyzer is
provided for the recycle gas.
Operating experience shows that the optimum water content in the recycle gas must be
within a range of 15 to 20 ppm (vol.). The associated chlorine level will then be about
0.5 ppm vol.
Above 50 ppm vol. of water, the reactor inlet temperature must be lowered to reduce
the chlorine elution (leaching) from the catalyst. The following figures are generally
accepted:
• > 50 ppm water
Temperature 
480°C
• > 100 ppm water
Temperature 
460°C
Below 10 ppm of water in the recycle gas the catalyst acidic function is enhanced. Water
injection must be used: 1 ppm wt in the feed results in an increment of 2 to 5 ppm vol. in
the recycle.
–
Remedies:
In case of waterupset:
 Adjust chlorine injection to make up for the increased chlorine loss.
 Decrease reactor inlet temperature as indicated above.
 Restore operating conditions upstream of the unit to reduce water
contamination.
 Check amount of oxygenated compounds in the feed.
 If it is known that large ingress of moisture cannot be prevented, consider
shutdown of the unit until the situation is remedied.
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• Halogens (chlorine, fluorine):
The maximum allowable amount in the unit feed is 0.5 ppm wt for each of them.
–
Mechanism:
The presence of chlorine as chloride in the feed modifies the acidic function of the
catalyst and promotes the hydrocracking reaction. Once chloride is eliminated, the
proper chloride balance of the catalyst can be restored.
The effect of fluorine as fluoride is similar but it is more difficult to remove from the
catalyst. It is very seldom to find fluoride in the reformer feed.
–
Effect of chlorine, fluorine contamination:
Hydrocracking reactions are enhanced:
 Lower reformate product yield.
 Higher LPG and C1 yields.
 Slightly higher octane.
 Decrease of hydrogen production.
–
Prevention and causes of contamination:
Chlorine and fluorine are sometimes present in crude as organic halides owing to the
field production techniques. They are normally eliminated in the pretreatment stage.
Note that if present in pretreatment feed in notable quantity (several wt ppm) they
provoke a huge corrosion in the cold part of the pretreatment reaction section.
B
Permanent poisons
Permanent poisons have been defined as contaminants which irreversibly damage the
catalyst.
–
Mechanism:
Most metals poison the metal function of the catalyst. Metal poisons tend to affect the
first reactor, then to break through and affect the 2 nd reactor. The first reactor is
typically the reactor where the poisoning is first detected.
–
Effect of metal contamination:
Metal contamination is characterized by:
• A dramatic decrease of the T in the first reactor associated with an increased
T in the second reactor, then if poisoning follows through, decrease of second
reactor T with an increased T in the third reactor, etc..
• A decrease in octane number.
• An increase in reformate yield.
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• A decrease of hydrogen production.
In addition, mechanical problems may result from the collection of corrosion products
(scale, rust...) in the first layer of the first reactor…
–
Prevention and causes of contamination:
The contaminants and the source of contamination are listed in Table 3 above.
Prevention consists of adequate hydrotreating and appropriate material selection to limit
corrosion.
Metal poisons are generally partially retained on the upstream hydrotreating catalysts.
However the retention capacity is limited and breakthrough may occur. Such a
breakthrough would result in a very harmful situation for the unit since these poisons
would not be eliminated by the catalyst regeneration.
It is very important to check periodically the metal content of the hydrotreater feed and
product. It enables to monitor the performance of the hydrotreater with regard to
demetallization and also to be warned of a possible metal breakthrough of the
hydrotreatment catalyst, providing the maximum metal retention of the hydrotreatment
catalyst is known. When the metal loading of the hydrotreater catalyst is nearing the
maximum metal retention, the hydrotreatment catalyst needs be replaced.
C
Coke
The coke which deposits on the catalyst is a temporary poison since its detrimental
effect is reversible through regeneration. Owing to its paramount importance in catalytic
reformers, coke formation is treated separately.
–
Mechanism:
Indane derivatives, polynuclear aromatics or naphthenes are the assumed precursors of
coke formation. They result either from slight amounts of polynuclear aromatics in the
feed (depending upon the nature of the crude and the end point of the feed) or from the
aromatics producing reactions of the reforming process itself.
Some diolefin
intermediates of reforming reactions are also potential coke precursors.
Coke deposit affects the catalyst activity by reducing the contact area between catalyst
and the reactants.
–
Prevention and causes of contamination:
Since coke formation is inherently associated with the reforming reactions, there is no
real way to avoid it. One can only minimize it.
Coke will be reduced by a decrease in reactor temperature (i.e. if severity is reduced) and
an increase in hydrogen circulation. Low system pressure, to the contrary, favors coke
formation.
Another parameter to watch to minimize coke is the feed end point, in order to limit
heavy polyaromatics amount. In European countries the maximum allowable feed end
point must not exceed 180°C (ASTM D86) as the marketed gasoline end boiling point is
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limited at 205°C. But in USA were the marketed gasoline and boiling point can reach
215°C the feed end boiling point can go up to 205°C (400°F). When the feed results from
a mixture of different streams, each stream must comply with the 180°C end point.
Note that in case of mixed feeds (use of imported naphtha, SR + cracked naphtha, etc…)
the final boiling point of the mixture does not give sufficient information. Each feed shall
be analyzed separatly as to know the final boiling point of each stream.
Very often mixed feeds are responsible of short catalyst cycle duration, even though the
final boiling point of the mixture stays within acceptable values.
Table 3 hereunder lists the main temporary and permanent poisons as well as their
acceptable level in reformer feed and their most likely source.
Table 4 hereunder lists the maximum acceptable limits on the catalyst.
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TABLE 3
Summary of maximum allowable impurities (wt) in the feed
Component
Type
Source
0.5 ppm max
Temporary
Crude
Organic nitrogen (as nitrogen) 0.5 ppm max
Temporary
Cracked Naphtha
Water or oxygenated
products
5.0 ppm max
Temporary
Contaminants
Chloride as chlorine
0.5 ppm max
Temporary
Crude
Fluoride as fluorine
0.5 ppm max
Temporary
Arsenic
5 ppb max
Permanent
Crude
S.R. or cracked
Naphthas or gas
condensates
Lead
5 ppb max
Permanent
Recycled slops
Mercury
1 ppb max
Permanent
Naphtha condensates
Copper
< detection limit
Permanent
Corrosion
Iron
< detection limit
Permanent
Corrosion
Silicon
< detection limit
Permanent
Additives (antifoaming)
Nickel
< detection limit
Permanent
Corrosion
Chromium
< detection limit
Permanent
Corrosion
Sodium
< detection limit
Permanent
Crude
Calcium
< detection limit
Permanent
Crude
Potassium
< detection limit
Permanent
Crude
Manganese
< detection limit
Permanent
Crude
Magnesium
< detection limit
Permanent
Crude
Sulfur (as sulfur)
Max. allowable (wt)
Note that ASTM D86 distillation End Boiling point must preferably be lower than 180°C.
Also diolefins shall be absent from the feed and olefins content kept at the lowest,
possibly less than 0.1 wt%, as these are coke precursors. This value of 0.1 wt%
corresponds to a bromine index equal to ~ 150
Operating with a feed having a content of impurities higher than the allowance will
lead either to the shortening of the cycle duration or to the decrease of the catalyst
life.
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TABLE 4
RG / PR SERIES CATALYSTS
CUMULATIVE POISONS LIMIT
(Beyond these limits catalyst performances start to deteriorate)
S wt ppm
= ~ 700
Pb wt ppm
= ~ 200
As wt ppm
= ~ 200
Zn wt ppm
= ~ 400
Co wt ppm
= ~ 400
Cr wt ppm
= ~ 400
Mo wt ppm
= ~ 400
Cd wt ppm
= ~ 400
Cu wt ppm
= ~ 400
Fe wt ppm
= ~ 5000
Si wt ppm
= ~ 400
Na wt ppm
= ~ 500
Ca wt ppm
= ~ 100
K wt ppm
= ~ 500
P wt ppm
= ~ 300
Mg wt ppm
= ~ 100
To be noticed that all poisons will be catched by catalyst of first reactor.
Every time that first reactor catalyst sample is available, it is recommended to check it for
contaminants and poisons.
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2.3.5 Catalyst distribution in reactors
Thermodynamics and kinetics have shown that there is an optimum operating
temperature range, approximately 450°C-520°C in order to simultaneously favor the rate
of the desirable reactions and limit the undesirable ones to an acceptable level. For each
specific case, the most appropriate operating temperature is selected taking into account
the feed quality (PNA, distillation range ...) and product requirement (octane).
Owing to the great endothermicity of the most important and desirable reactions
(naphthenes dehydrogenation and paraffins dehydrocyclization) this optimum
temperature cannot be sustained through out the whole catalyst volume. In addition,
dehydrogenation is also, by far, the fastest reaction, which means that the temperature
drops very sharply over the first part of the catalyst. In order to restore the catalyst
activity, when temperature has dropped to a certain level which depends upon the
reactions involved, the reactor feed is reheated. To achieve this, the catalyst is
distributed in several reactors (3 or 4) and intermediate heaters are provided.
Figures 2-4 and 2-5, illustrate this aspect, Figure 2-4 shows for a given feed (Paraffins:
45% LV, Naphthenes 45% LV, Aromatics 10% LV) the profile of the amount of P.N.A along
the catalyst volume. In this case there is no need for more than 10% of the catalyst in
the first reactor because the naphthenes dehydrogenation results in a temperature too
low to sustain the reaction any longer. The reactor effluent is reheated to allow for the
naphthenes dehydrogenation to continue and the paraffin dehydrocyclization to start.
Over the next 15% of catalyst, distributed in the 2nd reactor, temperature drops again to
a level where reheating is required to enable the paraffin dehydrocyclization to proceed.
The catalyst typical distribution in this case is:
• R1 =
10%
• R2 =
15%
• R3 =
25%
• R4 =
50%
In the case of 3 reactors typical distribution is:
• R1 =
15%
• R2 =
25%
• R3 =
60%
Each specific case, obviously, requires a specific catalyst distribution.
In a somewhat simplified but practical way, for operational guidance, the main reactions
take place in the various reactors can be represented in the following order:
• 1st reactor:
- Dehydrogenation
- Isomerization
REV. 4
• 2nd reactor:
– Dehydrogenation
– Isomerization
– Cracking
– Dehydrocyclization
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• 3rd and 4th reactor:
– Cracking
– Dehydrocyclization
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CATALYST DISTRIBUTION IN THE REACTORS
FIGURES 2-4 & 2-5
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Process variables
2.4.1 Independent variables
The process variables are:
• Pressure.
• Temperature.
• Space velocity.
• Hydrogen partial pressure or H2/HC recycle ratio.
• Quality of the feed.
The above are independent variables: each of them can be fixed by the operator within the operating range of the equipment - independently from the others.
For one set of independent variables, for same feed characteristics, there is only one
performance of the unit i.e. one set of values for:
• Product yields.
• Product quality (octane).
• Catalyst stability (coke make).
In this chapter we examine the effect on the unit performance of each independent
variable taken separately.
2.4.2 Pressure
Hydrogen partial pressure is the basic variable because of its inherent effect on reaction
rates. But for the ease of understanding total reactor pressure can be used. Reactor
pressure is most accurately defined as the average catalyst pressure. Due to catalyst
distribution in the reactors, it is usually close to the last reactor inlet pressure.
All the hydrogen producing reactions i.e. dehydrogenation, dehydrocyclization are
enhanced by low pressure.
The lower the pressure the higher the yields of both reformate and hydrogen for a given
octane number. This is the reason for minimizing unit pressure drop and operating at the
lowest practical pressure. Low pressure however increases the coke make.
Operator action on pressure is limited:
• Operating pressure rise is limited by equipment design pressure.
• Operating pressure lowering is limited by recycle compressor design power and
intake volume.
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2.4.3 Temperature
Catalyst activity is directly related to reactor temperature. Thus the most direct
operating variable available for the operator, to control product quality and yields, is the
reactor inlet temperature.
In a conventional semi-regenerative unit when all independent variables are steady, the
loss of activity of the catalyst caused by the coke deposit results in a decrease of the
product octane as well as the reformate yield and recycle gas purity. A slight
modification in reactor inlet temperature is used to compensate and maintain product
octane number, or:
• To process a different feed quantity.
• To process a different feed quality.
• To balance catalyst ageing, this occurs slowly over several years.
An increase of the reactor inlet temperature results in:
• An increased conversion of the non aromatic compounds of the feed mainly the
paraffins. But since the hydrocracking reaction is more favored than the
cyclization of paraffins, the end result is:
–
An increased octane but a decrease in reformate yield.
–
An increase of the coke deposit.
2.4.4 Space velocity
The space velocity is the amount of liquid feed, expressed in weight (or in volume) which
is processed in one hour, divided by the amount of catalyst, expressed in weight (or in
volume). Weight (volume) of feed and catalyst must be expressed with the same unit.
Weight Hourly Space Velocity:
Liquid Hourly Space Velocity:
WHSV =
LHSV =
Weight of feed (per hour)
Weight of catalyst
Volume of feed at 15 C (per hour)
Volume of catalyst
The inverse of the liquid hourly space velocity i.e. (LHSV)-1 is linked with the residence
time of the feed in the reactor. The space velocity then affects directly the kinetics of the
reforming reactions.
A decrease in the space velocity means an increased residence time, hence a higher
severity which results in increased octane, lower reformate yield, higher coke deposit.
When changing feed rate, an important recommendation derives from the above:
• Always decrease reactor inlet temperature first and decrease feed flow rate
afterwards.
• Always increase feed flow rate first and increase temperature afterwards.
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2.4.5 Hydrogen to hydrocarbon ratio and hydrogen partial pressure
The H2/HC ratio is the ratio of pure hydrogen in the recycle gas (mole/hour) to the feed
flow rate (mole/hour), at first reactor inlet.
H2
HC
=
Pure hydrogen (mole / hour) in recycle
Naphtha flow rate (mole / hour)
Hydrogen partial pressure is linked to the H2/HC ratio and total system pressure. Since
there is, in practice, little flexibility in the total pressure, hydrogen partial pressure is
mainly adjusted through recycle flow.
Recycle hydrogen is necessary in the reformer operation for purposes of catalyst
stability. It has the effect of sweeping the reaction products and condensable materials
from the catalyst and supplying the catalyst with readily available hydrogen. An increase
in H2/HC ratio will move the naphtha through the reactors at a faster rate and supply a
greater heat sink for the endothermic heat of reaction. The end result is an increased
stability.
A lower H2/HC ratio decreases the hydrogen artial pressure and increases coke
formation. Within the typical operating range, the H2/HC ratio has little influence on
product quality or yields. It is not a variable that the operator typically adjusts, it is set by
design based on an economic balance between equipment sizing i.e. recycle
compressors, fired heaters and the cycle duration.
Moreover, for a given unit, the amount of recycle is limited by the recycle compressor
characteristics (power, suction flow).
2.4.6 Feed quality
A
Distillation range
Light fractions have a poor naphthenic and aromatic content and consequently a high C 6
paraffinic content. Cyclization of C6 paraffins to aromatics is more difficult than
cyclization of C7 or C8 paraffins, as discussed in paragraph 2.2.2 C.
Hence, for a required octane number, the lighter the feed the higher the required
severity or, conversely, at constant severity, low initial boiling point results in lower
aromatic and hydrogen yields.
The restriction of benzene content in gasoline has resulted in selecting feed with IBP
above 100°C (210°F) to limit benzene precursors.
Heavy fractions have a high naphthenic and aromatic hydrocarbons content, thus they
need a lower severity to obtain good yields. But these fractions contain also polycyclic
compounds which produce a high coke deposit on the catalyst. High final boiling point of
the feed is favourable up to a certain level, however detrimental if above specified limits.
An end boiling point above 180°C (350°F) is generally not recommended (see also
paragraph 2.3.4.C).
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Chemical composition
The detailed chemical composition of the feed is determined by gas chromatography
analysis. This analysis is necessary to predict the aromatics and hydrogen production as
well as the severity of the operation.
Even if not sufficient for a complete prediction, an index of characterization of the
feedstocks related to the actual and potential aromatics content of the feed proves very
useful. N + 2A has long been used (N and A volume % of naphthenes and aromatics in the
feed). AXENS now uses 0.85 N + A which is found to be more representative.
The higher this index, the lower the severity of operation to meet the same product
specifications. The lower this index (i.e. the higher the paraffins content), the higher the
severity of operation to meet the same product specifications as the dehydrocyclization
of paraffins becomes important.
Note that cracked naphthas have a ratio naphthenes C6 nucleus / naphthenes C5 nucleus
much lower than SR naphthas.
Ratio NNC6 / NNC5
Number
of carbon
SR naphtha
Coker Naphtha
Hydrocracking
Naphtha
FCC naphtha
6
1.4 to 1.5
0.4 to 0.5
0.1 – 0.4
~ 0.1
7
1.6 to 1.7
0.2 to 0.3
~ 0.5
~ 0.3
8
~ 2.0
~ 0.6
~ 1.5
~ 0.5
Remember that aromatisation of NNC5 requires first an isomerisation in NNC6. If this
process is rather good for C7+ naphthenes it is only ~60% for the methylcyclopentane
while the cyclohexane is completly converted in benzene.
C
Impurities in the feed
The catalyst activity can be reduced, either temporarily or permanently by poisons
contained in the feed. Refer to paragraph 2.3.4.
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Summary
Table 5 summarizes the theoretical effect on the unit performance of each independent
process variable taken separately.
TABLE 5
Effect of process variables
Increased
RONC
Reformate
yield
H2 yield
Coke
deposit
Pressure
Temperature
Space velocity
H2/HC ratio
A + 0.85 N
Naphtha
End boiling point
Quality
Initial boiling point
Note: OFF-SPEC. Reformate must not be rerun to the reformer as:
a) It does not generate hydrogen and consequently the recycle gas hydrogen
purity decreases (
pp H2,
H2/HC ratio).
b) Aromatics desorption from catalyst takes more time than paraffins and
naphtenes desorption, thus less « room » left for the useful reactions.
Both effects lead to an increased catalyst coking.
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START-UP PROCEDURE FOR FRESH CATALYST
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Preparation of the unit - General considerations – Unit Dry-out
The following recommendations are of general type; hence do not apply to all
circumstances which can appear in an industrial process unit. Suitable modifications may
be necessary to accommodate each specific case.
Reaction section is isolated by blinds from the rest of the unit.
The present document gives some general and simple rules to be followed before the
start-up of a fresh catalyst or of a regenerated catalyst in an existing reformer.
• Should a spent catalyst have to be replaced, a combustion phase must be
carried out before dumping. This operation leads to a complete burning of coke
deposits. It also carries away all residuals, fines, oxides, sulfides, chlorides and
other deposits from all equipment to the reactors and to the separator drum.
Their elimination becomes easier before loading the new catalyst. This
procedure is mandatory and carried out by following only the combustion steps
with burn proof, without chlorine injection but maintaining the soda water
circulation. Oxychlorination and calcination do not need to be performed.
Particular attention should be taken regarding feed/effluent exchanger that
could contain some hydrocarbon and could not be exposed directly to high
concentration of oxygen.
• If a catalyst, fresh or regenerated, is to be loaded in a unit after water washing,
cleaning or pressure tests under water of equipements, the drying of this unit
must be done carefully before loading new catalyst. The operation is carried-out
in a close loop: heat exchanger, furnaces and reactors empty, coolers, separator
drum and recycle compressor. Should the catalyst be already loaded in dry and
clean reactors, those reactors must by-passed by installing suitable lines inbetween the furnaces to realize the close loop.
• In case of partial catalyst replacement, proceed as per normal regeneration
procedure paragraph 6.3.3 .

In case the unit is equipped with one dryer, this dryer needs to be regenerated
and kept under nitrogen to make it ready for the oil in. It will be by passed
during regeneration & reduction.
Typical unit drying procedure:
This step is mandatory prior loading new catalyst if water leak test have been
performed on equipments in the reaction section circuit (feed effluent exchanger).
Otherwise a complete oxychloration should be performed as per paragraph XYZ.
The elimination of water is achieved in circulating dry nitrogen under a 5 to 9 barg
pressure using the recycle compressor. The furnaces are fired and the outlet
temperature is increased up to 500°C. During such operation, water collected from the
reaction section is removed, after condensation in the trim cooler, at separator bottom
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and at all cold low points drains (see Figure 3-1). The quantity of water collected is
recorded precisely.
If presence of hydrocarbon in the circuit is suspected, little air could be added in the
system at 400°C to detect less than 1 vol% of O2 in the recycle gas. Then check that no
burning is taking place in equipments (feed/effluent exchanger, heater….). Absence of
CO2 is a good indicator that unit is well hydrocarbon free.
The unit is considered as “dry” when the water drained becomes lower than an
equivalent of 0.1 wt % of the total catalyst weight per hour. The unit is considered as
hydrocarbon free when CO2 is not produced in presence of 1 vol% of O2 at 400 – 500°C.
The furnace outlet temperature is then reduced at a rate of 50°C per hour till the
complete cooling of the catalytic section. This can be achieved by switching off the
furnaces at 200°C at reactors inlet, while maintaining the recycle compressor in service.
The atmosphere inside the reactors must be made suitable for man entry by switching
from nitrogen to air until authorization of refinery safety department is obtained to enter
the reactors.
3.2
Catalyst loading
This operation is carried out taking precautions highlighted hereafter:
• The catalyst is an expensive product and should be handled carefully, avoiding
any hazardous loss.
• The catalyst being very hygroscopic must be handled in such a way that the
minimum adsorption of water can be assured in spite of a final drying of the
catalyst which is done before introduction of the feed.
Before loading, it is recommended to check that all reactors are dry, clean and that all
internals are in the proper place and installed as recommended in the various drawings
and process data sheets.
The drums containing catalyst must be handled with care in order to minimize particle
attrition. Covers must be installed at the top of each reactor on temporary structures to
protect against rain. Loading of catalyst shall not be done during a period of rain or of
great level of humidity.
See Annexe11
3.3
Catalyst drying
This procedure, which must be applied before the catalyst reduction, ensures a good
elimination of water that has been reabsorbed by the catalyst during handling and
loading. It is carried out in the presence of chlorine and of oxygen. The catalyst drying
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does not replace the phase of oxychlorination which is part of the catalyst regeneration
procedure described in this document. It is only applicable to a brand new catalyst.
Note that catalyst dryring procedure is applicable only if no works followed by hydraulic
test took place in the reaction section (for instance repair or replacement of heater
tubes, etc…). In such a case, before catalyst loading, the reaction section shall be dryedout as per typical unit drying procedure.
After catalyst loading, the reaction circuit is established and placed under a 5 to 9 barg
pressure of nitrogen. After start-up of the recycle compressor, the oxygen concentration
in nitrogen is adjusted and maintained around 3 to 5 volume % during the whole drying
step. If presence of hydrocarbon in the circuit is suspected proceed to a first step with
less than 1 vol% of O2 in the recycle gas and check no burning is taking place in
equipments (feed/effluent exchanger, heater….). Absence of CO2 is a good indicator that
unit is well hydrocarbon free.
This oxygen concentration is followed regularly by laboratory analysis or by means of an
on-line analyzer.
Injection of air could become necessary to keep the oxygen concentration at the required
level.
Furnaces are fired and temperature of all catalytic beds is increased at a rate of 40°C per
hour up to 400°C (the average temperature between reactor inlet and outlet
temperatures when there are no thermocouples inside the reactors).
Between 400°C and 485°C it is preferable to reduce the heating rate at 25°C per hour, to
reduce the chloride leaching from the catalyst, HCl content at the oulet of the last
reactor needs to be below 20 ppm vol.
When the catalyst temperature reaches 350°C, an injection of a suitable pure chlorinated
compound is carried out (for example tetrachloroethylene, trichloroethylene,
trichloroethane etc…), for about four hours, so as to introduce chlorine (as Cl) equivalent
to 0.2 – 0.3 wt % of the total mass of catalyst. This injection is done at the first reactor
inlet while progressing normally with the temperature rise.
During the whole operation of drying, the water which has condensed and accumulated
in the separator drum and at all low and cold parts of the unit is drained every hour. The
quantity of water collected is recorded precisely. Separator temperature is kept as low
as possible.
The catalyst bed temperature is maintained at 485°C for two hours minimum, or until the
total quantity of water drained at low points is lower than 0.05 wt % of the catalyst per
hour. Then temperature is reduced to 200°C in catalyst beds at a rate of 50°C per hour.
Typical amount of water drained is less than 1 wt % of total inventory.
At this temperature level, furnaces and recycle compressor are shutdown
simultaneously.
If the reactors are equipped with sampling devices, catalyst must be sampled at the end
of the drying step. The analyses of chlorine on the catalyst shall be used to eventually
adjust the quantity of chloriding agent to be injected after oil-in.
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When unit is equipped with plate’s feed effluent exchanger like Packinox, it could be
recommended to decrease catalyst bed temperature below 180°C, at the outlet of the
last reactor, to avoid thermal stress of feed effluent exchanger when compressor will be
restarted. Check with exchanger manufacturer. Axens has no objection to decrease
catalyst bed temperature below 180°C. Additionally, to smooth down thermal shocks due
to potential cold gas waves, it must be considered restarting the compressor at the
minimum achievable driver speed and monitor last reactor outlet temperature until it
stabilizes. Then the normal operating recycle gas flow can be reached in complete safety
for internals.
3.4
Catalyst reduction
Before hydrogen introduction check that recycle gas is containing no more than 0.5 vol%
no less than 0.2 vol % of O2 and also less than 10 volppm CO and 500 volppm CO2
concentration. Repeat nitrogen purges as necessary.
Electrolytic hydrogen shall be used for the reduction of the metal oxides. The quantity of
hydrogen necessary to perform the reduction step is 25 to 30 times the volume of the
reaction section. This quantity takes into account the maintaining of the pressure in
reaction section up to the OIL-IN. As an alternative, hydrogen produced by a steam
reforming + PSA unit (99.9 vol. % of H2, H2O < 20 volppm and CO < 5 volppm) can also be
used. Both of them being dry and hydrocarbon free will allow to reduce the metals
oxides in a proper way. In several cases, as to ease the operation of recycle gas
compressor unable to operate with low molecular weight, it is possible to add dry
nitrogen to the pure hydrogen to adjust the molecular weight of the recycle gas. If H2 is
diluted with N2, the hydrogen purity must remain higher than 50 vol%.
If none of the above mentioned hydrogen gases are available it is possible to use
hydrogen rich gas from a reformer unit, mixed with nitrogen to minimize the impurity
level. The H2 purity of recycle gas (mixture of N2 + Hydrogen make-up) must be higher
than 50 vol% and C2+ content shall be limited to 2 vol% maximum. In that case, due to
the presence of hydrocarbons, catalyst coking will take place during the reduction
stage and lead to shorter cycle duration. In addition catalyst activity and stability will
also be affected as well as product yields.
Check for recycle gas composition for H2, N2 and HC. If pure hydrogen was used for
reduction, presence of HC will be a sign of potential catalyst contamination under its
oxide form.
Reactor inlet is increased up to 400°C at a rate of 40°C/h, then at 25°C/h up to 510°C. The
objective is to minimize chloride stripping during this step. Practically HCl content at the
oulet of the last reactor must stay below 20 ppm vol if not reduce heating rate at 15°C/h
The final reduction step starts when the average temperature between reactor inlet and
outlet is 500°C. As a reminder, Pt reduction may start with temperature as low as 200°C
and above, then it is important to keep the recycle gas compressor running any time to
evacuate the water produced by metals reduction. In case of recycle gas compressor
shutdown during this period, pressurization/depressurization cycles shall be performed
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to maintain a dry atmosphere within the reactors, avoiding metals sintering caused by
moisture presence.
The 500°C step is maintained until the quantity of water drained is lower than 0.05% of
the catalyst weight per hour (two consecutive measurements).
During this whole step, water formed by reduction is drained every hour from the
separator drum and at the various low and cold parts. The quantity of water collected is
recorded precisely. Separator temperature is kept as low as possible.
Hydrogen make-up may be necessary to maintain the pressure inside the reaction
section.
Under no circumstances should a neutralizing solution be used to protect cold parts of
the loop as this will lead to water saturated conditions and chloride removal which will
adversely affect metals dispersion.
If for any reason a started reduction has to be stopped for more than 24 hours another
oxidation step shall be performed. Indeed during the shutdown the catalyst will cool
down and adsorb water and if re-started in these conditions the reduction will be done
under too wet atmosphere this will lead to metals agglomeration and consequently is
going lead to shorten cycle length.
Note 1:
3.5
Oxygen removal before catalyst reduction. It is strongly recommended not to
use the ejector installed on top of the reaction section separator to pull
vacuum in the section as:
 In case of leak in the section, oxygen will be introduced.
 In case of lack of blinds, hydrocarbons can be introduced.
 In the particular case of cold wall reactor (internal insulation), contaminants
(sulfur, heavy hydrocarbons) will enter the section. In this last case it is
strictly prohibited to use the ejector.
Catalyst sulfiding
Sulfiding is performed to reduce temporarily the extra activity of the catalyst, especially
the hydrogenolysis function of the freshly reduced metallic phase which could lead to
cracking reactions with production of methane. This procedure (See figure 3-4) is based
on a reversible formation of sulfides.
This operation also allows the sulfiding of iron metal which could lay down in the reactor
during the catalyst regeneration.
The sulfiding step is carried out after reduction, when the catalyst beds temperature has
been reduced from 500°C down to 400°C at a rate of 50°C per hour.
The sulfiding step lasts for at least 3 hours at 400°C. The recommended sulfiding agent
is dimethyl disulfide (DMDS) which is an odorous liquid with a low vapor pressure.
The sulfur injection is carried out reactor by reactor starting with reactor n°2, then
reactor n°3 (then reactor n°4 if any) and finally with reactor n°1. This allows either to let
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more time for the sulfur to disperse in the last reactor and to oil-in the unit immediately
after reactor n°1 sulfiding.
For a fresh catalyst the quantity of DMDS to be introduced corresponds to 0.15 wt %
sulfur on catalyst (i.e. 150 kg of sulfur for 100 tons of catalyst equivalent to about 225 kg
of DMDS). For regenerated catalyst this quantity could be reduced to 0.04 – 0.08 wt %
sulfur on catalyst.
Before injection make sure that the sulfiding agent injection lines have been purged and
filled with sulfiding agent.
Every 5 minutes, the concentration of hydrogen sulfide at reactors outlet must be
determined using the usual method of Dräger tubes or equivalent. Such analyses are
used to make sure that each catalyst bed is well sulfided (~ 2 to 5 ppm vol.) and that the
excess of H2S is transfered from one reactor to the following one. If on two consecutive
measurements H2S breakthrough is noticed the injection is stopped.
If, in the course of the sulfiding step, an injection pipe happens to be plugged and it
becomes impossible to sulfide one given reactor, then the required quantity of DMDS for
this given reactor will be introduced at the inlet of the upstream reactor.
For the first reactor, another injection point should be found to introduce the required
quantity of sulphur.
During this step, remaining water is drained from the separator drum and at the various
low and cold parts.
 Due to presence of iron in first reactor you may need to inject more DMDS
in this reactor.
 For checking with Drager tube, make sure that the samples points at reactor
outlets are short and do not include sections of copper or carbon steel
piping.
Only sealed drums of sulfiding agents, not already open, can be used for sulfiding.
Note: –
3.6
DMDS is the selected sulfiding agent.
–
DMDS “evolution”, less odorous, can also be used.
–
Mercaptans can also be used but their handling is difficult.
–
TBPS could be acceptable.
–
H2S, too dangerous to handle shall be avoided.
–
CS2, too dangerous and responsible for catalyst coking is banned.
–
Polysulfides other than TBPS, responsible for catalyst coking, are banned.
Oil-in
Feed should be introduced to the unit within one hour of completing sulfiding.
The catalyst bed temperature being stabilized at 400°C, the reaction section pressure is
increased to a minimum of 8 barg at the separator drum by injection of start-up
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hydrogen. Swing open the blind on feed. The feed is introduced to the reactors at 60% of
the design rate in about 10 minutes. The furnaces are controlled to maintain 400°C at
the inlet each catalytic bed (around 405°C at furnace outlet). If exothermal phenomenon
is not observed (temperature increase higher than 5°C), the inlet reactor temperature
can be increased by 10°C during the first hour, then pushed to 460°C at a rate of 25°C per
hour.
During this period, a continuous injection of chlorinated agent is carried out. The
quantity to be introduced amounts to 10 wt. ppm calculated as Cl concentration in the
feed rate to the unit. If it is suspected (or checked when catalyst samplers are installed)
that the chlorine content of the catalyst is lower than 0.8 wt % the Cl concentration in
feed can be increased-up to 20 wt ppm. The chlorinated agent can be used as pure
component, or diluted in naphtha feed.
The production of gas brings an automatic increase of pressure which is adjusted at its
operating value. As soon as the pressure builds up start the H 2 rich gas compressor (if
any). The excess of gas is fed to the feed pretreater.
The temperature increase profile in function of the water content of the recycle gas is
described in the following paragraph 3.7.
If the feed pretreater cannot be in operation before the reformer start-up, it is necessary
to use dry and low sulfur content naphtha for this phase of start-up. Route the feed from
storage to the HDT stripper as to eliminate the water before entering the reformer unit.
As feed could pass through unusual circuit during this step, check the feed quality and
color before feed-in.
3.7
Startup phase
After feed injection, a phase of water removal develops for a period of 3 to 5 days during
which special precautions must be taken for temperature increase profile in function of
water and hydrogen sulfide concentrations in the recycle gas (See paragraph 3.7.5 and
figure 3-4).
For an efficient drying, the operator has advantage to bring the feed rate to its maximum
value in order to reduce the dehydration time, the risk of water pocket in the piping and,
in consequence, to minimize the period necessary to reach the desired octane level (for
information, it is recommended to decrease as fast as possible the water content in the
reformer feed below 5 wt. ppm). During this step, purge all drain low point where water
could be accumulated.
Please also note that a high feed rate brings a good repartition of flow through the
catalytic beds and an improved distribution of chlorine on the catalyst. Hence, the
exothermal reactions are minimized.
If the unit is equipped with a recycle gas dryer it shall be put on stream when the reactor
inlet temperatures are 450°C. Note that the dryer shall be regenerated before the
previous shutdown of the unit. When gas moisture at dryer outlet will equals inlet one or
when dryer inlet moisture falls under 30 vol. ppm whichever comes first, the dryer shall
be by-passed. Further regeneration of the dryer shall only take place when the moisture
REV. 4
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of the recycle gas is lower than 500 vol. ppm. When dryer is in service, reduce chloride
injection at 1 wtppm.
3.7.1 First temperature level: 460°C
This temperature level is maintained as long as the water content in the recycle gas is
over 200 vol. ppm and the hydrogen sulfide content higher than 10 vol. ppm.
The injection of chlorinated agent is maintained at a rate of 10 wt ppm chloride to feed
during this phase. If the unit is equipped with catalyst samplers, thus the chloride level of
the catalyst known, it is possible in case of low chloride content to increase the
chlorinated agent injection as to have up to 20 wt ppm chloride to feed.
3.7.2 Second temperature level: 470°C
The reactor inlet temperature is increased to 470°C and maintained at this level as long
as the water content in the recycle gas is above 100 vol. ppm and the hydrogen sulfide
concentration is higher than 5 vol. ppm. The injection of chlorinated agent is decreased
to 5 wt. ppm during this phase.
However, the rate of chloride injection and the place of introduction might be modified
in function of the chloride concentrations determined by analysis on the catalyst samples
which are collected from the reactor bottoms using the sampling systems (if this device is
installed).
3.7.3 Third temperature level: 480°C
The temperature is finally increased at inlet of each reactor to 480°C and maintained at
this level while keeping the water content in the recycle gas between 100 and 50 vol.
ppm and the hydrogen sulfide concentration between 5 and 2 vol. ppm. The chloride
injection is reduced to 3 wt. ppm in the naphtha feed.
3.7.4 High severity operation:  480°C
The reactor inlet temperature is then progressively increased to the required value to
reach the reformate octane number as soon as the recycle gas water content is below 50
vol. ppm and the hydrogen sulfide concentration is lower than 2 vol. ppm. It is
recommended to wait 30 vol. ppm moisture in the recycle gas before increasing the
reactor inlet temperatures above 500°C.
The chloride injection rate as a function of the recycle gas water content is defined in
more detail under the "normal operation" chapter. When catalyst sampling devices are
installed the catalyst of first and last reactor is analyzed for chloride content.
REV. 4
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3.7.5 WAIT increase summary
RIT°C
460 max
470 max
480 max
> 480
> 500
H2O in R/G ppm
> 200
< 200
< 100
< 50
< 30
H2S in R/G ppm
> 10
< 10
<5
<2
<1
Chloride injection rate wt ppm
10 – 5
5
3
2
0.2 - 1
Note:
 During this period, an increase of RiT’s implies a temporary rise of moisture
in the recycle gas, as at the new WAIT conditions more water is removed
from the catalyst. Consequently maintain anyway the new WAIT even
though the moisture in the recycle gas exceeds the allowance.
 If after an increase of RiT’s, you observe a drop of recycle gas purity, we
recommend to increase the feed rate in order to have higher space velocity
and reduce cracking.
REV. 4
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REACTION SECTION DRY OUT DIAGRAM
FIGURE 3-1
500°C
20°C / h
Water drain
< 10 L/h
400°C
50°C / h
20°C / h
200°C
40°C / h
Water drain
< 10 L/h
Water drain
< 10 L/h
Stop heaters
at # 200°C
Stop RG Compressor
Ready for
Catalyst loading
Duration for reaction section dry out about 48 hours
REV. 4
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CATALYST DRY OUT DIAGRAM
FIGURE 3-2
REV. 4
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CATALYST HEATING REDUCTION AND SULFIDING
FIGURE 3-3
REDUCTION:
Removal of O2/CO/CO2
Electrolytic hydrogen recommended
C2+ max 2%vol
Oil-in
500°C
50°C / h
40°C / h
Mini 4 hours
400°C
REDUCTION – 500°C
Duration 4 hours mini
And H2O less than
0.05% wt of catalyst per
hour
SULFIDING – 400°C
Duration 3 hours
0.15 wt%sulfur
Breakthrough 20 ppm
Reactors order 2-3-1
Reaction section under H2
RG Compressor in service
Start heaters
TOTAL DURATION ~20 hours
REV. 4
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TYPICAL REFORMING START-UP
FIGURE 3-4
Increase feed by 10% stroke
Increase temp. by 2°C stroke
480°C
470°C
460°C
Oil-in
10°C / h
10°C / h
H2O < 50 ppm
Cl in feed = 2 ppm
H2S < 2 ppm
25°C / h
H2O < 100 ppm
Cl in feed = 3 ppm
H2S < 5 ppm
60% of design feed
in 10 min
Cl in feed = 10 ppm wt
400°C
H2O < 200 ppm
Cl in feed = 5 ppm
H2S < 10 ppm
10°C / h
ESTIMATED DURATION ~ 5 / 6 days
REV. 4
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4
REV. 4
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NORMAL OPERATION
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4.1
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Change of feed rate
The modification of the feed rate, while maintaining a constant reformate octane level,
requires an adjustment of the reactor inlet temperature.
According to experience, it appears more simple to monitor the temperature at each
reactor inlet at the same value ("horizontal" profile). "Ascending" or "Descending"
profiles can be used (to handle very high naphtenic feed, heater skin temperature
problem, etc...).
4.1.1 Increase of feed rate
The increase of feed rate must be followed by an increase of reactor inlet temperature
using the indications given in the figure attached in appendices to chapter 4.5.
The simplified following rules can apply:
• increase of feed rate by 10%
 Increase of temperature by 2°C.
• increase of feed rate by 20%
 Increase of temperature by 3.5°C.
• increase of feed rate by 50%
 Increase of temperature by 8°C.
4.1.2 Decrease of feed rate
The reactor inlet temperature must be reduced at first. This operation is then followed
by a decrease of the feed rate according to information shown on the figure given in
appendices to chapter 4.5.
The simplified following guidelines can apply:
• decrease of feed rate by 10%  Decrease of temperature by 2°C.
• decrease of feed rate by 20%  Decrease of temperature by 4°C.
• decrease of feed rate by 50%  Decrease of temperature by 13°C.
The changes in feed rate and temperatures are accomplished in such a way that cracking
reactions and coke deposit on the catalyst are avoided.
Following such adjustments, it is necessary to check the reformate octane number after
stabilization of the operating conditions and to correct the eventual injections of water
and chlorine. The recycle ratio must be recalculated and possibly adjusted by varying on
the recycle compressor speed.
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4.2
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Change of reformate octane number
With a constant feed rate, the modifications of the reformate octane number are
obtained by changing the reactor inlet temperature.
The temperature increase per one octane number point depends on several factors:
operating pressure, recycle ratio, feedstock quality in terms of naphthenes and aromatics
content and mostly the state of catalyst activity.
Nevertheless, the operator can follow the indications given in Appendices attached to
chapter 4.5 « Operating Parameters », or follow the « rules-of-thumb » listed here
below:
• RON (clear):
85 to 90
 Temperature increase of ~ 1.8 °C per octane point.
• RON (clear):
90 to 95
 Temperature increase of ~ 2.4 °C per octane point.
• RON (clear):
95 to 100
 Temperature increase of ~ 2.8 °C per octane point.
• RON (clear):
100 to 102  Temperature increase of ~ 3.5 °C per octane point.
4.3
Recycle gas water content. Water and chloride injection
The adjustment of water content in the recycle gas at the separator is a determining
factor for a good operation of the reforming unit.
The recommended water concentration in the recycle gas must be between 15 and
25 volume ppm. This corresponds to water content in feed of about 3 - 5 wt ppm.
Normally a well operated HDT stripper provides a feed containing 1 to 3 wt ppm of
water.
An increase of concentration can indicate:
• A misoperation of the stripping column at the pretreater.
• The feeding of the reformer from a storage having not been water drained.
• The utilization of lines or equipments which have been either water washed or
pressure tested with water.
In such cases, an additional injection of chlorine must be performed
(tetrachloroethylene, trichloroethylene or trichloroethane are preferred) at a rate
indicated in the following table. If necessary, the reactor temperature shall be decreased.
Conversely if the unit is fed with a feedstock properly treated, an injection of water can
be carried out at a rate of 3 to 5 wt. ppm calculated on the feed rate expressed in weight.
Anyway never let the moisture in the recycle gas fall under 10 vol. ppm.
The water concentration is monitored by an on-line analyzer placed on the recycle gas.
The most common are of the "Beckmann" type with a phosphoric acid cell Ametek or
Dupont type. These analyzers, which are normally resistant when the water
concentration is below 700 ppm, show partial degradation in function with time and
their indications correspond to values getting lower and lower. Note that online
analyzers can be cross-checked with a portable analyzer (Meeco, Mitsubishi or Shaw).
REV. 4
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Before adjusting the water injection rate, it is recommended to check the analyzer cell as
per the methods given by the manufacturer.
Chloride level of the catalyst is maintained by injecting a chlorinated agent to the feed.
As to have an accurate control of the injection the chlorinated agent is diluted with
reformate.
Table indicating the approximate chloride content
of the feed in function of the water content
in the recycle gas
Recycle gas moisture
content,
volume ppm
Chloride content, (Cl),
weight ppm in feed
15 - 25
0.2 to 1.0
25 - 35
1.0 to 1.5
35 - 50
1.5 to 2.0
 50
3.0 (Decrease reactor inlet temperature to 480°C).
 100
5.0 (Decrease reactor inlet temperature to 460°C).
When the recycle gas contains more than 35 ppm water, such situation must be
temporary and no short term effects occur. In the range of 50 ppm, an increasing
production of LPG can be noticed against a decrease of the reformate yield, as the water
increase enhances temporarily the last reactor catalyst acidity.
The chloride concentration on the catalyst must stay in the following range:
• 0.9 wt %
to
1.1 wt %
If the unit is equipped with sampling device, it is recommended to collect a sample from
each reactor twice a month and eventually at a higher frequency during the temporary
changes of operation.
The normal analyses to be carried out on such samples are: chloride, coke, sulfur and
occasionally, contaminants, surfaces area, platinum dispersion.
As already mentioned in paragraph 2.3.3 “Catalyst mechanism”, for a moisture in recycle
gas in the normal range of 15 to 25 vol. ppm, the associated HCl should be approximately
traces (i.e 0.2 to 0.5 vol. ppm). HCl content in recycle gas is measured by dedicated
Dräeger tubes, scale 0 to 10 vol. ppm. Sampling point shall be drained until the
temperature is equal to the one of the recycle gas at compressor discharge.
Typically chloride injection in normal operating conditions shall be between 0.2 and 1.0
wt ppm related to feed. In the case of presence of nitrogen compounds in feed, chloride
injection shall be increased ( 0.25 wt. ppm of chloride for 0.1 wt ppm nitrogen above
0.2 wt. ppm in feed).
As a test for the proper chloride level, Axens recommends that the refiner increase
reactor inlet temperatures by 5 °C. If the product octane increases by roughly 1.5 - 2
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RONC, the catalyst chloride level is satisfactory. If the octane increase is greater than or
less than 1.5 – 2 RONC, the catalyst is possibly either over or under chlorided
respectively.
4.4
Sulfur content – hydrogen sulfide concentration
A properly hydrotreated feedstock must contain less than 0.5 wt ppm sulfur
If the accuracy of the analytical methods is not sufficient to monitor the content of sulfur
in feed, it is practical to measure the hydrogen sulfide in the recycle gas and in the
stabilizer off-gas.
H2S content in the recycle gas should not amount to more that 1.0 vol. ppm and should
normally amount less than 0.5 vol. ppm.
As already said in paragraph 2.3.4 for the low values of H2S in the recycle gas a more
accurate approach can be made by checking the H2S content of the stabilizer off-gas. For
instance 5 – 8 vol. ppm in stabilizer off-gas corresponds to about 1 vol. ppm in the recycle
gas.
4.5
Operating parameters
Paragraph 2.4 listed the process variables, i.e. the variables (pressure, temperature,
space velocity, H2/HC ratio) which according to the thermodynamics and the kinetics
have an impact on the reactions involved in the process. This rather theoretical
approach did not outline whether the operator could actually change the particular
variable. In the present chapter we will look at these variables again from a more
practical standpoint as operating parameters and how the operators can actually use
them to adjust the performance of the unit.
4.5.1 Pressure
Theoretically, the lower the pressure, the higher the reformate yield and hydrogen purity
- for a given space velocity - and feed characteristics.
The pressure at the last reactor inlet is generally considered as the most representative
for the purpose of this discussion.
However there is little flexibility since the unit and the recycle compressor are designed
for a given pressure. Lowering the operating pressure below the design pressure results
in higher pressure drop and is limited by the recycle compressor design driver power.
The low pressure which favors high yields, favors also coke build up.
4.5.2 Temperature
Temperature, together with space velocity (see hereafter) is the most important and
most used operating parameter.
By simply raising or lowering reactor inlet temperatures operators can raise or lower the
octane number of the product.
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Since all reactor inlet temperatures are not necessarily identical it is commonly accepted
to consider the weight average inlet temperature (WAIT) as representative of the reactor
temperatures.
Anyway WAIT flat profile is recommended for fixed bed reformer operation.
The WAIT is defined as follows:
WAIT =
wt of catalyst R1 x Ti1  + wt of catalyst R2 x Ti2  + wt of catalyst R3 x Ti3 
Ti1, Ti2, Ti3...
Total wt of catalyst
are inlet temperatures to reactors R1, R2, R3...
(wt of catalyst R.)
are the weight of catalyst in reactors R1, R2, R3,...
An increase of temperature (i.e. WAIT) has the following effects', assuming the space
velocity (i.e. the feed rate) and feed characteristics stay unchanged:
• Increases octane.
• Decreases the yield (of C5+ fraction).
• Decreases the H2 purity
• Increases the coke deposit.
At constant WAIT, the coke deposit and the ageing of the catalyst (caused by the
regenerations, the possible metal deposit and the unavoidable upsets) results in a slight
but steady loss of activity (i.e. of octane). An increase of temperature (WAIT) through the
cycle makes up for this activity loss.
Larger and temporary changes in temperature are required:
• To change octane- at constant feed quality and quantity.
• To change feed quantity and still maintain octane (see space velocity
hereunder).
• To change feed quality and still maintain octane (see feed quality).
The set of curves given in the appendix allow an estimate of the increment in the WAIT
which would result from increment in the selected parameter:
• Space velocity (WHSV)
figure 4-1
• Pressure
figure 4-2
• RON clear
figure 4-3
• Feed quality (A + 0.85N and MABP)
figures 4-4 and 4-5
Example: One wants to increase RON from 96 to 100. Refer figure 4-3 in appendix.
REV. 4
RON 96
 WAIT
=
-4.5
RON 100
 WAIT
=
5.5
Result
 WAIT
=
5.5 - (-4.5) = 10°C
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CONCLUSION: WAIT must be raised by 10°C.
WABT, weight average bed temperature, used for figures 4-2 and 4-3 is defined as
follows:

Ti1 + Ti1 '  +  catalyst wt R x Ti2 + Ti2 '  +  catalyst wt R x Ti3 + Ti3 ' 
 catalyst wt R1 x
 
 

2
3
2
2
2

 
 

WABT =
Total weight of catalyst
Ti1, Ti2, Ti3...
are inlet temperatures to reactors R1, R2, R3...
Ti1’, Ti2’, Ti3’...
are outlet temperatures for reactors R1, R2, R3...
Catalyst wt R1, R2, R3,... are the weight of catalyst for reactors R1, R2, R3,...
4.5.3 Space velocity
Space velocity has already been defined. It is the amount of liquid (expressed in weight
or in volume) which is processed per hour divided by the amount of catalyst (in weight or
in volume). The inverse of the space velocity is linked to the residence time (or contact
time) in the reactors.
Knowing the liquid flow rate (or the space velocity), plus the recycle flow and the
reactors operating conditions enables to calculate the actual flow in the reactor, hence
the contact time.
The lower the space velocity (i.e. the higher the contact time), the higher the severity,
assuming all other conditions unchanged.
Lowering the space velocity has, then, the same effects as increasing the temperature
i.e., it:
• Increases the octane.
• Decreases the product yields.
• Decreases H2 purity.
• Increases coke deposit.
If temperature increase is limited (by heater design duty or anything else) lowering space
velocity (i.e. decreasing flow rate) can give an additional boost to octane.
Operators must bear in mind that each time liquid feed rate is changed a temperature
correction must be applied if octane is to be maintained. When feed is increased,
temperature must be raised and conversely, when feed is reduced temperature must be
lowered.
When changing feed rate, an important rule is:
• For feed increase:
meet
Increase feed first, then adjust temperature increase to
octane
• For feed reduction: Lower temperature first, then adjust feed reduction to
meet
octane
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Example of temperature adjustment with space velocity (see figure 4-1 in appendix).
Original conditions
Increased feed
bbl/h
800
1000
m3/h
127.19
158.99
(API) lb/bbl
(55) 265.3
no change
Kg/m3
756.9
no change
lb/h
212,240.
265,300.
t/h
96.270
120.340
wt (lb)
117,910.
no change
(t)
53.480
no change
WHSV
1.8
2.25
WAIT reading (°C)
-2
+ 2.5
WAIT change (°C)
Base
+ 4.5
Feed vol.
Sp gr
Feed wt
Catalyst
CONCLUSION:
For an increase of WHSV from 1.8 to 2.25 the WAIT must be increased
by 4.5°C.
4.5.4 Hydrogen to hydrocarbon ratio
The H2/HC ratio is the ratio of pure H2 in the recycle (mole/h) to the feed flow rate
(mole/h).
The H2/HC ratio is calculated as follows:
R
H2
HC
Where R
REV. 4
=
xY
M
F
m
is the recycle flow in kg/h
M
is the recycle gas molecular weight
F
is the feed rate in kg/h
m
is the feed molecular weight
Y
vol. fraction of H2 in the recycle gas
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R
H2
HC
=
22.4
66/ 172
xY
F
m
is the recycle flow in Nm3/h
Where R
22.4
is the recycle gas molecular volume
F
is the feed rate in kg/h
m
is the feed molecular weight
Y
vol. fraction of H2 in the recycle gas
The recycle gas MW (M) is obtained by chromatographic analysis, as well as the H2 vol.
fraction (Y).
The feed MW (m) is obtained by chromatographic analysis or by correlation from its
distillation range and specific gravity.
Operators can change the H2/HC ratio by lowering or increasing the recycle compressor
flow.
The H2/HC ratio has no obvious impact on the product quality (octane) or yield. But a
high H2/HC ratio reduces the coke build up.
It is strictly recommended to operate with a H2/HC ratio equal to (or higher than) the
design figure.
4.5.5 Feed quality
The feed has always been hydrotreated with the object of removing the sulfur, the
nitrogen, the metals and in case of cracked naphthas the olefins and diolefins. It is
assumed the pretreated feed meets the specifications given in 2.4.6.
The feed quality, once hydrotreated, is mainly expressed by its chemical analysis and its
distillation range.
A
Chemical analysis
The chemical analysis of the feed will give for each carbon number (C6, C7 ... to C10) the
breakdown into Paraffins, Naphthenes, Aromatics (or PNA).
As the fastest reaction is the dehydrogenation of naphthenes into aromatics, the quality
index of a specific feed can be characterized by its initial concentration in naphthenes
and aromatics.
An index (A + 0.85N) where N and A are the volume percent of naphthenes and
aromatics in the feed, is used by AXENS to characterize the process feed.
A feed with a high (A + 0.85N) will be easier to process than one with a lower (A + 0.85N)
i.e.:
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• The same octane will be obtained at a lower severity (temperature) and the
product yield will be higher.
• Or for the same severity (temperature), the octane will be higher.
Typically, with a feed having a higher reforming index, the naphthenic hydrocarbon
content is generally higher, the endothermic reaction heat is increased and the feed flow
rate will be limited by the heater design duty. In the opposite case, with a paraffinic feed
the hydrogen purity of the recycle gas decreases and operation will be limited by the
recycle compressor capacity.
Figure 4-7 in appendix gives an estimate of the relationship between PNA and severity.
As with the other figures it is used to determine increments related to an existing base
case (see example), versus a new case.
B
Distillation range
One feed property that the Refiners can control is the distillation range. The naphtha
end point is controlled in the upstream unit distillation (crude unit fractionator for
straight run naphtha). The initial boiling point is controlled at the naphtha hydrotreater
splitter.
The feed distillation range in European countries where the end boiling point of the
marketed gasoline shall be lower than 205°C is generally as follows:
(°C)
(°F)
IBP
(Initial Boiling Point)
70-100
158-212
EBP
(End Boiling Point)
140-180
284-356
Lower IBP enables to include in the feed, components such as methylcyclopentane (BP
72°C) and cyclohexane (BP 80.7°C) which are benzene precursors.
In some cases however the benzene content in the gasoline is to be strictly limited. In
these cases the IBP must be raised to 82°C (180°F) or above.
A feed with a low IBP generally contains more C6 and lighter paraffins and requires a
higher severity, to obtain the same octane number than a feed with a higher IBP.
One must also consider the distillation end boiling point (EBP). High EBP means heavier
fractions which are richer in aromatics and naphthenes and thus are easier to process.
However the high boiling point fractions are potential sources of polynuclear aromatics
which are coke precursors. In addition high EBP of the feed leads to obtain a reformate
EBP higher than 200°C (problem with gasoline pool). EBP higher than 180°C (356°F) are
generally not recommended in European countries.
Figure 4-8 in Appendix gives an estimate of the change in severity associated with a
change in the distillation range. The distillation range is expressed by the mean average
boiling point (MABP) defined as follows:
REV. 4
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T 10% + 2 x T 50% + T 90%
4
Example of temperature adjustment for A + 0.85 N and RON changes
(See figures 4-4 and 4-3 in appendix)
Base case
Alternate
(A + 0.85 N) vol%
42
35
RON
97
96
 WAIT Figure 4-4
-4.8
+ 0.7
 WAIT Figure 4-3
- 2.3
-4.5
 WAIT from Figure 4-4 (A + 0.85 N)
Base
+ 5.5
 WAIT from Figure 4-3 (RON)
Base
- 2.2
CONCLUSION: WAIT must be increased by 3.3°C.
4.5.6 Butane content of the reformate
In the case a gas chromatography is not available the butane content of the reformate
can be estimated using figure 9-2.
4.5.7 Start of run WAIT calculation
S.O.R. WAIT is not really an Operating parameter but an indication of the catalyst activity.
Figure 4-6 gives the expected SOR WAIT versus RON C for different qualities of feedstock.
(0.85 N + A) using a fresh catalyst in the following operating conditions:
WHSV = 2 - reactor average pressure = 13.5 barg.
It is supposed that chloride level of the catalyst is between 0.9 to 1.1 wt %, that recycle
gas moisture, chloride and H2S contents are according to the recommendations.
After correction for pressure, WHSV and MABP the SOR WAIT can be compared to the
present WAIT. The  between both WAIT’s gives an idea of the catalyst activity.
Plotting on a graph the  WAIT allows to follow the loss of activity of the catalyst.
REV. 4
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4.5.8 Cycle length
The cycle length is defined:
• Either as the number of months of operation.
• or the number of cubic meters of feed treated per kilogram of catalyst or barrels
per pound.
It depends on:
• The operation severity:
–
WHSV,
–
RONC,
–
H2/HC,
–
Operating pressure.
• The nature of the feedstock:
–
Feed composition (0.85N+A vol. %),
–
End Boiling Point °C.
The influence of these parameters on the cycle length is given in the following figures:
Figure 4-7
Figure 4-8
Figure 4-9
Figure 4-10
Figure 4-11
Figure 4-12
REV. 4
:
:
:
:
:
:
WHSV
RONC
H2/HC
Operating pressure
Feed composition (A + 0.85Nvol. %)
End Boiling Point °C
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Application
The cycle length is equal to 10 months in the base case conditions. What will be the cycle
length in the new case?
Base case
New case
WHSV
1.5
2
RONC
94
96
H2/HC
6
5
Operating pressure
15
18
Feed composition (A + 0.85N vol. %)
30
40
End Boiling Point °C
160
170
Base case
New case
Figure 4-7
WHSV
Figure 4-8
RONC
Figure 4-9
H2/HC
Figure 4-10
PRESS.
Figure 4-11
A + 0.85N
Figure 4-12
EBP
1.78
1.3
1.62
1.34
1.28
1.0
1.0
1.22
0.64
1.0
1.1
1.0
The cycle length for new case will be equal to:
 1.30   1.34   1.0   1.22   1.0   1.0 
10 x 
 x
 x
 x
 x
 x   = 8.2 months
 1.78   1.62   1.28   1.0   0.64   1.1
4.5.9 Delta C5+ yield (wt %) for different feedstocks/versus RONC
Figure 4-13, allows you to forecast, for different feedstocks, the change in C 5+ yield which
will result from a variation of the reformate octane number.
The curves were plotted for an average reactor pressure of 13.5 barg. Other operating
pressures can be approximated by using similar trends as plotted in figure 4-13.
REV. 4
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APPENDICES ATTACHED TO
PARAGRAPH 4.5, OPERATING PARAMETERS
REV. 4
FIGURE 4-1
Wait correction for WHSV
FIGURE 4-2
Wait correction for pressure
FIGURE 4-3
Wait correction for RON Clear
FIGURE 4-4
Wait correction for (A + 0.85N)
FIGURE 4-5
Wait correction for MABP
FIGURE 4-6
SOR WAIT requirement versus feed composition
(A + 0.85N Vol. %)
FIGURE 4-7
Relative coking rate versus WHSV
FIGURE 4-8
Relative coking rate versus RON Clear
FIGURE 4-9
Relative coking rate versus recycle molal ratio H2/HC
FIGURE 4-10
Relative coking rate versus reactors average pressure
FIGURE 4-11
Relative coking rate versus feed composition
(0.85N+A vol. %)
FIGURE 4-12
Relative coking rate versus feed end boiling point °C
FIGURE 4-13
Delta C5+ yield (wt %) for different feedstocks /
versus RON Clear
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WAIT CORRECTION FOR WHSV
(TYPICAL HYDROTREATED STRAIGHT RUN NAPHTHA)
FIGURE 4-1
8
6
4
2
Delta Temperature (°C)
0
-2
-4
-6
-8
-10
-12
-14
-16
-18
0.5
1
1.5
2
WHSV in hour
REV. 4
2.5
3
-1
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WAIT CORRECTION FOR PRESSURE
(TYPICAL HYDROTREATED STRAIGHT RUN NAPHTHA)
FIGURE 4-2
8
6
Delta Temperature (°C)
4
2
0
-2
-4
-6
-8
10
12
14
16
18
20
22
24
26
28
30
32
34
36
38
Pressure in Kg/cm2
REV. 4
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WAIT CORRECTION FOR RON CLEAR
(TYPICAL HYDROTREATED STRAIGHT RUN NAPHTHA)
FIGURE 4-3
14
12
10
8
6
Delta Temperature (°C)
4
2
0
-2
-4
-6
-8
-10
-12
-14
-16
-18
88
90
92
94
96
98
100
102
104
Research Octane Number F1 clear
REV. 4
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WAIT CORRECTION FOR A + 0.85N
(TYPICAL HYDROTREATED STRAIGHT RUN NAPHTHA)
FIGURE 4-4
10
8
6
Delta Temperature (°C)
4
2
0
-2
-4
-6
-8
-10
-12
20
25
30
35
40
45
50
55
A + 0.85*N in vol%
REV. 4
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WAIT CORRECTION FOR ASTM 50%
(TYPICAL HYDROTREATED STRAIGHT RUN NAPHTHA)
FIGURE 4-5
12
10
8
6
Delta Temperature (°C)
4
2
0
-2
-4
-6
-8
-10
-12
-14
95
100
105
110
115
120
125
130
135
140
145
150
155
ASTM Distillation 50% °C
REV. 4
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SOR WAIT REQUIREMENT VERSUS FEED COMPOSITION
(A + 0.85N VOL. %)
FIGURE 4-6
525
520
25
515
35
510
45
WAIT °C
505
55
500
A+0.85N (vol%)
495
490
485
480
475
92
94
96
98
100
102
RON
For
REV. 4
P
WHSV
MABP
=
=
=
13.5 bar g
2
122°C
H2 / HC = 5 to 6
C6P - + C5N < 5 wt. %
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RELATIVE CYCLE LENGTH VERSUS WHSV
FIGURE 4-7
3.0
RELATIVE CYCLE LENGTH
2.5
2.0
1.5
1.0
0.5
0.8
1.2
1.6
2
2.4
2.8
3.2
WHSV in hour-1
REV. 4
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RELATIVE CYCLE LENGTH VERSUS RON CLEAR
FIGURE 4-8
2.2
2.0
1.8
RELATIVE CYCLE LENGTH
1.6
1.4
1.2
1.0
0.8
0.6
0.4
0.2
90
92
94
96
98
100
102
RON Clear
REV. 4
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RELATIVE CYCLE LENGTH VERSUS H2 / HC MOLAR RATIO
FIGURE 4-9
1.8
1.6
RELATIVE CYCLE LENGTH
1.4
1.2
1.0
0.8
0.6
0.4
2
3
4
5
6
7
8
H2/HC molar ratio
REV. 4
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RELATIVE CYCLE LENGTH
VERSUS REACTOR AVERAGE PRESSURE
FIGURE 4-10
2.0
1.8
RELATIVE CYCLE LENGTH
1.6
1.4
1.2
1.0
0.8
0.6
0.4
8
12
16
20
24
28
32
2
REACTOR AVERAGE PRESSURE Kg/cm g
REV. 4
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RELATIVE CYCLE LENGTH
VERSUS FEED COMPOSITION (A + 0.85N VOL. %)
FIGURE 4-11
1.5
1.4
1.3
RELATIVE CYCLE LENGTH
1.2
1.1
1.0
0.9
0.8
0.7
0.6
0.5
0.4
24
28
32
36
40
44
48
52
56
FEED COMPOSITION (A + 0.85N (vol%))
REV. 4
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RELATIVE CYCLE LENGTH
VERSUS FEED END BOILING POINT
FIGURE 4-12
1.3
1.2
RELATIVE CYCLE LENGTH
1.1
1.0
0.9
0.8
0.7
0.6
150
155
160
165
170
175
180
185
FEED END BOILING POINT °C
REV. 4
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DELTA C5+ YIELD (WT%) FOR DIFFERENT
FEEDSTOCKS / VERSUS RON CLEAR
FIGURE 4-13
A+0.85N (vol%)
0
-2
-4
47
-8
+
Delta C5 Yield (wt%)
-6
40
-10
-12
31
-14
25
-16
92
94
96
98
100
102
RON
REV. 4
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Troubleshooting
4.6.1 General
The main causes of concern related to process for the Refiners can be one or more of the
following:
• Unexpected decrease of octane number.
• Loss of reformate yield.
• Unexpected T reduction in the first catalytic bed.
• High hydrocracking rate.
4.6.2 Unexpected decrease of the octane number
We are not considering here the cases where the octane number decrease is easily
related to:
• A change in feedstock quality (more paraffins).
• A lowering of the reactor inlet temperature.
• An increased feed flow rate.
An unexpected decrease of the octane number may be caused by one or several of the
following:
• Leak in the feed/effluent exchangers.
• Presence of nitrogen in the feed.
• Presence of sulfur in the feed.
• Presence of metals in the feed.
• Low chloride content of the catalyst.
• Partial by-pass of the catalyst.
A
Leak in the feed / effluent heat exchangers
Since the feed is at higher pressure than the effluent, it may leak through the exchangers.
Relatively small leaks can lower the product octane number significantly.
How to detect feed / effluent exchangers leak
Presence of sulfur in the reformate allows you to suspect a leak at feed/effluent
exchangers (the small amount of sulfur present in feed 0.1 to 0.5 wt ppm is transformed
in H2S in the reactors and is mainly found in the recycle gas).
To ascertain a leak on these exchangers the following method can be followed.
REV. 4
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A-1 TEST WITH PHENOL
Small quantity of phenol is injected in the reformer feed. Phenol is dissociated in
the reactors (hydrocarbon and H2O) and shall not be found in the reformate if the
exchangers are not leaking. Phenol in feed and separator liquid is dosed by UV
detection (method IFP 9860).
Procedure is as follows:
1. Decrease the liquid level in separator down to the lower acceptable level.
Analyse feed before injection of phenol.
2. Inject 100 ppm phenol in reformer feed. Analyse feed 10 minutes after phenol
injection. Phenol concentration should be  100 wt ppm.
3. After 10 minutes, 20 minutes, 30 minutes, 40 minutes, 50 minutes, 60 minutes,
take a sample at separator bottom.
4. Stop phenol injection when the last sample at separator bottom has been
taken.
In case of exchangers leak, phenol will be detected in separator product. As the
method forecast use of 25 ml samples by safety each sample will be equal to 50 ml.
In the case you cannot perform by yourselves the analysis send well sealed sample
bottles to an independent laboratory or to AXENS laboratory.
A-2 TEST WITH HYDROCARBON TRACERS
A-2-1
Ratio cyclohexane / methyl cyclopentane in reformate.
cC 6
 0.1
M cC5
Above this ratio you can consider that the feed / effluent exchanger is
leaking. For instance a ratio of 0.3 means ~ 5% leak feed to effluent.
A-2-2
Ratio methylcyclohexane + 1Cis2 Dimethyl cyclopentane / 1 trans2
Dimethylcyclopentane in reformate.
McC6  1 cis2DMcC5
 1.7
1 trans2 DMcC5
Above this ratio you can consider that the feed / effluent exchanger is
leaking. For instance a ratio of 2.0 means ~ 5% leak feed to effluent.
Note: These ratios are valid if the components used for calculation are present in feed
for more than 1% each. Also they are valid for a reformate having a RON between
94 and 102.
B
Presence of nitrogen in the feed
The maximum allowable nitrogen content in the feed is 0.5 ppm wt expressed as NH3.
The effect to the catalyst, the causes and remedies have been described in section 2.3.4.
REV. 4
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C
87/ 172
Presence of sulfur in the feed
The maximum allowable sulfur content in the feed is 0.5 ppm wt expressed as sulfur
It is worth noting that sulfur affects primarily the octane number.
The effect, the causes and remedies have been covered in section 2.3.4.
D
Presence of metals in the feed
The presence of metal induces a permanent catalyst poisoning which results in a quick
and noticeable drop in the octane number. It shows mainly through a shift in the T
across the reactors. This will be discussed later (4.6.4).
E
Low chloride content of the catalyst
If the chloride content is too low, the catalyst activity is reduced and the octane is lower
than expected.
If catalyst sampling devices exist, it is possible to check and optimize the chloride levels.
The optimum content lies between 0.9 to 1.1 wt %.
Any shortfall in chloride level can be rectified by increasing the continuous injection level.
It is recommended to do this correction slowly as short large doses can create large
chloride gradients through the reactors, especially in dry units. It may take several weeks
to re equilibrate the catalyst chloride levels following a large deviation.
F
Partial by-pass of the catalyst
If for any reason, part of the catalyst is by-passed, the actual space velocity is higher than
expected and the octane will drop. A higher temperature would be required to make-up
for the octane loss.
The causes of by-passing can be multiple: damaged internals, build-up of fines or scales
at various places in the catalyst bed, catalyst slumping, etc...
A hazardous situation may occur because of channeling, due to the risk of exothermic
hydrocracking reactions taking place in the area where space velocity is very low.
By-passing and channeling are generally identified by increased pressure drop or uneven
and odd temperature profiles.
In such a case the reactor should be dumped, screened and reloaded.
4.6.3 Loss of product yield
A loss in product yield may be caused by one or several of the following:
• Presence of sulfur in the feed.
• Presence of metals in the feed.
• Too high chloride content of the catalyst.
REV. 4
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A
88/ 172
Presence of sulfur in the feed
The sulfur affects primarily the hydrogen rich gas production. Any increase in sulfur
above 0.5 ppm wt reflects immediately in a lower H2 yield and a reduced T across the
first reactor. Refer to 2.3.4. for the causes, effects and remedies.
B
Presence of metals in the feed
This will be discussed later (paragraph 4.6.4) in the item related to decrease T in the
first reactor.
C
Too high chloride content of the catalyst
A too high chloride content (above 1.1 wt. % as Cl) promotes hydrocracking reactions
which result in a lower products yield, a lower hydrogen purity, a lower hydrogen yield,
but a higher octane.
A too high chloride content will be detected through an increased yield of stabilizer
overhead liquid and gas and lower hydrogen recycle purity.
To correct it, chloride injection must be reduced, but not stopped unless severe over
chloriding occurs. It may take several weeks to re balance the chloride levels on the
catalyst.
4.6.4 Unexpected T reduction
The temperature difference T (inlet - outlet) in the various reactors is a very good
indication of the catalyst condition. Note that delta T absolute values are fluctuated
according feed rate, feed quality and recycle gas flow rate. It could be interested to
normalize the delta T of each reactor with the total Delta T for smoothing representation.
Obvious reasons for a T decrease are:
• Feedstock quality: naphthene dehydrogenation is highly endothermic. A lower
naphthene content shows up immediately on the T in the first reactor.
• H2/HC ratio increase (increased recycle). An increased recycle acts as a larger
heat sink for the heat of the reaction.
Other reasons of a T decrease are the consequence of upsets, mentioned earlier, such
as:
• Presence of sulfur in the feed.
• Presence of nitrogen in the feed.
• Too high chloride content (which favors hydrocracking - exothermic reaction).
A loss of T in the first reactor generally results from a sulfur upset in the unit feed. Poor
operation of the pretreatment reactor or of the stripper can then be suspected. Severity
must be reduced, according to the H2S content in the recycle (for H2S > 5 ppm volume,
reduce temperature to 480°C).
REV. 4
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In case of low T in reactors 2 to 4, hydrocracking can be suspected. If such is the case,
H2 purity drops and LPG production increases.
Generally hydrocracking does not occur in the first reactor because the average
temperature is lower.
But the most worrying aspect of a decreased T, when it cannot be related to one of the
above grounds is the result of a permanent poisoning with metals.
Lead, arsenic and other metals mentioned in paragraph 2.3.4 are very severe catalyst
poisons. The maximum allowable contents are shown in table 3 of paragraph 2.3.4.
Due to the very high poisoning activity of the metals, the catalyst in the first reactor is
affected first, and the T through this reactor decreases. The endothermic reactions
(naphthenes dehydrogenation) move then to the 2 nd reactor which exhibits an
unexpected high T.
Heavy metals generally result either from the crude processed or from a poor operation
of the upstream hydrotreating unit. Reprocessing of leaded gasoline (not a common
occurrence to day) can also cause catalyst poisoning.
Metals poisoning requires the dumping and replacement of the catalyst load.
4.6.5 High hydrocracking rate and risk of temperature runaway
Hydrocracking will be suspected when H2 purity drops, the amount of LPG and
stabilization off-gas increases with a simultaneous decrease of product yield.
Hydrocracking can be caused by:
• A fresh catalyst (too active) or an excessive acid function of the catalyst (i.e. high
chloride).
• A too high severity (either a too high temperature at the reactor inlet or a too
low space velocity).
• Poor distribution (channeling) in the reactor which results in local high severity
conditions.
The precautions to avoid hydrocracking are:
• Not to operate at capacity below the design turn-down ratio (generally 60%).
• Not to introduce feed on a fresh or regenerated catalyst before sulfiding.
• Not to introduce feed for the first time on a fresh or regenerated catalyst, at
temperature higher than 400°C. For a subsequent restart the temperature limit
can be increased to 430°C.
• When raising throughput, to raise flow first and increase temperature
afterwards.
• When lowering throughput, to lower temperature first and throughput
afterwards.
• To raise temperature always cautiously: proceed by increment of 2°C/h.
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• To monitor the differential pressure in the reactors. An increased differential
pressure is the sign of poor distribution.
If hydrocracking is suspected:
• Lower the reactor inlet temperature down to 400°C.
• Increase the feed rate to full capacity (to increase space velocity).
• Keep (or set) the recycle gas flow to the maximum.
• Consider catalyst dumping after next regeneration.
Note: Excessive hydrocracking will cause the reformate to become deep yellow. This
colour is due to the formation of polynuclear aromatics.
4.6.6 Stabilizer high pressure drop/bad fractionation
Ammonium chloride salts deposit on the upper cold trays of the column is generally
responsible for this bad behaviour.
Washing requirement can be decided when column becomes unstable, with considerable
fluctuations in the column top pressure and temperature as well as reflux drum and
column bottom levels. Consequently either reformate RVP or LPG C5+ content are
affected.
The following procedure for on-line washing of the stabilizer column can be adopted.
REV. 4

Line-up an empty reformate tank to receive reformate during stabilizer washing
period. Note that the reformate will be badly stabilized and will contain chlorinated
water.

Reduce the stabilizer bottom column down to 140°C. Stop the reflux pump when
reflux drum level is lower than 20%.

Reduce the pressure down to 14 kg/cm2g.

Connect treated water injection line to the reflux line downstream the FCV. Drain
the treated water line to remove iron oxide scales and other debris until the water
is clear.

Close the reflux FCV block valves and by-pass valve. At 140°C at stabilizer bottom
start to introduce the treated water to the column. Introduction of water shall be
very gradual to avoid pressure upsets. Then increase the water flow rate up to 20
wt. % of the feed rate.

Collect water sample from reformate to storage line, downstream the cooler, every
half an hour and check for chlorides and conductivity.

Continue the washing until the chloride level in water for two consecutive samples
is lower than 100 wt ppm.

Discontinue then the treated water injection and drain the reflux line for any holdup water.
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
Wait for 2 hours for the water in the stabilizer to drain and check the reformate
which must be free of water before starting the bottom temperature increase.

Drain the reflux drum, bottom, reflux pumps suction for any water.

Slowly increase bottom temperature and pressure, gradually, establish the reflux
and normalize the column conditions.

Check for reformate RVP and switch the reformate to rundown tank.
91/ 172
Alternatively, the column can be water washed or steamed at the next convenient
shutdown.
REV. 4
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REV. 4
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SHUTDOWN / RE-STARTUP PROCEDURES
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Normal shutdown
The procedure consists of switching-off the feedstock under hydrogen rich atmosphere in
order to reduce the cracking of heavy components in contact with a too hot catalyst.
• Reduce reactor inlet temperature to 460°C at a rate of 25°C per hour.
• Simultaneously, decrease the feed rate to 60% of its design value.
• Stop chlorinated agent and/or water injection to feed.
• Cut-off the feed and maintain the recycle gas circulation during 2 - 4 hours at
460°C and drain liquid hydrocarbon at separator.
• Cool down the catalytic beds to 350°C at a rate of 50°C per hour, then switch-off
furnaces.
If the unit has to be restarted within the next following 24 hours, shutdown the
compressor and leave the reaction section under pressure.
If the shutdown is planned for a longer period, the recycle gas circulation must be
maintained until the catalyst beds are totally cooled down. The recycle compressor is
then shutdown and the reaction section depressurized.
A nitrogen purge is then carried-out in the case maintenance work demands such
precaution.
5.2
Re-startup
The startup can then be carried out according to the following procedure:
• Increase the unit pressure to 8-12 barg, using reformer hydrogen rich gas or
electrolytic hydrogen. In case of hydrogen shortage the pressure can be
increased to the lowest pressure which permits a start of the recycle gas
compressor. A nitrogen make-up is also possible, but hydrogen concentration
shall stay above 60 vol. %.
• Start-up recycle gas compressor.
• Light all furnaces and increase the reactor temperature to 430°C at a rate of
40°C per hour.
• After temperature stabilization, inject feed at 60% of design capacity.
• Adjust reactor inlet temperature at a rate of 30°C per hour up to 460°C, then
20°C per hour up to 480°C and finally by steps of 10°C per hour in order to
obtain the desired reformate octane number, under the condition that moisture
in the recycle gas stays under 30 vol. ppm. On the contrary refer to section 3.7.
• Adjust the required feed rate.
• Resume chlorinated agent and/or water injection to feed
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Shutdown for regeneration
The procedure starts as for a normal shutdown:
• Reduce reactor inlet temperature to 460°C at a rate of 25°C per hour.
• Simultaneously, decrease the feed rate to 60% of its design value.
• Stop chlorinated agent and/or water injection.
• Switch-off the feed.
The subsequent operation consists of ensuring a maximum elimination of hydrocarbons:
• Under circulation of recycle gas, maintain the temperature at reactor inlet at
460°C during 2 - 4 hours and drain all liquid hydrocarbons which have
accumulated in the different cold low parts of the unit and at the separator
drum bottom. Keep same conditions until there is no more purge of liquid HC.
• Decrease the reactor inlet temperature to 350°C at a rate of 50°C per hour.
Switch-off the furnaces and stop the recycle compressor. If unit is equipped with
plate’s feed effluent exchanger like Packinox, it could be recommended to
decrease catalyst beds temperature at lower temperature to avoid thermal
stress when recycle gas compressor will be restarted.
• Decrease the reaction section pressure and purge with nitrogen in order to
eliminate hydrogen and hydrocarbons (sweeping, pressurization /
depressurization, or any other method). Purge the catalyst sampling device to
drain all liquid hydrocarbons they may have retained. The residual content of
hydrogen and of hydrocarbons must be below 1 volume % at the end of the
nitrogen purge operations (follow internal refinery safety policy). Isolate the
reaction section with blinds and proceed with the regeneration operation (see
chapter 6).
5.4
Emergency shutdown
The general guidelines to prevent damage to the catalyst while maintaining its activity
and to protect the main equipment are given here below.
Concerning catalyst preservation the operator must avoid:
• Catalyst bed temperatures exceeding 540°C, to keep grain temperature under
700°C as over this temperature there is a change of alumina structure.
• Presence of hydrocarbons without a sufficient hydrogen partial pressure which
results in a rapid coke deposit and the possible formation of « catalyst cakes »
which favour channeling.
• Quick depressurization of the reaction section which can damage the alumina
support and induce the formation of catalyst fines.
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The refinery safety regulations shall apply and have a priority and precedence over any
other considerations.
Note that reformer emergency shutdown provokes a pretreating unit shutdown if no
alternative supply of hydrogen is available.
5.4.1 Recycle compressor failure
Immediately:
• Switch-off all furnaces and stop the feed pump. Stop also chlorinated agent
and/or water injection pump. Close and isolate the feed valve and the pressure
regulation valve with the associated block valves.
If necessary, open the dampers in order to avoid overheating of the tubes.
• Keep the stabilizer at total reflux conditions.
• If the unit pressure level is sufficient to use again the compressor, start-up this
recycle compressor. When healthy increase the reaction section pressure to 8
barg by hydrogen make-up.
• If the pressure level is too low, carry out an injection of hydrogen to reach the
proper and minimum value necessary for the compressor operation. Start-up
the compressor, then increase the reaction section pressure to 8 barg by
hydrogen make-up.
5.4.2 Loss of feed
Either both pumps failure or more likely upstream units upset.
• Maintain the recycle gas flowrate and the pressure of the reaction section.
• Stop chlorinated agent and/or water injection pump
• If the shutdown is expected to be short, decrease the reactor inlet temperature
to 430°C and maintain this temperature level.
• Keep the stabilizer at total reflux conditions.
• Feedstock will be again injected at this temperature (proceed with normal startup procedure).
5.4.3 Other pumps failure
• Absorber feed pumps, if any.
• Stabilizer feed pumps, if any.
• Stabilizer reflux pumps or heater reboiler pumps.
• Treated water make-up or circulation to boiler pumps.
• etc...
This situation requires to stop the feed to the unit. Refer to paragraph 5.4.2 « Loss of
feed ».
REV. 4
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5.4.4 Utilities failure: fuel gas
All burners of all heaters will be shutdown.
Cut-off feed immediately and proceed as per loss of feed (paragraph 5.4.2) with however
the following difference: in order to insure the highest possible temperature for
hydrocarbons sweeping with the recycle gas, air to reaction heaters (primary and
secondary) must be closed.
5.4.5 Utilities failure: cooling water
In case of total failure the unit must be shutdown:
• Cut-off fire to all reactor heaters and cut-off feed.
• Keep the recycle gas compressor in service as long as possible without risk for
the machine, in order to cool down the heaters tubes and the catalyst.
• Cut-off fire to stabilizer reboiler.
5.4.6 Utilities failure: power supply
Normally recycle compressor is steam driven and will remain in operation except if the
power failure affects also the cooling water system.
Normally instrument power and instrument air shall neither be affected by a power
failure.
All other electric motors (pumps, cooler fans, etc...) are stopped.
• Close feed FCV and separator LCV block valves.
• Shut-off heaters firing (if not stopped by emergency shut down).
• Keep the recycle gas compressor in service as long as possible (take care with
the suction temperature).
• Maintain pressure in reaction section and stabilizer.
5.4.7 Utilities failure: HP steam
Normally the recycle gas compressor is steam driven and will shutdown.
For the shut down refer to paragraph 5.4.1 « Recycle gas compressor failure ».
5.4.8 Utilities failure: instrument/power or air
Both instrument air or instrument power failure will have the same effect:
• All control valves will fail to their safe position.
• Emergency safety valves (XV), if any, will close.
• Recycle compressor will stop.
• Cut-off fuel gas to all burners.
• Stop all pumps.
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• Close feed FCV and separator LCV block valves.
• Maintain pressure in reaction section and stabilizer.
5.4.9 Major leak - fire
The following are only guidelines which must be reviewed and complemented by the
refiner. They are considering the process view point.
In case of major leak the following actions must be taken:
• Cut-off fire to all heaters. Close the stack damper of all heaters and inject
snuffing steam.
• Shut off the feed pumps (from MCC).
• Depressurize the stabilizer and adsorber section if any.
• If possible keep the recycle compressor running for few more minutes to cool
down the heaters coils and sweep the catalyst from hydrocarbons, then shut-it
down.
• Depressurize the reaction section to flare.
• Decrease to a minimum the level in the vessels then stop all the pumps (from
MCC).
• When depressurized vessels cool down watch the pressure and inject nitrogen
to avoid vacuum.
In case of fire.
Generally speaking, all precautions to be taken and operations to be carried-out, are the
same as above and the fire fighting can take place.
Note: If leak occurs in one of the heaters, the hydrocarbons will ignite immediately. In
such a case let the fire develop in the heater box.
REV. 4
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REV. 4
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REGENERATION
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6.1
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General
For bimetallic platinum and rhenium catalysts the coke accumulation on the first reactor
is around 3 to 8 weight % at the end of a normal cycle. The end of a normal cycle is
defined as a C5+ yield decline higher than 2 wt. %.
The coke deposit increases in the successive reactors with a maximum in the last reactor
which can vary from 14 to 22 weight % at the end of a normal cycle.
Deviation of these coke deposit levels is an indication of mechanical and/or operating
problem. Flow by-passing in one reactor can displace the reactions to the next one,
resulting in an unusual coke deposit.
Excessive seal oil leak from the recycle gas compressor will result in a too high coke
deposit on the first reactor catalyst.
Poisoning of the first reactor catalyst will also shift the reactions to the following ones.
The following procedure applies for a normal regeneration carried out after the end of a
normal cycle of operation. It can be adapted if the catalyst has suffered from a
contamination by sulfur (see paragraph 6.5) and in case of poor performance of the
catalyst during the cycle.
As a general rule, when there is a suspicion of sulphur upset during the cycle, H2S in the
recycle gas above 2 ppm vol, it is advised to double the chloride injection versus normal
injection during the two last weeks of operation. This is in order to strip off sulphur from
the catalyst.
Then the unit is shutdown and hydrocarbons are purged as described in the procedure
"Shutdown for regeneration" as per paragraph 5.3.
During the whole regeneration, temperature at the separator shall be kept as low as
possible to minimize recycle gas moisture.
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Catalyst regeneration is composed of several phases summarized as per following table:
Standard Regeneration
Shutdown
Coke burning
Cooldown
Maintenance
Nitrogen Purge
Coke burn proof
Oxychlorination - calcination
Rincing - drying
Cooldown
Nitrogen Purge
Reduction
Sulfidation
Oil-in
REV. 4
Regeneration with sulphur removal
Sulfur removal
Shutdown
Coke burning
Cooldown
Maintenance
Nitrogen Purge
Sulfate removal
Cooldown – Nitrogen purge
Coke burn proof
Oxychlorination - calcination
Rincing - drying
Cooldown
Nitrogen Purge
Reduction
Sulfidation
Oil-in
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CATALYST REGENERATION AND REACTIVATION LOGIC DIAGRAM
FIRST COMBUSTION
RiT’s 400°C - T max 50°C
O2 0.1  0.5 vol. %
C2 Cl4 injection and Caustic solution
circulation at RiT’s 350°C
SECOND COMBUSTION
RiT’s 480°C - T max 30°C
O2 0.5  1.0 vol. %
C2 Cl4 injection in service
Caustic solution circulation
YES
MAINTENANCE
WORK and / or
CATALYST
SIEVING
NO
PROOF COMBUSTION
RiT’s 520°C - T max 0°C
O2 1.0 – 2.0 vol. %
C2 Cl4 injection in service
Caustic solution circulation
COOLING
RiT’s 520°C  50°C
Stop C2 Cl4 injection (350°C)
Stop caustic solution circulation(350°C)
O2  3.0 vol. %
OXYCHLORINATION
RiT’s 520°C - T max 0°C
O2 2.0  5.0 vol. %
C2 Cl4 injection in service
Caustic solution circulation
CALCINATION
RiT’s 520°C - T max 0°C
O2 6.0 to 8.0 vol. %
C2 Cl4 injection in service
Caustic solution circulation
MAINTENANCE WORKS
O2  21.0 vol. %
Works execution
O2  1.0 vol. %
RINSING / DRYING
RiT’s 520°C O2  3 vol. %
C2 Cl4 injection in service
Stop caustic solution circulation
Rinsing with semi water
Stop C2 Cl4 injection
Drying period
HEATING
RiT’s up to 520°C
O2 = 1.0 vol. %
C2 Cl4 injection at 350°C
Caustic solution circulation (350°C)
Note : Maintenance work has to be performed before
or after second burning step, never after calcination.
COOLING
RiT’s 520°C  200°C
O2  3 vol. %
Stop heaters – Stop RG compressor
O2  0.3 vol. % max
READY FOR REDUCTION
REV. 4
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Preparation of the unit
The reaction section is hydrocarbon free and put under slight positive nitrogen pressure.
The unit is then prepared for regeneration as follows:
• Turn all necessary blinds and isolate the unit from all pipe systems containing
hydrocarbons.
• Install gas sampling systems at reactor outlet and at high pressure separator.
The major part of the gases analyses are done with Drager Tube. These tubes,
although not having a high precision, will give acceptable data, if a correct
sampling system is used. It is mandatory that sample lines be as short as
possible to control HCl, H2S, SO2.
• Install an O2 and CO2 analyzers and connect to the different gas sampling
systems.

Note that higher the regeneration pressure, the shorter will be the
regeneration duration and the lower the chloride consumption, if equipment
allows.
• Prepare all lines for air injection and check the recording / reading of air flows.
In case of an external air compressor is used, it should be oil-free. Purge and
flush those pipes to eliminate water and various debris
• Check the pump and the pipes used for chlorinated agent injection as well as the
storage tank. Check the individual lines to each reactor. The chloride injection
device shall be installed as per figure 6-5.
Wash with demineralized water from upstream air coolers to separator. Drain at
separator bottom until total solid salts are below 2 wt%. Afterwards isolate the
recycle gas compressor by blinds, then wash with demineralised water the
separator, the compressor suction and discharge line. This will avoid recycle gas
compressor vibration problem and subsequent spurious shutdown caused by
the carry-over of salts to the compressor due to the high moisture content of
the recycle gas during the coke combustion. It is also wise to wash the
effluent/feed exchanger with hot water. If the unit is equipped with plates
exchanger (Packinox type) proceed to a chemical cleaning according
manufacturer instructions.
• Certain refiners wash the recycle gas compressor with condensates additioned
with sodium carbonates every regeneration after coke combustion (see with the
compressor manufacturer for exact procedure). Others prefer to dismantle the
compressor with same frequency.
• Install at the caustic solution inlet to process an injector as shown hereunder:
REV. 4
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• Prepare a caustic soda/demineralized water solution with a 4 wt. %
concentration of NaOH in the separator drum. Check and prepare the
circulation pump and its flow controller. The required flow is 25 to 35% of the
design feed flowrate to the unit, in order to obtain a good washing and a
sufficient neutralization of hydrochloric acid present in the effluent of
combustion. The caustic solution is injected upstream the effluent air coolers.
• Prepare an analytical program with control laboratory to monitor the complete
elimination of sodium carbonates and to predict the soda make-up into the
section when the solution does not show a sufficient neutralization effect.
• Set the unit under a nitrogen pressure at 1 bar below the pressure of the
regeneration air system while maintaining the normal flow direction through the
reactors. The reactor average pressure is adjusted between 5 and 15 barg
depending on the air supply pressure. Higher pressure will obviously reduce the
combustion period.
• Commission all effluent air cooler fin fans and check that cooling water is in
service at trim coolers.
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Coke combustion - oxychlorination - calcination
The combustion air is introduced at recycle compressor suction.
develops from first reactor to last reactor.
The combustion
For a reactor where coke burning develops, it is necessary to control carefully the inlet
temperature in order to avoid any run-away. Note that carbon and hydrogen in the
chlorinated agent will be oxidized to carbon dioxide and water with corresponding heat
release and oxygen consumption (see figure 6-3).
• Start the compressor.
• Make sure that the reactor effluent coolers are in their maximum service and
that the compressor flow rate is set at its maximum value compatible with the
regeneration operating conditions.
• Light the furnaces and increase the inlet reactor temperature to 400°C at a rate
of 50°C per hour.
• At 350°C, start the anti-corrosion loop circulation and begin injecting the
chlorinated agent to the 1st reactor. The injection rate should be such that the
molar ratio H2O/Cl is 20/1. See the attached Table 6-1 for separator drum
pressure and temperature. This injection is carried out until the end of the
calcination step.
Note: If there exist an air line to the last reactor equipped with a FI, parallel burning of
1st and last reactor can be undertaken. Chloride injection is done to first reactor
inlet with molar ratio of 20/1. If a line is available, additional chloride injection
with molar ratio of 30/1 should be put in place at last reactor inlet. Particular
care will be taken at the end of upstream reactor catalyst coke burn as to avoid
excessive O2 concentration at last reactor inlet by decreasing air flow-rate at last
reactor inlet.
Notes for caustic circulation:
REV. 4
A
Before establishing a caustic solution level in the separator circulate clear
water in the washing circuit and drain carefully this initial inventory (see
paragraph 6-2).
B
Make-up again clear water then inject caustic soda so as to have a
concentration of 4 wt. % maximum in the solution.
C
During the chloride injection monitor the concentration so as to have a pH of
the solution between 7.0 and 8.0 and density between 1.007 - 1.015.
D
Control carefully the caustic solution circulation flow (25 to 35% design feed
flow rate).
E
Do not exceed 5 wt. % of solids in the solution.
F
Alkalinity to be kept at 2 to 3 wt. % of NaOH equivalent.
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6.3.1 Coke combustion
FIRST COMBUSTION
• When all catalyst beds have reached a temperature of 400°C, introduce the
regeneration air at compressor suction and follow the evolution of temperature
inside the catalyst beds. If possible, all temperatures must be recorded.
• The regeneration air flowrate is controlled in such a way that the temperature
increase through the catalyst bed does not exceed 50°C, (outlet Rx temperature
-inlet Rx temperature).
The temperature of the combustion zone (flame front) shall not be higher than 500°C.
This can only be checked if bed thermocouples are available. Excessive exotherm is
controlled by decreasing the air flowrate in order to maintain a temperature at reactor
outlet not exceeding 450°C (the injection of air can be eventually stopped in case of
excessive exotherm above 60°C)
• In order to monitor the temperature increase, the oxygen concentration at
reactor inlet is controlled in the range of 0.5 volume % and under no
circumstances must exceed 0.8 volume %.
The burning generally starts with an oxygen concentration of 0.1 vol% in the gas entering
the reactor. Solid burn starts when oxygen concentration is established at 0.3 vol%. As
long as combustion takes place, the air flowrate must be adjusted in order to limit the
temperature increase through the catalyst bed at 50°C or to maintain the reactor outlet
temperature at 450°C max.
During the whole phase of combustion, the O2 and CO2 contents at reactor inlet and
outlet are regularly (every hour) monitored by either online or portable analyzers.
Adequate sampling devices must be installed for that purpose. Check with Draeger tube
of SO2 and HCl could be performed at each reactor’s outlet. An excess of SO2 (> 30 - 40
vol. ppm) could be the sign of catalyst contamination by sulphur. HCl breakthrough (> 20
vol. ppm) should be noticed.
When delta T between inlet and outlet becomes negative (-2/3°C) in the last reactor and
when the oxygen concentration at last reactor outlet equals the oxygen concentration at
reactor inlet, the combustion phase is considered complete. In the case parallel burning,
air to last reactor shall be slowly cut-off at the end of upstream reactor burning, so as to
have all combustion air injected at compressor suction. Simultaneously chloride injection
at last reactor (if any) will be stopped.
The combustion is considered complete when all reactor deltaT’s have been come back
at their base level for a period of minimum 4 hours and with oxygen and CO2
concentrations at reactor inlet’s and outlet’s within the same range.
SECOND COMBUSTION
When the first combustion is completed and the reactor temperatures are stable at
400°C, the inlet temperature is then progressively increased to 480°C in 4 hours with an
oxygen concentration of 0.5 to 1 volume % in the recycle gas, exotherm is not
representative during the temperature increase. Watch for the second combustion to
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take place. Wait stable inlet and outlet temperature to conclude that combustion is
ended.
A maximum temperature increase of 30°C through the catalyst bed is acceptable,
corresponding to a maximum temperature inside the reactors of 510°C. Where bed
thermocouples are available, individual point may not exceed 520°C.
PROOF BURNING
The maximum oxygen content in the gas at reactor inlet must be between 1 to 2 volume
%. When the second combustion is completed, the reactor inlet temperatures are
increased to 520°C in one hour. Keep RiT’s at 520°C during 2 hours. No post combustion
shall be noted at this stage. On the contrary reduce inlet temperature to 480°C.
Temperature increase can only be resumed when no temperature increase exists in the
reactors.
If no maintenance work is forecast proceed directly to oxychlorination step.
If maintenance work and/or catalyst unloading for screening are planned, it shall take
place at the end of carbon burn (first or second combustion step) and burn proof will be
done after unit restart up for catalyst reactivation. If so, stop chlorinated agent injection
and caustic solution circulation. Drain the caustic solution, wash with clear water, drain
at all low point then start cooling down the unit with the recycle gas compressor running
to a temperature at which the unit can be inspected and the catalyst unloaded. During all
the cooling down period maintain at least 3.0 vol. % of O2 in the recycle gas. Refer to
COOLING step in paragraph 6.3.3.
6.3.2 Catalyst sieving and reloading
Axens recommends catalyst unloading for screening after the combustion of the coke
during 3rd regeneration or after 5 years operation, whichever comes first. When
channeling is suspected (coke burning tailing), or abnormal pressure drop it is mandatory
to sieve the catalyst and inspect the concerned reactor internals.
Catalyst shall be unloaded from reactor bottom unloading nozzle. Unloading by “sucking”
(vacuum system) must be avoided as it leads to huge breakage of catalyst.
CATALYST SIEVING
Catalyst sieving machine is installed under the reactor to be unloaded. Sieved catalyst /
inert balls / fines dust and broken extrudates are recovered in drums.
Two upper grids (mesh opening  16 mm and  4 mm) allow the segregation of  3/4"
and  1/4" inert balls. The lower grid (mesh opening  1 mm) allows the elimination of
fines, dust and broken extrudates from the sieved catalyst.
Note that sieving a catalyst withdrawn from the bottom of the reactor leads to a
catalyst loss of  3-5%. On the contrary sieving catalyst unloaded by “sucking” can lead
to losses a high as 20%. In both cases fresh catalyst shall be available on site to
compensate the losses.
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CATALYST RELOADING
Contaminants affect mainly the catalyst of first reactor. It is not recommended to mix
first reactor catalyst with catalyst of others reactors.
General rules are:
-
Avoid mixing first reactor catalyst in downstream reactors as it is
the most contaminated catalyst
-
Compensate losses due to sieving with last reactor catalyst and
complete last reactor catalyst with new one.
For example:
If the catalyst of first reactor doesn’t need to be replaced, reload as follows:
• First reactor: Its own remaining catalyst plus catalyst from the last reactor or
fresh catalyst.
• Second reactor: Its own remaining catalyst plus catalyst from the last reactor.
• Third reactor: Its own remaining catalyst plus catalyst from the last reactor.
• Fourth reactor: Its own remaining catalyst plus new fresh catalyst make-up to
compensate the losses occurring during the sieving
If the catalyst of first reactor needs to be replaced, reload as follows:
• First reactor: Catalyst from the last reactor or new catalyst.
• Second reactor: Its own remaining catalyst plus catalyst from the last reactor.
• Third reactor: Its own remaining catalyst plus catalyst from the last reactor.
• Fourth reactor: Its own remaining catalyst plus new fresh catalyst make-up to
compensate the losses occurring during the sieving.
6.3.3 Catalyst reactivation
OXYCHLORINATION
If the unit has been shut down after coke combustion for maintenance work the unit is
first swept with nitrogen to reduce the oxygen concentration down to ~ 1 vol. % and then
the reaction section pressure increased to the regeneration pressure.
Then the recycle gas compressor is started-up, heaters are fired and reactor inlet
temperatures increased to 520°C (50°C/h). If unit was shutdown for maintenance work
and/or catalyst unloading for screening, burn proof step will be done before increasing
oxygen concentration.
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At 350°C the anticorrosion loop circulation is started up and the chloride injection
resumed (20/1 H2O/Cl molar ratio). Chloride injection shall be started only when oxygen
is present to promote chlorinating agent decomposition.
If unit was shutdown after the primary coke burning, resume operation at the second
burning step.
Bed average temperature 510°C, no sign of combustion:
• At 510°C stable average bed temperature, the oxygen concentration in the gas is
increased to 5 volume % in 2 hours while observing carefully for any restart of
the combustion (see figure 6-4). If necessary, reduce or stop the air injection.
Check that CO2 concentration stays lower than 3 - 5 vol. %. If needed put in
place nitrogen make-up.
• Maintain the conditions – bed temperature = 510°C and 5 vol. % O2 during 4
hours and check every hour that the HCl content in the gas at reactors outlet is
higher than 40 ppm vol. (Draëger tubes).
CALCINATION
• Increase the oxygen concentration up to 8 vol. % (for pressure < 5 kg/cm2g see
figure 6-2) in one hour.
• Note that some of the centrifugal recycle compressors have a maximum O 2
content on the process side of 5 vol% due to concerns about O 2 leaking into the
compressor lube system. 0.7 bar abs. O2 partial pressure is the lower limit for
obtaining good platinum redispersion in a semi-regeneration reformer. Please
refer to figure 6-2
• Then, maintain such conditions at bed temperature of 510°C for 6 hours without
exceeding an outlet reactor temperature of 520°C. Check that CO2 concentration
stays lower than 2 - 3 vol. % and do a nitrogen make up to purge if necessary.
RINSING / DRYING
• At the end of the calcination phase at bed temperature 510°C, stop the anticorrosion loop circulation and drain the caustic solution at all cold low points.
• At this step and before start to decrease RIT, it is very important to eliminate the
liquid water in order to avoid that catalyst catch water during cooling.
• The anti-corrosion loop will then be rinsed away with clear water then drained
carefully at all cold low points. Stop the chlorinated agent injection as soon as
the rinsing water make-up is stopped. The caustic solution purging, cleaning
with clear water and draining the clear water shall be done as quickly as
possible. During all these operations the temperature at reactors inlet is
maintained at 520°C. Temperature shall only be decreased when the total
clear water drained at all cold low points is lower than 0.05 wt % of the
catalyst per hour. This drying period should be carried-out in about 4 hours and
should not exceed 6 hours.
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COOLING
It is important to complete flushing and draining at 510°C before cooling. If not, catalyst
will absorb excessive water. During reduction, this water could be desorbed with
subsequent chloride loss and possible metal sintering.
While maintaining an oxygen concentration in the recycle gas higher than 3 vol. %, the
reactor outlet temperatures are decreased to 200°C in all catalyst beds at a rate of 50°C
per hour. When unit is equipped with plate’s feed effluent exchanger like Packinox, it
could be recommended to decrease catalyst bed temperature below 180°C to avoid
thermal stress when compressor will be restarted. Check with exchanger manufacturer.
Axens has no objection to decrease catalyst bed temperature below 180°C. At this
temperature, furnaces and recycle compressor are shutdown.
NITROGEN PURGE AND TURN BLINDS
The reaction section is depressurized and purged with nitrogen in order

to decrease oxygen down to 0.5 vol % but never below 0.2 vol % to protect
catalyst

to decrease CO below 10 vol. ppm and CO2 below 500 vol. ppm.

If the nitrogen purge follows a regeneration, purge the section until the SO2
content is lower than 25 vol. ppm.
Swing open the isolation blinds (except the one on the feed) as soon as possible. To avoid
any pollution of catalyst at this stage, it is mandatory to proceed by pressurization and
depressurization with dry nitrogen. (See note 1 paragraph 3.4).
As soon as the O2 content is lower than 0.5 volume %, nitrogen is displaced by electrolytic
hydrogen and the pressure of the reaction section increased to a minimum of 7 barg.
Immediatly after, continue with catalyst reduction (see paragraph 3.4), then catalyst
sulfiding (see paragraph 3.5).
Note 1: When the unit is equipped with catalyst samplers, catalyst samples should be
taken at the end of combustion phase. This could allow to finely adjust the
chlorine injection.
Note 2: If for any reason reduction cannot take place immediatly after the post
calcination drying the catalyst beds shall be cooled down to 180°C while
keeping an oxygen concentration ~ 3 vol. %. Nitrogen purge will take place only
when the forecast for reduction start-up is known.
Note 3: When equipments, which were by-passed during combustion and oxidation
steps, are put again in the circuit, particular attention has to be done regarding
presence of hydrocarbon in these equipments. When catalyst is under oxide
form, hydrocarbons are poisons for metal function.
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Emergency handling procedure for regeneration
Following broad approach (modified suitably for each specific site) will be used as
guideline.
1- If the caustic circulation fails and can not be restored immediately, stop
chloride injection then block in the air injection. Maintain the recycle gas
compressor operation and maintain the catalyst beds hot. Caustic make-up to
be stopped. If the duration is forecast to be long, decrease the reactor inlet
temperatures down to 400°C and maintain O2 concentration between 0.5 and
0.8 vol. %, resuming air injection.
2- If the catalyst temperatures become excessive, delta > 50°C, reduce the air
injection to all reactors. Investigate the incident thoroughly before resuming
the normal air injection.
3- If the recycle compressor fails, immediately stop chloride injection, block in the
air injection and caustic circulation. Cut-off fire to heaters. Try to maintain the
system pressure at 6 barg, with N2.
If the recycle compressor fails when first burn is in progress ensure that the O 2
content is in the range of 0.5 - 0.8% never above 1%. Inject N2 to maintain O2
content in this range if needed.
If the recycle compressor fails when final burn is in progress ensure that the O 2
content is in the range of 0.8 - 1.2% never above 1.5%. Inject N2 to maintain O2
content in this range if needed.
If the recycle compressor fails when calcination is in progress ensure that the
O2 content is in the range of 6-8 vol% never below 5%.
When recycle gas compressor is recovered, in each of the above cases ensure
that the O2 limits are not exceed before you increase the temperatures.
4- In all cases be sure that caustic does not find its way to the catalyst. Also do
not inject chloride when recycle compressor is stopped or when air is stopped
for long duration.
5- If the proceedings are to be halted for long duration before reduction, a
catalyst preservation procedure shall be adopted. See note 2 in the paragraph
6.3.3.
6- During steps with high oxygen concentration, in case of hydrocarbon
introduction in the reaction section (oil from compressor for example) a large
exotherm will take place. That could conduct to catalyst damage, potentially
irreversible depending of which temperature was reached. Immediately stop
air injection and cool down RiT’s to 400°C. Restart regeneration procedure at
burning step.
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Sulfur and Sulfate removal procedures
6.5.1 Sulfur removal procedure
This procedure is not included in the standard regeneration procedure.
Due to the low sulfur of reformer feed when multimetallic catalysts are used, the sulfur
removal should not be necessary. Check with Axens before proceeding.
Anyway, if due to a malfunction of the HDT unit, the reformer feed contained  1 wt ppm
total sulfur for a long period this procedure can be used. If the unit is equiped with
sampling devices, a total sulfur of more than 1000 wt ppm on the catalyst, means than
the sulfur removal shall be performed.
Contamination by sulfur can be reduced by circulating high purity hydrogen at high
temperature transforming sulfur compounds in H2S later on removed by a caustic
solution washing.
As chloride and sulfur occupy concurently the catalyst sites, addition of chloriding agent
during sulfur removal will help sulfur elimination.
As sulfur is transformed in sulfate during regeneration it is wise, especially with catalyst
at start of life, to remove the sulfur before the regeneration. Sulfates fixed on the
catalyst support are very difficult or impossible to remove.
It is important to perform the sulfur removal before coke combustion when the unit
suffered an important sulfur poisoning as to avoid an important formation of sulfates
during combustion. But it is to be considered that more coked the catalyst, more difficult
the sulfur removal.
After the feed has been switched off at 460°C, reactors inlet temperature is increased up
to 510°C and maintained at this value during two hours to strip the hydrocarbons
adsorbed on the catalyst and drain all hydrocarbons which have accumulated in the
different cold parts of the unit. Then reduce reactors inlet temperature to 400°C.
• Decrease the pressure of the reaction section to the minimum the recycle gas
compressor can accept.
• By injecting electrolytic hydrogen increase the pressure of the reaction section
to a pressure which can allow a smooth operation of the recycle gas compressor
with low MW gas. Repeat this operation if necessary as to have a hydrogen rich
gas containing less than 5 vol. % of C3+, then increase the pressure up to  10
bar at separator.
• Start the neutralization solution circulation (See paragraph 6.2).
• Start chlorinated agent injection to have a 20/1 H2O/Cl mole ratio (see table 61).
• Increase reactors inlet temperature to  530°C at a rate of 40°C/hour.
 inlet T + outlet T 

 = 525C
2


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• Maintain the pressure of the reactor by injecting electrolytic hydrogen
• Check H2S content at the outlet of each reactor. Plot the results on a graph.
• When H2S content at the outlet of all reactors is lower than 2 vol. ppm sulfur
removal is considered as complete.
• Stop chlorination agent injection and neutralizing circulation. Drain the
neutralizing solution, rinse the circuit with fresh waterand drain the cold parts of
the reaction section.
• Cool down the reaction section to 350°C at a rate of 50°C /hour. Switch-off the
furnaces and stop the recycle gas compressor.
• Decrease the pressure of the reaction section to flare and purge with nitrogen in
order to eliminate any remaining H2S, until residual hydrogen results lower than
0.1 vol. %. Isolate reaction section with blinds and proceed with the
regeneration procedure.
6.5.2 Sulfate removal procedure
If the sulfur removal is unsuccessful, either checked by sulfur analysis if the unit is
provided with catalyst samplers or if the H2S in recycle gas stayed very low during the
sulfur removal phase, sulfate removal shall be carried-out after the oxychlorination step.
Basically sulfate removal procedure is identical to sulfur removal one. To be noted that
during coke combustion part of the sulfur is eliminated under the form of SO2 while the
rest is transformed in sulfates and sulfur oxides.
The reaction section has the following conditions at the end of combustion:
• Recycle compressors in service, RiT’s = 510°C, O2 content in the recycle gas  1
vol. %, all reactors T’s negative, neutralization circulation in service.
The sequence is as follows:
• Stop chlorinated agent injection and neutralisation solution circulation.
• Rinse the circuit with fresh water, then drain carefully all cold low points.
• Decrease RiT’s from 510°C to 350°C.
• Cut combustible to fired heaters and stop the recycle gas compressor.
• Depressurize the section. Purge the section with nitrogen until O2 < 0.3 vol. %.
• Set blinds in reduction position.
• Introduce electrolytic hydrogen and follow exactly the sulfur removal procedure
(paragraph 6.5.1), the only difference being that RiT’s are decreased to 480°C,
in one hour instead of 350°C before switching-off the heaters and stopping the
recycle gas compressor.
• Pressurize the section with nitrogen up to the regeneration pressure.
• Start the recycle gas compressor and cut-on fire to the heaters.
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• Slowly inject air until obtaining 1 vol. %. Note that residual combustion can be
expected due to the carbon content of the chlorinating agent used during the
sulfate removal.
• Resume neutralisation solution circulation and chloriding agent injection.
• Increase RiT’s up to 510°C and follow catalyst reactivation procedure (paragraph
6.3.3)
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Miscellaneous
Flow meters reading correction:
A
Correction for gas given in volume
K = R x
1
B
1
MW d x P op x T d
MW op x P d x T op
Correction for gas given in weight
K = R x
2
2
MW op x P op x T d
MW d x P d x T op
Where:
Note:
K1
:
Corrected flow at 0°C and 1 ATM in Nm3/h.
K2
:
Corrected flow in kg/h.
R1
:
Instrument reading in Nm3/h.
R2
:
Instrument reading in kg/h.
MW d
:
Design molecular weight.
MW op :
Operating molecular weight.
Pd
:
Absolute design pressure in ATM.
P op
:
Absolute operating pressure in ATM.
Td
:
Design temperature in °K.
T op
:
Operating temperature in °K.
A molecular weight ~ 30 can be taken into consideration for the recycle gas flow
calculation during the regeneration if you have not a more accurate analysis.
Calculation of catalyst coke levels:
Assuming that coke molecular formula is CH1.3, the coke combustion can be represented
by the following equation:
CH1.3 +
2.65
O2  CO2 + 0.65 H2 O
2
Thus 0.7547 mole of coke is combusted for each mole of O2 consumed. This means that
it requires 2.23 Nm3 of oxygen to burn 1 kg of coke or 10.62 Nm 3 of air to burn 1 kg of
coke.
If the amount of air is known (air inlet minus losses) the quantity of coke combusted can
be calculated directly.
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On the contrary the following formula can apply:
 O2 inlet (vol. %) - O2 outlet (vol. %) 
x Recycle gas (Nm 3 / h)
O2 combustion rate = 

100


3
Total O2 consumption = O2 consumption rate (Nm /h) x combustion time (hours).
Coke combusted =
Total oxygen combustion (Nm 3 )
2.23 (Nm 3 / kg)
Example of coke calculation:
• Recycle gas
=
30 000 Nm3/h
• O2 concentration at 1st reactor inlet =
0.8 vol. %
• O2 concentration at last reactor outlet
=
• Combustion phase duration
12 hours
=
0.1 vol. %
Oxygen consumption:
 0.8 - 0.1
30 000 x 
x 12 = 2520 Nm 3

100


Combusted coke:
2520 Nm 3
= 1130 kg
2.23
It can also be calculated as follows:
The total quantity of air injected during the combustion phase multiplied by a coefficient
for air losses (0.9) corresponds to the total quantity of coke.
For instance 13 333 Nm3 of air corresponds to:
13333 x 0.9
= 1130 kg of coke
10.62
Repartition of the coke between the reactors can be calculated using the sum of hourly
positive T’s obtained in each reactor and the sum of hourly positive T’s for all reactors
during the combustion.
R1 T =
200°C coke =
R2 T =
600°C coke =
R3 T =
2200°C coke =
All Rx T =
3000°C coke =
1130 x 200
3000
1130 x 600
3000
1130 x 2200
3000
=
75 kg
% on catalyst =
=
226 kg
% on catalyst =
=
829 kg
% on catalyst =
1130 kg
% on catalyst =
75
1327
226
3000
829
5838
1130
10165
= 5.65
= 7.53
= 14.20
= 11.12
Eventhough these methods are not very accurate they give a good idea of the total coke
deposited during the cycle as well as the repartition of the coke.
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TABLE 6-1
Injection rate of chloride
during the phases of combustion and calcination
related to conditions at separator drum
Molar ratio H2O/HCl = 20/1
Temperature
°C
4
bar g
5
bar g
6
bar g
7
bar g
8
bar g
9
bar g
10
bar g
11
bar g
12
bar g
13
bar g
14
bar g
15
bar g
10
1,92
1,60
1,37
1,20
1,07
0,96
0,87
0,80
0,74
0,69
0,64
0,60
15
2,67
2,22
1,90
1,67
1,48
1,33
1,21
1,11
1,03
0,95
0,89
0,83
20
3,66
3,05
2,61
2,29
2,03
1,83
1,66
1,52
1,41
1,31
1,22
1,14
25
4,95
4,13
3,54
3,10
2,75
2,48
2,25
2,06
1,91
1,77
1,65
1,55
30
6,64
5,53
4,74
4,15
3,69
3,32
3,02
2,77
2,55
2,37
2,21
2,07
35
8,79
7,33
6,28
5,50
4,89
4,40
4,00
3,66
3,38
3,14
2,93
2,75
Expressed in Kg of chloride Cl per 10,000 Nm3/h of recycle gas.
Note: 1 Kg of chloride Cl =
0.68 liter of carbon tetrachloride
0.72 liter of tetrachloroethylene
0.85 liter of trichloroethylene
0.96 liter of trichloroethane
1.19 liter of dichloroethane
1.37 liter of dichloropropane
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WATER DIAGRAM IN AIR
FIGURE 6-1
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OXYGEN REQUIREMENTS DURING CALCINATION
FIGURE 6-2
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REGENERATION: COKE COMBUSTION
FIGURE 6-3
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REGENERATION: OXYCHLORINATION / CALCINATION
FIGURE 6-4
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CHLORINE INJECTION DEVICE
FIGURE 6-5
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7
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SAFETY AND HEALTH RECOMMENDATIONS
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7.1
123/ 172
General
A list of health and safety data sheets including the catalyst plus some of the chemicals
involved in the Reforming process is given below.
Regarding chemicals, the list mentions those which are specific to the AXENS process.
Consequently health and safety data related to well known hydrocarbons are not being
considered here.
The material safety data sheet of the catalyst provided by the manufacturer is attached.
Regarding the other material safety data, the Refiner is advised to request the last issue
of the following document: Regulated Hazardous Substances
published by: The Occupational Safety and Health
Organization
(OSHA)
US Department of Labor
7.2
List of health and safety data sheets
• RG / PR series
• Tetrachloroethylene
• Trichloroethylene
• Trichloroethane
• Dichloropropane
• DMDS
7.3
Catalyst safety data sheet
Updated MSDS are available on the following internet link:
http://www.quicksds.com/en/index.html
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8
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ANALYTICAL CONTROL
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8.1
125/ 172
Recommended methods and frequency
The frequencies given below routine related to a normal operation. During start-up and
test runs more analyses will be required. The following schedule is proposed for the
analysis.
The methods nominated "ASTM D-" are copyrighted by ASTM International ("ASTM"), 100
Barr Harbor Drive, West Conshohocken, PA 19428-2959 USA; http://astm.org/. The right
of use of these methods must be obtained from this organization.
8.1.1 Feed
Product
Feed
Analysis
Method
Frequency
Composition
IFP 9301
2 per week
Specific gravity
ASTM D 1298 / D 4052
1 per day
Distillation
ASTM D 86
2 per week
Molecular weight
IFP 9413
As required
Total sulfur content
ASTM D 5453 (1) / D 4045
1 per day
Doctor test
ASTM D 4952
As required
Water content
ASTM D 6304
As required
Oxygenated components
ASTM D 5599
As required
Total chlorine content
ASTM D 4929
1 per week
Total nitrogen content
ASTM D 4629
1 per day
Bromine index (2)
ASTM D 2710
As required
Arsenic content
IFP 9312
As required
Copper content
ASTM D 6732
As required
Iron content
ASTM D 5863 / D 5708
As required
Lead content
IFP 9406
As required
Mercury content
IFP 9606
As required
Nickel content
IFP 9507
As required
ICP (3) / AAS (4) (5)
As required
Other metal contaminants
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8.1.2 Products
Product
Method
Frequency
Composition
IFP 9302
2 per week
Distillation
ASTM D 86
1 per day
ASTM D 1298 / D 4052
1 per day
RON clear
ASTM D 2699 or carburane (6)
1 per day
MON clear
ASTM D 2700 or carburane (6)
1 per day
Reid Vapor Pressure
ASTM D 323 / D 5191
2 per week
Bromine index (2)
ASTM D 2710
As required
Existent gum
ASTM D 381
As required
Potential residue
ASTM D 873
As required
IFP 9908
1 per day
Draeger tube 6728041
As required
Water
ASTM D 5454
1 per week
HCl
Draeger tube CH 29501
1 per week
Ammonia
Draeger tube CH 20501
As required
Hydrogen production
Composition
IFP 9908
1 per day
LPG
LPG sampling method
ASTM 1265
As required
Composition
IFP 9909
1 per day
Density
ASTM D 1657
1 per week
Volatility
ASTM D 1837
1 per day
Vapor pressure
ASTM D 1267
As required
Copper strip
ASTM D 1838
As required
Total sulfur content
ASTM D 5453 (7) / D 3246 / D 4468
As required
Total chlorine content
ASTM D 4929 modified (8)
As required
ASTM D 2420 / D 4084
As required
ASTM D 6228
As required
Draeger tube CH 29501
As required
IFP 9909
1 per day
Stabilized reformate
Analysis
Specific gravity
Recycle gas
Composition
Hydrogen sulfide
H2S
Sulfur species
HCl
Fuel gas
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(1) ASTM D 5453 is recommended on usual basis. Total sulfur by ASTM D 4045 could be
used if necessary for cross-checking.
(2) Bromine index (mg Br2 / 100g) for low olefin content.
(3) ICP = Inductively Coupled Plasma.
(4) AAS = Atomic Absorption Spectrometry.
(5) Specification below available techniques quantification limit, back calculated on the basis of
spent catalyst analysis.
(6) Vinci Technology Licensee.
(7) ASTM D 5453 is recommended on usual basis. Total sulfur by ASTM D 3246 or D 4468 could be
used if
necessary for cross-checking.
(8) ASTM D 4929 part B: Oxidative combustion and microcoulometry detection used LPG injection
loop.
Product
Analysis
Method
Frequency
Regeneration gas
O2
IFP 9810 or Refinery GC Analyser
As required
CO
IFP 9810 or Refinery GC Analyser
As required
Draeger tube CH 19701
CO2
IFP 9810 or Refinery GC Analyser
As required
Draeger tube CH 23501
HCl
Draeger tube 6728181
As required
SO2
Refinery GC Analyser
As required
Caustic solution
Alcalinity
UOP 209
1 per shift
(regeneration)
Total solids
ASTM D 5907
1 per day
pH paper
1 per hour
pH
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8.1.3 Catalyst
Catalyst (12)
Carbon
310 CA 220 / ASTM D 3178 (9)
1 per week for spent catalyst
}As required for regenerated
310 CA 207 / IFP 9303 (10)
Chlorine content
catalyst
Na content
IFP 9508
As required
Sulfur content
310 CA 220 / IFP 9303 (10)
1 per week
ICP (3) or WDX RF (11)
As required
ASTM D 3663
As required
Other metals contaminants
Specific surface
(9) Recommended laboratory analyser: Eraly Analyser.
(10) To be used on usual basis. WDXRF method IFP 9303 could be used if necessary for
cross-checking. WDXRF = Wave-Length Dispersive X-Ray Fluorescence.
(11) Matrix matched calibration.
(12) From usual practice, only carbon, sulfur and chloride determination are performed
in refinery laboratory. Other determinations are seldom required and can be
performed in IFP laboratories or other specialised laboratories.
8.2
IFP Analytical methods
IFP 9301
Petroleum naphtha
Detailed analysis
Capillary gas chromatography
IFP 9302
Stabilized reformate effluent
Detailed analysis
Capillary gas chromatography
IFP 9303
Alumina-based catalysts
Determination of chlorine and sulfur
Wavelength dispersive x-ray fluorescence
IFP 9312
Petroleum products analysis
Determination of arsenic
Graphite furnace electrothermal atomic absorption spectrometry
IFP 9406
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Gasoline and naphtha analysis
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Determination of trace amounts of lead
Graphite furnace electrothermal atomic absorption spectrometry
IFP 9413
Gas condensates and crude oils up to 340 °C
Direct determination of the molar masses
Capillary gas chromatography
IFP 9507
Petroleum products
Determination of nickel and vanadium
Inductively coupled plasma atomic emission spectrometry
IFP 9508
Alumina-based catalysts
Determination of the sodium content
Atomic absorption spectrometry
IFP 9606
Liquid hydrocarbons
Determination of mercury content
Flameless atomic absorption spectrometry
IFP 9810
Catalytic cracked gas
Analysis of hydrogen, nitrogen, oxygen, carbon oxides, hydrogen
sulphide, ammonia and hydrocarbons
Gas chromatography
IFP 9908
Gas from reforming
Hydrogen and hydrocarbons (small content) analysis
Gas chromatography
IFP 9909
Gas from reforming
Hydrocarbons and hydrogen (small content) analysis
Gas cromatography
310 CA 207
Determination of chloride ions in used catalyst
Potentiometric titration
310 CA 220
Determination of total sulfur and carbon in used catalyst
Infra-red analysis of combustion products
These methods are available upon request, to AXENS customers.
REV. 4
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CATALYST HANDBOOK
9
REV. 4
130/ 172
MISCELLANEOUS
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CATALYST HANDBOOK
9.1
REV. 4
131/ 172
Chemicals specifications
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CATALYST HANDBOOK
132 / 187
COMPARATIVE CHARACTERISTICS OF THE CHLORIDING AGENTS
Name
Formula
Trichloroethylene
CHCl = CCl2
Source : Ins. Nat. Recherche Scientifique
Physical properties
 Aspect
 Detection limit
(ppm)
 Molecular weight
 Melting point
(°C)
 Boiling point
(°C)
 Specific gravity
(D20/4)
 Vapour density
 Vapour pressure (kPa)
(1 atm = 101,3 kPa)
1, 1, 1 - Trichloroethane
CH3 - CCl3
Tetrachloroethylene
CCl2 = CCl2
1, 1, 2, 2 - Tetrachloroethane
CHCl2 - CHCl2
Liquid - colorless - volatile
28
131,4
-87,1
86,7 (atmospheric pressure)
1,47
4,45
7,7 to 20°C
19,6 to 40°C
40,8 to 60°C
Liquid - colorless - volatile
100
133,4
-33
74 (atmospheric pressure)
1,32
4,6
13,3 to 20°C
26,6 to 36°C
53,2 to 54°C
Liquid - colorless - volatile
30
165,8
-22,4
121,2 (atmospheric pressure)
1,62
5,8
5,5 to 40°C
13,9 to 60°C
30,1 to 80°C
58,5 to 100°C
Liquid - colorless
1,5
167.9
-42,5
146,3 (atmospheric pressure)
1 ,60
5,8
0,7 to 20°C
25,2 to 100°C
> 120
(HCl ; Cl2 ; CO ; COCl2)
410
11
41
> 200
(HCl ; Cl2 ; CO ; COCl2)
537
8
10,5
> 140
(HCl ; Cl2 ; CO ; COCl2)
unknown
---
> 120
(HCl ; Cl2 ; CO ; COCl2)
unknown
unknown
unknown
-> 20
-> 100
> 140
--
T°C ambient
--
Fire hazard
 Normal conditions
 Fire case
Not flammable ; not explosive
decomposition : toxic gas
Not flammable ; not explosive
decomposition : toxic gas
Not flammable ; not explosive
decomposition : toxic gas
Not flammable ; not explosive
decomposition : toxic gas
Toxicity
 Average exposit. Value (ppm)
 Limit. Exposit. Value (ppm)
acute
75
200
acute
300
450
acute
50
unknown
Axens recommended
very toxic
1
5
Chemical properties
 Decomposition




(°C)
Autoignition
(°C)
LEL
(% vol.)
UEL
(% vol.)
HCL formation
with water
(°C)
with some metals (°C)
COCl2 = Phosgen (lethal gas).
* carcinogenic
* mortal risk by contact with skin
* not advisable to use
CATALYST HANDBOOK
133 / 187
DMDS:
Dimethyldisulfide is a pale yellow liquid with a foul odor. It contains 68% of sulfur and its
molecular weight is 94.
1.
Commercial specifications
Typical analysis
> 98%
99.5%
 Purity
0.30%
 Methylmercaptan < 1%
<
0.06%
 Moisture
< -5°C/-23°F
< -10°C/-14°F
 Cloud point
2.
General properties
 Freezing point
 Boiling point (1 atmosphere)
d 420CC = d3968CC
 Specific gravity
 Refractive index
 Viscosity
 Surface tension




nD20C = nD68C
(20°C = 68°F)
(20°C = 68°F)
Cryoscopic constant
Polar moment
Dielectric constant
Heat of vaporization
 Heat of combustion
 Flash point (open cup)
 Cubic expansion coefficient
(at 50°C = 122°F)
 Conductivity
 Compressibility coefficient
(at 22.7°C = 72.9°F and 100 bars = 1 450 psi)
 Compressibility coefficient
(at 100°C = 212°F and 100 bars = 1 450 psi)
 LEL
 UEL
 Auto ignition temperature
 Sulfur content
REV. 3
=
=
=
-84.7°C
109.6°C
1.0625
=
1.526
=
=
=
=
=
=
=
=
=
=
=
0.62 cP
33.6 dynes cm-1
0.0062 poundals per inch
0.030C-1
= 0.054 F-1
1.95 D
9.6
9.6 mth mole-1
183.8 BTU/lb
665.8 mth mole-1
12 750 BTU/lb
16°C
= 61°F
0.0011
=
=
=
5.2 10-7 -1 cm-1
13.2 10-7 -1 inche-1
0.45
=
0.70
=
=
=
=
1.1% in air
16.1% in air
> 300°C (> 570°F)
68 wt %
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= -121°F
= 229°F
Catalyst Handbook Rev 4.docx
CATALYST HANDBOOK
134 / 187
Nitrogen :
Purity
99.5 mol% minimum
O2 and noble gases
0.4 mol% maximum
CO + CO2
40 ppm maximum
H2O
20 ppm maximum
Hydrogen:
Electrolytic grade for reduction:
Purity
99.5 mol% minimum
CO + CO2
20 ppm maximum
O2
nil
H2S
nil
Chloride
nil
H2O
5 ppm maximum
Reformer for sulphate removal:
REV. 3
Purity
75 mol% minimum
C2+
2 mol% maximum
CO + CO2
nil
O2
nil
H2S
5 ppm maximum
Chloride
5 ppm maximum
H2O
20 ppm maximum
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9.2
REV. 3
135 / 187
TBP - ASTM Boiling range transformation
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CATALYST HANDBOOK
136 / 187
EDMISTER METHOD FOR
TBP – ASTM BOILING RANGE TRANSFORMATION
FIGURE 9-1
Amplitude of ASTM curve segments, °C
Amplitude of TBP curve segments, °C
ASTM – TBP CORRELATION
Segments of the
distillation curve, % vol.
ASTM 50% point, °C
REV. 3
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CATALYST HANDBOOK
137 / 187
Correlation TBP – ASTM
Often the ASTM D86 distillation is known, not the TBP.
Curve can be used to obtain from the ASTM D86 the TBP distillation data.
For a reformer feed having the following ASTM D86:
IBP
= 82°C
10% = 99°C
30% = 110°C
50% = 125°C
70% = 143°C
90% = 166°C
EBP = 185°C
The TBP curve can be obtained as follows:
1.
2.
3.
4.
5.
6.
7.
ASTM 50%
ASTM (30% - 50%)
ASTM (10% - 30%)
ASTM (IBP - 10%)
ASTM (70% - 50%)
ASTM (90% - 70%)
ASTM (EBP - 90%)
=
=
=
=
=
=
=
125°C
- 15°C
- 11°C
- 17°C
18°C
23°C
19°C
Curve







TBP 50%
TBP 50%
= ASTM + 1°C
= 125°C + 1°C
=
126°C
TBP (30% - 50%)
TBP 30%
= 26°C
= 126°C - 26°C
=
100°C
TBP (10% - 30%)
TBP 10%
= 22°C
= 100°C - 22°C
=
78°C
TBP (IBP - 10%)
TBP IBP
= 32°C
= 78°C - 32°C
=
46°C
TBP (70% - 50%)
TBP 70%
= 26°C
= 126°C + 26°C
=
152°C
TBP (90% - 70%)
TBP 90%
= 31°C
= 152°C + 31°C
=
183°C
TBP (EBP - 90%)
TBP EBP
= 21°C
= 183°C + 21°C
=
204°C
The TBP corresponding to the above ASTM D86 is:
REV. 3
IBP
=
46°C
10%
=
78°C
30%
=
100°C
50%
=
126°C
70%
=
152°C
90%
=
183°C
EBP
=
204°C
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CATALYST HANDBOOK
9.3
REV. 3
138 / 187
Reformate RVP versus butane content
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REFORMATE RVP VERSUS BUTANE CONTENT
FIGURE 9-2
REV. 3
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10 TECHNICAL ASSISTANCE SERVICES FOR FIXED BED
UNIT
REV. 3
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CATALYST HANDBOOK
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Technical Assistance
Axens compared to some of our competitors is not only a well-known catalyst
manufacturer, but also a process design company. These two aspects allow us to provide
advices, indeed it is straightforward that the operating procedures, current operations,
equipment and catalysts are interrelated.
Based on this double knowledge Axens offers a wide range of technical assistance for
both fixed bed unit and continuous unit.
Here below you will have an overview of the various areas in which Axens can provides
an expertise.
10.1 Catalyst performance estimation on site
Axens offers proprietary software in order to estimate the predicted unit performance.
This software is easy to use as the interface looks like windows one.
This allows you to compare the actual versus the predicted performance and to check if
there are any discrepancies.
10.2 Unit follow up at Axens offices
Based on large experiences both for fixed bed unit and continuous units Axens can
provide assistance and advices by analysis of operation data. Our technical assistance
department will review these data and make comments and remarks
Both aspects of the unit follow up, on site and at Axens offices’, help us in order to
reduce the time needed for corrective actions.
10.3 Catalyst analysis
During the life of the catalyst, Axens can perform chemical and physical analysis. For
example, we can check the presence of poison, sulfur, coke, and chloride on the catalyst.
These analysis allow us to give you advises for the catalyst replacement and better
operations.
Moreover if we suspect any difficulties during oxychlorination and reduction steps we
can measure the metals dispersion.
Regarding the physical analysis we, mainly, can control the surface area and the
mechanical resistance of the catalyst. The last point is a major issue for continuous unit,
as a low mechanical resistance will increase the operating cost of the unit.
Moreover, the catalyst sampling device recommended by Axens can be installed on
existing reactors to collect catalyst samples during operation. As a consequence, for fixed
bed unit, we are not obliged to wait for regeneration of the catalyst in order to have
samples.
Finally Axens can also performed analysis of the feed and the reformat and octane
measurements.
REV. 3
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CATALYST HANDBOOK
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10.4 Follow up of the regeneration and the start up
During regeneration and unit start up, we supply on-site assistance from a technically
trained start up engineer.
This advisor will assist and review the operating procedure.
Moreover, Axens recommends, when the reactors are reloaded to review internals to
check the gas ditribution.
10.5 Catalyst optimisation performance
During the life of the catalyst, Axens will be available to discuss the optimisation of the
catalyst performance. Indeed thank to your research and industrial experience we can
propose improvements taking into account any modifications in the operating
conditions.
For example for radial fixed bed units we can propose Texicap™ in order to improve the
gas distribution and to make the loading easier and shorter. For cycle length
improvement we can propose the dense loading Catapac™.
10.6 Operator training
The starting point for your operational staff should begin with theoretical and
operational classroom training, process simulation and on-site visits of units.
Customer tailored training courses are conceived, prepared and jointly executed by
Axens and the IFP School, an international center for graduate level training for
petroleum industry.
Training programs can be carried out either at an Axens site or at a specified refinery. We
are committed to offer additional training sessions to your staff as required, during the
operational lifetime of your units.
10.7 Training simulators
Axens’ Performance Programs Business Unit conceives generic version training
simulators for the operational staff. For further information, please consult our web site
www.axens.net.
REV. 3
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CATALYST HANDBOOK
143 / 187
10.8 Catapac – Texicap – Catalyst sampler
REV. 3
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Catalyst Handbook Rev 4.docx
CATAPAC
ULTRA HIGH DENSITY LOADING SYSTEM
FOR UP TO 25% MORE REACTOR OR
ADSORBER CAPACITY
A KEY TO SUCCESSFUL CATALYSIS OR
ADSORPTION: SURFACE AREA
Whether it be catalysis or adsorption, an
important key to successful processing is the
surface available for mass transfer. By
increasing the amount of surface area in a
given volume, bed performance increases
accordingly. This can be done by increasing the
vessel volume; however, because vessels are
expensive, good design requires the vessel to
be as small as practical. There are other ways.
One can change various aspects of the
material; its size, shape, structure or
composition or, using the very cost-effective
Catapac, you can put more material in the
same volume.
Schematic of a bed loaded conventionally
A CHANGE FROM TRADITIONAL WAYS
Traditional methods for loading catalyst or
adsorbents in a vessel involve things like
“socks” or chutes, raking (leveling the bed) or
just plain dumping. The Catapac system breaks
with the old ways and puts up to 25% more
material in the same volume. This is like getting
five vessels for the price of four, or, for existing
installations, getting 25% better performance
for a tiny fraction of the cost of new
equipment.
Moreover, Catapac performs the operation
more gently, safely and quickly than
conventional loading methods.
Schematic of a “Catapacked” bed
Axens Process Licensing
TM
As shown in the preceding illustrations, loading
with Catapac distributes the particles uniformly
throughout the bed. Large voids are eliminated
and flow is distributed in smaller, more
homogeneous channels.
As Catapacking
reduces void volumes considerably, operators
will notice that bed settling and the problems it
causes disappear. Bed performance improves.
EFFECTIVE FOR A VARIETY OF PARTICLES AND
BED SIZES
Catapac works on all kinds of fixed beds, radial
as well as axial flow. It handles all particle sizes
and shapes that are commonly available in
extrudates, beads, and even grain. Beds ranging
in diameter from 50 cm to 7 meters are loaded
with ease.
The gain in bed density depends on particle
shape and homogeneity, but even spherical
shapes or beads show packing improvement.
The bed density improvement is largely due to
Catapac’s gentle, even distribution and
orientation action.
REDUCED BREAKAGE AND FINES
Other loading methods create fines because the
particles are subjected to the crude mechanical
actions that raking or dumping involve. Broken
catalyst or adsorbent fines are undesirable
because they cause pressure drop to build up
and reduce bed performance due to channeling.
Catapac’s technique minimizes breakage and
fines production because the particles are
gently sprinkled into place from a short
distance.
Walking on the bed, or leveling it using a rake
or board are expressly avoided because that
would degrade bed performance.
FAST, SAFE, EASY TO USE
Loading is quick thanks to Catapac’s
lightweight, high flow loader and because there
is no one inside to worry about. Vessel entry
has often led to safety or health incidents;
Catapac loading operations are conducted
outside the vessel. No electrical set-up other
than lighting is needed.
To check loading progress, the Catapac loader is
simply stopped and restarted after the check.
No catalyst mound builds up. Thus, significant
maintenance cost savings are obtained while
increasing operating productivity.
AXENS - YOUR REFERENCE FOR REVAMPS OR
NEW INSTALLATIONS
Optimizing reactors and adsorbers for top
performance has long been an Axens strength.
Installed in all dense-loaded reactors and
adsorbers designed by Axens, Catapac is a good
example of our technology at work with over
300 loading operations and 7500 tons loaded in
recent years.
Catalysts and adsorbents are other areas where
Axens excels. Through them, you have the best
available catalysts and adsorbents for your new
or existing plant.
Axens maintains its enviable position in
technology services by constantly creating,
improving and testing its catalysts, processes
and equipment in what is one of the world’s
most comprehensive R&D centers.
011CATA
Axens Process Licensing
BETTER FLOW DISTRIBUTION
Process Licensing
89 bd Franklin Roosevelt, BP 50802
92508 Rueil-Malmaison Cedex – France
Tel: + 33 1 47 14 21 00
Fax: + 33 1 47 14 24 98
www.axens.net
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650 College Road East, Suite 1200
Princeton, NJ 08540 - USA
Tel: 1 609 243 8700
Fax: 1 609 987 0204
www.axens.net
Axens Process Licensing
TEXICAP
TM
15% MORE PERFORMANCE FROM
FIXED-BED RADIAL-FLOW REACTORS
INCREASED PERFORMANCE OVER
CONVENTIONAL RADIAL BED DESIGN
Fixed-bed radial-flow reactors are used when
low pressure drop is critical to performance,
such as in reforming units. Most fixed-bed
radial reactor designs “waste” the top 15% or
so of the catalyst bed.
Compared with original flow conditions,
reactor pressure drop will be lower, now that
the reactants flow through a larger catalyst
bed. Or, viewed differently, the reactor can
accept 15% more feed.
This is done intentionally to prevent reactants
from short-circuiting the catalyst after the bed
settles. When the catalyst settles, a gap
appears between the top of the bed and cover
plate. Were it not for the shroud extending
down into the bed, some reactants would pass
preferentially through the gap, short-circuiting
the catalyst.
This practice does not permit the full
utilization of the expensive reforming catalyst
(containing platinum) in the reactor.
Responding to the problem, Axens developed
and patented an extremely cost-effective way
to recover unused bed volume.
As shown in the figure, the conventional metal
shroud assembly is replaced by Texicap, a
flexible flowguide that moulds to the shape of
the top of the bed, regardless of the amount of
the settling. There is no further need to have
dead space and the formerly dormant catalyst
section can now go to work.
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Fax: + 33 1 47 14 24 98
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650 College Road East, Suite 1200
Princeton, NJ 08540 - USA
Tel: 1 609 243 8700
Fax: 1 609 987 0204
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OTHER REVAMP CONSIDERATIONS
The flowguide is an engineered composite of
refractory fibers and fillers. It is strong and
impermeable; furthermore, it contains no
asbestos.
Texicap has been proven industrially with great
customer satisfaction. It withstands a
reformer's severe regeneration conditions as
well as its hydrogen and hydrocarbon
atmospheres. The first application, installed
years ago, is still as good as new.
Significant low cost improvements carried out
inside existing radial reactors are both
possible and easily implemented. In fact, a
combination of Axens technologies makes it
possible to add as much as 40 percent to a
reactor’s performance – at a fraction of the
reactor replacement cost. A smart first step
would be to ensure that the reactor is loaded
with catalyst using the Catapac dense loading
system.
EASE OF INSTALLATION
CATAPAC CATALYST LOADING
The maintenance department will like Texicap
when they find out that installing or removing
it takes about a tenth of the manpower and
time it took to remove and replace the shroud
and cover assembly. Texicap eliminates the
problems associated with metal shrouds, such
as seized or broken bolts, heavy metal panels
to coax into place and long hours spent in
cramped quarters. Manipulating Texicap
couldn't be easier; one man can usually do it in
an hour.
Payout is incredibly fast.
Depending on the catalyst shape, you can put
significantly more catalyst into a reactor bed
compared with the old “sock” loading
procedure. For radial and axial flow beds
alike, the catalyst is placed uniformly and
gently, reducing potential channeling as well.
It works on all kinds of materials and shapes
from cereal grain to catalysts; it's very
effective for extrudates (up to 25% higher bed
density), but good results have also been
obtained on beads. Your safety and
maintenance personnel will be pleased to
know that there's no need for anyone to be
inside the reactor during loading.
REFERENCES
There are over fifty Texicap installations todate. Customer feedback has been excellent.
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TOUGH AND DURABLE FLOWGUIDE
Process Licensing
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Princeton, NJ 08540 - USA
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Fax: 1 609 987 0204
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Axens Process Licensing
CATALYST
SAMPLING
SYSTEM
ON-LINE SYSTEM WITHSTANDS THE SEVERE
OPERATING CONDITIONS OF REFORMER
UNITS
SPECIFICALLY DESIGNED FOR REFORMER FIXED
BED REACTORS AND CCR REGENERATORS
The equipment includes the following items:
The rugged Axens sampling system has proven its mettle
in the harsh conditions encountered in catalytic
reformers. Getting catalyst samples without interfering
in the operation is mandatory for good reformer
performance providing plant personnel with vital
operating information.




Sampling head that takes uniform, constant
volume catalyst samples
Pipe that allows removal of the catalyst sample
from the reactor. The body length is adjustable to
attain the optimum head position in the catalyst
bed.
Drive system that controls the head rotation
Sample receiving system equipped with Axens'
special-design isolation valves.
DESCRIPTION
Special design considerations were taken to ensure that
the system operates on demand in the aggressive
reforming environment that includes hydrogen and
hydrocarbon at temperatures upwards of 500°C.
Process Licensing
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Tel: + 33 1 47 14 21 00
Fax: + 33 1 47 14 24 98
www.axens.net
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Tel: 1 609 243 8700
Fax: 1 609 987 0204
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Axens Process Licensing
USAGE IN REFORMING UNITS
SAMPLING HEAD
During normal operations, the catalyst sampling system
is used to control catalyst chlorine levels, to schedule
regenerations, to analyze catalyst carbon content, to
determine the nature of any catalyst contamination or
poisoning that may have occurred, and to detect
abnormal carbon deposition.
The head shape and clearance as well as its surface
hardness due to a special surface treatment have been
developed by Axens to withstand severe operating
conditions. Moreover, each system is supplied with a
head test certificate issued by Axens which ensures its
reliability.
During regeneration, the sampling equipment is used
during the oxychlorination stage to control the catalyst
chlorine content.
SAMPLING PROCEDURE
In the standby position, the sampling head in the
reactor is empty. During the sampling process, the
following procedure is applied:
 The receiving pot is
purged.
 Under hydrogen and
hydrocarbon
atmosphere,
the
receiving pot is opened
to the reactor.
 Sampling head is filled
with catalyst using the
sampler drive system.
 The catalyst flows to the
receiving pot.
 The receiving pot is
isolated by valves.
 The catalyst is purged
and the drain valve is
opened.
The catalyst sample is
available for analysis and
the operation can be
repeated as needed.
Slide Valve Installation
INSTALLING A CATALYST SAMPLER ON AN
EXISTING REACTOR
If the reactor has not been previously fitted with an
appropriate sampling system nozzle, the sampler can
be passed through the catalyst unloading nozzle. To
unload the catalyst bed, the sampling system can be
removed by means of a slide-valve that Axens can
provide for this purpose. This valve also is used to
control catalyst flow during unloading.
To-date, more than eighty catalyst sampling systems
are operating around the world. The sampler is
available under license from Vinci Technologies.
Process Licensing
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92508 Rueil-Malmaison Cedex – France
Tel: + 33 1 47 14 21 00
Fax: + 33 1 47 14 24 98
www.axens.net
IFP North America, Inc.
650 College Road East, Suite 1200
Princeton, NJ 08540 - USA
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Fax: 1 609 987 0204
www.axens.net
CATALYST HANDBOOK
151 / 166
11 TYPE OF REACTORS – CATALYST LOADING
REV. 3
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11.1
Type of reactors
A
Radial reactor - figures 11-1 / 11-2
152 / 166
The loading philosophy for a radial reactor proves to be somewhat more complex than
for a downflow one. The major concern for a radial reactor is partial bypassing of the
catalyst. To prevent this, allowance must be made for the settling that will occur during
the run, called the slump allowance. Further, the catalyst must be loaded to a level such
that regardless of which path the feed will take to the center-pipe, it will always traverse
a minimum of 1 bed thickness. This is called the seal allowance. The slump allowance is
determined by taking 5 percent of the bed height. Therefore, the total catalyst bed
depth equals (1.05 x basic catalyst depth) + seal allowance. The seal allowance usually
provides for one half a radial bed thickness above the top row of holes in the centerpipe.
Note also that if there is an alumina balls support under the catalyst bed a seal allowance
of one half the radial bed thickness shall be provided from the bottom of the scallops to
the upper level of alumina balls support. During loading, care should be taken not to get
any inerts or catalyst in the scallops. This will accumulate at the bottom of the scallops
and cause maldistribution of flow.
Note that if the catalyst is loaded with a dense loading apparatus, as the arrangement of
catalyst extrudates within the reactor is horizontal leading to an almost maximum
loading density, there will be a little bed settling. Slump allowance can be reduced to 2%
of the bed height.
Note also that to increase the volume of catalyst for a given radial fixed bed reactor,
AXENS has patented a device: a flexible flow guide which allows to use the normally
wasted seal catalyst necessary on top of the catalyst bed of this type of reactor (See
figure 11-3).
B
Cylindrical down flow reactor - figure 11-4
Of the three reactor types commonly employed in reformer service, the downflow (or
axial) is the least complicated to load. For an existing system, the available loading
volume is determined, and catalyst is loaded to the required level. A minimum 100 mm
layer of 20 mm alumina balls should be allowed on top of the catalyst. A minimum 450
mm between the top of the inert layer and the bottom of the inlet distributor should be
allowed.
C
Spherical downflow reactor - figure 11-5
The difficulty with loading a spherical reactor turns out to be a practical consideration. It
is very difficult to level the bed due to the large diameters typically encountered,
especially at the level of the great circle of the sphere. Since it is usually desirable to load
symmetrically above and below the great circle, the spherical reactor requires a lot of 
20 mm alumina balls at the bottom of the reactor.
REV. 3
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RADIAL REACTOR
CATALYST LOADING DIAGRAM
FIGURE 11-1
REV. 3
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RADIAL REACTOR
CATALYST LOADING DIAGRAM
FIGURE 11-2
REV. 3
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RADIAL REACTOR
USING FLEXIBLE FLOW GUIDE
(TEXICAP)
FIGURE 11-3
REV. 3
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CYLINDRICAL DOWNFLOW REACTOR
CATALYST LOADING DIAGRAM
FIGURE 11-4
REV. 3
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SPHERICAL DOWFLOW REACTOR
CATALYST LOADING DIAGRAM
FIGURE 11-5
REV. 3
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11.2
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Catalyst loading procedure
Preliminary
When the reaction section drying out is complete and nitrogen has been replaced by air,
the reactors can be opened for inspection of the internals.
If needed, the reactors will be brushed and vacuum cleaned before catalyst loading
starts.
The catalyst loading period must be as short as possible to minimize the risk of moisture
entering the reaction system. If necessary, the job will be done in shifts.
Catalyst loading will be interrupted in case of rain or snow, or efficient protection must
be set in place.
Catalyst is delivered either in steel drums or cloth big bags. The catalyst must be handled
with care to avoid breakage of the extrudates. During the loading, the loading sleeves
will be adjusted to allow a free fall no higher than 1 meter above the catalyst level.
Catalyst drums must not be rolled. Big bag shall be handled very carefully.
Catalyst loading period by itself lasts between 2 and 5 days according to the amount of
catalyst to be loaded, the number of reactors and their type, the work force and last but
not least the weather. This figure does not take into account neither the reactors
opening nor their boxing-up, but includes internals installation. In this period of time
there is no provision for internals repair (scallops, center pipe, etc..).
Note also that dense loading does not make the overall loading any faster or longer. On
the contrary use of Texicap instead of conventional steel covering system can save pup to
20% of the time.
Equipment and personnel
The list below is based on the following assumption: catalyst loading will be carried out
with 1 crane.
• 1 telescopic crane capable of lifting approximatively 3 tons, 5 meters above the
reactor upper manhole (See figure 11-6).
• 1 forklift to handle the catalyst drum pallets.
• Stationary hopper equipped with Ø 8 to 10" nozzles to be installed on the
reactor upper manhole (See figure 11-11).
• 2 mobile hoppers, each will a capacity of about 5 to 6 catalyst drums (See figure
11-10).
• 1 structure (scaffolding and timbers) to unload the drums into the mobile
hoppers (See figure 11-8).
• 1 safety harness, rope ladders, portable oxygen analyzer, low voltage lighting,
dust masks, air masks, goggles, plastic sheets to protect drums and reactor in
case of rain.
REV. 3
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To serve the above mentionned equipment the following personnel is necessary:
• 1 foreman,
• 1 crane operator,
• 1 fork lift operator,
• 1 team of 4 riggers at ground level for catalyst handling and loading into the
mobile hoppers.
• 1 team of 6 fitters for vessels opening closing and for catalyst loading.
Note: This personnel shall be permanently present during the loading operation. The
total number of personnel will be obtained by multiplying by the number of shifts,
if the loading takes place on 24 hours basis.
Catalyst packaging, handling and storage
Catalyst can be delivered either in drums or in big bags. Catalyst must be handled with
care to avoid breakage of the catalyst extrudates. During the loading, the loading sleeve
shall be adjusted to allow a free fall no higher than 1 meter, above the catalyst level.
Catalyst drums must not be rolled and big bags must be carefully handled to avoid
catalyst attrition.
Catalyst delivered in sealed still drums shall be stored in an enclosed and ventilated
warehouse. Storage time under such a condition is two years without any damage for
the catalyst and for the packaging.
Outdoor storage is possible, provided original packaging is stored under water-tight
cover and raised above the naturally flood-free ground: Such storage is not
recommended and should be minimized to reduce as much as possible the risk of water
damage.
Catalyst delivered in big bag shall normally be loaded in the reactor as soon as it arrives
on site. Short duration storage can be envisaged provided, it is in a covered and dry
warehouse.
PACKAGING
DRUM
Description
Steel drum (UN standard). Air tight cover.
Thickness: 0.8 mm Tare: 14.5 kg
Dimensions
Diameter: 0.6 m – Height: 0.9 m
Standard capacity: 217 liters.
Netweight
105 kg
BIG BAG
Description
REV. 3
External skin: fiber glass spliced propylene
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Internal skin: 80 micron polyethylene
REV. 3
Dimensions
Square section 80 x 89 cm.
Cylinder diameter: 110 cm.
Height: empty 160 cm, full: 120 to 150 cm.
Netweight
700 kg
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Typical catalyst loading report
REACTOR LOADING REPORT
PLANT TYPE:
CLIENT & JOB NUMBER:
LOCATION:
UNIT NUMBER:
REACTOR:
DATE:
SHEET OF
CATALYST TYPE:
Drum
Lot
N°
Drum
N°
Net
wt
1
2
3
4
5
6
7
8
9
10
11
12
13
14
15
16
17
Drum
Lot
N°
Drum
N°
Net
wt
18
19
20
21
22
23
24
25
26
27
28
29
30
31
32
33
34
Drum
Lot
N°
Drum
N°
Net
wt
35
36
37
38
39
40
41
42
43
44
45
46
47
48
49
50
51
TOTAL WEIGHT LOADED:
LOADING SUPERVISOR NAME:
NUMBER OF DRUMS:
SIGNATURE:
LOADING DENSITY:
Note: Bar code plastic label attached to the seal securing the drum closure handle shall
be kept for reference.
REV. 3
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CATALYST HANDLING WITH DRUMS
FIGURE 11-6
REV. 3
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FILLING-UP THE HOPPER AT GROUND LEVEL FROM DRUMS
FIGURE 11-7
REV. 3
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CATALYST HANDLING WITH BIG BAG
FIGURE 11-8
REV. 3
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GROUND LEVEL PREPARATION FOR BIG BAG LIFTING
FIGURE 11-9
REV. 3
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CATALYST LOADING - MOBILE HOPPER
FIGURE 11-10
REV. 3
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CATALYST LOADING - STATIONARY HOPPER
FIGURE 11-11
REV. 3
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CATALYST LOADING - SLIDE VALVE
FIGURE 11-12
Note: As to avoid the sleeve full of catalyst falling accidentally on the worker standing
on the catalyst bed, in addition to the clamp the sleeve shall be secured by a rope.
REV. 3
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SOCK LOADING AXIAL REACTOR
(UNDER AIR)
FIGURE 11-13
REV. 3
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SOCK LOADING RADIAL REACTOR
FIGURE 11-14
REV. 3
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CATAPAC DENSE LOADING
FIGURE 11-15
REV. 3
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