Note : Les descriptions sont présentées dans la langue officielle dans laquelle elles ont été soumises.
` 1326464
1 --
K 9792
HEAVY OIL CRACKING PROCESS
'~
The invention relates to a new process for the
conversion of a heavy oil fraction, especially a heavy
oil fraction containing a limited amount of asphaltenic
constituents, into lighter components.
In the refining process of crude oil to final
products, heavy fractions, e.g. fractions boiling
between 370-520 C, are usually processed in cracking
processes such as fluidized catalytic cracking,
hydrocracking and thermal cracking, in order to convert
these high boiling fractions into more valuable lighter
fractions.
At the present moment there is a growing demand
for middle distillates, i.e. kerosene and gas oil,
especially high quality middle distillates. Kerosene
usually has a boiling point between about 150 and about
270 C and is mainly used for jet fuel. A major quality
parameter for kerosene and related to the burning
properties thereof is the smoke point. Gas oil usually
has a boiling point between about 250 and about 370 C
and is mainly used as fuel for compression-ignition
engines. Important quality parameters comprise its
ignition quality as expressed by the cetane number and
its cold flow properties as expressed by the cloud
point.
As indicated above three main cracking processes
are used in oil refining.
Fluidized catalytic cracking is usually performed
at a relatively low pressure (1.5 to 3 bar), and at
relatively high temperatures (480-600 C) in the
presence of an acidic catalyst (for instance zeolite
containing catalysts). The reaction is carried out in
.. - ~ . .
.
1326~6~
-- 2
the absenoe of hydrsgen and the residence time of the
feed is very short (0.1-10 seconds). During the
reaction a large amount of carbonaceous materials
(coke) is deposited onto the catalyst (3 to 8 %w of the
feed). Continuous regeneration of the catalyst by
burning-off coke is therefore necessary. The products
obtained in this process contain relatively large
quantities of olefi~s, iso-paraffins and aromatics
boiling in the gasoline range. Thus, a major product
obtained by fluidized catalytic cracking is a gasoline
component of good quality. Further, light cycle oils
boiling in the kerosene range and some heavy cycle oils
boiling in the gas oil range and above are obtained,
both of a moderate to low quality for use as kerosene
and gas oil.
Hydrocracking is usually performed at a relatively
high hydrogen pressure (usually 100-140 bar) and a
relatively low temperature (usually 300 to 400 C). The
catalyst used in this reaction has a dual function:
acid catalyzed cracking of the hydrocarbon molecules
and activation of the hydrogen and hydrogenation. A
long reaction time is used (usually ~.3 to 2 l/l/h
liquid hourly space velocity). Due to the high hydrogen
pressure only small amounts of coke are deposited on
the catalyst which makes it possible to use the
catalyst for 0.5 to 2 years in a fixed bed operation
without regeneration. The products obtained in this
process are dependent on the mode of operation. In one
mode of operation, predominantly naphtha and lighter
products are obtained. The naphtha fraction contains
paraffins with a high iso/normal ratio, making it a
valuable gasoline blending component. In a mode
directed to heavier products, kerosene and ga~ oil are
mainly obtained. In spite of the extensive
hydrogenation, the quality of these products is
1326~64
moderate only, due to the presence of remaining
aromatics together with an undesired high iso/normal
ratio of the paraffins amongst others.
Thermal cracking is usually performed at a
5 - relatively low or moderate pressure (usually 5 to 30
bar) and at a relatively high temperature (420-520 C)
without catalysts and in the absence of hydrogen. A
long reaction time is used (residence time normally
2-60 minutes). The middle distillates obtained from
thermal cracking of high boiling distillates are of
good quality as far as the ignition properties are
concerned. The high content of olefins and heteroatoms
~especially sulphur and nitrogen), however, requires a
hydrofinishing treatment. A major problem in thermal
cracking, however, is the occurrence of condensation
reactions which lead to the forming of polyaromatics.
The cracked residue from thermal cracking, therefore,
is of a low quality (high viscosity and high carbon
residue after evaporation and pyrolysis, expressed for
instance by its Conradson Carbon Residue (CCR)
content).
It is known from U.S. patent specification
4,017,380 to subject a residual oil to a
hydrodesulphurization treatment and to use the thus
deactivated hydrodesulphurization catalyst as a fixed
or packed (non-fluid) bed of catalytically inert and
non porous solids in a hydrovisbreaking process, which
process has to be carried out in upward flow. It is
stated categorically in said U.S. patent specification
that the use of a hydrotreating catalyst in down flow
operation under visbreaking conditions would only tend
towards undesired aftercracking without increasing the
yield of the desired middle distillate product.
A new cracking process has now been found which is
especially suitable for the conversion of heavy oil
132S464
; - 4 -
fractions containing a low amount of asphaltenic
constituents into middle distillates of good guality
boiling in the range of 150-370 C, i.e. kerosene and
gas oil. The new process, hydrocatalytic thermal
cracking, (HCTC), is performed at a relatively high
temperature (400-550 c)~ It i8 carried out under a
moderate hydrogen pressure (lo to 60 bar) in the
presence of a non-acidic, hydrogen activating catalyst.
Notwithstanding the relatively low hydrogen
pressure the process of the present invention shows an
extremely low rate of coke formation. Undisturbed
operation in a fixed bed reactor can be readily
achieved for a period of at least 1000 hours. Depending
on the specific conditions applied, even substantial
longer operation times are possible. In this respect it
is remarked that lowering the hydrogen pressure in a
conventional hydrocracking process immediately would
lead to deactivation of the catalyst by basic nitrogen
and carbonaceous deposits, thus limiting the run
length.
The present invention thus relates to a process
for the conversion of a heavy oil fraction into lighter
frar-tions, comprising passing a heavy oil fraction
having a low content of asphaltenic constituents
; 25 together with a hydrogen containing gas stream through
a reaction zone containing a non-acidic, hydrogen
activating catalyst at a temperature of 400-550 C,
preferably 410-530 C, more preferably 440-510 C, and
a hydrogen partial pressure of 10-60 bar, preferably
20-40 bar.
The molecular weight reduction is essentially
effected by thermal cracking of feedstock molecules.
Thus, in contrast with catalytic cracking and
hydrocracking, the novel process does not depend on the
pres~nce of acidic sites on the catalyst, which should
~3264~4
remain active during the cracking cycle or life of the
catalyst. Due to the presence of hydrogen even at
relatively moderate pressure, only very small amounts
of coke are deposited on the catalyst, thus making it
possible to operate in a fixed bed mode (e.g. swing
reactor) or a moving bed mode ~e.g. bunker flow
reactor).
The middle distillates obtained are of good
quality due to the high amount of n-paraffins and the
low amount of olefins, in ~pite of the presence of a
certain amount of aromatic compounds. The hydrogen
consumption of the process is relatively low, as the
aromatic compounds are hardly hydrogenated. A further
advantage is the fact that, dependent on the catalyst,
the ~ulphur present in the feed can be converted for a
6ubstantial part into hydrogen sulphide, thus resulting
in a product, containing a relatively small amount of
sulphur.
The bottom material, i.e. material boiling above
the boiling point of the middle distillate products,
has excellent properties (viscosity, carbon residue and
sulphur content) and can be used as a valuable fuel oil
blending component. Further, said heavy material is
unexpectedly an excellent feedstock for a fluidized
catalytic cracking reaction. When compared with a usual
feedstock for a fluidized catalytic cracking reactor,
for example a straight run flashed distillate, the
gasoline yield and quality are similar. When compared
with the bottom material obtained from a distillate
thermal cracking reactor as feedstock for a fluidized
catalytic cracking process a much higher ga~oline yield
is obtained.
When compared with a usual thermal cracking
proces~ a comparable middle distillate product is
obtained, provided ~hat the thermal cracking product is
1326~64
subjected to an additional hydrofinishing treatment.
The quality of the unconverted fraction of the produc'c
obtained by the new process, however, is much better
than the quality of the unconverted fraction of thermal
cracking. Due to the presence of activated hydrogen
during the reaction the heavy fraction resulting from
the present process has a low viscosity, a low content
of polyaromatic compounds and a low sulph~r content.
When compared with a usual catalytic cracking
process the HCTC-process does not depend on the
presence of acidic sites on the catalyst. The HCTC
process can be suitably carried out in the substantial
or even complete absence of acidic sites in the
catalyst. Thus, feeds containing a substantial amount
of basic nitrogen and/or sulphur containing compounds
can be processed without difficulties. Due to the
presence of activated h~drogen only very small amounts
of coke are deposited on the catalyst, while in
fluidized catalytic cracking large amounts of coke are
deposited on the catalyst, making continuous
regeneration of the catalyst necessary. The products
obtained by the present process are predominantly
middle distillates of good quality together with a
heavy, unconverted fraction of relatively good quality.
The major product obtained by fluid catalytic cracking
is a gasoline hlending component together with a
smaller amount of light cycle oil of moderate to low
quality as aromatic csmpounds form the larger part of
this light fraction. During the process according to
the present invention hardly any hydrogen transfer
~ reactions, resulting in the formation of (poly)aromatic
compounds and paraffins from naphthenes and olefins,
occur.
With regard to the usual hydrocracking process the
process according to the present invention does not
1326~64
,~ ,
-- 7 --
depend on the presence of acidic sites on the catalyst.
Therefore, HCTC is relatively insensitive to feedstock
impurities, especially (basic) nitrogen and carbon
residue (CCR). As the process according to the present
5 invention can be carried out during a substantial
period at relatively low hydrogen pressure investment
costs are considerably lower when compared with a
conventional hydrocracking process. The hydrogen
- consumption in the HCTC-process is relatively low. With
regard to the iso/normal ratio of the paraffins it may
be remarked that due to the radical type of cracking in
the HCTC-process the iso/normal ratio of the paraffins
is low, which is favourably for the ignition quality of
the gas oil. The classic hydrocracking process results
in a high iso/normal ratio due to the carbonium ion
reaction mechanism, thus unfavourably affecting the
- quality, of the middle distillates, especially the
ignition quality of the gas oil.
A suitable feed for the HCTC-process according to
the present invention is a heavy oil fraction having a
low content of asphaltenic constituents. Vacuum
distillates, and/or deasphalted oils of any source and
almost limitless as far as the sulphur and nitrogen
content is concerned can be used. Suitably the content
of asphaltenic constituents in the feed is less than
3~w, preferably less than 2%w, more preferably less
than 1.5%w, and most preferably less than 1.0%w. Under
the asphaltenic constituents mentioned hereinbefore
"C7-asphaltenes" are meant, i.e. the asphaltenic
fraction removed from the heavy oil fraction by
precipitation with heptane. The feed may contain a
substantial amount of carbon residue tCCR), ~uitably
below 15%w, preferably below 10%w, more preferably
below 6%w. The amount of sulphur in the feed is
suitably below 10%w, preferably below 6%w, more
. ., .~
1326464
preferably below 4%w. The amount of nitrogen iOE
suitably below 6%w, preferably below 4~w.
Very suitably a vacuum di~tillate or flashed
distillate can be used as feed having a boiling range
substantially between 350 and 580 C, preferably
between 370 and 520 C. Another very suitable feed is a
deasphaltized residual oil (DA0), for instance a
propane, butane or pentane deasphalted long or short
residue.
Also synthetic distillates and/or synthetic
deasphalted oils, which are available in for instance
complex refineries, are suitable feeds for the present
process. A very suitable source for producing such
synthetic feeds comprises the so-called hydroconversion
process of residual oil fractions, for instance short
residue. Such a hydroconversion process preferably
comprises a hydrodemetallization step, followed by a
hydrodesulphurization/hydrodenitrogenation step and/or
a hydrocracking step. It is remarked that usually
synthetic flashed distillates or synthetic deasphalted
oils are processed in a catalytic cracking process.
However, this results mainly in the production of
gasoline but no kerosene or gas oil of acceptable
guality is obtained. Conventional hydrocracking of such
feeds i5 hardly possible due to the very refractory
nature of the nitrogen compounds present and the need
for low nitrogen-feeds in the hydrocracking process.
Hydrogen conver~ion processes such AS H-oil, LC-fining
and Residfining can also be used for the production of
the above-indicated synthetic feeds.
Another very suitable feed for the HCTC-proce~s
originates from a visbreaking process of for instance
short residue. Upon thermally cracking a heavy residue
followed by flashing or distillation of the product, a
distillate can be obtained substantially boiling in the
`
1326464
`: g
range between 350 and 520 C which i6 an excellent
feedstock for the process according to the present
invention.
Mixtures of relatively heavy and relatively light
feedstocks, e.g. a DA0 and a flashed distillate, may be
used advantageously in view of reduced coke formation.
The hydrocatalytic thermal cracking process is
suitably carried out at a reaction temperature of
400-550 C, preferably 410-530 C, more preferably
` lO between 440-510 C, most preferably at about 450 C. It
will be appreciated that a higher conversion will be
obtained when the temperature is higher, as the rate of
thermal cracking of hydrocarbons will be faster at
higher temperatures. To obtain the same conversion rate
a (slightly~ higher temperature should be used for a
feedstock which is more difficult to crack thermally,
for instance a feedstock rich in cyclic compounds, than
for a feedstock which cracks more easily.
The space velocity of the feed in the novel HCTC
process is suitably chosen between 0.1 to 10 l/l/h,
preferably between 0.5 to 6 l/l/h, more preferably
between l.0 to 5 l/l/h.
; The hydrogen partial pressure under which the
HCTC-process is carried out suitably lies between 10-60
bar, preferably 20-40 bar, more preferably about 25
bar. The total pressure in the reactor usually will be
between 15 and 65 bar, and is preferably between 25 and
45 bar, more preferably about 30 bar. In this respect
it is remarked that the hydrogen partial pressure at
the reactor inlet usually will be 3-10 bar higher than
at the outlet of the reactor.
The catalysts to be used in the process according
to the present invention should contain a hydrogen
activating function. Suitable catalysts comprise one or
more group IVa, group VIb or group VIII metals.
1326464
. -- 10 --
Suitably supports such as silica, alumina, aluminium
phosphates, spinel compounds, titania and zirconia can
be used. Conventional Group VIb and VIII metal
combinations can be employed. It is remarked that the
term "non-acidic" in this specification relates to the
substantial absence of one or more active acidic sites
in the catalyst which are able to accelerate the
cracking reaction of hydrocarbons via carbonium ion
chemistry. Under initial reaction conditions some
acidic sites may be present. However, these acidic
sites rapidly deactivate due to coke formation and
basic nitrogen adsorption whilst the hydrogen
activating function remains substantially unchanged.
When the catalyst comprises a group VIII noble
metal the use of palladium or platinum is preferred.
When the catalyst comprises a group IVa metal
preferably tin is used. When the catalyst comprises a
group VIb metal, preferably molybdenum, chromium or
tungsten is used. When a group VIII non-noble metal is
used, preferably iron, cobalt or nickel is used.
It has been found that very good results can be
obtained using ~o-based catalysts, in particular with
catalysts containing silica as carrier and having a
surface area between 125 and 250 m2~g. The use of such
catalysts allows good hydrodesulphurization activity
together with a low coke make.
; Preferred catalysts are those catalysts which show a distinct but limited hydrodesulphurization activity.
~hese catalysts show a very low coke formation together
with relatively good product properties for the middle
distillate fraction. Preferably the second order rate
constant of the hydrodesulphurization reaction under
the HCTC conditions lies between O.l and l.0, more
preferably between 0.2 and 0.5 l/(h.%S), defined under
1326464
-- 11 --
stationary conditions at 450 C and using Kuwait
flashed distillate.
The hydrogen/feed ratio of the process according
to the present invention may be varied over a wide
range. A suitable hydrogen/feed ratio lies between
50 Nl~kg and 5000 Nl/kg, especially between lO0 Nl/kg
and 2000 Nl/kg. It is preferred to use a hydrogen/feed -
ratio between lO0 and 500 Nl/kg, more preferably
between 200 Nl/kg and 400 Nl/kg. Using these preferred
low hydrogen/feed ratios the coke laydown on the
catalyst is surprisingly very low. Furthermore, a high
cracking conversion is obtained. When compared with a
conventional hydrocracking process the hydrogen/feed
ratio is significantly lower for the process of the
present invention, which is beneficial for process
economics. The usual hydrogen/feed ratio in hydro-
cracking operations lies between 700 and 1500 Nl/kg.
Generally in hydroprocessing high hydrogen/feed ratios
are necessary to suppress coke-formation and to improve
the conversion rate. Surprisingly, in the HCTC process
a low gas rate is not only possible but also beneficial
with respect to both coXe formation and conversion.
In a preferred embodiment the above described
preferred hydrogen/feed ratio of 200 to 400 Nl/kg is
used in combination with a catalyst comprising a group
VIII noble metal, preferably palladium and/or platinum.
The use of the above-indicated hydrogen/feed ratio in
combination with the indicated catalyst resulted in a
very low coke rate, while the amount of 6ulphur on the
catalyst was also surprisingly low.
In another preferred embodiment of the invention
the hydrogen containing stream comprises a mixture of
hydrogen and hydro~en sulphide. Carrying out the
HCTC-reaction with a mixture of hydrogen and hydrogen
sulphide leads to an increase of both conversion level
~326~64
- 12 -
- and the selectivity to middle distillates. The amount
of hydrogen sulphide in the mixture present in the
reactor is suitably up to 50% (v/v~ of the amount of
hydrogen. Preferably an amount of hydrogen sulphide is
used between l and 30%, more preferably between 5 and
25%, and most preferably about 10%.
The HCTC reaction according to the present
invention is suitably carried out in a fixed bed mode,
e.g. a trickle bed downflow reactor. In view of
periodical catalyst regeneration, preferably two or
more fixed bed are used, operated in a swing-operation.
- The HCTC reaction is conveniently carried out in an
upflow fixed bed reactor, especially when relatively
light feedstocks are used. Application of an upflow
reactor in that case will result in a reduced rate of
coke deposition on the catalyst, thereby increasing the
possible run lenght between two catalyst regenerations.
~he reduction of the amount of coke on the catalyst in
the upflow mode can be 50% or more when compared with
the downflow mode. Other preferred modes of operation
the process according the present invention are moving
bed operations, e.g. a bunker flow reactor, and an
ebullated bed operation~
The products produced in the HCTC process can
either be used as such or can be subjected to further
treatment. It is possible, for instance, to subject
part or all of the product(s) obtained to a
desulphurization treatment, in particular a
hydrodesulphurization treatment, to ad~ust the sulphur
amount of the product to the desired amount. A further
possibility comprises subjecting part or all of the
(hydrodesulphurized) product to a hydrofini6hing
treatment, optionally before or after distillation of
the ~hydrode~ulphurized) product. It 19 also possible
:
.
1326~64
- 13 -
to recycle at least part of the unconverted material
present in the product to the HCTC reactor.
Catalyst regeneration is suitably carried out by
burning off the carbonaceous material deposited on the
catalyst using an oxygen and/or steam containing gas.
In case of a fixed bed ~e.g. a swing bed) the catalyst
regeneration may be carried out in the cracking reactor
itself. In case of for instance a bunker flow reactor,
the regeneration is typically carried out in a separate
regenerator.
The invention is illustrated by the following
Examples, although the invention is not limited to
these Examples.
Exam~le 1
Catal~st screening exPerimentS
A Kuwait flashed distillate was subjected to the
hydrocatalytic thermal cracking process according to
the present invention. The feed properties are
described in Table I.
The reaction was carried out in an isothermally
operated microflow trickle bed downflow reactor. The
catalysts were prepared by conventional pore volume
impregnation techniques, unless stated otherwise.
Commercial available carriers (silica or alumina) were
used (catalysts 1 to 12 and 19-21). Commercially
available catalysts, either as such or slightly
modified, were used in exper$ments 13 to 18. The
carrier properties are described in Table II. Inorganic
precursors were used to prepare catalysts 1-12 and
19-21 (e.g. metal nitrates, ammonium molybdate).
Chloride precursors were omitted. Tin was deposited as
an organometallic compound (e.g.
tin(II)2-Qthylhexanoate). NiMo/SiO2 was prepared via a
deposition-precipitation technique as described in e.g.
3s British patent specification 2,189,163. Before use, the
~3264~4
- 14 -
catalysts were calcined at 350-450 C (except for
NiMo/SiO2 catalysts), followed by crushing to smaller
particles (30-80 mesh). An overview of the catalyst
formulations is given in Table III.
Prior to exposing the catalysts to reaction
conditions a sulphiding procedure is applied.
Especially in the case of molybdenum-containing
catalyst~ this is a preferred embodiment as otherwise
during the first hours of the experiment sometimes
excessive coke formation occurred.
Two sulphiding procedures have been applied. The
first one consisted of heating the catalyst together
with a sulphur-containing feedstock and hydrogen at a
rate of 75 C/h to 375 C and keeping the temperature
- 15 constant overnight. Subsequently, the temperature was
increased to 400 C, kept constant for 6 h, increased
to 425 C and again kept constant overnight followed by
heating to 450 C. Another sulphiding and start-up
procedure was applied making use of H2S. Exposing the
catalyst to a mixture of H2~H2S (7/1 v/v) at 10 bar,
the temperature was increased at a rate of 75 C/h to
375 C. Next, the feedstock was introduced and the
temperature was increased at a rate of 75 C/h. It was
checked that both procedures lead to identical catalyst
l 25 performance in terms of e.g. coking. With noble metal
catalysts it turned out that also reduction with
hydrogen prior to testing gave satisfactory results.
The reactions were carried out at 450 C and a
total pressure of 30 bar. The space velocity (LHSY) was
about 1.0 l/l/h. The H2/feed ratio was between 850 and
1100 Nl/kg. The reaction time varied between 170 and
220 hours.
Analvses and data handling
The liquid product was analyzed for the boiling
point distribution using TBP-GLC. Moreover, GLC
132~64
- 15 -
analysis of the off-gas was carried out. On basis of
these analyses conversions and selectivities were
calculated. The conversion has been defined as the net
removal (%) of material boiling above 370 C. The
product slate was split up into gas (C1-C4), naphtha
~C5-150 C), middle distillates (150-370 C) and coke.
The selectivities (%) have been calculated as the
amount of the product in question, divided by the total
amount of products (material boiling below 370 C and
coke). Hydrogen consumptions were calculated on basis
of CNE (Combustion Mass Spectrometric Element) analyses
of both the feedstock and the liquid product and of the
gas analyses. The hydrodesulphurization (HDS) activity
(second order rate constant) was determined from the
sulphur content of liquid product. The results of the
experiments are summerized in Table IV.
The product properties of the middle distillates
obtained are described in Tables IVa and IVb. The
product properties of the bottom fractions are
described in Table IVc.
.,
,
~32~164
- 16 -
TABLE I
PROPERTIES OF KUWAIT FLASH~D DISTILLATE FEEDSTOCK
=====================================_==========z==================
Specific gravity, d 70/4 0.8858
Sulphur %w 2.95
Nitrogen (total) %w 0.0680
Nitrogen (basic) %w 0.0250
Carbon %w 84.76
Hydrogen %w 12.11
Ramsbottom carbon test %w 0.55
Viscosity at 100 C cSt 7.312
Aromatics (W) mmol/100 g
Mono 55.6
Di 27.3
Tri 23.9
Tetra 10.5
Tetra+ 12.5
TBP/GLC
IBP C 353
10 %w recovered at C 395
20 %w recovered at C 412
30 %w recovered at C 426
40 %w recovered at C 439
50 $w recovered at C 450
60 %w recovered at C 462
70 %w recovered at C 476
80 %w recovered at C 491
90 %w recovered at C 515
96 %w recovered at C 523
FBP C --
1~26~64
- 17 -
TABLE II
:
PROPERTIES OF CATALYST CARRIERS
==========================--===================================z===
Carrier A1203 sio2
Shape 0.8 mm extrudates1.5 mm spheres
Pore volume, 0.76 0.85
ml/g (H20)
Pore volume, n.d. o.g~
ml/g (N2)
Surface area, 247 263
m2/g (N2)
Pore diameter, n.d. 19
nm (N2)
Pore diameter, 8.0 14.1
nm (Hg)
__________________ _______________________________________________
~ n.d. = not determined
-^`` 1326464
- 18 -
TABLE III
SURVEY OF EXPERIMENTAL CATALYSTS USED FOR
HCTC OF KUWAIT FLASHED DISTILLATE
============_=============_======
Catalyst Composition PV SA Bulk
number%w on carrier ml/g m2/g density
(g/ml)*
______________________________________._____________________________
1 0.3Pt/0.4Sn/2Cs/A1203 - - 0.53
2 0.2Mo/2.5Sn/A1203 - - 0.65
3 2Mo/2.5Sn/A1203 - - 0.66
4 1.2Ni/2.5Sn~A1203 - - 0.61
1.2Co/2.5Sn/A1203 - - 0.62
6 0.26Mo/3.2Sn/SiO2 - - 0.50
7 3.9Co/11.8Mo/2.4Sn/A1203 - - 0.75
8 0.3Pt/A1203 - - 0.60
9 0.3Pt/2.5Sn/A1203 - - 0.62
4.7Ni/16.2Mo/SiO2 - - 0.66
11 0.4Ni/2.0V/Sio2 - - 0.40
12 4.0Mo/SiO2 - - 0.50
13 2.7Ni/13.2Mo/3P/A12030.5 149 0.73
14 3.2Ni/9.lMo/A1203 0.56 164 0.74
3.2Ni/9.lMo/2C6/Al203 - - 0.78
16 3.2Co/9.6Mo/A1203 0.5g 221 0.71
17 3.2Co/9.lMo/A1203 0.56 165 0.80
18 3.2Co/9.lMo/2Cs/A1203 - - 0.78
19 0.3Pt/0.3Pd/A1203 - - 0.66
2.5Sn/A1203 - - 0.66
21 lOMo/2.5Sn/A1203 - - 0.74
_________________ ___________________________________________ _____
* Bulk density of crushed catalyst
PV - pore volume determined with water
SA = surface area determined with nitrogen (BET method)
132646~
-- 19 --
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-- 1326464
- 21 -
TABLE IVa
PRODUCT QUALITY ASPECTS OF KEROSENE FRACTIONS (150-250 C)
Catalyst number 16 _12 19 ~
Yield on feed, %w 10.6 10.9 16.9 13.2
Specific gravity, d 20/4 0.8130 0.8044 0.8077 0.8030
Mid boiling point, C 208 205 207 204
Freezing point, C -57 n.d. n.d. n.d
Smoke point, mm 15.5 18.5 22.0 19.0
Sulphur, %w 0.025 0.620 0.602 0.697
Carbon, ~w 86.77 86.01 85.80 85.91
Hydrogen, %w 13.21 13.55 13.69 13.59
Ozone number, mmol/g 0.38 n.d. 1.6 n.d.
n.d. = not determined
* H2/feed ratio 250 Nl/kg
1326464
- 22 -
TABLE IVb
PRODUCT QVALITY ASPECTS OF ÇAS OIL FRACTIONS (250-370 C)
Catalyst number 16 12 19 _ 11
Yield on ~eed, %w28.6 25.231.1 26.1
Specific gravity, d 20/4 0.8842 0.8897 0.8893 0.8874
Mid boiling point, C333 324 324 321
Cloud point, C 0 -3 1 -6
Aniline point, C 59.5 n.d.n.d. n.d.
Sulphur, %w 0.241 2.512.26 2.79
Carbon, %w 87.75 85.7685.80 85.50
Hydrogen, ~w 12.05 12.0712.06 12.15
Ozone number, mmol/g0.39 n.dØ91 n.d.
Cetane number 43 47 47 47
n.d. = not determined
~32 64 64
-- 23 --
TAB:iLE IVc
PRODUCT QUALITY ASPECTS OF BOTTOM FRACTIONS (>370 C)
CatalYst number 16 12 19 11
Yield on feed, ~w 51.154.1 35.6 49.3
Specific gravity, d 70/4 0.85720.88430.9116 0.9174
Viscosity (60 C), cSt 67.70 15.45 19.98 23.26
Mid boiling point, C 4 2 04 2 4 423 432
Sulphur, %w 0.2992 .192.67 3.41
Carbon, %w 87.3885.7886.71 85.92
Hydrogen, %w 12.4211.8411.12 11.22
RCT, %w 0.150.19 1.24 0.87
Example II
Effect of Pressure
Using the same general reaction conditions as described in
Example I the effect of the total pressure was investigated. The
results are summarized in Table V.
-` 132~64
-- 24 --
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1326464
- 25 -
Example III
Comparison of catalysts with different HDS-act vities
Using the same general reaction conditions as
described in Example I, the relationship between the
hydrodesulphurization activity and the coke selectivity
of some catalysts was studied. The results are
summarized in Table VI.
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-` 1326~6~
- 27 -
Example IV
Effect of temperature
Using the same general conditions as described in
Example I, the effect of the temperature was
investigated. Catalyst No. 16 was used in experiments
1, 2 and 3, catalyst No. 12 was used in experiment 4.
The results are summarized in Table VII.
TABLE VI I
Experiment number 1 2 3 4
Temperature deg.C390 425 4~0 500
Runhours h213 165 171 166
H2/feed Nl/kg 900 900 goo 1800
Conversion %19.1 43.1 52.8 51.8
Selectivity Cl-C4 ~ 2.6 5.0 7.5 13.8
Selectivity C5-150% 6.4 10.9 9.3 18.6
Selectivity 150-250 % 19.0 22.4 24.5 22.5
Selectivity 250-370 ~ 71.7 61.4 58.2 44.8
Selectivity Coke % 0.15 0.18 0.34 0.52
Total H2 cons. ~wof 0.64 0.55 0.67 n.d.
Exam~le V
Effect of run length
Using the same general reaction conditions as
described in Example I, the effect of the run length at
a low H2/feed ratio was investigated. Catalyst 11 was
used for all experiments. The results are summerlzed in
Table VIII.
,~ .
1326~64
- 28 -
TABLE VIII
Experiment number 1 2 3 4
H2/Feed Nl/kg 280 270 230 230
Runhours h23 47 208 425
Conversion %64.0 66.2 57.4 49.5
Selectivity C1-C4 % 3.8 4.S 4.4 3.7
Selectivity C5-150 % 19.8 20.8 18.0 18.1
Selectivity 150-250 % 27.0 26.8 27.6 26.3
Selectivity 250-370 % 49.3 47.9 49.9 51.9
Selectivity Coke % 0.09 0.04 0.02 0.02
Total H2 cons. %wof 0.23 0.17 0.28
K2 HDS l/(h.%WS)0.16 0.13 0.10 0.07
Example VI
Effect of catalyst comPosition on sul~hur deposition
Using the same general reaction conditions as
described in Example I, the effect of the catalyst
composition on sulphur deposition was investigated. The
results are summerized in Table IX.
1321~64
- 29 -
TABLE IX
Experiment number1 2 3 4 5
Catalyst number 1 1 11 11 _ 19
Gas rate, Nl/kg240 910 230 1020 250
Run time, h 190 196 208 138 209
Conversion of >370 C 63 52 57 45 64
Coke on spent
catalyst, %w
Carbon + hydrogen5.9 14.5 3.9 12.1 14.3
Sulphur <0.6 0.3 7.4 2.0 0.6
ExamPle VII
E~fect of feedstock
Using the same general reaction conditions as described in
Example I, three different feedstocks were compared. The feedstock
properties are described in Table X. The Kuwait flashed distillate is
described in more detail in Table I. The Xuwait deasphalted oil i~ a
butane-deasphalted short residue. The Maya synthetic flashed
distillate has been produced by hydrodemetallization and
hydroconversion of Maya short residue, followed by flashing. The
results of the experiments are summerized in Table XI.
TABLE X
Feedstock Kuwait Kuwait Maya synthetic
flashed deasphalted flashed
distillate oil distillate ____
bp > 370 C, %w 98 lO0 88
Sulphur, %w 2.95 4.15 0.42
Nitroyen, %w0.07 0.21 0.35
RCT, %w 0.55 6.12 1.8
H/C (atomic)1.71 1.63 1.59
~` 1326~
- 30 -
TABLE XI
Feedstock Kuwait FD Kuwait DAOMaYa SFD
Catalyst number 16 14 16
Net conversion of
bp > 370 C material, % 52.8 47.9 29.0/43.9*
Selectivity to products, ~w
1 C4 7.5 9 4 6.5
C5-150 C 9.3 13.2 8.3
150-370 C 82.7 77.084.9
coke 0.34 0 340 37
Total liquid product:
sulphur content, %w <0.2 0.8 ~0.2
nitrogen content, %w 0.05 0.18 0.30
~2 consumption, %w on feed 0.7 1.2 _ n.d.
* at LHSV = O.5 1.1 l.h 1
n.d. = not determined
Example VIII
Effect of H2S in hydrogen feed
Using the same general reaction conditions as
described in Example I, the effect of H2S in the
hydrogen feed was investigated. Catalyst No. 3 was used.
The results are summarized in Table XII.
-- 1326464
-- 31 --
TABLE XI I
Experiment 1 _ 2
Fresh gas composition
% vol H2 100 90
% vol H2S o 10
Net conversion of 370 C+ 46.5 51.3
Selectivities, %wof
C1 C4 (gas) 6.0 4.9
C5-150 C (naphtha) 13.5 12.6
150-370 C (mid. dist.) 80.5 82.5
Example IX
Use of HCTC-unconverted material in fluidized catalYti
cracking
An HCTC experiment was carried out at 450 C, 30
bar pressuref an H2/feed ratio of 900 Nl/kg, and a LHSV
of 1.0 l/l/h using catalyst No. 16 and using Kuwait
flashed distillate as feed. The unconverted material,
i.e. the fraction boiling above 370 C was used as feed
for a fluidized catalytic cracking (FCC) reaction. The
FCC-unit was operated at constant coke yield and
stripper efficiency. A second experiment was carried out
using Kuwait flashed distillate. The feed-properties and
the FCC yields are summarized in Table XIII.
-- 1326464
- 32 -
- TABLE XIII
Feedstock: Unconverted material Kuwait Plashed
of HCTC ex~eriment Distillate
_
Properties:
Specific gravity, d70/4 0.8572 0.8858
Sulphur, ~w 0.30 2.95
RCT, %w 0.15 0.55
Viscosity, cSt 67.70 (60 C) 7.312 (loO C)
Mid boiling point, C 420 420
~CC-yields (% weight on feed):
C2 2.0 3.3
C3 5.5 5.6
C4 9.7 10.0
C5-220 C 48.8 47.4
LCO 19.4 18.6
HCO 8.7 9.1
Coke 6.0 6.0
Example X
Use of thermallY cracked flashed distillate as feed for HC-
TC
Thermally cracked flashed distillate originating from
Arabian heavy ~eedstock was used as feed for an HCTC
experiment carried out at a temperature of 450 C, a
pressure of 30 bar, a LHSV of 1.0 l/l/h, a H2/feed ratio of
250 Nl/kg and a run length of 161 hrs, usinq catalyst No.
12. Feedstock properties: specific gravity ~d 70/4):
0.9139, sulphur (%w): 2.22, nitrogen (%w) 0.31, RCT: 0.4
(%w). The net conversion was 42.7%. The results are
summarized in Table XIV.
-- 1326~6~
- 33 -
TABLE XIV
Selectivities, %
Cl C4 3'9
C5-150 C 11.5
150-250 C 21.6
250-370 C 62.9
Coke 0.05
Total H2 cons. ~wof 0.22
K2 HDS l/(h.%wS) 0.24
ExamPle XI
Use of feed containing enhanced amount of ashaltenes
A Kuwait long residue was used as feed for an HCTC
experiment carried out at a temperature of 450 C, a
pressure of 50 bar, a LHSV of 1.0 l/l/h, a H2/feed ratio
of 1000 Nl/kg and a run length of 50 hours, using
catalyst No. 12. Feedstock properties: specific gravity
(d 70/4): 0.9139, sulphur (%w): 3.69, nitrogen (%w):
0.15, metals (ppm): 42, RCT (%w): 5.1, C7-asphaltenes
(%w): 2.4. The net conversion was 45~. Selectivities
(%): C1-C4: 8.0, C5-150 C: 11.1, 150 C-370 C: 80.1,
coke: 0.7.
