Note: Descriptions are shown in the official language in which they were submitted.
F-4149 ~L3~)Z3~
PROCESS FOR HYDROTREATING CATALYTIC CRACKING FEEDSTOCXS
This invention relates to the catalytic cracking of
petroleum oils and more particularly, to a process for improving
catalytic cracking processes by hydrotreating the feedstock.
Catalytic cracking is an established and widely used
process in the petroleum refining industry for converting oils and
residua of relatively high boiling point to more valuable lower
boiling products including gasoline and middle distillates such as
kerosene, jet fuel and heating oil. The pre-eminent catalytic
cracking process now in use is the fluid catalytic process (FCC) in
which the pre-heated feed is brought into contact with a hot
cracking catalyst which is in the form of a fine pcwder, typically
with a particle size of 10 - 300 microns, usually about 100 microns,
for the desired cracking reactions to take place. During the
cracking, coke is deposited on the catalyst and this results in a
loss of activity and selectivity. The coke is removed by
continuously removing the coked, spent catalyst from the cracking
reactor and oxidatively regenerating it by contacting it with air in
a regenerator. The combustion of the coke is a strongly exothermic
reaction which, besides removing the coke, serves to heat the
catalyst to the temperatures appropriate for the endothermic
cracking reaction. The process is carried out in an integrated unit
comprising the cracking reactor, the regenerator and the appropriate
ancillary equipment. The catalyst is continuously circulated from
reactor to regenerator and back to the reactor with the circulation
rate being adjusted relative to the feed rate of the oil to maintain
a heat balanced operation in which the heat produced in the
regenerator is sufficient for maintaining the cracking with the
circulating, regenerated catalyst being used as the heat transfer
13~Z3~
F-4139 --2--
medium. Typical fluid catalytic cracking processes are described in
the monograph Fluid Catalytic Cracking with Zeolite Catalysts~
Venuto, P. B. and Habib, E. T., Marcel Dekker Inc., N.Y. 1979, to
which reference is made for a description of such processes. As
described in the monograph, the catalysts which are currently used
are based on zeolites, especially the large pore synthetic
faujasites, zeolites X and Y, which have generally replaced the less
active, less selective amorphous and clay catalysts formerly used.
Another catalytic cracking process still used in the
industry is the moving, gravitating bed process, one form of which
is known as Thermofor Catalytic Cracking (TCC) which operates in a
similar manner to FCC but with a downwardly moving gravitating bed
of a bead type catalyst, typically about 3 - 10 mm in diameter.
Fixed bed units have now been replaced by moving or fluidized bed
lS units of the FCC or TCC type.
The feed to the catalytic cracker can generally be
characterized as a high boiling oil or residuum, either on its own
or mixed with other fractions, usually of a high boiling point. The
most common feeds are gas oils, that is, high boiling, non-residual
oils with an initial boiling point usually above about 230C (about
450F), more commonly above about 345C (about 650F), with end
points of up to about 620C (about 1150F). Typical gas oil feeds
include straight run (atmospheric) gas oil, vacuum gas oil and coker
gas oil; residual feeds include atmospheric residua, vacuum residua
and residual fractions from other refining processes. Oils from
synthetic sources such as Fischer-Tropsch synthesis, coal
liquefaction shale oil or other synthetic processes may also yield
high boiling fractions which may be catalytically cracked either on
their own or in admixture with oils of petroleum origin.
The ease with which any given cracking feed is cracked and
the selectivity for the desired products depends partly upon the
composition of the feed. Virgin petroleum stocks which have not
previously been subjected to cracking tend to crack relatively
easily because they possess long chain alkyl groups which, by a
13~Z3Z9
F-4139 ~~3~~
process of dealkylation which occur readily during cracking, form
lower boiling products. The aromatic residues which are left
following dealkylation are more highly refractory so that catalytic
cracking cycle oils, e.g. LCO or HCO, generally require severe
hydrotreating to saturate them before they can be cracked to any
significant extent under conventional conditions. Besides improving
crackability, especially of these highly aromatic cycle oils,
hydrotreating has been recognized as useful for other purposes,
including demetallation and, above all for desulfurization and
denitrogenation, both of which are desirable to improve product
quality, catalyst selectivity and aging rate as well as reducing
emissions, principally SOx, from the regenerator. Use of a feed
hydrotreater has been reported to result in higher conversion and
gasoline yield, lower coke make per pass, more favorable light gas
distribution, higher isobutane yield, and lower contaminant content
in products and unit off-gases: cf. the Venuto/Habib monograph and
the Oil and Gas Journal, 19 May 1966, 131-139; 14 Oct 1974, 99-110;
21 July 1975, 53-58.
In general, the hydrotreating of catalytic cracking
feedstocks has been carried out at relatively low pressures, below
about 7000 kPa (about 1,000 psig) and in most cases below about
5,500 kPa (about 785 psig) as the use of higher pressures does not
enhance the desulfurization activity which has, for the most part,
been the principal objective of cracking feed hydrotreating.
Denitrogenation ha$ generally followed with the desulfurization to
some degree, depending upon the composition of the feed and the
severity of the processing although more severe processing
conditions are required for nitrogen removal than sulfur removal.
The hydrogenation of aromatics is known to be favored by the use of
higher hydrogen partial pressures although the response of aromatic
compounds differs according to their composition: condensed ring
aromatics may be hydrogenated at lower pressures than non-condensed,
polycyclic aromatics and although the hydrogenation of both is
favored by the use of relatively low temperatures (since
~3~232~
~-4139 --4-~
hydrogenation is an exothermic reaction) the optimum temperature for
hydrogenating each of these types varies (Oil and Gas Journal 21
July 75~ 53~58)r The use of lower temperatures, however, does not
favor sulfur or nitrogen removal because these reactions require a
certain measure of cracking (an endothermic reaction) and the same
is true of demetallation although, in most cases, a relatively mild
treatment will suffice for metals removal. The choice of conditions
for the hydrotreating of catalytic cracking feeds has therefore
represented a compromise between different competing factors without
a significant attempt at optimizing the conditions in a way which
permits the greatest advantages to be obtained at minimum cost. In
the case of hydrotreating, minimum cost implies minimum hydrogen
consumption and maximum catalyst cycle life.
We have now found that the hydrotreating of catalytic feeds
may be optimized in an unexpected manner which permits the cracking
operation to be operated under more favorable conditions with better
product ~uality and reduced coke make and which, moreover, enables
hydrogen consumption in the hydrotreating step to be minimized
without adversely affecting the catalyst aging rate. At the same
time, sulfur and nitrogen removal take place readily so as to ensure
satisfactorily low contaminant level in the products and a low level
of pollutants in the emissions, particularly of SOx, from the
cracking unit.
The basis of the present invention is our finding that in
the hydrotreating of catalytic cracking feeds, the saturation of
aromatics is more important than nitrogen removal and that if
conversion, i.e. bulk conversion to lower boiling products, in the
hydrotreater is minimized, the hydrogenation of cracking fragments
will be correspondingly reduced. This, in turn, will result in a
lower hydrogen consumption, most of which will be represented by a
reduction in the quantity of hydrogen-rich dry gas produced during
the hydrotreating step. Thus, by operating under appropriate
conditions, the hydrogen is used more efficiently -- to ncrease
aromatic saturation without increasing dry gas production -- and
13~23Z~
F-4139 ~~5~~
this enables the cracking operation to be carried out under more
favorable conditions with improved product quality and distribution.
As will be explained in greater detail below, the present
catalytic cracking feed hydrotreating process is carried out under
conditions which may generally be characterized as those of high
pressure, low temperature and low severity with limited conversion
of not more than 20 volume percent. Pressures are typically above
about 10,000 kPa (about 1435 psig) and preferably above about 11,000
kPa (about 1,580 psig). Temperatures, as indicated by the
start-of-cycle (SOC) temperature are typically below about 390C
(about 734F) and preferably below 365C (about 690F). At these
low temperatures, space velocity (LHSV) may be relatively 1QW~
typically at about 1 hr 1 or even less; although desulfurization
would proceed at higher space velocities, typically up to 2 hr 1,
catalyst aging would be accelerated so that, in general, space
velocity will be below 2 and in most cases below 1.5. End of cycle
(EOC) temperatures are generally to be limited to 425C (800F) and
preferably, 415C (780F).
In the accompanying drawings:
Fig. lA is a graph relating the effect of hydrotreating
severity on coke make during cracking and on cracking
conversion;
Fig. lB is a graph relating the effect of hydrotreating
severity on gasoline yield from cracking;
Fig. 2 is a graph relating the catalyst:oil ratio during
cracking against the resulting conversion;
Fig. 3A is a graph relating the coke make during cracking to
conversion during cracking at different hydrotreating
pressures and denitrogenation levels;
Fig. 3B is a graph relating the coke make during cracking
to gasoline yield from cracking at different
hydrotreating pressures and denitrogenation levels;
Fig. 4 is a graph relating hydrogen consumption to
hydrotreating severity at different pressures.
13~JZ3;~9
F-4139 --6--
The present process provides for improvements in the
catalytic cracking of hydrocarbon feeds such as those conventionally
subjected to catalytic cracking processes, e.g. gas oils, long and
short resids, and other high boiling fractions such as those
previously described. However, the process is primarily applicable
to non-residual feeds since it is with these that the optimal
results are obtained. Generally, the feed will have an initial
boiling point of at least 205C (about 400F), more usually at least
315C (about 600F) and in most cases at least 345C (about 650F).
End points will depend upon other processing constraints,
particularly the ability to handle the residual fractions left after
removing high end point distillates~ Generally, end points will be
from 510C (about 950F) to 565C (about 1,050F) although lower and
higher values may be encountered.
The high boiling feeds which are to be subjected to
catalytic cracking are generally characterized by a relatively high
content of aromatic components including polycyclic aromatics and
fused ring aromatics together with paraffins, naphthenes and
heterocyclics, the relative proportions of each being dependent on
the origin of the feed and its previous processing history.
Naphthenes will mostly be dicyclo- and tricyclo-paraffins,
especially at higher boiling points and the proportion of dinuclear
and polynuclear aromatics will also increase with increasing boiling
point. The paraffin content is generally less than 30 volume
percent so that aromatics and naphthenes will together constitute at
least 70 volume percent of the feed, with the majority of this being
aromatics. A number of contaminants will also be present,
principally sulfur, nitrogen, oxygen and metallic impurities in
amounts dependent upon the source of the feed. In general, sulfur
will be present in amounts from about 0.1 to 3 weight percent
although very heavy crudes, e.g. Boscan (Venezuela) may yield feeds
which may be as high as about 5 weight percent sulfur. Nitrogen
will generally be present in amounts from 0.05 - 0.5 weight percent
(500-5000 ppmw) although some synthetic feeds such as shale oils may
13~'2~Z9
F-4139 ~~7~~
have higher contents. A number of different metals may be present,
including sodium and heavier metals, of which the most significant
from the view of the catalytic cracking process are nickel and
vanadium since they exert the greatest deactivating effects on the
cracking process. Combined nickel and vanadium contents will
generally be from about 5 to 1,000 ppmw, with many being from 20 to
lOO ppmw, according to source. Various formulae have been developed
to express the metals content of the feed in a single figure,
weighting the content of each of the metals present according to its
deactivating effect on the cracking catalyst. For example, refer to
the nickel equivalent formula of Oil and Gas Journal 23 Oct 1961,
143 where
Nickel equivalents = Ni ~ V ~ Fe ~ Cu
-4.~ 7~T l~Z~
(see also U.S. Patent No. 4,376,038).
These contaminants should desirably be reduced to low levels
prior to the catalytic cracking operation not only because they may
adversely affect the functioning of the cracking catalyst but also
because they may enter the cracked products or be emitted as a
pollutant from the cracking process, e.g. Sx from the regenerator
stack. Metals, especially nickel and vanadium, should be reduced to
a low level, usually below 10 ppmw nickel equivalents, preferably
below 1 or 2 ppmw nickel equivalents. Various methods exist for
doing this. For example, sodium may be removed by a desalter and
metals may be removed by coking but another method commonly used is
hydroprocessing over a catalyst, as described above. This is the
method employed in the present process. Generally, the conditions
may be described as being elevated temperature and pressure with a
catalyst which combines acidic and hydrogenation functionality. The
present process departs from prior practice, however, in its careful
selection of conditions so as to favor desulfurization and
denitrogenation while, at the same time promoting aromatics
saturation. Although nitrogen is known to affect the cracking
process adversely it has been found, surprisingly, that aromatics
13~1;Z3
F-4139 --8--
saturation was more important than nitrogen removal and that it was
possible to obtain better operation of the cracking step as
aromatics saturation increased. Although hydrogen consumption is
increased by the aromatics saturation which is achieved at the high
pressures employed in the process, the benefits achieved are more
than sufficient to outweigh the added costs provided that the
conditions are controlled so as to maintain the desired reaction
equilibria.
In general, the conditions used in the present process may be
characterized as being of low temperature, high pressure, low space
velocity and low conversion. Because one objective of the process
is to saturate aromatics at the expense of denitrogenation,
conversion is to be limited because under more severe conditions,
hydrogen does not go into the aromatics as efficiently (the term
"conversion" is used here in its specialized sense to mean a bulk
conversion, that is, a conversion to lower boiling components).
Based on the discovery that suitable choice of reaction conditions
can lead to increased aromatics saturation without, at the same
time, correspondingly increased nitrogen removal, the process can be
optimized so as to improve the processing in the cracking step at
minimum cost. The production of distillate in the hydrotreating
step is minimized, i.e. conversion to lower boiling products is
maintained at a low level so that the hydrogen which is added during
the hydrotreating is, for the most part, carried over to the
subsequent cracking step in order to facilitate the operation of the
cracking unit. If conversion during the hydrotreating exceeds
desirable values not only is valuable hydrogen lost in the form of
dry gas but also as hydrogen-rich naphtha and distillate, so that
the cracking step operates on a relatively hydrogen-deficient feed,
with consequent deterioration in its operation: coke make increases
and gasoline selectivity decreases. The production of low boiling
distillates in the hydrotreating step is also undesirable because
the naphtha product, being straight chain paraffinic, is low in
octane. Thus, the maximum total conversion, i.e. conversion in the
23Z~
F-4139 ~ 9
hydrotreating and cracking steps combined, is obtained by minimizing
conversion in the hydrotreating step so as to obtain the most
favorable, selective conversion in the cracking step. For any given
feed, there is an optimum in hydrogen consumption which varies
according to the composition of the feed. Thus, for any selected
feed, the optimum hydrotreating conditions (to maximize hydrogen
input and minimize conversion) should be selected empirically
according to factors including, principally, aromatic content and
nitrogen level.
An unexpected characteristic of the present process is that an
improved process is obtained by designing for higher pressure.
Conventionally, the objective in hydrotreater design is to minimise
hydrogen consumption and this is generally achieved by operating at
lower pressure. The principal objective in the present process,
however, is to minimise light gas production and this is attained by
high pressure, low temperature operation. When these factors are
coupled with the preferred denitrogenation levels, noting that
different temperatures are required at different pressure for a
constant denitrogenation level, the final cracking process is
significantly improved.
The hydrogen partial pressures used in the present hydrotreating
step are generally in excess of 10,000 kPa (about 1,435 psig) and
preferably above about 11,000 kPa (about 1,580 psig). The upper
limit on the pressure is set mainly by equipment limitations which,
in turn, are dictated by the cost of fabricating vessels for high
pressures. In most cases, the total system pressure in the
hydrotreating step will not exceed about 30,000 kPa (about 4,335
psig) and in most cases will be below about 20,000 kPa (about 2,900
psig). The use of higher hydrogen partial pressures tends to
increase aromatics saturation and therefore is preferred because a
higher level of saturation improves the cracking performance.
However, the use of higher pressures will lead to increased hydrogen
consumption with consequent higher costs and heavier demands on the
hydrogen plant. Also, saturation beyond a certain level produces
~L3~
F-4139 --10--
decreasing benefit since at these higher pressures relatively less
hydrogen goes into saturation of polynuclear aromatics. Thus, a
balance may need to be struck between attainment of highest process
objectives and the benefits conferred by this.
s In general, pressures between about 11,000 and 15,000 kPa (about
1435-2160 psig) will give the optimum benefit in the process as a
whole with pressures of about 12000 kPa giving particular benefit.
The temperature is to be selected in part according to the
hydrogen pressure used because higher temperatures may be tolerated
lo at the higher hydrogen partial pressures which tend to inhibit
catalyst aging. However, the thermodynamics of the desired
aromatics saturation process -- which is exothermic -- provide a
strong incentive for maintaining the temperature at as low a level
as possible, consistent with achieving a severity high enough to
achieve the degree of desulfuriza~ion which is actually desired.
Because the use of higher pressures does not favor desulfurization
activity, the level of desulfurization achieved will be dependent on
the severity of the hydrotreating step, as indicated in part by the
temperature. Another factor relevant to the final choice of
temperature is the conversion during the hydrotreating step. As
mentioned above, one objective is to minimize conversion to naphtha
and dry gas because this represents a loss of valuable hydrogen to
the cracking process where its presence can produce the greatest
improvement. A certain level of conversion is, however, inevitable
and is also necessary if desulfurization is to be achieved. The
temperature used, therefore, will depend upon the characteristics of
the feed (especially its cracking characteristics, degree of
unsaturation and sulfur content) as well as catalyst characteristics
(especially its aging rate and the balance of its
cracking/hydrogenation activities).
In terms of numerical values, the temperature of the
hydrotreating step will be between 345 and 455C (about 650 and
850F), preferably between 365 and 455C (about 690 and 850F),
with the lower figure being applicable to the start-of-cycle (SOC)
~ 3~Z3~
F-4139 --11--
temperature~ The preferred upper temperature limit is about 410C(about 770F) for the end of the cycle (EOC), preferably 400C
(750~F) after which catalyst regeneration is required. The actual
end of cycle temperature will, however, depend upon the
characteristics of the feed and the catalyst, especially its
hydrogenation activity, since the run should be terminated when the
saturation of the aromatics in the feed reaches an unacceptably low
level by reason of the increased temperature. (Temperatures
referred to are average catalyst bed temperatures, entire bed
averaged axially, edge effects excluded).
The aging rate of the catalyst and the concomitant necessity to
increase processing temperature as the cycle progresses is therefore
a factor in the selection of temperature at any given time in the
cycle. The aging rate, in turn, will be dependent on severity and
hydrogen partial pressure. Severity, in turn will depend upon space
velocity. Because the present process seeks to attain a high degree
of aromatics saturation with denitrogenation as only a secondary
objective, lower space velocities are desired at the lower
temperatures used to favor the exothermic saturation reactions.
Thus, in the present process, space velocity (LHSV) will be
generally below 2 hr 1, more ~enerally in the range 0.5 - 1.5
hr l, with optimum values of about 1 hr l. In this respect, the
present process represents a departure from prior FCC feed by
hydrotreating practice which used relatively high space velocities
of about 2 hr 1 (LHSV) for desulfurization as the main objective.
At these severities, desulfurization will be sufficient to meet FCC
Sx emission restrictions, equivalent to about 85% desulfurization
of a feed with about 1 to 2 weight percent sulfur; in most cases,
desulfurization will be at least 95% with nitrogen removal at 60 -
80%. Cycle durations of 1 to 2 years are envisaged at these
severities, particularly if the preferred mode of operation at
temperatures entirely below about 390C ~about 730F) are employed.
The hydrotreating catalyst used may be conventional in type,
comprising an acidic or non-acidic support, e.g. silica, alumina,
~3~Z3Z~
F-4139 --12--
silica-alumina or a crystalline aluminosilicate, preferably a large
pore zeolite such as natural faujasite, zeolite X, zeolite Y or a
composite of these materials such as zeolite X in an amorphous
silica or silica-alumina matrix, together with a hydrogenation
component which is typically a transition metal or metals, usually a
base metal of Groups VA, VIA or VIIIA of the Periodic Table (IUPAC
Table). In order to maintain conversion at a low level during the
hydrotreating step, a low acidity support should be used. Base
metals are usually preferred for the hydrogenation component because
of their low cost and resistance to poisoning by contaminants but
noble metals such as platinum or palladium could be used. Preferred
base metals are vanadium, chromium, cobalt, nickel, molybdenum and
tungsten; combinations of metals such as cobalt-molybdenum,
nickel-cobalt, nickel-molybdenum may be used to advantage. If the
feed to the hydrotreater contains significant quantities of nitrogen
as a contaminant, it will generally be preferred to avoid the use of
cobalt because the catalytic activity of this metal is inhibited by
nitrogen but other metals and combinations, e.g. Ni-Mo, may be used.
The effluent from the hydrotreating step is then passed to the
catalytic cracking step. The catalytic cracking is conventional in
type and may be a fluid catalytic cracking (FCC) operation or a
gravitating moving bed operation, e.g TCC or Houdriflow. Cracking
conditions are selected according to the characteristics of the
cracking feed from the hydrotreater together with any recycle and,
of course, according to the desired products, product distribution
and the characteristics of the catalyst and of the unit. The
hydrotreating of the feed in the manner described above provides a
cracking feed of improved crackability and with reduced contaminant
level. Because the use of high hydrogen pressures tends to saturate
the feed aromatics to a greater extent, the cracking proceeds more
readily to cracking products which are lower in molecular weight and
with a relatively higher hydrogen content than would be obtained
with the use of low to moderate hydrotreating pressures. The
cracked products include significantly higher levels of high octane
13~ ~ Z3Z9
F-4139 --13--
gasoline. By contrast, mild hydrotreating of the feed using lower
pressures tends to produce cracked products with a higher proportion
of low cetane distillate because of the higher aromatic content of
the feed. Also mild hydrotreating conditions may lead to production
of an olefinic gasoline which, although high in octane, may be
considered undesirable by regulatory anthorities because of its
asserted effects on the environment.
Example 1 (Cases A - K)
The effects of hydrotreating FCC feed under different conditions
were evaluated as described below.
A base feed of California/Alaskan (88:12) crude origin gas oils
(mixed vacuum gas oil, coker gas oil) was prepared. Compositions of
the feeds are shown in Table 1 below.
13~23Zg
F-4139 --14--
Table 1
Feed Composition
. . .
Base
Feed Components, Vol % Feed
Atmospheric Gas Oil
Vacuum Gas Oil (527C EP) 78
Coker Gas Oils 22
Gravity, API 16.8
Sulfur, wt. pct. 1.28
Nitrogen, ppmw 5100
Basic Nitrogen, ppmw 1900
Ni/V, ppmw 1.5/2.4
CCR, wt. pct. 0.9
345CI Aromatics, wt. pct. 68.0
Distillation, Wt. Pct.
250C- 4.8
250-345C 15.3
345Cl 79.9
The feed was hydrotreated under varying conditions using a commercial
Ni-Mo/A12O3 (3 pct Ni, 18 pct Mo) hydrotreating catalyst. The
hydrotreating conditions used are given in Table 2 below.
~3~'23;Z~
F-4139 --15--
Table 2
Hydrotreating Conditions for FCC Feed Preparation
Average De-N
Pressure, Reactor Wt. Pct. (2)
Case DOS(l) kPa (psial Feed LHSV Temp C(F) Total Basic
A 812065 (1750) Base 0.9 365 (690) 53 56
B 1512065 (1750) Base 2.2 400 (752) 61 68
C 258650 (1240) Base 1.0 384 (723) 61 68
D 448650 (1240) Base 1.8 411 (772) 65 77
F 3612065 (1750) Base 0.5 377 (710) 75 81
G 4112065 (1750) Base 1.0 401 (753) 80 89
H 1312065 (1750) Base 1.5 363 (685) 37 36
I 298650 (1240) Base 1.0 374 (706) 45 50
J 468650 (1240) Base 1.0 421 (790) 78 94
L 318650 (1240) Base 2.4 418 (784) 49 48
Notes:
(1) Days on stream
(2) 250C+ (480F~) product relative to feed
Cases A and B were designed to simulate start- and end-of-cycle
hydrotreating performance at 12065 kPa (1750 psia) hydrogen pressure and
60% denitrogenation. Cases C and D were designed for this same
denitrogenation level, but at 8650 kPa (1240 psia) hydrogen. Some
unavoidable variation in denitrogenation was observed. Several other
cases (F-K) examined various denitrogenation severities at both pressure
levels. Case L reflects hydrotreater operation at reduced hydrogen
circulation at 1240 psia hydrogen.
Each of the hydrotreater products was distilled to nominally
250C (480F) to produce the FCC feed.
The properties of each hydrotreated FCC feed are detailed in
Table 3 below.
F-4 149 - --16-- 13~
~¦æ ~ 0 ~ 0 ~
o o ~ _ -- ~ ~ o _ g~
- o -
O O ~ _ U~
I¦æ ~ 8 æ ~ ~ 2 ~
o O u~ _ ~ O
o ~ ~ o o o ~
O O ~ - ~ ~ æ
1~ -oo~
o o ~ _ ~ _ ~
o ~ ` ~ ~ #
O O _ -
o ~ ~ 2 0 ~ 0 ~ 3
~1~
' 1~ ~ ~ ~ o o o
~ o o ~ o
Ln ~ ~ ~ ~ $ O C~ O O ~ ~ ~
o o
Ln ~ o ~ ~ o Q O O
O O
n ~ ~
~ ~ Q~ ~` ~ ~ Sri ~ ~ o ~n ~I ~o g ~ ~ ~ ________
:~ ~ ct~ QO ~ ) O '7 ~ L~ n
O O~ _ r~ ~ ~ o~ -- _ r~
~n 3
~ ~`~ ~ ~ _ 3 ~ ~3 ? ~n 2 ~ 'n~ ~ ~ cr~ , _ _ _ ~ _
c . .~n ~ ~~ O ~i o ~ o ~ n Lr~
~ ~ ~ I`n ~ ~ Q ~n
o ~ U e ~
.~.~ n .Lr~ I~ ~ 8 3~ ~ ~ o ~ s
13t~2;~Z~
F-4139 --17--
These data show a wide range of hydrotreated FCC feed
properties at both hydrotreating pressures. At the same
denitrogenation severity, the higher pressure hydrotreating cases
gave significantly more aromatics saturation, particularly at high
s denitrogenation. Some variation in the amount of 480-650F in the
feed was observed (particularly in Case J); however, for the most
part, the variation was apparently caused by laboratory distillation
differences, rather than actual yield difference caused by
hydrotreating severity or raw feed differences.
The hydrotreated feeds were catalytically cracked in an FCC
pilot plant having a capacity of 80 l. day 1 (0.5 BPD) using an
equilibrium cracking catalyst having the properties given in Table 4
below.
~3~ tZ3Z~
F-4139 --18--
Table 4
FCC Catalyst Properties
Density, g~cc
Packed 0.91
Loose 0.80
Real 2.78
Particle 1.31
Pore Volume, cc/g 0.41
Surface Area, m /g 97
Composition
A1203, wt. pct. 46.4
SiO2, wt. pct. 42.0
Re203, wt. pct. 2.4
Ni, ppmw 390
V, ppmw 805
Fe, ppmw 0.6
Cu, ppmw 54
Sb, ppmw 1.0
FAI Test, Clean Burned (1) 174-6933
Conversion, vol pct. 61.7
Gasoline, vol pct. 54.9
C4, v? pct. 11.9
Dry Gas, wt pct. 3.3
Coke, wt. pct. 0.87
CLETG0, wt. pct. 0.39
Hydrogen Factor, mol/mol 54
Note:
(1) Fluid Activity Index
~3~3Z~
F-4139 --19--
In the cracking tests, two to eight material balances were
taken for each feed. Catalyst-to-oil (C/O) ratio was varied from
about 5 to about 12 weight catalyst/weight oil. Riser top
temperature was held constant at nominally 530C (990F) throughout
the study, although the effect of riser top temperature was
evaluated with two additional balances at 520C (968F) and 541C
(1006F) with the Case A feed. All o~her variables including
oil-to-riser temperature were held essentially constant.
- Spent and regenerated catalyst samples were taken twice
during each four hour material balance. Equilibrium catalyst was
added liberally to maintain catalyst activity. In addition, three
complete catalyst change-outs were performed during the study. As a
result, FAI variation over the entire study was less than 5
numbers. The results are presented in summary form in the
accompanying drawings and the following discussion.
The various parameters examined in this study, including:
denitrogenation severity, hydrotreating pressure, FCC feed cut
point, FCC riser top temperature and added light atmospheric gas oil
in the feed, are discussed below.
Denitrogenation Severity
In order to establish the effect of denitrogenation
severity, several feed cases were grouped together as follows:
Denitro~enation, % Ca
O Raw
40~50 H, I, L
55-65 A-D
75-80 F, G
On a constant coke basis, an increase in hydrotreater severity
yields higher conversion and gasoline yields, as shown in Figures lA
and lB which show the effect of denitrogenation selectivity on coke
selectivity (Fig. lA) and gasoline selectivity (Fig. lB) in the FCC
process. The increase in gasoline yield at 5.5% coke is about 18
vol % at 40-50% denitrogenation versus raw and up to 32 vol % at
13~23~9
F-4139 --20--
75-80% denitrogenation. This increase in gasoline yield at constant
coke is accompanied by corresponding decreases in gases, light cycle
oil, and main column bottoms. Although minor amounts of scatter are
observed, the FCC yields are mainly a function of FCC conversion,
independent of hydrotreating severity (Figures 1 - 7), indicating
that denitrogenation is not, in itself, the most significant yield
determining factor in the process.
Hydrotreating serves the dual purpose of removing nitrogen
compounds, which act as temporary catalyst poisons, and saturating
aromatics (including many containing nitrogen compounds) which act
as coke precursors. Hydrotreating also reduced the CCR from 0.8% in
the raw feeds to 0.1-0.2%. The removal of CCR reduces the
"additive" coke formed in the FCC, that is, the coke that is
deposited on catalyst directly from CCR in the feed. Additive coke,
defined as 85% of the CCR, does not, however, make a major
contribution to the overall coke make nor does it explain why the
coke yield is much higher for the raw feeds than the hydrotreated
feeds. The removal of coke precursors, by saturating aromatics and
removing nitrogen, is probably the more significant reason for the
reduced coke make (or conversely, greater conversion at constant
coke). In general, increased denitrogenation is accompanied by
increased aro~atics saturation and, thus, at higher denitrogenation
the overall coke make is less, despite similar additive coke yields.
While nitrogen-containing compounds can be considered as
coke precursors (most are contained in aromatic rings), the major
effect of nitrogen, reducing catalyst activity, can be seen more
clearly in Fig. 2 which is a plot of FCC conversion against the
catalyst:oil ratio at differing denitrogenation levels.
As more nitrogen is removed by hydrotreating, less active sites on
the cracking catalyst are poisoned by basic nitrogen, thus, allowing
greater conversion at a given catalyst:oil ratio.
Hydrotreating severity obviously has a large effect on
conversion at constant coke, but surprisingly FCC yields can be
described solely as a function of FCC conversion, independent of
'~U'~
F-4139 --21--
hydrotreating severity except at conversions approaching 100-CA,
where CA is the aromatic carbon in the feed when the gasoline
yield drops off due to overcracking. Case J, with high
denitrogenation but relatively little aromatics saturation at the
lower hydrotreating pressure, best exhibits this trend.
Hydrotreating Pressure Effect
Although high and low pressure cases are grouped together
in the above section, differences in FCC performance for feeds
hydrotreated to the same nitrogen level, but at different
hydrotreater pressures were noted. The difference can best be seen
when comparing Cases B and C at 60% denitrogenation and Cases G and
J at 80% deni~rogenation. Case B and G are hydrotreated at 1750
psia H2, Cases C and J at 1250 psia H2. At constant coke, the
differences between Cases B and C indicate a relatively small effect
of pressure, as shown in Figures 3A and 3B which show the coke and
gasoline selectivities at differing pressures and hydrotreating
severities.
These studies show that hydrotreating the FCC feed greatly
increases the FCC conversion and gasoline yield at constant coke.
In general, the benefits of hydrotreating increase with increasing
hydrotreater severity, i.e. denitrogenation and at higher
denitrogenation severity, the benefits of high pressure operation of
the hydrotreater are most significant in terms of superior
conversion and gasoline yield.
At modest denitrogenation levels, the degree of aromatics
saturation is nearly the same (CA is 19% for Case B; 21% for Case
C) but at higher hydrotreating pressures, FCC conversion and
gasoline selectivity increase notably: Case C at 80%
denitrogenation shows an increase in FCC conversion and gasoline
selectivity at constant coke relative to the 60% denitrogenation
cases. Case J at the lower hydrotreating pressure, conversely shows
a decrease in conversion and gasoline selectivity for the same
increase in denitrogenation severity. This phenomenon can easily be
~3~2;~
F-4139 --22--
explained by the increase in aromatics content in the Case J feed
(CA= 26%) at the high HDT temperature required to achieve 80%
denitrogenation at 1 LHSV. Case G, on the other hand, showed
additional aromatics saturation (CA= 18~) at its milder
s hydrotreating temperature. Case J also shows a slight shift in
yields at constant conversion relative to the other cases, probably
because conversion for Case J ap~roaches 100-CA. The difference
in aromatics saturation, and subsequently, in FCC g~soline
selectivity conclusively establishes higher pressure, and higher
severity, hydrotreating as the preferred conditions.
Example_2
The base feed was subjected to hydrotreating and catalytic
cracking as described in Example 1 at three temperature ranges:
345-365C to simulate SOC
390 to simulate mid cycle
410C to simulate EOC
Space velocity was varied at each combination of temperature and
pressure to cover the range of denitrogenation projected for the
design. Once-through hydrogen was utilized with most data collected
at 535 n.l.l. (3000 SCF/Bbl) hydrogen circulation.
Three different small scale (100cc catalyst volume) fixed
bed down-flow isothermal reactors were used in this study. The
catalyst used was the same as in Example 1, presulfided with 1.38% S
SRGO at 2860 kPa (400 psig) and temperatures of 230 to 345C (450
to 650F). The results are discussed below.
The results showed that at SOC temperatures (345-365C)
the hydrotreating pressure has no significant effect on the yield of
345Cl products from the cracker but that at the higher
hydrotreating temperatures, higher operating pressures reduce
conversion in the hydrotreating step. For example, at the EOC
temperature of 410C and at 75% denitrogenation, an advantage of
about 5 weight percent in the 345Ci yield is observed at the higher
pressure of 12065 kPa (1750 psig) as compared to operation at 8650
~3~23Z~3
F-4139 --23--
kPa (1250 psig). This, in turn, causes higher cycle average yield
of FCC feed as well as a higher quality FCC feed. Because the use
of the higher hydrogen pressure tends to reduce hydrotreating
conversion, severities in excess of 90% denitrogenation are needed
s at 12065 kPa before extensive hydrocracking occurs.
Below 80% denitrogenation, the 250C+ (480F~) yields were
more sensitive to pressure than the 345CI (650F) yields. At low
SOC temperatures, there is no apparent effect of pressure, whereas
at 390 (730F) and 410C (790F), the 250C+ (480F~) yields remain
greater than 90 wt percent at all reasonable severity levels but
decrease significantly at lower pressures. The effect of operating
pressure on light gas make was apparent even at 365C (690F) and
becomes more pronounced as temperature was increased. At 410C
(790F) and 75% HDN, more than twice as much light gas production
was observed at 8720 kPa (1250 psig) than at 12065 kPa (175~ psig).
This is inefficient use of hydrogen in the hydrotreater. Since
overall hydrogen consumptions are only a weak function of operating
pressure, less hydrogen is being utilized to upgrade the FCC feed at
constant denitrogenation and the lower operating pressures. The
lower quality 345Cl (650F+) hydrotreated product was also
reflected in FCC performance as lower conversion and gasoline yield.
Incremental C5-250C (C5-480F) naphtha was not a
function of operating pressure at 345 (650F) and 365C (690F) but
was apparent at 390C (730F) and is significant at 410C (770F).
This is caused by the "rolling down" of higher boiling hydrocarbons
as conversion of 345CI (650F~) increases with hydrotreating
temperature, particularly at the lower pressures. At 12065 kPa
(1750 psig) and 75% denitrogenation yield of C5-250C (C5-480F)
remains essentially constant at 8-10 wt % as temperature increases
from 345C to 410C (650 to 770F). Over this same temperature
range, 8720 kPa (1250 psig) operation gives an increase from 8 to 16
wt % of C5-250C (C5-480F).
The yield of 250-345C (480-650F) distillate is less
sensitive to pressure. No effect is seen at 345C (650F) and 365C
~3~32g
F-4139 --24--
(690F), and differences of only about 2 wt % yield are observed at
390C (735F) and 410C (770F). This means that the increasing
spread in 345C+ (650F~) yields between the two pressures as
temperature increases is realized mainly as 250C- (480F-)
product. Thus, there is a direct loss of FCC feed regardless of
whether 250CI (480F+) or 345C+ (650F+) is sent to the cracker.
Hydrogen Consumption
Hydrogen consumption was found to be a function of
temperature as well as denitrogenation level except at 12065 kPa
(1750 psig) where a temperature effect is not apparent, probably due
to thermodynamic limitations in saturating aromatics at the lower
pressure levels.
The composite curve for hydrogen consumption at all
temperatures and pressures (Figure 4) indicated 133 n.l.l. 1 (750
SCF/Bbl) consumption at 60~ hydrodenitrogenation and 153 n.l.l. 1
(860 SCF/Bbl) at 75~ hydrodenitrogenation. A rapid increase in
consumption is observed above 90% hydrodenitrogenation where
significant hydrocracking can occur.
