Note : Les descriptions sont présentées dans la langue officielle dans laquelle elles ont été soumises.
1'~
Prior Art
There are a number of continuous cyclical proc-
esses employing fluidized solid techniques in which carbon-
aceous materials are deposited on the solids in the reac-
tion zone and the solids are conveyed during the course
of the cycle to another zone where carbon deposits are at
least partially removed by combustion in an oxygen-contain-
ing medium. The solids from the latter zone are subse-
quently withdrawn and re-introduced in whole or in part
to the reaction zone. Among such processes are fluid
coking, fluid dehydrogenation, and fluid catalytic
cracking. -
One of the more important processes of this
nature is the fluid catalytic crackLng process for the
conversion of relatively high-boiling hydrocarbons to
lighter hydrocarbons boiling in the heating oil and
gasoline (or lighter rang~. The hydrocarbon feed is
contacted in one or more reaction zones with the partic-
ulate cracking catalyst maintained in a fluidized s~ate
under conditions suitable for the conversion of hydro-
carbons. ¦ ~ -The gaseous effluent from the reaction zone is ~ ¦
passed to a product recovery zone while the ca~alyst is
generally passed to a stripping zone for removal of
strippable hydrocarbons from the particles. The
stripped catalyst is subsequently introduced into a
fluidized regeneration zone where non-strippable
carbonaceous material is contacted with an oxygen-
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containing gas, for example air, under conditions such
that a major portion of the carbon on the catalyst
particles is removed there~rom by combustion.
Generally, the regeneration is done in a
regeneration zone comprising one or more dense beds
locaked in the bo~tom portion of the regeneration zone
a~d a single rather large disengaging space which is
positioned above and in connection wi~h the dense bed.
Provisions are made ~or recovering and returning
1~ catalyst entrained in the flue gas effluent passing
from the dense bed. This is generally accomplished by
passing this effluent flue gas containing entrained
catalyst through cyclones located in the disengaging
space.
Superficial velocities within such
regeneration zones are generally within the range of -
about 0.4 to 2 metre per second with 0.4 to 1 metre
per second being the more common range. Residence time
of the catalyst within the regeneration zone is
generally in the 2 to 5 minute range with 2 to 3 being
the more com~on, while thè residence time of gas is
generally within the range of 10 to 20 seconds. Such
regeneration zones having generally the configuration
discussed above and the superficial velocity limitation
are referred to herein as conventional regeneration
zones,
~ It is the present refinery practice to
operate such conventional regeneration zones to
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essentially preclude conversion of CO to CO2. The
reason for this practice is to avoid possible thermal
damage to separation means located in the disengaging
space where there is lit~le heat sink to absorb the
5 - heat of reac~ion. Specifically, it is the usual
practice to control the oxygen-containing gas stream
introduced to such regeneration zone directly
responsive to a small predetermined temperature
differential between the regeneration gas outlet,
or the disengaging space and the dense bed of the
regeneration zone to minimize excess oxygen therein
and to essentially eliminate afterburning of CO to CO2
in the upper disengaging portion of the regeneration
zone. Such practice produces a small amount of oxygen
in the flue gas, generally in the range of~0.1 to 1%
oxygen. This practice is exemplified by Pohlenz U.S.
Patents 3rl61,583 and 3,206,393. Present industry
practice is to direct the flue gas which contains
approximately equal molar amounts of CO2 and CO, either
directly to the atmosphere or to a CO boiler where it
is used as fuel to make steam. In other applications,
such as disclosed in Campbell U.S. Patent 3,363~993,
the flue gas is fired as fuel in a heater which is used
to preheat the fresh feed charged to the reaction zone
of the fluid catalytic cracking process.
In fluid catalytic cracking processes com-
prising a reaction zone and a regeneration zone, fresh
catalyst must be added periodically or continuously
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iO7881~5
to replace catalyst which has been lost from the process
or which has become catalytically inactive. Although
the efficiency of cyclones and other such equipment ~or
the recovery of solid catalyst particles is usually
very high, some catalyst is always physically lost
from the regeneration zone. With time, catalyst within
the system loses activity and there~ore b~comes
effectively lost because the cumulative effects of
ex~osure to conta~inant metals contained in the fresh
feed, high temperatures, and steam. For these reasons,
it is necessary or desirable to add fresh make-up
catalyst to maintain the desired total inventory at the
desired equilibrium activity of the catalyst, Typical
daily make-up catalyst rates, for a fluid catalytic
cracking process, are between about 0.5 to 2.0% of
total catalyst inventory in the system with the average
being about 1% of inventory per day. Because the
aforementioned catalysts are relatively expensive, it
is evident that processes requiring large inventories
and therefore large catalyst make-up rates tend to be
less attractive economically. As a result every effort
is made to reduce the initial catalyst inventory
investment and that portion of overall operating costs
associated with catalyst make-up.
Since most of the catalyst inventory is
containea within the regeneration zone, it is the
present practice to employ operating conditions in the
regeneration zone that favor high ~arbon burning rates,
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thereby permitting lower regenerator inventories. Such
desired operating conditions are high partial pressure
of oxygen and tempera~ure. There have been, therefore,
recent industry trends toward higher pressure and higher
temperature regenerators ~or this reason of obtaining
higher burning rates. Previously the preferred pres-
sure range had been from about 10 to about 25 psig. and
preferred temperatures had been in the 1100-1150 F.
range. Pressures in the range of 3-4 atm. and temper-
atures in the range of 621 to 677C. or higher are now
rather common. Although some inventory reduction has
been achieved, limitations on these approaches have -
been imposed by higher equipment costs due to hig~er
pressures, by increased catalyst deactivation due to
long residence time in the high temperature regener-
ators, and by the terminal velocity above which the
catalyst cannot be maintained in the lower part of the -
regeneration zone as a dense bed.
Further attempts to reduce catalyst
inventories and make-up rates have been made by
effecting a staged regeneration within a regeneration
vessel. U.S. Patents 3,494,858 (E. C. Luckenbach~ and
3,563,911 (R. W. Pfeiffer and L. W. Garret, Jr.) are
such examples.
U.S. Patent 3,494,858 discloses a counter-
current regeneration process in which spent catalyst
is partially regenerated in a first fluidized bea with
partially spent regeneration gas, further regenerated
- - -
~l~78~3{)5 .
in a transfer line regeneration zone with fresh regener-
ation gas and then passed to a second fluidiæed bed
wherein further regeneration may or may not take place
with partially spent regeneration gas. Preferred super-
ficial velocities are in the xange of 0.7 to 0.9 m/sec.
and preferred temperatures are about 5g3 to 635C. No
afterburning of CO to CO2 is mentioned, but "substan-
tially complete" removal of carbon to a level below
0.2~ and possibly as low as 0.1~ is contemplated.
Substantially completely regenerated catalyst may then
be stripped of high oxygen-containing gases in a
separate stripper zone with steam or flue gas.
U.S. Patent 3,563,911 discloses a two-stage
regeneration process in which spent catalyst is
partially regenerated in a first dense bed with a first
oxygen-containing gas stream and then further regener-
ated in a seccnd dense bed ~ith a second ox~ygen-
containing gas stream. A common dilute phase i9 super-
imposed above both dense beds. Preferably, superficial
velocities are maintained in the range of about 0.6 to
1.4 metre per second and the preferred temperature
range is from about 607C. to 732C. It is desirable
in the process of this invention to control by means
of a flue gas and last bed temperature differential the
amount of oxygen-containing gas admitted to the regen- -
eration zone such that only a small amount of CO after-
f burning takes place. This is consistent with present
industry practice on single-stage regeneration zones.
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Examples are presented which indicate that at gas
velocities of 0.7 and 1.4 me~re per second, some degree
of inventory reduction over that of a particular type
of single-stage regeneration can be achieved by staging.
The process of our invention employs higher
velocities and oxygen con~entrations than those
presently used in conventional regeneration processes.
By admitting fresh regeneration gas to provide or the
essentially complete combustion of CO rather than the
present practice of limitin~ the fresh regeneration gas
to essentially preclude afterburning, higher oxygen
concentrations result and higher temperatures and coke
burning rates are obtained. Catalyst regeneration can
therefore be completed in a shorter amount of time. By -
employing in the process of our invention higher veloc-
ities than those presently used, catalyst is transported
from a first dense bed to a dilute phase riser and then
to a second dense bed. Thus dramatic catalyst inventory
reductions are therefore possible because of both the
higher oxygen concentration and the higher superficial
velocity. Additional advantages are lower catalyst
make-up rates and improved regeneration and catalyst ;
stability. Additionally, the combustion of CO in a
manner to recover at least a portion of the heat of -
combustion of CO eliminates an air pollution problem
without the need for a CO boiler and also permits
reduced feed preheat requirements.
It is an additional feature of our process
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that the fluid catalytie cracking catalyst contains
catalytically effective amounts o~ a CO eonversion
promoter. The use of the promoter permits either the
same rate of CO conversion to occur at a temperature
as much as 55C. or more lower than that required with
no CO conversion promoter or a faster rate of CO
conversion to occur at a particular temperature than
that which would occur at the same temperature without
the use of a CO eonversion promoter. It is this latter
advantage whieh is of partieular commereial importance.
Without a CO conversion promoter, uneven dispersion of
fresh regeneration gas within the dense-phase eatalyst
bed often requires higher regeneration zone tempera-
tures or higher fresh regeneration gas rates than
desired to maintain a sufficiently fast rate of CO
conversion so that essentially complete converslon of
CO takes place within the regeneration zone. Increas-
ing the regeneration zone temperature may require that
torch oil be burned in the regeneration zone or may
require that increased amounts of slurry oil be
recycled back to the hydrocarbon reaction zone so that
the spent catalyst will contain more coke which can be
burned in the regeneration zone to increase the tempera-
ture. Increased fresh regeneration gas rates, besides
~5 using blower capacity, often overloaded cyclone
separation devices and produced hlgher amounts of flue ,
gas particulate emissions (catalyst) than allowed by
air pollution regulations. The use of the CO conver-
.
~\
1C3 7~38~5
sion promoter permits the elimination of torch oil or
increased slurry oil recycle rates and a reduction in
the amount of excess fresh regeneration gas and thus
gives back to the re~iner more FCC process flexibility.
U.S. Patent 2,436,927, assigned to the
predecessor of my assignee, has recognized that axides
of metals fx`om the first transition series of the
Periodic Table comprising copper, chromium, manganese,
cobalt and nickel may be either included as a component
of the cracking catalyst or supported on a suitable
carrier and employed as catalysts for the oxidation of
carbon monoxide in the regeneration zones of fluid
catalytic cracking processes. That patent is concerned
with a method of preventing "af-terburning" by
eliminating the possibility of combustible mixtures of
oxygen and carbon monoxide occurring in the light or
dilute phase region of the regen~ration zone. This is
the antithesis of the use to which such catalysts is put
to in the present process. In the process described in
U.S. Patent 3,808,121, coke and CO oxidation are
accomplished in an FCC regeneration zone by employing
two separate catalysts of different particle size and
composition: a hydrocarbon conversion catalyst and a
CO oxidation catalyst. In that process the CO
oxidation catalyst is maintained within a conventional-
type regeneration zone and does not pass out o~ that
zone to the hydrocarbon reaction zone. In contrast,
the present process employs one catalyst, possessing
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1C3'781~0~ j
both hydrocarbon conversion and CO conversion capabil-
ities, in a unique flow arrangement. We have discov-
ered that oxides of metals selected from the group
consisting of the noble and some of the non noble metals
when incorporated on or into an FCC catalyst catalyze
essentially complete conversion of CO to CO2 in the
dilute-phase transport riser which is provided in our
process for such conversion.
SUMMARY OF THE INVENTION
It is accordingly, a broad objective of the
process of our invention to provide a regeneration proc-
ess for the oxidation of coke from a spent fluidizable
catalyst and for the essentially complete catalytic
L5 conversion of CO, produced by the oxidation of coke,
to CO2 to produce regenerated catalyst and spent
regeneration gas.
In brief summary, our invention is, in one
embodiment, a process for the regeneration of a coke-
contaminated fluid catalytic cracking catalyst removed
from a hydrocarbon reaction zone and containing cata-
lytically effective amounts of a CO conversion promoter,
and the catalytic conversion of carbon monoxide, result-
ing from the oxidation of coke from said catalyst, to
carbon dioxide which process comprises the steps of:
passing the catalyst and a fresh free oxygen-containing
regeneration gas to a first dense bed of fluidized
particulate catalyst in a regeneration zone and
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oxidizing coke in the first dense bed maintained at
- oxidizing conditions to produce partially regenerated
catalyst, containing said CO conversion promoter, and
partially spent regeneration gas con~aining CO; passing
the partially regenerated catalyst and partially spent
regeneration gas to a dilute phase transpor-t riser and
therein catalytically conver~ing at conversion condi-
tions, including the presence of the CO conversion
- promoter, carbon monoxide to carbon dioxide to produce
spent regeneration gas; separating regenerated cata-
lyst from spent regeneration gas; and, introducing said
regenerated catalyst to a second dense bed of partic- _
ulate materiaL from which the regenerated catalyst is
returned to the hydrocarbon reaction zone.
Other embodiments and objects of the present -
invention encompass details about-catalysts, operating
features and operating conditions all of which are
hereinafter disclosed in the following discussion of
each of these facets of our invention.
DESCRIPTION OF THE DRAWING
Having thus described the invention in brief
general terms, reference is now made to the schematic
drawinqs in order to provide a better understanding of
the present invention.
It is to be understood that the drawings are
shown only in such details as are necessary for an
understanding of the invention and that various minor
items such as valves, bleed and dispersion steam lines,
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107B8VS
instrumentation and other control means have been
omitted therefrom for the sake of simplicity. Further-
more the scope and spirit of the claims appended hereto
are not to be limited thereby.
The drawings of this specification include
Figure 1, which depicts schematically the side view of
a specific apparatus suitablP for carrying out the
process o-E our invention, and Figure 2 and Figure 3,
which depict side views of alternative apparatus also
suitable for carrying out this invention. Figuxes 2 and
3 include the same primary features as Figure l; Figure
3 is of particular interest because it indicates how the
present invention might be applied to an existing
regeneration vessel.
Figure 1 shows a regeneration apparatus 100
basically containing a first dense bed 1, a dilute
phase transport riser 2, a catalyst and regeneration-
gas separation means 3 and 4, a disengaging space 5 and
a second dense bed 6. First dense bed 1 is shown at
the lower portion of the figure and is connected to one
end of a vertically extending dilute-phase riser 2
through a transition region 28. Catalyst and regener-
ation-gas separation means 3 is attached to the outlets
7 of transport riser 2. Separated regeneration gas
leaving separation means 3 passes into disengaging
space 5 then into separation means 4 where it exits at
26 into plenum chamber 27. Regeneration gas leaves the
plenum chamber 27 and the regeneration zone 100 via out-
1071~380:j ~
let 8 and 8'. Separated catalyst from separation means .
3 and 4 is directed toward second dense bed 6.
Spent catalyst withdrawn from a hydrocarbon :
reaction zone (not shown) and containing catalytically
effective amounts at a CO conversion pxomoter is intro- .
duced through inlet line 9 to the first dense bed 1 .
having a level indicated at 10 which is located in a
transition region 28 positioned between the first dense ..
bed 1 and the transport riser 2. Fresh reyeneration gas ..
is introduced via line 11 into dense bed 1 through : :
distributing device 12 which allows the fresh regener-
ation gas to be more readily dispersed within the dense
bed 1. Typically the distributing device can be a -.
metal plate containing holes or slots or preferably can ~
be a pipe grid arrangement, both types of which are
quite familiar to those skilled in the art~ Oxidation
of the carbonaceous deposits takes place in dense bed 1
and regeneration gas and fluidized catalyst are carried
out of bed 1 through the transition region 28 and into
transport riser 2 wherein CO is converted at conve.sion
conditions, including the presence of the CO conversion
promoter, to CO2 and wherein at least a portion of the
heat of combustion of CO is transferred to the catalyst. .
Transport riser 2 is vertically positioned
having its inlet at the lower portion an~ its outlet
means 7 near its top portion. outlet means 7 may be
single or multiple openings located at or near the
uppermost portion of transport riser 2 which allow cata-
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~C~788~5
lyst and regeneration gas to pass out of transport riser
2, As shown in Figure 1, catalyst and regenera~ion gas
separation means 3 is attached to the outlets 7 of trans-
port riser 2. Separation means 3, typically a cyclone
separation means, is used to achieve a substantial .
separation o~ regeneration gas and entrained catalyst
which pass out of transport riser 2. Although the Fig-
ure 1 shows only 1 such cyclone attached to the trans-
port riser 2, it is contemplated that from 1 to 4
cyclones could be so positioned. Entrained catalyst and
regeneration gas pass via outlet 7 into the separation
means 3 out of which regeneration gas substantially free
of catalyst passes through outlet 13 while catalyst ,
passes through diple~ 14 directed toward second dense
bed 6. Optionally separation means 3 could be omitted, .
leaving regeneration gas and catalyst exiting ou~let
means 7 to directly enter disengaging space 5. Some
separation of regeneration gas and catalyst would be
achieved but not as efficiently as with cyclone separa- .
tion means 3 as shown.
Separation means 4, also typically a cyclone .
separation means, has an inlet 16 which receives .
regeneration gas and any entrained catalyst located in
the disengaging space 5. Regeneration gas and any
entrained catalyst are substantially separated from :
each other with the regeneration gas passing out of the
separation means 4 at outlet 26 into plenum chamber 27, :
and then out of the regeneration zone via regeneration .
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1~378~3~5 -
.
gas outlet 8 and 8'. Catalyst separated from the
regeneration gas is passed via dipleg 17 downward
toward second dense bed 6 having a level as indicated at
15. Although only one such separation means ~ is shown,
obviously more than one could be employed.
An external combustible fluid such as ~uel gas
or a liquid hydrocarbon stream may be admitted to the
transport riser 2 through line 20 via distributor 21.
The burning of such a fluid may be necessary to assist
in the startup of the process or to increase t~e temper-
ature within the dilute phase transport zone 2 suffi-
ciently to initiate CO oxidation or it may be desirable
to increase the temperature of the catalyst particles
passing through the riser.
Additionally, a second stream of fresh
regeneration gas may be admitted to the transpor~ riser
throu~h line 18 via distributor 19 for the purpose of
supplying needed oxygen to support burning of the
external combustible fluid.
Catalyst passing through diple~s 14 and 17
discharges in a downward direction toward the second
dense bed 6.
Second dense bed 6 is so positioned in
relationship to the first dense bed 1 as to maintain
a sufficient-head of re~enerated catalyst necessary to
overcome the sum of the pressure drop in regeneratea
catalyst exit line 22 and control valve 23 plus the
pressure maintained in the hydrocarbon reaction zone
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(not shown) which is connected to line 22. Although
Figure 1 shows the level 15 of second dense bed 6
positioned above the level 10 of irst dense bed 1, the
beds may be placed at other locations with resp~ct to
each other so long as this head is provided. Further-
more, the level 15 of catalyst within the second dense
bed 6 may be so controlled as to provide the desired
residence time within the bed.
It should be noted that second dense bed 6
need not be a stripper as indicated in Figures 1, 2 and
- 3. The te~m second dense bed as used in this specifi-
cation means regenerated catalyst maintained in dense
phase for head and sealing purposes prior to being
returned to the hydrocarbon reaction zone.
The regenerated catalyst in second dense bed
6 moves in a downward direction and at least a portion
eventually passes out of the regeneration vessel and is
returned to the hydrocarbon reaction zone via conduit
22. Additionally, a portion of the regenerated catalyst
can be returned to first dense bed 1 to increase the
temperature of that bed. Also located on conduit 22
is a valve 23 which may be used to control the rate of
withdrawal of regenerated catalys~ from bed 6. Typi-
cally, valve 23 is a slide valve and is operated by a
reactor temperature controller or level controller.
A stripping medium may be admitted to the
second dense bed 6 through line 24 via distributor pipe
25 to strip from the regenerated catalyst adsorbed and
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.
i
interstitial re~eneration g~s. Generally the stripping
medium ~7ill be superheated steam.
It is anticipated in the process of this
invention that most of ~he catalyst within the regener-
ation zone will be contained in the first dense bed with
the smaller portion contained in the second dense bed.
More specifically, when skeam stripping is employed
within the second dense bed 6, the second dense bed
volume will be so designed such that the catalyst resi-
dence time within the bed is less than one minute and
preferably less than 30 seconds.
Shown in Figure 2 is the side view o~ an
alternate apparatus 200 in which this invention may be
practiced. Primary components are: a first aense bed
201, a dilute phase transport riser 202, a transition
region 228, a catalyst and regeneration gas separation
means 203 and 203' and 204, a disengaging space 205, and
a second dense bed 206. Pirst dense bed is again sho~m
at the lower portion of the figure and is connected to
one end of a vertically extending dilute phase riser 202
through a transition region 228. Catalyst and regener-
ation gas separation means 203 and 203' are attached to
outlets 207 and 207' of transport riser 202. Separated
regeneration gas leaving separation means 203 and 203'
via outlets 213 and 213' passes into disengaging space
205, then into separation means 204, and finally exits
the regeneration zone via line 208. Separated catalyst
from separation means 203 and 203' and 204 is directed
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1~78~305
via diplegs 214, 214' and 217 to second dense bed 206.
Spent catalyst withdrawn from a hydrocarbon
reaction zone and containing catalytically effective
amounts of a CO conversion promoter is introduced
through inlet line 209 to the first dense bed 201.having
a level indicated at 210 which is within the transition
region 228. Fresh regeneration gas is introduced via
line 211 into dense bed through distributing device 212
which allows the fresh regeneration gas to be more
readily dispersed within the dense bed 201. Typically
the distributing device can be a metal plate containing
holes or slots or preferably can be a pipe grid arrange-
ment, both types of which are quite familiar to those
skilled in the art. Oxidation of the carbonaceous
deposits takes place in dense bed 201 and regeneration
gas and fluidized catalyst are carried out of bed 2û1 :;
through the transition region 228 and into transport
riser 202 wherein CO is converted in the presence of a
CO conversion promoter and wherein at least a portion of
the heat of combustion of CO is transferred to the
regenerated catalyst.
Transport riser 202 is vertically positioned
having its inlet at the lower portion and its outlet ::
means 207 and 207' near its top portion. Outlet means - :
207 and 207' may be openings located at or near the
uppermost portion of transport riser 202 which allow
catalyst and regeneration gas to pa,s out of transport
riser 202. As shown in Figure 2 catalyst and regener-
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10~880~
ation gas separation means 203 and 203' typica~ly
yclones, are attached to the outlets 207 and 207' of
transport riser 202 and are used to achieve a substantial
separation of regeneration gas and entrained catalyst
passing out of transport riser 202 so that the material
present in the disengaging space 205 i5 essentially
regeneration gas with very little entrained catalyst
present. Entrained catalyst and regeneration gas pass
from riser 202 via outlets 207 and 207' into the
separation means 203 and 203' where regeneration gas
substantially free of catalyst passes out of the
separation means through outlets 213 and 213' and
catalyst passes through diplegs 214 and 214' in a
downward direction toward second dense bed 206 having
a level or interface 215. Preferably diplegs 214 and
214' pass into second dense bed 206 below the second
dense bed catalyst level or interface 215. As men-
tioned in the Figure 1 description, separation means
203 and 203' could be omitted, leaving regeneration gas
and catalyst exiting outlet means 207 and 207' to
directly enter disengaging space 205. Some separation
of regeneration gas and catalyst would be achieved but
not as efficiently as with cyclone separation means 2Q3
and 203' as shown.
2~ Separation means 204, also typically a cyclone
separation means, has an inlet 216 which receives
regeneration gas and any entrained catalyst located in
the disengaging space 205. Regeneration gas and any
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entraine~ catalyst are substantially separatea from each
other with the regeneration gas passing out of the
separation means 204 and out of the regeneration zone
via regeneration gas outlet 208. Catalyst separated
from the regeneration gas is passed via dipleg 217
down to the second dense bed 205.
An external combustible fluid such as fuel
gas or a liquid hydrocarbon stream may be admitted to
the transport riser 202 through line 220 via distributor
221 for the purpose of increasing the temperature of
the regeneration zone upon initial startup or i~creasing
the temperature within the dilute phase transport zone 2
sufficiently to initiate C0 oxidation, or for increasing
the temperature of the catalyst particles passi~g
through the riser. Additionally, a second stream of
fresh regeneration gas may be admitted to the transport
riser through line 218 via distributor 219 as in Figure :
1, for the purpose of supplying needed oxygen to support ~ ~
burning of the external combustible fluid. ~ ~:
Catalyst passing through diplegs 214 and 214'
- and 217 discharges in a downward direction toward the
second dense bed 206.
This second dense bed 206 is so positioned in
relationship to the first dense bed 201 as to ma~ntain
a sufficient head of regenerated catalyst necessary to
overcome the sum of the pressure drop in regenerated
catalyst exit line 222 and control valve 223 plus the
pressure maintained in the hydrocarbon reaction zone
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88~S
(not sho~n) ~7hich is connected to line 222. Although
- Figure 2 shows the second dense bed 206 positioned above
the first dense bed 201, the beds may be placed in other
locations with respect to each other ~s lon~ as this
head is provided. As a consideration independent of the
head requirement the level 215 of the dense bed can be
controlled as required to provide the desired catalyst
residence time within the bed.
The regenerated catalyst in second dense bed
206 moves in a downward direction and eventually at
least a portion of the catalyst passes out of the
regeneration vessel and is returned to the hydrocarbon
reaction zone via conduit 222. Additionally, a portion
of the regenerated catalyst can be returned from dense
bed 206 to dense bed 201 to increase the temperature in
bed 201. Also located on conduit 222 is a valve 223
which may be used to control the rate of withdrawal of
regenerated catalyst from bed 206. Typically valve 223
is a slide valve and as mentioned is generally operated
by a reactor temperature controller or level controller.
A stripping medium may be admitted to the
second dense bed 206 through lines 224 and 224' via
distributors 225 and 225' to strip from the regenerated
catalyst adsorbed and interstitial regeneration gas.
Generally the stripping medium will be superheated - :.
steam.
As pr~viously mentioned, it is anticipated in
the process of this invention that most of the catalyst
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~ 7 ~ ~ 5
within the regeneration zone will be contained in the
first dense bed with the smaller portion contained in
the second dense bed. More specifically when steam
stripping is employed within the second dense bed, the
second dense bed volume will be so designed such *hat the
catalyst residence time within the bed is less than 1
minute and preerably less than 30 seconds.
Figure 3 shows the side view of another
alternate apparatus 300 suitable for practicing this :
invention. Specifically, Figure 3 represents a possible
modification of existing regenerators to produce an
apparatus sui~able for carrying out the process of this ¦
invention. Basically an existing regeneration vessel ¦
303 is modified to provide a first dense bed 301, a !~
dilute phase transport riser 302, a second dense bed 306, 1 -
and the associated equipment for riser 302 and bed 306. :
The portions of Figure 3 generally function the same as
those previously described for Figure 1 and Figure 2.
Briefly, spent catalyst withdrawn from a
hydrocarbon reaction zone is introduced through înlet
line 309 into a first dense bed o~ catalyst 301 having
a level or interface indicated at 310 located within
transition region 328 positioned between first dense
bed 301 and transport riser 302. Fresh regeneration gas
is introduced at line 311 into dense bed 301 through
distributing device 312. Oxidation of coke contained on
the catalyst takes place in dense bed 301 and regener-
ation gas and fluidized catalyst are swept out of bed
, .
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iC~7880S
301 through transition region 328 and into transport
riser 302 wherein CO oxidation takes place in the
presence of the catalyst containing a CO conversion
promoter and wherein at least a portion o the heat of
S combustion of Co is transferred to the catalyst. -
A combustible fluid may be admikted to the
transport riser 30~ through line 320 via distributor ¦
321 and additionally a second stream of fresh regener-
ation gas may be admitted to the riser through line 318
via distributor 319 for reasons previously described.
Entrained catalyst and regeneration gas then -
pass out of transport riser 302 via riser outlet means
307 into disengaging space 305. Preferably outlet means
307 is so placed that it ejects entrained catalyst and
regeneration gas in a downward direction to reduce the
extent of catalyst entrained within the disengaging
space 305.
Separation means 304, typically a cyclone
separation means, has an inlet 316 and receives regener~
ation gas and any entrained catalyst from the disen-
gaging space 305. Regeneration gas and any entrained
catalyst are substantially separated from each other
with the regeneration gas passing out of the separation
means 304 and out of the regeneration zone 303 via out-
let 308. Separated catalyst is passed via dipleg 317
in a downward direction t~ward a second dense bed 306.
Interface 315 defines the boundary between the second
dense bed 306 and the disengaging space 305.
-24-
1~7~
:
Regenerated catalyst in bed 306 moves in a
do~nward direction and at least a portion of the catalyst
leaves the bed and the regeneration zone through conduit
322 and is returned to the reaction zone. A portion of
the regenerated catalyst can also be returned from dense
bed 306 to dense bed 301 to increase the temperature in
dense bed 301. The rate of catalyst with~raw is
controlled by valve 323 which is generally opexated by ;
a reactor temperature or level contro31er.
The stripping medium may be admitted to the
second dense bed 306 through line 324 via distributor
325 to effect stripping of regeneration gas from regener-
ated catalyst.
DESCRIPTION OF T~IE INVENTION
,
At the outset, the definition of various terms
used herein will be helpful to an understanding of the
process of our invention.
The term "afterburning" as generally under-
stood by those skilled in the art means the incomplete
oxidation of CO to CO2 within the regeneration zone or
the flue gas line. Generally afterburning is charac-
terized by a rapid temperature increase and occurs dur- -~
ing periods of unsteady state operations or process
"upset". It is, therefore, usually o~ short duration
until steady state operations are resumed.
In contrast to afterburning, the term
"essentially complete conversion of CO" shall refer to
.'
-25-
- ~
:
380~ ~ ~
the intentional, sustained, controlled, and essentially
complete combustion of Co to CO2 within the regeneration ! :
zone. "Essentially complete" shall mean that the CO
concentration is the spent regeneration gas (hereinafter
defi~ed) has been reduced to less than about 1000 ppm
and more preferably less than 500 ppm.
The term "spent catalyst" as used in this
specification means catalyst withdrawn from a hydro-
carbon conversion zone because of reduced activity
caused by coke deposits. Spent catalvst passing into
the regeneration zone can contain anywhere from a few
tenths up to about 5 wt. % of coke, but typically in
FCC operations spent catalyst will contain from about -
0.5 to about 1.5 wt. % coke. -
The term "regenerated catalyst" as used in
this specification shall mean catalyst from which at -~
least a portion of coke has been removed. Regenerated
catalyst produced by our process will generally contain
less than 0.5 wt. ~ coke and more typically will contain
from about 0.01 to about 0.15 wt. % coke.
The term "regeneration gas" as used in this
specification shall mean, in a generic sense, any gas
which is to contact catalyst or which has contacted
catalyst within the regeneration zone. Specifically,
the term "fresh regeneration gas" shal1 includs free-
oxygen-containing gases such as air or oxygen enriched
or ~eficient air which pass into the regeneration zons
to allow oxidation of coke from the spent catalyst and
-26-
~\ : ::
~C~78~(35 - `
essentially co~plete conversion of CO. Usually the
fresh regeneration gas will be air. Free-oxygen shall
refer to uncombined oxygen present in a regeneration gas.
The term "partially spent regeneration gas"
shall refer to regeneration gas which has contacted
catalyst within the dense-phase bed and which contains a
reduced quantity of free-oxygen as compared to resn
regeneration gas. Partially spent regeneration gas will
generally contain several volume percent each of nitro-
gen, free-oxygen, carbon monoxide, carbon dioxide, and
water. More speciically, the partially spent regener- --
ation gas will generally contain from about 7 to about
14 vol. % each of carbon monoxide and carbon dioxide.
The term "spent regeneration gas" shall mean
regeneration gas which contains a reduced concentration
of CO as compared to partially spent rege~eration gas.
Preferabl~ the spent regeneration gas wi l contain less
than about 1000 ppm of CO and more typically and prefer-
ably less than about 500 ppm CO. Free oxygen, carbon
dioxide, nitrogen, and water will also be present in the
spent regeneration gas. The free-oxygen concentration
of the spent regeneration gas will generally be greater
than 0.1 vol. ~ of the spent regeneration gas.
T'ne terms "dense-phase" and "dilute-phase"
are commonly used terms in t'ne art of FCC to generally
characterize catalyst densities in various parts of the
regeneration zone. While the demarkation density is
somewhat ill-defined, as the term "dense-phase" is used
,
-27-
,~,i
,:J .
' ' ' .
1~788~5
herein, it shall refe~ to regions within the re~ener-
ation zone where the catalyst density is greate~ ~han
about 80 kg~m3 and as "dilute-phase" is used herein it
refers to regions where the catalyst density is less
than about 80 kg/m3. Usually the dense-phase density
will be in the xange of from about 80 to about 560 kg/m3
or more and the dilute-phase density will be much less
than 80 kg/m3 and in the range of from about 1.6 to
about 80 kg/m3. Catalyst densities within FCC vessels
are commonly measured by measuring pressure or head
differences across pressure taps installed in the ves-
sels and spaced at known distances apart. -
The unique features and advantages of the
process of our invention will be made clearer by refer-
ence first to a typical fluid catalytic cracking process
with particular emphasis on the features of the conven
tional regeneration zone of such process.
In a typical FCC process flow, finely divided
-regenerated catalyst which leaves the regeneration zone
contacts a feed stock in a lower portion of a hydrocar-
- bon reaction zone. While the resultant mixture passes
up through the reaction zone conversion of the feed to
lighter products and to coke deposited on the catalyst
occurs. The effluent from the riser is discharged into
a disengaging space where additional conversion can take
place. The hydrocarbon vapors, containing entrained
catalyst, are then passed through one or more cyclone
separation means to separate any spent catalyst from
-28-
::
:
1078~30S :
:'
., ~ . .
the hydrocarbon vapor stream. The separated hydrocarbon
vapor stream is passed into a fractionation zone known
in the art as the main column wherein the hydrocarbon
efluent is separated into such typical fractions as
light gases and gasoline, light cycle oil, heavy cycle
oil, and slurry oil. Various ractions ~rom the main
column can be recycled along with the ~eed stock to the
reaction riser. Typically, fractions such as light
gases and gasoline are further separated and processed
in a gas concantration process located downstream of
the main column. Some of the fractions from the main
column as well as those recovered from the gas concen-
tration pxocess may be recovered as final product
streams. The separated spent catalyst passes into the
lower portion of the disengaging space and eventually
leaves the hydrocarbon reaction zone after passing
through a stripping means in which a stripping gas,
usually steam, countercurrently contacts the spent cata-
lyst purging adsorbed and interstitial hydrocarbons from
the catalyst. The spent catalyst then passes into a
regeneration zone along with a fresh regeneration gas
stream wherein combustion of coke produces a regenerated
catalyst containing a reduced quan~ity o coke and a
flue gas containing approximately equal amounts of
carbon monoxide and carbon dioxide along with water,
nitrogen, and perhaps a small quantity of oxygen. Typi-
cally the spent catalyst contains from 3.5 up to 1 or
more wt. % coke while the freshly regenerated catalyst
29
'7~ S
contains less than ~bout 0.5 and more typically 0.2 to
0.4 wt. % coke. Usually, the fresh regeneration gas
passed into the regeneration zone is air but in some
instances the air stream may be either enriched or
deficient in oxygen. Flue gas is separated from
entrained regenerated catalyst by cyclone separation
means located within the regeneration zone and separated
flue gas passes from the regeneration zone either to a
carbon monoxide boiler wherein the chemical heat of car-
bon monoxide is recovered outside of the regeneration
zone by combustion as a fuel for the production of steam
or directly to the atmosphere. Regenerated catalyst
which was separated from the flue gas is returned to the
lower portion of the regeneration zone in which is main-
tained a dense bed of catalyst. Regenerated catalyst
leaves this dense bed and as previously mentioned,
contacts the feed stock in a hydrocarbon reaction zone.
Generally, regenerated catalyst is not stripped of
entrained flue gas prior to contacting the feed.
In a typical conventional regeneration zone,
the spent catalyst is introduced into and is maintained
in the bottom portion of the zone in one or more dense
beds by limiting the superficial velocity vf the
incoming fresh regeneration gas. Regenerated catalyst
is withdrawn from the same bottom portion of the regen-
eration zone. The superficial veiocity is limited to
the transport velocity, that is, the velocity past whic~
the catalyst would be carried out of the dense bed to
-30-
~78805
the cyclones. Typical velocities are therefore less than
about 0.9 metre per second with 0.5 to 0.8 bein~ the
usual range.
Most of the total catalyst inventory of a fluid
catalytic cracking process is contained in the regener-
ation zone. In view of the present industry trend to~ard ¦
short contact time hydrocarbon reaction zones, an even
larger percentage ~f the total catalyst inventory is
contained~in-the regeneration zone. Thè determination
of the inventory in a typical conventional regeneration
zone is based upon the feed rate to the FCC process (or
more specifically to the coke yield from that feed rate)
and the superficial velocity. This coke yield antici- I ;
pated from a desired feed rate determines the rate of
the fresh regeneration gas to the regeneration zone.
This gas rate at a limiting superficial velocity then
determines the cross-sectional area of the regeneration
zone. With a known catalyst density and height of the
dense bed, the inventory of the regeneration zone, and
for practical purposes for the FCC process, is fixed.
Catalyst residence times which result for these conven-
tional regeneration zones are generally from about 2 to
5 minutes with about 2 to 3 being the general range.
With the above description as a reference
point, our process can be briefly described as a regen-
eration process for the oxidation of coke from a coke-
contaminated fluid catalytic cracking catalyst and for
the essentially complete catalytic conversion of CO,
-31-
~ .
, : . ~ . - : - . .:
1(:178805
produced by the oxidation of coke, to Co2 to produce
regenerated catalyst and spent regeneration gas in which:
spent catalyst removed from a reaction zone and contain-
ing catalytically effective amounts of a CO conversion
promoter and fres~ regeneration gas pass into a first
dense bed wherein coke is oxidized to produce partially
spent regeneration gas and regenerated catalyst; par-
tially spent regeneration gas and regenerated catalyst
pass from the first dense bed to a dilute phase transport
riser wherein the catalytic conversion of CO to CO2 takes
place in the presence of the CO conversion promoter to
produce spent regeneration gas and wherein at least a
portion of the heat of combustion of CO is transferred
to the regenerated catalyst; regenerated catalyst and
spent regeneration gas leaving the transport riser are
separated; and, the regenerated catalyst passes to a
second dense bed of particulate material from which
regenerated catalys~ is returned to the hydrocarbon
reaction zone.
Thus, the unique features distinguishing the
`process of our invention from conventional regeneration
processes are apparent. One such feature is that the
superficial velocity of the fresh regeneration gas is
not limited, as it is in conventional regeneration
zones, to the critical velocity. Another feature is
that CO, produced by the oxidation of coke, is essen- --
tially completely converted to CO2 an~ ~t least a
portion of the heat of combustion of CO is rec~veEe~
. . ,
-32-
~'7~8~)~ ~
within the regeneration zone by being transferred to the
regenerated catalyst. Furthermore, this C0 conversion
is cataly~ed by a C0 conversion promoter which is an
integral part o the fluid cracking catalyst thereby
permitting a suficiently fast rate oE C0 conversion to
be maintained at a lower temperature or a lo~er resh
regeneration ~as rate or both than required without a C0
conversion promoter. The advantages offered by these
features of our invention are further described below.
Since it is not intended in the process of our
invention that catalyst in the first dense bed remain in
that bed, the superficial velocity of tne fresh regener-
ation gas into that bed is not limited to the critical
velocity. In the first dense bed superficial velocities
will be in the range of about 0.9 to 3 metre per second
so that catalyst can be carried from the first dense bed
into the dilute phase transport riser. Velocities
contemplated for t'ne ailute phase transport riser will
be in the range of about 3 to 7.5 metre per second.
Since velocity is not only not limited to the
critical velocity, but is in fa~ct 2 to 3 times the
critical velocity, dramatic catalyst inventory reduc-
tions can now be achieved. As previously des~ribed,
regeneration 20ne cat-alyst inventories are directly
related to superficial velocities employe~ within the
regeneration zone. Catalyst inventories using the
process of this invention will be about 40 to 60 per- ~
cent of those present single or multistate regeneration -
-33-
iO78805
prQcesses~ As an example, a moderately sized FCC process
of the type presently in industry use will contain about
150 tons of catalyst; by using the regeneration process
of this invention in an FCC process of the same size a
refiner could save the initial investment represented by
at least 75 tons of catalyst.
Catalyst make-up rates requixed to make up
losses and maintain activity will also be reduced because
such rates tend to be a percentage of the total catalyst
inventory. Conversely, as previously mentioned, heavier,
more contaminated feed stocks could be charged to an FCC
process employing this invention without requiring a
catalyst make-up rate in excess of that presently accepted
on an FCC process using conventional regeneration tech-
niques. Feed stocks would no longer have to be limited
to relatively clean vacuum gas oils containing limited
quantities of Conradson carbon, metals, and nitrogen
compounds. Xigher molecular weight feed stocks contain-
ing higher amounts of these contaminants could better be
tolerated with no additional economic penalty. Because
such feed streams require less processing, the refiner
will realize an additional savings.
Better solid-gas contact and reduced residence
times of both catalyst and regeneration ~as are other
advantages of the higher velocity. Higher velocities
will produce more turbulent flow with better mixing and
hence will result in more efficient regeneration. Be-
-34-
~L~37~3~3VS I ;
- cause of better gas-solid contact, higher oxygen partial
pressure, and higher temperatures the rate of coke com-
bustion will be increased. Since coke will now be
.
removed in a shorter amount of time, the catalyst resi-
dence time can be reduced. Catalyst residence times can
be reduced ~rom the present 2 to 5 minutes to less than
2 minutes and regeneration gas residence times can be
reduced from about 20 seconds to less than 10 seconds.
With shorter catalyst exposure time and lower temper-
atures made possible by the use of a C0 conversion pro-
moter catalyst activity will be prolonged and reduced
make-up rates can result.
From the standpoin~ of savings to be effected,
another important result of shorter catalyst residence
time is that it may permit the stripping of flue gas
components from the regenerated catalyst. It is present
industry practice to~ strip only spent catalyst; spent
catalyst from a reaction zone i5 stripped of absorbed
and interstitial hydrocarbons before the catalyst is
sent to the regeneration zone for the purpose of recover-
ing valuable gasoline and light hydrocarbons that would
otherwise be burned in the regeneration zoneO Although
it is known that prolonged exposure to steam will deacti-
vate catalyst, steam is generally the preferred strip-
ping medium. The small amounts used, however, and the
short catalyst residence time and relatively low temper-
ature within the hydrocarbon reaction zone minimize any
deactivation. In spite of the fact that flue gas com--
-35-
\
~'71313~5
ponents are entrained by re~enerated catalyst into the
hydrocar~on reaction zone and hence become part of a
product stream, steam stripping of regenerated catalyst
has generally not been practiced because of the longer
ca~alyst residence time in the regeneration zone and the
large catalyst inventory which is generally contained in
a single dense bed. Exposure of this quantity of catalyst
to steam for this longer period of time would increase
the catalyst deactivation rate. Steam would again be the
preferred stripping medium rather than an inert gas
because it can be condensed and separated before reaching
the product recovery section of the FCC process. The
- following example and more detailed description will make
clear the advantage of stripping regenerated catalyst.
The effluent from the hydrocarbon reaction zone
of an FCC process actually contains not only hydrocarbons
but steam from spent catalyst stripping and flue gas com-
ponents in amounts from about 1 to 3 grams per thousand
grams of catalyst circulated in the FCC process. In a
typical FCC process of modest size about 1,350,000 kg/hr.
of catalyst will be circulated ana will therefore entrain
about 2,000 kg~hr. of flue gas components into the
reaction zone. This means that on a volume basis, the
reactor effluent will contain about 1,600 m3~hr. or
38,400 m3/day of flue gas components to be processed down-
stream of the reaction zone. The total hydrocarbon reac-
tion zone effluent is direated to the main column where
it is separated into gas and unstabilized gasoline as an
-36-
~t~88~5
overhead product stream and various siae cut product
streams. This overhead product stream containing the flue
gas components, light hydrocarbon gases, steam and gaso-
line is directed first to a main column overhead condenser
where steam and gasoline are condensed and then to a main
column overhead receiver.
In order to recover and separate light hydro-
carbons and to stabilize the gasoline, the gas and unsta-
bilized gasoline from the receiver are sent to a gas
concentration process consisting primarily of a compres-
sor, absorber columns, and fractionation columns along
with associated equipment. To recover the light hydro-
carbons, th,e gas stream containing flue gas components
frbm the receiver is first compressed to about 11.2 to
18 atm. before being directed to the absorber columns.
Light hydrocarbons, primarily C3's and C4's, are absorbed
by heavier liquid hydrocarbon streams in the absorber
columns leaving all unabsorbed lean gas containing the
flue gas components as a product stream from the last
absorber. A typical analysis of the off gas leaving the
last absorber is as follows:
i(37B805
Table No. 1
Typical Absorber Off Gas Analysis
Composition ` Mol.
Carbon dioxide 3.8
Oxygen + argon 0.3
Nitrogen 22.8 s
Carbon monoxide 4.5 i~
Hydrogen 8.9
Methane 23.6
Ethylene 11.8
E~hane 12.5
Propylene 4.8
Propane 1.4
Isobutylene ~ l-butene0.7
Cis ~ 2 - butene 0.4
Trans - 2 - butene 0.5
Isobutane 1.7
Norrnal butane 0.5
Isopentane 0.7
Normal pentane 0.2
Total C6~ and C5 olefins 0.9 ..
100.O
As can be seen from the breakdown, the entrained flue gas
components, nitrogen, carbon monoxide, and carbon dioxide
constitute about one-third of this product stream, ' '.. ~.'~
(31.1%~. Because of the short catalyst residence time
employed in the process of our invention, steam stripping
of the regenerated catalyst may be employed whereby most
of this material will be carried out of the regeneration
zone rather than becoming a substantial part of this
product stream. Without these components, considerable -
savings can be realized by using a less expensive gas
plant, that is, one having smaller compressors and
smaller absorbers along with other g~s:ha~d1ing-e~ip-
ment. As well, the absorber off gas which is generally
used as a fuel will have a higher heat content.~ ¦
-3~-
:1.
1~88V5
It is an additional unique feature of the pro-
cess of our invention that C0 produced by the oxidation
` o coke is essentially completely converted to C~2 in a -
dilute phase transport riser and that at least a portion
of the heat of combustion of C0 is transferred to the
catalyst passing through the transport riser to the
second dense bed. Practical economical advantag~s o~ the
feature are that it alleviates an air pollution problem 1-
without the necessity of a CO boiler and that chemical
heat at combustion of C0 is recovered within and used
within the FCC process itself. The hea~ of combustion of
CO is transferred to the catalyst within the short resi-
dence time transport riser to produce regenerated catalyst
at a higher delivery temperature. Thus the regenerated
catalyst which has passed from a first dense bed through
the transport riser to a second dense bed and then to the
hydrocarbon reaction zone contains an amount of heat
greater than that provided by the oxidation of coke alone
and permits a reduction in hydrocarbon feed prehea~.
Normally feed to the hydrocarbon reaction zone is pre-
heated in a furnace fired with external fuel to a~out 204
to 371C. to provide additional heat input before
contact with the regenerated catalyst. When our regener-
ation process is incor~orated into an FCC process, fe~d
- 25 preheat can usually be reduced or eliminated.
It is, furthermore, a feature of this process
that this CO conversion is catalyzed by catalytically
effective amounts of a CO conversion promoter contained
-39-
.,
105 .
-
as an integral part of the ~luid cracking catalyst
- employed in the FCC process. This feature is o~ part-~c-
ular importance for instance whén low coke yields are pro~
duced by the FCC process. This may occur for example when
a light feed stock is being charged to the FCC process or
when mild operating conditions are being employed. When
low amounts of coke are oxidiæed within the regeneration
zone, the temperature produced may be insufficient to
produce the desired rate of CO conversion to CO2. It may,
therefore, be necessary to ~ire torch oil within the
regeneration zone to increase the temperature. Without
the use of a CO conversion promoter, uneven dispersion of
fresh regeneration gas within the regeneration zone may
also require higher temperature or higher fresh regener-
ation gas rates than desired to maintain a sufficiently
fast rate of CO conversion. If ~he rate of CO conversion
is too slow, essentially complete conversion of the CO
will not be achieved within the transport riser and may
occur farther downstream of the riser where the presence
of catalyst has not been provided to serve as a heat sink.
The rate of coke oxidation is not per se
affected by employing a fluid catalytic cracking catalyst
containing a CO conversion promoter, but the rate of CO
conversion is increased. With a CO conversion promoter,
the kinetic rate constant for CO conversion to CO2 may
be increased typically from 2 to 5 times or more. Thus,
a faster rate of CO conversion can be obtained in the
presence of a CO conversion promoter at a given regener-
-40-
- : ~
7~8~3~
.,
- ation zone temperature than can be ohtained without the
promoter. Conversely, the same rate o~ CO conversion can
be obtained at a lower regeneration zone tempera~ure than
that required without a CO conversion promoter.
Suitable catalysts for use in the process of
our invention shall broadly be those which are in a physi-
cal form capable of fluidization in an FCC process, which
possess hydrocarbon cracking activity and which contain
catalytically effective amounts of a CO conversion pro-
moter~
While suitable CO conversion promoters can
comprise one or more oxides selected from the transition
metals (defined as metals of Group IB through VIIB and
Group VIII of the Periodic Table of the Elements) and
the rare earth metals, particularly suitable CO con~ersion
promoters will comprise one or more oxides selected from
the group consisting of noble metals. Even more
preferred, the C~ conversion promoter wi~l comprise p'aL-
inum oxide or palladium oxide or both. The CO conversion
~0 promoter can be incorporated as a component of the crack-
ing catalyst into any of the well-known amorphous FCC
catalysts comprising silica and/or alumina or into any of
the "molecular sieve"-containing FCC catalysts by methods
known to the catalyst manufacturing art such as co-precip-
itating or cogelling therewith or by impregnating with an
aqueous solution of a thermally decomposable salt and
heating to dry and decompose the salt. Suitable "molec-
ular sieves" include both naturally occurring and
-41-
:~7B~I()S
synthetic aluminosilicate materials known to the art such
as faujasite, mordenite, chabazite, zeolite X, and zeolite
Y to name a few. Catalytic effective amounts of a CO
conversion promoter shall generally mean such:amounts o~
a promoter as will increase the kinetic rate constant of
CO conversion to CO2. We have found that when the CO
conversion promoter comprises one or more oxid~s o~ the
noble metals a very small amount o~ its promoter catalyzes
CO conversion with substantially no harmful resul~s in
either product properties or yields obtalned from the
hydrocarbon reaction zone. Such amounts will preferably
be within tne range of from about 0.5 to about 200 wt.
ppm. of the total catalyst and more preferably within the
range of from about 0.5 to about 100 wt. ppm. of the total
catalyst.
Oxidizing conditions within the first dense bed
shall include a temperature within the range of from about
621C. to about 760C., dependina primarily upon the
amount of coke on spent catalyst entering the first dense
bed and secondarily upon wh~ther a portion of the hot
regenerated catalyst is returned directly from the second
dense bed to the first dense bed. More preferably, the
temperature will be within the range of from about 649C.
to about 704C. and if necessary, regenerated catalyst ¦
can be returned from the second dense bed to the first ¦
dense bed to maintain the temperature within this range.
Operating temperatures within the transport riser will
generally be higher than those within the first dense bed
and will typically be within the range of from about
-42-
.
`. :
- - . . . ..
-: . .
~L~'7~380~
677C. to about 774C. with a range of from 677C.:to
732C, being more preferred.
Pressures contemplated for use in the process of
this invention are from about normal atmospheric pressure
up to about 4.a atm. with the prefe~red range being from
about 2 to about 4.4 atm. Because of the improved gas-
solid contact due to higher velocities and the higher coke
burning rate because of higher temperatures and higher
oxygen concentrations, lower pressures can be employed in
the process of this invention with no penalty in coke
burning capacity. For this reason the process may
reverse the present industry trend toward higher pressure
regeneration processes. Less expen~sive vessels and air
blower or compressors having lower pressure ratings
therefore can be used.
One or more gas-solids separation means may be
utilized to separate regeneration gas from entraine~
regenera~ed catalyst. Preferred separation means will be
cyclone separators whose design and construction is well
known to the art. A single cyclone may be used but
preferably more than one of these cyclones will be used
in parallel or in series flow arrangements to effect the
desired degree of separation.
The following examples are presented to
illustrate some of the features and advantages of our
process as compared first with a conventional (non-CO
burning) regeneration process and second as compared
with a CO-burning regeneration process in which the
catalyst did not contain a CO conversion promoter. The
-43-
.i ,
1(3'788~5
examples are intended for illustration purposes only and
we do not intend to unreasonably limit the scope and
spirit of the claims attached hereto by specific refer-
ence to sizes, flow rates, pressures or analyses.
EXA~lPLE I
This example presents results obtained when
spent molecular sieve-containing FCC catalysts were
regenerated with air by a conventional commercial regen-
eration process and by a commercial process containing
features of this invention. Results of the regener-
ations are presented in Table 2 as Test 1 and Test 2
respectively.
The coke contained about 10.1 wt. ~ hydrogen,
and coke on spent catalyst was ab-out ~9 wt. %. ~he
conventional regeneration zone contained a single dense
bed in the bottom portion of the zone with a large
dilute phase disengaging space positioned above the
dense bed. Flue gas from this regeneration zone was
burned in a CO boiler. Analysis of the flue gas ~spent ;
regeneration gas) shown in Test 1 of the example was for
a sample removed from the flue gas line before the CO
boiler. In the regeneration process which produced
Test 2 results CO was essentially completely combusted
within the dilute phase transport riser of the regener-
ation zone. The flue gas was sampled just before it was
vented to the atmosphere. I
A comparison of the results of Test 1 and Test
2 first of all shows ~he higher oxygen concentration and
-44-
.
.
, . . . . . .
1~3'7138~
higher temperature of our invention. The inlet oxygen
concentration to each process is the same (air) but our .
process shows a higher outlet oxygen concentration of ..
1.8 vol. % compared to 0.2 vol. ~. The average oxygen
concentration within our regeneration zone is therefore
higher than is that of the conventional regeneration
zone.
10'78805
Table No. 2
Regeneration Process` Comparison
Test 1 Tes~ 2
Non-CO Burning CO Burning
Temperatures, C.
Dense bed 643 677
Dilute phase 641 --
Transport riser -- 741
Flue gas 674 738
Second dense bed -- 727
Pressure, atm. 2.7 2.4
Net dry air to regenerator,
kg/hr. 105,600 105,100
Dry air/coke, gm/gm 11.13 14.51
Coke yield, wt. % of fresh
- feed (at 75% reaction
zone conversion) 6.61 5.05
Coke on regenerated cata-
lyst, wt. % 0.2 0.02
Vessel size, diameter, m7.5 4.8
-Catalyst residence time, min. 3 0.9
Gas residence time, sec.15.5 5.5
Superficial velocity, m/sec.
Dense bed 0.72 1.5
Dilute phase 0.72 6.0
Flue gas analysis, vol.
C2 9.2 14.9
Argon 1.1 1.0
N2 79.2 82.3
CO 1~.2 0*
2 0.2 1.
- CH4 0.1 0
Regenerated catalyst stripping no yes
Feed preheat, C. 211 136
Catalyst inventory, tons60 35
*Actual reading was 270 ppm. vol.
~'-'.
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10'78805
The conventional regeneration process produced
a regenerated catalyst tempexature of 643C. while in our
process, a regenerated catalyst temperature of 727C.
resulted because at least a portion of the heat of CO
S conversion was transferred to the regenerated catalyst in
the transport riser. By recycling a portion of this
hotter regenerated catalyst back to the first dense bed,
the temperature within the dense bed for Test 2 has been
increased over the dense bed temperature for Test 1. The
higher regenerated catalyst temperature of our invention
has also reduced the catalyst circulation rate to the
hydrocarbon reaction zone thereby reducing the coke yield
at the same conversion level from 6.61 wt. g to 5.05
wt. %. This results in more liquid product yield from
the hydrocarbon reaction zone used in combination with
our process.
Because of the higher coke burning rate, -~
(caused by the hi~her temperature and higher oxygen
concentration) and because of the higher superficial ~ I
velocity and the lower coke yield, our process used
slightly less air at a lower pressure than did the
conventional process and produced regenerated catalyst
having a lower weight percent coke and flue gas contain-
ing substantially no CO. Thus, by converting the CO to
C02 within the transport riser and transferring at least
a portion of the heat to the regenerated catalyst in the
riser an air pollution problem has been essentially
eliminated without re~uiring a CO boiler and at
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iO ;'8805 1 .
the same time, the feed preheat requirement has been
reduced by about 75C. Furthermore, the inventory of our
process is about 40% less than that of the conventional
process and therefore, the size of the regeneration
vessel is smaller.
Although not shown, it is anticipated that
because of the smaller inventory and shorter residence
time the catalyst make-up rates over a period of time :~
will be less for our process. As indicated in Table 2,
the regenerated catalyst was stripped in the present
invention but not in the conventional process. Although
not shown in the table, the advantaye of this stripping
would be apparent in a comparison of equipment sizes
required in the gas concentration plant.
EX2~1PLE II
This example shows the advantages realized on I -
one particular FCC process unit when a catalyst contain-
ing a CO conversion promoter comprising about 10 wt. ppm.
of platinum oxide was employed. Data obtained before
and after the use of the catalyst illustrating the
advantages are shown in Table No. 3 as Tests 1 and 2
respectively.
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1078~305
Table No. 3
Regeneration ~ith and Without CO Conversion Promoter
Test 1 Test 2
Without CO With CO
Conversion Conversion :
Promoter Promoter .
Reyeneration Zone Temps., C.
cyclones 761 705
regenerated catalyst759 705
first dense bed 737 682
Air heater outlet 379 171 :
Feed preheater outlet 308 271 .
Slurry recycle to
hydrocarbon conv. zone, m3/hr17.2 10.5 .
Flue Gas Analysis, Vol. ~
C2 14.6 15.0 :-
- 2 2.8 2.6 ~ .
CO ~500 vol. ppm.~500 vol. ppm. -:
Before the use of the catalyst containing the .-
GO CO conversion promoter, it was necessary to fire the air
heater continuously, use the feed preheater, and increase
the amount of slurry oil recycled back to the hydrocarbon .
conversion zone to increase the regeneration zone temper-
ature to a point ~7here the flue gas leaving the regener- .
ation zone contained a desired CO concentration of less
than 500 vol. ppm. As shown in Test 1, the air heater
temperature was 379C., the feed preheater temperature . .
was 308C. and the slurry recycle was 17.2 m3/hr. These
operating conditions produced regeneration zone temper- :
atures in the range of about 737C. to 761C. at which .
temperatures the CO concentration in the flue gas was
less than 500 vol. ppm.
The data for Test 2 shows that by using an FCC
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~3'7~3805
catalyst containing a CO conversion promoter comprising
platinum~oxide the desired CO concentration of less than
500 vol. ppm. was achieved at regeneration zone temper-
atures 55C. lower than those of Test 1. This lower
regeneration ~one temperature requirement permitted a
reduction of from 379C. to 171C. in the air heater
temperature, a reduction of from 30~C. to 271C. in the
feed preheater temperatures and a reduction oE from 17.2
m3/hr. to 10.5 m3/hr of slurry oil recycle. These
reductions permitted a savings in utility costs and a
slight increase in the fresh feed rate.
In addition to permitting lower temperatures
the use of a catalyst containing a CO conversion pro-
moter in other FCC process units has permitted a reduc-
tion in the fresh regeneration gas ~air) rate needed for
a particular feed rate to achieve a desired CO concen-
tration in the flue gas. In an FCC process unit being
operated near the design limits of cyclone s~parators
such a reduction may in turn reduce particulate emissions
from the unit. Where cyclone separation efficiency is
not a problem, other units may be able to increase the
fresh feed rate at a fresh regeneration gas rate no
higher than that previously used with a lower fresh feed
rate but without a catalyst containing a CO conversion
promoter.
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