FIELD OF THE INVENTION
[0001] This invention relates to the production of olefins and aromatics from hydrocarbon
feedstocks. More particularly, the invention relates to the production of olefins
and aromatics by catalytically cracking alone or cracking and dehydrogenating a hydrocarbon.
Most particularly the invention relates to a process for cracking hydrocarbons in
the presence of an entrained stream of catalytic heat carrying solids at short residence
times to preferentially produce olefins having three or more carbon atoms and/or to
produce aromatics, specifically benzene.
BACKGROUND OF THE INVENTION
[0002] It has long been known that naturally occurring hydrocarbons can be cracked at high
temperatures to produce valuable olefinic materials, such as ethylene and propylene.
[0003] The growth in the propylene based plastics market relative to the ethylene based
plastics market has made it desirable to improve the propylene yield when cracking
hydrocarbons to olefins.
[0004] In addition, higher order olefins, e.g., C₄ olefins, are important precursors for
providing high octane blending components, i.e., C₄'s are precursors to MTBE production
and alkylation.
[0005] However, when heavy hydrocarbons feedstocks are non-catalytically cracked to olefins
it's virtually impossible to achieve the desired co-product ratios to fit market needs,
i.e., propylene to ethylene yield ratios are rarely greater than 0.55. Higher ratios
are attainable only at low hydrocarbon conversion which represents a significant processing
penalty in terms of recycle costs and feed degradation. One well-known non-catalytic
cracking process is pyrolysis which typically takes place in the presence of steam
at high temperatures. The mechanism by which pyrolysis to olefins is achieved is explained
in terms of a free radical mechanism.
[0006] At high temperatures, radical initiation takes place by homolysis of a carbon - carbon
bond. Once initiated, the free radicals undergo two principal reactions. They are
(1) scission at the beta position of the radical and (2) abstraction of a hydrogen,
resulting in termination of the reaction.
[0007] The scission at the beta position will continue to the point where a methyl radical
will be formed at 90 percent frequency. The methyl radical will then abstract a hydrogen
atom from another molecule to form methane and another free radical. Ethylene and
methane are the principal products from such free radical pyrolysis reactions. Only
about 10 percent of the time will a longer radical abstract a hydrogen from a molecule
to form C₃ to C₇ paraffins and olefins. Thus, thermal cracking results in high yields
of ethylene relative to higher order olefins with the higher order olefins occurring
principally as a result of hydrocarbon branching in the initial hydrocarbon feedstock.
[0008] One effort at producing increased production of C₃ and higher olefins is directed
to subjecting a light hydrocarbon comprising at least one alkane to cracking conditions
in the presence of hydrogen sulfide and a solid contact material comprising silica
(Kolts, United States Patent No. 4,471,151). The contact material employed, such as
silica gel, preferably has a high surface area i.e. at least 50 m ²/gm. Typical H₂S
concentrations of 0.1 to 10 mole percent based on the alkane feed are employed in
the process. It is theorized in Kolts that the improvement in cracking is due to the
high surface area material which acts as a catalyst to decompose H₂S. The result is
increased conversion levels with improved selectivity to desired products. However,
the improved selectivity to propylene was demonstrated only when cracking n-butane.
[0009] The solid contact material employed in Kolts is suitable only for fixed bed operations
and not for fluidized bed environments due to its very low mechanical stability. Thus,
the solid catalyst of Kolts continues to have the drawbacks of typical catalytic dehydrogenation
catalysts designed for fixed beds. These are larger size, diffusion limited catalysts
incapable of continuous regeneration in a circulating loop system.
[0010] A fluidized catalytic cracking (FCC) unit may also be employed to catalytically produce
C₃ and higher compounds. The FCC unit uses acidic cracking catalysts to increase the
production of C₃ to C₇ compounds through a carbonium ion mechanism compared to the
free radical pyrolysis reaction mechanism. However, the acidic cracking activity of
the catalysts, in addition to promoting cracking and isomerization, promotes rapid
hydrogen transfer resulting in high yields of paraffins rather than olefins. Further,
the nature of the catalytic cracking unit itself favors the shift to paraffins.
[0011] The typical definition of residence time in a catalytic cracking operation is the
time the feedstock is in contact with the catalyst itself. This definition is acceptable
if the temperatures are low such that thermal reactions do not occur to any appreciable
extent. However, thermal and catalytic reactions proceed in parallel. While catalyst
separation will terminate the catalytic portion of the reaction, the thermal reactions
(pyrolysis) will continue until the temperature is reduced to a level where the rate
of reaction is insignificant (quench). In this situation, the total kinetic residence
time can be defined as the time from the introduction of the hydrocarbon into the
system to the quenching of the effluent including the separation of the solids from
the reaction. Total conversion is thus the summation of the catalytic reaction (time
in contact with the catalyst) and the thermal reactions (time at the reaction temperature).
[0012] The typical FCC reaction environment has relatively long residence times including
time for solids separation (normally greater than one second) and does not include
a quench. Cracking takes place at lower temperatures under these longer residence
times. Conversion is achieved at these lower temperatures due to the extended contact
with the catalyst. Thermal reactions are minimized at these lower temperatures thus
eliminating the need for quenching the effluent. While increased C₃ and higher compounds
are produced in comparison to pyrolysis, the effluent will have a disproportionately
high concentration of paraffins due to the increased hydrogen transfer activity. The
favored conditions for olefin production, specifically higher temperatures and shorter
residence times, are difficult to achieve especially when processing light feedstocks
such as LPG and naphthas which require proportionately higher temperatures to initiate
and sustain the reaction (either catalytic or thermal).
[0013] The above processes all improve the cracking of hydrocarbons to olefins. However,
these processes suffer either from high capital and operating costs associated with
fixed bed operations and hydrogen sulfide dilution, or result in low yields of the
desired olefins. In addition, the use of hydrogen sulfide as a diluent raises environmental
and health concerns because of its extremely high toxicity.
[0014] It has now been found that the higher order olefins, i.e. propylene, butenes, etc.
can be obtained in high yields by the cracking of hydrocarbons in the presence of
an acidic cracking catalyst alone or in combination with a noble metal oxide dehydrogenation
catalyst in a short residence time fluidized solids cracking environment. This short
residence time is achieved by a combination of a low residence time reactor, a very
short residence time separation system, and a product quench.
[0015] It is therefore an object of the present invention to provide a process in which
hydrocarbons can be catalytically cracked to produce olefins and aromatics.
[0016] It is another object of the present invention to provide a process for preferentially
cracking hydrocarbons to obtain C₃ to C₅ olefins and/or aromatics.
[0017] It is another object of the present invention to provide a process in which a hydrocarbon
may be cracked to a variety of desired products by altering the catalyst system in
the process.
[0018] It is a further object of the present invention to provide a reaction system including
a quenching step for preferentially cracking hydrocarbons to obtain C₃ to C₅ olefins
and/or aromatics while avoiding the thermal degradation of products.
SUMMARY OF THE INVENTION
[0019] The present invention relates generally to a process for preferentally cracking hydrocarbons
to obtain olefins, preferably C₃ to C₅ olefins, and aromatics at the acid sites of
catalyst solids and, optionally, catalytically dehydrogenating the resulting paraffin
isomers to thereby produce olefins.
[0020] Acidic catalytic cracking of hydrocarbons proceeds by a carbonium ion mechanism unlike
the free radical mechanism of thermal cracking. The carbonium ion is formed by the
abstraction of a hydride ion from the carbon - hydrogen bond. The abstraction of the
hydride ion and the creation of a carbonium ion is catalyzed by the acid sites on
the catalyst solids.
[0021] Carbonium ion cracking also occurs at the beta position thereby leading to the formation
of an olefin and a primary carbonium ion. The primary carbonium ion undergoes a rapid
ionic shift (isomerization) to produce a secondary or tertiary carbonium ion. This
coupled with the beta cracking rule leads to the formation of propylene in high yields
without the concurrent production of significant amounts of ethylene. Any ethylene
found in the product is the result of the competitive free radical cracking route.
In addition to providing the carbonium ion mechanism for isomerization, the acidic
sites on the catalyst promote hydrogen transfer. Thus, while the thermodynamic equilibrium
conditions at the temperatures contemplated in the invention favor olefins over paraffins,
the increased hydrogen transfer activity may result in a disproportionately high paraffin
yield. This is especially true for the branched isomers such as isobutylene. In these
cases, if a specific dehydrogenation catalyst is used in combination with an acidic
cracking catalyst, the yield distribution can be shifted toward the thermodynamic
equilibrium and higher concentrations of the desired olefins can be obtained.
[0022] For the purpose of this invention, the kinetic residence time is defined as the total
time from the point where the hydrocarbon is introduced to the reactor zone to the
point where the cracked products are quenched, including the intermediate separation
step. This distinguishes the present process from other processes where measurement
of the residence time is terminated prior to the point of separation and quench. This
is especially important since the catalytic cracking of hydrocarbons always proceeds
in parallel with pyrolysis. The extent to which products are formed catalytically
or thermally is a function of catalyst activity, catalyst loading, catalyst residence
time, reaction temperature profile, and the total kinetic residence time in the thermal-catalytic
environment. For example, mild acidic catalytic activity at higher temperatures could
be used to shift diolefin production to paraffins and olefins without substantially
altering the ratio of the carbon products obtained by pyrolysis. Alternatively, very
highly active acidic cracking catalysts could be used at significantly lower temperatures
to minimize the thermal route and maximize the acidic catalyst product distribution.
Further, it has been found that catalytic dehydrogenation catalysts can be used in
combination with the acidic cracking catalysts to shift the reaction in favor of olefin
production.
[0023] The present invention is particularly well suited for cracking hydrocarbon feedstocks
such as C₄-C₇ paraffins, naphthas, and light gas oils to higher order olefins, i.e.,
having three to five carbon atoms and/or to aromatics. However, it should be noted
that the process has general applications for cracking the entire range of hydrocarbons
from light distillates to heavy resids.
[0024] The process of the present invention proceeds by delivering a preheated hydrocarbon
feedstock and steam to the top of a downflow tubular reactor. Simultaneously, hot
catalyst solids are introduced to the top of the reactor and the combined stream of
hydrocarbon, steam and catalyst solids pass through the reactor zone, a separation
zone, and a quench zone where the hydrocarbon undergoes cracking at low severity and
short residence times and the effluent is stabilized to prevent product degradation.
[0025] The tubular reactor is operated at a temperature of about 900-1500°F, preferably
1000°-1300°F and at a pressure of about 10-100 psia with a total kinetic residence
time of about 0.05 to 2.0 seconds, preferably about 0.10 to 0.5 seconds.
[0026] After separation from the cracked effluent the catalyst solids are stripped of residual
hydrocarbon, regenerated and reheated in a transfer line and returned to the tubular
reactor to continue the cracking process.
[0027] The present invention is particularly well adapted foruse in a short residence time
fluidized solids cracking apparatus and in a short residence time separation apparatus,
as described in United States Letters Patent Nos. 4,370,303 to Woebcke et al, and
4,433,984 to Gartside et al, and pending U.S. Serial No. 084,328 to Gartside et al
each of which is incorporated herein by reference.
[0028] The specific catalyst solids and the catalyst to hydrocarbon ratio are chosen based
on the feedstock characteristics and the product distribution desired. Catalyst activity
and catalyst loading will define operating temperatures at the short residence times
employed in the present invention and thus determine the split between the catalytic
and thermal reactions. The catalyst type, either acidic cracking alone or in combination
with noble metal oxide dehydrogenation, will further determine the product distribution
between olefins and paraffins.
DESCRIPTION OF THE DRAWING
[0029] The process of the present invention will be better understood when considered with
the following drawings, wherein:
FIGURE 1 is a schematic view of the process scheme of the present invention;
FIGURE 2 is a cross-sectional elevational view of the reactor feeder employed in the
apparatus of the present invention;
FIGURE 3 is a cross-sectional elevational view of the separator employed in the present
invention;
FIGURE 4 is a sectional view through line 4-4 of FIGURE 3.
FIGURE 5 is a schematic view of an optional quenching process scheme of the present
invention;
FIGURE 6 is an elevational view of one embodiment of the overall system of the present
invention;
FIGURE 7 is a cross-sectional elevational view of one embodiment of the reactor and
gas-solids separator employed in the present invention;
FIGURE 8 is a sectional plan view through line 8-8 of FIGURE 7; and
FIGURE 9 is a schematic elevational view of another embodiment of the solids regeneration
assembly employed in the present invention.
DETAILED DESCRIPTION OF THE INVENTION
[0030] As has been previously indicated, the process of the present invention is directed
to a means for cracking hydrocarbon feedstocks in the presence of catalytically active
heat carrying solids for the purpose of producing olefins with a high selectivity
especially towards C₃ to C₅ olefins and/or aromatics.
[0031] The hydrocarbons contemplated as feedstocks include the high boiling distillate gas
oils, atmospheric gas oils, naphthas, and C₄-C₇ paraffins. However, it should be noted
that the process has general applications for catalytically cracking a wide range
of hydrocarbons to produce the desired olefins and/or aromatics.
[0032] Referring to the drawings and first to FIGURE 1, the process of the present invention
can be performed in a short residence time fluidized solids cracking system 1, hereinafter
QC system, incorporating a tubular reactor 2, a reactor feeder 4, a separator 6, a
quench means 24 and a solids stripper 8.
[0033] The system 1 also includes means for regenerating the catalyst solids separated from
the cracked product after the reaction. The system shown illustratively includes an
entrained bed heater 10 wherein the catalyst solids can be regenerated and reheated,
a transport line 12 and a fluid bed vessel 14 wherein the solids are stripped of combustion
gases and again distributed to the reactor 2.
[0034] In operation, hot catalyst solids from the fluid bed vessel 14 enter the reactor
feeder 4 and are admixed with steam entering through a line 16. The hydrocarbon feed
is delivered through a line 18 to a preheater 20, then through a line 22 to the upper
region of the tubular reactor 2. The preheated hydrocarbon feed along with the catalyst
solids and steam from the reactor feeder 4 are passed through the tubular reactor
2. Intimate mixing of the hot catalyst solids, steam and preheated hydrocarbon occurs
in the reactor and cracking proceeds immediately.
[0035] Upon exiting the tubular reactor 2 the cracked hydrocarbon effluent and steam are
immediately separated from the catalyst solids in the separator 6 and the cracked
effluent product passes overhead through the quench area 24 where the cracked product
is immediately quenched with steam or a light hydrocarbon delivered to the quench
area 24 through a quench line 26. This reduces the temperature of the mixture below
the point where substantial thermal reactions occur. Alternatively, the cracked product
exiting the tubular reactor 2 and separated from the catalyst solids in the separator
6 may be quenched by passing the entire mixture over a bed of solids with catalytic
(dehydrogenation) activity. Since dehydrogenation is an endothermic reaction, the
flowing mixture will be cooled as the reaction proceeds. This can be used with the
introduction of steam to improve the reaction conditions. As seen in FIGURE 5, the
preferred method of quenching in this manner includes the use of a catalyst reactor
25, in which the bed of the catalytic solids are contained, located immediately downstream
of the separator 6, where quenching occurs in the previous embodiment.
[0036] The quenched product is passed through a cyclone 28 where small amounts of entrained
catalyst solids are removed and delivered through a line 30 to the solids stripper
8 where they are combined with the bulk of the stripped solids delivered from the
separator 6 through a line 32. In the solids stripper 8, the catalyst solids are striped
of residual hydrocarbon by steam, nitrogen or other inert gases delivered to the solids
stripper 8 through a line 34.
[0037] The catalyst solids, which have accumulated carbon or coke deposits from the tubular
reactor 2 are then passed to the entrained bed heater 10. Air delivered to the heater
10 through a line 36 is mixed with the stripped catalyst solids in the heater 10 and
the mixture is fed into the transport line 12 for conveying the catalyst solids back
to the fluid bed vessel 14. In the presence of air from the line 36, the carbon deposits
on the catalyst solids are removed by combustion to provide the heat necessary for
the cracking reaction. If additional fuel is required it may be added into the entrained
heater 10 from a fuel source (not shown).
[0038] In essence, the process of the present invention is conducted by delivering a hydrocarbon
such as naphtha, atmospheric gas oil or mixtures thereof, through the line 18 to
the preheater 20 wherein the temperature of the hydrocarbon is elevated to about 800-900°F.
Simultaneously, catalyst solids from the fluid bed vessel 14 are delivered to the
reactor feeder 4 (best seen in FIGURE 2) where they are admixed with steam supplied
through the line 16 and delivered to the reactor at a temperature in the range of
1000-1600°F. The catalyst solids to the hydrocarbon feed ratio ranges from 1 to 60:1
based on weight depending on the particular catalyst utilized. The water vapor/hydrocarbon
feed ratio is in the range of 0 to 1.0, preferably 0.0 to 0.3.
[0039] Optionally, the catalytic cracking process may be initiated by injecting an alkane
such as ethane into the tubular reactor 2, via injection line 16, to form olefins
and free radicals. This will tend to increase isomerization by forming carbonium ions
and stabilize the heavier hydrocarbon formation by competing with the free radicals
formed as well. Such alkanes are added just upstream of the hydrocarbon feed 22.
[0040] A suitable catalyst solids for the present invention may be one of the generally
available supports having acid properties such as, silica gel, alumina, clays, etc.
The catalyst solid can have associated therewith other catalytically active material.
Alternatively, the catalyst system employed may be a conventional zeolitic FCC catalyst
or one of the high activity ZSM-5 or rare earth zeolitic catalysts. The catalyst employed
may also include a dehydrogenation catalyst consisting of one of the noble metal oxides
such as the oxides of iron, chromium, platinum, etc. on a suitable support such as
silica alumina. Alternately, the catalyst could be a mixture of the aforementioned
catalysts to achieve specific yield distributions.
[0041] The composite hydrocarbon feedstock is elevated to 800 to 1100°F. and the catalyst
solids are heated to 1200 to 1700°F in the tubular reactor 2. The ratio of solids
to hydrocarbon is set by heat balance and desired solids catalytic activity.
[0042] The cracked effluent product and catalyst solid effluent from the tubular reactor
2 flow directly into separator 6 (best seen in FIGURE 3) where a separation into a
gas product phase and a catalyst solid phase is effected. The gas product is removed
via the line 24, while the catalyst solids enter the solids stripper 8 through the
line 32. An in-line quench of the gas product is provided in quench area 24 through
the quench line 26. Cold solids, water, steam, light hydrocarbons, and recycle oils
are examples of suitable quench materials. Alternatively, quenching takes place in
the catalyst reactor 25 (see FIGURE 5) by passing the product over a catalyst bed,
the additional reaction being without the presence of solids.
[0043] The total residence time from the point of hydrocarbon introduction to the tubular
reactor 2 to the point of quench in the quench area 24, optionally comprising a catalyst
reactor 25, is preferably about 0.1 to 0.3 seconds.
[0044] In the solids stripper 8 the catalyst solids are stripped of gas impurities by a
stream of steam, nitrogen or inert gas delivered through the line 34. Vapors are removed
from the solids stripper 8 through the line 30.
[0045] The stripped catalyst solids are removed from the stripper 8 through a line 38. The
catalyst solids which have accumulated carbon from the tubular reactor 2 are passed
to the entrained bed heater 10 where air is delivered through a line 36 to provide
the necessary atmosphere for regenerating the catalyst solids. The catalyst solids
are entrained in the heater 10 and returned to the fluid bed vessel 14 through the
transport line 12 where the catalyst solids continue to regenerate. In addition, the
regeneration of the catalyst solids raises the temperature of the catalyst solids
to about 1200 to 1700°F prior to delivery of the catalyst to the fluid bed vessel
14.
[0046] Details of the reactor feeder 4 are more fully described in United States Letters
Patent No. 4,338,187 to Gartside et al., which is incorporated herein by reference.
The reactor feeder of Gartside et al. has the capability of rapidly admixing hydrocarbon
feed and catalyst solids. As seen in FIGURE 2, the reactor feeder 4 delivers catalyst
solids from a solids receptacle or fluid bed vessel 70 through vertically disposed
conduits 72 to the tubular reactor 2 and simultaneously delivers hydrocarbon feed
to the tubular reactor 2 at an angle into the path of the catalyst solids being discharged
from the conduits 72. An annular chamber 74 to which hydrocarbon is fed by a single
entry comprising a toroidal feed line 76 terminates in angled openings 78. A mixing
baffle or plug 80 also assists in effecting rapid and intimate mixing of the hydrocarbon
feed and the catalyst solids. The edges 79 of the angled openings 78 are preferably
convergently beveled, as are the edges 79 at the reactor end of the conduits 72. In
this way, the gaseous hydrocarbon stream from the chamber 74 is angularly injected
into the mixing zone and intercepts the catalyst solids phase flowing from the conduits
72. A projection of the gas would form a cone shown by dotted lines 77, the vortex
of which is beneath the flow path of the solids. By introducing the gaseous hydrocarbon
phase angularly, the two phases are mixed rapidly and uniformly, and form a homogeneous
reaction phase.
[0047] The mixing of a solid phase with a gaseous phase is a function of the shear surface
between the solids and gas phases, and the flow area. A ratio of shear surface to
flow area (S/A) of infinity defines perfect mixing while poorest mixing occurs when
the solids are introduced at the wall of the reaction zone. In the system of the present
invention, the gas stream is introduced annularly to the solids which ensures high
shear surface. By also adding the gas phase transversely through an annular feed means,
as in the preferred embodiment, penetration of the phases is obtained and even faster
mixing results. By using a plurality of annular gas feed points and a plurality of
solid feed conduits, even greater mixing is more rapidly promoted, since the shear
surface to flow area ratio for a constant solids flow area is increased. Mixing is
also a known function of the length to diameter ratio of the mixing zone. A plug creates
an effectively reduced diameter D in a constant length L, thus increasing mixing.
[0048] The plug 80 reduces the flow area and forms discrete mixing zones. The combination
of annular gas addition around each solids feed point and a confined discrete mixing
zone greatly enhances the conditions for mixing. Using this preferred embodiment,
the time required to obtain an essentially homogenous reaction phase in the reaction
zone is quite short. Thus, this preferred method of gas and solids addition can be
used in reaction systems having a residence time below 1 second, and even below 100
milliseconds. Because of the environment of the tubular reactor 2 and the reactor
feeder 4, the walls are lined with an inner core 81 of ceramic material.
[0049] The separator 6 of the QC system, as shown in FIGURE 3, can also be relied on for
rapid and discrete separation of product and catalyst solids discharging from the
tubular reactor 2. The inlet to the separator 6 is directly above a right angle corner
90 at which a mass of catalyst solids 92 collect within a chamber 93. An optional
weir 94 downstream from the right angle corner 90 facilitates accumulation of the
mass of solids 92 especially when run on small scale rather than commercial scale
production. The gas outlet 24 of the separator 6 is oriented 180° from a separator
gas-solids inlet 96 and the solids outlet line 32 is directly opposed in orientation
to the gas outlet 24 and downstream of both the gas outlet line 24 and the weir 94.
[0050] In operation, centrifugal force propels the catalyst solids to the wall opposite
inlet 96 of the chamber 93 while the gas portion having less momentum, flows through
the vapor space of the chamber 93. Initially, catalyst solids impinge on the wall
opposite the inlet 96 but subsequently accumulate to form a static bed of solids 92
which ultimately form in a surface configuration having a curvilinear arc of approximately
90° of a circle. Solids impinging upon the bed 92 are moved along the curvilinear
arc to the solids outlet 95, which is preferably oriented for downflow of solids by
gravity. The exact shape of the arc is determined by the geometry of the particular
separator and the inlet stream parameters such as velocity, mass flowrate, bulk density,
and particle size. Because the force imparted to the incoming solids is directed against
the static bed 92 rather than the separator 6 itself, erosion is minimal. Separator
efficiency, defined as the removal of solids from the gas phase leaving through the
outlet 97, is therefore, not affected adversely by high inlet velocities, up to 150
ft./sec., and the separator 6 is operable over a wide range of dilute phase densitites,
preferably between 0.1 and 10.0 lbs./ft.³. The separator 6 of the present invention
achieves efficiencies of about 90%, although the preferred embodiment, can obtain
over 97% removal of catalyst solids.
[0051] It has been found that for a given height H of the chamber 93, efficiency increases
as the 180° U-bend distance between the inlet 96 and the outlet 97 is brought progresively
closer to the inlet 96. Thus, for a given height H the efficiency of the separator
6 increases as the flow path decreases and, hence, residence time decreases. Assuming
an inside diameter D
i of the outlet 96, the distance CL between the centerlines of the inlet 96 and the
outlet 97 is preferably not greater than 4.0(D
i), while the most preferred distance between said centerlines is between 1.5 and 2.5(D
i). Below 1.5(D
i) better separation is obtained but difficulty in fabrication makes this embodiment
less attractive in most instances. Should this latter embodiment be desired, the separator
6 may require a unitary casting design because the inlet 96 and the outlet 97 would
be too close to one another to allow welded fabrication.
[0052] It has been found that the height H should be at least equal to the value of 1.5
x D
i or 4 inches in height, whichever is greater. Practice teaches that if H is less than
D
i or 4 inches the incoming stream is apt to disturb the bed solids 92 thereby reentraining
solids in the gas product leaving through the outlet 97. Preferably the height H is
on the order of twice D
i to obtain even greater separation efficiency. While not otherwise limited, it is
apparent that too large a height H eventually merely increases residence time without
substantive increases in efficiency. The width W shown in FIGURE 4 of the flow path
is preferably between 0.75 and 1.25 times D
i, most preferably between 0.9 and 1.10 (D
i).
[0053] The outlet 97 may be of any inside diameter (Dog). However, velocities greater than
75 ft./sec. can cause erosion because of residual solids entrained in the gas. The
inside diameter Dog of the outlet 97 should be sized so that a pressure differential
between the solids stripper 8 shown in FIGURE 1 and the separator 6 exists such that
a static height of solids is formed in the solids outlet line 32. The static height
of solids in the solids outlet line 32 forms a positive seal which prevents gases
from entering the solids stripper 8. The magnitude of the pressure differential between
the solids stripper 8 and the separator 6 is determined by the force required to move
the solids in bulk flow to the solids outlet 95 as well as the height of solids in
the line 32. As the differential increases the net flow of gas to the solids stripper
8 decreases. Solids, having gravitational momentum, overcome the differential, while
gas preferentially leaves through the gas outlet 97. Preferably, the inside diameter
Dog of the gas outlet 97 is the same as the inside diameter of the inlet 96, when
one outlet is employed, to provide outlet velocity less than or equal to inlet velocity.
[0054] FIGURE 4 shows a cutaway view of the separator 6 along section 4-4 of FIGURE 3. It
is essential that longitudinal side walls 101 and 102 be rectilinear, or slightly
arcuate as indicated by the dotted lines 101a and 102a. Thus, the flow path through
the separator 6 is essentially rectangular in cross-section having a height H and
width W as shown in FIGURE 4. The embodiment shown in FIGURE 4 defines the geometry
of the flow path by adjustment of the lining width for the walls 101 and 102. Alternatively,
baffles, inserts, weirs or other means may be used. In like fashion the configuration
of the walls 103 and 104 transverse to the flow path may be similarly shared, although
this is not essential.
[0055] The separator shell and manways are preferably lined with erosion resistant linings
105, which may be required if solids at high velocities are encountered. Typical commercially
available materials for erosion resistant linings include Carborundum Precast Carbofrax
D, Carborundum Precast Alfrax 201 or their equivalent. A thermal insulation lining
106 may be placed between the shell and the lining 105 and between the manways and
their respective erosion resistant linings when the separator 6 is to be used in high
temperatures service.
[0056] The details of the separator 6 are more fully described in United State Letters Patent
No. 4,288,235 which is incorporated herein by reference.
[0057] An alternative embodiment of the apparatus employed to carry out the present invention
is disclosed in the aforementioned Gartside et al. Serial No. 084,328, incorporated
herein by reference.
[0058] Referring to the drawings and particularly to FIGURES 6-9, there is described a system
202 comprising a reactor system 204, a solids regeneration assembly 208 and a solids
delivery system 210.
[0059] The reactor system 204, best seen in FIGURE 7, includes a convergent mixing section
211, an elongated reaction section 212, a divergent section 213 downstream of the
elongated reaction section 212, a separator 206 and a quench system 207 (shown in
FIGURE 8). The mixing sections 211 are formed with a plug section 214 shown in cross-section
as having an arcuate lower surface 215. A horizontally disposed plate 217 is arranged
over the plug section 214 in spaced-relationship with the plug section 214 to form
solids inlet passages 219 to the interior of the mixing section 211. The solids inlet
passages 219 are configured in cross-section with a right angle turn and terminate
in a reactangular openings 225 through which the particulate solids enter the mixing
section 211, in the form a curtain of solids 226. The horizontal openings 225 are
directly above each hydrocarbon feed inlet. Venturi configured passages 203 extend
from the solids inlet passages 219 to the hydrocabon feed inlets 228.
[0060] Steam plenums (not shown) are arranged along each longitudinal edge of the horizontal
opening 225 to deliver pre-acceleration gas (steam) through nozzles (not shown) into
the curtain of solids 226 passing through the horizontal openings 225. A gas delivery
line (not shown) is provided to deliver gas, usually steam or light hydrocarbon, under
pressure to the nozzles. The nozzles are arranged at a downward angle of 45° to the
horizontal. The pre-acceleration gas is delivered to the plenums at pressures of 3
to 5 psi above the pressure in the reactor and discharges through the nozzles at the
same relative pressure at a velocity of about 150 feet per second. The pre-acceleration
gas accelerates the flow of solids through the horizontal openings 225 from a nominal
three to six feet per second to approximately 50 feet per second for the mix of solids
and pre-acceleration gas. A more detailed description is found at Gartside et al Serial
No. 084,328.
[0061] The hydrocarbon feed inlets 228 are located on the reactor wall arranged either normal
to the solids curtain 226 or at an angle upwardly of 30° into the solids curtain 226.
The hydrocarbon feed is delivered to a manifold 223 through a line 224. The feed inlet
nozzles 228 are fed with hydrocarbon from the manifold 223. As seen in FIGURE 7, the
feed inlet nozzles 228 are diametrically opposed from each other in the same horizontal
plane. The mixing zone 211 of the reactor is rectangular with the configuration making
a transition to a tubular reactor at the elongated reaction section 212.
[0062] The feedstock entering the mixing zone 211 through nozzles 228 immediately impinge
the solids curtains 226 and the desired mixing of feed and hot particulate solids
occurs. With the opposing set of nozzles 228, the opposing feed jets and entrained
solids from the solids curtain 226 will be directed by the arcuate contour 215 of
the plug section 214 and impact with each other at approximately the vertical centerline
of the mixing zone 211. When a gas-liquid mixed phase hydrocarbon is fed through the
nozzles 228, the nozzles 228 are arranged at an angle normal or 90° to the solids
curtain 226. When the hydrocarbon feed is a gas, the nozzles 228 are arranged at an
upwardly directed angle of 30° into the solids curtain. The quantity of solids entering
the mixing zone 211 of the reactor system 204 through the horizontal inlets 219 is
controlled in large part by the pressure differential between the mixing zone 211
of the reactor system 204 and the chamber 231a above the solids reservoir 218 in a
solids control hopper 231 directly above the horizontal inlets 219. Pressure probes
233 and 235 are located respectively in the mixing zone 211 of the reactor system
204 and the control hopper chamber 231a to measure the pressure differential. Gas
(steam) under pressure is delivered through a line 230 to the control hopper chamber
231a to regulate the pressure differential between the mixing zone 211 of the reactor
system 204 and the control hopper chamber 231a to promote or interrupt flow of the
solids from the solids control hopper 231 to the mixing zone 211.
[0063] As best seen in FIGURE 7, the separator 206 is comprised of a mixed phase inlet 232,
a horizontal chamber section 234, a plurality of cracked gas outlets 236 and particulate
solids outlets 238. The basic principles relating to relative diameters (Di, Dog,
Dos), chamber height (H) and length (L) recited in the first embodiment described
herein are applicable herein. The separator 206 is arranged in combination with the
elongated cracking zone 212 and divergent section 213 of the reactor system 204. The
divergent section 213 terminates in the separator mixed phase inlet 232 which is centrally
disposed at the top of the horizontal section 234. As a result of the configuration
of the composite reaction system including the separator 206, a solids bed 242 develops
on the floor 240 of the horizontal section 234 with the cross-sectional profile 243
of the bed 242 forming a curvilinear arc over which the mixed phase gas and solids
travel. The expansion of solids and cracked gas in the divergent section 213 enhances
heat transfer and limits the velocity of the solids-gas mixture entering the separator
206.
[0064] The solids are sent to the lateral ends 246 of the horizontal section 234 and discharge
downwardly through the solids outlets 238. The cracked gases follow a 180° path and
after separation from the solids discharge through gas outlets 236 that are located
on the top of the horizontal section 234 intermediate the lateral ends 246. The plurality
of solids outlets 238 and gas outlets 236 provide simultaneously for both minimum
time in the separation zone and maximum solids-gas separation.
[0065] The separation or quench system 207 also includes a conventional cyclone separator
250 directly downstream of each gas outlet 236, as best seen in FIGURE 8. The entry
line 254 to each cyclone separator 250 is arranged at an angle of 90° to the gas outlet
236 with the cyclone separator 250 vertically disposed in the system. The cyclone
separators 250 serve to collect the remaining entrained particulate solids from the
cracked gas discharged from the separator 206. A dipleg line 249, returns the particulate
solids to the regeneration assembly 208 and the cracked gas is sent for downstream
processing through the gas outlet 251.
[0066] Each cyclone entry line 254 extending from the cracked gas outlet 236 to the cyclone
250 is provided with a direct quench line 252. Quench oil, usually the 100-400°F cut
from a downstream distillation tower is introduced into the cyclone 250 through the
direct quench line 252 to terminate the reactions of the cracked gas.
[0067] As best seen in FIGURE 9, the regeneration assembly 208 is comprised of a stripper
253, control hopper 255, entrained bed heater 258, a lift line 257, and a rengerated
solids vessel 260.
[0068] The stripper 253 is a tubular vessel into which the particulate solids from the separator
206 are delivered through solids outlet legs extending from the separator solids outlets
238 and from the cyclone diplegs 249. A ring 262 having nozzle openings 264 is provided
at the bottom of the stripper 253. A stripping gas, typically steam is delivered to
the ring 262 for discharge through the nozzles 264. The stripping steam passes upwardly
through the bed of particulate solids to remove impurities from the surface of the
particulate solids. The stripping steam and entrained impurities pass upwardly through
the particulate solids in the stripper 253 and discharge through a vent line (not
shown) to the cracked gas line.
[0069] The stripped solids are accumulated in the control hopper 255 for eventual delivery
to the entrained bed heater 258. The control hopper 255 is a collection vessel in
which solids enter through a standpipe 266 and from which an outlet line 273 extends
to deliver solids to the entrained bed heater 258. The assembly of the control hopper
255 and the standpipe 266 provides for a slumped bed solids transport system. The
pressure differential maintained between the slumped bed surface 268 in the control
hopper 255 and the exit 270 of the outlet line 273 determine the solids flow rate
between the control hopper 255 and the entrained bed heater 258. A line 272 is provided
to selectively introduce steam under pressure into the control hopper 255 to regulate
the pressure differential. Probes 267 and 269 are placed respectively in the control
hopper 255 and entrained bed heater 258 to monitor the pressure differential and regulate
a valve 265 in the steam line 272.
[0070] The entrained bed heater 258 is essentially tubular in configuration. An array of
distinct fuel nozzles 261 fed by fuel lines 263 are arranged essentially symmetrically
on the lower inclined surface 275 of the entrained bed heater 258. Pressurized air
enters the entrained bed heater 258 through a nozzle 277 arranged to direct the air
axially upwardly through the entrained bed heater 258. The air jet provides both the
motive force to lift the solids particles upwardly through the entrained bed heater
258 to the rengerated solids vessel 260 and the air necessary for combustion. The
fuel is ignited by contact with the solids in the presence of air.
[0071] The combustion gas/solids mixture moving upwardly through lift line 257 enters the
regenerated solids vessel 260 tangentially, preferably, perpendicular to the lift
line to separate the combustion gases from the solids. As shown in FIGURE 6, the vessel
260 has a distube 285 in the gas outlet nozzle 286 to provide cyclonic movement which
improves the separation efficiency of the system.
[0072] The regenerated solids vessel 260 is a cylindrical vessel provided with a standpipe
271, seen in FIGURE 7, extending to the reactor hopper 231. Again the structure of
the regenerated solids vessel 260 provides for accumulation of a slumped bed 281,
seen in FIGURE 9 above which pressure can be regulated to enable controlled delivery
of the regenerated particulate solids to the reactor hopper 231.
[0073] The upper solids collection vessel 260 seen in FIGURES 6, 7 and 9 contains a stripping
section as the lower portion with a stripping ring 279 and form a part of the solids
deliver system 210. Above ring 279, the solids are fluidized; below the ring 279 the
solids slump and are fed to the standpipe 271 shown in FIGURE 7. The standpipe 271
feeds the slumped bed in the control hopper 231 as best seen in FIGURE 7. Solids flow
into the reactor hopper 231 through the standpipe 271 to replace solids that have
flowed into the reactor 204. Unaerated solids (slumped solids) will not continue to
flow into the reactor hopper 231 once the entrance 282 to the hopper 231 has been
covered. Thus, the position of the entrance 282 defines the solids level in hopper
231. As solids flow from hopper 231 via the pressure differential between the vapor
space in the chamber 231a above the bed 218 and the mixing zone 211, the entrance
282 is uncovered allowing additional solids to flow into the hopper 231.
[0074] One embodiment of the process of the present invention as shown in the accompanying
FIGURE 1 is illustrated by the following comparative example (Table I) wherein a light
FCC naphtha is cracked employing conventional tubular pyrolysis, conventional catalytic
cracking at typical FCC residence times of greater than 1 second using moderately
active catalysts, catalytic cracking with high activity catalysts at short residence
times for FCC units (0.9 seconds), and very short residence time cracking plus quench
(QC system) with a similar high activity catalyst. Two cases employing the high activity
catalyst are shown to illustrate the effect of residence time on olefin yields.
TABLE I
|
Conventional Coil Pyrolysis |
Cat Cracker Conventional Catalyst |
Catalytic Cracker High Activity Catalyst |
QC with High Activity Catalyst |
Example: |
A |
B |
C |
D |
Feedstock: |
Light FCC Naphtha |
Residence Time (sec): |
|
|
|
|
Reactor |
0.3 |
1.0 |
|
|
Total (to quench) |
0.3 |
2.0* |
0.9* |
0.15 |
Reactor Temperature |
816 |
565 |
510 |
540 |
Conversion, Wt% |
65 |
28 |
50 |
56 |
Product Yield, Wt% |
|
|
|
|
CH₄ |
13.4 |
|
3.0 |
0.8 |
Total C2's |
15.2 |
2.6 |
13.0 |
11.2 |
C₃H₆ |
11.4 |
5.6 |
10.2 |
19.2 |
C₃H₈ |
0.6 |
4.1 |
11.6 |
8.8 |
Total C₄'s |
10.4 |
13.4 |
7.7 |
14.4 |
C₃H₈/C₃H₆ ratio** |
0.05 |
0.73 |
1.14 |
0.46 |
* no quench |
** paraffin.olefin ratio |
[0075] Referring to Table I, Example A illustrates the yields obtainable using conventional
pyrolysis operated at typical thermal cracking temperatures and residence times. Example
B illustrates a conventional catalytic riser reactor employing typically longer residence
times and lower temperatures than the pyrolysis Example A. As seen, the conventional
catalytic conversions are substantially lower than those obtained in the pyrolysis
Example A. The lower conversion is a result of the lower temperature operation (565°C
vs. 816°C) with insufficient catalytic activity for this relatively light feedstock.
Even at these low conversions however, the total C₃ and C₄ compounds are high relative
to the pyrolysis case as a result of the carbonium ion mechanism. Further, the ratio
of C₃ paraffins to C₃ olefins is substantially increased due to hydrogen transfer
activity of the acidic cracking catalyst.
[0076] Example C illustrates the product yields which will be obtained by employing high
activity acidic catalysts at low FCC residence times or high QC residence times without
quenching. The selected operating conditions of Example C will result in a suppression
of the methane and ethylene yields compared to the pyrolysis system of Example A.
The conversion is increased relative to Example B even at lower temperatures (510°C
vs. 565°C) due to the increased activity. There is a significant increase in the total
C₃ production as a result of the acidic cracking cataylst (21.8 vs. 12.0) but the
C₄ yields decrease due to the increased conversion. Further, due to the longer residence
times, there is a significant amount of hydrogen transfer as evidenced by the unacceptably
high C₃ paraffin to olefin ratio compared to either Example A or B.
[0077] Example D illustrates the dramatic improvement in olefin yields that will be obtained
by employing the process of the present invention in a very short residence time QC
system. As seen, there is about a 100% improvement in C₃ olefin yields when the reactor
temperature is increased about 30°C and the total kinetic residence time, i.e., cracking
reaction plus separation plus quench, is reduced to about 0.15 seconds. In addition,
the paraffins to olefin ratio is reduced to less than half that obtained in the longer
residence time Example C. The paraffin to olefin ratio for this case is higher than
for the pyrolysis case at a similar residence time as a result of the hydrogen transfer
activity of the catalyst. The methane yield, however, is further suppressed below
the lower level of Example C and the C₄ yields are improved by almost 100% indicating
less secondary cracking due to the quenching and short residence time reaction.
[0078] In another embodiment of the present invention, the QC system may be adapted to enhance
the production of aromatics, and specifically benzene. Table II illustrates the use
of the QC system to enhance aromatics production from the cracking of n-hexane at
a fixed 70% conversion. Two examples of the QC system, one using a highly active catalyst,
the other a deactivated zeolitic catalyst, are compared to conventional pyrolysis.
TABLE II
QC REACTOR |
|
Pyrolysis Reactor |
High Activity Catalyst |
Deactivated Catalyst |
Example: |
1 |
2 |
3 |
Temperature °C |
735 |
550 |
740 |
Total Residence Time (Sec) (including quench) |
0.2 |
0.2 |
0.2 |
Conversion |
70 |
71 |
82 |
Wt% Feed |
|
|
|
CH₄ |
6.3 |
7.0 |
9.8 |
C₂H₄ |
35.0 |
8.4 |
35.2 |
C₃H₆ |
19.8 |
27.1 |
4.6 |
C₃H₈ |
nil |
19.6 |
2.1 |
Aromatics |
nil |
2.1 |
20.9 |
C₃H₈/C₃H₆ |
nil |
0.72 |
0.46 |
[0079] Referring to Table II, Example 1, it is seen that a hydrocarbon feedstock conversion
at 70% will be obtained in conventional pyrolysis at a reactor outlet temperature
of 730°C and a residence time of 0.2 second. As indicated, pyrolysis produces significant
amounts of olefins but only trace amounts of paraffins and aromatics. Similar results
are obtained when using a completely inert solid, such as pure alumina, in a QC cracking
environment.
[0080] In the QC system using a high activity acidic cracking catalyst, 70% hydrocarbon
conversion can be obtained at a reduced temperature of 550°C and a residence time
of 0.2 second (Example 2). Ethylene production will be suppressed while the yields
of C₃ olefins and paraffins enhanced. Furthermore, only small amounts of aromatics
are produced.
[0081] If instead of a highly active catalyst, a deactivated zeolitic catalyst is used,
a completely different yield spectrum is obtained (Example 3). Zeolitic catalyst deactivation
is usually a result of prolonged exposure to high temperatures and steam causing the
zeolite matrix to collapse. This results in a significant reduction in catalyst surface
area and hence catalyst activity. Typically in an FCC unit which uses zeolitic catalysts,
the catalytic solids are withdrawn and fresh catalyst is added to maintain activity.
Such "spent" solids are suitable for use as deactivated catalysts.
[0082] The reaction temperatures required are similar to those required to achieve pyrolysis
conversion due to the low activity. However, quite unexpectedly, there is a substantial
increase in aromatics production (specifically benzene) and a corresponding decrease
in C₃ and C₄ olefin production. The ethylene yields are similar to those from pyrolysis
given the predominance of the free radical cracking reactions at these temperatures.
However, the deactivated catalyst provides enhanced aromatization activity at these
higher temperatures and thus aromatics are formed at the expense of the C₃ and C₄
olefins and paraffins.
[0083] These unexpected results will thus enable an operator to vary the operating conditions
of the QC system to either select high C₃ to C₅ olefin production or high aromatic
production in accordance with the present invention, depending on the desired product,
the available feedstocks and the choice of catalyst.
[0084] Specific embodiments of the invention have been described and shown in the above
examples to illustrate the application of the inventive principles. The invention
in its broader aspects is not limited to the specific described embodiments and departures
may be made therefrom within the scope of the accompanying claims without departing
from the principles of the invention and without sacrificing its chief advantages.
[0085] In another embodiment of the invention a dehydrogenation catalyst is combined with
an acidic cracking catalyst.
TABLE III
MINAS NAPHTHA CRACKING |
|
Coil Pyrolysis |
Moderate Activity Acidic Cracking Catalyst |
|
Residence Time (Sec) |
0.2 |
0.2 |
|
Temp C |
827 |
746 |
|
|
|
|
Yield with Catalyst |
|
|
|
Yield from Coil Cracking |
Yield, wt% |
|
|
|
CH₄ |
12.5 |
9.5 |
0.76 |
C₂H₄ |
23.0 |
17.3 |
0.75 |
C₂H₆ |
3.7 |
2.8 |
0.76 |
C₃H₆ |
13.4 |
13.7 |
1.20 |
C₃H₈ |
0.5 |
3.0 |
1.20 |
C₄H₆ |
4.2 |
0.7 |
|
C₄H₈ |
4.5 |
7.3 |
1.53 |
C₄H₁₀ |
0.7 |
6.7 |
|
|
62.5 |
61.0 |
|
[0086] The example shown in Table III uses a Minas naphtha feedstock and compares cracking
both catalytically and thermally. The catalytic case requires a lower temperature
to achieve the given conversion thus will have in this case only 75% of the thermal
products (C₁ and C₂ compounds). The carbonium ion cracking will shift the yield spectrum
to favor C₃ and C₄ compounds.
TABLE IV
ISO/NORMAL C₄ YIELDS FROM MINAS NAPTHA CRACKING (REFERENCE TABLE III) |
|
Coil Pyrolysis |
Acidic Cat Only |
Acidic Cat plus* Dehydrog Cat |
I-C₄H₈ |
1.5 |
1.84 |
5.02 |
N-C₄H₈ |
3.0 |
5.44 |
6.71 |
I-C₄H₁₀ |
0.23 |
4.43 |
1.25 |
N-C₄H₁₀ |
0.47 |
2.28 |
1.00 |
Iso (% olefin) |
87 |
29 |
80 |
Nor (% olefin) |
87 |
70 |
87 |
* Mixture of Sn and Cr oxides on an alumina support |
[0087] As shown in Table IV, use of acidic catalyst alone results in a very significant
increase in total iso C₄'s (paraffins plus olefins) due to the ionic nature of the
cracking. However, most of the iso compounds appear as iso-paraffins whereas the thermodynamic
equilibrium exclusively favors the production of olefins, not paraffins. In the case
of the coil pyrolysis, both the iso and normal C₄'s are 87% olefinic indicating a
reasonable approach to equilibrium. For the catalyst case however, the normal compounds
are only 70% olefinic and the iso compounds only 29% olefinic. This is because the
hydrogen transfer activity of the catalyst results in a "new" equilibrium relationship
based on reaction kinetics rather than thermodynamics. Iso compounds show a much greater
tendency to undergo hydrogen transfer in the presence of the acidic catalyst than
normal compounds.
[0088] When noble metal oxide catalysts are added to the flowing acidic cracking catalysts,
the paraffins produced at the acidic sites can be dehydrogenated to their olefinic
counterpart. The extent to which this occurs is dependent upon the concentration and
activity of the dehydrogenation catalyst. In Table IV, a dehydrogenation catalyst
consisting of oxides of tin (Sn) and chromium (Cr) is mixed with the acidic cracking
catalyst to achieve an 80% approach to equilibrium for the iso compounds and a corresponding
87% approach for the normal compounds. As can be seen, the production of the valuable
C₄ olefins, both normal and iso, are signficantly increased. The isobutylene production
from the same feed is increased by a factor of over 3 and the normal butene by a factor
of over 2. The use of mixed catalyst systems provides an additional product distribution
flexibility for the catalytic process. Rather than admix the dehydrogenation catalyst
with the acidic cracking catalyst and follow the mix reaction with a quench, using
steam for example, the dehydrogenation catalyst can be located in a packed bed within
a catalyst reactor 25 located downstream of the primary separation in separator 6.
The paraffins formed by contact with the acidic catalyst will be dehydrogenated to
their olefin counterpart.
1. A process for catalytically cracking a hydrocarbon feedstock to produce olefins,
aromatics or a combination thereof comprising:
(a) introducing the hydrocarbon feedstock to a cracking reactor;
(b) simultaneously delivering a hot acidic cracking catalyst solids to the cracking
reactor;
(c) catalytically cracking the hydrocarbon feedstock with the heat supplied by the
hot catalyst solids to form a cracked product;
(d) separating the cracked product from the hot catalyst solids; and
(e) quenching the separated cracked product;
wherein the residence time of the hydrocarbon feedstock from step (a) through step
(e) is in the range of from about 0.1 to 2.0 seconds.
2. The process of Claim 1, wherein the residence time is from about 0.05 to 0.5 seconds.
3. The process of Claim 1 further comprising:
(f) delivering the separated catalyst solids to a stripper to remove residual cracked
gas products;
(g) combusting the separated catalyst solids to thereby remove carbon deposits and
to heat the stripped catalyst solids to thereby form regenerated catalyst solids;
and
(h) transporting the regenerated catalyst solids to the cracking reactor.
4. The process of Claim 1, wherein a portion of the catalyst solids is comprised of
a dehydrogenation catalyst.
5. The process of Claim 4, wherein the dehydrogenation catalyst is selected from noble
metal oxides on an inert carrier.
6. The process of Claim 1, wherein the temperature of the catalytic cracking reaction
is from about 900 to 1500°F., and the weight ratio of catalyst solids to hydrocarbon
feedstock is between 1 and 60.
7. The process of Claim 6, wherein the temperature of the catalytic cracking reaction
is from about 1000°F to 1300°F, and the residence time is 0.1 to 0.3 seconds.
8. The process of Claim 1, wherein the olefin has 3 to 5 carbon atoms and the aromatic
compound is has 6 to 8 carbon atoms.
9. The process of Claim 1, wherein the hydrocarbon feedstock is selected from C₄-C₇
paraffins, naphthas and light gas oils.
10. The process of Claim 1, wherein the active acidic catalyst solid is a zeolitic
catalyst.
11. The process of Claim 10, wherein the catalyst support is selected from the group
consisting of silica gel, silica-alumina, clays or a mixture of any of the foregoing.
12. The process of Claim 11, wherein the cracked product is primarily mono-aromatics
and said acidic catalyst solids are thermally deactivated.
13. The process of Claim 1, comprising delivering the hydrocarbon feed stream and
hot catalyst solids to a tubular thermal regenerative cracking reactor through a reactor
feeder having vertical passages communicating with the tubular regenerative cracking
reactor and the solids in a hot solids vessel, providing localized fluidization of
the solids above the vertical passages, and delivering the hydrocarbon feed to the
tubular thermal regenerative reactor at an angle to the path of the catalyst solids
entering the thermal regenerative reactor.
14. The process of Claim 1, comprising separating the hot catalyst solids and the
cracked product gases in a separator wherein the catalyst solids and cracked product
gases enter the separator through a separator inlet and reverse direction ninety degrees
and then the product gases reverse direction another ninety degrees to effect a one
hundred eighty degree reversal in direction from the entry direction and then the
catalyst solids continue in the path oriented ninety degrees from the catalyst solids
cracked product gas separator inlet and thereafter, the path of the catalyst solids
is directed downwardly and the separated product gases are quenched.
15. The process of Claim 1, comprising separating the catalyst solids and cracked
gases in a separator comprising a chamber for rapidly disengaging about 80% of the
catalyst solids from the incoming mixed phase stream, said chamber having approximately
rectilinear longitudinal side walls to form a flow path of height H and width W approximately
rectangular in cross section, said chamber also having a mixed phase inlet of inside
width Di a gas outlet and a solids outlet, said inlet being at one end of the chamber and
disposed normal to the flow path of height H which is equal to at least Di, or 4 inches, whichever is greater, and the width W is from 0.75 Di to 1.25 Di said solids outlet being at the opposite end of the chamber and being suitably arranged
for downflow of discharged solids by gravity, and said gas outlet being therebetween
at a distance no greater than 4D from the inlet as measured between respective centerlines
and oriented to effect a 180° change in direction of the gas whereby resultant centrifugal
forces direct the catalyst solids in the incoming stream toward a wall of the chamber
opposite to the inlet forming thereat and maintaining an essentially static bed of
solids, the surface of the bed defining a curvilinear path of an arc of approximately
90° of a circle for the outflow of solids to the solids outlet.