Note: Descriptions are shown in the official language in which they were submitted.
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HEAVY OIL CRACKING METHOD
CROSS REFERENCE TO RELATED APPLICATIONS
[0001] This application claims priority from US Provisional Application Serial
No.
61/223,885 filed July 8, 2009.
BACKGROUND OF THE INVENTION
[0002] The invention is directed to a method for cracking heavy oil. More
particularly, the invention is directed to a method for cracking heavy oil
using staging
to convert the heavy oil feed to a cracked pitch product.
[0003] A recent review article [Hulet (2005)] examined the key features and
configurations of short residence time cracking processes developed over the
past
25 years. This work succinctly summarized the promise, key features; and
challenges of short residence time processes:
"There is a strong economic incentive for considering short residence time
cracking
processes. Not only do such processes increase the yields of the more valuable
liquid
and gaseous products, but more compact designs would also decrease capital
costs.
Careful control of the vapor residence times appears to be crucial in order to
prevent
secondary cracking and yet allow for maximum cracking of the feedstock. Rapid
and
thorough mixing of the feedstock with the heat source, not just creating a
uniform
dispersion, is also a key design aspect to consider. Finally, rapid and
complete
separation must also be carefully considered; again, to help control product
residence
time and avoid secondary cracking but also from a heat balance point of view."
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[0004] These insights suggest short residence time cracking processes should
rapidly heat the feed, rapidly separate products to control residence time of
all
products, rapidly cool the reaction products to avoid secondary cracking, and
make
productive use of the thermal energy. The most successful short residence time
processes meet all these criteria. However, none of these processes fully meet
all
these criteria while treating asphaltic residual oils. Fully meeting these
criteria with
asphaltic residual oils is the goal for more effective deasphalting, thermal
cracking,
and hydrocracking processes.
[0005] Fluid catalytic cracking (FCC) is undoubtedly the most common and
commercially successful short residence time cracking process. Typically, the
FCC
process intimately contacts a gas oil boiling range hydrocarbon feedstock with
hot
catalyst particles in an entrained flow, short-residence, riser reactor to
produce more
valuable cracked products, particularly gasoline and olefins, and less
desirable dry
gas and coke by-products. The FCC process developers have used improved feed
nozzle designs to increase feed heating rate. The FCC feed nozzles improve the
uniformity of the initial contact between the carbonaceous feed and the hot
regenerated catalyst, which increases the feed heating rate and decreases the
yield
of the undesirable dry gas and coke FCC produces. For example, US Pat. No.
6,387,247, summarizes a long standing effort to use feed injection nozzle
improvements to increase the feedstock heating rate and improve the overall
FCC
reactor performance.
[0006] FCC process developers have also identified approaches to control the
residence time. For example, US Pat. No. 6,979,360, teaches methods for
conducting short contact time hydrocarbon conversions in a FCC reactor with
the
rapid inertial separation of gas and solid FCC reactor products. US Pat. No.
6,616,900 extends this concept by using a staged FCC riser reactor with
interstage
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product removal. US Pat. No. 5,762,882, teaches methods to remove reaction
products from the spent catalyst via vaporization. With distillate feeds, the
by-product coke production is roughly in balance with FCC process heat
requirement. The higher coke yield associated with more asphaltic residual FCC
feeds has a large adverse effect on the process performance. US Pat. No.
4,415,438 teaches the use of a thermally stable catalyst and high catalyst
regeneration temperature to increase the heavy oil feed heating rate and
decrease
coke yield. US Pat. No. 5,271,826 achieves a similar result by increasing the
regenerated catalyst to feed ratio to achieve an elevated riser initial
temperature and
then adding a quench liquid to temperature the riser temperature. Canadian
Patent
No. 2,369,288, teaches thermal cracking of residual oil feed with inert solids
in an
FCC reactor-type short contact time reactor to eliminate catalyst deactivation
problems, but also results in an inferior product yield distribution,
including coke
production in excess of the process heat requirement. US Patent Application
Publication No. 2006/0042999 teaches deasphalting of the FCC heavy oil feed to
decrease the catalyst deactivation rate due to metals and coke precursors in
the
feed. US Pat. No. 6,171,471 teaches the combination of mild hydrocracking and
deasphalting of the residual oil FCC feed to decrease the catalyst
deactivation rate
due to metals and coke precursors. Despite these efforts to decrease the FCC
process coke yield with residual asphaltic feeds, the coke yield far exceeds
the
amount required to preheat the regenerated catalyst or inert solids.
[0007] As a result, similar principles were used to develop successful short
contact time fluid coking processes that maximize the conversion of asphaltic
residual oil feeds to distillates. The fluid coking process typically
comprises partial
combustion of coke particles, rapid heating of the residual oil feed by
intimate
contact with a fluidized bed of hot coke particles, rapid separation of
entrained
by-product coke from the vapor product using cyclones, and a quench system to
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rapidly cool the vapor product to minimize secondary thermal cracking. US Pat.
No.
2,881,130 teaches atomization and distribution of the residual oil feed to
prevent
bogging of the fluidized coke bed and to increase the residual oil feed
heating rate.
The fluid coking process developers have also identified methods to control
residence time. For example, US Pat. No. 4,816,136 teaches sequentially
contacting
the feed with higher temperature coke particles in a riser reactor and then in
a
fluidized coke bed that operates at a lower temperature to increase the
distillate
yield. US Pat. No. 5,658,455 describes a method for a short vapor residence
time
reactor to minimize secondary thermal cracking reactions. US Pat. No.
4,497,705
teaches methods to solvent refine the recycle heavy oil to selectively remove
less
carbonaceous species to decrease undesirable secondary cracking of these
valuable
products. US Pat. No. 4,587,010 teaches methods to strip valuable products
from
the coke prior to regeneration via partial oxidation. The fluid coking has
many
common features with the FCC process and has several advantages for treating
carbonaceous residual oil feeds. The fluid coking process eliminates the rapid
catalyst deactivation problem and need to burn very large quantities of coke
that are
associated with the FCC process treating asphaltic residual oil feeds. Since
both the
fluid coking and FCC processes require that all liquid products are produced
by
vaporization, neither process can operate with a short residence time for
unconverted asphaltic residual oils.
[0008] US Pat. No. 3,393,133 teaches high temperature and short residence time
distillation processes to maximize distillate yield with minimum degradation
of the
residual oil due to thermal cracking reactions. However, solvent extraction
[Altgelt
(1994)] is the preferred method to produce residual oil fractions with much
higher
equivalent normal boiling points and essentially no thermal cracking
degradation.
Successful solvent refining processes have been developed, e.g. US Pat. No.
4,810,367, to continuously produce deasphalted oil, resin, and asphaltene
streams
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from an asphaltic residual oil feed. These processes require a large number,
but
simple and reliable, unit operations to contact and separate the residual oil
from the
solvents. In addition, the solvent and residual oil separation steps have a
significant
steam heat requirement. As a result, this process technique is particularly
useful in
petroleum refineries, where the required solvents, steam, and maintenance
infrastructure are readily available. US Pat. No. 6,357,526 teaches a solvent
extraction field upgrader method to produce a deasphalted oil synthetic crude
product and an asphaltic fuel to produce steam for bitumen extraction. For
this
remote application, a flash deasphalting process has many potentially
desirable
features. A very high temperature, very short contact time flash unit
operation could
produce the deasphalted oil and asphaltic streams in a compact and single unit
operation without the need for a solvent. The thermal energy input could be
used to
separate the deasphalted oil and asphaltene products, produce steam for
bitumen
extraction, and produce a hot asphaltene stream that can be burned without the
need for reheating or pelletization. Unfortunately, the earlier processes do
not
provide any method that can heat the bitumen feed, separate the deasphalted
oil
and asphaltene products, and cool the separated deasphalted oil and asphaltene
products sufficiently rapidly to avoid excessive thermal cracking and
degradation of
the deasphalted oil product.
[0009] Resid hydrocracking is the most well established method to convert
asphaltic materials to less carbonaceous materials and to reduce the metals
and
coke precursor concentration in the unconverted asphaltene species. US Pat.
No.
2,987,465 first introduced the ebullated bed hydrocracking reactor concept. An
ebullated hydrogenation reactor utilizes up-flow of the carbonaceous asphaltic
residual oil and hydrogen feeds to contact an expanded bed of particulate
hydrotreating catalyst and or entrained colloidal hydrotreating catalyst. The
expanded
bed hydrotreating catalyst bed is much less susceptible to plugging than the
previous
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fixed catalyst bed designs. US Pat. No. 5,164,075 teaches methods to produce
colloidal heavy oil catalysts that are particularly effective for
hydrogenating asphaltic
species. US Pat. No. 6,511,937 teaches methods to recover and recycle
colloidal
heavy oil hydrocracking catalysts. All these resid hydrocracking process
simultaneously hydrogenate and thermally crack the residual oil. US Pat. No.
4,427,535 identifies a fundamental limitation with this approach. The thermal
cracking reactions have higher activation energies than the hydrogenation
reactions.
Hydrogenation reactions retard polymerization reactions that produce coke
precursors and ultimately coke. Therefore, a resid hydrocracker must be
operated at
a much lower temperature than a FCC or fluid coking unit in order to maintain
operability.
[0010] US Patent Application Publication No. 2005241993 teaches the use of a
colloidal molybdenum sulfide catalyst to increase the hydrogenation rate,
particularly
the rate of hydrogenation of asphaltic species. This innovation increases
hydrogenation rate of asphaltic species and the maximum operable temperature,
but
does not alter the basic nature of this hydrocracker operating temperature
limitation.
Hydrogen donor diluent cracking processes substantially eliminates this
temperature
limitation by hydrogenating a naphthenic distillate at moderate temperatures
to
produce a hydrogen donor solvent and then thermal cracking the residual oil in
the
presence of the donor solvent to substantially reduce the rate of coke
precursor
formation. US Pat. No. 4,698,147 teaches simultaneously increasing the
hydrogen
donor diluent cracking temperature and decreasing the contact time to
monotonically
increase the maximum resid conversion to distillates. US Pat. No. 4,002,556
teaches that high temperature and short contact time hydrogen donor diluent
cracking also substantially reduces the hydrogen consumption required to
maintain
operability. These high temperature and short contact time benefits seem to be
only
limited by the practical limits on the feed heating rate and the product
cooling rate.
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The hydrogen donor diluent process requires sufficient pressure to maintain a
liquid
phase hydrogen donor solvent at the elevated thermal cracking temperatures.
[0011] The invention is able to address the concerns of these earlier
processes
by using a staging process to convert heavy oil into a cracked distillate.
SUMMARY OF THE INVENTION
[0012] The invention provides for advantages to the process of converting
residual heavy oil. The invention improves control over the reactor
composition. The
staging process provides for prompt removal of cracked distillate product to
inhibit
secondary cracking and increases the solubility of coke precursor in the
liquid phase.
Thermal efficiencies are also realized by the invention.
[0013] The invention further provides for a method for cracking heavy oil
comprising feeding residual heavy oil first to a heating stage comprising an
atomizer,
a contactor and an inertial vapor-liquid separator, and then to at least one
stage
selected from the group consisting of temperature maintenance stage comprising
a
reactor and an inertial vapor-liquid separator, and a second heating stage
comprises
an atomizer, a contactor and an inertial vapor-liquid separator
[0014] The invention provides for a method of cracking residual heavy oil
comprising at least one heating stage and at least one or more of a stage
selected
from the group consisting of a heating stage and a temperature maintenance
stage.
[0015] The heating stage is accomplished by three unit operations comprising
an
atomizer, a contactor and an inertial vapor-liquid separator.
[0016] The temperature maintenance stage is accomplished by two unit
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operations comprising a reactor and an inertial vapor-liquid separator.
[0017] The method of cracking residual heavy oil can be accomplished by any
combination of the two stages can be employed so two heating stages may be
employed or two temperature maintenance stages or one or more of each stage.
Thus numerous combinations of heating stages and temperature maintenance
stages can be employed as long as there is more than one stage used in the
cracking process.
[0018] The heating stage starts with a feed of a heating gas formed from the
reaction products of an oxygen-containing compound and a fuel. This reaction
occurs in a heating pressurized combustor. The heating gas is fed to the
atomizer
along with a residual heavy oil feed.
[0019] The now atomized residual heavy oil is fed to a contactor where the
residual heavy oil is maintained for sufficient residence time to achieve the
desired
conversion of residual oil to distillates via thermal cracking reactions.
[0020] The products from the contactor are then fed to the inertial vapor-
liquid
separator where the products are separated into vapor and liquid portions.
[0021] The vaporized gas and liquid distillate products are removed from the
inertial vapor-liquid separator and the remaining residual oil liquid is fed
to the
second stage of the method, the temperature maintenance stage.
[0022] The first step of the temperature maintenance stage is the reactor step
where the reactor receives the liquid product from the inertial vapor-liquid
separator
of the heating stage. The reactor will provide sufficient residence time to
achieve the
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desired cracking.
[0023] The cracked reaction products are then fed to an inertial vapor-liquid
separator where they will mix with the reaction products from a cracking
pressurized
combustor. The liquid portion and vapor portion are separated and the gas and
cracked distillate product recovered.
BRIEF DESCRIPTION OF THE DRAWINGS
[0024] Fig. 1 is a schematic view of a pressurized combustor.
[0025] Fig. 2A is a schematic view of a conical atomizer.
[0026] Fig. 2B is a schematic view of a cylindrical atomizer.
[0027] Fig. 3 is a schematic view of a vapor-liquid contactor.
[0028] Fig. 4 is a schematic view of an inertial vapor-liquid separator.
[0029] Fig. 5 is a schematic view of a two stage heating and thermal cracking
process according to the invention.
DETAILED DESCRIPTION OF THE INVENTION
[0030] The invention will now be described in more detail and with reference
to
the drawing figures.
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[0031] The pressurized combustor on Figure 1 mixes a fuel 1 with an oxidant 2
in
a burner 3. The fuel may be any hydrocarbon and/or hydrogen. The fuel is
preferably a H2 and CO synthesis gas that is produced by conventional
gasification
of the pitch by-product (see Stream 40 on Figure 5). The oxidant 2 may be any
mixture of air, 02, and steam. The burner 3 utilizes conventional ignition and
flame
monitoring methods to maintain a stable flame 4. The combustion reactions are
substantially completed within a pressurized shell 5. The pressurized shell 5
may be
advantageously fitted with internal insulation 6 to decrease heat loss and the
pressurized combustor shell 5 temperature. The pressurized combustor product
gas
7 would typically be in the 5 to 20 bar pressure range, 1400-1800 C
temperature
range, and 0 to -5% excess 02.
[0032] Typical heavy oil atomizers are illustrated on Figures 2A and 2B.
Figure
2A illustrates a conical convergent-divergent nozzle atomizer and Figure 2B
illustrates a cylindrical convergent-divergent nozzle atomizer. The
pressurized
combustor product gas 7 is fed to the nozzle motive plenum 8. Similar and well
established methods can be used to design both types of nozzles in terms of
the
motive gas plenum 8 temperature, pressure, and composition, the nozzle throat
9
area, the expansion angle 10, and nozzle 11 discharge area.
[0033] In the case of the conical nozzle (see Figure 2A), the atomization
nozzle
heavy oil feed 12 is directed by the heavy oil conduit 13 to the conical
nozzle mixing
chamber 14. The heavy oil atomization effectiveness is a complicated function
of the
conical gas jet 15 kinetic energy and temperature, conical gas jet
intersection angle
16, and conical nozzle mixing chamber 14 volume and shape. Generally, the
conical
atomization nozzle efficiency increases with increasing the conical gas jet 15
kinetic
energy and temperature, conical gas jet intersection angle 16, and decreasing
conical nozzle mixing chamber 14 volume. The atomized heavy oil product 17
from
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the conical nozzle is fed to the vapor-liquid contactor or reactor 18 (see
Figure 3).
[0034] The vapor-liquid contactor or reactor 18 is illustrated on Figure 3.
This
vessel is considered a vapor-liquid contactor if its primary function is to
provide the
atomized heavy oil sufficient residence time to approach the temperature of
the gas
phase that is substantially below design thermal cracking temperature,
typically
between 650 C and 850 C. This vessel is considered to be a reactor if the
heavy oil
temperature is near the design thermal cracking temperature. The vapor-liquid
contactor or reactor 18 is typically a cylindrical pressure vessel with a feed
inlet
nozzle 19 and a product outlet nozzle 20. The vapor-liquid contactor or
reactor 18
may be advantageously equipped with internal insulation 21 to decrease the
heat
loss and the vapor-liquid contactor or reactor 18 shell temperature. The
diameter of
the vapor-liquid contactor or reactor 18 diameter is set to provide highly
turbulent
flow with a reasonable pressure drop. The vapor-liquid contactor or reactor 18
length
is set to achieve the desired gas less atomized heavy oil temperature
difference
when operating in the vapor-liquid contactor operating mode or the desired
thermal
cracking extent of reaction prior to separating the cracked vapor product from
the
unconverted heavy oil when operating in the reactor operating mode.
[0035] The preferred inertial vapor-liquid separator is a cyclone. Figure 4
illustrates the key features of the cyclone inertial vapor-liquid separator.
The general
shape of the cyclone inertial separator 22 is a truncated cone with the
smaller
diameter at the feed end 23 and the larger diameter at the product end 24. The
product end 24 diameter would generally be between 1.1 and 1.4 times the feed
end
23 diameter to ensure a low liquid inventory and residence time in the vapor-
liquid
separator. Internal insulation (not shown on Figure 4) may be advantageously
used
to decrease inertial separator heat loss and wall temperature. A rectangular
conduit
25 is preferably used to shape the inertial vapor-liquid separator feed 26.
The inertial
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vapor-liquid separator is advantageously fitted with a weir 27 to minimize
short-
circuiting of liquid droplets in the feed 26 to the vapor outlet 28. A
distillate quench
stream 29 may be advantageously used to decrease the vapor product 30
temperature to between 300 and 4000C to control the rate of thermal cracking
reactions in the vapor product 30.
[0036] An atomizer is advantageously used to increase the interfacial area of
the
hydrocarbon quench liquid 29. The hydrocarbon quench liquid preferably has a
normal boiling point less than 350 C. The liquid product conduit 32 is used to
remove the liquid product 33 from the inertial vapor-liquid separator. A
liquid product
conditioning stream 34 may be advantageously introduced via the liquid product
conditioning stream conduit 35 to achieve the desired liquid product 33
properties.
For example, one could advantageously use steam as the liquid product
conditioning
stream 34 to minimize the quantity of the vapor in the feed stream 26 that
exits in the
liquid product stream 33 and cools the liquid product stream 33 to control the
thermal
cracking rate. Alternatively, one could advantageously use the pressurized
combustor product gas 7 on Figure 1 to control the quantity of the vapor in
the feed
stream 26 that exits in the liquid product stream 33 and heat the stream 33 to
compensate for heat losses and the endothermic thermal cracking heat of
reaction.
Typically, the inertial separator feed 26, the liquid product conditioning
stream 34
would all have either a clockwise or counterclockwise flow. One can also
advantageously improve liquid drainage and decrease the vapor-liquid separator
volume by using a product end insert 36. The relative dimensions in Figure 4
are
only intended to provide rough guidance for the inertial vapor-liquid
separator. The
vapor residence time in the inertial vapor-liquid separator would typically be
between
1 and 5 milliseconds.
[0037] This process description will describe the options to use these unit
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operations for a staged thermal cracking with the aid of Figure 5.
[0038] The stage cracking method (See Figure 5) seeks to maximize the
conversion of the residual heavy oil feed 37 to substantially uncracked
distillate
products 38 and cracked distillate products 39 and minimize the pitch product
yield.
The staged heavy oil cracking system on Figure 5 utilizes two heating stages
and two
thermal cracking stages. This staged method consists of at least one
preheating and
one thermal cracking stage. A more typical arrangement would consist of two
heating and two thermal cracking stages, as shown on Figure 5. Up to four
heating
and cracking stages can be advantageously used. A heating stage would typical
consist of a conical atomizer (see Figure 2A), a contactor vessel (see Figure
3), and
a cyclone inertial vapor-liquid separator (See Figure 4). The residual heavy
oil feed
37 and 1St heating stage atomizing and heating gas 42 are introduced into the
1st
heating stage atomizer 41.
[0039] The asphaltic residual oil feed 37 typically contains more than 25
weight
percent of species with normal boiling points greater than 975 F (524 C), more
preferably greater than 50 weight percent, and most preferably greater than 75
weight percent. The residual oil feed typically has micro carbon and heptane
insoluble contents between 5 and 40 weight percent. The more carbonaceous
feeds
typically exhibit higher micro carbon and heptane insoluble values. Typical
sources
for residual oil feed include petroleum atmospheric or vacuum residual oil,
oil sands,
bitumen, tar sand oils, coal tar, pyrolysis tars, or shale oils.
[0040] The residual oil feed 37 may also be partially hydrogenated prior
entering
atomization-heating nozzle 41. The optional residual oil feed hydrogenation
step
typically has a hydrogen consumption between 100 and 1500 standard cubic feet
per
petroleum barrel (SCF/bbl) or between 95 and 285 gram moles per cubic meter,
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more preferably between 150 and 1000 SCF/bbl, most preferably between 200 and
700 SCF/bbl. The optional hydrogenation step is preferably performed in an
ebullated bed hydrotreating reactor with conventional cobalt-molybdate or
nickel-
molybdate on activated alumina supported hydrotreating catalyst and/or
colloidal
hydrogenation catalyst system, preferably a molybdenum sulfide colloidal
catalyst.
[0041] The residual oil feed 37 is preheated in a conventional fired heater
(not
shown on Figure 5) to the maximum temperature consistent with modest thermal
cracking rates, typically in the 350 to 450 C temperature range.
[0042] The 1st heating stage atomizing and heating gas 42 is typically
produced in
the heating stage pressurized combustor using air oxidant 44 with optional
steam
additions to prevent sooting and a convert fuel 45. Alternatively, one could
use an
02-steam oxidant. The 02 flow rate in the oxidant 44 feed to the pressurized
combustor 43 is typically between 90% and 100% of the theoretical flow rate
required to convert the hydrogen and/or carbon in the fuel feed 45 to CO2 and
H2O.
The 1st heating stage atomizing and heating gas 42 typically is in the
temperature
range of 1300-1800 C and the pressure range of 5 to 20 bar. The heating stage
pressurized combustor fuel 45 may be any hydrocarbon fuel or hydrogen. The
preferred fuel is a H2 and CO synthesis gas produced by gasification of the
pitch
product 40.
[0043] The 1St heating stage atomizer 41 (see Figure 2A) uses the 1st heating
stage atomizing and heating gas 42 to atomize the residual heavy oil feed 37
and 1st
heating stage contactor 46 (See Figure 3) provides sufficient residence so
that the oil
droplets approach the gas temperature, typically within 10-20 C, which would
generally be accomplished within a 1-10 millisecond time frame.
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[0044] The 1st heating stage contactor product 47 is promptly fed to the 1st
heating stage inertial separator 48 (see Figure 4). The 1st heating stage
inertial
separator 48 rapidly (typically less than 5 milliseconds) separates the vapor
and
liquid portions of the 1st heating stage contactor product 47. The 1st heating
stage
inertial separator quench 49 promptly reduces the 1st heating stage inertial
separator
distillate product 50 to about 350 C to control the extent of thermal cracking
of the
distillate product 50.
[0045] The operation of the 2"d heating stage atomizer 52, contactor 54, and
inertial separator 56 are similar to the operation of the corresponding unit
operations
in the 1st heating stage with some minor exceptions. First, the ratio of the
flow rate of
the atomizing and heating gas for the 1st stage 42 is typically substantially
higher
than the flow rate for the 2nd stage 53. Generally, these flow rates are
adjusted such
that the temperature of the 2"d heating stage inertial separator liquid
product 59 has
reached the desired thermal cracking temperature, typically in the 650 to 850
C
temperature range, and the vaporized distillate flow rates in streams 50 and
58 are
roughly equivalent. In addition, the residence time in the 2"d heating stage
contactor
54 is less than the 1st heating stage contactor in order to control the extent
of thermal
racking. As a result, the temperature difference of the vapor and liquid
portion of the
2"d heating stage contactor product 55 is greater than the 1st heating stage
contactor
product 47.
[0046] 1st cracking stage reheat and purge gas 60 is typically produced in the
cracking stage pressurized combustor 61 using a steam-02 oxidant 62 and a
convent
fuel 63. The 02 flow rate in the oxidant 62 feed to the cracking pressurized
combustor 61 is typically between 90% and 100% of the theoretical flow rate
required to convert the hydrogen and/or carbon in the fuel feed 63 to CO2 and
H2O.
The 1st cracking stage reheat and purge gas 60 typically is in the temperature
range
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of 1000-1600 C and the pressure range of 2-10 bar. The lower temperature and
pressure of the 1st cracking stage reheat and purge gas 60 relative to 1St
heating
stage atomizing and heating gas 42 results from the fact that this stream must
only
provide sufficient thermal energy to compensate for modest heat losses and
heat of
cracking in order to maintain the desired thermal cracking temperature. This
lower
heating requirement also eliminates the need for an atomizer. The cracking
stage
pressurized combustor fuel 63 may be any hydrocarbon fuel or hydrogen. The
preferred fuel is a H2 and CO synthesis gas produced by gasification of the
pitch
product 40. A water quench stream 64 may be advantageously used to obtain the
desired 1St cracking stage reheat and purge gas 60 temperature.
[0047] The 1St cracking stage reheat and purge gas 60 in fed to the terminal
reheating inertial separator 56 as shown on Figure 4 and described above in
order to
minimum the nitrogen and uncrack distillate content of the 2nd heating stage
inertial
separator liquid product 59. A minority, typically less than 20%, of the 1St
cracking
stage reheat and purge gas exits 2nd heating stage inertial separator 56 via
stream
58. The 2"d heating stage inertial separator liquid product 59 is promptly fed
to the
1 st cracking stage reactor 65. The diameter of the 1 st cracking stage
reactor 65 is
adjusted to provide a highly turbulent flow regime in the 1st cracking stage
reactor 65
to minimize liquid hold-up on the wall and to achieve reasonable heat transfer
from
the gas to the liquid phase. The length of the 1st cracking stage reactor 65
is
adjusted to provide sufficient residence time to achieve the desired
conversion. The
1St cracking stage reactor product 66 is feed to the 1St cracking stage
inertial
separator 67, which is operated in the same manner at the 2"d heating stage
inertial
separator 56. Likewise, the 2"d cracking stage is operated in the same manner
as
the first stage except a pitch product 78 quench, usually water, can be
advantageously used to reduce the pitch product 40 temperature to about 350 C
and
minimize further degradation.
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[0048] As noted in Figure 5, the second stage cracking reheat and purge gas
enters cyclone 67 through line 71. Gas and cracked pitch product will exit
cyclone 67
through line 70 to reactor 72 where it will continue through line 73 to
cyclone 74. The
second cracking stage distillate quench is fed through line 75 to cyclone unit
74.
This feed is accompanied by the feed from reactor 72 which enters cyclone 74
through line 73. Gas and cracked distillate product leave through line 76 to
line 39
and exit the system. The gas and cracked pitch product will exit through line
77
where it is quenched by a pitch product quench entering through line 78. The
now
quenched gas and cracked pitch product will exit the system through line 40.
[0049] While this invention has been described with respect to particular
embodiments thereof, it is apparent that numerous other forms and
modifications of
the invention will be obvious to those skilled in the art. The appended claims
in this
invention generally should be construed to cover all such obvious forms and
modifications which are within the true spirit and scope of the present
invention.
-17-
