Note: Descriptions are shown in the official language in which they were submitted.
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CATALYTIC HYDROGENATION PROCESS UTILIZING MULTI-STAGE
EBULLATED BED REACTORS
BACKGROUND OF INVENTION
This invention pertains to improved catalytic hydrogenation of heavy
hydrocarbonaceous feedstocks utilizing catalytic multi-stage ebullated bed
reactors for producing desired lower boiling hydrocarbon liquid products. It
pertains particularly to such catalytic multi-stage hydrogenation processes
having increased catalyst loading and liquid volume together with reduced gas
hold-up in each reactor, and thereby provides improved performance efficiency
for the processes.
In conventional catalytic hydrogenation processes for heavy hydrocarbon
feedstocks utilizing multi-stage ebullated bed reactors, the hydrogen gas
recycle
rate in each reactor is usually kept relatively high to assure that excess
hydrogen gas exists in the catalyst bed to provide the necessary chemical
hydrogenation reactions with the feedstock. However, such excess hydrogen
flow requires relatively high superficial gas velocities in the reactor(s),
which
results in less available volume for the reacting liquid and increased gas
hold-up
in the reactor. Because the feedstock hydrogenation and hydrocracking
reactions occur predominantly in the liquid phase, this conventional practice
has
the result of undesirably reducing the percentage of feedstock liquid being
exposed to and reacted with the catalyst in the reactor, and undesirably
reduces
process performance. Also, for known catalytic ebullated bed type reactors
which utilize internal gas/liquid separation devices, the volume of catalyst
in a
particular size reactor is undesirably limited.
Many prior art patents have been directed to various improvements in catalytic
hydrogenation processes for heavy hydrocarbon feedstocks utilizing catalytic
ebullated bed reactors, and have disclosed various operational parameters for
such reactors. For example, U.S. Patent No. 3,183,180 to Schuman et al., U.S.
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4,217,206 and U.S. 4,427,535 to Nongbri et al disclose hydrogenation of
petroleum residua using catalytic single stage ebullated bed reactors having
internal gas/liquid separation, and U.S. 4,576,710 to Nongbri et al and U.S.
4,853,111 to MacArthur et al disclose use of such catalytic two-stage
reactors.
Other prior art patents have disclosed hydrogenation process improvements
utilizing catalyses having various compositions and pore structures, and
specific
reaction conditions based on characteristics of the feedstocks. However, a
need
still remains for providing a comprehensive improved catalytic multi-stage
ebullated bed reactor system which is capable of producing improved
hydrogenation process performance efficiencies.
SUMMARY OF INVENTION
This invention provides an improved catalytic multi-stage hydrogenation
process
for treating heavy hydrocarbonaceous feedstocks and producing desired lower
boiling hydrocarbon liquid products with enhanced process performance. For
this improved hydrogenation process, we have discovered that a more efficient
catalytic multi-stage ebullated bed reactor system having improved performance
results can be achieved by maximizing the catalyst loading and also providing
increased reactor liquid residence time in each reactor, by utilizing reduced
catalyst space velocity and reduced superficial gas velocity which are
maintained within desired critical ranges in each reactor. These process
improvements result in desirably increasing the liquid hold-up volume percent
and reducing excessive gas hold-up volume percent in each of the reactors.
These desirable reaction results are accomplished by providing such increased
volume percent of particulate catalyst and lower catalyst space velocities in
each
reactor by utilizing an external gas/liquid separator, in combination with
utilizing
lower superficial upward gas velocities and reduced gas hold-up in each
reactor,
while providing a desired outlet hydrogen partial pressure and desired level
of
hydrogenation or hydroconversion as selected for any particular feedstock.
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For this invention, the catalytic ebullated bed reactor construction
arrangement
for the first stage reactor does not include an internal gas/liquid separation
device, but instead utilizes an efficient external gas/liquid separator.
Utilizing
such external gas/liquid separation results in an increased volume of
particulate
catalyst being provided in a particular size reactor and reduces the catalyst
space velocity, which is defined as the volumetric rate of feedstock processed
per unit weight of fresh catalyst in the reactor. For such commercial size
reactors having outside diameter of 12-14 ft. and a height of 50-60 ft., a
vertical
distance of 5-10 ft. should be maintained between the ebullated bed maximum
expansion level and the reactor outlet conduit, so as to avoid any carryover
of
catalyst from the reactor. Also, operating conditions for each of the two-
staged
catalytic ebullated bed reactors are selected so that the upward superficial
gas
velocity is maintained within a desired critical range, and the gas hold-up
volume
percentage in each reactor is beneficially reduced, which consequently permits
more reactor liquid to be in contact with the catalyst bed, so that the
reactor
performance as well as the overall process performance results are enhanced.
This invention is useful for processing heavy hydrocarbonaceous feedstocks
and providing overall hydroconversions in the range of 50-100 vol.% to produce
desired lower boiling hydrocarbon liquid products.
Therefore, to summarise, the present invention concerns a process for
catalytic
multi-stage ebullated bed hydrogenation of heavy hydrocarbonaceous
feedstocks for producing lower boiling hydrocarbon liquids and gases, the
process comprising:
(a) feeding a heavy hydrocarbonaceous liquid feedstock together with
hydrogen gas into a first stage catalytic ebullated bed reactor at liquid
space velocity of 0.2 - 2.0 volume of feed per hour per volume of reactor
(Vf/hrNr,) and at catalyst space velocity of 0.03 - 0.33 bbl/day/lb catalyst,
providing upward superficial gas velocity of 0.02-0.30 ft/sec while
maintaining reaction temperatures of 700-850 F, and 800-3,000 psi
hydrogen partial pressure at the reactor outlet, and producing a first stage
reactor effluent material;
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(b) phase separating the first stage effluent material into a gas portion and
a
first liquid portion, and passing the first liquid portion to a second stage
catalytic ebullated bed reactor maintained at the reaction conditions
similar to those of step (a), and producing a second stage reactor effluent
material;
(c) phase separating the second stage effluent material into a gas and a
second liquid portion;
(d) fractionating said second liquid portion to produce a medium-boiling
hydrocarbon liquid fraction product having normal boiling range of 400-
650 F and a vacuum bottoms fraction material having a normal boiling
point above about 650 F; and
(e) recycling said vacuum bottoms fraction material directly to said first
stage
catalytic ebullated bed reactor to provide a recycle volume ratio of the
vacuum bottoms material to fresh feedstock of 0-1.0/1, so as to achieve
between 50 and 100 vol.% conversion of the 975 F+ fraction in the
feedstock to lower-boiling hydrocarbon liquid and producing increased
yields of said medium-boiling hydrocarbon liquid product.
The invention also concerns a process for catalytic multi-stage ebullated bed
hydrogenation of heavy hydrocarbonaceous feedstocks and for producing lower
boiling hydrocarbon liquids and gases, the process comprising:
(a) feeding a heavy hydrocarbonaceous liquid feedstock containing 50-90
vol.% 975 F+ residua together with hydrogen gas into a first stage
catalytic ebullated bed reactor at liquid space velocity of 0.2-2.0 volume of
feed per hour per volume of reactor (Vf/hrNr,) and at catalyst space
velocity of 0.03-0.33 bbl/day/Ib catalyst, providing upward superficial gas
velocity of 0.02-0.25 ft/sec while maintaining reaction temperatures of
700-850 F, and 800-3,000 psi hydrogen partial pressure at the reactor
outlet, and producing a first stage reactor effluent material containing gas
and liquid portions;
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(b) phase separating the first stage effluent material into a gas portion and
a
first liquid portion, and passing the first liquid portion on to a second
stage
catalytic ebullated bed reactor maintained at the reaction conditions of
step (a), and producing a second stage reactor effluent material;
(c) phase separating the second stage effluent material into gas and a
second liquid portion, and withdrawing the second liquid portion;
(d) fractionating said second liquid portion to produce a medium boiling
hydrocarbon liquid fraction product having normal boiling range of 400-
650 F, a vacuum gas oil having a normal boiling range of 650-950 F, and
a vacuum bottoms material having a normal boiling temperature above
about 950 F; and
(e) recycling said vacuum bottoms fraction material directly back to said
first
stage catalytic ebullated bed reactor to provide a recycle volume ratio of
the vacuum bottoms material to fresh feedstock of 0-0.7/1, so as to
achieve 50-100 vol.% conversion of the 975 F+ fraction in the feedstock
to lower boiling hydrocarbon liquid and producing increased yields of said
medium-boiling hydrocarbon liquid product containing low sulfur and
nitrogen.
The broad and preferred characteristics for the hydrocarbonaceous feedstocks
and the reactor broad and preferred operating condition ranges for which this
invention is useful are provided in Table 1 below:
TABLE 1
FEEDSTOCK AND REACTOR OPERATING CONDITIONS
Condition Broad Preferred
Feedstock Residua Content, vol.% 975 F+ 30 - 100 50 - 90
Feedstock CCR*, wt.% 1- 50 10 - 40
Feedstock Nickel plus Vanadium, Wppm Up to 1,000 100 - 800
Reactor LHSV**, hr 1(per Reactor Stage) 0.2 - 2.0 0.4 - 1.2
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Reactor Temperature, F 700 - 850 750 - 840
Reactor Total Pressure, Psig 1,000 - 4,000 1,500 - 3,000
Reactor Outlet Hydrogen Partial Pressure, Psi 800 - 3,000 1,000 - 2,500
Reactor Superficial Gas Velocity, fps 0.02 - 0.30 0.025 - 0.20
Catalyst Space Velocity, BPD/Lb (per Stage) 0.03 - 0.33 0.04 - 0.20
Catalyst Replacement Rate, Lb/Bbl (per Stage)0.05 - 0.5 0.1 - 0.4
Catalyst Bed Expansion, % 25 - 75 35 - 50
Vacuum Bottoms Recycle Rate, Vr/Vfeed 0-1 0.2 - 0.7
Cutpoint of Vacuum Bottoms Recycle, F 650+ 900+ ,
* CCR = Conradson carbon residue.
** LHSV = Liquid hourly space velocity in each reactor, as defined as
volumetric
fresh feed rate divided by reactor total volume.
In the process, the fresh feedstock together with hydrogen are introduced into
a
first stage catalytic ebuiiated bed reactor, which does not contain an
internal
gas/liquid phase separator device. The catalyst bed is expanded by 25-75
percent above its settled level by the upflowing liquid and gas streams, and
is
maintained within the broad operating conditions of 700-850 F temperature,
800-3,000 psig hydrogen partial pressure at the reactor outlet, liquid hourly
space velocity of 0.20-2.0 volume fresh feed per hour per volume of reactor
(Vf/hrNr) and at catalyst space velocity of 0.03 - 0.33 barrel feed per day
per
pound fresh catalyst in the reactor. Because of the lower catalyst space
velocity
and superficial gas velocity being utilized in the reactor, the reacting
liquid
votume percentage is increased and gas hold-up volume is desirably reduced.
The first stage reactor usually hydroconverts 30-95 vol.% of the fresh heavy
feedstock and any recycled residua material to a lower boiling hydrocarbon
effluent material.
The first stage reactor effluent material is phase separated in an external
gas/liquid separator, a gas fraction is removed, and a sufficient portion of
the
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remaining liquid is recycled to the reactor to maintain the desired 25-75%
catalyst bed expansion therein. The remaining liquid fraction is passed
together
with additional hydrogen to a second stage catalytic ebullated bed type
reactor.
The second stage ebullated bed reactor is operated similarly to the first
stage
reactor and typically is maintained at 0-50 F, lower temperature in the broad
range of 700-850 F (370-455 C) and 0.20-2.0 (Vf/hrNr) space velocity, so as to
effectively further hydragenate the remaining unconverted residual material
therein. The second stage reactor usually further hydroconverts 30-95 vol.% of
the remaining residua feed material to lower boiling hydrocarbon materials.
From the second stage reactor, the effluent material is passed to various
gas/liquid separation and distillation steps, from which gases and low-boiling
hydrocarbon liquid product and distillation vacuum bottoms fraction materials
are
removed. If desired for achieving higher percentage conversion of the
feedstock,
a portion of the vacuum bottoms fraction material boiling above at least 650 F
(343 C) temperature and preferably boiling above about 900 F(482 C) can be
recycled back to the first stage catalytic reactor inlet at a recycle volume
ratio to
the fresh feedstock of 0-1.0/1, and preferably at 0.2-0.7/1 recycle ratio for
further
hydroconversion reactions therein.
Particulate catalyst materials which are useful in this hydrogenation process
may contain 2-25 wt.% percent total active metals selected from the metals
group consisting of cadmium, chromium, cobalt, iron, molybdenum, nickel, tin,
tungsten, and mixtures thereof deposited on a support material selected from
the group consisting of alumina, silica and combinations thereof. Also,
catalysts
having the same characteristics may be used in both the first stage and second
stage reactors, or each reactor may use catalysts having different
characteristics. Useful particulate catalysts will be in the form of beads,
extrudates or spheres and have broad and preferred characteristics as shown in
Table 2 below:
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TABLE 2
USEFUL CATALYST CHARACTERISTICS
Catalyst Characteristic Broad Preferred
Particle Diameter, in. 0.025-0.083 0.030-0.065
Particle Diameter, mm 0.63-2.1 0.75-1.65
Bulk Density, lb/ft3 25-50 30-45
Particle Crush Strength, lb/mm 1.8 min. 2.0 min.
Total Active Metals Content, wt.% 2-25 5-20
Total Pore Volume, cm2/gm* 0.30-1.50 0.40-1.20
Total Surface Area, m2/gm 100-400 150-350
Average Pore Diameter, Angstrom** 50-350 80-250
* Determined by mercury penetration method at 60,000 psi pressure.
** Average pore diameter calculated by ADP = 4 Pore Volume x 104
Surface Area
Catalysts having unimodal, bimodal and trimodal pore size distributions are
useful in this process. Preferred catalyses should contain 5-20 wt.% total
active
metals consisting of combinations of cobalt, molybdenum and nickel deposited
on an alumina support material.
This improved process for catalytic multi-stage hydrogenation of heavy
hydrocarbonaceous feedstocks advantageously provides enhanced
performance results by utilizing increased catalyst loading and liquid volume
percent together with reduced gas hold-up in each of the multiple staged
reactors with external gas/liquid separation. Such enhanced performance
efficiency is manifested principally by providing better utilisation of the
reactor
volume for any particular desired hydroconversion result. This process is
generally useful for catalytic hydrogenation and hydroconversion of heavy
petroleum crudes, topped crudes, and vacuum residua, bitumen from tar sands,
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for coal hydrogenation and liquefaction, and for catalytic co-processing
coal/oil
blends to produce lower boiling, higher value hydrocarbon liquid products.
BRIEF DESCRIPTION OF DRAWINGS
This invention will be described further with the aid of the following
drawings, in
which:
Fig. 1 is a schematic flow diagram of an improved catalytic two-stage
hydrogenation process for heavy hydrocarbonaceous feedstocks for producing
desired lower-boiling liquid and gas products according to the invention;
Fig. 2 is a graph generally showing the typical general relationship
between catalyst space velocity for a catalytic ebullated bed reactor and
feedstock hydrodesulfurization results for the reactor;
Fig. 3 is a graph of experimental data generally showing the relationship
between superficial gas velocity in a catalytic ebullated bed reactor and gas
hold-up volume percentage in the reactor for various superficial liquid
velocities;
and
Fig. 4 is a graph generally showing the effect of reactor gas hold-up
volume percent on hydrodesulfurization results particularly in a second stage
catalytic reactor.
DESCRIPTION OF INVENTION
The present invention is now described in more detail for a hydrogenation
process utilizing an improved catalytic two-stage ebullated bed reaction
system
for treating heavy hydrocarbon feedstocks. For the process as shown by Figure
1, a pressurized heavy hydrocarbon feedstock such as petroleum vacuum
residua containing 30-100 vol.% 975 F+ residua and preferably 50-90 vol.% is
provided at 10 and combined with hydrogen at 12. A heavy vacuum bottoms
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recycle liquid can be added at 13, and the combined stream at 14 is
pressurized
and fed through flow distributor la upwardly into first stage catalytic
ebullated
bed reactor 16 containing ebullated bed 18. The total feedstock to reactor 16
consists of the fresh hydrocarbon feed material at 10 plus any recycled vacuum
bottoms material at 13. The recycle volume ratio of the vacuum bottoms
material
to the fresh oil feedstock is in the range of 0-1.0/1, and preferably is 0.2-
0.7/1
recycle ratio, with the higher recycle ratios being used for achieving higher
overall percentage conversion of the feedstock residua.
The first stage reactor 16 contains an ebullated bed 18 of particulate
supported
type catalyst having the form of beads, extrudates, spheres, etc., and is
maintained within the range of broad and preferred operating conditions as
shown in Table 1 above. The physical level of catalyst at 18a in the reactor
is
higher than for typical ebullated-bed reactors. This is because the usual
internal
recycle cup device which occupies a significant portion of reactor height, is
not
provided for separating the reactor liquid and vapor portions within the
reactor
16. Instead, an external or interstage phase separator 20 is provided between
the first and second stage catalytic reactors to effectively separate the
reactor
liquid and vapor effluent portions. Removal of the usual internal recycle cup
separator results in more catalyst and a higher level for the expanded
catalyst
bed in the reactor and desirably provides for a lower catalyst space velocity,
which contributes to the higher levels of performance for the reactors. A
vertical
height distance "h" of 5 -10 ft. is maintained between the maximum bed
expansion lever and the inlet of reactor outlet conduit 19 to prevent
carryover of
catalyst particles from the expanded bed 18.
From first stage reactor 16, overhead effluent stream 19 is withdrawn and
passed to the external phase separator 20. From separator 20, a vapor stream
21 is removed and passed to gas purification section 42. Also, a liquid stream
22
is withdrawn, and a sufficient flow is recirculated through conduit 24 by
ebuilating pump 25 back to the reactor 16 to expand the catalyst bed 18 by the
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desired 25-75 percent above its normal settled bed height. For the first stage
reactor 16, particulate catalyst material is added at connection 17 at the
desired
replacement rate, and can be used catalyst withdrawn from second stage
reactor 30 at connection 36, and usually treated at unit 38 as desired to
remove
undesired particulate fines, etc. at 37. Fresh make-up catalyst can be added
to
catalyst bed 18 as needed at connection 17a, and an equivalent amount of
spent catalyst is withdrawn from catalyst bed 18 at connection 17b.
The typical general relationship between reactor catalyst space velocity and
reactor performance results is illustrated in Figure 2, which shows the effect
of
lower catalyst space velocities on hydrodesulfurization performance for
ebullated-bed reactors having equal total volumes, hydrocarbon feedrates,
reaction temperatures and catalyst replacement rates. Figure 2 clearly shows
the improvement in first stage reactor desulfurization performance provided by
lower catalyst space velocities, resulting mainly from use of an external
gas/liquid separation device instead of the usual internal separation device
and
for nominal residue conversion levels between about 65 and 90 vol.%. The
hydrocarbon liquid feedstock and hydrogen both react in contact with the
catalyst in the reactor ebullated bed to form lower boiling components which
have lower contaminant levels than the feedstock.
The hydrogen gas provided at 12 to the first stage reactor 16 is mainly
recycled
unreacted hydrogen having purity in the range of 85-95 vol. percent and some
essentially pure make-up hydrogen as needed. For this improved process, the
hydrogen feed rate to the first stage reactor and to the subsequent staged
reactors is established at a minimum required level, which provides at each
reactor outlet a required hydrogen partial pressure which is determined based
on characteristics for a particular feedstock, the catalyst characteristics,
the
desired level of reaction severity, and the product quality objectives.
Typically,
the required hydrogen feed rate to a catalytic reactor is expressed as a
multiple
of the quantity of hydrogen chemically consumed in the reactor, and such
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hydrogen rate is usually in the range of 2.0 to 5.0 times the chemical
hydrogen
consumption therein. Minimizing hydrogen gas feed rate in the catalytic
ebullated-bed reactor(s) results in lower gas hold-up of hydrogen and
hydrocarbon vapor evolved therein, and provides longer liquid residence time
and enhanced liquid phase kinetics at the catalyst surface. The longer reactor
liquid residence time is explained by the following relationship:
Reactor Liquid Residence Time = Volume of Reactor Occupied by Liguid *
Liquid Hourly Space Velocity **
10 * Reactor volume occupied by Liquid = Volume total - Volume occupied by Gas
-
Volume occupied by solid (catalyst)
** Volumetric rate of fresh liquid feed divided by reactor total volume
The volume percent of hydrogen gas hold-up in the catalytic ebullated-bed
reactor including hydrocarbon vapors generated therein, is primarily related
to
the reactor superficial gas velocity, with increased upward superficial gas
velocity resulting in an increased gas hold-up volume percentage in the
reactor.
Experimental data showing this relationship between the upward superficial gas
velocity and gas hold-up volume percent in catalytic ebullated-bed reactors is
shown in Figure 3. The measured gas hold-up volume percent in the reactor is
shown as a function of the reactor superficial gas velocity at three different
levels of reactor liquid upward superficial velocity. The superficial gas
velocity for
up flowing hydrogen gas clearly has the primary effect on gas hold-up volume
in
the reactor, with a secondary effect being due to different superficial liquid
upward velocities for the feed liquid in the reactor
Regarding the need for providing a sufficient quantity of reactant hydrogen
gas
in the reactor for desired chemical consumption therein, recent laboratory
studies at gas hold-up percentages less than about 5-10 vol.% have clearly
shown that this is a sufficient hydrogen quantity. Gas hold-up in excess of
about
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-5 vol.% has usually been a consequence of scale-up of small size experimental
catalytic ebullated-bed reactors to commercial size reactors (i.e., for taller
reactors having lower length/diameter ratios than for slender laboratory scale
reactors), and result in a less efficient reaction system because the liquid
residence times and gas hold-up volumes are usually adversely affected. The
present invention advantageously minimises this excessive hydrogen gas and
hydrocarbon vapor hold-up volume percentage in the reactor, so as to provide
the enhanced reaction kinetics and higher overall levels of process
performance
for the reactor system.
This relationship of catalytic reactor performance such as percent
hydroconversion, hydrodesulfurization, etc. of the heavy hydrocarbon feedstock
to the percentage of gas hold-up in an ebullated bed reactor is further
illustrated
in Figure 4. This comparison was made for catalytic ebullated bed reactors
having equal total volumes, hydrocarbon feedrates, reaction temperatures and
catalyst replacement rates. The results indicate that for reduced gas hold-up
in a
second stage reactor, the hydrodesulfurization results are significantly
increased
for various overall hydroconversion levels of 65 vol.% and 90 vol.% for the
feedstock.
As mentioned above, the first stage reactor effluent stream 19 is passed to
the
interstage separator 20, which has two main functions: (a) to provide an
ebullating recycle liquid stream back to the first stage reactor with minimal
gas
entrainment, and (b) to provide a liquid feed stream to the second stage
reactor
having a minimal vapor content. The effect of the function (b) is reduced gas
hold-up in the second stage reactor and the same reaction benefits as
described
for the first stage reactor. The liquid feed to the second stage reactor 30
contains the unconverted residue from the original feedstock, and
hydroconversion fractions which normally boil above about 600 F (316 C).
Recycled hydrogen, together with fresh make-up hydrogen at 45 is added as
stream 32 to the second stage reactor 30, the hydrogen gas rate being selected
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stream 32 to the second stage reactor 30, the hydrogen gas rate being selected
so as to result in a minimal hydrogen partial pressure at the reactor 30
outlet as
needed to meet processing and product objectives as described above. Relative
to typical hydrogen gas rates previously used, the gas rate provided at 32 to
the
second stage reactor 30 for this invention is substantially lower. This
results in
lower gas hold-up volume percentages in the reactor, greater liquid residence
time, and a more efficient reactor system. In this situation, the gas hold-up
is
reduced from about 27 to 12 vol. percent, which results in an improvement in
second stage desulfurization results from 65 to 70 wt.% based on the fresh
feedstock.
Also from the external phase separator 20, a liquid portion 26 from the liquid
stream 22 provides liquid feed material upwardly through flow distributor 27
into
ebullated bed 28 of the second stage catalytic ebullated bed reactor 30. The
catalyst bed 28 is expanded by 25-75% above its settled height by the
upflowing
gas and liquid therein. Reactor liquid is withdrawn from an internal phase
separator 33 through conduit 34 to recycle pump 35, and is reintroduced
upwardly through the flow distributor 27 into the ebullated bed 28 to maintain
the
desired catalyst bed expansion therein.
The second stage catalytic reactor 30 with ebullated catalyst bed 28 is
operated
within the broad and preferred conditions as shown in Table 1 above, and
maximises resid hydrogenation reactions which occur therein. The second stage
reaction temperature is preferably 0-50 F lower than that of the first stage
reactor. Recycle and fresh hydrogen is provided at 32 to the second stage
reactor 30, so that a minimal but adequate level of hydrogen partial pressure
of
1,000-2,500 psi is maintained at the reactor 30 outlet.
The catalyst particles in ebullated beds 18 and 28 have a relatively narrow
size
range for uniform bed expansion under controlled upward liquid and gas flow
conditions. While the useful catalyst size range is between 0.025 and 0.083
inch
effective diameter, including beads, extrudates, or spheres, the catalyst size
is
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reactor, the density of the catalyst particles, and the lifting effect of the
upflowing
liquid and hydrogen gas are important factors in providing the desired 25-75
percent expansion and operation of the catalyst beds. If desired, used
particulate catalyst may be withdrawn from the second stage reactor bed 28 at
connection 36 and fresh catalyst is added at connection 36a as needed to
maintain the desired catalyst volume and catalytic activity therein. This used
catalyst withdrawn at 36, which has relatively low metal contaminant
concentration, can be passed to a treatment unit 38 where it is washed and
screened to remove undesired fines at 37, and the recovered catalyst at 39 can
provide used catalyst addition at 17 to the first stage reactor bed 18,
together
with any fresh make-up catalyst added at connection 17a as needed. From the
second stage reactor 30, an effluent stream is removed at 31 and passed to a
phase separator 40. From separator 40, a hydrogen-containing gas stream 41 is
passed to the gas purification section 42 for removal of contaminants such as
C02, H2S, and NH3 at vent 43. Purified hydrogen at 44 is recycled back to each
catalytic reactor 16 and 30 as desired as the hydrogen streams 12 and 32
respectively, while fresh hydrogen is added at 45 as needed.
Also from the separator 40, a liquid fraction 46 is withdrawn, pressure
reduced
at 47 to 0-100 psig, and is introduced into fractionation tower unit 48. A
gaseous
product stream is removed at 49 and a light hydrocarbon liquid product
normally
boiling between 400-650 F is withdrawn at 50. A bottoms nominal 650 F+
fraction is withdrawn at 52, reheated at heater 53, and passed to vacuum
distillation step at 54. A vacuum gas oil liquid product is removed overhead
at
55. Vacuum bottoms stream 56, which has been hydrogenated in the second
stage catalyst reactor 30, can be recycled back as stream 13 to the first
stage
catalytic reactor 16. The recycle volume ratio for vacuum bottoms stream 56 to
fresh feed at 10 can be 0-1.0/1, and preferably should be 0.2-0.7/1 for
achieving
hydroconversion of the feedstock exceeding about 70 vol. percent. It is
pointed
out that by utilizing this two stage catalytic hydroconversion process, the
thermal
reactions and catalytic activity in each stage reactor can be effectively
matched
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reactions and catalytic activity in each stage reactor can be effectively
matched
and enhanced. The remaining unconverted vacuum bottoms material not being
recycled at 13 is withdrawn at 57 as a net product.
This invention will now be described further by use of the following example,
which is intended to be illustrative only and should not be construed as
limiting
the scope of the invention.
Example
To demonstrate the process advantages of this invention, analyses of four
commercial ebullated-bed reactor cases have been developed and are
presented below. The basis for these comparative cases is the catalytic two-
stage ebullated bed reactor processing of a typical Arabian light/heavy vacuum
residual feedstock and providing 65 and 90 vol.% hydroconversion of the
1050 F+ vacuum residua fraction and with a high percentage level of
desulfurization. The vacuum residual feedstock has inspection analyses as
shown in Table 3 below.
TABLE 3
FEEDSTOCK ANALYSES
Characteristic Value
Residue Content (1050 F+), vol.% 92
Gravity, API 4.7
Sulfur, wt.% 5.3
Conradson Carbon Residue, wt.% 24.6
Nickel plus Vanadium, Wppm 222
Two conventional process base cases No. 1 and 3 which do not incorporate
features of the present invention and two improvement cases No. 2 and 4 which
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the process performance advantages of the invention. The cases No. 1 and 2
comparisons are both for a moderate 65 vol.% overall hydroconversion of the
1050 F+ vacuum residua fraction, and the cases No. 3 and 4 comparisons are
both for a high 90 vol.% overall hydroconversion of the residua fraction to
lower
boiling hydrocarbon products. These examples are based on actual laboratory
and commercial data at either identical or similar reaction and operating
conditions, including the feedstock and catalyst characteristics. The
operating
conditions for the four comparison cases are provided in Table 4 below.
TABLE 4
10 REACTOR OPERATING CONDITIONS
Case No. 1 2 3 4
Desired Overall Conversion, V% 65 65 90 90
First Stage Reactor
LHSV, V/hrN 0.60 0.60 0.60 0.60
Reactor Temperature, F 814 814 844 844
Catalyst SV, BPD/Lb 0.106 0.085 0.127 0.108
Catalyst Replacement Rate, Lb/Bbl 0.123 0.123 0.175 0.175
Superficial Gas Velocity, Ft/Sec 0.114 0.105 0.107 0.096
Reactor Gas Hold-Up, V% 20.1 18.5 18.9 16.9
Hydrogen
Chemical Consumption, SCF/Bbl 758 869 1147 1114
Inlet Circulation Rate, X Consumption 3.6 3.0 3.1 2.7
Inlet Purity, vol.% 94.9 92.0 96.1 92.0
Partial Pressure Inlet, Psia 2505 2505 2505 2505
Partial Pressure Outlet, Psia 2415 2046 2166 1967
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16
Second Stage Reactor
LHSV, V/hrN 0.60 0.60 0.60 0.60
Reactor Temperature, F 814 814 844 844
Catalyst SV, BPD/Lb 0.106 0.106 0.127 0.127
Catalyst Replacement Rate, Lb/BbI 0.123 0.123 0.175 0.175
Superficial Gas Velocity, Ft/Sec 0.153 0.069 0.144 0.079
Reactor Gas Hold-Up. vol.% 26.9 12.1 25.4 13.9
Hydrogen
Chemical Consumption, SCF/Bbl 512 481 733 806
Inlet Circulation Rate, X Consumption 2.1 3.0 2.0 2.7
Inlet Purity, vol.% 85.3 92.0 92.0 92.0
Partial Pressure Inlet, Psia 2206 2355 2180 2480
Partial Pressure Outlet, Psia 1964 1963 1949 1958
For the base cases No. 1 and 3, the catalytic ebullated-bed two-stage reactors
are operated at typical pre-invention conditions including a high feed rate of
hydrogen entering the first stage reactor, the upward ebullation liquid flow
being
provided from an internally located recycle cup or gas/liquid separator, and
with
cell of the first stage reactor effluent material (vapor + liquid) being
passed
directly to the catalytic second stage reactor. The superficial gas velocities
in the
first and second staged reactors are about 0.11 and 0.15 ft/s respectively,
and
result in undesirably large gas hold-up volumes of 18-20 vol.% and 25-27 vol.
%
respectively in the first and second staged reactors. For the two improvement
Cases No. 2 and 4, the improved results for the present invention utilizing
the
same reactor total volume and liquid hourly space velocity as for the
respective
base Cases No. 1 and 3 are demonstrated. For the first stage reactor, the
catalyst volume is increased and the catalyst space velocity is decreased by
15-
20 percent due to elimination of the internal recycle cup or gas/liquid
separator
from the reactor upper portion. The first stage gas hold-up volume is reduced
by
CA 02412923 2006-06-29
17
8-11 percent primarily because a lower hydrogen gas circulation rate and a
lower hydrogen partial pressure at the reactor outlet are utilized.
More significantly, in the second stage reactor the gas hold-up volume is
reduced by 45-55 percent. This reduction in second stage gas hold-up volume
percentage is due to the use of interstage gas/liquid separation, and the use
of a
reduced minimal hydrogen gas recirculation rate. This reduction in the second
stage reactor gas hold-up volume becomes available for providing increased
reactor liquid volume and increases the effective liquid residence time in the
second stage reactor by 20-30 percent. The comparative process performance
for hydroconversion and desulfurization for the Cases No. 1 and 2. and for
Cases No. 3 and 4 are shown in Table 5 below.
TABLE 5
PROCESS COMPARATIVE PERFORMANCES
Difference Difference
Case No. -1 -2 1-2 -3 -4 3--4
First Stage Reactor
1050 F+ Conversion, vol.% 45.8 44.5 -1.3 73.0 72.5 -0.5
Desulfurization, M.% 70.0 70.4 +0.4 63.2 65.0 +1.8
Second Stage Reactor
1050 F+ Conversion, vol.% 35.4 41.4 +6.0 63.0 68.7 +5.7
Desulfurization, wt.% 65.1 69.9 +4.8 53.3 58.9 +5.6
Overail Results
1050 F+ Conversion, vol.% 65.0 67.5 +2.5 90.0 91.4 +1.4
Desulfurization, wt.% 89.5 91.1 +1.6 82.8 85.6 +2.8
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It is noted that the level of first stage reactor residue conversion for the
comparative cases shows a slight decrease due to higher solids hold-up.
However, first stage desulfurization is increased slightly due to the higher
catalyst loading and lower gas hold-up volume percentage in the reactor. Also,
as a primary result of the significantly lower gas hold-up in the second stage
reactor, i.e. from 26.9 to 12.1 vol.%, at moderate 65% conversion and from
25.4
to 13.9 vol.% at the hiaher 90% conversion, the process overall percent
hydroconversion is increased from 65 to 67.5 vol.% for the moderate 65 vol.%
conversion cases, and from 90.0 to 91.4 vol.% for the high 90 vol.% conversion
case. The increase in overall desulfurization from 89.5 to 91.1 wt.% in the
moderate conversion cases and from 82.8 to 85.6 md.% in the high conversion
cases is a direct result of the increase in the second stage desulfurization.
It
should be noted that the moderate 65 vol.% conversion cases utilized a
particulate catalyst having a unimodal pore size distribution, and the high
conversion cases utilized a catalyst having a bi-modal pore size distribution
which results in a somewhat lower desulfurization level.
Although this invention has been described broadly and also in terms of
preferred embodiments, it will be understood that modifications and variations
can be made to the process which are all within the basic scope of the
invention
as defined by the following claims.
