U.S. patent application number 09/248740 was filed with the patent office on 2002-01-31 for multistage moving-bed hydroprocessing reactor with separate catalyst addition and withdrawal systems for each stage, and method for hydroprocessing a hydrocarbon feed stream.
Invention is credited to SCHEUERMAN, GEORGIEANNA L..
Application Number | 20020011428 09/248740 |
Document ID | / |
Family ID | 46253030 |
Filed Date | 2002-01-31 |
United States Patent
Application |
20020011428 |
Kind Code |
A1 |
SCHEUERMAN, GEORGIEANNA L. |
January 31, 2002 |
MULTISTAGE MOVING-BED HYDROPROCESSING REACTOR WITH SEPARATE
CATALYST ADDITION AND WITHDRAWAL SYSTEMS FOR EACH STAGE, AND METHOD
FOR HYDROPROCESSING A HYDROCARBON FEED STREAM
Abstract
A method, and a reactor, for hydroprocessing a hydrocarbon feed
stream through multistage moving catalyst beds contained within a
single onstream reactor vessel, with separate catalyst addition and
withdrawal systems for each of the multistages of moving catalyst
beds. The reactor contains two or more different and distinct
moving catalyst beds for any hydroprocessing application. The
method includes serially passing, without leaving the reactor
vessel, at least a partially treated hydrocarbon stream from one
hydroconversion reaction zone containing a moving catalyst bed with
a first set of catalytic characteristics to another hydroconversion
reaction zone containing a moving catalyst bed with a second set of
catalytic characteristics that differ in catalytic abilities from
the first set of catalytic characteristics.
Inventors: |
SCHEUERMAN, GEORGIEANNA L.;
(MORAGA, CA) |
Correspondence
Address: |
CHEVRON CORPORATION LAW DEPARTMENT
PATENT DIVISION
PO BOX 6006
SAN RAMON
CA
94583-0806
US
|
Family ID: |
46253030 |
Appl. No.: |
09/248740 |
Filed: |
February 10, 1999 |
Related U.S. Patent Documents
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Application
Number |
Filing Date |
Patent Number |
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09248740 |
Feb 10, 1999 |
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08639054 |
Apr 24, 1996 |
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5916529 |
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08639054 |
Apr 24, 1996 |
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08497638 |
Jun 30, 1995 |
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5599440 |
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08497638 |
Jun 30, 1995 |
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08235043 |
Apr 29, 1994 |
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08235043 |
Apr 29, 1994 |
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08215254 |
Mar 21, 1994 |
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5409598 |
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08215254 |
Mar 21, 1994 |
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08014847 |
Feb 8, 1993 |
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5302357 |
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08014847 |
Feb 8, 1993 |
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07727656 |
Jul 9, 1991 |
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07727656 |
Jul 9, 1991 |
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07381948 |
Jul 19, 1989 |
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5076908 |
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Current U.S.
Class: |
208/89 ; 208/147;
208/149; 208/150; 208/163 |
Current CPC
Class: |
C10G 49/14 20130101;
B01J 8/008 20130101; B01J 8/22 20130101; B01J 8/1881 20130101; B01J
8/003 20130101; B01J 2208/00884 20130101; C10G 49/002 20130101;
B01J 8/125 20130101 |
Class at
Publication: |
208/89 ; 208/149;
208/150; 208/163; 208/147 |
International
Class: |
C10G 045/00; C10G
035/10 |
Claims
I claim:
1. A reactor comprising a reactor vessel having an internal
cylindrical wall with an internal reactor cylindrical wall
diameter; a first catalyst bed support means coupled to and
supported by said internal cylindrical wall of said reactor vessel
such as to essentially entirely span across said internal reactor
cylindrical wall diameter for supporting a first catalyst moving
bed in a first hydroconversion reaction zone having a first
physical and/or catalytic properties and for allowing a hydrocarbon
feed stream to pass therethrough while preventing catalyst
particles from a first catalyst moving bed to exit therethrough; a
first catalyst outlet means, communicating through said internal
cylindrical wall of said reactor vessel with a first catalyst
moving bed supported by said first catalyst bed support means, for
withdrawing a first generally spent catalyst from a first
hydroconversion reaction zone; a first catalyst inlet means,
communicating through said internal cylindrical wall of said
reactor vessel with a first catalyst moving bed supported by said
first catalyst bed support means, for introducing a first generally
fresh catalyst into a first hydroconversion reaction zone; a second
catalyst bed support means coupled to and supported by said
internal cylindrical wall of said reactor vessel such as to
essentially entirely span across said internal reactor cylindrical
wall diameter for supporting a second catalyst moving bed in a
second hydroconversion reaction zone having a second physical
and/or catalytic properties and for allowing a hydrocarbon feed
stream to pass therethrough while preventing catalyst particles
from a second catalyst moving bed to exit therethrough; a second
catalyst outlet means, communicating through said internal
cylindrical wall of said reactor vessel with a second catalyst
moving bed supported by said second catalyst bed support means, for
withdrawing a second generally spent catalyst from a second
hydroconversion reaction zone; and a second catalyst inlet means,
communicating through said internal cylindrical wall of said
reactor vessel with a second catalyst moving bed supported by said
second catalyst bed support means, for introducing a second
generally fresh catalyst into a second hydroconversion reaction
zone.
2. The reactor of claim 1 additionally comprising a first catalyst
moving bed supported by said first catalyst bed support means and
having a first physical and/or catalytic properties; and a second
catalyst moving bed supported by said second catalyst bed support
means and having a second physical and/or catalytic properties.
3. The reactor of claim 2 wherein said first physical and/or
catalytic properties differ from said second physical and/or
catalytic properties.
4. The reactor of claim 3 additionally comprising a first conduit
means, communicating through said internal cylindrical wall of said
reactor vessel, for introducing a hydrocarbon feed stream into the
reactor vessel; and a second conduit means, communicating through
said internal cylindrical wall of said reactor vessel, for passing
a hydrocarbon product out of said reactor vessel.
5. The reactor of claim 4 wherein said first catalyst bed support
means allows an untreated hydrocarbon feed stream from said first
conduit means to pass therethrough while preventing catalyst
particles from said first catalyst moving bed to exit therethrough;
and said second catalyst bed support means allows an initially
treated hydrocarbon feed stream from said first catalyst moving bed
to pass therethrough while preventing catalyst particles from said
second catalyst moving bed to exit therethrough.
6. The reactor of claim 3 wherein said second catalyst bed support
means receives an untreated hydrocarbon feed stream and allows an
initially treated hydrocarbon feed stream from said second catalyst
moving bed to pass therethrough while preventing catalyst particles
from said second catalyst moving bed to exit therethrough; and said
first catalyst bed support means allows a treated hydrocarbon feed
stream from said first catalyst moving bed to pass therethrough
while preventing catalyst particles from said first catalyst moving
bed to exit therethrough.
7. A method for hydroprocessing a hydrocarbon feed stream
comprising the steps of: (a) disposing a first catalyst bed in a
first hydroconversion reaction zone of a reactor vessel, wherein
said first catalyst bed contains a first set of catalytic
properties; (b) disposing a second catalyst bed in a second
hydroconversion reaction zone of said reactor vessel of step (a),
wherein said second catalyst bed contains a second set of catalytic
properties; (c) flowing through said first catalyst bed in said
first hydroconversion reaction zone a hydrocarbon feed stream to
produce an initially-treated hydrocarbon feed stream; and (d)
flowing through said second catalyst bed in said second
hydroconversion reaction zone said initially-treated hydrocarbon
feed stream to produce a treated hydrocarbon feed stream.
8. The method of claim 7 additionally comprising withdrawing at
least partially spent first catalyst from said first catalyst bed
simultaneously with said flowing step (c); and adding fresh first
catalyst to said first catalyst bed simultaneously with said
flowing step (c).
9. The method of claim 7 additionally comprising withdrawing at
least partially spent second catalyst from said second catalyst bed
simultaneously with said flowing step (d); and adding fresh second
catalyst to said second catalyst bed simultaneously with said
flowing step (d).
10. The method of claim 8 additionally comprising withdrawing at
least partially spent second catalyst from said second catalyst bed
simultaneously with said flowing step (d); and adding fresh second
catalyst to said second catalyst bed simultaneously with said
flowing step (d).
11. The method of claim 7 wherein said flowing step (d) comprises
flowing said initially-treated hydrocarbon feed stream through said
second catalyst bed without said initially-treated hydrocarbon feed
stream having left or exited said reactor vessel of said step
(a).
12. The method of claim 10 wherein said flowing step (d) comprises
flowing said initially-treated hydrocarbon feed stream through said
second catalyst bed without said initially-treated hydrocarbon feed
stream having left or exited said reactor vessel of said step
(a).
13. The method of claim 7 wherein said first set of catalytic
properties are different from said second set of catalytic
properties.
14. The method of claim 12 wherein said first set of catalytic
properties are different from said second set of catalytic
properties.
15. The method of claim 14 additionally comprising demetallizing
said hydrocarbon feed stream of said step (c) in said first
catalyst bed; and hydrotreating said initially-treated hydrocarbon
feed stream of said step (d) in said second catalyst bed to remove
at least one chemical element from said initially-treated
hydrocarbon feed stream.
16. The method of claim 13 additionally comprising demetallizing
said hydrocarbon feed stream of said step (c) in said first
catalyst bed; and hydrotreating said initially-treated hydrocarbon
feed stream of said step (d) in said second catalyst bed to remove
at least one chemical element from said initially-treated
hydrocarbon feed stream.
17. The method of claim 15 wherein said at least one chemical
element is selected from the group consisting of nitrogen, sulfur,
and mixtures thereof.
18. The method of claim 16 wherein said at least one chemical
element is selected from the group consisting of nitrogen, sulfur,
and mixtures thereof.
19. A reactor comprising a reactor vessel having an internal
cylindrical wall with an internal reactor cylindrical wall
diameter; a first distributor plate assembly coupled to and
supported by said internal cylindrical wall of said reactor vessel;
a first catalyst bed support assembly coupled to and supported by
said internal cylindrical wall of said reactor vessel such as to
define a first plenum chamber between said first distributor plate
assembly and said first catalyst bed support assembly and for
supporting a first catalyst moving bed in a first hydroconversion
reaction zone having a first set of physical and/or catalytic
properties and for allowing a hydrocarbon feed stream to pass
therethrough from said first plenum chamber while preventing
catalyst particles from a first catalyst moving bed to exit
therethrough; a first catalyst outlet means, communicating through
said internal cylindrical wall of said reactor vessel with a first
catalyst moving bed supported by said first catalyst bed support
assembly, for withdrawing a first generally spent catalyst from a
first hydroconversion reaction zone; a first catalyst inlet means,
communicating through said internal cylindrical wall of said
reactor vessel with a first catalyst moving bed supported by said
first catalyst bed support assembly, for introducing a first
generally fresh catalyst into a first hydroconversion reaction
zone; a second distributor plate assembly coupled to and supported
by said internal cylindrical wall of said reactor vessel; a second
catalyst bed support assembly coupled to and supported by said
internal cylindrical wall of said reactor vessel such as to define
a second plenum chamber between said second distributor plate
assembly and said second catalyst bed support assembly and for
supporting a second catalyst moving bed in a second hydroconversion
reaction zone having a second set of physical and/or catalytic
properties and for allowing a hydrocarbon feed stream to pass
therethrough from said second plenum chamber while preventing
catalyst particles from a second catalyst moving bed to exit
therethrough; a second catalyst outlet means, communicating through
said internal cylindrical wall of said reactor vessel with a second
catalyst moving bed supported by said second catalyst bed support
assembly, for withdrawing a second generally spent catalyst from a
second hydroconversion reaction zone; and a second catalyst inlet
means, communicating through said internal cylindrical wall of said
reactor vessel with a second catalyst bed supported by said second
catalyst bed support assembly, for introducing a second generally
fresh catalyst into a second hydroconversion reaction zone.
20. The reactor of claim 19 additionally comprising a first
catalyst moving bed supported by said first catalyst bed support
assembly and having a first set of physical and/or catalytic
properties; and a second catalyst moving bed supported by said
second catalyst bed support assembly and having a second physical
and/or catalytic properties.
21. The reactor of claim 20 wherein said first set of physical
and/or catalytic properties differ from said second set of physical
and/or catalytic properties.
22. The reactor of claim 21 additionally comprising a first conduit
means, communicating through said internal cylindrical wall of said
reactor vessel, for introducing a hydrocarbon feed stream into the
reactor vessel; and a second conduit means, communicating through
said internal cylindrical wall of said reactor vessel, for passing
a hydrocarbon product out of said reactor vessel.
23. The reactor of claim 22 wherein said first catalyst bed support
assembly allows an untreated hydrocarbon feed stream from said
first conduit means to pass therethrough while preventing catalyst
particles from said first catalyst moving bed to exit therethrough;
and said second catalyst bed support assembly allows an
initially-treated hydrocarbon feed stream from said first catalyst
moving bed to pass therethrough while preventing catalyst particles
from said second catalyst moving bed to exit therethrough.
24. The reactor of claim 21 wherein said second catalyst bed
support assembly receives an untreated hydrocarbon feed stream and
allows an initially-treated hydrocarbon feed stream from said
second catalyst moving bed to pass therethrough while preventing
catalyst particles from said second catalyst moving bed to exit
therethrough; and said first catalyst bed support assembly allows a
treated hydrocarbon feed stream from said first catalyst moving bed
to pass therethrough while preventing catalyst particles from said
first catalyst moving bed to exit therethrough.
Description
[0001] This application is a continuation-in-part of co-pending
application entitled CATALYST, METHOD AND APPARATUS FOR AN
ON-STREAM PARTICLE REPLACEMENT SYSTEM FOR COUNTERCURRENT CONTACT OF
A GAS AND LIQUID FEED STREAM WITH A PACKED BED, Ser. No.
08/497,638, filed Jun. 30, 1995; which application is a
continuation of application Ser. No. 08/235,043, filed Apr. 29,
1994; which application is a continuation in part of Ser. No.
08/215,254, filed Mar. 21, 1994, now U.S. Pat. No. 5,409,598, dated
Apr. 25, 1995; which application is a divisional of Ser. No.
08/014,847, filed Feb. 8, 1993. now U.S. Pat. No. 5,409,598, dated
Apr. 12, 1994; which application is a continuation application of
Ser. No. 07/727,656, filed Jul. 9, 1991; which application is a
divisional application of application Ser. No. 07/381,948, filed
Jul. 19, 1989, now U.S. Pat. No. 5,076,908 dated Dec. 31, 1991.
Benefit of the earliest filing date is claimed, especially with
respect to all common subject matter.
BACKGROUND OF THE INVENTION
[0002] 1. Field of the Invention
[0003] The present invention broadly relates to a multistage moving
bed hydroprocessing reactor with separate catalyst addition and
withdrawal systems for each stage.
[0004] More particularly, the present invention relates to, a
method of, and apparatus for hydroprocessing where two or more
different (i.e. different in physical and/or catalytic properties)
and distinct moving bed catalysts are employed in a single onstream
reactor for any hydroprocessing application, and where catalyst
particles are added and withdrawn from each of the two or more
different and distinct moving bed catalyst while the single
onstream reactor is hydroprocessing a hydrocarbon feed stream. The
present invention provides a method and a reactor which combines
the advantages of a moving-bed reactor with the advantages of
layering different catalyst within a single reactor, preferably and
optionally while maintaining continuous or intermittent replacement
of catalyst for plug-like flow of each of the catalyst bed through
the reactor. Such plug flow with high processing rates is obtained
by selecting the size, shape and density of the catalyst particles
to prevent ebullation at the design flow rate so as to maximize the
amount of catalyst in the vessel during normal operation and during
catalyst transfer. Catalysts are selected by measuring bed
expansion with hydrocarbon, hydrogen and catalyst at the design
pressures and flow velocities within the available reaction volume
of the reactor.
[0005] 2. Description of the Prior Art
[0006] Hydroprocessing or hydrotreatment to remove undesirable
components from hydrocarbon feed streams is a well known method of
catalytically treating such heavy hydrocarbons to increase their
commercial value. "Heavy" hydrocarbon liquid streams, and
particularly reduced crude oils, petroleum residua, tar sand
bitumen, shale oil or liquified coal or reclaimed oil, generally
contain product contaminants, such as sulfur, and/or nitrogen,
metals and organo-metallic compounds which tend to deactivate
catalyst particles during contact by the feed stream and hydrogen
under hydroprocessing conditions. Such hydroprocessing conditions
are normally in the range of 212 degree(s) F to 1200 degree(s) F
(100 degree(s) to 650 degree(s) C.) at pressures of from 20 to 300
atmospheres. Generally such hydroprocessing is in the presence of
catalyst containing group VI or VIII metals such as platinum,
molybdenum, tungsten, nickel, cobalt, etc., in combination with
various other metallic element particles of alumina, silica,
magnesia and so forth having a high surface to volume ratio. More
specifically, catalyst utilized for hydrodemetallation,
hydrodesulfurization, hydrodenitrification, hydrocracking etc., of
heavy oils and the like are generally made up of a carrier or base
material; such as alumina, silica, silica-alumina, or possibly,
crystalline aluminosilicate, with one more promoters or
catalytically active metal(s) (or compound(s)) plus trace
materials. Typical catalytically active metals utilized are cobalt,
molybdenum, nickel and tungsten; however, other metals or compounds
could be selected dependent on the application.
[0007] Because these reactions must be carried out by contact of a
hydrogen-containing gas with the hydrocarbon feed stream at
elevated temperatures and pressures, the major costs of such
processing are essentially investment in vessels and associated
furnaces, heat exchangers, pumps, piping and valves capable of such
service and the replacement cost of catalyst contaminated in such
service. Commercial hydroprocessing of relatively low cost feed
stocks such as reduced crude oils containing pollutant compounds,
requires a flow rate on the order of a few thousand up to one
hundred thousand barrels per day, with concurrent flow of hydrogen
at up to 10,000 standard cubic feet per barrel of the liquid feed.
Vessels capable of containing such a reaction process are
accordingly cost-intensive both due to the need to contain and
withstand corrosion and metal embrittlement by the hydrogen and
sulfur compounds, while carrying out the desired reactions, such as
demetallation, denitrification, desulfurization, and cracking at
elevated pressure and temperatures. Pumps, piping and valves for
handling fluid streams containing hydrogen at such pressures and
temperatures are also costly, because at such pressures seals must
remain hydrogen impervious over extended service periods of many
months and years. It is also cost-intensive to insure that all
additional reactor vessels (e.g. fixed bed reactors, etc.) are
obtained in order to catalytically process hydrocarbon feed streams
from an initial reactor vessel where certain catalytic
hydroprocessing (e.g. hydrodemetallation) is/are to be performed to
one or more other reactor vessel(s) where additional catalytic
hydroprocessing (e.g. hydrodenitrification) is/are to be
performed.
[0008] Further, hydroprocessing catalyst for such one or more
reactor vessel(s), which typically contains catalytically active
metals such as titanium, cobalt, nickel, tungsten, molybdenum,
etc., may involve a catalyst inventory of 500,000 pounds or more at
a cost of $2 to $4/lb. Accordingly, for economic feasibility in
commercial operations, the process must handle high flow rates and
the one or more reactor vessel(s) should be filled with as much
catalyst inventory as possible to maximize catalyst activity and
run length. Additionally, the down-time for replacement or renewal
of catalyst must be as short as possible. Further, the economics of
the process will generally depend upon the versatility of the
system to handle feed streams of varying amounts of contaminants
such as sulfur, nitrogen, metals and/or organic-metallic compounds,
such as those found in a wide variety of the more plentiful (and
hence cheaper) reduced crude oils, residua, or liquified coal, tar
sand bitumen or shale oils, as well as used oils, and the like.
[0009] It is known to use a series of individual reactor vessels
stacked one above the other, with fluid flow either co-current or
counterflow to catalyst. In such a process, catalyst moves by
gravity from the upper vessel to a lower vessel by periodically
shutting off, or closing, valves between the individual vessels. In
a counterflow system, this permits removal of catalyst from the
lowermost or first stage vessel, where the most contaminated, or
raw, feed stock, originally contacts the catalyst. In this way,
most of the major contaminating components in the hydrocarbon
stream are removed before the hydrocarbon material reaches major
conversion steps of the process performed in higher vessels of the
stacked series. Thus, most of the deactivating components of the
feed stream are removed before it reaches the least contaminated
catalyst added to the topmost vessel. However, such systems require
valves suitable for closing off catalyst flow against catalyst
trapped in the line. Hence, valve life is relatively short and
down-time for replacement or repair of the valves is relatively
costly. Also, such series of individual reactor vessels are costly
since each respective reactor vessel must be purchased
separately.
[0010] As particularly described in U.S. Pat. No. 5,076,908 to
Stangeland et al, a substantially packed-bed type reactor system is
an upflow type reactor system including multiple reaction zones of
packed catalyst particles having little or no movement during
normal operating conditions of no catalyst addition or withdrawal.
In the substantially packed-bed type reactor system of Stangeland
et al., when catalyst is withdrawn from the reactor during normal
catalyst replacement, the catalyst flows in a downwardly direction
under essentially plug flow or in an essentially plug flow fashion,
with a minimum of mixing with catalyst in layers which are adjacent
either above or below the catalyst layer under observation.
[0011] As particularly distinguished from prior known methods of
on-stream catalyst replacement in hydroprocessing, the method and
apparatus in U.S. Pat. No. 5,076,908 to Stangeland et al. more
specifically provides a system wherein plug flow of the catalyst
bed is maintained over a wide range of counterflow rates of a
hydrocarbon feed stream and hydrogen gas throughout the volume of
the substantially packed catalyst bed. Such packed catalyst bed
flow maintains substantially maximum volume and density of catalyst
within a given reactor vessel's design volume by controlling the
size, shape and density of the catalyst so that the bed is not
substantially expanded at the design rate of fluid flow
therethrough. The proper size, shape and density are determined by
applying coefficients gained during extensive studying of bed
expansion with hydrocarbon, hydrogen and catalyst at the design
pressures and flow velocities as particularly described below.
However, Stangeland et al. does not teach or suggest a method
and/or a reactor for hydroprocessing a hydrocarbon feed stream
through a single onstream reactor having two or more different
(i.e. different in physical and/or catalytic properties) and
distinct moving bed catalysts, where catalyst particles are added
and withdrawn from each of the two or more different and distinct
moving bed catalysts while the single onstream reactor is
hydroprocessing the hydrocarbon feed stream.
[0012] The prior art does not disclose or suggest two or more
different and distinct moving bed catalysts in a single onstream
reactor; nor does the prior art disclosed or suggest the above
enumerated and pertinent features of either the total system or
significant portions of such a system in U.S. Pat. No. 5,076,908 to
Stangeland et al, as disclosed by the following patents:
[0013] Jacquin et al. U.S. Pat. No. 4,312,741, is directed toward a
method of on-stream catalyst replacement in a hydroprocessing
system by controlling the feed of hydrogen gas at one or more
levels. Catalyst, as an ebullated bed counterflows through the
reactor but is slowed at each of several levels by horizontally
constricted areas which increase the hydrogen and hydrocarbon flow
rates to sufficiently locally slow downward flow of catalyst. While
local recycling thus occurs at each such stage, rapid through-flow
of fresh catalyst, with resultant mixing with deactivated or
contaminated catalyst, is suppressed. The ebullating bed aids
simple gravity withdrawal of catalyst from the vessel. Improvement
of the disclosed system over multiple vessels with valves between
stages is suggested to avoid the risk of rapid wear and
deterioration of valve seals by catalyst abrasion.
[0014] Kodera et al. U.S. Pat. No. 3,716,478, discloses low linear
velocity of a mixed feed of liquid and H.sub.2 gas to avoid
expansion (or contraction) of catalyst bed. By low linear velocity
of fluid upflow, gas bubbles are controlled by flow through the
packed bed, but the bed is fluidized by forming the bottom with a
small cross-sectional area adjacent the withdrawal tube. This
assists discharge of catalyst without backmixing of contaminated
catalyst with fresh catalyst at the top of the single vessel. The
range of the bed level in the vessel is from 0.9 to 1.1 of the
allowable bed volume (.+-.10%) due to fluid flow through the bed. A
particular limitation of the system is that flow of the fluids
undergoing catalytic reaction is restricted to a rate that will not
exceed such limits, but must be adequate to ebullate the bed
adjacent the catalyst withdrawal tube. Alternatively, injection of
auxiliary fluid from a slidable pipe section is required. The
patentees particularly specify that the diameter of the lower end
of the vessel is smaller to increase turbulence and ebullation of
catalyst adjacent the inlet to the catalyst withdrawal line.
Fluidization of catalyst is accordingly indicated to be essential
to the process. However the disclosed gas flow rates are well below
commercial flow rates and there is no suggestion of temperatures or
pressures used in the tests or the size, density or shape of the
catalyst.
[0015] Bischoff et al, U.S. Pat. No. 4,571,326, is directed to
apparatus for withdrawing catalyst through the center of a catalyst
bed counterflowing to a liquid hydrocarbon and gas feed stream. The
system is particularly directed to arrangements for assuring
uniform distribution of hydrogen gas with the liquid feed across
the cross-sectional area of the bed. Such uniform distribution
appears to be created because the bed is ebullating under the
disclosed conditions of flow. Accordingly, considerable reactor
space is used to initially mix the gas and hydrocarbon liquid feeds
in the lower end of the vessel before flowing to other bottom feed
distributors. The feeds are further mixed at a higher level by such
distributor means in the form of "Sulzer Plates" or a "honeycomb"
of hexagonal tubes beneath a truncated, conical, or
pyramidal-shaped funnel screen. The arrangement may include an open
ramp area parallel to the underside of the screen between the tube
or plate ends. Further, to maintain gas distribution along the
length of the catalyst bed, quench gas is supplied through
upflowing jets in star-shaped or annular headers extending across
middle portions of the vessel. The arrangement for withdrawal of
spent catalyst requires ebullation of at least the lower portion of
the bed. As noted above, added vessel space for uniform mixing of
hydrogen and feed before introducing the fluids into an ebullated
bed, as well as an ebullating bed, increases the required size of
the hydroprocessing vessel, increases catalyst attrition, increases
catalyst bed mixing and substantially increases initial, and
continuing operating costs of the system.
[0016] Bischoff et al. U.S. Pat. No. 4,639,354, more fully
describes a method of hydroprocessing, similar to U.S. Pat. No.
4,571,326, wherein similar apparatus obtains uniform ebullation
through the vertical height of a catalyst bed, including a quench
gas step.
[0017] Meaux U.S. Pat. No. 3,336,217, is particularly directed to a
catalyst withdrawal method from an ebullating bed reactor. In the
system, catalyst accumulating at the bottom of a vessel and
supported on a flat bubble-tray may be withdrawn through an
inverted J-tube having a particular ratio of the volume of the
short leg of the J-tube to the longer leg. The diameter of the
J-tube is suited only to flow of catalyst from a body of catalyst
ebullated by the upflowing hydrocarbon feed and gas.
[0018] U.S. Pat. Nos. 4,444,653 and 4,392,943, both to Euzen, et
al., disclose removal systems for catalyst replacement in an
ebullating bed. In these patents, the fluid charge including
hydrocarbon containing gas is introduced by various arrangements of
downwardly directed jets acting laterally against or directly onto
the conical upper surface of the bed support screen or screens.
Alternatively, the feed is introduced through a conical screen
after passing through a distributor arrangement of tortuous paths
or a multiplicity of separate tubes to mix the gas and liquid feed
over the conical screen. Such arrangements use a considerable
volume of the pressure vessel to assure such mixing.
[0019] U.S. Pat. Nos. 3,730,880 and 3,880,569, both to Van der
Toorn, et al., disclose a series of catalytic reactors wherein
catalyst moves downwardly by gravity from vessel to vessel through
check valves. As noted above, such valves require opening and
closing to regulate the rate of flow, or to start and stop catalyst
transfer, with catalyst in the valve flow path. Feed of process
fluids is either co-current or countercurrent through the catalyst
bed.
[0020] Van ZijllLanghaut et al. U.S. Pat. No. 4,259,294, is
directed to a system for on-stream catalyst replacement by
entrainment of the catalyst in oil pumped as a slurry either to
withdraw catalyst from or to supply fresh catalyst to, a reactor
vessel. Reacting feed is suggested to be either co-current or
countercurrent with catalyst flow through the reactor. Valves
capable of closing with catalyst in the line, or after back-flow of
slurry oil, are required to seal off the catalyst containing vessel
at operating temperatures and pressures from the receiving reactor
vessel, or isolate the catalyst receiving lock hopper from the
withdrawal section of the vessel.
[0021] Carson U.S. Pat. No. 3,470,900, and Sikama U.S. Pat. No.
4,167,474, respectively illustrate multiple single bed reactors and
multi-bed reactors in which catalyst is replaced either
continuously or periodically. The feed and catalyst flow
co-currently and/or radially. Catalyst is regenerated and returned
to the reactor, or disposed of. No catalyst withdrawal system is
disclosed apart from either the configuration of the internal bed
support or the shape of the vessel bottom to assist gravity
discharge of catalyst.
[0022] One of the basic principles and teachings of Stangeland et
al. in U.S. Pat. No. 5,076,908, is that by specifically selecting
the size, shape, and density of the catalyst pellets, combined with
appropriate control of process liquid and gas velocities, random
motion and backmixing of the catalyst can be minimized, and
plugflow characteristics of the catalyst downward and the liquid
and gas flows upward, maximized. Stangeland et al. economically
utilizes space within a hydroprocessing vessel over a wide range of
processing rates without substantial random motion or ebullation of
a packed bed of catalyst during high counterflow rates of the
hydrocarbon feed and a hydrogen containing gas through the packed
bed, while maintaining continuous or intermittent replacement of
catalyst for plug-like flow of the bed through the vessel. Such
plug flow with high processing rates is obtained by Stangeland et
al. by selecting the size, shape and density of the catalyst
particles to prevent ebullation and bed expansion at the design
flow rate so as to maximize the amount of catalyst in the vessel
during normal operation and during catalyst transfer. Catalysts are
selected utilizing data gained while studying catalyst bed
expansion, such as in a large pilot plant run, with hydrocarbon,
hydrogen and catalyst at the design pressures and flow velocities
within the available reaction volume of the vessel. Catalyst is
removed from the bed by Stangeland et al. through laminar flow of
the catalyst particles in a liquid slurry system in which the
liquid flow line is uniform in diameter, and substantially larger
than the catalyst particles, throughout the flow path between the
reactor vessel and a pressurizable vessel including passageways
through the flow control valves.
[0023] However, the method and reactor disclosed by Stangeland et
al. in U.S. Pat. No. 5,076,908, as well as the method(s) and
reactor(s) taught by the above-identified prior art patents
relating to U.S. Pat. No. 5,076,908 to Stangeland et al, all
suggest that additional reactor vessels are needed to further
process hydrocarbon products produced by a reactor vessel
containing a single catalyst bed. The reactor vessel(s) of the
prior art allow the employment of one catalyst bed in a single
reactor. If two or more catalyst beds are needed and they are not
to be mixed together, then two or more separate moving-bed reactors
placed in series are required. Separate reactors for separate
purposes are expensive. Therefore, what is needed and what has been
invented is a method and a reactor that is capable of containing
two or more separate and distinct moving catalyst beds wherein each
of separate and distinct moving catalyst has a different catalytic
purpose (e.g. one moving catalyst bed would be for
hydrodemetallation; a second moving catalyst bed would be for
hydrodenitrification; a third moving catalyst bed would be for
hydrodesulfurization; etc.).
SUMMARY OF THE INVENTION
[0024] The present invention accomplishes its desired objects by
providing a reactor comprising a reactor vessel having an internal
cylindrical wall with an internal reactor cylindrical wall
diameter; and a first catalyst bed support means coupled to and
supported by the internal cylindrical wall of the reactor vessel
such as to essentially entirely span across the internal reactor
cylindrical wall diameter for supporting a first catalyst moving
bed in a first hydroconversion reaction zone having a first
physical and/or catalytic properties and for allowing a hydrocarbon
feed stream to pass therethrough while preventing catalyst
particles from a first catalyst moving bed to exit therethrough.
The reactor also comprises a first catalyst outlet means,
communicating through the internal cylindrical wall of the reactor
vessel with a first catalyst moving bed supported by the first
catalyst bed support means, for withdrawing a first generally spent
catalyst from a first hydroconversion reaction zone; and a first
catalyst inlet means, communicating through the internal
cylindrical wall of the reactor vessel with a first catalyst moving
bed supported by the first catalyst bed support means, for
introducing a first generally fresh catalyst into a first
hydroconversion reaction zone. A second catalyst bed support means
is coupled to and supported by the internal cylindrical wall of the
reactor vessel such as to essentially entirely span across the
internal reactor cylindrical wall diameter for supporting a second
catalyst moving bed in a second hydroconversion reaction zone
having a second physical and/or catalytic properties and for
allowing a hydrocarbon feed stream to pass therethrough while
preventing catalyst particles from a second catalyst moving bed to
exit therethrough. The reactor further includes a second catalyst
outlet means, communicating through the internal cylindrical wall
of the reactor vessel with a second catalyst moving bed supported
by the second catalyst bed support means, for withdrawing a second
generally spent catalyst from a second hydroconversion reaction
zone; and a second catalyst inlet means, communicating through the
internal cylindrical wall of the reactor vessel with a second
catalyst bed supported by the second catalyst bed support means,
for introducing a second generally fresh catalyst into a second
hydroconversion reaction zone.
[0025] The reactor additionally comprises a first catalyst moving
bed supported by the first catalyst bed support means and having a
first physical and/or catalytic properties; and a second catalyst
moving bed supported by the second catalyst bed support means and
having a second physical and/or catalytic properties. The first
physical and/or catalytic properties differ from the second
physical and/or catalytic properties. A first conduit means,
communicating through the internal cylindrical wall of the reactor
vessel is provided for introducing a hydrocarbon feed stream into
the reactor vessel; and a second conduit means, also communicating
through the internal cylindrical wall of the reactor vessel is
provided for passing a hydrocarbon product out of the reactor
vessel. The first catalyst bed support means allows an untreated
hydrocarbon feed stream from the first conduit means to pass
therethrough while preventing catalyst particles from the first
catalyst moving bed to exit therethrough; and the second catalyst
bed support means allows an initially treated hydrocarbon feed
stream from the first catalyst moving bed to pass therethrough
while preventing catalyst particles from the second catalyst moving
bed to exit therethrough. Alternatively, the second catalyst bed
support means receives an untreated hydrocarbon feed stream and
allows an initially treated hydrocarbon feed stream from the second
catalyst moving bed to pass therethrough while preventing catalyst
particles from the second catalyst moving bed to exit therethrough;
and the first catalyst bed support means allows a treated
hydrocarbon feed stream from the first catalyst moving bed to pass
therethrough while preventing catalyst particles from the first
catalyst moving bed to exit therethrough.
[0026] The present invention further accomplishes its desired
objects by broadly providing a method for hydroprocessing a
hydrocarbon feed stream comprising the steps of:
[0027] (a) disposing a first catalyst bed in a first
hydroconversion reaction zone of a reactor vessel, wherein the
first catalyst bed contains a first set of catalytic
properties;
[0028] (b) disposing a second catalyst bed in a second
hydroconversion reaction zone of the reactor vessel of step (a),
wherein the second catalyst bed contains a second set of catalytic
properties;
[0029] (c) flowing through the first catalyst bed in the first
hydroconversion reaction zone a hydrocarbon feed stream to produce
an initially-treated hydrocarbon feed stream; and
[0030] (d) flowing through the second catalyst bed in the second
hydroconversion reaction zone the initially-treated hydrocarbon
feed stream to produce a treated hydrocarbon feed stream.
[0031] The method additionally comprises withdrawing at least
partially spent first catalyst from the first catalyst bed
simultaneously with the flowing step (c); and adding fresh first
catalyst to the first catalyst bed simultaneously with the flowing
step (c). The method further provides for withdrawing at least
partially spent second catalyst from the second catalyst bed
simultaneously with the flowing step (c); and adding fresh second
catalyst to the second catalyst bed simultaneously with the flowing
step (c). The flowing step (d) comprises flowing the
initially-treated hydrocarbon feed stream through the second
catalyst bed without the initially-treated hydrocarbon feed stream
having left or exited the reactor vessel of the step (a).
[0032] The method additionally includes demetallizing the
hydrocarbon feed stream of the step (c) in the first catalyst bed;
and hydrotreating the initially-treated hydrocarbon feed stream of
the step (d) in the second catalyst bed to remove at least one
chemical element from the initially-treated hydrocarbon feed
stream. The at least one chemical element is selected from the
group consisting of nitrogen, sulfur, and mixtures thereof.
[0033] The present invention yet further accomplishes its desired
objects by providing a reactor comprising a reactor vessel having
an internal cylindrical wall with an internal reactor cylindrical
wall diameter; a first distributor plate assembly coupled to and
supported by the internal cylindrical wall of the reactor vessel;
and a first catalyst bed support assembly coupled to and supported
by the internal cylindrical wall of the reactor vessel such as to
define a first plenum chamber between the first distributor plate
assembly and the first catalyst bed support assembly and for
supporting a first catalyst moving bed in a first hydroconversion
reaction zone having a first set of physical and/or catalytic
properties and for allowing a hydrocarbon feed stream to pass
therethrough from the first plenum chamber while preventing
catalyst particles from a first catalyst moving bed to exit
therethrough. The reactor also comprises a first catalyst outlet
means, communicating through the internal cylindrical wall of the
reactor vessel with a first catalyst moving bed supported by the
first catalyst bed support assembly, for withdrawing a first
generally spent catalyst from a first hydroconversion reaction
zone; and a first catalyst inlet means, communicating through the
internal cylindrical wall of the reactor vessel with a first
catalyst moving bed supported by the first catalyst bed support
assembly, for introducing a first generally fresh catalyst into a
first hydroconversion reaction zone. A second distributor plate
assembly is coupled to and supported by the internal cylindrical
wall of the reactor vessel; and a second catalyst bed support
assembly is also coupled to and supported by the internal
cylindrical wall of the reactor vessel such as to define a second
plenum chamber between the second distributor assembly and the
second catalyst bed support assembly and for supporting a second
catalyst moving bed in a second hydroconversion reaction zone
having a second set of physical and/or catalytic properties and for
allowing a hydrocarbon feed stream to pass therethrough from the
second plenum chamber while preventing catalyst particles from a
second catalyst moving bed to exit therethrough. The reactor
further includes a second catalyst outlet means, communicating
through the internal cylindrical wall of the reactor vessel with a
second catalyst moving bed supported by the second catalyst bed
support assembly, for withdrawing a second generally spent catalyst
from a second hydroconversion reaction zone; and a second catalyst
inlet means, communicating through the internal cylindrical wall of
the reactor vessel with a second catalyst bed supported by the
second catalyst bed support assembly, for introducing a second
generally fresh catalyst into a second hydroconversion reaction
zone.
[0034] The reactor additionally comprises a first catalyst moving
bed supported by the first catalyst bed support assembly and having
a first set of physical and/or catalytic properties; and a second
catalyst moving bed supported by the second catalyst bed support
assembly and having a second physical and/or catalytic properties.
The first set of physical and/or catalytic properties differ from
the second set of physical and/or catalytic properties. A first
conduit means, communicating through the internal cylindrical wall
of the reactor vessel, for introducing a hydrocarbon feed stream
into the reactor vessel; and a second conduit means, communicating
through the internal cylindrical wall of the reactor vessel, for
passing a hydrocarbon product out of the reactor vessel. The first
catalyst bed support assembly allows an untreated hydrocarbon feed
stream from the first conduit means to pass therethrough while
preventing catalyst particles from the first catalyst moving bed to
exit therethrough; and the second catalyst bed support assembly
allows an initially treated hydrocarbon feed stream from the first
catalyst moving bed to pass therethrough while preventing catalyst
particles from the second catalyst moving bed to exit therethrough.
Alternatively, the second catalyst bed support assembly receives an
untreated hydrocarbon feed stream and allows an initially-treated
hydrocarbon feed stream from the second catalyst moving bed to pass
therethrough while preventing catalyst particles from the second
catalyst moving bed to exit therethrough; and the first catalyst
bed support assembly allows a treated hydrocarbon feed stream from
the first catalyst moving bed to pass therethrough while preventing
catalyst particles from the first catalyst moving bed to exit
therethrough.
[0035] In a preferred embodiment of the invention, the catalyst
comprises a plurality of catalytic particulates having a mean
diameter ranging from about 35 Tyler mesh to about 3 Tyler mesh;
and a size distribution such that at least about 90% by weight of
the catalytic particulates have an aspect ratio less than about 2.0
and a diameter ranging from R.sub.1 to about R.sub.2, wherein:
[0036] (1) R.sub.1 has a value ranging from about {fraction (1/64)}
inch to about 1/4 inch,
[0037] (2) R.sub.2 has a value ranging from about {fraction (1/64)}
inch to about 1/4 inch, and
[0038] (3) a value of a ratio R.sub.2/R.sub.1 ranges from about 1.0
to about 1.4.
[0039] The catalyst may be employed in any hydrogenation process.
Preferably, the catalyst is for producing a plug-flowing
substantially packed bed of hydroprocessing catalyst during
hydroprocessing by contacting a substantially packed bed of
hydroprocessing catalyst with an upflowing hydrocarbon feed stream.
More particularly, when the catalytic particulates are disposed in
a hydrocarbon reaction zone, a substantially packed bed of
hydroprocessing catalyst is produced; and when a hydrocarbon feed
stream flows upwardly through the substantially packed bed of
hydroprocessing catalyst, plug-flowing commences when a volume of
the catalytic particulates is withdrawn from a bottom of the
hydrocarbon reaction zone. As used herein "catalyst" includes other
particles which interact with a feed stream, such as sorbents, or
other fluid contact bodies. The catalyst is disposed in a reaction
zone and a hydrocarbon feed stream is flowed upwardly through the
catalyst for hydroprocessing the hydrocarbon feed stream.
[0040] The catalytic particulates have a size distribution such
that a maximum of about 2.0% by weight of said catalytic
particulates have a diameter less than R.sub.1. The catalytic
particulates also have a size distribution such that a maximum of
about 0.4% by weight of the catalytic particulates have a diameter
less than R.sub.3, wherein R.sub.3 is less than R.sub.1 and the
value of the ratio R.sub.1/R.sub.3 is about 1.4. The catalytic
particulates have a maximum attrition of about 1.0% by weight of
the catalytic particulates through a diameter having a value of
R.sub.1; and the catalytic particulates have a maximum attrition of
about 0.4% by weight of the catalytic particulates through a
diameter having a value of R.sub.3, wherein R.sub.3 is less than
R.sub.1 and the value of the ratio R.sub.1/R.sub.3 is about
1.4.
[0041] From the foregoing summary it will be apparent that several
significant factors contribute directly to the present invention
accomplishing its desired objects, and to the efficient use of a
given process reactor vessel to assure non-ebullating, plug-like
flow of a body of catalyst particles therethrough while being
contacted by a counter-flowing hydrocarbon fluid stream of gas and
liquid at maximum space-velocity. Among such significant factors
are: (i) the size, volume and density characteristics of such
catalyst particles at preselectable flow velocities and pressures
of the hydrocarbon fluid stream; (ii) control of catalyst bed
ebullation and/or levitation during hydrocarbon fluid and hydrogen
flows; (iii) laminar flow of the catalyst particles during movement
into and out of the catalyst moving bed for replacement (or
regeneration or rejuvenation) without bed ebullation or levitation;
(iv) and for a preferred catalyst support means, concentric annular
feed of alternate rings of the gas and liquid components of the
hydrocarbon feed uniformly into the full moving catalyst bed, which
is capable of recovering promptly from upset or pressure changes in
the reactor vessel to restore such alternate rings of gas and
liquid over process runs of extended length (e.g. several thousand
hours); and (v) redistribution of the gas components along the
axial length of the moving bed.
[0042] It is another object of the present invention to broadly
provide a method for hydroprocessing a hydrocarbon feed stream
through two or more essentially downwardly plug-flowing
substantially packed bed of hydroprocessing catalyst within two or
more hydroconversion reaction zones within a single onstream
reactor.
[0043] These, together with the various ancillary objects and
features which will become apparent to those skilled in the art as
the following description proceeds, are attained by this invention,
a preferred embodiment as shown with reference to the accompanying
drawings, by way of example only, wherein:
BRIEF DESCRIPTION OF THE DRAWINGS
[0044] FIG. 1 is a schematic view of a typical hydroprocessing
vessel to which the present invention is particularly directed for
on-stream catalyst replacement during continuous plug-like flow of
catalyst through up-flowing liquid hydrocarbon feed and gas
streams;
[0045] FIG. 2 is a bottom plan view of the concentric and radial
catalyst bed support means for a truncated conical or pyramidal
screen, taken in the direction of arrows and along the plane of
line 2-2 in FIG. 1.;
[0046] FIG. 3 is an elevational cross-sectional view of the support
means and screen taken in the direction of arrows and along the
plane of line 3-3 in FIG. 2;
[0047] FIG. 4 is a partial elevational view of an alternate form of
a laminar flow arrangement for withdrawing deactivated catalyst
particles from the reactor bed;
[0048] FIG. 5 is a cross-sectional plan view of the reactor vessel
taken in the direction of arrows and along the plane of line 5-5 in
FIG. 1 showing a preferred form of gas redistribution and quench
system over a central portion of the catalyst bed;
[0049] FIG. 6 is a perspective view, partially in cross-section, of
one of the quench or redistribution shed units shown in FIG. 5;
[0050] FIG. 7 is a perspective view of a preferred arrangement of
two tiers of shed units of FIG. 5 at a given level in the catalyst
bed;
[0051] FIG. 8 is a partial cross-sectional view illustrating a
catalytic bed with a plurality of superimposed layers with respect
to each other before commencement of a plug-flow;
[0052] FIG. 9 is a partial cross-sectional view illustrating a
catalytic bed which is moving downwardly in a plug-flow
fashion;
[0053] FIG. 10 is a partial cross-sectional view of the reactor and
a partial perspective view of another embodiment of the catalytic
support means;
[0054] FIG. 11 is a partial cross-sectional view of the reactor and
the catalytic support means of FIG. 10 which includes a plurality
of annular mixture zones under the substantially packed bed of
hydroprocessing catalyst with each annular mixture zone containing
a liquid hydrocarbon component and a hydrogen-containing gas
component and wherein the annular mixture zones are concentric with
respect to each other and are coaxial with respect to the reactor
and the substantially packed bed of hydroprocessing catalyst;
[0055] FIG. 12 is the partial cross-sectional view of the reactor
and support means in FIG. 11 with the inert pellets, and
illustrating ribs or spokes secured to an imperforate center plate
and supporting a plurality of segmented plates;
[0056] FIG. 13 is another cross-sectional view of the reactor and
support means as similarly illustrated in FIG. 12 with a bed of
inert pellets having a liquid hydrocarbon component and a
hydrogen-containing gas component flowing around the inert pellets
for entering the annular mixture zones;
[0057] FIG. 14 is a schematic diagram of another embodiment of the
present invention having a moving-bed reactor containing two
separate hydroconversion reaction zones;
[0058] FIG. 15 is schematic diagram for another embodiment of the
present invention depicted in FIG. 14; and
[0059] FIG. 16 is another schematic diagram for another embodiment
of the present invention depicted in FIGS. 14 and 15.
DETAILED DESCRIPTION OF THE INVENTION INCLUDING PREFERRED
EMBODIMENTS OF THE INVENTION
[0060] Referring in detail now to the drawings, and initially more
particularly to FIG. 1, a hydroprocessing system is shown embodying
the method of the present invention to increase substantially both
the continued catalytic activity of a single volume or single bed
of catalyst 10 and the efficient use of a single reactor vessel of
a given reactor volume, such as reactor vessel 11. Vessel 11, as
indicated by the thickness of its cylindrical side wall 12 and
domed closure heads, or ends, 13 and 14, is designed to react a
hydrogen containing gas mixed with a liquid hydrocarbon stream at a
pressure of up to about 300 atmospheres (about 4500 lbs per square
inch) and up to about 650.degree. C. (about 1200.degree. F.). Such
reaction gas and a feed stream of hydrocarbon liquids are
preferably premixed and introduced as a single stream through
bottom head 13 by line 16.
[0061] To assure maximum catalytic benefit during the
hydroprocessing of the hydrocarbon feed stream and the
hydrogen-containing gas, it is essential that vessel 11 contain as
much catalyst as possible within the design volume of vessel 11.
Accordingly as indicated, support means 17 for catalyst bed 10 is
placed as low as possible in vessel 11 while assuring full and
adequate dispersion of the hydrogen phase within the liquid
hydrocarbon stream. At the same time, the upper limit of bed 10 is
near the top of domed head 14, while providing an adequate space 21
for disengaging any entrained catalyst from the resulting products
withdrawn through center pipe 18. To insure that catalyst is not
entrained into product fluids exiting through center pipe 18, a
screen 15 may be installed in space 21 above a bed surface 20
defining the top of the catalyst bed 10. Fresh catalyst is then
added to bed surface 20 through pipe 19 extending through screen
15. Desirably, the upper level or top of the catalyst bed 10,
designated as the bed surface 20, is preferably controlled on a
continuous basis by gamma ray absorption measurement made possible
by a gamma ray source 22 and gamma ray detector 24 positioned in
close proximity to the bed surface 20 of catalyst bed 10. Such a
gamma ray source may be in the form of radioactive isotopes, such
as Cesium 137, disposed inside the reactor in a specially designed
well. Alternatively the source can be an electrically controllable
source, such as a thermal neutron activated gamma ray generator.
Detector 24 may be in the form of an ionization tube,
Geiger-Mueller tube or a scintillation detector. Suitable sources
and detectors are manufactured by Ronan Engineering Co., Texas
Nuclear and other vendors. By detecting the level of surface 20, it
is possible, in accordance with the invention, to insure that the
catalyst inventory is maintained at the optimum level and that the
reactor is never overfilled.
[0062] Overfilling the reactor increases the chance that catalyst
particles will be crushed in the isolation valves in the transfer
lines when they are closed, at the end of each transfer. Bed level
control is also needed to confirm that ebullation of the bed is
minimized and that undesirable excursions from the design flow rate
for hydrogen and hydrocarbon feed flowing upwardly through bed 10
are avoided for the selected catalyst. To this end, the size,
shape, and density of catalyst particles supplied to the bed are
selected in accordance with the designed maximum rate of flow of
the feed streams to prevent such ebullation. Such control assures
that bed 10 progressively moves down through vessel 11 in layers as
by a plug flow. A "plug flow" of the catalyst bed 10 is illustrated
in FIGS. 8 and 9 and may be best described as when a lowermost
volumetric layer A is removed, the next volumetric layer B flows
downwardly to replace the lowermost volumetric layer B and assumes
a new position as a lowermost volumetric layer B. The removed
lowermost volumetric layer A is replaced with an upper volumetric
layer J. The procedure is again repeated (as best shown by the
dotted line representations in FIG. 9) by removing the lowermost
volumetric layer B and causing the next volumetric layer C to flow
downwardly in a plug-like fashion to replace the lowermost
volumetric layer B and assume a new position as a lowermost
volumetric layer C. The removed lowermost volumetric layer B is
replaced with an upper volumetric layer K. The procedure may be
continually repeated to define a downwardly plug-flowing catalyst
bed 10 which is moving in direction of arrow W in FIG. 9.
[0063] The procedure to determine whether or not a catalyst bed 10
is plug-flowing may be by any suitable procedure. For example, in a
preferred embodiment of the present invention wherein metals (e.g.
vanadium) are being removed from a hydrocarbon feed stream, the
catalyst bed 10 is plug-flowing if a catalytic sample (e.g. 15
catalytic particulates) from withdrawn catalyst is analyzed and it
is found through elemental metal analysis that the catalytic sample
has a uniform high metal load, preferably at least about 1.5 times
more than the average metal load of the catalyst bed 10, and more
preferably at least about 2.0 times more than the average metal
load of the catalyst bed 10. Those possessing the ordinary skill in
the art can determine the average load of the catalyst bed 10 from
the total amount of metals removed from the hydrocarbon feed
stream, the weight of the catalytic bed 10, etc.
[0064] It is to be understood that whenever the specification or
the claims states or mentions any type of catalyst movement or
catalyst bed 10 movement (e.g. "removing", "moving", "supplying",
"replacing", "delivering", "flow", "flowing", "transfer",
"transferring", "addition", "adding", "admixing", etc.) for any
type or mixture of catalyst without stating or mentioning the
basis, the basis for such type of catalyst or catalyst bed movement
may be on any type of basis, such as "intermittent basis",
"periodic basis", "continuous basis", "semi-continuous basis", etc.
Thus, by way of example only, removal of lowermost volumetric
catalytic layers and addition of upper volumetric catalytic layers
may be on a "periodic basis", "a continuous basis", or even "a one
time basis", all without affecting the spirit and scope of the
present invention(s). It is to be also understood that the
"removal" or "withdrawal" of catalyst and the "addition" or
"replacement" of catalyst are mutually exclusive of each other and
may be performed simultaneously or at different times without
affecting the spirit and scope and of the present invention(s).
Preferably, the "addition" or "replacement" of catalyst is
performed after the "removal" or "withdrawal" of catalyst and after
the catalyst bed 10 has moved downwardly into a non-moving state or
non-moving posture. Catalysts are selected utilizing data acquired
by measuring bed expansion, such as in a large pilot plant run,
with hydrocarbon, hydrogen and catalyst.
[0065] To further assure that plug flow continues throughout the
full length of the bed, and particularly at the bottom portion, bed
support means 17 is particularly characterized by the truncated
polygonal or conical configuration of support means 17.
[0066] As shown in the preferred embodiment of FIGS. 2 and 3, and
as best seen in FIG. 2, support 17 includes a series of annular
polygons, approaching the form of annular rings, formed by a
plurality of segment plates 27 between radial ribs or spokes 26
extending from imperforate center plate 25 to sidewall 12 of vessel
11. As shown in FIG. 3, spokes 26 may be any suitable geometric
shape, such as rod-like (see FIG. 3) or substantially flat plates
(see FIG. 10), which divide the circumference of the vessel into
many segments (eight in this case) and similarly support the ends
of outer octagonal ring 23 of support means 17 formed by annular or
circumferential plates 27. In each case, radial ribs or spokes 26,
and annular segment plates 27 form a plurality of concentric rings,
or annular polygons which support conical, or pyramidal, perforated
plate or screen 28. Thus screen 28 is permeable to both gas and
liquid rising from the lower portion of vessel 11.
[0067] In one preferred embodiment of the particular merit of the
concentric annular polygons as illustrated in FIG. 3, the
interconnected plates 27 form a plurality of ring-like structures
extending generally axially parallel to the sidewall 12 with the
radial ribs or spokes 26 radially extending towards the sidewall 12
of reactor vessel 11. The mixture of the hydrocarbon liquid feed
and hydrogen gas that is to enter the catalyst bed 10 separates by
gravity into radially alternate gas and liquid rings, made up of
adjacent segments between each pair or radial spokes 26. Thus, both
phases flow upwardly through alternate concentric annular passages
under screen 28. The preferential separation of gas from liquid in
each ring includes an annular cap segment of gas overlying an
adjacent lower annular segment filled with liquid. Hence, both
fluids have equal, and angularly adjacent, access to the bed
through screen 28. The plurality of alternate annular rings of
hydrogen gas and hydrocarbon liquid assure even and equal feed of
both phases across the full cross-sectional area of screen 28 into
bed 10. Among other factors, we have particularly found that this
configuration insures even and equal distribution across the full
cross-sectional area of the catalyst bed. Such equal distribution
across the full diameter of the bed 10, permits a quiescent flow
section to form directly above center plate 25 which truncates
conical bed support means 17. This decreases substantially
potential local ebullation or eddy currents from being induced in
the catalyst bed at the point of catalyst withdrawal through inlet
30 of inverted J-tube 29 to assure localized laminar flow of
catalyst and liquid from within bed 10.
[0068] Uniform feed of the mixture of the hydrocarbon feed stream
and hydrogen is particularly facilitated to the inlet side of
plates 27 of support means 17 through plenum or inlet chamber 33
enclosed between support 17 and circular plate member 31, which
extends across the full cross-sectional area of vessel 11. The
circular plate member 31 defines a grid-like structure for
supporting a permeable screen 6 having one or more openings, as
best shown in FIGS. 11, 12 and 13. As further best shown in FIGS.
11, 12 and 13, the permeable screen 6 supports a bed 3 of a
plurality of inert pellets 4 (e.g. alumina pellets) which are sized
not to pass through the openings in the permeable screen 6, to
prevent eddy currents in the plenum chamber 33, and to keep bubbles
of hydrogen-containing gas diffused within the hydrocarbon feed
streams. Plate 31 includes a multiplicity of similar large diameter
tubes 32 forming openings through plate 31. Each tube is several
inches in diameter and extends axially to a similar depth, say on
the order of 4 to 6 inches, below plate 31. Tubes 32 provide equal
access to the mixture of hydrogen and hydrocarbon feed stream into
plenum chamber 33. Even distribution of the incoming feed stream
into bottom header 35 from feed line 16 may also be assisted by
deflector plate 34 to assure that oversized bubbles of hydrogen
that may be contained in the feed stream will be equally
distributed across the full cross-sectional area of plate 31 and
equally distributed to each of tubes 32 for flow into plenum
chamber 33. The length of tubes 32 may be selected to form a
suitable gas head under plate 31 to suppress surges in the feed
streams entering header 35.
[0069] As noted above, the vertical, transverse width or axial
length of plates 27 which set off each individual annular and
radial segment, provide equal access to both hydrogen and liquid
feed into catalyst bed 10, and stepped under screen 28 so that they
effectively form rings of gas and hydrocarbon feed alternately
across the full diameter at the inlet side of catalyst bed 10. In
this way, no single area of the inlet to catalyst bed 10 becomes a
segregated or preferential, flow path for either gas or the liquid.
Further, if pressure surges result in full wetting of screen 28 by
the liquid phase, recovery of gas flow is assisted by the aerial
breadth of each segment between plates 27 and radial plates 26.
[0070] In another preferred embodiment of the particular merit of
the concentric annular polygons as illustrated in FIGS. 10, 11, 12
and 13, there is seen a liquid hydrocarbon component LH and a
hydrogen-containing gas component HG (hydrogen-containing gas
bubbles) entering as an LH--HG mixture into the plenum chamber 33
from tubes 32. The LH--HG mixture is introduced into the plenum
chamber 33. In this preferred embodiment of the present invention,
the annular or circumferential plates 27 are secured to and are
supported by the radial ribs or spokes 26, each of which has a
vertical or transverse width that is essentially equal to the
vertical or transverse width of the annular or circumferential
plates 27. The radial ribs or spokes 26 also function as a means
for reducing a size of hydrogen-containing gas bubbles, especially
over-size hydrogen-containing gas bubbles from the
hydrogen-containing gas component HG. Those skilled in the art will
readily recognize that the number of radial ribs or spokes 26
employed will depend on a number of factors, such as the
anticipated number of over-size hydrogen-containing gas bubbles in
the upwardly flowing hydrocarbon feed stream, the weight of the
catalyst bed 10, etc. The interconnected plates 27 and radial ribs
or spokes 26 form a web or web-like structure defining a plurality
of annular mixture zones, generally illustrated as MZ in FIGS. 10
and 11. The annular mixture zones MZ are essentially continuous or
are generally endless annular mixture zones MZ, and may contain or
be subdivided into any reasonable desired number of mixture zones
(or sub-mixture zones), such as MZ.sub.1, MZ.sub.2, MZ.sub.3,
MZ.sub.4, MZ.sub.5, and MZ.sub.6 in FIGS. 10 and 11. Each of the
individual mixture zones MZ.sub.1, MZ.sub.2, MZ.sub.3, MZ.sub.4,
MZ.sub.5, and MZ.sub.6 is for all practical purposes an annularly
continuous or endless mixture zone of uniform thickness, excepting
a periodic interruption by radially ribs 26, which are relatively
narrow vis-a-vis the spaced distance between any pair of contiguous
ribs 26-26. As evident in FIGS. 10, 11, 12 and 13, concentric with
mixture zone MZ.sub.1 and as a partial bottom to same is
imperforate center plate 25, which is preferably spaced from and
off of the plate 31 and the screen 6 such that inert pellets 4 may
be supported by the screen 6 and the plate 31 immediately
underneath the imperforate center plate 25. Mixture zone MZ.sub.1
is essentially a cylindrical annular mixture zone with an open top
and boundaries defined by the space between a plurality of
interengaged and coupled plates 27.sub.1s and the perimeter of the
imperforate center plate 25.
[0071] The plurality of annular mixture zones MZ (or the annularly
continuous or endless mixture zones MZ.sub.2s, MZ.sub.3s,
MZ.sub.4s, MZ.sub.5s, and MZ.sub.6s) under the catalyst bed 10 are
concentric with respect to each other and are coaxial with respect
to the reactor vessel 11 and the catalyst bed 10. The plates 27 may
be radially spaced from each other at any suitable distance
(preferably of uniform distance) to assist in accomplishing the
desired objects of the present invention; however, preferably the
plates 27 are radially spaced from each other at a generally
uniform thickness or distance that ranges from about 1 inch to
about 4 feet, more preferably from about 6 inches to about 3 feet,
most preferably from about 1 foot to about 2 feet. The radially
spaced relationship between and among the plates 27 generally
defines a uniform thickness for each of the mixture zones (i.e.
MZ.sub.2s, MZ.sub.3s, etc.). It is to be understood that while the
plurality of annular mixture zones MZ is represented in FIGS. 2, 3,
10, 11, 12 and 13 as being a plurality of non-circular
geometric-shaped zones (e.g. octagonal in FIG. 2), the spirit and
scope of the present invention includes that the plurality mixture
zones MZ may comprise any geometric-shaped zones including not only
polygonal-shaped zones, but also a plurality of concentric circular
mixture zones, etc., all of which would also be concentric with
respect to each other and coaxial with respect to the reactor
vessel 11 and/or the catalyst bed 10 (or the hydroconversion
reaction zone).
[0072] Therefore, the plates 27 function to form generally uniform
thick and essentially circular bands of concentric hydrocarbon feed
streams that are also coaxial with respect to the catalyst bed 10.
By way of example only and as best shown in FIGS. 2 and 10, angular
mixture zone MZ.sub.2 is defined by the eight (8) interengaged or
intercoupled plates 27.sub.1s and the eight (8) interengaged or
intercoupled plates 27.sub.2s. The eight (8) plates 27.sub.1s and
the eight (8) plates 27.sub.2s each form an annulate boundary for
the essentially circular band of hydrocarbon feed stream in mixture
zone MZ.sub.2. Because the spacing or distance between plates
27.sub.1s and 27.sub.2s is generally circumferentially uniform, the
thickness or size of the essentially circular band of hydrocarbon
feed stream in mixture zone MZ.sub.2 is essentially uniform
transversely and/or equal in transverse or horizontal cross
section. Similarly, mixture zone MZ.sub.6 is defined by the eight
(8) interengaged or intercoupled plates 27.sub.5s and the eight (8)
interengaged or intercoupled plates 27.sub.6s, the combination of
which form annulate boundaries for the essentially circular band of
hydrocarbon feed stream in mixture zone MZ.sub.6. As was previously
similarly indicated for plates 27.sub.1s and 27.sub.2s, because the
spacing or distance between plates 27.sub.5s and 27.sub.6s is
generally circumferentially uniform, the thickness or size of the
circular band of hydrocarbon feed stream in mixture zone MZ.sub.6
is essentially uniform transversely and/or equal in transverse or
horizontal cross section. Plates 27.sub.2, 27.sub.3, 27.sub.4, and
27.sub.5 similarly functionally interengage and intercouple to
define annulate boundaries for mixture zones MZ.sub.3, MZ.sub.4,
and MZ.sub.5. As indicated and as best shown in FIG. 2, ribs 26
extend radially from imperforate center plate 25 and planarly
represent visually pie-shaped segments. Between any pair of
contiguous ribs 26-26, the lengths of the respective plates 27
increase from plate 27.sub.1 to or towards plate 27.sub.6 while the
widths are essentially the same as best shown in FIG. 3. Thus,
plate 27.sub.2 is longer than plate 27.sub.1 while possessing the
identical approximate width. Likewise: plate 27.sub.3 is longer
than plate 27.sub.2, plate 27.sub.4 is longer than 27.sub.3, plate
27.sub.5 is longer than plate 27.sub.4, and plate 27.sub.6 is
longer than plate 27.sub.5, while all the plates 27 simultaneously
have generally the same width or the same longitudinal extension
below the screen 28 (see FIG. 3). Thus, the vertical dimensions or
the widths of the plates 27 (i.e. the structural extensions of the
plates 27 that are generally parallel to the longitudinal axis of
the reactor vessel 11 and/or the catalyst bed 10 therein) are
generally equal. All plates 27 are preferably spaced such that the
hydrocarbon feed stream flows parallel to the longitudinal axis of
the catalyst bed 10 before contacting and entering the same. Both
the upper edges and lower edges of plates 27.sub.1s, 27.sub.2s,
27.sub.3s, 27.sub.4s, 27.sub.5s, and 27.sub.6s are all at a
different level or height, as best shown in FIGS. 3 and 11. The
mixture zones MZ differ from a plurality of tubes, conduits, or
pipe-like passages for introducing an essentially complete or
essentially total integral cylindrical hydrocarbon feed stream into
the catalytic bed 10. As best shown in FIGS. 3 and 11, the upper
and lower edges of plates 27.sub.1s are at a different level or
height than the upper and lower edges of plates 27.sub.2s which are
at a different level or height than the upper and lower edges of
plates 27.sub.3s. Similarly, the upper and lower edges of plates
27.sub.3s are at a different level or height than the upper and
lower edges of plates 27.sub.4s which are at a different level or
height than the upper and lower edges of plates 27.sub.5s. The
upper and lower edges of the latter are at a different level or
height than the upper and lower edges of plates 27.sub.6s.
[0073] After the LH--HG mixture enters and flows through the screen
6 into the plenum chamber 33, the flowing LH--HG mixture enters
into each of the generally continuous annular mixture zones
MZ.sub.2s, MZ.sub.3s, etc. for dividing or separating the flowing
LH--HG mixture into a plurality of flowing generally continuous
annular LH--HG mixtures, which have been designated LH--HG.sub.2,
LH--HG.sub.3, LH--HG.sub.4, LH--HG.sub.5 and LH--HG.sub.6 in FIG.
11. As was previously indicated, the mixture zone MZ.sub.1 is also
basically an annular or cylindrical shaped mixture zone defined by
the space between the perimeter of the imperforate center plate 25
and intercoupled segmented plates 27.sub.1s and receives
hydrocarbon feed stream (i.e. hydrocarbon liquid feed and/or
hydrogen gas) in and through the space by which it is being
defined. In a preferred embodiment of the present invention and as
best shown in FIG. 13, before the flowing LH--HG mixture enters
into each of the generally continuous annular mixture zones
MZ.sub.1s, MZ.sub.2s, MZ.sub.3s, etc. the LH--HG mixture flows
around the plurality of inert pellets 4 in zig-zag fashions for
reducing the possibility of eddy currents and for keeping bubbles
of hydrogen gas diffused within the liquid hydrocarbon and
preventing agglomeration of same into larger size bubbles. The
hydrocarbon feed stream entering into mixture zone MZ.sub.1 is
designated LH--HG.sub.1. The plurality of LH--HG mixtures (i.e.
LH--HG.sub.1, LH--HG.sub.2, etc.) pass through the screen 28 and
respectively enter into the catalyst bed 10 from each of the
mixture zones (i.e. MZ.sub.1s, MZ.sub.2s, MZ.sub.3s, etc.) at a
flow rate such as not to ebullate, levitate or expand the catalyst
bed 10 upwardly and/or towards the screen 15 and the domed head 14
by more than 10% by length beyond substantially the full axial
length of the bed catalyst 10 in a packed bed state, such as the
packed bed state reflected in FIG. 8. The plurality of generally
continuous annular LH--HG mixtures flow upwardly through screen 28
and into the catalyst bed 10. The catalyst bed 10 in the present
invention preferably comprises catalyst particles which are
substantially the same and/or uniform size, shape, and density and
which are selected in accordance with the average optimum velocity
of the hydrocarbon feed stream (i.e. a mixture of a liquid
hydrocarbon component LH and a hydrogen-containing gas component
HG, or the continuous annular LH--HG mixtures) flowing into the
plenum chamber 33 and subsequently into and through the plurality
of mixture zones MZ.sub.2s, MZ.sub.3s, etc. The rates of flow of
the plurality of the respective LH--HG mixtures (i.e. LH--HG.sub.1,
LH--HG.sub.2, etc.) from the respective mixture zones MZ.sub.1s,
MZ.sub.2s, etc., and thus also the flow rates of the hydrocarbon
feed stream into plenum chamber 33 from and through line 16, are
all to be controlled in an amount and to an extent sufficient to
maintain expansion or levitation of the catalyst bed 10 to less
than 10% by length over or beyond substantially the full axial
length of the bed 10 in a packed bed state. More preferably, the
expansion of the substantially packed bed of catalyst is limited to
less than 5%, most preferably less than 2% or even less than 1%, by
length over or beyond substantially the full axial length of the
bed 10 in a packed bed state. Ideally the expansion of the
substantially packed bed of catalyst is limited to essentially 0%
by length.
[0074] The flow rate of the hydrocarbon feed stream through line 16
is to be at a rate not substantially greater than the optimum rate
of flow. The optimum rate of process fluid flow through the
substantially packed bed of catalyst will vary from process unit to
process unit based on several factors including oil and hydrogen
feed characteristics, catalyst specifications, process objectives,
etc. Based on catalyst particles having substantially the same
and/or uniform size, shape and density, the flow rate of the
hydrocarbon feed stream preferably ranges from about 0.01 ft/sec to
about 10.00 ft/sec and more preferably from about 0.01 ft/sec to
about 1.00 ft/sec. Similarly and/or likewise and further based on
the catalyst particles having substantially the same and/or uniform
size, shape, and density, the flow rate of the continuous annular
LH--HG mixtures (i.e. the summation of the flow rates for
LH--HG.sub.1 through LH--HG.sub.6 from mixture zones MZ.sub.1s
through MZ.sub.2s respectively in FIG. 11) is to be at a rate not
substantially greater than the optimum rate of flow, preferably
ranging from about 0.01 ft/sec to about 10.00 ft/sec, and more
preferably from about 0.01 ft/sec to about 1.00 ft/sec. The
specific flow rate would depend as indicated on a number of
variables, such as the particular application (e.g. demetallation
or desulfurization etc.) of the hydroprocessing process. The
specific flow rates however would be at any suitable rate
controlled in an amount and to an extent sufficient to limit
expansion of the substantially packed bed of catalyst to less than
10% by length over or beyond a substantially packed bed of
hydroprocessing catalyst in a packed bed state.
[0075] In a preferred embodiment of the invention and for such a
flow rate for the hydrocarbon feed stream and for such a flow rate
for the continuous annular LH--HG mixtures, the catalyst particles
preferably have the substantially same and/or uniform size, shape
and density in order to obtain over the desired demetallization
and/or desulfurization and/or denitrification of the liquid
hydrocarbon component LH in the hydrocarbon feed stream (i.e.
LH--HC mixture) into produced hydrogen upgraded product fluids that
are being withdrawn from the reactor vessel 11 through the center
pipe 18. At the above indicated flow rates for the hydrocarbon feed
stream flowing through line 16, and for the flow rates for the
generally continuous annular LH--HG mixtures (i.e. LH--HG.sub.1,
LH--HG.sub.2, etc.), the produced upgraded product fluids are being
preferably withdrawn through the center pipe 18 from the reactor
vessel 11 at a rate ranging from about 0.01 ft/sec to about 10.00
ft/sec and more preferably from about 0.01 ft/sec to about 1.00
ft/sec. The withdrawal rate(s) of the produced upgraded product
fluids is not to be greater than the optimum rate of flow and will
also vary from process unit to process unit based on several
factors including oil and hydrogen feed characteristics, catalyst
specifications, process objectives, etc. The specific withdrawal
rate(s) would be any suitable withdrawal rate, controlled in an
amount and to an extent sufficient to prevent and/or limit
expansion of the substantially packed bed of catalyst to less than
10% (more preferably less than 5%, most preferably less than 2% or
even less than 1%) by length over or beyond substantially the full
axial length of the bed 10 in a packed bed state.
[0076] The arrangement in inlet distributor 31 for uniformly
distributing hydrogen gas and liquid hydrocarbon feed as shown in
FIG. 3 may be modified by lengthening or shortening tubes 32,
forming uniformly distributed cylindrical passageways into plenum
chamber 33. A particular advantage of using tubes 32, as compared
to merely perforations or holes of adequate diameter, lies in the
formation of a gas pocket under plate 31 in the areas around the
individual tubes 32. We have found that this is desirable because
such a gas pocket trapped beneath tray or plate 31 provides
pressure surge dampening, which may result from flow changes of the
mixture of hydrogen and liquid being supplied to the reactor
vessel. However, the length of the tubes 32 is maintained as short
as reasonably possible to so function. Again, this is because of
the desirability of utilizing as little as possible of all
processing space available in vessel 11 for anything but contacting
the feed streams with conversion catalyst. A particular advantage
to using tubes 32, as compared to a combination of tubes and
perforations, is that the designed flow distribution pattern is
maintained over a wider range of flow rates. With tubes and
perforations, gas normally flows up the perforations and liquid
flows up the tubes. However, gas will find new flow paths through
the tubes if the gas flow increases or the perforations become
plugged, resulting in undesigned and potentially undesirable flow
patterns.
[0077] To further assist in maintenance of plug-like flow of
catalyst bed 10 throughout its axial length, there is additionally
provided in a preferred form or embodiment of the invention a
plurality of axially spaced apart hydrogen gas redistribution or
hydrogen gas-quenching stages 39 within bed 10. In the arrangement
of FIG. 1, the location of one of the gas redistribution stages 39
is illustrated by the single inverted angle member 40 extending
transverse to the axis of bed 10. The details of quench system 39
are best seen in FIGS. 5 to 7 where a plurality of inverted
V-shaped sheds 40 (i.e. inverted angle members 40) are equally
distributed over at least one transverse row extending generally
across the cross-sectional area of vessel 11. As shown in FIG. 6
and in FIG. 7, a gas injection line 42 feeds an elongated tube 41
extending through each individual shed 40 from a header 44 and
branch lines 45 supplying the individual tubes 41. Desirably, but
not necessarily, a second tier of sheds 40 is axially spaced above
the first tier, with the sheds 40 in each tier being positioned at
90 degree(s) to the other tier, as shown in FIG. 7. Construction of
an individual shed 40 is best seen in FIG. 6, wherein distribution
pipe 41 includes a plurality of discharge holes 48, desirably
proportioned to give equal distribution of hydrogen gas along the
full length of tube 41. Desirably, holes 48 are on the top side of
tube 41 so that gas leaving the tube is forced to flow downwardly
within shed 40 to join gas rising from bed 10 under the area
enclosed by the V-sides 49 of shed 40. Preferably, the full length
of each skirt formed by sides 49 includes equally spaced slots 50
to exhaust both rising gas from bed 10 and quench gas entering from
line 42. A particular value of the present arrangement is that gas
which may have become channeled in a portion of the bed below the
quench system can be redistributed across the full cross-sectional
area of the bed to further avoid generation of local hot spots,
eddy currents, or ebullation, within the upper portion of bed
10.
[0078] In accordance with another significant aspect of the present
invention, FIG. 1 shows a catalyst replacement system, which in
general comprises a series of lock chambers for transferring fresh
catalyst into bed 10 through a pair of pressure lock chambers,
including charging vessel 60 and supply vessel 70. A similar series
of lock chambers, including discharge vessel 80 and disposal vessel
90, transfer catalyst out of bed 10. If necessary, a single pair of
vessels could be used to charge and discharge the catalyst,
although the piping and sequencing procedure would be more complex.
In both cases, transfer flow is specifically designed to be as a
liquid slurry and laminar to avoid undue abrasion of catalyst
particles going into reactor vessel 11 and to avoid abrupt
agitation of the overlying bed of particles, with consequent
ebullation and eddying of catalyst or fines in bed 10, when
catalyst is withdrawn through inlet 30 of J-tube 29 at the bottom
of reactor vessel 11.
[0079] To achieve laminar flow for supply of catalyst from charging
vessel 60 to the top of reactor vessel 11 or for catalyst removal
from the bottom of bed 10 to discharge vessel 80, it is essential
that the pressure differential between reactor vessel 11 and
vessels 60 or 80, be accurately controlled as by detecting the
pressure differences between supply line 61 or discharge line 82
and reactor vessel 11. The pressure difference is best zero when
shut-off valves 64 or 84 are first opened or closed. The pressure
differences between vessel 11 and line 61 is measured by gage 63
and pressure detectors 62 and 65. Differential pressure gage 83 and
detectors 81 and 85 serve a similar function to control transfer of
catalyst through valve 84 from the bottom of reactor vessel 11 to
discharge vessel 80.
[0080] With reference particularly to supply of catalyst from
vessel 60, it will be understood, of course, that the vessel 60 is
capable of being brought to a slightly higher pressure than the
operating pressure of reactor vessel 11, and closely controlled to
assure that catalyst supplied to vessel 60 from supply vessel 70 is
by laminar flow. For this purpose, as indicated, vessels 70 and 60
are at atmospheric pressure, catalyst is first introduced into
supply vessel 70 by way of funnel 100 through line 101 and valve
102, and nitrogen is preferably flushed through supply vessel 70
through line 106 and/or line 71 to eliminate air and moisture that
may be present on the catalyst. Either before or after catalyst is
introduced, vessel 70 is charged with a distillate hydrocarbon
stream, preferably gas oil, to provide the necessary slurrying
liquid for mixing and transporting catalyst. This may either be
through funnel 100, valve 102, and line 101, or through line 104,
valve 105 and line 106. Valve 102 is then closed and the catalyst
is then preferably heated to dehydrate and eliminate water from the
catalyst. It is to be understood that whenever the specification or
the claims states, mentions, or implies "mixing" or "admixing" or
"commingling", or any of the like, including of any type(s) of
catalyst, such stated, mentioned, or implied verbiage means within
the spirit and scope of the present invention any type of "mixing"
or "admixing" or "commingling", or any of the like, including any
incidental mixing or any otherwise non-thorough/non-homogeneous
mixing. Preferably, however, any type of "mixing" or "admixing" or
commingling", or any of the like, will be essentially thorough
and/or essentially homogeneous.
[0081] An important requirement is that before transferring liquid
to the charging vessel 60, the pressure in supply vessel 70 must be
equalized to that in charging vessel 60, assuming, of course, that
isolation valve 64 between vessel 60 and the reactor vessel 11 is
closed, and also that valves 67, 68 and 78 are closed. With valves
64, 67, 68, 78 and 102 closed and pressure equalized between the
vessels 60 and 70, transfer valve 75 may be opened to provide the
same diameter path for the catalyst slurry to flow throughout the
path from J-tube 71 to vessel 60. The transfer is closely
controlled by regulating the nitrogen gas flow rate and pressure
introduced from line 104 through valve 105. The pressure and flow
rate are just sufficient to assure the desired laminar flow of
catalyst into inlet 72 of J-tube 71 and thus upwardly through line
76 and into charging vessel 60, which forms a catalyst charging
vessel. Laminar flow to transfer catalyst through J-tube 71 is
entirely in the liquid phase, with the catalyst as a slurry in the
gas oil. Transfer of all catalyst is assisted by the funnel shape
of bottom 79 of vessel 70, and the position of intake 72 to J-tube
71 at the apex of bottom 79. If all the catalyst in vessel 70 is
transferred to vessel 60, flush oil from vessel 70 will naturally
clear all the catalyst out of line 76. However, to assure that all
such catalyst has passed through valve 75 (so that valve 75 need
not close on hard, abrasive catalyst with potential danger of
scoring the valve 75 or the valve seat therein) additional flush
fluid is preferably introduced from line 77 through valve 78 to
clear line 76, either back into vessel 70, or forward into vessel
60.
[0082] With catalyst thus loaded into vessel 60, a similar
procedure is used for transferring catalyst under laminar flow
conditions as a liquid slurry into reactor vessel 11 through supply
pipe 61 for distribution to the top 20 of bed 10. If desired, of
course, a deflector plate (not shown) may be used to distribute
catalyst evenly across the top of catalyst bed 20. However, we have
found that such a distribution aid is not required. In the transfer
of catalyst from the charging vessel 60 to reactor vessel 11, it
will be understood that the pressure in vessel 60 is brought to the
pressure of reactor vessel 11. This is done by injecting process
hydrogen through valve 67. The oil should be heated to a
temperature as close as possible to the temperature of reactants in
vessel 11, without vaporizing the oil. We have found this to be
particularly important to minimize any disturbance of the
hydroprocessing process when fresh catalyst is added to an onstream
reactor vessel, such as reactor vessel 11. Once these requirements
are met, valve 64 should be opened for transfer. The actual laminar
transfer of the liquid slurry is controlled by valve 67 throttling
the flow and pressure of hydrogen admitted from line 66. After
transfer of the catalyst, valve 68 in flush line 69 is opened
briefly to assure that any catalyst left in lines 61 and 19 is
cleared before valve 64 is closed, for the reasons noted before.
Excess hydrogen pressure in vessel 60 may be relieved in a
controlled manner via a suitable bleed line running back to the
common hydrogen source (not shown) of the hydroprocessing
system.
[0083] Substantially continuous or intermittent transfer of
deactivated catalyst for regeneration or disposal from the bottom
of catalyst bed 10 in reactor vessel 11 to discharge vessel 80 is
controlled in the same manner. As in all transfer of catalyst
throughout the system of the present invention depicted in FIG. 1,
the flow path from inlet 30 of J-tube 29, through line 82,
including the bore of valve 84, is uniform in cross-sectional area
and diameter. Similarly, transfer from discharge vessel 80 to
disposal vessel 90 is through inlet 89 of J-tube 86 to discharge
outlet 98 of line 92, including valve 94, into vessel 90.
Deactivated catalyst is transferred laminarly from the bottom of
the catalyst bed 10 as a slurry in the hydrocarbon feed stream
which, as previously mentioned, comprises the liquid hydrocarbon
feed stream or a mixture of hydrocarbon liquid feed and
hydrogen-containing gas. Typically, the catalyst is transferred
essentially in the liquid hydrocarbon feed stream (i.e. the liquid
component of the hydrocarbon feed stream).
[0084] In general the diameter of these laminar flow passageways
are at least five times, and may be as high as fifty or more times,
the diameter of the individual particles to be passed therethrough.
In this connection to avoid jamming or obstruction, the inlets 72,
109, 30, 89 and 99 into their respective tubes 71, 108, 29, 86 and
96 are not flared or otherwise restricted, or perforated, so that
all flow is solely and directly through the full and equal bore of
such inlets. In the case of catalyst removal from reactor vessel
11, inlet 30 of tube 29 is positioned at and over unperforated
center plate 25 of catalyst support screen means 17, so that it is
out of the direct flow of any hydrogen gas stream rising through
the innermost annular passageway formed by walls 27 and radial ribs
or spokes 26. This assures that flow into entry 30 is substantially
a liquid only slurry mixture with catalyst particles. Such a
mixture at laminar flow conditions produces maximum carrying
capacity of the fluid. Additionally, the external dimensions of the
circular bend or arc portion of the J-section of the tube 29 is
several times the diameter of inlet 30 and the connected flow path,
including the downwardly directed portion. The portion of tube 29
above inlet 30 is many times shorter and smaller in volume than the
remainder of J-tube 29, down to, and including, control valve 84. A
particular advantage of keeping this portion of tube 29 small is to
avoid the necessity of forcing substantial amounts of catalyst back
into the bed 11 against the gravity head of catalyst bed 10 when
that portion of the line is cleared at the end of each
transfer.
[0085] Desirably, during periods when the catalyst is not being
transferred, a small amount of hydrogen may be continually bled
through valve 88 into bed 10 through J-tube 29 to assure that
catalyst particles do not clog entry 30. This avoids potential
build up of coke at entry 30 of pipe 29. Such an arrangement
assures that catalyst can be withdrawn by laminar flow without
artificially fluidizing or levitating bed 10 directly adjacent to
J-tube entry 30.
[0086] Because gravity drainage of catalyst by an opening through
the center of the catalyst support screen means 17 is not required
in the present arrangement, as in the prior art, it is possible to
operate the entire system without use of solids handling valves.
Accordingly, each of the transfer valves in the present arrangement
are preferably conventional ball valves formed with a single
through bore in a rotatable ball. Specifically, we have found that
conventional valves used to feed and control flow of hydrocarbons,
catalyst and hydrogen, into and out of the vessel 11, must seal
against high pressure differentials between the vessel and the
transfer vessels. For this service, a solid satellite,
spherical-ball valve having a through bore of the same diameter as
the inlet and outlet lines to the valve and metal-to-metal seals,
provides superior service when used in the catalyst transfer lines
for carrying out the method of the present invention. Further,
their commercial cost and ready availability for such severity of
service makes them most useful economically, both for initial
installation and for service replacement. Valves manufactured by
The Kaymr and Mogas Companies, called full-port valves are
particularly useful in the present embodiment. Further, the
arrangement permits transfer of catalyst almost exclusively in a
liquid phase which substantially reduces abrasion or comminution of
catalyst particles during transfer. Additionally, exclusion of
entrained gas substantially improves the efficiency of liquid
transfer of catalyst particles and further reduces potential damage
to the catalyst.
[0087] FIG. 4 illustrates a partial view of the bottom of pyramidal
catalyst bed support means 17 showing an alternate system for
transferring catalyst in a laminarly flowing liquid. In this
embodiment, an L-valve is formed by vertical tube 54 and horizontal
tube 52 for withdrawing catalyst particles from the bottom of bed
10. As shown, intake 56 is preferably directly above the central,
non-perforated, section 25 of the truncated pyramid formed by
support means 17. While such an arrangement is less preferred than
that shown in the embodiment of FIG. 1, such an arrangement is made
suitable by the fact that the slurry of liquid and catalyst can be
made to flow only under laminar flow conditions. With either the
J-tube of FIG. 1, or the L-valve of FIG. 4, arrangements, the
pressure in discharge vessel 80 is brought up to equal that in
reactor vessel 11. Valve 84 is opened and catalyst flow is
controlled, as seen in FIG. 1, by regulating flow through valve 93.
Such flow decreases the gas-pressure in discharge vessel 80 and
line 82 sufficiently to induce a laminar flow of catalyst particles
from vessel 11 when transfer valve 84 is opened. After valve 84 has
been flushed with vacuum gas oil through valve 88 and line 87 and
then closed. The pressure in vessel 80 is then reduced to a lower
pressure (about 50 psig or less). The residuum is drained from
discharge vessel 80 through drain line 120, below J-tube 86 and
conical screen 121. Flush oil is then sent in through valve 93 to
wash residuum off the catalyst and to cool the catalyst. The
discharge vessel 80 can be drained and filled as many times as
needed. The pressure in disposal vessel 90 is made equal to that in
vessel 80 and valve 94 is opened. The flow and pressure are then
controlled through valve 110 to induce laminar flow of catalyst
through J-tube 86 and into disposal vessel 90. Valve 94 is flushed
with flush oil through valve 107 and closed. The contents of the
disposal vessel 90 is preferably washed and cooled with flush oil
which is then drained through drain line 122 below conical screen
123. The spent catalyst contents of the disposal vessel 90 is then
washed with water if desired through valve 110. The disposal vessel
90 should be purged of any hydrogen by sending in nitrogen gas also
through valve 110. Finally, disposal vessel 90 is nearly
depressurized and the catalyst is dumped using water as the carrier
fluid through J-tube 96 by nitrogen flow through valve 110 to
control the rate of catalyst flow in discharge pipe 124.
[0088] Continuing to refer to the drawings for other preferred
embodiments of the present invention, a method is provided for
maximally occupying a reactor volume with a substantially packed
bed of hydroprocessing catalyst (e.g. catalyst bed 10) during
hydroprocessing by contacting the substantially packed bed of
hydroprocessing catalyst with an upflowing hydrocarbon feed stream
having a liquid component and a hydrogen-containing gas component.
The substantially packed bed of hydroprocessing catalyst,
preferably comprising the catalyst of the present invention as more
particularly described below under the subtitle "The Catalyst", is
disposed in a reactor zone (or reaction zone or zone for reaction)
contained within a reactor volume (e.g. the entire internal
volumetric space available within the reactor vessel 11) such that
the substantially packed bed of hydroprocessing catalyst occupies
from about 50% by volume to about 98% by volume of the reactor
volume; more preferably from about 50% by volume to about 98% by
volume; most preferably at least about 90% by volume or from about
90% by volume to about 95% by volume of the reactor volume. Stated
alternatively, hydroprocessing catalyst is disposed or otherwise
positioned within a reactor volume such that hydroprocessing
catalyst occupies from about 50% by volume to about 98% by volume
of the reactor volume; more preferably from about 50% by volume to
about 98% by volume; most preferably at least about 90% by volume
or from about 90% by volume to about 95% by volume of the reactor
volume. "Reactor volume" (or the entire internal volumetric space
available within the reactor vessel 11) means or may be generally
defined as the volumetric space within the reactor vessel 11 (or
any similar hydroprocessing reactor vessel), including the
summation or addition of the following internal volumes: (i) an
internal volume within the reactor vessel 11 represented by a
volume (or internal cylindrical volume or main body volume of the
reactor vessel 11) spanning or extending from an upper tangent line
180 (see FIG. 8) to a lower tangent line 182 and generally
illustrated as arrow TL in FIG. 8; and (ii) an internal volume
within the upper dome closure end 14 (or hemispherical head) of the
reactor vessel 11; and (iii) an internal volume within the lower
dome closure end 13 (or hemispherical bottom) of the reactor vessel
11. A "tangent line" is known to those skilled in the art as a
plane (i.e. horizontal plane) taken generally along the junctures
of the sidewall 12 (which is essentially a straight upright wall)
of the reactor vessel 11 with the upper and lower dome closure ends
14 and 13 respectively.
[0089] A hydroprocessing feed stream including a liquid component
and a hydrogen-containing gas component upflows into the
substantially packed bed of hydroprocessing catalyst at a rate of
flow such that expansion of the substantially packed bed of
hydroprocessing catalyst is limited to less than 10% by length
beyond a substantially full axial length of the substantially
packed bed of hydroprocessing catalyst in a packed bed state. A
volume of the hydroprocessing catalyst is withdrawn from the
reactor zone to commence essentially plug-flowing downwardly of the
substantially packed bed of hydroprocessing catalyst within the
reactor zone; and hydroprocessing replacement catalyst is added to
the essentially plug-flowing downwardly, substantially packed bed
of hydroprocessing catalyst at a rate to substantially replace the
volume of the withdrawn hydroprocessing catalyst. The procedure may
be repeated as many times as desired, even continuously repeated
during continual hydroprocessing.
[0090] Another method is provided for hydroprocessing a hydrocarbon
feed stream that is upflowing through a hydroconversion reaction
zone having a substantially packed bed of catalyst which comprises
forming a plurality of annular mixture zones under a
hydroconversion reaction zone having a substantially packed bed of
hydroprocessing catalyst such that each of the annular mixture
zones contains a hydrocarbon feed stream having a liquid component
and a hydrogen-containing gas component and wherein the annular
mixture zones are concentric with respect to each other and are
coaxial with respect to the hydroconversion reaction zone. The
hydrocarbon feed stream from each of the annular mixture zones is
introduced into the substantially packed bed of hydroprocessing
catalyst to commence upflowing of the hydrocarbon feed stream from
each of the annular mixture zones through the substantially packed
bed of the catalyst.
[0091] Considering the range of hydroconversion systems and/or
hydroconversion reaction zones which could benefit from the
preferred embodiments of the present invention, one skilled in the
art will appreciate the variety of catalysts, having a variety of
physical properties and elemental compositions, which could be used
in such a range of systems. It is within the spirit and scope of
the present invention to encompass these systems employing
catalysts having a size, shape and density which vary widely from
system to system. However, it is important for the present
preferred embodiment that the catalyst particles be of uniform
and/or same size, and shape (same density when in fresh catalyst
state) within a single hydroconversion reaction zone of a
hydroconversion system, in order to achieve the desired catalyst
and hydrocarbon flow patterns within the hydroconversion reaction
zone. A detailed description of the preferred catalyst is presented
below under the subtitle "THE CATALYST". It is to be understood
that whenever the specification or the claims states, mentions, or
implies "fresh catalyst", such stated, mentioned, or implied "fresh
catalyst" means within the spirit and scope of the present
invention any type of catalyst having any usable catalyst life or
activity (e.g. regenerated catalyst, rejuvenated catalyst,
partially fouled catalyst obtained from any source, etc.).
Preferably, "fresh catalyst" means a type of catalyst that has
never been used before and is obtained directly from a manufacturer
with the lowest desired density and the highest desired catalyst
life or activity.
[0092] A hydroconversion system and/or a hydroconversion reaction
zone of a present preferred embodiment of the present invention
contains a catalyst which is described in detail below under the
subtitle "The Catalyst", and may also be operated as a fixed bed
(i.e. a catalyst bed which does not expand), a moving bed, an
ebullated bed, an expanded bed or a fluidized bed configuration. A
moving bed system is preferred.
[0093] By "moving bed", as used herein, is meant a reaction zone
configuration in which a catalyst is added at one end of a catalyst
bed in an intermittent or substantially continuous manner and is
withdrawn at the other end in an intermittent or substantially
continuous manner. A "moving bed" also includes a "plug-flow" or
"plug flowing" catalyst bed 10 or substantially packed bed of
catalyst. As previously indicated, when any type of catalyst or
catalyst bed 10 movement is mentioned, stated, or implied, the
spirit and scope of the present invention includes such type of
movement on any type of basis or in any manner (e.g. "periodic",
"fully continuous", "non-continuous" etc.) without the necessity of
having to specifically mention the type of basis or manner.
Preferably, catalyst is added at the top of the reaction zone and
withdrawn at the bottom. In the type of moving bed to which the
present preferred embodiment is directed, the catalyst particles in
the bed are substantially in contact with one another and plug-flow
downwardly. The catalyst bed is not significantly expanded when
process fluids (e.g., liquid and gas) passes through it. It has
essentially the character of a fixed bed except for maybe a slight
expansion upwardly and for the addition and removal of catalyst. As
the term is used herein, a "moving bed" is not the same as a
"fluidized bed", "ebullating bed" or "expanded bed". In fluidized
beds, the flow rate of a single phase fluid, relative to the
particles of the catalyst, is fast enough so that the catalyst
behaves like a fluid, with particles circulating throughout the bed
or even being carried out of the bed with the products. Ebullating
or expanded beds are very similar to fluidized beds, except that
the relative rate of flow of two phase fluids (e.g., liquid and
gas) is regulated to expand the catalyst bed in random motion
between 110% and 140% of the height of the catalyst in a "slumped"
or packed state. The typical ebullating bed reactor will have a
mass of solid particles whose gross volume in the reaction vessel
is at least 10 percent larger when feed is flowing through it, as
compared to the stationary mass with no feed flowing through it.
Although the particles in the bed do not necessarily circulate as
if they were fluids, they are separated from one another and go
through random motion.
[0094] Several advantages ensue from use of a moving bed reactor.
By establishing and maintaining appropriate gas and liquid
velocities in packed bed type reactors, just below the threshold of
inertia that would cause the catalyst bed to fluidize and/or
"channel" and/or lift the catalyst into random motion, the uniform
catalyst characteristics described above will allow the catalyst to
migrate downward through the reactors in a predictable plugflow
manner, as catalyst batches are withdrawn from the reactor bottom.
And further, by maintaining plug flow catalyst movement downward
within the reactors (e.g. reactor 11), the catalyst within the
reactors can be maintained in layers having differing activity
levels and reaction rates. The number of catalyst layers depend on
the frequency of catalyst addition and withdrawal, and the amount
added and withdrawn in any given period of time. Typically,
however, the number of different aged catalyst layers within the
reactor (e.g. reactor 11) will be in the range from 10 to 60.
[0095] Intermittent or continuous catalyst additions and withdrawal
may be used. Catalyst replacement rates can range from several
percent of the charge per day to several percent of the charge per
week, depending on the reactor size, catalyst metals loading
capacity, feed rate, and feed composition and processing
objectives. Fresh catalyst is introduced into the downstream end of
the catalyst bed (e.g. catalyst bed 10), and a corresponding volume
of deactivated catalyst is removed from the upstream end of the
catalyst bed, at a rate which is sufficient to maintain the actual
overall average level of catalytic upgrading activity of the bed as
a whole at or above the selected minimum average activity level. By
"upstream" end of the catalyst bed (e.g. catalyst bed 10), as used
herein, is meant the end of the moving bed into which the heavy
hydrocarbonaceous feed is introduced. By "downstream" end of the
catalyst bed is meant the end of the bed from which the process
effluent is recovered. In a normal gravity flow system, the
catalyst is added and effluent removed at the top of the vessel
(the downstream end). Spent catalyst is withdrawn and feed
introduced at the bottom (the upstream end).
[0096] In a particularly important application of the present
invention, catalyst is continuously added at the top of the reactor
(e.g. reactor 11) to the slowly moving bed (e.g. bed 10), and spent
(and deactivated catalyst) catalyst is continuously withdrawn from
bottom of the slowly moving bed. The deactivated catalyst is
removed from the reactor (e.g. reactor 11) after it has been
deactivated to a substantially lower level of activity than an
acceptable minimum average level of activity of the overall
catalyst bed. This allows more efficient and complete use of the
catalyst activity, e.g. its metals capacity, for such feed
upgrading functions as demetallation. As previously indicated,
spent (and deactivated) catalyst is withdrawn from the bottom of a
reactor in a hydrocarbon liquid. One of the features of the present
invention is that the hydrocarbon liquid that is withdrawing and
transporting catalyst is the liquid hydrocarbon component LH which
is intended to flow upwardly through the bed of catalyst but has
not. Thus, one of the features of the present invention is that the
hydrocarbon liquid for transporting spent (and deactivated)
catalyst is an unconverted liquid hydrocarbon component LH or a
partially converted liquid hydrocarbon component LH or a mixture of
both; and the transporting hydrocarbon liquid (i.e. the liquid
hydrocarbon component LH) has not passed entirely upwardly through
the catalyst bed.
[0097] The product from the method of the present invention exits a
reactor (e.g. reactor vessel 11) and is normally subjected to
further conventional refinery processing. All or part of the
product can be passed to a conventional, fixed bed upgrading
operation, such as a hydrodesulfurization operation. Part of the
product stream can be recycled, either for further catalytic
treatment or as a diluent. Treatment of heavy feeds by catalytic
demetallation according to the present process followed by fixed
bed desulfurization is particularly effective, but all or part of a
demetallized product from the countercurrent demetallation reaction
zone can also be processed in a countercurrent moving bed
desulfurization reaction zone.
[0098] The present preferred embodiments of the present invention
are applicable to hydroconversion reaction zones for hydrocracking,
hydrodemetallization, hydrotreating, hydrodesulfurization,
hydrodenitrification, hydrofinishing and the like, all of which
catalytically upgrade a heavy hydrocarbonaceous oil that represents
the liquid hydrocarbon stream or liquid hydrocarbon feed stream
(i.e. the liquid hydrocarbon component LH). By "heavy" liquid
hydrocarbon stream, as used herein and as previously indicated, is
meant liquid hydrocarbon stream at least 50 volume percent of which
boils above about 2040 C and which preferably contains a
substantial fraction boiling above about 343.degree. C. and
particularly preferably above about 510.degree. C. Preferred liquid
hydrocarbon streams are residual fractions and synthetic crudes.
They can be derived from crude petroleum, from coal, from oil
shale, from tar sand bitumen, from heavy tar oils, and from other
synthetic sources. The present invention is advantageously employed
to refine highly refractory and contaminated liquid hydrocarbon
streams. The liquid hydrocarbon stream may be substantially free
from finely divided solids such as shale fines, sand or the like.
Alternatively, the liquid hydrocarbon stream may contain a
substantial concentration (e.g. about 1 weight percent or more) of
finely divided solids. As previously indicated, the liquid
hydrocarbon stream (i.e. the liquid hydrocarbon component LH) is
preferably premixed with any type of hydrogen-containing gas (i.e.
the liquid hydrocarbon component HG) which is preferably hydrogen,
before being introduced into the reactor vessel 11 as a single
stream or hydrocarbon stream. The mixing ratios of the liquid
hydrocarbon stream to the hydrocarbon containing gas may be any
suitable ratio, well known to those artisans possessing the
ordinary skill in the art.
[0099] Typically, a heavy hydrocarbonaceous oil employed as a
hydrocarbon feed stream in the present invention contains
undesirable metals. Undesirable metals which are often present in
hydrocarbonaceous feeds notably include nickel, vanadium, arsenic,
and iron. These metals deactivate conventional, fixed bed catalysts
(such as fixed bed hydroprocessing catalysts) and also rapidly and
irreversibly deactivate catalysts when high metals level feed are
charged directly to conventional units. These metals are often
present as organo-metallic compounds. Thus, the use of the
terminology "iron, nickel, arsenic or vanadium compounds" means,
those metals in any state in which they may be present in the
hydrocarbon feed stream in the process of the present invention,
either as metal particles, inorganic metal compounds, or an
organo-metallic compounds. Where amounts of metals are referred to
herein, the amounts are given by weight based on the metal itself.
For maximum efficiency in such a countercurrent demetallation
process, the hydrocarbon feed stream should have levels of
undesirable metals greater than about 150 ppm by weight of the
hydrocarbon feed stream, preferably greater than about 200 ppm by
weight of the hydrocarbon feed stream, and more preferably greater
than about 400 ppm by weight. Although nickel, vanadium, arsenic,
and iron are the usual metal contaminants, other undesired metals,
such as sodium and calcium, may also contribute to the metals
content of the hydrocarbon feed stream for purposes of catalytic
demetallation upgrading processing.
[0100] Catalytic upgrading conditions (e.g. catalytic
desulfurization conditions, catalytic hydrogenation conditions such
as designed for asphaltenes saturation, catalytic denitrification
conditions and catalytic hydrocracking conditions, etc.) employed
in the hydroconversions reaction zones within the reactor vessel 11
for preferred embodiments of the present invention all include a
reaction temperature generally in the range of from about
230.degree. C. to about 480.degree. C., a pressure generally in the
range of from about 30 to about 300 atmospheres, a hydrogen rate
ranging from about 1000 to about 10,000 standard cubic feet per
barrel of feed, and a liquid hourly space velocity (LHSV) in the
range of from about 0.20 hr-1 to about 10 hr-1. For feed
demetallation upgrading, the temperatures and pressures within the
reaction zone can be those typical for conventional demetallation
processing. The pressure is typically above about 500 psig (514.7
psia; 35.5 bar). The temperature is typically greater than about
315.degree. C., and preferably above 371.degree. C. Generally, the
higher the temperature, the faster the metals are removed; but the
higher the temperature, the less efficiently the metals loading
capacity of the demetallation catalyst is used. While demetallation
reaction can be conducted in the absence of added hydrogen,
hydrogen is generally used and therefore requires full and equal
distribution into the moving bed along with any gases evolving from
the feed.
[0101] In carrying out a process of the preferred embodiments of
the present invention, a minimum average level of catalytic feed
upgrading activity for the countercurrently moving catalyst bed
(e.g. catalyst bed 10) as a whole is selected for the particular
catalytic upgrading reaction. For a moving bed (e.g. catalyst bed
10) in a demetallation reaction system, for example, the minimum
average upgrading activity level for the catalyst bed is one which
removes the necessary amount of metals from the hydrocarbon feed
stream when it passes through the moving bed at demetallation
conditions. Similarly, for a desulfurization reaction system, the
moving catalyst bed (e.g. catalyst bed 10) removes the necessary
amount of sulfur from the hydrocarbon feed stream when it passes
through the moving bed at desulfurization conditions. Thus, as will
be apparent to those skilled artisans, the minimum average
upgrading activity level for a particular reaction system will
depend on the desired degree of a contaminant, such as metals,
sulfur, nitrogen, asphaltenes, etc., which the refiner desires to
remove from the heavy oil feed. The degree of demetallation or
desulfurization (or etc.) will typically be set by economics and
the downstream processing that the heavy feed will undergo.
Further, according to preferred embodiments of the present
invention, the actual average level of catalytic upgrading activity
for the catalyst bed (e.g. catalyst bed 10) as a whole is measured.
Measurement of the actual average level of upgrading is made by
determining the extent to which the hydrocarbon feed stream is
being upgraded in the countercurrent moving bed system. For
example, when upgrading involves demetallation, demetallation
activity is measured by a determination of the residual
concentration of metals remaining in the liquid effluent stream
from the moving bed system. When upgrading involves
desulfurization, desulfurization activity is, analogously, measured
by a determination of the residual concentration of sulfur
remaining in the liquid effluent from the reaction system. Overall
catalyst bed upgrading activity for other upgrading reactions is
measured in a similar manner by determining the residual amount of
the containment which is to be removed by the process. In the
present preferred embodiments, the rate at which catalyst is
removed from the reaction zone, and the rate of catalyst
replacement to the reaction zone, is established by a number of
economic and operating factors, which include maintaining a desired
average level of catalytic upgrading activity.
[0102] Referring in detail now to FIGS. 14-16 for another
embodiment of the present invention and wherein similar parts of
the present invention depicted in FIGS. 14-16 are identified by
like reference numerals that depict the same similar parts of the
invention depicted in FIGS. 1-13, there is seen a pair of moving
catalyst bed 10-10 contained in the reactor vessel 11. Each moving
catalyst bed 10 is in communication with the pipes 19 wherein fresh
catalyst is added onto the bed surface 20. Each moving catalyst bed
10 is also in communication with the discharge line 82 wherethrough
at least partially spent catalyst is withdrawn. In FIG. 14, pipe 19
passes through domed head 14 for adding fresh catalyst to the top
moving catalyst bed 10. The center pipe 18 in FIG. 14 receives a
hydrocarbon feed which flows downwardly through the top or upper
moving catalyst bed 10, and subsequently downwardly through the
bottom moving catalyst bed 10 for discharge through line 16 which
is functioning as a discharge line instead of a feed line.
Similarly, center pipe 18 in FIG. 14 is functioning as a feed line
instead of a discharge line.
[0103] Referring particularly to FIGS. 15 and 16 center pipe 18 is
functioning as a discharge line wherethrough treated hydrocarbon
streams exit the reactor vessel 11. Pipe 19 passes through center
pipe 18 in a coaxial manner in order to deposit fresh catalyst upon
the bed surface 20 of the top catalyst bed 10. In FIG. 15 another
pipe 19 passes through the domed head 14, through the entire top
moving catalyst bed 10, through the catalyst support means 17, and
then through the circular plate member 31 (i.e. a distributor plate
assembly with depending hollow tubes 32) for depositing catalyst
upon the bed surface 20 of the bottom or lower moving catalyst bed
10. In FIG. 16 the second pipe 19 passes through the cylindrical
side wall 12 of the reactor 11.
[0104] The discharge line 82 has various embodiments. In FIG. 14
the discharge line passes through the bottom domed head 13 for
removing at least partially spent catalyst from the bottom or lower
moving catalyst bed 10. In FIGS. 15 and 16 the discharge line 82
passes generally coaxially into line 16 and through the circular
plate member 31 and through the imperforate center plate 25 of the
catalyst support means 17 for removing at least partially spent
catalyst from the bottom or lower moving catalyst bed 10. In FIG.
15 a second discharge line 82 passes through the bottom domed head
13, through the bottom or lower moving catalyst bed 10, and
terminates into the J-tube 29 after passing through the circular
plate member 31 and the catalyst support means 17 for removing at
least partially spent catalyst from the top or upper moving
catalyst bed 10. In FIG. 16 the discharge line 82 passes through
the cylindrical side wall 12, through the circular plate member 31,
and then through the imperforate center plate 25 of the catalyst
support means 17, which supports the top or upper moving catalyst
bed 10, for withdrawing at least partially spent catalyst from the
top moving catalyst bed 10. In FIG. 14 the discharge line 82 also
passes through the cylindrical side wall 12 for withdrawing
partially spent catalyst from the top moving catalyst bed 10.
[0105] FIG. 14 is a general schematic diagram of the present
invention, generally illustrating the reactor vessel 11 having two
separate hydroconversion reaction zones including a pair of the
moving catalyst beds 10-10. Each of the hydroconversion reaction
zones has an inlet (i.e. pipe 19) and an outlet (i.e. discharge
line 82) for respectively charging the respective moving catalyst
beds 10 with fresh catalyst and for respectively discharging at
least partially spent catalyst from the respective moving catalyst
beds 10. The hydrocarbon feed stream in FIG. 14 enters the reactor
vessel 11 through the center pipe 18 and treated hydrocarbon
product fluids exit the reactor vessel 11 through line 16. No
catalyst enters or leaves with the hydrocarbon feed stream or with
the hydrocarbon product stream; instead catalyst enters (i.e.
through pipe 19) and leaves (discharge line 82) through the
separate connections provided for the catalyst. Although FIG. 14
illustrates a cocurrent moving-bed, feed and catalyst downflow, the
present invention may also be used for a countercurrent moving-bed,
upflow feed and downflow catalyst, as best shown in FIGS. 15 and
16.
[0106] FIGS. 15 and 16 illustrate a more preferred embodiment of
the present invention. Both FIGS. 15 and 16 illustrate a
countercurrent moving catalyst bed reactor system with the two
separate moving catalyst beds 10-10. Hydrocarbon feed stream enters
the reactor vessel 11 through line 16, and after hydrotreatment
within the two moving catalyst beds 10-10, a hydrocarbon product
stream is withdrawn through center pipe 18. As previously
indicated, fresh catalyst is added to both moving catalyst beds
10-10 via pipes 19-19 and exits the respective moving catalyst beds
10 through discharge lines 82. The preferred catalyst support means
17 (which includes screen 28) has been described in detail above.
Screen 28 of the catalyst support means 17 allows gas and liquid to
flow through them but does not allow catalyst to exit through them.
Therefore, screens 28 of the catalyst support means 17-17 keep the
catalyst within its appropriate hydroconversion zone.
[0107] In FIGS. 15 and 16, each of the catalyst support means 17 is
a conical shaped catalyst support means 17 (including conical
shaped screens 28) for directing catalyst in a plug-flow fashion
towards discharge lines 82. The angle of each of the conical shaped
catalyst support means 17 including the screens 28 associated
therewith, is preferably between about 10.degree. and about
80.degree., more preferably between about 25.degree. and about
70.degree.. The circular plate members 31-31 including the
depending tubes 32 function as a distributor assembly for
distributing both gas and liquid into the plenum chambers 3-33 in
the upflow type reactor vessels 11 in FIGS. 15 and 16.
[0108] It is to be understood that the specific designs for reactor
internals within the reactor vessels 11 in FIGS. 14-16 are not
critical to the present invention. Any internals capable of
providing multiple hydroconversion reaction zones, each supplied
with catalyst displacement and replacement systems would accomplish
the purpose of the present invention. Designs similar to FIGS. 15
and 16 could be devised for cocurrent flow of hydrocarbon feed
streams through catalyst. Obviously different reactor internals
would be required, but the basic overall design would remain the
same.
[0109] It is also to be understood that while only two
hydroconversion reaction zones (i.e. the pair of catalyst beds
10-10) are shown in FIGS. 14-16, the spirit and scope of the
present invention would include more then two hydroconversion zones
(i.e. more then two catalyst beds 10). Therefore, if one was
interested in using three or more moving catalyst beds 10 within
the single moving bed reactor vessel 11, three or more
hydroconversion zones could be provided for the catalyst.
[0110] Each of the moving catalyst beds 10 would consist of
catalyst particles of different catalytic characteristics, such as
different physical and/or catalytic properties. Considering the
range of hydroconversion reaction capabilities which each of the
moving catalyst beds 10 could possess, one skilled in the art will
appreciate the variety of catalysts having a variety of physical
properties and elemental compositions, which could be used for each
of the moving catalyst beds 10. It is within the spirit and scope
of the present invention to encompass hydroconversion reaction
systems employing catalysts possessing such shape and density which
vary widely from system to system. Each of the moving catalyst beds
10 would typically have a particular hydroconversion reaction (e.g.
hydrometallation, hydrodenitrification, hydrodesulfurization; etc.)
capability that differs from the particular hydroconversion
reaction capability of the other moving catalyst bed(s) 10. Thus,
by way of example only, the catalyst particles in the bottom or
lower moving catalyst bed 10 could possess hydrodemetallation
capabilities, where as the catalyst particles in the top or upper
moving catalyst bed 10 could posses hydrodenitrification and/or
hydrosulfurization and/or any hydrotreating capabilities. A broad
and general description of any suitable catalyst for any of the two
or more moving catalyst beds 10 is described below under the
subtitle "THE CATALYST", especially for a preferred embodiment of
the present invention where each of the two or more moving catalyst
beds 10 is a "plug-flowing" catalyst bed 10 possessing different
catalytic capabilities (i.e. different physical and/or catalytic
and/or elemental properties).
[0111] Continuing to refer to FIGS. 14-16 for a method for
hydroprocessing a feed stream, the top catalyst bed 10 and the
bottom catalyst bed 10 are respectfully disposed in the respective
hydroconversion reaction zones of the reactor vessel 11. As
previously indicated, the top or upper first catalyst bed 10
contains a first set of catalytic properties or characteristics
that differ from the catalytic properties and characteristics of
the bottom catalyst bed 10. A hydrocarbon feed stream flows through
the bottom catalyst bed 10 to produce an initially-treated
hydrocarbon feed stream. From the bottom catalyst bed 10, the
initially-treated hydrocarbon feed stream flows (i.e. flows through
distributor plate 31 into plenum chamber 33 and subsequently from
the plenum chamber 33 through the catalyst support means 17)
through the upper catalyst bed 10 to produce a treated hydrocarbon
feed stream, more specifically a hydrocarbon product that exits
through center pipe 18. At least partially spent catalyst is
withdrawn from both of the catalyst beds 10-10 through the
discharge lines 82 simultaneously with the flowing of a hydrocarbon
stream through the catalyst beds 10-10. Similarly, fresh catalyst
is added to both of the catalyst beds 10-10 through pipes 19 also
simultaneously with the flowing of the hydrocarbon stream through
the catalyst beds 10-10.
[0112] The initially-treated hydrocarbon feed stream flows from the
lower catalyst bed 10 directly into the upper catalyst bed 10
without having left or exited the reactor vessel 11. Preferably,
the lower or bottom catalyst bed 10 demetallizes a hydrocarbon feed
stream in producing the initially-treated hydrocarbon feed stream,
and the top catalyst bed 10 hydrotreats (i.e. hydrodenitrifies
and/or hydrodesulfurizes and/or hydrodemetallizes) the
initially-treated hydrocarbon feed stream to remove at least one
chemical element from the initially-treated hydrocarbon feed
stream. The at least one chemical element would be selected from
the group consisting of nitrogen, sulfur, or any other element, and
mixtures thereof.
The Catalyst
[0113] In a preferred embodiment of the invention, the catalyst
which is charged to the reactor vessel 11 preferably satisfies the
following four main criteria: (i) the catalyst has the appropriate
catalytic activity and life for the particular application (e.g.
demetallation, denitrification, hydrodesulfurization, etc.); (ii)
the catalyst has physical properties which minimize its random
motion in the reactor vessel 11; (iii) the catalyst has physical
properties which minimize catalyst loss both in the catalyst
transfer steps and in the reactor vessel 11; and (iv) the catalyst
is sufficiently uniform in size and shape and density to prevent
classification by size in normal operation.
[0114] The catalyst in the present invention preferably has the
appropriate catalytic activity and life for the specific
application (e.g. demetallation, denitrification,
hydrodesulfurization, etc.). For example, if the catalyst is to be
used for demetallation, it should have sufficient HDM activity and
metals loading capacity (i.e. life) to meet the target
demetallation without the use of uneconomic amounts of catalyst.
The metals loading capacity of the catalyst is preferably greater
than about 0.10 grams of metal per cubic centimeter of catalyst
bulk volume and is more preferably greater than about 0.20 grams
metal per cubic centimeter of catalyst bulk volume. The catalyst
properties which most affect catalytic activity and metals loading
capability are: pore structure (pore volume and pore size
distribution); base material (e.g. alumina versus silica);
catalytic metals (amount, distribution, and type (nickel,
molybdenum, cobalt, etc.)); surface area; and particle size and
shape. If the catalyst is to be used for denitrification and/or
desulfurization or any other hydrotreating purposes, the same
foregoing principles and criteria would apply.
[0115] The catalyst in the present invention also preferably has
physical properties which minimize catalyst lifting into random
motion in the upflow type reactor vessel 11. Since one of the
benefits of the present invention is the countercurrent contacting
that is achieved between the reactants and catalyst, it is
preferred to maintain plug flow of the catalyst downwards through
the entire length of the reactor vessel 11. The catalyst properties
which are critical to minimizing or preventing catalyst expansion
are: catalyst particle density (highest particle density possible
is preferred while still meeting catalytic activity and metals
loading requirements); particle size (largest size practical is
preferred); skeletal density (higher skeletal density is preferred
to reduce skeletal buoyancy); and size uniformity. One of the
salient features of the present invention is that the catalyst will
not expand into random motion in the reactor vessel 11, but will
still move rather easily during flow transportation.
[0116] Under actual process conditions within the reactor,
significantly smaller catalysts could rise to the top while
significantly larger catalysts could migrate to the bottom. This
intervenes with optimal plug flow movement of catalyst. For this
reason, size specifications for the catalysts of the present
invention are narrower than those for fully packed or fixed bed and
ebullated bed catalysts.
[0117] The catalyst of the present invention should further have
physical properties which minimize catalyst loss in the catalyst
transfer steps and in the reactor vessel 11. Breakage or attrition
of the catalyst in either the transfer steps or in the reactor
vessel 11, can have significant adverse effects on the performance
of the reactor system itself and on any downstream equipment or
processing unit. The following catalyst properties are critical to
catalyst loss: catalyst attrition (minimum attrition is absolutely
required); catalyst crush strength (maximum crush strength is
required without producing a catalyst which is very brittle and
might suffer from excessive attrition); catalyst size and shape
(spherical catalyst are preferred since they move more easily and
have no rough or sharp edges to break off); and fines content
(minimum fines is an absolute requirement to avoid adverse effects
in the reactor vessel 11 and downstream equipment).
[0118] The catalyst is sufficiently uniform in size and shape and
density to prevent classification by size in normal operation.
Generally, narrow specifications are required for the catalyst to
prevent classification by size. Specific catalyst size is selected
so that it is near the point of being expanded into random motion,
but not to the point of expansion into random motion per se or
ebullation.
[0119] All of the four main criteria for the selection of the
catalyst of the present invention are important and are not
independent or mutually exclusive of each other. The four main
criteria must be balanced against each other to optimize the
catalyst for the specific application. For example, to minimize
catalyst expansion into random motion we would prefer a large and
very dense catalyst. This is contrary to the properties we might
want for a residuum demetallation application where we need a small
particle with low density diameters. These competing needs must be
balanced to ensure minimum catalyst expansion or ebullation while
achieving adequate catalytic activity and metals loading
capability, minimum attrition and minimum classification by
size.
[0120] Because there are competing catalyst requirements and
because each application is unique, the catalyst for the present
invention may be any suitable catalyst that is capable of assisting
in the operation of the invention and assisting in accomplishing
the desired objects of the invention.
[0121] The catalyst of the present invention unexpectedly produces
a plug-flowing substantially packed bed (i.e. catalyst bed 10) of
hydroprocessing catalyst during hydroprocessing by contacting a
substantially packed bed of hydroprocessing catalyst with an
hydrocarbon feed stream (i.e. a liquid component and a
hydrogen-containing gas component) that is upflowing at a rate
controlled in an amount and to an extent sufficient to limit
expansion of the substantially packed bed of hydroprocessing
catalyst to less than 10% by length beyond a substantially full
axial length of the substantially packed bed of hydroprocessing
catalyst in a packed bed state. More preferably, the expansion of
the substantially packed bed of hydroprocessing catalyst is limited
to less than 5%, most preferably less than 2% or even less than 1%
, by length beyond a substantially full axial length of the
substantially packed bed of hydroprocessing catalyst in a packed
bed state. The rate of flow of the hydrocarbon feed stream may be
any suitable rate controlled in an amount and to an extent
sufficient to limit the expansion of the substantially packed bed
of hydroprocessing catalyst, preferably the rate of flow is at a
rate ranging from about 0.01 ft/sec. to about 10.00 ft/sec.
[0122] The catalyst of the present invention more specifically
unexpectedly produces a plug-flowing substantially packed bed of
hydroprocessing catalyst when a volume of the hydroprocessing
catalyst is withdrawn or transferred under preferably laminar flow
conditions from the bottom of the substantially packed bed of
hydroprocessing catalyst while, and simultaneously to, the
substantially packed bed of hydroprocessing catalyst maximally and
optimally occupies at least about 50% by volume, preferably from
about 80% by volume to about 98% by volume (i.e. the entire
internal and/or inside available volume) of the reactor vessel 11.
The substantially packed bed of hydroprocessing catalyst of the
present invention maximally and optimally occupies a volume within
the reactor vessel 11 that is larger or greater than a volume of a
bed of catalyst in an ebullating reactor vessel that has
substantially the same entire internal and/or inside available
volume as the reactor vessel 11 and wherein the volume of the bed
of catalyst in the ebullating reactor vessel is in a "slumped" (or
packed) catalyst bed condition or state. Typically, a bed of
catalyst in an ebullating reactor vessel in a "slumped" catalyst
bed condition occupies approximately up to less than about 75% by
volume (maximum) of the entire internal and/or inside available
volume of the ebullating reactor vessel. Thus, the substantially
packed bed of hydroprocessing catalyst preferably maximally and
optimally occupies at least about 75% by volume, preferably from
about 80% by volume to about 98% by volume of the entire internal
and/or inside available volume of the reactor vessel 11. Most
preferably, the substantially packed bed of hydroprocessing
catalyst of the present invention maximally and optimally occupies
from about 75% by volume to about 95% by volume of the entire
internal and/or inside available volume of the reactor vessel
11.
[0123] The catalyst of the present invention furthermore
specifically unexpectedly produces the plug-flowing substantially
packed bed of hydroprocessing catalyst when the volume of the
hydroprocessing catalyst is withdrawn or transferred in the
hydrocarbon feed stream under preferably laminar flow conditions
from a central portion or section of the substantially packed bed
of hydroprocessing catalyst and at a lowermost or bottommost
section thereof and below the entry points of the plurality of
annular mixture zones MZ containing the hydrocarbon feed stream
(i.e. a liquid component and a hydrocarbon-containing gas
component). As previously indicated, when the volume of the
hydroprocessing catalyst of the present invention is withdrawn or
transferred to commence plug-flow, it is transferred or withdrawn
preferably laminarly in the liquid component of the hydrocarbon
feed stream and is removed from above and in proximity to an
impervious zone (i.e. imperforate center plate 25) of the bed
support means 17 and substantially out of the flow path of the
LH--HG mixtures (i.e. LH--HG.sub.2, LH--HG.sub.3, etc.) emanating
out of the mixture zones MZ (i.e. MZ.sub.2s, MZ.sub.3s, etc.). The
particular volume (or amount) of catalyst that is withdrawn at any
desired time from the bottom of the substantially packed bed of
hydroprocessing catalyst may be any suitable volume or amount which
accomplishes the desired objectives of the present invention.
Preferably, such as by way of example only, the particular volume
or amount of catalyst that is withdrawn at any desired time is a
volume or amount ranging from about 0.10% by weight to about 25.00%
by weight of the substantially packed bed (i.e. catalyst bed 10).
The rate of withdrawal of a particular volume (or amount) of
catalyst may also be any suitable volume or amount which
accomplishes the desired objectives of the present invention, such
as a rate of withdrawal where the flow rate of the catalyst (e.g.
the catalyst in the hydrocarbon feed stream) ranges from about 0.1
ft/sec. to about 20 ft/sec., more preferably from about 0.1 ft/sec.
to about 10 ft/sec., and at a catalyst concentration ranging from
about 0.10 lbs catalyst/lb. catalyst slurry (i.e. weight of
hydroprocessing catalyst plus weight of hydrocarbon feed stream) to
about 0.80 lbs catalyst/lb. catalyst slurry, more preferably from
about 0.15 lbs catalyst/lb. catalyst slurry to about 0.60 lbs
catalyst/lb. catalyst slurry. As previously indicated, the
withdrawn catalyst may be conveniently replaced by introducing a
volume of fresh catalyst through the top of the reactor vessel 11
onto the catalyst bed 10. The replacement or catalyst addition rate
may be any suitable replacement or catalyst addition rate which
will accomplish the desired objects of the present invention, such
as a flow replacement rate of the replacement catalyst (i.e. the
replacement catalyst in the hydrocarbon refined stream (e.g. gas
oil)) ranging from about 0.1 ft/sec. to about 20 ft/sec., more
preferably from about 0.1 ft/sec. to about 10 ft/sec., and at a
catalyst replacement concentration ranging from about 0.10 lbs.
replacement catalyst/lb. catalyst slurry (i.e. weight of
replacement catalyst plus the hydrocarbon refined stream (e.g. gas
oil) as the slurrying medium) to about 0.80 lbs replacement
catalyst/lb. catalyst slurry, more preferably from about 0.15 lbs
catalyst/lb. catalyst slurry to about 0.60 lbs catalyst/lb.
catalyst slurry.
[0124] In a preferred embodiment of the present invention, the
catalyst of the present invention comprises an inorganic support
which may include zeolites, inorganic oxides, such as silica,
alumina, magnesia, titania and mixtures thereof, or any of the
amorphous refractory inorganic oxides of Group II, III or IV
elements, or compositions of the inorganic oxides. The inorganic
support more preferably comprises a porous carrier material, such
as alumina, silica, silica-alumina, or crystalline aluminosilicate.
Deposited on and/or in the inorganic support or porous carrier
material is one or more metals or compounds of metals, such as
oxides, where the metals are selected from the groups Ib, Vb, VIb,
VIIb, and VIII of the Periodic System. Typical examples of these
metals are iron, cobalt, nickel, tungsten, molybdenum, chromium,
vanadium, copper, palladium, and platinum as well as combinations
thereof. Preference is given to molybdenum, tungsten, nickel, and
cobalt, and combinations thereof. Suitable examples of catalyst of
the preferred type comprise nickel-tungsten, nickel-molybdenum,
cobalt-molybdenum or nickel-cobalt-molybdenum deposited on and/or
in a porous inorganic oxide selected from the group consisting of
silica, alumina, magnesia, titania, zirconia, thoria, boria or
hafnia or compositions of the inorganic oxides, such as
silica-alumina, silica-magnesia, alumina-magnesia and the like.
[0125] The catalyst of the present invention may further comprise
additives which alter the activity and/or metals (and/or nitrogen
and/or sulfur and/or any other elements) loading characteristics of
the catalyst, such as but not limited to phosphorus and clays
(including pillared clays). Such additives may be present in any
suitable quantities, depending on the particular application for
the hydroconversion process including the applied catalyst.
Typically, such additives would comprise essentially from about
zero (0)% by weight to about 10.0% by weight, calculated on the
weight of the total catalyst (i.e. inorganic oxide support plus
metal oxides).
[0126] Although the metal components (i.e. cobalt, molybdenum,
etc.) may be present in any suitable amount, the catalyst of the
present invention preferably comprises from about 0.1 to about 60
percent by weight of metal component(s) calculated on the weight of
the total catalyst (i.e. inorganic oxide support plus metal oxides)
or and more preferably of from about 0.2 to about 40 percent by
weight of the total catalyst, and most preferably from about 0.5 to
about 30 percent by weight of the total catalyst. The metals of
Group VIII are generally applied in a minor or lesser quantity
ranging from about 0.1 to about 30 percent by weight, more
preferably from about 0.1 to about 10 percent by weight; and the
metals of Group VIB are generally applied in a major or greater
quantity ranging from about 0.5 to about 50 percent by weight, more
preferably from about 0.5 to about 30 percent by weight; while as
previously mentioned above, the total amount of metal components on
the porous inorganic support is preferably up to about 60 percent
by weight (more preferably up to about 40 percent by weight) of the
total catalyst. The atomic ratio of the Group VII and Group VIB
metals may vary within wide ranges, preferably from about 0.01 to
about 15, more preferably from about 0.05 to about 10, and most
preferably from about 0.1 to about 5. The atomic ratios would
depend on the particular hydroprocessing application for the
catalyst and/or on the processing objectives.
[0127] The groups in the Periodic System referred to above are from
the Periodic Table of the Elements as published in Lange's Handbook
of Chemistry (Twelfth Edition) edited by John A. Dean and
copyrighted 1979 by McGraw-Hill, Inc., or as published in The
Condensed Chemical Dictionary (Tenth Edition) revised by Gessner G.
Hawley and copyrighted 1981 by Litton Educational Publishing
Inc.
[0128] In a more preferred embodiment for the catalyst, the oxidic
hydrotreating catalyst or metal oxide component carried by or borne
by the inorganic support or porous carrier material is molybdenum
oxide (MoO.sub.3) or a combination of MoO.sub.3 and nickel oxide
(NiO) where the MoO.sub.3 is present in the greater amount. The
porous inorganic support is more preferably alumina. The Mo is
present on the catalyst inorganic support (alumina) in an amount
ranging from about 0.5 to about 50 percent by weight, preferably
from about 0.5 to about 30 percent by weight, most preferably from
about 1.0 to about 20 percent by weight, based on the combined
weight of the inorganic support and metal oxide(s). When nickel
(Ni) is present it will be in amounts ranging up to about 30
percent by weight, preferably from about 0.5 to about 20 percent by
weight, more preferably from about 0.5 to about 10 percent by
weight, based on the combined weight of the catalyst inorganic
support and metal oxide(s). The oxidic hydrotreating catalyst or
metal oxide component may be prepared by any suitable technique,
such as by depositing aqueous solutions of the metal oxide(s) on
the porous inorganic support material, followed by drying and
calcining. Catalyst preparative techniques in general are
conventional and well known and can include impregnation, mulling,
co-precipitation and the like, followed by calcination.
[0129] The catalyst has a surface area (such as measured by the
B.E.T. method) sufficient to achieve the hydroprocessing objectives
of the particular application. Surface area is typically from about
50 sq. meters per gram to about 300 sq. meters per gram, more
typically from about 75 sq. meters per gram to about 150 sq. meters
per gram.
[0130] The catalyst mean crush strength should be a minimum of
about 5 lbs. Crush strength may be determined on a statistical
sample of catalytic particulates. For example, a fixed number (say
30 catalyst particles) are obtained from a statistical lot
comprising a plurality of catalyst particles that are to be
employed in the hydrogenation process of the present invention.
Each catalyst particle is subsequently disposed between two
horizontal and parallel steel plates. A force is then applied to
the top steel plate until the disposed catalyst particle breaks.
The force applied to break the catalyst particle is the crush
strength. The test is repeated for the remaining catalyst
particles, and a mean crush strength is obtained. Preferably
further, no more than about 35% by wt. of the catalyst particles or
catalytic particulates have a mean crush strength of less than
about 5 lbs.; more preferably, no more than about 15% by wt. of the
catalyst particles or catalytic particulates have a mean crush
strength of less than about 5 lbs; and most preferably, no more
than about 0% by wt.
[0131] The catalyst of the present invention comprises a plurality
of catalytic particulates having a uniform size, which is
preferably spherical with a mean diameter having a value ranging
from about 35 Tyler mesh to about 3 Tyler mesh, more preferably
ranging from about 20 Tyler mesh to about 4 Tyler mesh, and most
preferably from about 14 Tyler mesh to about 5 Tyler mesh. The
Tyler mesh designations referred to herein are from a table
entitled "Tyler Standard Screen Scale Sieves" in the 1981 Edition
of Handbook 53, published by CE Tyler Combustion Engineering, Inc.,
50 Washington St., South Norwalk, Conn. 06856.
[0132] Likewise, the preferred catalyst particle has a uniformly
smooth and rounded surface. Preferred shapes include, for example,
spheres, spheroids, egg-shaped particles and the like. More
preferably, the catalyst of the present invention, is a rounded
particle having an aspect ratio of less than about 2.0, more
preferably equal to or less than about 1.5, most preferably about
1.0. More preferably, the catalyst of the present invention has a
size distribution such that at least about 90% by weight (more
preferably at least about 95% by weight; most preferably at least
about 97% by weight) of the catalytic particulates contained within
the catalyst bed 10 have an aspect ratio of less than about 2.0,
more preferably less than about 1.5, and most preferably about 1.0.
As used herein, "aspect ratio" is a geometric term defined by the
value of the maximum projection of a catalyst particle divided by
the value of the width of the catalyst particle. The "maximum
projection" is the maximum possible catalyst particle projection.
This is sometimes called the maximum caliper dimension and is the
largest dimension in the maximum cross-section of the catalyst
particle. The "width" of a catalyst particle is the catalyst
particle projection perpendicular to the maximum projection and is
the largest dimension of the catalyst particle perpendicular to the
maximum projection.
[0133] The catalyst should have a particle size distribution such
that the catalyst bed 10 expands under the conditions within the
reactor vessel 11 to less than 10% by length (more preferably less
than 5% and most preferably less than 1% by length) beyond a
substantially full axial length of the substantially packed bed of
the hydroprocessing catalyst in a packed bed state. In order to
maximize reactor throughput, the catalytic particulates have a
narrow size distribution. The catalyst employed in the
hydrogenation process of the present invention broadly comprises a
size range or size distribution such that at least about 90% by
weight, preferably at least about 95% by weight, more preferably,
at least about 97% by weight, of the catalytic particulates in the
catalyst bed 10 have a diameter ranging from R.sub.1 to R.sub.2,
wherein: (i) R.sub.1 has a value ranging from about {fraction
(1/64)} inch (i.e. the approximate opening size of a 35 mesh Tyler
screen) to about 1/4 inch (i.e. the approximate opening size of a 3
mesh Tyler screen); (ii) R.sub.2 also has a value ranging from
about {fraction (1/64)} inch (i.e. the approximate opening size of
a 35 mesh Tyler screen) to about 1/4 inch (i.e. the approximate
opening size of a 3 mesh Tyler screen); and (iii) the ratio
R.sub.2/R.sub.1 has a value greater than or equal to about 1 and
less than or equal to about 1.4 (or about the square root of 2.0).
More preferably, the catalytic particulates in the catalyst bed 10
have a diameter ranging from R.sub.1 to R.sub.2 wherein R.sub.1 and
R.sub.2 each has a value ranging from about {fraction (2/64)} inch
(i.e. the approximate opening size of a 20 mesh Tyler screen) to
about {fraction (12/64)} inch (i.e. the approximate opening size of
a 4 mesh Tyler screen), most preferably from about {fraction
(3/64)} inch (i.e. the approximate opening size of a 14 mesh Tyler
screen) to about {fraction (9/64)} inch (i.e. the approximate
opening size of a 5 mesh Tyler screen), and wherein the ratio
R.sub.2/R.sub.1 has a value ranging from about 1.00 to about 1.4
(or about the square root of 2.0).
[0134] The catalyst employed in the hydrogenation process of the
present invention also broadly comprises a size range or size
distribution such that a maximum of about 2.0% by weight (more
preferably a maximum of about 1.0% by weight and most preferably a
maximum of about 0.5% by weight or less) of the catalyst particles
or catalytic particulates has a diameter less than R.sub.1. The
catalyst also has a size range or size distribution such that a
maximum of about 0.4% by weight (more preferably a maximum of about
0.2% by weight and most preferably a maximum of about 0.1% by
weight or less) of the catalyst particles or catalytic particulates
have a diameter less than R.sub.3, wherein R.sub.3 is less than
R.sub.1 and the value of the ratio R.sub.1/R.sub.3 is about 1.4 (or
about the square root of 2.0). The catalyst particles or catalytic
particulates of the catalyst preferably have a maximum attrition of
about 1.0% by weight (more preferably a maximum of about 0.5% by
weight and most preferably a maximum of about 0.25% by weight or
less) of the catalyst particles or catalytic particulates through a
diameter (i.e., a Tyler screen opening) having a value of R.sub.1,
and a further maximum attrition of about 0.4% by weight (more
preferably a maximum attrition of about 0.2% by weight and most
preferably a maximum attrition of about 0.1% by weight or less) of
the catalyst particles or catalytic particulates through a diameter
(i.e., again a Tyler screen opening) having a value of R.sub.3
wherein R.sub.3 again (as stated above) is less than R.sub.1 and
the value of the ratio of R.sub.1/R.sub.3 is about 1.4 (or about
the square root of 2.0). [Note that the attrition procedure is
specified in ASTM D 4058-87. However, in the standard method, the
fines are removed through a 850T (.sup..about.20 mesh) screen. In
the present method, the screen is an opening equal to the minimum
catalyst size desired for the particular application, as more
specifically defined by the value of R.sub.1 and R.sub.3.] Thus, by
way of example only, for a catalyst with a specified size range of
about 10 to about 12 Tyler mesh, one would specify up to about 2.0%
by wt. fines (more preferably up to about 1.0% by wt.) MAX through
12 Tyler mesh and up to about 0.4% by wt. (more preferably up to
about 0.2% by wt.) MAX through 14 Tyler mesh. Similarly, for a
catalyst with a specified size range of about 6 to about 8 Tyler
mesh, one would specify up to about 2.0% by wt. fines (more
preferably up to about 1.0% by wt. fines) MAX through 8 Tyler mesh
and up to about 0.4% by wt. fines (more preferably up about 0.2% by
wt. fines) MAX through 10 Tyler mesh. For the catalyst with the
specified size range of about 10 to about 12 mesh, one would
specify an attrition of up to about 1.0% by wt. (more preferably up
to about 0.5% by wt., most preferably up to about 0.25% by wt.) MAX
through 12 Tyler mesh and up to about 0.4% by wt., (more preferably
up to about 0.2% by wt., most preferably up to about 0.1% by wt.)
MAX through 14 Tyler mesh. Similarly further, for catalyst with the
specified size range of about 6 to about 8 Tyler mesh, one would
specify an attrition of up to about 1.0% by wt. (more preferably up
to about 0.5% by wt., most preferably up to about 0.25% by wt.) MAX
through 8 Tyler mesh and up to about 0.4% by wt. (more preferably
up to about 0.2% by wt., and most preferably up to about 0.1% by
wt.) MAX through 10 Tyler mesh.
[0135] The specific particle density of the catalyst particles is
determined by the requirements of the hydroconversion process. For
the present invention it is preferred that the catalyst particles
have a uniform density. By "uniform density" is meant that the
density of at least about 70% by weight, preferably at least about
80% by weight, and more preferably at least about 90% by weight, of
the individual catalyst particles do not vary by more than about
10% from the mean density of all catalyst particles; and more
preferably the individual catalyst particles do not vary by more
than about 5% from the mean density of all of the particles. In a
preferred embodiment of the present invention the catalyst (i.e.
fresh catalyst) has a particle density ranging from about 0.6 g/cc
to about 1.5 g/cc, more preferably from about 0.7 g/cc to about 1.2
g/cc, most preferably from about 0.8 g/cc to about 1.1 g/cc. After
the catalyst has at least been partially spent, the particle
density would range from about 0.6 g/cc to about 3.0 g/cc, more
preferably from about 0.7 g/cc to about 3.0 g/cc and most
preferably from about 0.8 g/cc to about 3.0 g/cc. The particle size
determination will remain substantially the same as defined above.
Fines and attrition may increase during hydroprocessing.
[0136] While the catalyst of the present invention may be any
catalyst as defined above, we have discovered that the more
preferred catalyst for optimally accomplishing the objectives of
the present invention comprises in combination the following
properties: (i) a porous inorganic oxide support; (ii) one or more
catalytic metals and/or additional catalytic additives deposited in
and/or on the porous inorganic oxide support; (iii) a crush
strength at least about 5 pounds force; (iv) a uniform size ranging
from about 6 to about 8 Tyler mesh sizes; (v) a fines content up to
about 1.0 percent by weight through 8 Tyler mesh and up to about
0.2 percent by weight through 10 Tyler mesh; (vi) an attrition up
to about 0.5 percent by weight through 8 Tyler mesh and up to of
about 0.2 percent by weight through 10 Tyler mesh; (vii) a
generally uniform spherical shape; and (viii) a uniform density
ranging from about 0.7 g/cc to about 3.0 g/cc. We have discovered
unexpectedly that the more preferred catalyst having or containing
the immediate foregoing combination of properties, unexpectedly
produces in an optimal fashion the plug-flowing substantially
packed bed (i.e. catalytic bed 11) of hydroprocessing catalyst
which is simultaneously expanding to less than 10 percent by length
(more preferably less than 1% by length) beyond a substantially
full axial length of the substantially packed bed of
hydroprocessing catalyst in a packed bed state while (and
simultaneously with) the substantially packed bed of
hydroprocessing catalyst maximally and optimally occupying from
about 50% percent by volume to about 98% percent by volume (i.e.
the entire internal and/or inside available volume or reactor
volume) of the reactor vessel 11.
[0137] The particular type of porous base material or inorganic
oxide support, the particular type of catalytic metal, the pore
structure, the catalyst surface area and catalyst size, would all
depend on the intended specific application (e.g. demetallation,
desulfurization, denitrification, etc.) of the catalyst. Generally,
the more preferred catalyst comprises a porous inorganic oxide
support selected from the group consisting alumina, silica, and
mixtures thereof, and has a surface area ranging from about 75
square meters per gram to about 150 square meters per gram. The
preferred catalyst comprises catalytic metal(s), present as
oxide(s) deposited in and/or on the porous inorganic support.
Oxide(s) of the catalytic metal(s), or the metallic oxide component
of the preferred catalyst, is selected from the group consisting of
molybdenum oxide, cobalt oxide, nickel oxide, tungsten oxide, and
mixtures thereof, and comprises from about 0.5 to about 50 percent
by weight, more preferably from about 0.5 to about 30 percent by
weight, of the total catalyst (i.e. inorganic oxide support plus
metal oxide(s)). The more preferred catalyst further comprises a
general uniform spherical shape having a mean diameter ranging from
about 20 Tyler mesh to about 4 Tyler mesh. While a spherical shaped
catalyst is the more preferred catalyst, an extrudate may be
employed if it is very strong, say having a crush strength over 5
lbs. of force. The absolute size of the catalyst may vary from
application to application, but the more preferred catalyst has the
narrow size distribution as previously stated above.
[0138] From the foregoing discussion it will be clear to the
skilled practitioner that, though the catalyst particles of the
present process have a uniform size, shape, and density, the
chemical and metallurgical nature of the catalyst may change,
depending on processing objectives and process conditions selected.
For example, a catalyst selected for a demetallation application
with minimum hydrocracking desired, could be quite different in
nature from a catalyst selected if maximum hydrodesulfurization
and/or hydrodenitrification and/or hydrocracking are the processing
objectives. The type of catalyst selected in accordance with and
having the properties mentioned above, is disposed in any
hydroconversion reaction zone. A hydrocarbon feed stream is passed
through the catalyst, preferably passed through such as upflow
through the catalyst, in order to hydroprocess the hydrocarbon feed
stream. More preferably, the catalyst is employed with the various
embodiments of the present invention.
EXAMPLES
[0139] The following examples are exemplary examples of process
runs, conducted in accordance with various method steps of the
present invention and employing the apparatus in accordance with
various preferred embodiments of the present invention. The
following set-forth examples are given by way of illustration only
and not by any limitation, and set-forth a best mode of the
invention as presently contemplated. All parameters such as
concentrations, flow rates, mixing proportions, temperatures,
pressure, rates, compounds, etc., submitted in these examples are
not to be construed to unduly limit the scope of the invention.
Example I
[0140] In a semi-commercial scale residuum conversion pilot plant
operating at 100-200 BPD, the catalyst transfer procedure, as
described above, was demonstrated more than 50 times. During each
transfer, about 2 cubic feet of catalyst was moved into and out of
the reactor vessel running continuously at typical residual
desulfurization (RDS) conditions. Transfer rates up to 16 cubic
feet per hour of catalyst were accomplished through pipes with an
inside diameter 8 times larger than the catalyst diameter. Plug
flow movement of the catalyst and the absence of bed ebullation
were proven using radioactively tagged catalyst particles
incorporated in the test bed.
[0141] Among the significant features of the invention specifically
demonstrated in such runs were that: (1) ball valves, such as those
made by the Kaymr and Mogas companies, can be used to isolate the
RDS reactor from the catalyst transfer vessels, and to transfer
catalyst particles without using solids handling valves, (2) the
catalyst bed level and thus ebullation can be adequately monitored
using a gamma-ray source and detector, (3) J-tubes (all with upward
flow sections substantially shorter than the downward flow paths)
will satisfactorily transfer catalyst particles, without local
ebullation, by laminar fluid flow, (4) use of fluid feed inlet
distributor means with a conical support and concentric annular
segment plates prevents ebullation at the base of the catalyst bed
and provides adequate radial distribution of gas and liquid, by
forming concentric alternate rings of gas and liquid (5)
substantial differences were shown where bed ebullation (expansion)
occurred with one catalyst as compared to no significant bed
ebullation (expansion) with another catalyst using the same size
and shape but with a lower density, and (6) transfer of catalyst
into , and out of, a bed traveling downwardly by gravity in a
reactor vessel while continuously operating a hydroprocessing
system to react a gas containing hydrogen and a feed stream of
hydrocarbon liquids flowing as a single stream from the bottom of
the bed, will permit countercurrent flow without separation during
upward passage through and out of the top of the reactor vessel,
and (7) as the result of intermittent catalyst discharge, the
catalyst bed moves countercurrently down through the reactor in
plug-like flow.
Example II
[0142] From the foregoing tests in an apparatus under flow
conditions described in Example I, the effectiveness of the
foregoing mechanical and hydraulic factors were validated for
performing hydrotreating processing with hydrocarbon and hydrogen
streams counterflowing through a moving bed of catalyst particles,
as follows:
[0143] In a pilot plant operating at up to 4 BPD hydrocarbon feed
and hydrogen at 2200 PSI, catalyst bed expansion measurements were
made at commercial-scale flow velocities with beds of catalysts of
different sizes, shapes, and densities as indicated in Table I.
Each type of catalyst was tested separately. Bed ebullation
(expansion) was measured using a gamma-ray source and detector
means mounted to detect 10% bed expansion. Table I shows flow
velocities required to produce 10% bed expansion with several
catalysts at a standard hydrogen recirculation rate of 5000 SCFB.
These results confirmed the bed expansion results from the
semi-commercial scale plant of Example I.
[0144] Table II is a similar set of runs using beds of three of the
same catalyst particles as those tested under conditions shown in
Table I except that the liquid viscosity, liquid density and
pressure of the hydrocarbon feed stream and gas were lower in Table
II than Table I to match a different set of commercial operating
conditions. From Tables I and II the effect of catalyst particle
size, density and shape are clearly indicated for different flow
conditions for the liquid and gas components of the feed. The
design feed rates for a hydrocarbon treating process were
calculated by standard scaling procedures to indicate the values in
MBPD (thousands of barrels per day) through a reactor vessel
containing a catalyst bed 11.5 feet in diameter.
[0145] In general catalyst for commercial use would be selected on
the basis of levitation or ebullation at a selected rate which is
substantially higher than normal design feed rate, say up to 100%
greater. Additionally, these tests indicate that some commercial
catalysts will not lift at reasonable design feed rates if the
particles have a high degree of uniformity and are sufficiently
strong to maintain their integrity during movement into and out of
the reactor vessel, without attrition or breakage.
1TABLE I CATALYST BED EXPANSION STUDY TEST RESULTS 2200 PSI
Hydrogen and Flush Oil Liquid Density 51 lb/ft 3 Viscosity 1.1 cp
Gas Density 0.49 lb/ft 3 Viscosity 0.016 cp Flow rates for 10% Bed
Expansion @ 5000 SCFB H.sub.2 Effective Liquid Gas Relative
Skeletal Particle Density in Oil Velocity Velocity MBPD in 11.5
Catalyst Size Shape Density Density (1) Ft/Min Ft/Sec Ft. ID
Reactor A 1 Cylinder 2.69 1.05 0.55 0.46 0.11 13 B 1.6 Quadralobe
3.55 1.03 0.56 0.60 0.14 17 C 2 Cylinder 3.61 1.60 1.05 0.46 0.11
13 D 3.2 Sphere 2.33 0.60 0.21 0.32 0.07 9 E 3.2 Sphere 3.63 0.83
0.47 1.38 0.33 40 F 3.2 Cylinder 3.58 1.37 0.89 1.38 0.33 40 (1)
Effective Density in Oil = Density of the Particle in Oil with
Buoyancy Forces Includes = (Skeletal Density) (vol % Skeleton) +
(Oil Density) (vol % Pores) - Oil Density
[0146]
2TABLE II CATALYST BED EXPANSION STUDY TEST RESULTS With Hydrogen
and Hydrocarbon at 1000 PSI Liquid Density 48 lb/ft 3 Viscosity
0.56 cp Gas Density 0.23 lb/ft 3 Viscosity 0.013 cp Flow rates for
10% Bed Expansion @ 5000 SCFB H.sub.2 Effective Liquid Relative
Skeletal Particle Density Velocity Gas Velocity MBPD in 11.5
Catalyst Size Shape Density Density in Oil (1) Ft/Min Ft/Sec Ft. ID
Reactor C 2 Cylinder 3.61 1.60 0.53 0.13 15 E 3.2 Sphere 3.63 0.83
1.38 0.33 40 F 3.2 Cylinder 3.58 1.37 1.50 0.50 60
Example III
[0147] In a 4 foot diameter vessel a "cold model" was operated
using up to 8000 BPD water and 275 SCFM air. The features of the
inlet liquid and gas distributor as well as the hydrogen gas
redistribution and quench stages, described above and shown in the
drawings were scaled and tested. Flow measurements and underwater
photography proved that distribution of the inlet gas and liquid
was uniform across the full cross-sectional area of the catalyst
support screen in the vessel. Redistribution of the rising gas
through the inverted V-shaped sheds was shown to be surprisingly
effective even when gas was intentionally maldistributed below the
redistributor stages.
Summary of Test Results for Examples I,II and III
[0148] Briefly, these test results show that the present invention
makes possible substantially continuous flow of uniformly
distributed hydrogen and hydrocarbon liquid across a densely packed
catalyst bed to fill substantially the entire volume of a reactor
vessel by introducing the fluids as alternate annular rings of gas
and liquid at a rate insufficient to levitate the bed and with the
catalyst selected with a density, shape and size adequate to
prevent lifting of the bed at the desired feed rates. (Catalysts
are selected by measuring bed expansion in a large pilot plant run
with hydrocarbon, hydrogen, and catalyst at the design pressures
and flow velocities). At the desired flow rate, such catalyst
continually flows in a plug-like manner downwardly through the
vessel by introducing fresh catalyst at the top of the bed by
laminarly flowing such catalyst in a liquid stream on a periodic or
semicontinuous basis. Catalyst is removed by laminarly flowing
catalyst particles in a liquid stream out of the bottom of the bed.
Intake for such flow is out of direct contact with the stream of
gas flowing through the bed and the flow path is substantially
constant in cross-sectional area and greater in diameter by several
times than the diameter of the catalyst particles.
Example IV
[0149] A plurality of catalytic particulates were charged into a
reaction zone contained within a reactor, such as reactor vessel
11. The plurality of catalytic particulates formed a catalyst bed
(such as catalyst bed 10 in FIGS. 1, 8 and 9). The catalyst bed was
supported in the reactor by a truncated conical bed support similar
to the support that is generally illustrated as 17 in FIGS. 8, 9
and 11-13. An inlet distributor, such as circular plate member 31
in FIGS. 1 and 11 with the multiplicity of tubes 32, extended
across a full cross-sectional area of the reactor underneath the
truncated conical bed support to form a plenum or inlet chamber
between the inlet distributor and the truncated conical bed
support, as generally illustrated in FIGS. 8, 9 and 11-13. The
truncated conical bed support for the catalyst bed included a
series of annular polygons that included a plurality of segmented
plates (such as segmented plates 27 in FIGS. 2 and 3) connected to
or formed with radial spoke members such as members 26 FIGS. 10-13.
The plurality of segmented plates, each having a thickness of about
10 inches and a width of about 1.5 inch, were secured to 8 radial
spoke members. The interengaged segmented plates and radial spoke
members formed a web-like structure to produce essentially
annularly continuous mixture zones for receiving a flow of
hydrocarbon feed stream, and were overlayed with a screen having
screen openings with a mean diameter that was smaller than the
catalytic particulates. Each mixture zone underneath the screen had
a generally circumferentially uniform thickness.
[0150] The catalytic particulates comprised an alumina porous
carrier material or alumina inorganic support. Deposited on and/or
in the alumina porous carrier material was an oxidic hydrotreating
catalyst component consisting of NiO and/or MoO.sub.3 and/or
P.sub.2O.sub.5. The Mo was present on and/or in the alumina porous
carrier material in an amount of about 3% by wt., based on the
combined weight of the alumina porous carrier material and the
oxidic hydrotreating catalyst component(s). The Ni was present on
and/or in the alumina porous carrier material in an amount of about
1% by wt., based on the combined weight of the alumina porous
carrier material and the oxidic hydrotreating catalyst
component(s). The surface area of the catalytic particulates was
about 120 sq. meters per gram.
[0151] The plurality of catalytic particulates were generally
spherical with a mean diameter having a value ranging from about 6
Tyler mesh to about 8 Tyler mesh and an aspect ratio of about 1.
The mean crush strength of the catalytic particulates was about 5
lbs. force. The metals loading capacity of the catalyst or
plurality of catalytic particulates was about 0.3 grams of metal
per cubic centimeter of catalytic particulate bulk volume.
[0152] The catalytic particulates had a size distribution such that
98.5% by weight of the catalytic particulates in the catalyst bed
had a diameter ranging from R.sub.1 to R.sub.2 wherein: (i) R.sub.1
had a value of about 0.093 inch (i.e. the approximate opening of an
8 mesh Tyler screen); (ii) R.sub.2 had a value of about 0.131 inch
(i.e. the approximate opening size of a 6 mesh Tyler screen); and
(iii) the ratio R.sub.2/R.sub.1 had a value equal to about the
square root of 2.0 or about 1.414. The size distribution of the
catalytic particulates was also such that a maximum of about 1.0%
by weight of the catalytic particulates had a diameter less than
R.sub.1. The catalyst further also had a size distribution such
that a maximum of about 0.2% by weight of the catalytic
particulates had a diameter less than R.sub.3, wherein R.sub.3 was
less than R.sub.1 and the value of the ratio R.sub.1/R.sub.3 was
about the square root of 2.0 or about 1.414.
[0153] The catalytic particulates of the catalyst had a maximum
attrition of about 0.5% by weight of the catalytic particulates
through a diameter (i.e. a Tyler screen opening) having the value
of R.sub.1, and a further maximum attrition of about 0.2% by weight
of the catalytic particulates through a diameter (i.e. a Tyler
screen opening) having the value of R.sub.3 wherein R.sub.3 again
was less than R.sub.1 and the value of the ratio of R.sub.1/R.sub.3
was about the square root of 2.0 or about 1.414. Stated
alternatively, for the catalytic particulates with the specified
size range or distribution of about 6 to about 8 Tyler mesh, the
specified attrition for the catalytic particulates was up to about
0.5% by weight MAX through 8 Tyler mesh and up to about 0.2% by
weight MAX through 10 Tyler mesh.
[0154] The catalytic particulates had a maximum fines content of up
to about 1.0% by wt. through 8 Tyler mesh and up to about 0.2% by
wt. through 10 Tyler mesh. Stated alternatively, for the catalytic
particulates with the specified size range or distribution of about
6 to about 8 Tyler mesh, the specified fines content for the
catalytic particulates was up to about 1.0% wt. fines MAX through 8
Tyler mesh and up to about 0.2% by wt. fines MAX through 10 Tyler
mesh. The catalytic particulates had a uniform density such that
mean density of the catalytic particulates were about 0.9 g/cc.
[0155] The liquid component of the hydrocarbon feed stream was a
heavy atmospheric residuum wherein at least 95% by volume of which
boiled above about 343.degree. C. and wherein a substantial
fraction (e.g. 50% by volume) boiled above about 510.degree. C. The
"heavy" hydrocarbon feed had an undesirable metal content of about
90 ppm by weight of the "heavy" hydrocarbon feed. The
hydrogen-containing gas of the hydrocarbon feed stream was
essentially 97% pure hydrogen and was mixed with the heavy
atmospheric residuum stream in a mixing ratio of 623 liters of
hydrogen-containing gas at standard conditions per liter of heavy
atmospheric residuum in order to form the hydrocarbon feed
stream.
[0156] The hydrocarbon feed stream was passed through the inlet
distributor and introduced into the plenum chamber of reactor at a
flow rate ranging from about 0.1 ft/sec. to about 1.00 ft/sec. The
hydroprocessing pressure and temperature within the reactor were
about 2300 psig. and about 400.degree. C. respectively. From the
plenum chamber of the reactor the hydrocarbon feed stream entered
into the annular continuous mixture zones and was uniformly fed
through the screen and into the catalyst bed such as not to induce
local ebullation or eddy currents in the catalyst bed, especially
in proximity to the conical bed support which was overlayed with
the screen.
[0157] The catalyst bed in the reactor contained a plurality of
axially spaced apart hydrogen gas redistribution (or hydrogen
gas-quenching) assemblies (see FIGS. 5 and 7 as illustrative of the
hydrogen gas-quenching assemblies). As the hydrocarbon feed stream
flowed upwardly through the catalyst bed, hydrogen gas was emitted
from the hydrogen gas redistribution assemblies, which
redistributed any hydrogen-containing gas that had become channeled
in a portion of the catalyst bed below (or in close proximity to)
the hydrogen gas redistribution assemblies and further avoided
generation of local hot spots, eddy currents or ebullation in the
upper part (especially above the hydrogen gas redistribution
assemblies) of the catalyst bed.
[0158] The liquid hydrocarbon feed stream exited the reactor at a
withdrawal flow rate of about 3.6 ft/sec. and had been upgraded
such that it contained a metal content of about 3 ppm by wt. of the
liquid hydrocarbon feed stream. As the hydrocarbon feed stream
flowed upwardly through the catalyst bed, a gamma ray source in the
catalyst bed in combination with a gamma ray detector on the
reactor (see for example gamma ray source 22 in the catalyst bed 10
competing with the gamma ray detector 24 on the reactor vessel 10
in FIG. 1) detected that the catalyst bed expanded less than 10% by
length over or beyond substantially the full axial length of the
catalyst bed in a packed bed state.
[0159] After the reactor was on stream for about 1 weeks,
approximately 7.25 cubic meters (or about 3.3% by weight of the
catalyst bed) of catalytic particulates were laminarly withdrawn in
the hydrocarbon feed stream through a J-tube (such as J-tube 29 in
FIG. 1) at a flow rate of about 3.6 ft/sec. The withdrawn catalyst
in the hydrocarbon feed stream had a concentration of about 0.5
lbs. catalyst/lb. catalyst slurry (i.e. weight of withdrawn
catalyst plus weight of hydrocarbon feed stream). When and/or as
the volume of catalytic particulates were withdrawn or transferred
from the bottom of the catalyst bed, the catalyst bed (i.e. a
substantially packed bed of catalyst) began to plug-flow. The
withdrawn catalyst was replaced by introducing a comparable volume
of fresh replacement catalyst through the top of the reactor. The
fresh replacement catalyst was slurried in a hydrocarbon refined
stream (e.g. gas oil) and was introduced into the reactor at a flow
catalyst replacement rate of about 3.6 ft/sec., and at a catalyst
replacement concentration of about 0.5 lbs. replacement
catalyst/lb. catalyst slurry (i.e. weight of replacement catalyst
plus the hydrocarbon refined stream (e.g. gas oil) as the slurrying
medium).
[0160] While the present invention has been described herein with
reference to particular embodiments thereof, a latitude of
modification, various changes and substitutions are intended in the
foregoing disclosure, and it will be appreciated that in some
instances some features of the invention will be employed without a
corresponding use of other features without departing from the
scope of the invention as set forth.
* * * * *