Background of the Invention
The present invention relates to the upgrading of distillate fuels,
and more particularly to a process for removing sulfur, nitrogen and aromatic
compounds from distillate fuels using a multi-bed catalytic reactor.
When distillate fuels are produced from hydrocarbon feed, it is necessary
to remove sulfur and nitrogen and to saturate the aromatics in order that the distillate
fuel meets environmental standards and has a sufficiently high cetane number. When
hydrocarbon feeds for producing distillate fuels are subjected to desulfurization,denitrogenation
and dearomatization in a catalytic reactor, a significant temperature increase
takes place across the reactor bed due to the exothermic heats of reaction. One
known technique for compensating for this temperature rise in a multi-bed reactor
is to use interbed gas and/or liquid quenching. In this type of process, the quenching
fluid is introduced through distributors into a mixing device, known as a quench
box, which is located between adjacent catalyst beds. The quench gas or liquid
provides direct contact cooling of the reaction mixtures. When this technique is
used, quench gas, which usually is recycle gas, is injected in a quantity sufficient
to readjust the hydrogen partial pressure to the desired level. Quenching liquid,
which usually is a recycle liquid, is then used to provide for the remaining cooling
There are several disadvantages to the above-mentioned direct contact
cooling system. The use of recycle gas and/or liquid for direct contact cooling
requires the use of relatively large compression equipment for supplying cooling
gas. Furthermore, when cooling liquid is used, a larger catalyst volume may be
required than would be needed if no liquid quench were used.
Summary of the Invention
An object of the invention is to provide a process for producing
distillate fuel using a multi-bed reactor which uses an efficient cooling system
for cooling the reaction mixture.
Another object of the invention is to provide an improved method
for controlling the temperature in a multi-bed catalytic reactor which is used
for producing distillate fuels.
A further object of the invention is to provide a method for producing
distillate fuels in which quenching liquid is not required.
Another object of the invention is to provide a method for producing
distillate fuels so that a compressor of reduced size can be used.
The invention in a preferred form is a process for making distillate
fuels from a distillate hydrocarbon feed. The process comprises introducing distillate
hydrocarbon feed and hydrogen gas into the top of a first reaction zone of a multi-bed
reactor in order to produce a first reaction mixture. The first reaction mixture
is removed from the multi-bed reactor at the bottom of the first reaction zone
and is cooled in a first heat exchanger using the hydrocarbon feed as the cooling
medium. Hydrogen gas is injected into the cooled first reaction mixture in order
to increase the hydrogen partial pressure and reduce the density of the first
reaction mixture while further cooling the first reaction mixture. The cooled first
reaction mixture containing injected hydrogen gas is then introduced into a second
reaction zone of the multi-bed reactor in order to produce a second reaction mixture.
Injection of hydrogen gas after cooling in the heat exchanger creates
a density differential of the two phase mixture and promotes hydraulic circulation.
The hydrogen gas preferably is injected in an amount sufficient to provide for
about half of the total cooling requirement for the first reaction mixture prior
to introduction into the second reaction zone.
In a particularly preferred form, the process of the invention further
includes the steps of removing the second reaction mixture from the multi-bed reactor,
cooling the removed second reaction mixture in a second heat exchanger, injecting
hydrogen gas into the second reaction mixture upon cooling in order to reduce the
density of the second reaction mixture and to cool the second reaction mixture,
and introducing the cooled second reaction mixture containing injected hydrogen
gas into a third reaction zone of the multi-bed reactor in order to obtain a third
Brief Description of the Drawing
The Figure is a process flow diagram showing a preferred embodiment
of a system for producing distillate fuels in accordance with the present invention.
Detailed Description of the Invention
The process and system of the invention are particularly adapted
for use in hydrogenating hydrocarbons to produce distillate fuels. The feed typically
has about 10% by volume boiling point of from about 300°F (149°C) to 500°F (260°C),
and about 90% by volume boiling point of at least about 500°F (260°C) and no more
than about 750°F (399°C).
A representative example of a distillate hydrocarbon feed which may
be hydrogenated in accordance with the present invention has the following characteristics:
16.8 - 29.9
H/C Atomic Ratio
1.4 - 1.9
Sulfur, wt. %
0.2 - 1.2
Nitrogen, wt. %
0.01 - 0.1
FIA. vol. %
35 - 60
1 - 4
Initial Boiling Point
It is to be understood, however, that the scope of the present invention
is not to be limited to this particular distillate hydrocarbon feed.
Referring now to the Figure, the hydrogenation of the hydrocarbon
feed takes place in a reactor 10 having a first reaction zone 12, a second reaction
zone 14, and a third reaction zone 16. The first and second reaction zones 12 and
14 are packed with fixed beds of hydrogenation catalysts supported on partitions
18 and 20, respectively. The third reaction zone 16 is packed with a fixed bed
of a hydrogenation catalyst which may be a noble metal or a non-noble metal hydrogenation
catalyst and is supported on a partition 22. Partitions 18, 20 and 22 have outlet
collectors 24, 26 and 28, respectively connected thereto to provide for the removal
of the reaction mixture from the reaction zone in which the particular outlet collector
is positioned. Space 30 is provided between the partition 18 and the upper end
of reaction zone 14, and space 32 is provided between partition 20 and the upper
end of reaction zone 16, in order to allow for removal of the hot reaction mixture
out of the reactor 10 and insertion of the cooled reaction mixture combined with
hydrogen gas back into the reactor 10 in a manner described further below.
A fresh hydrocarbon feed which is to be upgraded to distillate fuel
enters the reactor system in line 34 and is mixed with hydrogen from line 11 in
line 35. The hydrocarbon feed and hydrogen gas mixture is preheated in heat exchangers
56 and 44 to the reaction temperature, i.e. a temperature of about 550 - 750°F
(288-399°C), and is subsequently transferred in line 40 to the reactor 10. Stream
36 transfers the feed mixture from heat exchanger 56 to heat exchanger 44. Bypass
line 38 around heat exchanger 44 is used, if necessary, to prevent excessive preheating
of the feed mixture. The feed enters the top of the reactor 10. The feed passes
downwardly through the catalyst bed 13 in the first reaction zone 12 under conditions
in which a substantial amount of the sulfur, nitrogen and aromatic compounds are
hydrogenated to form the desired diesel fuel products. Preferably, the first reaction
zone 12 is operated at a temperature of from about 550°F (288°C) to about 750°F
(399°C), more preferably from about 600°F (316°C) to about 710°F (377°C), and at
a pressure of from about 600 psig (41.8 atm.) to about 2,000 psig (137.1 atm.),
more preferably from about 750 psig (52.0 atm.) to about 1,500 psig (103.1 atm.)
and at a hydrogen partial pressure from about 510 psig (35.5 atm) to about 1700
psig (116.5 atm), more preferably from about 640 psig (44.5 atm) to about 1275
psig (87.6 atm) and at an LHSV of from about 0.3 hr. to about 2.0 hr.-1.
The effluent from the first reaction zone 12 is a two-phase mixture of a liquid
phase and a gas phase. The liquid phase is a mixture of the higher boiling components
of the fresh feed. The gas phase is a mixture of hydrogen, inert gaseous impurities,
and vaporized liquid hydrocarbons of a composition generally similar to that of
the lower boiling components in the fresh feed. The liquid-gas reaction mixture
from the first reaction zone 12 enters space 30 between the first and second reaction
zones, at which point the mixture is removed from the first reaction zone 12 through
outlet collector 24 and is removed from the reactor in line 42. The reaction mixture
in line 42 enters heat exchanger 44 in which it is cooled, using the feed stream,
to a temperature of about 600 - 660°F (316 - 349°C). The cooled reaction mixture
exits the heat exchanger 44 in line 46. A stream of hydrogen gas is introduced
into the cooled reaction mixture through line 48 to maintain the required hydrogen
partial pressure at the inlet to the second reaction zone and reduce the density
of the cooled reaction mixture to achieve circulation. At the same time, the gas
further cools the reaction mixture to a temperature of about 550 - 700°F (288 -
399°C). The reaction mixture with newly injected hydrogen gas is then transferred
via line 52 back to space 30 between the first and second reaction zones. The mixture
which is injected at space 30 then passes downwardly to the second reaction zone
14. No liquid or gas quench tray is needed in space 30 or space 32.
In the second reaction zone 14, the reaction mixture and the newly
injected hydrogen gas move downward together through catalyst bed 15 in order to
hydrogenate additional aromatics. At the entrance to the second reaction zone 14,
the reaction mixture with hydrogen gas preferably has a temperature of about 550
- 750°F (288-399°C). The hydrogen gas partial pressure is preferably about 600
- 1200 psig (41.8-82.5 atm.). The increased hydrogen partial pressure and cooler
temperature are favorable for shifting chemical equilibrium towards saturated compounds,
therefore providing for higher aromatics saturation. Preferably, the second reaction
zone 14 is operated at a temperature of from about 550°F to about 750°F, more preferably
from about 600°F (316°C) to about 710°F (377°C), at a pressure of from about 600
psig (41.8 atm.) to about 2,000 psig (137.1 atm.), preferably from about 750 psig
(52.0 atm.) to about 1,500 psig (103.1 atm.), and at an LHSV of from about 0.3
hr.-1 to about 2.0 hr. -1.
The liquid-gas reaction mixture from the second reaction zone 14
is removed from the second reaction zone 14 through outlet collector 26, is withdrawn
from the reactor 10 through line 54, and is transferred to a heat exchanger 56
in which it is cooled, using the fed stream, to a temperature of about 600 - 660°F
(316 - 349°C). The cooled reaction mixture is removed from the heat exchanger 56
in line 58 and is mixed with hydrogen gas from line 60 in order to maintain the
required hydrogen partial pressure at the inlet to the third reaction zone and
to further reduce the density of the reaction mixture to achieve circulation.
The cooled reaction mixture with newly injected hydrogen gas is then returned to
the reactor in line 64 and is inserted in space 32. The reaction mixture with hydrogen
gas then moves downward through the third reaction zone 16 in which remaining aromatics
are hydrogenated. In addition to saturating the aromatics, the reaction of the
hydrocarbon reaction mixture with hydrogen gas in the third reaction zone 14 acts
to strip dissolved H2S and NH3 impurities from the liquid
effluent, thereby improving the hydrogen partial pressure and, as a result, enhancing
the catalyst's kinetic performance.
The reaction mixture containing newly injected hydrogen gas which
enters third reaction zone 16 has a hydrogen partial pressure which is about 600
- 660 psig. The third reaction zone 16 with a catalyst bed 17 preferably is operated
at a temperature of from about 550°F (288°C) to about 700°F (371°C), more preferably
from about 600°F (316°C) to about 675°F (357°C), at a pressure and LHSV with approximately
the same pressure range and LHSV range as the first and second reaction zones.
The liquid and gas effluent from the third reaction zone 16 accumulates
at outlet collector 28 and is removed from the reactor 10 in line 66. This line
contains the distillate fuel product. This product may then be processed further,
such as by distillation, to remove any impurities.
The process of the invention are efficient in that no liquid quench
is required and gas quench is minimized to that which is needed to obtain the desired
hydrogen partial pressure. The injected hydrogen gas typically provides for half
of the total cooling of the reaction mixture which is required before the reaction
mixture is sent to the next reactor bed.
As mentioned above, the catalysts in the first and second zones preferably
comprise non-noble metals. As representative examples of such catalysts, there
may be mentioned nickel, Raney nickel, cobalt-molybdenum, nickel-molybdenum, and
nickel-tungsten. The catalyst in the third reaction zone may comprise a noble metal
or non-noble metal, as indicated above. Examples of noble metal catalysts include,
but are not limited to, platinum and palladium.
The catalyst is preferably supported on a support such as, but not
limited to, alumina, silica or combination thereof.
A representative example of a distillate fuel which is obtained according
to the present invention is the following:
28 - 32
Color (ASTM D-1500)
H/C Atomic Ratio
< 5-500 ppm
< 5 ppm
FIA. vol. %
30 - 40
250 - 325 (121 - 163° C)
340 - 445 (171 - 229° C)
475 - 540 (246 - 282° C)
600 - 660 (316 - 349° C)
680 - 740 (360 - 393° C)
It is noted that the process of the present invention can be used
to produce distillate fuels having sulfur concentrations as low as 5 - 50 ppm,
aromatics concentrations as low as 5 - 20%, and color specifications of 0.5 - 1
by ASTM D-1500.
This invention simplifies the reactor internals by eliminating the
need for quench box trays at the inlet to each reaction zone. This invention can
also be applied to counter-current reactor configurations. The last reaction zone
will be counter-current.