U.S. patent number 3,828,474 [Application Number 05/328,805] was granted by the patent office on 1974-08-13 for process for producing high strength reducing gas.
This patent grant is currently assigned to Pullman Incorporated. Invention is credited to Orlando J. Quartulli.
United States Patent |
3,828,474 |
Quartulli |
August 13, 1974 |
PROCESS FOR PRODUCING HIGH STRENGTH REDUCING GAS
Abstract
This invention provides a process for producing a high strength
reducing gas suitable for reducing metallic ores such as iron ore.
The process is a multi-step process using a C.sub.3 to C.sub.15
hydrocarbon such as liquid naphtha as the starting material. The
first step of the process comprises gasifying the hydrocarbon by
passing a preheated mixture of the hydrocarbon and steam through a
bed of a reforming catalyst to produce a gas consisting essentially
of methane, hydrogen, carbon oxides and steam. Carbon dioxide is
then removed from this gas mixture and the resulting gas is steam
reformed in the presence of a reforming catalyst to produce
reducing gas comprising hydrogen and carbon monoxide.
Inventors: |
Quartulli; Orlando J. (London,
EN) |
Assignee: |
Pullman Incorporated (Chicago,
IL)
|
Family
ID: |
23282518 |
Appl.
No.: |
05/328,805 |
Filed: |
February 1, 1973 |
Current U.S.
Class: |
48/214A; 48/213;
48/127.7; 252/373 |
Current CPC
Class: |
C01B
3/38 (20130101); C10G 49/007 (20130101); Y02P
30/40 (20151101); Y02P 30/446 (20151101) |
Current International
Class: |
C01B
3/38 (20060101); C01B 3/00 (20060101); C10G
49/00 (20060101); C01b 002/14 () |
Field of
Search: |
;48/197R,196R,214,213,211 ;252/373 |
References Cited
[Referenced By]
U.S. Patent Documents
Foreign Patent Documents
|
|
|
|
|
|
|
1,195,428 |
|
Nov 1967 |
|
GB |
|
820,257 |
|
Sep 1959 |
|
GB |
|
Primary Examiner: Bashore; S. Leon
Assistant Examiner: Kratz; Peter F.
Claims
Having thus described our invention, we claim:
1. A process for producing a reducing gas from a hydrocarbon
feedstock having an average of from three to 15 carbon atoms, which
comprises the steps of:
a. gasifying the hydrocarbon by passing a preheated mixture of the
hydrocarbon and steam through a bed of a reforming catalyst in a
gasification reactor to produce an effluent gas comprising methane,
hydrogen, carbon oxides and steam;
b. removing carbon dioxide from said effluent gas; and then
c. reforming the resulting purified gas in the presence of steam
and a reforming catalyst in reforming reactor to produce a gas
containing at least 88% of hydrogen and carbon monoxide.
2. A process according to claim 1, wherein the hydrocarbon
feedstock is selected from the group consisting of liquid naphtha
liquified propane and liquified butane and mixtures thereof.
3. A process according to claim 1, wherein the mixture of
hydrocarbon and steam is introduced into the gasification reactor
at a temperature of a least 660.degree.F.
4. A process according to claim 1 wherein gasification is carried
out under a pressure of from 100 psig. to 650 psig.
5. A process according to claim 1, wherein carbon dioxide removal
is effected by solvent absorption.
6. A process according to claim 5, wherein the solvent is selected
from the group consisting of monoethanolamine, potassium promoted
carbonate solution and unpromoted potassium carbonate solution.
7. A process according to claim 1, wherein the gas is heated to a
temperature of from 600.degree. to 1,200.degree.F before reforming
in step (c).
8. A process according to claim 1 wherein the gas leaves the
reformer in step (c) at a temperature of from 1,650.degree. to
2,100.degree.F.
9. A process according to claim 1, wherein the reforming in step
(c) is carried out under a pressure of from atmospheric pressure to
250 psig.
10. A process according to claim 1 wherein the reforming catalyst
in the reforming reactor comprises an alkali metal promoted nickel
catalyst in the inlet zone of the reformer and an unpromoted nickel
catalyst in the outlet zone of the reformer.
11. A process according to claim 5 wherein a hydrogen-containing
stream from said solvent absorption step is employed in a
hydrodesulfurization of said hydrocarbon feedstock prior to the
gasification step.
12. A process according to claim 5 wherein a CO.sub.2 -containing
stream from the solvent absorption step is recycled with a
hydrogen-containing stream and employed in a hydrodesulfurization
of said hydrocarbon feedstock prior to the gasification step.
13. A process for producing a reducing gas suitable for the direct
reduction of metallic ores, comprising the steps of:
a. vaporizing a hydrocarbon feedstock selected from the group
consisting of liquid naphtha, liquefied propane, liquefied butane
and mixtures thereof;
b. gasifying the vaporized hydrocarbon feedstock in the presence of
steam in a bed of reforming catalyst in a fixed-bed gasification
reactor to produce an effluent stream comprising methane, hydrogen,
carbon oxides and steam;
c. removing carbon dioxide from said effluent stream by solvent
absorption;
d. reforming the purified effluent stream from step (c) in a
tubular reactor in the presence of about 1.1 to about 1.4 Mols of
steam per atom of carbon in said purified effluent stream and in
the presence of a catalyst comprising nickel on a support material
at an inlet temperature in the range of from about 600.degree.F to
about 1,200.degree.F;
e. recovering from step (d) a high strength reducing gas containing
at least 90 percent hydrogen plus carbon monoxide and less than 10
percent water and carbon dioxide;
f. passing said reducing gas at an exit temperature in the range of
from about 1,650.degree.F to about 1,200.degree.F to a unit for
reducing metallic ores.
Description
BACKGROUND OF THE INVENTION
This invention is concerned with a process for producing a reducing
gas suitable for reducing metallic ores such as iron ore.
Previous reforming processes for the production of high strength
reducing gas for iron ore reduction utilize natural gas which is
steam reformed in a single facility at various pressures and steam
flows. With a natural gas feed it is possible to achieve H.sub.2 +
CO concentrations in the reformed gas of the order of 88 percent
and greater by operation at relatively low steam input in
conjunction with the use of a supported nickel catalyst.
It is well known in the art that to obtain a high strength reducing
gas without any cooling of the reformed gas or without removal of
CO.sub.2 it is necessary to operate the reformer at a low ratio of
steam per atom of carbon in the feed, otherwise known as the
steam-carbon ratio. For instance, in the case of methane or natural
gas, the reaction proceeds according to the following equation:
CH.sub.4 + H.sub.2 O = CO + 3H.sub.2 ( 1)
The stoichimetric requirement of steam per carbon atom for the
above equation is 1.0. However, other reactions can occur within
the reformer such as the carbon monoxide disproportionation
reaction
2 CO = CO.sub.2 + C (2)
and the cracking reaction:
CH.sub.4 = C + 2H.sub.2 ( 3)
reactions (2) and (3) are associated with deposition of carbon on
the catalyst which in turn causes deactivation of catalyst and, in
extreme cases, plugging of the catalyst bed. To avoid carbon
deposition it is necessary to introduce large amounts of steam if a
low activity reforming catalyst such as an unpromoted nickel
catalyst is used, or to increase the steam-carbon ratio slightly
above the stoichiometric level if a high activity nickel catalyst,
such as an alkali metal promoted catalyst is used. In the former
case steam-carbon ratios of from 2 to 3 to 1 are required depending
on operating conditions. In the latter case, it is possible to
operate the reformer at a relatively low steam-carbon ratio. For
example, if it were assumed that reaction (1) is carried out at a
steam-carbon ratio of 1.10 to 1.0 at conditions of 1840.degree.F
and 25 psig., the concentration of H.sub.2 + CO in the reducing gas
would be of the order of 95 percent and better. A typical
composition of the reducing gas at these conditions assuming a
methane feed gas would be:
Component Mol % ______________________________________ H.sub.2
73.09 CH.sub.4 0.35 CO 23.75 CO.sub.2 0.45 H.sub.2 O 2.36 Total
100.00 % H.sub.2 + CO 96.84
______________________________________
It will be noted that at conditions of 25 psig. and 1840.degree.F
the effluent from the reformer consists largely of H.sub.2 + CO,
small amounts of CO.sub.2 and H.sub.2 O, and residual methane, the
amount of which is dependent upon the amount of nickel catalyst
provided in the reformer.
If it is desired to operate the reformer in the same manner with
higher molecular weight hydrocarbon feeds such as liquid naphtha,
it has not been possible to achieve the high degree of H.sub.2 + CO
obtainable with natural gas feeds. Because of the higher carbon
number of the naphtha feed, it was necessary to introduce greater
amounts of steam into the process in order to suppress carbon
deposition as represented by equations (2) and (3). For example,
with a non-promoted reforming catalyst, steam-carbon ratios of from
3 to 10 to 1.0 would be required depending on feedstock and
operating pressure. However, even with a promoted reforming
catalyst it is also necessary to use relatively larger amounts of
steam albeit less than that required for conventional nickel
catalysts.
Despite operation at relatively low steam-carbon ratio in
conjunction with a promoted nickel reforming catalyst, the content
of H.sub.2 + COin the reformed gas is only about 80 percent, which
is generally considered to be unsuitable for a high efficiency iron
ore reduction operation. The basic reason for the low H.sub.2 + CO
content of the reducing gas is the diluent effect of the additional
reforming steam and associated CO.sub.2 produced in the reforming
operation. While removal of steam from the reformer effluent would
yield a high strength reducing gas, such a step would be uneconomic
because it would be necessary to cool the reformer effluent for
water condensation and removal after which it would have to be
reheated prior to introduction into the iron ore reduction
facility. Cooling and reheating requires costly heat exchange
equipment and associated large high temperature piping which would
make the process unattractive. In addition, the water condensation
step increases pressure drop and, thus, requires that the reformer
operation be carried out at much higher pressure necessitating an
increase in either the operating temperature or the steam-carbon
ratio, or both, in order to meet the residual methane
requirement.
The object of the present invention is to provide a multi-stage
process wherein hydrocarbons having an average of from 3 to 15
carbon atoms can be converted into a high strength reducing gas
suitable for reducing iron ore.
SUMMARY
According to the present invention, there is provided a process for
producing a reducing gas from a hydrocarbon feedstock having an
average of from 3 to 15 carbon atoms, which comprises the steps
of:
a. gasifying the hydrocarbon by passing a preheated mixture of the
hydrocarbon and steam through a bed of a reforming catalyst in a
gasification reactor to produce an effluent gas comprising methane,
hydrogen, carbon oxides and steam;
b. removing carbon dioxide from said effluent gas; and then
c. reforming the resulting gas in the presence of steam and a
reforming catalyst to produce a gas containing generally from 85 to
99 percent by volume hydrogen and carbon monoxide.
BRIEF DESCRIPTION OF THE DRAWINGS
FIGS. 1 to 3 are flow diagrams showing alternative embodiments of
apparatus for carrying out the process of this invention.
DETAILED DESCRIPTION OF THE INVENTION
Several reactions occur simultaneously in the gasification stage
which is carried out as an adiabatic process. These include:
CnHm + n H.sub.2 O = n CO + (n + m/2) H.sub.2 (4)
co + 3 h.sub.2 = ch.sub.4 + h.sub.2 o (5)
co + h.sub.2 o = co.sub.2 + h.sub.2 (6)
other reactions including hydrocracking are known to occur in the
gasification stage. Reactions (4) to (6) are the basic reactions
occuring in the system. Reaction 4, conversion or gasification of
the hydrocarbon to CO and hydrogen is endothermic. Reaction 5,
covering methanation of CO, and reaction 6, the CO shift
conversion, are exothermic. In the gasification reaction, steam is
consumed in the process in the production of CO and hydrogen.
However, as this reaction occurs, the highly exothermic reactions 5
and 6 proceed at a rapid rate with the net result that there is a
rise in the temperature across the reactor. The gas so produced
from the overall exothermic conversion consists essentially of
methane, carbon dioxide, hydrogen and steam, with a small amount of
carbon monoxide. The composition of the product gas from this
gasification step is determined by the temperature, pressure, and
steam ratio chosen for the equilibrium which involves a number of
reactions notable of which are the aforementioned steam -- CH.sub.4
equation (1) and CO shift conversion reaction (6).
The formation of methane is favored by low temperature and high
operating pressure as represented by the equilibrium expression
defined by equation (1).
K = CO .times. H.sub.2.sup.3 .times. p.sup.2 /CH.sub.4 .times.
H.sub.2 O .times. .SIGMA.N.sup.2
where CO, H.sub.2, CH.sub.4, and H.sub.2 O denote molal flows of
each component; .SIGMA.N the total molal flow of wet gas; and P,
the pressure at the outlet of the reactor in atmospheres
absolute.
The concentration of CO and CO.sub.2 is influenced by the quantity
of steam introduced into the system and the operating temperature
which are represented by the equilibrium expression as defined by
equation (6).
K = CO .times. H.sub.2 O/CO.sub.2 .times. H.sub.2
By removing the excess carbon in the form of CO.sub.2 from the
effluent gas from the gasification stage it is possible to simulate
a feed having essentially the same characteristics as a methane
feed in terms of carbon-hydrogen ratio and theoretical hydrogen
which is easily reformable in a subsequent reformer operation.
Removal of CO.sub.2 from the steam permits operation at essentially
the same conditions as a natural gas reformer and thus, enables
production of a reducing gas having high concentrations of H.sub.2
and CO.
The starting material for the process of the invention is a
hydrocarbon fraction having an average of from 3 to 15 carbon
atoms. Suitable hydrocarbons include liquid naphtha fractions,
fractions rich in saturated hydrocarbons such as pentanes, hexanes
and heptanes and similar light hydrocarbons including liquified
propane gas and butane. Preferably the fractions are those having
an end point of up to 350.degree. - 420.degree.F.
The gasification step of the process comprises preheating a mixture
of the hydrocarbon feedstock and steam, preferably in a
steam/hydrocarbon weight ratio of at least 1.5:1, and passing the
gaseous mixture through a bed of reforming catalyst to produce a
methane rich gas. The steam/hydrocarbon ratio increases with the
molecular weight of the hydrocarbon so that, for example, a
suitable ratio for liquified propane gas is 1.6:1, for light
naphtha is 1.8:1 and for heavy naphtha is 2.0:1. A suitable
catalyst comprises nickel, cobalt, iron or mixtures thereof on a
support, such as an unpromoted nickel catalyst or a supported
nickel catalyst promoted with alkali metal such as sodium or
potassium. Suitable supports include alumina, magnesia, zirconia
and mixtures thereof. Such catalysts are disclosed in U.S. Pat.
Nos. 3,567,411; 3,433,609 and 3,201,214.
The steam-hydrocarbon mixture is preferably introduced into the
catalyst bed at a temperature of at least 660.degree.F. The
reaction is exothermic and the catalyst bed is generally maintained
by the heat of reaction at a temperature of from 750.degree. to
1,020.degree.F. The reaction is preferably carried out at an
elevated pressure, suitably of from 100 psig. to 625 psig.
The gasification step may be carried out in a catalytic rich gas
(CRG) reactor as developed by the Bristish Gas Council. Details of
such a reactor are described in I.G.E. Journal, June 1969, pages
375 to 396. Additional processes for the production of methane rich
gases are disclosed in Hydrocarbon Processing, April 1971, pages
97, 98, 99, 102, and 110.
The methane-rich gas resulting from the gasification step is
treated for removal of carbon dioxide. This is suitably effected by
a regenerative circulating solvent system with preferred solvents
being 20 percent monoethanolamine or promoted or unpromoted
potassium carbonate solution.
The gas stream after removal of carbon dioxide is then subjected to
catalytic steam reforming in a reformer furnace. The gas stream
mixed with additional steam is heated to the inlet temperature of
the furnace and then introduced into a furnace such as tubular
reformer which contains a reforming catalyst. Suitable catalysts
are those already described with reference to the gasification step
but it is particularly preferred to use an alkali metal promoted
nickel catalyst in the upper inlet zone of the furnace and an
unpromoted nickel catalyst in the lower outlet zone. Processes for
steam reforming to produce high strength reducing gas are described
in detail in Belgium Patent No. 765,920 issued Oct. 19, 1971 and in
a paper entitled "HIGH GRADE REDUCING GAS FOR METALLURGICAL
APPLICATIONS" by Finneran et al, 30th. Iron Making Conference,
A.I.M.E. April 20, 1971, Pittsburgh, Pa.
The preferred temperature conditions in the reformer are an inlet
temperature of from 600.degree. to 1,200.degree.F, preferably from
900.degree. to 1,000.degree.F, and an outlet temperature of from
1,650.degree. to 2,100.degree.F, preferably from 1,750.degree. to
1,850.degree.F. The pressure in the reformer is suitably from
atmospheric to 250 psig., preferably from 10 to 150 psig. The
preferred steam to carbon ratio is in the range of from about 1.1
to about 1.4 mols. of steam per atom of carbon in the feed.
The gas leaving the reformer at the outlet temperature thereof is
suitably delivered directly to an iron ore reduction plant. The
reducing gas generally contains at least 88 percent of carbon
monoxide and hydrogen.
Among the advantages of the process of this invention is that the
reforming catalyst in the reformer furnace is maintained in a high
state of activity due to the presence of hydrogen in the feed. High
hydrocarbon conversion should therefore be maintained throughout
the life of the catalyst. Furthermore, the danger of carbon
deposition in the reformer is minimized by virtue of the relatively
high hydrogen content of the inlet gas and the size of the reformer
furnace is somewhat smaller than that required for natural gas
reformers of equivalent capacity due to the lower reformer duty
resulting from the presence of hydrogen in the inlet gas.
Referring to the accompanying drawings,
FIG. 1 is a flow diagram of apparatus for carrying out the process
of this invention disclosing a fixed bed gasification reactor such
as a Gas Council CRG reactor for the primary gasification stage.
Purified (i.e., desulfurized) naphtha is delivered through conduit
1 to a heat exchanger 2 located in the effluent circuit of a CRG
reactor 3. The partially heated stream passing through conduit 4 is
mixed with steam supplied through conduit 5, which steam can be
produced within the process. The combined steam-naphtha mixture
flows through conduit 6 to a preheater furnace 7 where it is heated
to the inlet conditions of the CRG reactor 3. If the naphtha feed
is supplied as a vapor it is not necessary to provide the heat
exchanger 2, in which case the combined steam-naphtha mixture flows
directly to the preheater furnace 7. In the CRG reactor 3 the
naphtha is completely converted to a methane rich gas comprising a
mixture of hydrogen, methane and carbon oxides in the presence of a
steam-reforming catalyst. The gas mixture is normally in
equilibrium with the unreacted steam at the outlet temperature of
the reactor. Examples of commercially available catalysts
particularly suited for the CRG gasification stage include British
Gas Council Type A and B CRG nickel catalysts.
The gas mixture from CRG reactor 3 then flows through conduit 9 to
the preheat exchange 2 giving up a portion of the heat to the
incoming naphtha feed, after which it flows to heat recovery
equipment 10 which may comprise a waste heat boiler, a reboiler for
supplying heat to the CO.sub.2 stripper of the CO.sub.2 removal
system, or any suitable system for recovery of waste heat.
Following vapor-liquid separation in separator 11, the disengaged
methane rich gas flows to a CO.sub.2 absorber 12 where it is
contacted countercurrently with a lean regenerative solvent
introduced through inlet 13 for removal of the bulk of the CO.sub.2
in the stream. CO.sub.2 rich solvent leaves the absorber 12 at the
bottom outlet 14 and flows to a solvent regeneration system (not
shown). Alternate prior art CO.sub.2 removal processes are
disclosed in Hydrocarbon Processing, April 1971, pages 96, 101,
103, 104, and 117.
The overhead from the CO.sub.2 absorber, consisting mainly of
methane, hydrogen and small amounts of carbon monoxide, then flows
through conduit 15 to the convection section of a tubular reformer
furnace 16 where it is preheated to an intermediate temperature
after which it is joined with steam at 17 and further preheated to
the inlet temperature of the reformer catalyst tubes. The combined
steam-gas mixture is essentially equivalent to a normal natural
gas-steam mixture in terms of carbon-hydrogen ratio, theoretical
hydrogen, and steam-carbon ratio and thus is easily reformable in
the radiant section of the reformer furnace which consists of
parallel connected reformer tubes manifolded for equal distribution
of inlet gas and externally heated by fuel burners 19. The reformer
tubes contain a supported nickel catalyst. The preferred
arrangement of catalyst consists of an alkali metal promoted nickel
reforming catalyst containing 0.1 to 10 wt. percent alkali as the
metal in the upper zone of the reformer tubes and a high activity
nickel catalyst in the lower zone. The ratio of alkali metal
promoted catalyst to high activity catalyst varies with the
requirements of the process. As a rule, the volumetric ratio of
alkali metal promoted catalyst to the total catalyst volume is from
0.2 to about 0.8 to 1.0, the actual ratio employed depending on the
inlet temperature of the reformer, the nature of the feed, and the
temperature profile throughout the reforming tube. Examples of
commercially available catalysts particularly suited for the second
stage reformer include:
Upper zone: Imperial Chemical Industries, ICI 46-1
Lower zone: Imperial Chemical Industries, ICI 57-1
The gas leaving the reformer furnace flows into an outlet header 20
at essentially reformer catalyst outlet temperature and is
delivered to an iron ore reduction facility (not shown). The
facility may be a direct reduction unit or a blast furnace. The
content of reductants in the reformed gas is of the order of 88
percent or better, preferably 88 to 99 percent CO + H.sub.2.
Typical conditions and compositions for the process represented by
the FIG. 1 flowsheet are given in Table 1 for a 365.degree.F end
point naphtha feed. In this case, the incoming naphtha feed
emanated from a desulfurization unit (not shown) in which a recycle
gas containing hydrogen, was used for removal of sulfur from
naphtha by catalytic hydro-desulfurization.
TABLE 1
__________________________________________________________________________
COMPOSITION OF PROCESS STREAMS (VOL.%) PRO- CESS FEED TO CRG
EFFLUENT FROM FEED TO CO.sub.2 OVERHEAD FROM FEED TO REFORMER
EFFLUENT FROM LOCA- REACTOR CRG REACTOR ABSORBER CO.sub.2 ABSORBER
FURNACE REFORMER FURN. TION
__________________________________________________________________________
COMP- ONENT WET DRY WET DRY WET DRY WET DRY WET DRY WET DRY
__________________________________________________________________________
NAPH- 7.62 75.00 -- -- -- -- -- -- -- -- -- -- THA CH.sub.4 0.36
3.51 32.37 58.85 57.56 58.85 74.15 75.99 40.94 75.99 0.32 0.33
H.sub.2 2.17 21.29 9.45 17.17 16.79 17.17 21.63 22.18 11.94 22.18
74.03 76.14 CO 0.01 0.10 0.70 1.26 1.24 1.26 1.59 1.63 0.88 1.63
22.41 23.05 CO.sub.2 0.01 0.10 12.50 22.72 22.22 22.72 0.20 0.20
0.11 0.20 0.47 0.48 H.sub.2 O 89.83 -- 44.98 -- 2.19 -- 2.43 --
46.13 -- 2.77 --
__________________________________________________________________________
TOTAL 100.00 100.00 100.00 100.00 100.00 100.00 100.00 100.00
100.00 100.00 100.00 100.00 TEMP., .degree.F 840 975 - 990 150 -
170 150 - 170 950 - 1000 1850 PRESS. psig. 475 - 500 450 -475 200 -
300 200 -300 75 - 100 25
__________________________________________________________________________
FIG. 2 shows a modified version of the FIG. 1 process flowsheet in
which recycle hydrogen from the CRG reactor outlet is delivered to
a naphtha hydro-desulfurization system comprising a preheater and a
catalytic facility in which a combination of cobalt-molybdate and
zinc oxide catalysts or nickel-molybdate and zinc oxide catalysts
is provided. These are considered conventional catalysts for
removal of sulfur compounds from naphtha and other similar
hydrocarbon feeds.
In using the apparatus of FIG. 2, the naphtha feed has a sulfur
compound content at an intermediate level (say 50 ppm) and the
sulfur is removed by treatment with nickel-molybdate and zinc oxide
catalysts or similar systems. For much higher sulfur contents, it
is necessary to provide additional desulfurization facilities in
the design.
Referring to FIG. 2, liquid naphtha is delivered through conduit
100 and joined at inlet 101 with a recycle gas stream in an amount
equivalent to the requirements of the desulfurization facility. The
recycle gas stream supplies the hydrogen required to convert
organic sulfur to H.sub.2 S in the downstream desulfurization step.
The combined naphtha-recycle gas mixture passes through conduit 102
to a convection bank coil shown generally by reference numeral 103,
after which the preheated mixture flows through conduit 104 to a
heat exchanger 105 for further preheating by recovery of heat from
the CRG reactor effluent circuit. The mixture flows through conduit
106 to desulfurizer heater 107 for superheating to the required
inlet conditions of a desulfurizer 109. The desulfurizer 109 is
provided with a bed of nickel-molybdate catalyst which serves to
catalyze the destructive hydrogenation of organic sulfur compounds
to H.sub.2 S, which in turn is absorbed by contact with a layer of
zinc oxide absorbant provided in the lower portion of the
desulfurizer.
The desulfurized naphtha-recycle gas mixture leaves the
desulfurizer through outlet 110 and is joined at inlet 111 with
steam produced within the process. The combined mixture flows
through conduit 112 to a heater 113 where it is heated to the inlet
conditions of a fixed bed CRG gasification reactor 115. In the CRG
reactor 115, the naphtha is converted to a mixture of hydrogen,
methane, carbon dioxide and small amounts of carbon monoxide. The
CRG effluent containing the converted gases and unreacted steam,
flows through conduit 116 to the feed preheat exchanger 105, after
which it is fed to heat recovery equipment 117 which may include a
number of items such as a steam generator, a boiler feedwater
heater and/or a reboiler for supplying heat to a CO.sub.2 stripper
(not shown) provided in the CO.sub.2 removal system. Other items
may be included in the CRG effluent circuit for removing heat. The
type of heat recovery system has no bearing, however, on the basic
design of this invention.
Following disengagement of water from converted gases in a
separator 118, the converted gases flow to a CO.sub.2 absorber 119
and are contacted countercurrently with a lean regenerative solvent
introduced through inlet 120 for removal of the bulk of the
CO.sub.2 contained in the stream. CO.sub.2 rich solvent at the
bottom of the CO.sub.2 absorber flows through outlet 121 to a
solvent regeneration system (not shown) and is subsequently
returned, after CO.sub.2 stripping, to the CO.sub.2 absorber.
The overhead from the CO.sub.2 absorber, consisting mainly of
methane and hydrogen plus small amounts of carbon monoxide, then
flows through conduit 123 to tubular reformer furnace 122 and is
processed in the same manner as that described with reference to
FIG. 1. Before delivery to the reformer furnace, a slipstream is
taken through conduit 101 from the absorber overheat, which serves
to supply the recycle hydrogen required for the desulfurization
step. The slipstream gas is compressed in a compressor 124 prior to
delivery to the desulfurization system.
Other means may be provided for supplying recycle hydrogen for the
desulfurization step such as a small separate reformer furnace
which may be located in a slipstream circuit of the CRG reactor and
serves to convert the methane in the CRG effluent to hydrogen which
is then returned to the desulfurization facility. Similarly, other
systems can be employed for desulfurization such as a cobalt
molybdate catalyst in conjunction with zinc oxide treatment. For
large amounts of sulfur in the feed, the desulfurization design may
be modified to include use of an H.sub.2 S stripper which serves to
remove H.sub.2 S by fractionation rather than by zinc oxide
absorption. For this design, only nickel-molybdate or
cobalt-molybdate catalyst is provided. The catalyst serves to
destinctively hydrogenate organic sulfur to H.sub.2 S, which in
turn is removed in the H.sub.2 S stripper.
FIG. 3 shows a further modification of the FIG. 1 and 2 flowsheets
wherein recycle hydrogen is obtained from the outlet of the
reformer furnace. The flowsheet also shows a CO.sub.2 slipstream
taken from the overhead of the CO.sub.2 stripper which can be
recycled for adjusting the CO/CO.sub.2 ratio in the recycle
hydrogen gas in the interest of improving sulfur removal or for
suppressing carbon deposition of sulfur treatment catalysts due to
CO disproportionation.
Referring to FIG. 3, liquid naphtha is delivered through conduit
200 and heated in a reformer furnace coil shown generally by
reference numeral 201. The preheated stream is joined with a
recycle stream flowing through conduit 203 in an amount equivalent
to the hydrogen required for naphtha desulfurization. The combined
mixture then flows through conduit 204 to heat exchanger 205 for
further preheating to desulfurizer heater inlet conditions. The
mixture is then superheated in a desulfurizer heater 207 to the
required inlet conditions of the desulfurizer 209 which contains
nickel-molybdate and zinc oxide catalysts. The desulfurized stream
in outlet 210 is joined with steam introduced through conduit 211,
the steam being produced within the process. The mixture flows
through conduit 212 to a heater 213 in which it is heated to the
inlet conditions of a CRG reactor 215 for conversion to hydrogen,
methane, carbon dioxide and carbon monoxide. The CRG effluent
containing converted gas and unreacted steam, flows through outlet
216 to the feed preheat exchanger 205 after which it is delivered
to heat recovery equipment 217.
Following vapor-liquid separation in a separator 218, the converted
gas is treated for CO.sub.2 removal in an absorber 219 contact with
a regenerative solvent supplied from a CO.sub.2 stripping facility
220. The CO.sub.2 stripping facility incorporates a CO.sub.2
stripper and reboiling and overhead condensing equipment.
The overhead from the CO.sub.2 absorber, consisting mainly of
methane and hydrogen plus small amounts of carbon monoxide, then
flows through conduit 221 to a reformer furnace 230 for conversion
to high strength reducing gas. A slipstream is taken through
conduit 222 from the reducing gas product circuit and recycled to
the desulfurization system along with a small amount of CO.sub.2
product taken from the CO.sub.2 stripping facility through conduit
223. Although not considered essential, the slipstream CO.sub.2
product may be used for adjustment of the CO/CO.sub.2 ratio of the
feedstream flowing to desulfurization equipment. The combined
slipstreams are compressed in a recycle compressor 225 and
ultimately joined with the incoming naphtha feed.
A bypass circuit 226 may be provided for diverting a portion of the
CRG reactor outlet gas directly to the reformer furnace in the
event that a downward adjustment of the content of reducing gas in
the reformer effluent is required.
SUMMARY OF ADVANTAGES OF THE INVENTION
1. The reducing gas produced by the process described herein may be
passed directly to any type of direct reduction unit. Direct
reduction of iron ore is discussed in the Journal of Metals,
December 1958, pp. 804-809, and in Metals Progress, January 1960,
pp. 111-115.
2. A high content reducing gas having concentrations of H.sub.2 +
CO of 88 percent and greater can be achieved with a wide range of
naphtha and similar feeds. The concentration of H.sub.2 + CO in the
reducing gas with a naphtha feed, therefore, is the same as or
higher than that obtained from a natural gas feed.
3. The reforming catalyst in the reformer furnace is maintained in
a high state of activity due to the presence of hydrogen in the
feed entering the catalyst tubes. High hydrocarbon conversion,
therefore, is expected to be maintained throughout the like of the
reforming catalyst.
4. The danger of carbon deposition in the reformer tubes is
minimized by virtue of the relatively high hydrogen content of the
reformer inlet gas.
5. The size of the reformer furnace is somewhat smaller than that
required for natural gas reformers of equivalent capacity due to
the lower reformer duty resulting from the presence of hydrogen in
the inlet gas.
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