Note : Les descriptions sont présentées dans la langue officielle dans laquelle elles ont été soumises.
Z O O~ 9
422-333
MM:570
P~ODUCTION OF ~IYDROGEN FROM
HYDROC~RBON~CEOUS FEEDSTOCK
Field of the Invention
The invention relates to tlle production of a hydrogen
containing gas from hydrocarbonacoeus feedstock. In
particular, the invention relates to the preparation of
hydrogen containing gas by a process which involves the
catalytic partial oxidation of hydrocarbonaceous feedstocks
to produce hydrogen-rich synthesis gas, which is further
processed to remove non-hydrogen components.
Description of the Prior Art
~ 1ydrocarbonaceous feedstocks, such as natural gases
recovered from sites near petroleum deposits, are a
convenient source of hydrogen. Typically, such gases
contain, as their major constituents, methane, ethane,
propane and butane. They may also include low-boiling
liquid hydrocarbons. Hydrocarbanceous feedstock can be
converted to hydrogen by first converting it into a
synthesis gas containing a major amount of hydroges),
together with minor amounts of carbon monoxide, car~on
dioxide and methane. The synthesis gas is treated to remo~-
~carbon oxides, residual methane and nitrogen ~if present) ~--
produce a gas consisting essentially of hydrogen.
The most commonly employed method for converting
hydrocarbonaceous feedstocks to synthesis gas has been
catalytic steam reforming. Steam reforming involves an
endothermic reaction exemplified for methane by the
equation:
CH4 ~ HzO ---> CO ~ 311z (l)
In this process, the hydrocarbonaceous feedstock is reacte~l
with steam in the presence of a catalyst, usually a
;~:004~9
nickel-containing catalyst, at a temperature between about
1200F (650C) and 1900F (1040C). The hydrocarbons react
with steam under these conditions to produce carbon monoxid~-
and hydrogen. Catalytic steam rerormirlg i5 an expensive
process to carry out. Not only is the steam refonner and
its nickel-containing catalyst very expensive, but also, tl
reactions are highly endothermic. Consequently, a greal:
deal of energy must be provided to drive the reaction.
Occasionally, air or oxygen is provided to the
reforming reaction in order to provide energy through
partial oxidation of hydrocarbons. Air or oxygen reformill~
is usually performed as a secondary reforming step to reduce
unreacted methane (methane slippage) to less than one
percent by volume. Vpon exiting the primary steam reformer,
lS the unreacted methane is converted in the secondary steam
reformer by the in~ection of air or oxygen, whereby the hea~
of reaction is supplied by the combu6tion of methane,
hydrogen and carbon monoxide.
Synthesis gas production may also be carried out
autothermally in an autothermal reactor by adding an oxidant
such as air. The endothermic heat of reaction is supplied
by the exothermic combustion reactions:
C~ + 20z ----> CO~ + 2~120 (2)
Hz -~ ~02 -----> H~O ( 3)
CO ~ ~ Oz ----> COz ~4)
The autothermal reactor typically consists of two
catalyst beds, the first bed providing a high ou$1et
temperature sufficient for steam reforming in the second
bed. Alternatively, the reactants can be partially re~orme!l
in a steam reforming furnace and enter the autothermal
reactor at a temperature sufficiently high to ignite
spontaneously with the entering oxygen, thus producing a
higher temperature sufficient for reforming in the
downstream catalyst bed. Autothermal reforming generally
200i~2~9
takes place at relatively low ~hroughputs. The yrocess i9
carried out at space velocities of the order of f3, 000 hr-
~to 12,000 hr-l. "Space velocity~ can be defined as ~he
volumetric hourly rate of throughput per volu~e of
catalyst. All figures quoted herein refer to the vol~lmetri~
hourly rate at standard conditions of temperature and
pressure. A disadvantage of the autothermal reforming
process for hydrogen production is that it consumes a
portion of the generated hydrogen in the combustion reactio~
that is used to provide the heat of reaction for reforming.
The foregoing procedures for producing synthesis gas
have the drawbacks of requiring expensive catalysts; large
volumes of catalyst; relatively low rates of throughput;
equipment that is expensive and, in some cases, takes up
lS excessive amounts of space; and, in some cases, requires
unacceptably large amounts of energy to drive the process.
Partial oxidation of hydrocarbonaceous feedstocks
represents one alternative to steam reforming in the
production of synthesis qas. Most of the partial oxidaton
processes that have been employed commercially are
non-catalytic processes. Non-catalytic partial oxidation
reactions however, are relatively inefficent. They operate
at high temperatures, i.e., in the range of 2,200F (1200C)
to 2,800F (lS40C) and require large amounts of oxygen.
Typically, the oxygen-to-carbon ratio required in
non-catalytic partial oxidation is greater than 0.8:1 and
often greater than 1:1. Furthermore, free carbon is
produced which is removed in a later step.
U.S. Patent 4,390,347 issued to Dille at al. describes
a process for the production of synthesis gas by the
non-catalytic partial oxidation of a liquid
hydrocarbonaceous fuel. The hydrocarbonaceous feedstock is
reacted with a free oxygen-containing gas in the yresellce o f:
steam at an autogenously maintained temperature within the
2004~
range of 1700F (930C) to 3000F (1650C) at a pressure in
the range of about 1 to 23 atmospheres (1 to 23 bar)
absolute. The oxygen-to-carbon molar ratio i8 said to be
from 0.7:1 to 1.5:1.
S U.S. Patent 3,890,113 issued to Child et al, describes
the production of a methane-rich stream in which
non-catalytic partial oxidation of a hydrocarbonaceous
feedstock is carried out in the pres~nce of steam and
oxygen. The ratio of free oxygen in the oxidant to carbon
in the feedstock is in the range of 0.8:1 to 1.5:1. The
product synthesis gas is subjected to a water gas shift
reaction to increase the amount of hydrogen in the gas.
U.S. Patent 3,927,998 issued to Child et al., relates
to the production of a methane rich stream by the partial
oxidation of a hydrocarbonaceous fuel employing a steam to
fuel weight ratio of 2.2:1 to 2.9:1 and an oxygen-to-carbon
molar ratio of 0.8:1 to 0.84:1. The partial oxidation is
carried out in the absence of catalysts. The synthesis gas
is cooled and water, carbon dioxide, carbon and other
impurities are removed. The hydrogen and carbon monoxide in
the gas are reacted in a catalytic methanation zone to
produce a methane-rich stream.
Conversion efficiency of oxidation processes can
generally be improved by the use of catalysts; but where the
oxidation process in only partial, i.e. with insufficient
oxygen to completely oxidize the hydrocarbon, then the
catalyst is sub~ect to carbon deposit and blockage. Carbon
deposits can be avoided by using expensive catalyst
materials in generally uneconomical processes. For example,
U.S. Patent 4,087,259, issued to Fu~itani et al., describes
employment of a rhodium catalyst in a process wherein liquid
hydrocarbonaceous feedstock is vaporized and then partially
oxidized in contact with the rhodium catalyst at a
temperature in the range of 690 to 900C with optional
2()0~9
-- s
steam added as a coolant at a rate not more than 0.5 by
volume relative to the volum~ Or the Iiqllid llydt-)car~ol- in
terms of the equivalent amount of water. The r~lodium
catalyst enables partial oxidation without causing
deposition of carbon, but at temperatures greater thall
900C, thermal decomposition ensues producing ethylene or
acetylene impurities. When steam is added, the c~uan~ity of
hydrogen produced is increased while the yield of carbon
monoxide remains constant due -to catalytic decomposition of
the steam to hydrogen gas and oxygen. ~ "LHSV" (Liquid
Hourly Space Velocity) from 0~5 to 25 l/hour is disclosed;
particularly, a high yield from partial oxidation of
gasoline vapor, without steam, is produced at a temperature
of 725C and at a LHSV of 20, and with steam, is produced a~
temperatures of 700C and ~00C and at a L}ISV of 2.
In order to obtain acceptable levels of conversion
using catalytic partial oxidation processes of the prior ar~
it has been necessary to use space velocities below about
12,000 hr-1. For example, U.S. Patent No. 4,522,~94,
issued to l~wang et al., describes the production of a
hydrogen-rich gas to be used as fuel for a fuel cell. The
process reacts hydrocarbon feed with steam and an oxidant in
an autothermal reformer using 2 catalyst ~ones. The total
hourly space velocity is between 1,960 hr-' and 18,000
hr-~. Because the prior art processes must be carried out
at low space velocity, catalytic partial oxidation reactors
of the prior art have had to have large catalyst beds in
order to achieve the throughputs desired in commercial
operation. This increases the si~e and cost of the p;3l-~ial
oxidation reactor.
It is an object of the present invention to provi(3e a
process for the production of a hydrogen rich gas from
hydrocarbonaceous feedstock which is energy efficient, is
capable of using low cost catalysts and employs relatively
Z00~2~9
small equipment volume to achieve commercially acceptable
throughput.
It is a further ob~ect of the invention to provide a
process for the production of a hydrogen rich gas from
hydrocarbonaceous feedstock Wit]l a relatively low oxygen
demand, thereby increasing throughput of hydrocarbonaceous
feed.
These and other objects of the invention are achieved
by a process which is described below.
Summary of the Invention
The invention provides a process for producing hydroge
from hydrocarbonaceous feedstocks in which the feedstock is
converted into hydrogen by a process scheme which is highly
cost efficient and utilizes equipment that is less costly
and space-consuming than that employed in prior art
processes.
This invention provides a process for the production oI
a hydrogen rich gas in which hydrogen-rich synthesis gas is
generated by the catalytic partial oxidation of a
hydrocarbonaceous feedstock such as natural gas Wit}l an
oxidant stream that is preferably oxygen-rich under
temperature and steam conditions avoiding produc~ion of ~re
carbon; converting the carbon nonoxide in the synthesis gas
to carbon dioxide by subjecting the synthesis gas to a wate
shift gas reaction; removing carbon dioxide f~om the gas
stream; and recovering hydrGgen-rich gas.
In one embodiment, the invention provides a process foJ
producing hydrogen from hydrocarbonaceous ~eedstock whi.ch
comprises:
(a) introducing to a catalytic partial oxidation zone
an essentially completely mixed gaseous mixture ol
a hydrocarbonaceous feedstock, oxygen or arl
~oo~
oxygen-containing gas and, optionally, steam in
which the steam-to-carbon molar ratio is from 0:1
to 4.0:1 and the oxygen-to-carbon Inolar ratio is
from 0.4:1 to O.rJ:l, said mixture being intro~uce-l
to the catalytic partial oxidation zorle at a
temperature not lower than 200F (93C) below its
catalytic autoignition temperature;
(b) partially oxidizing the hydrocarbonaceous
feedstock in the catalytic partial oxidation zone
under temperature and steam conditions avoiding
formation of free carbon by passing the mixture
through a catalyst capable of catalyzing the
partial oxidation of the hydrocarbons, said
catalyst having a ratio of geometric surface area
to volume of a~ least 5 cmZ/cm3 and a volume
corresponding to a space velocity of between
20,000-1 and 500,000-1, thereby producing
synthesis gas containing hydrogen, carbon monoxide
and carbon dioxide;
(c) separating hydrogen in the synthesis gas to
produce a hydrogen product stream.
Brief Description of the Drawinqs
Fig. 1 is an elevated cross-section view of a reactor
having at its input a mixer and distributor suitable ~or
introducing the reactants to the catalyst bed.
Fig. 2 is an enlarged elevational cross-section view Or
a broken-away portion of the mixer and distributor ol Fig.
1. .
Fig. 3 is a top view of a broken-away quar~er sect iOIl
of the mixer and distributor of Fig. 1.
Fig. 4 is a bottom view of a broken-away quarter
section of the mixer and distributor of Fig. 1.
20042~9
-- 8 --
Fig. 5 is a diagrammatic elevational cross-sectional
illustration of a broken-away portion of the mixer an~
feeder of Figs. 1 and 2 showing critical dimensions.
Fig. 6 is a flow diagram in schematic form illustratir
S a process of the invention for hydrogen production.
Fig. 7 is a yraph plotting oxygen-to-carborl molar ratio
vs. steam-to-carbon molar ratio for three different
operating temperatures at an operating pressure of 400 psig
(2760 KPa).
Fig. 8 is a graph plotting the hydrogen-to-carbon
monoxide molar ratio in the catalytic partial oxidation
reaction product vs. the steam-to-carbon molar ratio for
three different operating temperatures at an operating
pressure of 400 psig (2760 KPa).
Fig. 9 is a graph plotting the volume ~ methane in the
catalytic partial oxidation product vs. the steam-to-car~on
molar ratio for three different operating tPmperatures at an
operating pressure of 400 psig (2760 KPa).
Fig. 10 is a graph plotting the volume % carbon dioxid~
in the catalytic partial oxidation product vs. the steam-to-
carbon molar ratio for three different operating
temperatures at an operating pressure of 400 psig (2760
KPa).
Fig. 11 is a graph plotting the molar ratio of total
hydrogen and carbon monoxide in the product to total
hydrogen and car~on in the feedstock vs. the steam-to-carl)c
molar ratio for three different operating temperatures a~ a
operating pressure of 400 psig (2760 KPa).
Fig. 12 is a process flow diagram of a hydrogen
producing plant using catalytic partial oxidation in
accordance with the invention.
'~ O ~ ~tZ ~9
Detailed Description of the Preferred Embodimellts
The process of the present invention can be used to
produce hydrogen from any gaseous or low-boiling
hydrocarbonaceous feedstock. Typically, the
hydrocarbonaceous feedstock is a gas containing principally
methane, such as natural gas having the following
approximate composition: methane, 93%; ethane, 5%; propane
1.5~; butane and higher hydrocarbons, 0.5~.
In general, the process of the invention involves tlle
steps of catalytic partial oxidation of hydrocarbonaceous
feedstock at a space velocity in the range from 20,000 hour-
1 to 500,000 hour-1 under temperature and steam conditions
avoiding production of free carbon to produce synthesis gas;
treatment of the resultant synthesis gas to remove
components other than hydrogen (e.g., carbon oxides) and to
recover a carbon dioxide stream; and recovery of a
hydrogen-rich gas. Preferred methods for carrying out these
steps are described below.
CatalYtic Partial Oxidation
The catalytic partial oxidation of hydrocarbonaceous
feedstock is carried out according to a process described i--
copending, commonly assigned U.S. application Serial No.
085,160 filed Auyust 14, 1987 in the names of M. Dunster an~
J.D. ~orchnak.
One particular aspect of the invention is the
substantial capital cost savings and/or advantageous
operating economy resulting from the employment of catalyti
partial oxidation to produce the raw synthesis gas employed
in the hydrogen producing process. This is made possil~)e
the discovery that catalytic partial oxidation performed at
a temperature, as measured at the exit of the catalytic
reaction zone, equal to or greater than a minimum non-
zooL~z~
- lo -
carbon-forming temperature selected as a linear functio
which includes a range from 1600F (~70C) to 1900F
(1040C) corresponding to a range of the steam-to-car~on
molar ratio from 0.4:1 to 0:1 and at a space velocity in thl-
range from 20,000 hour-1 to 500,000 hour-l produces
essentially no free carbon deposits on the catalyst.
Further, it is found that products of the partial catalytic
oxidation in the process of the invention consist
essentially of hydrogen, carbon monoxide and carbon dioxide
at oxidation temperatures equal to or greater than t~e
minimum temperature, rhodium catalysts are not required to
prevent carbon formation. For example in Fig. 7, dotted
line 25 represents a generally linear function which, at a
steam/carbon ratio of 0, corresponds to a minimum partial
oxidat~on temperature of about 1900F (1040C), and at a
steam/carbon ratio of 0.4 corresponds to a minimum partial
oxidation temperature of about 1600F (870C); favorable
catalytic partial oxidation without producing free carbon
occurs at temperatures and steam/carbon ratios equal to or
greater than points on the line. Further, lower minimum
temperatures at corresponding steam/carbon ratios greater
than 0.4 can be extrapolated from the linear function
represented by line 25.
In the catalytic partial oxidation step of the process
of the invention, reactant gases are introduced to the
catalytic partial oxidation reacton zone, i.e. the catalyst
bed, at an inlet temperature not lower than 200F (93C)
below the catalytic autoignition temperature of the feed
mixture. Preferably the reactant gases are introduce(l a~ a
temperature at or above the autoignition temperature o[ t~le
mixture. The reactants should be completely mixed prior ~o
the reacton. Introducing the thoroughly mixed reactall~
~ases at the proper temperature ensures that the partial
oxidation reactions will be mass transfer controlled.
zoo~
Consequently, the reaction ~ate is relatively indeyendent of
catalyst activity, but deper.~ent on the surace-area-
to-volume ratio o~ tlle cata';sL. IL is F)ossi~ o ~ e ~ly
of a wide variety of materi21s as a catalyst, pr~vided that
the catalyst has the desire~ surEace-area-to-volume ratio.
It is not necessary that th~ catalyst have specific
catalytic activity ror stea. reforming. Even materiais
normally considered to be nc~-catalytic can promote the
production of synthesis gas herein when used as a catalyst
in the proper configuration. The term "catalyst~, as used
herein, is intended to encor?ass such materials.
The catalytic partial cxidation step can be understood
with reference to the figures. The catalytic partial
oxidation zone is typically the catalyst bed of a reactor
such as that illustrated in Fig. 1. ~s shown in Fig. 1, a
reactor 100 for partially o~.idizing a gaseous feedstock
includes an input mixing anc distributor section indicated
generally at 30. The mixer and distributor 30 mixes the
feedstock with an oxidant ar.~ distributes the mixture to the
entrance of a catalytic reactor section indicated generally
at 32 wherein the feedstock is partially oxidized to produce
a product which is then passed through the exit section
indicated generally at 34.
The reactor includes ar. outer shell 40 of structural
metal such as carbon steel ~-ith a top 42 secured thereon by
bolts (not shown) or the li~;e. A layer 43 of insulation,
such as 2300F (1260C) ~PC~ ceramic fiber insulation, is
secured to the inside of the upper portion of the shell 40
including the top 42. In t~.e lower portion of the mixing
section 30, in the reactor section 32 and outlet sectioll 34
there are secured layers 46, 48 and 50 on the inside of tlle
shell. The layer 46 is a castable or equivalent ilusul~lLi~n
such as 2000F (1090C) ceramic insulation. The layer 48 i
also a castable or equivalent layer of insulation but
200`'a~9
!l
- 12 -
containing 60% alumina for withstanding 3000F (1650C).
The internal layer 50 is a refractory or equivalent layer
such as 97% alumina Wit)l ceramic anchors or 97~ alumina
brick for withstanding the interior environment of the
reactor section.
qlhe catalytic reactor sectior. 32 contains one or more
catalyst discs 54. As shown, the reactor contains a
sequence of discs 54 separated by high alumina rings 5~
between each adjacent pair of discs. The stack is supported
by a grill with high alumina bars 56. A sample port 60 is
formed in the lower end of the reaction section and has a
~ube, such as type 309 stainless steel tube 62, extending
below the bottom refractory disc 54 for withdrawing samples
of the product.
The outlet section 34 is suitably formed for being
connected to a downstream heat recovery boiler (not showrl)
and/or other processing equipment.
The catalyst comprises a high surface area material
capable of catalyzing the partial oxidation of the
hydrocarbonaceous feedstock. The catalyst is in a
confi~uration that provides a surface area to volume ratio
of at least 5 cm2/cm3. Preferably, the catalyst has a
geometric surface area to volume ratio of at least 20
cm2/cm3. While there is no strict upper limit of surface
area to volume ratio, it normally does not exceed about 40
cm 2 / cm3
The catalyst disc 54 can be, for example, a monolithic
structure having a honeycomb type cross-sectional
configuration. Suitable monolithic structures of this type
are produced commercially, in sizes smaller than those use~l
in the process of the invention, as structural substrates
for use in the catalytic conversion of automobile exhausts
and as catalytic combustion chambers of ~as turbines or for
catalytic oxidation of waste streams. Typically, the
2004Z~9
- 13 -
monolithic structure is an extruded material containillg a
plurality of closely packed channels running t~lrou~h the
length of the structure to form a honeycomb structure. ~he
channels are typically square and may be packed in a densit~
as high as 1,200 per square inch (190 per cmZ) of cross
section. The monolithic structure can be constructed of an~
of a variety of materials, including cordierite
(MgOtAl203/SiO2), Mn/MgO cordierite (Mn-MgO/Al203/~iO2),
mullite (Al203/SiO2), mullite aluminum titanate
~A120~/SiO2-(Al,Fe)203/TiO2), zirconia spinel
(ZrO2/MgO/~1203), spinel (MgO/~1203), alumina (~1203) ~n~l
high nickel alloys. The monolithic catalyst may consist
solely of any of these structural materials, even thougl
these materials are not normally considered to have
catalytic activity by themselves. Using honeycombed
substrates, surface area to volume ratios up to 40 cm~/cm-~
or hiqher can be obtained. Alternatively, the monolithic
substrate can be coated with any of the metals or metal
oxides known to have activity as oxidation catalysts. These
include, for example, palladium, platinum, rhodium,
irridium, osmium, ruthenium, nickel, chromium, cobalt,
cerium, lanthanum and mixtures thereof. Other metals which
can be used to coat the catalyst disc 54 include noble
metals and metals of groups IA, II~, III, IV, VB, VIB, or
VIIB of the periodic table of elements.
The catalyst discs 54 may also cohsist of structural
packing materials, such as that used in packing absorption
columns. These packing materials generally comprise ~hirl
sheets of corrugated metal tightly packed together to form
elongate channels runniny therethrough. The structural
packing materials may consist of corrugated sheets of metals
such as high temperature alloys, stainless steels, chromium,
magnanese, molybdenum and refractory materials. These
materials can, if desired, be coated with metals or metal
200~
- 14 -
oxides known to have catalytic activity for the oxidation
reaction, such as palladium, platinum, rhodium, irridium,
osmium, ruthenium, nickel, chromiu~, cobalt, cerium,
lanthanum and mixtures thereof.
The catalyst discs 54 can also consist of dense wire
mesh, such as high temperature alloys or platinum mesh. If
desired, the wire mesh can also be coated with a metal or
metal oxide having catalytic activity for the oxidation
reaction, including palladium, platinum, rhodium, irridium,
osmium, ruthenium, nickel, chromium cobalt, cesium,
lanthanum and mixtures thereof.
The surface area to volume ratio of any of the
aforementioned catalyst configurations can be increased by
coating the surfaces thereof with an aqueous slurry
lS containing about 1% or less by weight of particulate metal
or metal oxide such as alumina, or metals of groups IA, IIA,
III, IV, VB, VIB and VIIB and Eiring the coated surface at
high temperature to adhere the particulate metal to the
surface, but not so high as to cause si.ntering of the
surface. The particles employed should have a BET
(Brunnauer-Emmett-Teller) surface area greater than about 10
m2~gram, preferably greater than about 200 m2/gram.
In the practice of the invention, a gaseous mixture of
hydrocarbonaceous feedstock, oxygen or an oxygen-containing
gas, which can be air, oxygen-enriched air, or other
oxygen-rich gas, and, optionally, steam is introduced into
the catalytic partial oxidation zone at a temperature not
lower than 200F (93C) below its autoignition temperature.
Preferably, the oxidant is an oxygen-rich gas. The term
"oxygen-rich gas" as used herein refers to a ga~ containin~
at least 70 mol. % oxygen, preferably at least 90 mol.
oxygen, and which can be pure oxygen. Preferably, t}le
gaseous mixture enters the catalytic partial oxidation zone
at a temperature equal to or greater than its autoignition
20042~9
- 15 -
temperature. It is possible to operate the reactor in a
mass transfer controlled mode wi~h the reactants entering
the reaction zone at a temperature somewhat below the
autoignition temperature since the heat of reaction will
provide the necessary energy to raise the reactant
temperature within the reaction zone. In such a case,
however, it will generally be necessary to provide heat
input at the entrance to the reaction zone, for example by a
sparking device, or by preheating the contents of the
reactor, including the catalyst, to a temperature in excess
of the autoignition temperature prior to the introduction ol
the reactants to initiate the reaction. If the reactant
temperature at the input to the reaction zone is lower than
the autoiginition temperature by more than about 200F
(93C), the reaction becomes unstable.
When the reactant mixture enters the catalytic partial
oxidation zone at a temperature exceeding its autoignition
temperature, it is necessary to introduce the mixture to the
catalyst bed immediately after mixing; that is, the mixture
of hydrocarbonaceous feedstock and oxidant should preferably
be introduced to the catalyst bed before the autoignition
delay time elapses. It is also essential that the gaseous
reactants be thoroughly mixed. Failure to mix the reactants
throughly reduces the quality of the product and can lead t~
overheating. A suitable apparatus for mixing and
distributing the hydrocarbonaceous feedstock and oxygen or
oxygen-containing gas so as to provide thorough mixing and
to introduce the heated reactants into the reaction zone in
a sufficiently short time is illustrated in Figs. 1-5 and
described in more detail in copending commonly assigned
patent ~pplication Serial No. 085,159, filed August 14, 1987
in the names of J.D. Korchnak, M. Dunster, and J.ll. Mantell.
Referring to Fig. 1, one of the feed gases, i.e.
hydrocarbonaceous gas or oxygen-containing gas, is
200~ 9
I'
- 16 -
introduced into the input section 30 through a first inlet
port 66 through the top 42 wtlicll communicates to an upper
feed cone 68 which forms a first chamber. The cone 6a is
fastened by supports 69 in the top 42. The other feed gas
is introduced into the illpUt section 30 through second
inlets 70 extending througll side ports of the shell 40 and
communicating to a second chamber 72 which is inl;erposed
between the upper chamer 68 and the inlet of the catalyst
reaction section 32. ~ ring 73 mounted on the central
portion of an upper wall 75 of the chamber 72 sealingly
engages the lower edge of the cone 68 so that the wall 82
forms a common wall between the upper chamber 68 and lower
chamber 72. The chamber 72 has an upper outer annular
portion 74, see also Figs. 2 and 3, which is supported on
lS the top surface of the refractory layer 50. A lower portion
of the chamber 72 has a tubular wall 76 which extends
downward in the refractory sleeve 50. The bottom of t~e
chamber 76 is formed by a cast member 78.
Optionally, steam can be introduced into either or both
. of the hydrocarbonaceous feedstock and oxygen or
oxygen-containing gas. The gases are fed to the reactor in
relative proportions such that the steam-to-carbon molar
ratio i8 in the range from 0:1 to 4.0:1, is preferably in
the range from 0.8:1 to 3.0:1, and is most preferably in the
range from 1.0:1 to 2.0:1. The oxygen-to-carbon ratio is in
the range from 0.4:1 to 0.8:1 and is preferably in the range
from 0.45 to 0.65.
The reactant mixture preferably enters the catalytic
reactor section 32 at a temperature at or above its
autoignition temperature. Depending on the particular
proportions of reactant gases, the reactor operating
pressure and the c`atalyst used, this will generally be
between about 550F (288C) and 1,100F (593C).
Preferably, hydrocarbonaceous feedstock and steam are
~o~z~
- 17 -
admixed and heated to a temperature in the range from 650F
(343C) to 1,200F (650C) prior to passage through inlet
port(s) 70 or 66. Oxygen or oxygen-containing gas, such as
air, i8 heated to a temperature in the range from 150F
(65C) to 1200F (650C) and passes through the other inlet
port(s) 66 or 70.
Referring to Fig. 2, the mixing and distributLng means
comprises a plurality of elon~ ted tubes 80 having upper
ends mounted in the upper wall 75 of the chamber 72. The
lumens of the tubes at the upper end communicate with the
upper chamber 68. The bottom ends of the tubes 80 are
secured to the member 78 with the lumens of the tubes
communicating with the upper ends of passageways 84 formed
vertically through the member 78. Orifices 86 are formed in
the walls of the tubes 80 for directing streams of gas from
the chamber 72 into the lumens of the tubes 80. The inlets
66 and 70, the cone 68, the supports 69 are formed from a
conventional corrosion and heat resistant metal while the
chamber 72, tubes 80 and member 78 are formed from-a
conventional high temperature alloy or refractory type
material.
The number of tubes 80, the internal diameter 90 (see
Fig. 5) of the tubes 80, the size and number of the orifices
86 in each tube are selected relative to the gas input
velocities and pressures through inlets 66 and 70 so as to
produce turbulent flow within the tubes 80 at a velocity
exceeding the flashback velocity of the mixture. The
minimum distance 92 of the orifices 86 from the bottom end
of the tube 80 at the opening into the diverging passageways
84 is selected to be equal to or greater than that required
for providing substantially complete mixing of the gas
streams from chambers 68 and 72 under the conditions of
turbulence therein. The size of the internal diameter 90 of
the tubes 80 as well as the length 94 of the tubes is
Z00~9
- 13 -
designed to produce a sufficient pressure drop in the gas
passing from the chamber 68 to the reaction chaml~er so as to
provide for substantially uniform gas flow through the tubes
80 from the chamber 68. Likewise the size of the orifices
86 is selected to provide sufficient pressure drc~p between
the chamber 72 and the interior of the tubes 80 relative to
the velocity and pressures of the gas entering through
inlets 70 so as to provide substantially uniform volumes of
gas flows through the orifices 86 into the tubes 80.
The diverging passageways in the member 78 are formed
in a manner to provide for reduction of the velocity of the
gas to produce uniform gas distribution over the inlet of
the catalyst. The rate of increase of the cross-section of
the passageway 84 as it proceeds downward, i.e., the angle
98 that the wall of the passageway 84 makes with the
straight wall of the tubes 80, must generally be equal to or
less than about 15 and preferably equal to or less than 7
in order to minimize or avoid creating vortices within the
passageways 84. This assures that the essentially
completely mixed gases, at a temperature near to or
exceeding the autoignition temperature, will pass into the
catalyst bed in a time preferably less than autoignition
delay time. The configuration of the bottom end of the
passageways 84, as shown in Fig. 4, is circular.
2~ The catalytic partial oxidation reaction is preferably
carried out in the catalytic reaction section 32 at a
pressure greater than 100 psig (690 KPa), more preferably at
a pressure greater than 250 psig (1720 KPa). Generally the
catalytic partial oxidation is performed at a temperature,
as measured at the exit of the catalyst,-in the range from
1400F (760C) to 2300F (1260C). Preferably, the
catalytic partial oxidation temperature, as measured at the
exit, is in the range from 1600F (870C) to 2000F
(1090C). At temperatures below about 1400F (760C),
200'~Z~9
- 19 -
I uneconomic concentrations of methane are produced, and at
temperatures above 2300F (1260C), excessive amounts of
oxygen or oxygen-containing gas are required.
Essentially little or no reforming reactions are
employed in the process of the invention; that is, the
process of the invention relies essentially solely on
partial oxidation and the water gas shift reaction
CO ~ H~O ---~ CO2 + H2 (5)
to convert hydrocarbonaceous feedstock to synthesis gas.
Catalytic partial oxidation of uniformly premixed feedstock
and oxygen does not require any reforming reactions to take
place. The catalyst is selected to promote the partial
oxidation reaction, and not necessarily any reforming
reac~ion. The steam reforming reaction (equation l)
generally requires a low space velocity, i.e. generally
below about 12,000 hour-1, and the employment of space
velocities above 20,000 hour- in the present process
precludes efficient steam reforming of the feedstock. It is
believed that increased hydrogen production, above that
attributable solely to partial oxidation, is due more to the
wa~er gas shift reaction (equation 5) than to the steam
reforming reaction (equation 1).
The product ga~ exiting the outlet section 34 consists
essentially of hydrogen, carbon oxides, i.e. carbon
monoxide and carbon dioxide, methane, water vapor and any
inert components (e.g. nitrogen or argon) introduced with
the feedstock or oxidant. Trace amounts of C2 and higher
hydrocarbons may be present in the product gas. As used
herein "trace amounts" means less than about 0.1% by weight.
Removal of Carbon Oxides
The synthesis gas exiting the cata~ytic partial
oxidation zone is cooled to a temperature from about 350F
(175C) to about 750F (400C) using conventional heat
exchange methods, either by heating the hydrocarbon and
2004Z~9
- 20 -
steam feedstock, heating the oxidant s~ream, supe~heating
steam, raising steam in a boiler, preheating boiler
feedwater or a combination thereof.
The first step in the removal of carbon oxides is t~-e
conversion of carbon monoxide to carbon dioxide by the water
gas shift reaction
CO + ~20 ---> COz + 1-~2 ( 5)
in which carbon monoxide is reacted with water to produce
carbon dioxide and hydrogen. The water gas shift reaction
is known, and suitable equipment for carrying out the
reaction is commercially available. The water gas shift
reaction can be carried out in two stages, i.e. a high
temperature shift and a low temperature shift. In this
procedure, the synthesis gas is first reacted with water
vapor at a temperature in the range from about 580F (304C)
to 750F (400C) and a pressure in the range from about 10
atm to 40 atm (1000 KPa to 4000 KPa), followed by reaction
at a temperature in the range from about 350F (175C) to
500F (260C) and a pressure in the range from about 10 atm
to 40 atm (1000 to 4000 KPa). ~lternatively the water gas
shift reaction can be carried out in a single stage, low
temperature tubular, steam-raising reactor shift vessel. In
this procedure, the water vapor and synthesis gas are
reacted at a temperature in the range from about 350F
(175C) to 500F (260C) and a pressure in the range from
about 10 atm to 40 atm (1000 to 4000 ~Pa). The exit stream
from the water gas shift reaction zone has a carbon monoxide
content less than about 1.0%.
Essentially all of the remaining carbon monoxide can be
converted to carbon dioxide by catalytic selective
oxidation. In this procedure, the exit stream from t~e
water gas shift reaction zone, after heat removal to re<luce
its temperature to a range from about 100F (38C) to 250F
(120C), is reacted with air in the presence of a catalyst
200~Z~
- 21 -
that is highly selective for the oxidation of carbon
monoxide under conditions in which little or na hydrogen is
oxidized. The catalytic selective oxidation procedure is
known in the art and described by U.S. Patents No.
3,216,782, No. 3,216,783 and No. 3,631,073. Suitable
process equipment for carrying out the selective oxidation
procedure is commercially available for example, under the
trademark Selectoxo.
Alternatively to the catalytic selective oxidation
procedure, remaining carbon monoxide can be methanated using
known procedures. Since methanation reacts each mole of
carbon monoxide with three moles of hydrogen, this procedure
consumes some of the hydrogen product. This may be
acceptable, depending on the end use intended for the
product gas.
Any other procedures known in the art for removing
carbon monoxide can be employed to remove traces of carbo
monoxide from the gas stream.
After conversion of carbon monoxide to carbon dioxide,
carbon dioxide is removed from the gas stream and recovered
using known procedures such as, for example, passing it
through a countercurrent stream of a liquid absorbent
medium, such as potassium carbonate, MEA, DEA, etc., which
absorbs the carbon dioxide. Commercial processing units for
carbon dioxide removal are available for example, under the
trademarks Selexol, Amine Guard, and Benfield. These
processes absorb the carbon dioxide into a chemical or
physical absorption medium at relatively high pressure and
low temperature, allowing other gases to pass through
essentially unchanged. The chemical or physical absorbent
is then regenerated by pressure let down into a lower
pressure vessel and, if a chemical absorbent is used,
stripped of carbon dioxide by a counter current stream of
steam. The carbon dioxide gas is discharged from the top of
2004Z~ ~
- 22 -
the regenerator and the absorbent returned to the absorber
to recover more carbon dioxide.
Additional HYdroqen Enrichment
Additional steps may be employed to increase the
proportion of hydrogen in the product gas. In particular,
when air or oxygen-enriched air is used as an oxidant in the
catalytic partial oxidation step, the synthesis gas contains
nitrogen which can be removed from the product stream.
Any known procedure for removing nitrogen from a gas
stream can be employed. A preferred method for removing
nitrogen is pressure swing adsorption. Pressure swing
adsorption involves the adsorption of components to be
removed at high pressure followed by their desorption at low
pressure. The process operates on a repeated cycle having
two basic steps, adsorption and regeneration. Not all the
hydrogen is recovered as some is lost in the waste gas
during the regeneration stage. By careful selection of the
frequency and sequence of steps within the cycle, however,
the recovery of hydrogen can be maximized and the ratio of
hydrogen to nitrogen in the product effluent gas can be
strictly controlled to give the desired ratio.
Regeneration of the adsorbent is carried out in three
basic steps: (a) The adsorber i8 depressurized to the low
pressure. Some of the waste components are desorbed during
this step. (b) The adsorbent is purged at low pressure,
with the product hydrogen removing the remaining waste
components. (c) The adsorber is repressurized to adsorption
pressure ready for service.
The waste gases evolved during regeneration are
collected in a waste gas surge drum and then used as fuel.
Pressure swing absorption can also be used to remove
residual carbon dioxide, methane, water vapor and trace
contaminants such as H2S.
20~)4~L9
- 23 -
Another suitable method for removing nitrogerl from the
gas stream i8 by cryogenic s~paration, a procedure whereby
gases are fractionated accoriing to their liquification
temperatures. Commercially ~vailable cryogenic separators
can be employed to remove ni~rogen from the gas stream.
Using the procedures of the invention, a hydrogen
product stream can be obtain~d that contains at least 75
mol. ~ hydrogen, preferably at least 90 mol. % hydrogen, and
most preferably at least 95 mol. % hydrogen.
The process of the inv~tion is illustrated
schematically in a preferred embodiment in the flow diagram
of Fig. 6. ~Iydrocarbonaceous feedstock such as a natural
gas is fed, together with steam and oxygen-rich gas, to a
catalytic partial oxidation reactor 100 such as illustrated
in Fig. 1 ~ The resulting sy~thesis gas exits the catalytic
partial oxidation reactor 108 at a temperature from about
1600F (870C) to 1900F (10;0C) and is passed through a
heat exchanger at step 102 t~Ireduce its temperature to
between 350F (175C) and 75~F (400C). The gas is then
passed to a water gas shift _eaction zone 104 where carbon
monoxide and water vapor are reacted in contact with a shift
catalyst to convert carbon ~noxide to carbon dioxide and to
produce additional hydrogen. The quantity of gas contacted
with the shift catalyst will depend on the required quality
of the product gas stream. If the gas is intended for
reducing purposes only, for example, it may be acceptable to
eliminate the water gas shif t reaction step. The exit gas
from the water gas shift reastion zone 10~ is passed throug}
a heat exchanger at step 106 to reduce its temperature to
between 80F (27C) and 200~ (93C). The gas is then fed
to a carbon dioxide removal Init at step lG8, where the gas
stream is contacted with a c~untercurrent stream of carbon
dioxide absorbent, which abs~rbs the carbon dioxide froln the
gas stream. The hydrogen pr~duct stream may, if desired, be
200~X~9
- 24 -
sub~ected to further processing steps, for example, nitrogen
removal if air is employed in the oxidant stream of the
catalytic partial oxidation step 100.
When compared to present day commercial processes, the
process of the present invention employing the catalytic
partial oxidation route using, as oxidant, a stream
containing in excess of 70 mole percent of oxygen, as
described herein, offers the following advantages:
1. The high cost steam reforming furnace is eliminated
when compared to the conventional commercial route.
2. Low oxygen consumption when compared to
conventional partial oxidation.
3. Low water consumption when compared to steam
reforming.
4. Low cost when compared to catalytic partial
oxidation routes which use air or enriched air to
produce a nitrogen rich synthesis gas at the exit of
the catalytic partial oxidation reactor.
5. Reduced area requirement when compared to the steam
reforming route (particularly suitable for offshoLe
applications).
6. High efficiency when compared to conventional
commercial hydrogen prod4ction prosesses and when
compared to catalytic partial oxidation using air or
enriched air containing less than 70 mole percent of
oxygen.
7. Lower in capital cost than all present commercial
processes.
In another embodiment of a hydrogen producing process
in accordance with the invention as illustrated in Fig. 12,
a natural gas stream 200 is passed through coil 202 of a
fired heater 204 to preheat th,e hydrocarbonaceous feedstock.
The heater 204 is heated by combustion of waste fuel 206
with air stream 20~. From the heater 202, the feedstock is
200`~
- 25 -
fed in line 210 to desulphurization vessel 212. Feedstock
desulphurization depends on the quantity and tyl)es of
sulphur present in the feedstock. 1~ydrogell sul~>llide is most
economically removed simply by passing the feedstock over a
bed of zinc oxide absorbent. ~rhe hy~rogen sulphide present
reacts with the zinc oxide to produce zinc sulphide. Tlle
natural gas feedstock 200 is heated to a temperature in the
range from 400F (200C) to 750F (400C) and passed tl~rougl
the z~nc oxide absorbent contained in the desulphurization
vessel 212 to remove sulfur compounds.
The desulphurized natural gas feedstock 214 is mixed
with steam 216 to achieve a steam to carbon molar ratio in
the range from 1.0:1 to 1.5:1. The mixture 218 of steam and
feedstock is further heated in coil 220 of the heater 204 to
approximately 1100F (590C) and passed over line 222 to one
inlet of the catalytic partial oxidation (CP0) reactor 100
where the feedstock-steam mixture is thoroughly mixed with
oxygen 224 and fed to the catalytic oxidation zone as
described above. The main reactions taking place within the
CPO reactor 100 are the partial oxidation reactions:
C~H2~+2 + ~/2O2 ---> nCO + (n+l)H2 (6)
and the water gas shift reaction
C0 + 1~20 ---> CO2 + 112 ( S )
as described above. The exit temperature of the effluent
226 from the CPO reactor is approximately 1750F (954C).
The reactor effluent 226 is first cooled in CPO boiler
22~ by raising steam from water flow from steam drum 229,
before being passed tllrougll line 230 to water quencll
evaporator or desuperheater or cooler 233, and then through
line 234 to low temperature shift reactor 236. In order to
~:00~2~
- 26 -
avoid reducing conditions within the shift reactor 236,
which would result in catalyst deterioration by reduction o~
the shift catalyst and formation of carbon, and methanation
of carbon oxides, and also in order to improve carbon
nomoxide shift conversion, the steam to gas ratio is
increased upstream of t~e shift reactor by the injec~io-l of
water. Additional steam raising and the addition of steam
would be an alternative to water injec~ion.
The water gas shift reaction is exothermic and in orde~
to recover the heat of reaction effectively, the low
temperature shift reactor 236 is constructed as a tubular
steam-raising, isothermal reactor. The favorable low
reaction temperature within catalyst filled tubes 238 of th~
reactor 236 is maintained by raising steam from water flow
from the drum 229. Alternatively, the CPO reactor effluent,
which has a high concentration of carbon monoxide, could be
passed through a high temperature shift reactor, and the
resulting higher concentration of carbon monoxide could be
accepted or the high temperature shift could be followed by
a low temperature shift, or the CPO reactor effluent could
be passed through a multibed low temperature shift reactor
with interbed cooling.
The plant illustrated in Fig. 12, employs the common
steam drum 229 for the CPO boiler 228 and the low
temperature steam raising tubular shift reactor 236. Outpu~
240 of the steam drum 229 provides the steam 216 for mixing
with the feedstock as well as providing steam in line 242
which is passed through coil 244 of the heater 204 to
produce superheated steam 246 which can be used to drive
turbines, or be used for any other purpose.
Layout or other plant process considerations may cause
the employment of separate steam drums to be more cost
effective. Efficiency improvement may also be obtained by
using a separate steam drum for the CPO boiler and ralsing
20(~2i,9
- 27 -
steam at a higher pressure. The use of a high temperature
shift catalyst would also allow the exothermic heat of
reaction of the shift reaction to be recovered as higher
pressure steam. ~lternatively to the preheatirlg of air and
feedstock and the superheating of steam in the fired heater,
such preheatinq and superheatiplg can be carried out by heat
exchange with the CPO effluent downstream from the reactor
100 .
The output 250 of the lowl temperature shift reactor 236
is passed through heat exchanger 252 which heats boiler feed
water stream 254, through line 256 to heat exchanger 258
which preheats water stream 260, through line 262 to water
cooled condenser 264, and then through line 266 to knockout
drum 268 where condensed water 270 is separated from the gas
stream 272. The condensate 270 together with make-up
demineralized water 274 form the water stream 260 which,
after passing through the heater 264, passes through line
276 to deaerator 278 where air and dissolved gases are
removed. Steam input 280 to column section 282 of the
deaerator 278 assists in removal of dissolved gases. A
portion of the steam input 280 is supplied over line 284
from the steam line 240, and the remaining portion is
supplied from gas output 286 of blowdown drum 288 receiving
blowdown stream 290 from the steam drum 229. Deaerated
water 302 is raised in pressure by pump 304 to line 306
which forms the spray feed 232 and the boiler feed 254.
The separated gas stream 272 from the separator 268 is
passed to a hydrogen recovery or purification unit 292 ln
which hydrogen product stream 1294 is separated from was~e
gas stream 296 passing to surge drum 298 which supplies ~he
waste fuel 206 and excess waste fuel 300. The waste ~uel
will typically have a low calorific value which is not
suitable for combustion in a conventional burner, and so the
burner for heater 204 is an "adiabatic combustor" or a
2()0~Z3~
i
- 28 -
catalytic combustor with the flue gases heating the process
stream coils in a convection heat transfer duct. ~ fuel
pressure of 5 to 10 psig (34 to 6~ KPa) is suffioient for
the adiabatic combustor and the catalytic combustor.
Alternatively, natural gas is mixed with the was~e fuel and
a duct type burner with a natural gas pilot is used to
combust the mixture with air. The hydrogen recovery unit
292 is a pressure swing adsorption (PSA) unit, cryogenic
separation unit, or prism separation unit.
Pressure swing adsorption involves the adsorption of
components to be removed at high pressure followed by their
desorption at low pressure. The process operates on a
repeated cycle having two basic steps, adsorption and
regeneration. Not all the hydrogen is recovered as some is
lost in the waste gas during the regeneration stage. By
careful selection of the frequency and sequence of steps
within the cycle however, the recovery of hydrogen is
maximized. ~egeneration of the adsorbent is carried out in
three basic steps. First, the adsorber is depressurized to
the low pressure. Some of the waste components are desorbed
during this step. Second, the adsorbent is purged at low
pressure, with the product hydrogen removing the remaining
waste components. Thirdly, the adsorber is repressurized to
adsorption pressure ready for service. A typical pressure
swing adsorption system suitable for the present process is
available under the trademark "l~y Siv" from Union Carbide.
Such a system may include up to ten individual beds. At any
one time three of the vessels would be in the adsorption
stage and the other seven vessels would be in various stages
of regeneration.
A cryogenic separation unit first removes carbon
dioxide by conventional bulk carbon dioxide removal
techniques prior to cryogenic separation of hydrogen. Fina~
traces of carbon dioxide and water vapor must be removed in
zoo~ 9
- 29 -
a molecular sieve unit in order to prevent blockage of the
cold box by solidification of these components. From the
molecular sieves, the feed sas is cooled by returning
product hydrogen and fuel g2s streams to condense nitrogen,
methane and carbon monoxide before being let down in an
expansion turbine to approximately 30 psig (210 KPa~. ~ftel
reheating, by cooling the ir.let stream, the product hydrogen
is recompressed to the required pressure in a centrifugal
compressor. The condensed gases are let down to fuel gas
pressure and also used to cool the incoming feed gas before
being passed to the fired heater for use as fuel. The fuel
gas is available at approxi~ately 20 psig (140 KPa).
A prism separation unit makes use of the principle of
selective permeation throug}: membranes. Permeation of gas
molecules though a membrane is in two stages, the first
being dissolution into the ~smbrane structure and the second
being diffusion through the ~embrane. With a mixture of
gases, different components ~ill permeate at different
rates, even when the partial pressure driving forces are
equal. The hollow-fiber me~brane consists of a porous
asymmetric polymer substrate and a polymer coating of high
permeability. The selective permeation characteristics of
the system allow faster gases such as water vapor, hydrogen
and carbon dioxide to be se~arated from the slower gases
such as methane, nitrogen, argon, oxygen, carbon monoxide
and hydrocarbons. The faster gases permeate through the
hollow fiber and are removed a low pressure while the slowe~
gases concentrate upstream of the membrane and are removed
essentially at inlet pressure. ~s carbon dioxide and
hydrogen both permeate the ~smbrane at similar speeds, it is
necessary to remove carbon dioxide by a bulk carbon dioxide
removal method prior to passing the gases to the prism
separators. The operating pressure is too low to achieve
efficient hydrogen recovery and the feed gas must therefore
20~)4~
- 30 -
be compressed to approximately 575 psig (3960 KPa) before
being passed to the separators. The product from the first
separator stage is available a~ 100 psig (689 KPa) but the
purity is low. Recompression to 625 psig (4310 KPa) and a
second stage of prism separation is thereforelnecessary to
yield the final product at about 200 psig (1380 KPa).
The process of the invention can be fuLt~er understood
with reference to the follo~ing examples, which are intented
to illustrate the invention and are not intended to limit
its scope in any way.
EX~''PLE I
Natural gas is converted to synthesis gas in a
catalytic partial oxidation reactor of the construction
shown in Fig. 1. There are included nine catalyst discs 54,
each having a diameter of 30 inches (0.76m) and a thickness
of io inches (0.25m). The discs are formed from a honeycomb
monolith of cordierite material with a geometric surface
area of approximately 25 cm2/cm3. A high surface area
alumina layer is deposited on the cordierite to serve as a
support upon which finely dispersed catalytic metal
components are distended. The catalytic metal components
are approximately 50% by weight platinum and 50~ by weight
palladium. Space velocity of the catalyst is 97,000 hr.-
Natural gas (>95~ methane) is mixed with steam at
various steam-to-carbon molar ratios, heated and supplied
through 10-inch (25.4 cm) diameter inlet 66 at a presure of
400 psig ~2760 KPa). Oxygen containing gas is heated and
supplied through two 8-inch (20.3 cm) inlets 70 at a
pressure of 410 psig (2830 KPa). The diameter of the lower
portion 76 of the chamber 72 is 27 inches (0.68m) Witll tlle
diameter of the upper portion 74 being 36 inches (0.91m).
There are 261 tubes 80 having 0.5 inch (12.7mm) internal
diameters and having lengths of 20 inches (0.51m). Six
orifices B6 of 0.123-inch (3.12 mm) diametex are formed in
~0(~2~9
- 31 -
each tube with four of the orifices evenly spaced around
each tube at a distance of 4 inches (0.102m) nbove the lower
end of the tube and with the remaining two orifices ormed
opposite each other at a distance 6 inches (0.152m) above
the lower end of the tube. The bottom member 7~ has a
thickness of 5 inches (0.127m) and the passageway sections
84 are conical with upper diameters of 0.5 inches (12.7 mm)
and lower diameters of 1.75 inches (44.5mm). Pressures
within the chambers 68 and 72 are maintained at essentially
the inlet pressures.
The temperature of the mixed reactant gases is 1,100F
(593C). Fig. 7 shows oxygen consumption for the catalytic
partial oxidation process, as a function of steam-to-carbon
molar ratio, for reaction temperatures of 1,600F (870C),
1,750F (950C) and 1,900F (1040C) and an operating
pressue of 400 psig (2760 KPa). It can be seen from the
graph that oxygen consumption, expressed as oxygen-to-carbon
molar ratio, is relatively low for the process of the
invention as compared with present commercial partial
oxidation processes. The dashed line 25 in Fig. 8
represents the conditions under which carbon deposits are
formed.
Fig. 8 shows the molar ratio of hydrogen, as H2, to
carbon monoxide in the product as a function of
steam-to-carbon ratio for reaction temperatures of 1,600F
(870C), l,750F (950C) and 1,900F (1040C).
Figs. 9 and 10, respectively, show the amounts oE
methane and carbon dioxide, as volume ~, in the product as
Eunction of steam-to-carbon ratio for reaction temperatures
oE 1,600F (870C), 1,750F (950C) and 1,900F (1040C).
Fig. 11 shows the effective H2 production of the
process, expressed as total moles of H2 and carbon monoxide
in the product divided by total moles of 112 and carbon in
the feedstock.
2OOJ L~tZ~9
- 32 -
Example II
An example of the process performed by the hydrogen
plant of Fig. 12 is illustrated in the followin~ TABLES I,
II and III which contain mole/hour, mole percent, and
parameters of pressure, temperatllre, water/steam flow, and
heat transfer for the process. The molesthour are lb
molesthour (0.4563 kg moles/hour). The example employs
pressure swing adsorption of the hydrogen recovery unit 292
to produce the hydrogen product from the synthesis gas.
!
2004Z~
-- 33 --
O~ ~ O ~ O
~3 0U7 ~ : oo 0 ~~ ,. ,. o o~
0 0,~ ~~ 0 ~a~
O ~ 1 U~ l ~O ~C) 0 r ~ CO
E- 0~t~`I~ 0~D1'--1~ ~ ~O
o ~a) o ~3 ~ U~ ,~ 0
. . . .. . .. .
~1 0 3~` Or`~ ~I)co ~1
0, 0 ~ ~ O
0~` ~ ~ C`lU~
O ~ 0
r~ O O
.
O O
~ a~ 0
t~ O O
~_) O O
~ 0
O ~~ ~ 0
~O'D V~ ~
lC`.l~I N~ ~i o
Ul~
J C`l 0 0 0 a~ ~ 0
m u' ~ ~ ~ 0 a~ 0 u~ O
~ a) O O O O O O ~ ~ o
r~
~ o 0o o o o ~ ~
3, ....
l 0
o o u~ o u) o
~ 0 0
ta o o ~o t-- ~ v)
0 ~ ~ 0
~ ~ ~ ~ a:~ o ~u~
o
t~ o o r~ r~ ~ 0 ~ o ul O
0 0 ~ ~ a~ 1`
~ 0 v~
C`l U) ~D V
:~
d
Z 07 o ~ ~ o d ~ :~ ô ~o 0 ô ô
O Id O ~`3 ~ t~l ,1 t ~ o u~ ,~ o o ~ ~ o
~ ~ ~ U) ~ J~ V ~ V ~ ~-~
H (d d d d d d ::1 d ~rl d ~ d d d bO ~ d d ~1 d ~ ~ bO
O `- o ~ d --¢ `~
Z O ~ ¢ 1
2()0~
-- 34 --
O ~~1 ~~ O ~ a) ~ o
O ~
O
O
a) ~ '~
:~ O O
t~ O O
O
O O
O o O
-- ~ ¢ U~ ,~ O O ~ '~ ~
O O OO O O O O
W ~
J P' o ~ ~ ~ ~ ~ ~ co
-- ~ o o o o o o ~ ,~ o
.1 ~ ~ O
~ ,~ o
~ o o ~ I`CD a~ I` o eo u~
t~ o o Ul ~o U~ ~ o U~
U) ~Du~ 0
o c~
~ ~ 0 ~o O
a
H ~ C`~ u~ VV :~ 10 3 ~ o d ~`
u ~ ~ K ~' '' ~ ~ a
200~2~
~ ~ o ~ ~ ~ ,, ~ ~
a
K
~d
~, o o 0
z ~
0 ` '~ o
0 .
_ ~
-- Q)
~ 0 ~ o a0~ J
a
O O O 0 0~ 0 0 0 ~1: 0 ~ U~ ~ O O ~O O
o ~D 0 00 0 U~ ~ O ~1 t~ ~ o t~ ~ o
~I ~~ o~ o~ 1~ C ~ U~ t~
o Vl o o V~ C V~ oO V~ V~ a~
0 ~o o
O ~ ~ 7 0 3
o~ e . ~
~O~X~ 9
-- 36 --
-
~o
u~ d
d t~
0
E~ ~
v
~a
V p
Z;
0 g~
~ _
CO
Z ~ o
S~
z a
o ~ "~
P- ~
0
Ul ~
~ ~ o o CO o o o
m a ~1 0 " ~ u~ O O
~D ~ ~ O ~ O
~ ~ ~ O ~ ~ U)
a ~ o u~ v) o ~ o
~, O
~J O ~
~ ~D ~ O
V C~
0 ~ V
~ ~ O ~ ~ ~ V
Z ~ 3 0
O ~
0 0 d a d ~ J
I v ~`I 3 3 ~) ~ V ~
~a o o bO n~ r~ O
a ~ 0~
~oo~2~
- 37 -
Since many modifications, variations and changes in
detail may be made in the above described embodiments
without departing from the scope and spirit of the
invention, it is intended that all matter described in the
foregoing description and shown in the accompanying drawings
be interpreted as illustrative and not in a limiting sense.
