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1
2002.08.sin 2.3
R&D on Hydrogen Production by Autothermal Reforming
(Shinnen Satte Laboratory)
� Takashi Suzuki, Katsumi Miyamoto, Shuichi Kobayashi,
Noriyuki Aratani, Tomoyuki Yogo
1. R&D Objectives
The purpose of the present R&D is to develop high-efficiency hydrogen production technology
through reforming of naphtha fraction to kerosene fraction, including GTL oil, etc. Another aim is
to investigate the constituents that make up the raw material hydrocarbons and reforming
characteristics. The R&D is scheduled to take place over 5 years from JFY2000 to JFY2004. As
part of this R&D program, the Shinnen Satte Laboratory will endeavor to develop technological
processes involving the autothermal reforming method, to investigate the reforming
characteristics of naphtha fraction to kerosene fraction, including GTL oil, etc., and to develop
desulfurization catalyst and processes essential for removing sulfur compounds in everything
from naphtha to kerosene fraction.
2. R&D Contents
The contents of R&D conducted are summarized below.
• Design of autothermal reforming process
• Basic establishment of Autothermal Reforming - Fuel Processing System
• Development of autothermal reforming catalyst
• Investigation of hydrocarbon constituents in fuel and of autothermal reforming
characteristics
• Development of catalytic desulfurization process
3. R&D Results
3.1 Design of Autothermal Reforming Process
3.1.1 Basic Establishment of Autothermal Reforming - Fuel Processing System
An autothermal reforming - fuel processing system (hereinafter ATR-FPS) was investigated for
the purpose of evaluating autothermal reforming reactions. A flow diagram of the basic process
in autothermal reforming is presented in Figure 3.1.1. The ATR-FPS is comprised of a reform
reactor + CO shift reactor + CO selective remover. Air equivalent to 3 times the amount of
oxygen required for CO combustion after CO shift reaction (O2/CO = 1.5) is supplied at the inlet
to the CO selective remover. In addition, heat balance was determined from changes in
enthalpy arising from compositional changes, on the assumption that in reactions after CO shift
reaction, methanation reaction does not advance, and that discharge heat loss from the system
is zero.
2
Ra
w
mate
ria
l
Re
form
er
Sh
ift
CO
re
mo
va
l
Figure 3.1.1: Flow Diagram of Autothermal Reforming Process
3.1.2 Establishment of Oxygen/Carbon Ratio and Steam/Carbon Ratio
In determining the oxygen/carbon (hereinafter O2/C) ratio and the steam/carbon (hereinafter
S/C) ratio supplied to ATR-FPS, O2/C and S/C were adjusted by equilibrium calculation so that
the total of heat absorption/generation from each reactor and of heat exchange required for
reaching each reaction temperature becomes 0-1 [kJ/mol-crude oil] in the entire ATR-FPS with
autothermal reforming + CO shift + CO selective removal. System internal pressure is 1.033
kg/cm2 and each reactor is isothermal. Taking normal decane as an example, the relationship
between S/C and O2/C obtained by equilibrium calculation when the autothermal reforming
reaction temperature was set at 650°C is shown in Figure 3.1.2. In the figure, when O2/C is 0.45,
if S/C is set to 2.37, the heat balance of ATR-FPS as a whole becomes –0.36 kJ/mol-nC10.
In evaluating catalytic activity or comparing the autothermal reforming reactivity of each fuel oil,
standard conditions of O2/C and S/C must be set up. In making these settings, the following two
setting standards must be satisfied because with the current power generation system, with
PEMFC built in, unless the CO concentration in gas at the CO shift reactor outlet is less than 0.5
vol%, it will be difficult to have the CO concentration reach under 10 ppm at the CO selective
remover at a later stage. Another reason is that it is important to have the concentration of
hydrogen in generated gas as high as possible.
(1) CO concentration in CO shift outlet gas must be less than 0.5 vol%.
(2) O2/C must be as low as possible.
Taking normal decane as an example, Figure 3.1.3 shows the relationship between O2/C
obtained by equilibrium calculation and CO concentration at shift reaction outlet when the
autothermal reforming reaction temperature was set at 650°C. It was clarified that when O2/C
was set at 0.45, the two aforesaid standards are satisfied. It was also confirmed that similar
results obtain with raw material hydrocarbon other than normal decane.
3
To
tal h
eat
bala
nce (
kJ/m
ol-nC
10)
CO
concentr
ation (
vol%
) aft
er
shift
(250°C
)
CO concentration upper limit:
Figure 3.1.2: O2/C Ratio vs S/C Ratio
Obtained by Equilibrium
Calculation
(n-decane, 650°C)
Figure 3.1.3: O2/C Obtained by
Equilibrium Calculation vs
CO Concentration at Shift
Reaction Outlet
(n-decane, 650°C): CO
Concentration Upper Limit at CO
Selective Remover Inlet After
Shift (estimated value)
3.1.3 Investigation of Hydrogen Production on 1 Nm3/hr Scale
With normal decane equivalent to kerosene fraction taken as the raw material, with throughput
at 240 ml/hr, and with O2/C set at 0.45 and S/C at 2.37 as indicated in Section 3.1.2, hydrogen
production at 1 Nm3/hr was implemented using a fixed-bed, flow-type unit for evaluating the
activity of heat-resistant reactor equipment. Used as the catalyst was autothermal reforming
catalyst A (hereinafter ATR catalyst A), obtained as a result of investigation of partial oxidation
reaction and preliminary investigation of autothermal reforming conducted in the previous fiscal
year. Using for reactions the gas generated after autothermal reforming reaction, reforming
reactivity was evaluated from the gasification rate that indicates the percentage of conversion of
normal decane to CO, CO2 and C4 hydrocarbon or below, and from the C2-C4 hydrocarbon
selection rate that indicates the percentage of conversion to hydrocarbon of C4 or below.
Because CO is denatured to H2 of equivalent mol volume in the post-stage shift reactor of an
autothermal reformer, H2 + CO yield and H2 + CO production volume were used to evaluate
hydrogen production.
Shown in Figure 3.1.4 is the temperature distribution of catalytic layer in autothermal reform
reaction. Heat generation can be seen at the upper catalytic layer, and it was confirmed that
oxidation reaction progresses on the catalyst.
Next, the impact of reaction temperature was investigated. Gasification rate and C2-C4
hydrocarbon selection rate are shown in Figure 3.1.5, while hydrogen yield and hydrogen
production volume are shown in Figure 3.1.6. Gasification rate increases as temperature rises,
and with normal decane at around 700°C, gasification of 97% or more took place. It was also
recognized that when the reaction temperature rises, the types of hydrocarbon produced without
complete reforming tend to decrease. In this way it was confirmed that as the reaction
temperature rises, hydrogen yield and hydrogen production volume increase, gradually
approaching the equilibrium value.
4
Te
mp
era
ture
(°C
)
Upper catalyst layer Middle catalyst layer Lower catalyst layer
Time (hr)
Figure 3.1.4: Catalyst Layer Temperature Distribution (650°C) in
Autothermal Reforming Reaction at 1 Nm3/hr
Gasific
ation r
ate
(%
)
Reaction temperature (°C)
Gasification rate (%)
C2-C4 selection rate
C2-C
4 s
ele
ction r
ate
(%
)
(H2 +
CO
) yie
ld (
%)
(H2 +
CO
) P
roduction V
olu
me (
Nm
3/h
r)
Reaction temperature (°C)
Equilibrium value
Equilibrium value
Figure 3.1.5: Impact of Reaction
Temperature on
Gasification Rate and
C2-C4 Hydrocarbon
Selection Rate
Figure 3.1.6: Impact of Reaction
Temperature on Hydrogen
Yield and Hydrogen
Production Volume
3.2 Basic Design of Autothermal Reforming - Fuel Processing System
3.2.1 ATR-FPS Basic Design
(1) ATR-FPS outline and design conditions
ATR-FPS is a hydrogen production process that uses the autothermal reforming method. In the
latest investigation, two units each of shift and selective oxidation reactor were installed so that
CO shift or selective removal would be accomplished more securely. The ATR-FPS flow
process is shown in Figure 3.2.1.
5
Autothermal reform type hydrogen production
Reactor system
Flow sheet
Figure 3.2.1: ATR-FPS Flow Diagram
(2) Raw material specifications
In the latest investigation, the crude oil composition ratio used was
n-decane:n-butyl-cyclohexane:diethyl benzene (= 64:16:20 mol%); crude oil molecular formula
was C11.3H22.3; crude oil molecular volume was 158.55; crude oil density (25°C) was 756.6 kg/m3,
and crude oil flow volume was 380 ml/hr (= 287.6 g/hr = 1.813 gmol/hr). Steam and air supplied
for reforming were each set at O2/C = 0.45 and S/C = 2.44, the conditions for attaining heat
balance. In addition, as opposed to CO concentration at the shift reactor outlet, O2/CO was set
to 1.5 (3 times the theoretical ratio) at the first CO selective remover and to 0.15 at the second
reactor (1/10 the volume at the first reactor).
(3) Substance balance
Table 3.2.1 presents substance balance as calculated based on the conditions given in Section
3.2.1 (2). The following was also clarified.
Hydrogen was produced at approximately 950 Nl/hr from raw material kerosene at 380 ml/hr
(approx. 290 g/hr).
The composition of hydrogen-rich gas product was H2 (42.83 mol%): CH4 (0.015 mol%): CO2
(20.66 mol%): N2 (36.49 mol%), with the ratios of H2, CO2 and N2 at roughly 6:3:5.
6
Table 3.2.1: ATR-FPS Substance Balance
Kerosene + Steam + Air Kerosene
(4) Heat Balance
(a) Heat recovery system
Optimization of ATR-FPS heat recovery was attempted using pinch technology. It was
confirmed that all the heat required (e.g., raw material vaporization, steam vaporization)
could not be replenished by means of waste heat recovery. The heat balance obtained
is shown in Table 3.2.2 and heat recovery for ATR-FPS is given in Figure 3.2.2. As
Table 3.2.2 indicates, whereas the total of received heat is 962.2 kcal/hr, that of heat
given off is 1140.7 kcal/hr, and since the volume of heat given off is greater than that of
heat received, it is apparent that heating by an outside source is not required.
According to pinch technology shown in Figure in 3.2.2, however, when a temperature
differential (10°C) has been secured at pinch point (60°C), some 34.0 kcal/hr of heat,
equivalent to roughly 4% of the total heat volume from heating (962.2 kcal/hr), is
required from an external source.
7
Table 3.2.2: Heat Balance
Heat exchanger Inlet
temperature
(°C)
Outlet
temperature
(°C)
Heat
volume
(kcal/hr)
Crude oil vaporization 25 202 56.8
Steam vaporization 25 102 555.8
Fuel oil + steam +
heated air
102 450 349.6
Received
heat
Total 962.2
ATR to LTS-1 750 250 577.4
LTS-1 to LTS-2 320 200 126.6
LTS-2 to PROX-1 205 150 56.8
PROX-1 to PROX-2 192 150 43.9
From PROX-2 154 40 336.0
Emitted
heat
Total 1,140.7
Figure 3.2.2: ATR-FPS Heat Recovery
3.3 Development of Autothermal Reforming Catalyst
3.3.1 Experiment Method
A fixed-bed, flow-type micro reactor was used for reactions. Deep desulfurized kerosene (Sf <
0.5 ppm) was used as raw material oil. At prescribed volumes, calculated in accordance with
Section 3.1.2, deep desulfurized kerosene and water were supplied via fluid feed pump, and air
was supplied by thermal mass flow. Used as reactor was an electric oven heated to a
prescribed temperature. ATR catalyst A served as the standard catalyst. In the current fiscal
year, plans call for preparation of ATR catalysts B-E, to which tertiary ingredients have been
added to ATR catalyst A in order to have catalyst that offers greater steam reforming activity, in
an effort to achieve autothermal reforming at high activity. Carrier was obtained by kneading two
types of oxide with binder, and then after molding, by sintering at a prescribed temperature
under air flow. The carriers obtained were globular in shape, measuring 2-4 mm in diameter and
90-110 m2/g in surface area. For active metal retention, metallic salt water solution was
impregnated in obtained carrier, dried and adjusted. Prior to reaction, pretreatment took place
for 3 hrs at 700°C under hydrogen current. Used in evaluating the reaction was CO + CO2
selection rate, which gives the percentage of carbon in raw material hydrocarbons that has been
transformed into CO and CO2.
8
The gas obtained was analyzed using TCD or FID gas chromatogram; an oxygen combustion
type carbon analyzer was used to analyze the volume of carbon accumulated on catalyst after
the reaction.
3.3.2 Autothermal Reforming Reaction with Each Type of Catalyst
Figure 3.3.1 presents the results of a comparison of the dependency on reaction temperature of
autothermal reforming reaction activity with each type of catalyst. The catalysts can be arranged
in sequence as follows.
Catalyst B > Catalyst C > Catalyst A (standard) > Catalyst D > Catalyst E
At 750°C or above, a CO + CO2 selection rate of virtually 100% was exhibited with all the
catalysts.
The temperature distribution of each catalytic layer at this time is shown in Figure 3.3.2, with
600°C taken as sample reaction temperature. With all the catalysts, it was confirmed that the
temperature rises sharply near the inlet and declines as you go to the lower layers. Dissanayake
et al.1)
and Groote et al.2)
report that in autothermal reforming reaction with methane, a complete
oxidation reaction of methane takes place first, followed by water vapor reforming, CO2
reforming, and aqueous gas shift reaction. It is believed that in the latest results as well,
oxidation reaction advanced over the upper catalyst layer. When the catalysts are arranged in
order of highest temperature at the catalyst layer inlet, the following sequence obtains.
Catalyst E > Catalyst D > Catalyst C > Catalyst A (standard) > Catalyst B
This sequence manifests a trend virtually opposite that of the activity sequence. It shows that
when the catalyst inlet temperature is high (that is when the heat given off is great), the
complete oxidation reaction takes precedence over the reforming reaction (steam reforming
reaction or CO2 reforming reaction). In catalyst with low autothermal reforming activity, the
reforming reaction tends to advance less easily than the complete oxidation reaction.
Conversely, in catalyst with high autothermal reforming activity, the complete oxidation reaction
and reforming reaction both advance easily. The fact that the complete oxidation reaction and
reforming reaction take place at nearby locations on the catalyst suggests that the temperature
at catalyst inlet is kept relatively low. With catalyst A (standard) and catalyst C, however, this
trend is reversed, and the factors determining catalyst layer temperature distribution are not just
the catalyst’s complete oxidation activity or reforming activity. Such things as thermal
conductivity and filling density are also contributing factors.
In the development of autothermal reforming catalyst, it is not only high activity that is important
but also curtailment of carbon precipitation. Of the catalysts considered in Figure 3.3.1, ATR
catalyst A and ATR catalysts B and C, which manifest higher activity than ATR catalyst A, were
used in reactions that took place for 16 hrs at 600°C and the volumes of precipitated carbon
thereafter were compared as shown in
9
Figure 3.3.3. It was recognized that the volume of precipitated carbon with ATR catalyst B,
which exhibited the highest activity among the catalysts compared, was the smallest, and it
became evident that the volume of precipitated carbon with ATR catalyst A can be reduced by
about 35%. With ATR catalyst C, which manifested higher activity than ART catalyst A,
approximately 1.2 times greater carbon precipitation was noted as compared to ATR catalyst A.
This shows that escalation of activity by the addition of tertiary constituents does not always
match with curtailment of carbon precipitation. It is conjectured that in the design of autothermal
reforming catalyst, escalation of activity and curtailment of carbon precipitation must be
considered from separate standpoints.
CO
+ C
O2 s
ele
ction r
ate
(%
)
Catalyst layer outlet temperature (°C)
Catalyst A
Catalyst B
Catalyst C
Catalyst D
Catalyst E
(Deep desulfurized kerosene, LHSV = 1，O2/C = 0.45, S/C = 2.437)
Distance (cm) from catalyst layer outlet
Catalyst A
Catalyst B
Catalyst C
Catalyst D
Catalyst E C
ata
lyst
layer
tem
pera
ture
(°C
)
Outlet
Inlet
Figure 3.3.1: Comparison of
Autothermal Reforming
Reactivity with Each Type
of Catalyst
Figure 3.3.2: Catalyst Layer Temperature
Distribution in Autothermal
Reforming Reaction with
Each Type of Catalyst
Pre
cip
ita
ted
ca
rbo
n v
olu
me
(m
ass%
)
Catalyst A Catalyst B Catalyst C
Figure 3.3.3: Comparison of Precipitated Carbon Volume after
Reaction with Each Type of Catalyst (600°C, 16 hrs)
10
3.4 Investigation of Hydrocarbon Constituents in Fuel and of Autothermal Reforming
Characteristics
3.4.1 Experiment Method
A fixed-bed, flow-type micro reactor was used for reactions. The hydrocarbon compounds
shown in Table 3.4.1 were used as raw material, together with standard ATR catalyst A. So as
to compare the autothermal reforming reaction characteristics among each hydrocarbon, the
most ideal reaction temperature was one at which side reactions such as combustion or
decomposition could be suppressed as much as possible. Representative of each hydrocarbon,
deep desulfurized kerosene and normal hexane were taken as raw materials. Reaction was
initiated under the same conditions as for autothermal reforming reaction, using a micro-reactor
without catalyst, and the presence of side reactions in the reactor was confirmed. As shown in
Figure 3.4.1, because the C2-C4 hydrocarbon selection rate is zero at 600°C or below, it was
confirmed that no side reactions take place inside the reactor. Consequently, autothermal
reforming reaction was implemented at 600°C, where there would be no side reactions. Taking
the carbon mol flow volume included in the raw material hydrocarbons as standard, the rate was
0.71 mol/hr-C. O2/C and S/C were determined in accordance with Section 3.1.2. Reforming
reactivity was evaluated using a virtual speed constant k (CO + CO2) covering CO + CO2
production, determined on the assumption of a primary reaction and CO + CO2 selection rate
that correlates with the hydrogen selection rate targeted.
Table 3.4.1: Model Hydrocarbon Compounds Equivalent to Naphtha -
Kerosene Fractions Used
Hydrocarbon
count
n-P i-P O N A
LN 6 n-hexane 2,2
dimethylbutane
1-hexane cyclohexane benzene
m-xylene HN 8 n-octane 2,2,4-trimethylp
entane
1-octane ethylcyclohexane
ethylbenzene
10 n-decane n-butylcyclohexane diethylbenzeneKERO
Other n-dodecane,
n-hexadecane
1,2,4-trimethyl-
benzene
C2
-C4 s
ele
ctio
n r
ate
(%
)
Reaction temperature (°C)
(Reaction conditions: O2/C = 0.45, S/C heat balance conditions, LHSV = 1.0 to 1.6)
Desulfurized kerosene
Figure 3.4.1: Temperature Dependency vs Side Reaction in Reactor
11
3.4.2 Reforming Reactivity of Normal Paraffin Due to Differences in Carbon Number
Figure 3.4.2 presents a comparison of relative reforming activities for normal paraffin due to
differences in carbon number. It can be seen that as carbon number increases, relative
reforming activity declines. A comparison of precipitated carbon volume at this time is shown in
Figure 3.4.3. Because the precipitated carbon volume increased together with an increase in
carbon number, relative reforming activity and precipitated carbon volume exhibited similar
trends with ATR catalyst A. Figure 3.4.4 gives the selection rates of C2-C4 hydrocarbons
produced at this time. It can be seen that the C2-C4 hydrocarbon selection rates become greater
as the carbon number increases. This fact suggests that when the carbon chain becomes large,
even though it decomposes midway, unreformed hydrocarbon increases. What is more, the bulk
of unreformed hydrocarbons are olefins. That precipitated carbon volume increases together
with an increase in carbon number can be ascribed to the fact that unreformed olefins condense
on catalyst surface, making it easy for carbons to be formed.
Re
lative
activity k
(C
O +
CO
2)
/-
C number
C d
ep
ositio
n w
eig
ht
(ma
ss%
)
C number
Figure 3.4.2: Comparison of
Autothermal Reforming
Reactivity of Normal
Paraffin Due to
Differences in Carbon
Number
Figure 3.4.3: Comparison of Precipitated
Carbon Magnitude on
Catalyst Due to Differences
in Carbon Number
C2
-C4
se
lectio
n r
ate
(%
)
C number
Figure 3.4.4: Selection Rate of C2-C4 Hydrocarbon in Normal Paraffin
Due to Differences in Carbon Number
12
3.4.3 Reforming Reactivity of Model Hydrocarbons Equivalent to Gasoline Fraction
Figure 3.4.5 presents a comparison of relative reforming activity for each model hydrocarbon
compound of normal paraffin, isoparaffin, olefin, naphthene and aroma; the carbon number of
model hydrocarbon equivalent to gasoline fraction was taken as 8. A comparison of precipitated
carbon volume at this time is shown in Figure 3.4.6. In each case, the sequence of relative
reforming activity and precipitated carbon volume becomes as follows.
Isoparaffin = Naphthene = Normal paraffin > Aromatic > Olefin
It is believed that factors due to hydrocarbon structure play a large role in these sequences, but
the details will have to be further investigated in the future.
Re
lative
activity k
(C
O +
CO
2)
/-
C d
ep
ositio
n w
eig
ht
(ma
ss%
)
Figure 3.4.5: Comparison of
Autothermal Reforming
Reactivity Due to
Differences in
Hydrocarbon Type
(Carbon No.: 8)
Figure 3.4.6: Comparison of Precipitated
Carbon Volume on Catalyst
Due to Differences in
Hydrocarbon Type (Carbon
No.: 8)
13
3.4.4 Reforming Reactivity of Each Model Hydrocarbon Type Equivalent from Naphtha
to Kerosene Fraction
Respecting the reforming reactivity of each model hydrocarbon equivalent to fractions from
naphtha to kerosene, Figure 3.4.7 presents a comparison of relative reforming activity,
organized by carbon number, and Figure 3.4.8 gives a comparison of carbon deposition weight.
The relative reforming activity of normal paraffin is relatively high at low class but at high class, it
tends to become lower than that of other hydrocarbons. Regardless of the carbon number, the
reforming activity of naphthene was high in comparison to other hydrocarbons. With aromatic
compounds, a clear correlation with carbon number could not be confirmed. On the contrary,
structural factors such as substituent position or chain length are suspected. For isoparaffin and
olefin, the trends in relative reforming activity could not be determined. Carbon deposition
weight was small with naphthene and isoparaffin, but large with olefin. The deposited carbon
weight with normal paraffin was small in comparison to other hydrocarbons at low class, the
same as relative reforming activity, and large at high class. Among aromatic compounds, the
deposited carbon weight did not exhibit a clear correlation with carbon number, because of
substituent reactivity or structural factors such as electron polarization in aromatic rings.
Relative reforming activity also exhibited different trends. These results indicate that among the
hydrocarbons equivalent to fractions from naphtha to kerosene, the reforming reactivity is
dependent upon carbon number in some cases, as in normal paraffin, but this factor is not
adequate for explaining all hydrocarbons. To explore the details in greater depth, such things as
hydrocarbon physical properties, chemical properties and interactions with catalyst will have to
be further investigated.
Re
lative
activity k
(C
O +
CO
2)
/-
C number
C d
ep
ositio
n w
eig
ht
(ma
ss%
)
C number
n-paraffin
Naphthene
Aroma
Olefin
i-paraffin
n-paraffin
Naphthene
Aroma
Olefin
i-paraffin
Figure 3.4.7: Carbon No. vs Relative
Reactivity
Figure 3.4.8: Carbon No. vs Precipitated
Carbon Magnitude
3.5 Development of Catalytic Desulfurization Process
3.5.1 Experiment Method
Almost all the kerosene in circulation in the market contains a sulfur component at the level of
20-60 massppm. As a sample for desulfurization reaction, kerosene on the market at 50
massppm was prepared. Properties are indicated in Table 3.5.1.
14
At the oil refinery, the hydrodesulfurization reactor operates with hydrogen partial pressure at
2.0-3.0 MPa for light fraction and at a high pressure of 10.0 MPa or greater for heavy fraction.
Nevertheless, with the small-scale fuel cell power generation system, including power for
household use, atmospheric pressure must be considered in terms of the High-Pressure Gas
Control Law, the Electric Utility Law, and so on. In the present research, therefore, the following
two points were assumed for desulfurization reaction.
1. The raw material is kerosene on the market (with 50 massppm sulfur component).
2. The reaction pressure is atmospheric pressure.
The purpose of the initial investigation, therefore, was to determine desulfurization performance
of active metals at atmospheric pressure. The constituents and configurations of catalysts used
in evaluating activity are listed in Table 3.5.2. Desulfurization catalyst A is a regular
extrusion-mold-type hydrodesulfurization catalyst. Desulfurization catalyst B was prepared by
having precious metal, the active metal, retained in globular-shaped alumina, the carrier.
Desulfurization catalyst C was obtained by molding base metal and zinc oxide in
columnar-shaped tablets.
A microreactor was used in evaluating the hydrodesulfurization activity of these three catalysts.
Prior to the hydrodesulfurization reaction, desulfurization catalyst A filled in the reactor
underwent preliminary sulfurization for 2 hrs at 350°C via 5% hydrogen sulfide/hydrogen gas;
desulfurization catalysts B and C underwent hydrogenation pretreatment reduction for 2 hrs at
350°C under conditions of hydrogen flow.
3.5.2 Impact of Catalytically Active Metals
Figure 3.5.1 presents the results of an evaluation of the hydrodesulfurization activity of each
catalyst in kerosene on the market with the H2/oil ratio at 100 and the reaction temperature at
250°C. The desulfurization activity of these three catalysts with LHSV = 10 can be expressed in
relative terms from reaction speed constant. When the activity of desulfurization catalyst A is
taken as 100, that of desulfurization catalyst B becomes 250 and that of desulfurization catalyst
C, 773, revealing that the activity of desulfurization catalyst C is the highest.
Table 3.5.1: Properties of Raw Material Kerosene Used
IBP 152 Density g/cm3 0.791
10% 170 Aroma % 22.4
20% 177 Olefin % 0.2
30% 185
Composition
Saturation % 77.4
40% 192 Sulfur component massppm 50
50% 201
60% 210
70% 221
80% 234
90% 248
Distillation (°C)
EP 275
15
Table 3.5.2: List of Catalysts for Investigation of Kerosene Hydrodesulfurization
Catalyst Constituent Configuration
Catalyst A Hydrodesulfuization catalyst
on market
Extrusion molded products
Catalyst B Precious metals Globular shape
Catalyst C Base metals Columnar shaped tablets
Re
lative
activity v
alu
e
Catalyst A Catalyst B Catalyst C
Figure 3.5.1: Evaluation Results for Hydrodesulfurization Activity
3.5.3 Impact of Catalyst Adjustment Conditions
Desulfurization catalyst C is catalyst in which base metal components and metal oxides have
been molded into columnar tablets. In order to investigate the impact of catalyst preparation
conditions on hydrodesulfurization activity, four types of catalyst (Table 3.5.3) with different
active metal load weight were prepared by means of the impregnation and co-precipitation
methods. Using these catalysts, tests to evaluate activity were conducted with LHSV = 0.25 and
reaction temperature at 300°C so as to clarify initial deterioration in activity. The results are
shown in Figure 3.5.2. The performance of catalyst prepared by the impregnation method was
such that the Sp value exceeded 0.2 massppm in catalyst with low metal retention magnitude
for about 100 hrs and in catalyst with high metal load weight, for about 450 hrs. In catalyst
prepared by the co-precipitation method, on the other hand, the Sp value did not exceed 0.2
massppm for up to 500 hrs irrespective of the active metal magnitude. In order to determine the
amount of hydrogen required for reaction in the catalyst system, an investigation was made in
which the hydrogen supply volume was modified over 100 hrs after reaction startup. The results
are shown in Figure 3.5.3. With the H2/oil ratio at 50 or above, the Sp value was low irrespective
of catalyst preparation method, but when the ratio fell below 50, it was found that the catalyst’s
desulfurization activity drops sharply. Given this fact, it is conjectured that the H2/oil ratio must
be at least 50 in this catalyst system.
Table 3.5.3: List of C-type Catalysts
Catalyst name Preparation method Metal load weight
Catalyst C-1 Impregnation Low
Catalyst C-2 Impregnation High
Catalyst C-3 Co-precipitation Low
Catalyst C-4 Co-precipitation High
16
Catalyst C-1 Catalyst C-2 Catalyst C-3 Catalyst C-4
Figure 3.5.2: Evaluative Tests of Hydrodesulfurization Reaction Life
Catalyst C-2 Catalyst C-4
Figure 3.5.3: Impact of H2/Oil Ratio on Hydrogenation Desulfurization
Reaction
4. Synopsis
4.1 Design of Autothermal Reforming Process
(1) Basic Establishment of Autothermal Reforming Type Hydrogen Production System
An autothermal reforming - fuel processing system (ATR-FPS) was set up from supply of
raw materials to after selective removal of CO. On the assumption that there is no
discharge heat loss, a method was established for calculating reaction conditions so that
the system as a whole achieves heat balance.
(2) Investigation of Hydrogen Production on 1 Nm3/hr Scale
Hydrogen production on the 1 Nm3/hr scale by the autothermal reforming method was
confirmed, as was the impact of reaction temperature on the autothermal reforming
reaction.
4.2 Basic Design of Autothermal Reforming - Fuel Processing System
The basic design of the autothermal reforming - fuel processing system (ATR-FPS) was
completed in consideration of raw material specifications, substance balance and heat balance.
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4.3 Development of Autothermal Reforming Catalyst
Autothermal reforming catalyst B, superior to the current autothermal reforming catalyst A in
activity and in coking resistance, was discovered.
4.4 Investigation of Hydrocarbon Constituents in Fuel and of Autothermal Reforming
Characteristics
Comparisons were made of the autothermal reforming characteristics of typical constituents
comprised of naphtha to kerosene fraction. It was confirmed that the reforming reactivity is
higher, the lower the class of hydrocarbon. In addition, with naphtha fraction, the reforming
reactivity of saturated hydrocarbon is the highest, followed in sequence by that of aroma and
olefin.
4.5 Development of Catalytic Desulfurization Process
The reaction conditions required for hydrodesulfurization were established, and desulfurization
catalyst B-2 was developed. This catalyst exhibits high activity such that the sulfur concentration
in produced oil is 0.2 massppm or less. Service life evaluation tests confirmed that the durability
of catalyst B-2 is 500 hrs.
5. Bibliography
1. D. Dissanayake et al., J. Catal., 132 (1991) 117
2. A. M. D. Groote et al., Appl. Catal. A, 138 (1996) 245
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