Synthesis of Fe–Co–Ce/Zeolite A-3 Catalysts and their Selectivity to Light Olefins for Fischer–Tropsch Synthesis in Fixed-Bed Reactor

Article abstract of DOI:10.1007/s10562-018-2647-2

Abstract: In the present work, Fe–Co–Ce/Zeolite A-3 catalysts were prepared, using the impregnation method and applied in Fischer–Tropsch Synthesis (FTS). The catalyst’s performance was investigated by changing the catalyst’s composition and process conditions. Increased pressure from 2 to 10 and H₂/CO ratio from 1 to 4 were accompanied by decreased olefin to paraffin proportion and increased C4+ production. Decline in the GHSV led to decrease in methane and light olefin generation, and a significant increase in the number of carbon in the products. The catalysts were characterized by SEM, BET, STA and XRD. Catalyst composition was changed with different loading of support (Zeolite A-3) and metals (Fe, Co, Ce). Finally, 45%Fe–45%Co–10%Ce/60%Zeolite A-3 catalyst was judged to be the best composition to obtain the optimal operation conditions (high selectivity towards light olefins and low selectivity towards methane, simultaneously). Graphical Abstract: [Figure not available: see fulltext.].

Full text of DOI:10.1007/s10562-018-2647-2

Catalysis Letters
https://doi.org/10.1007/s10562-018-2647-2
Synthesis of Fe–Co–Ce/Zeolite A-3 Catalysts and their Selectivity
to Light Olefins for Fischer–Tropsch Synthesis in Fixed-Bed Reactor
Reza Nohtani1 · Ali Akbar Mirzaei1 · Amir Eshraghi1
Received: 11 November 2018 / Accepted: 13 December 2018
© Springer Science+Business Media, LLC, part of Springer Nature 2019
Abstract
In the present work, Fe–Co–Ce/Zeolite A-3 catalysts were prepared, using the impregnation method and applied in Fischer–
Tropsch Synthesis (FTS). The catalyst’s performance was investigated by changing the catalyst’s composition and process
conditions. Increased pressure from 2 to 10 and H2/CO ratio from 1 to 4 were accompanied by decreased olefin to paraffin
proportion and increased C+ production. Decline in the GHSV led to decrease in methane and light olefin generation, and a
4
significant increase in the number of carbon in the products. The catalysts were characterized by SEM, BET, STA and XRD.
Catalyst composition was changed with different loading of support (Zeolite A-3) and metals (Fe, Co, Ce). Finally, 45%Fe–
45%Co–10%Ce/60%Zeolite A-3 catalyst was judged to be the best composition to obtain the optimal operation conditions
(high selectivity towards light olefins and low selectivity towards methane, simultaneously).
Graphical Abstract
Keywords Catalyst composition · Reaction conditions · FT reaction · Light olefins selectivity
1 Introduction
* Ali Akbar Mirzaei
mirzaei@hamoon.usb.ac.ir
The abundant human use of non-renewable energy sources
such as coal, petroleum and natural gas has led to con-
cerns about environmental change, global warming and
1
Department of Chemistry, Faculty of Sciences, University
of Sistan and Baluchestan, 98135-674 Zahedan, Iran
Vol.:(0123456789)
1 3

R. Nohtani et al.
living organisms. The problem has made renewable energy
an attractive topic and better supported by governments.
On the other hand, these non-renewable resources will be
exhausted in the near future. So, it is necessary to use
renewable sources and green fuels [1]. Synthesis gas
produced by biomass gasification is used as feed in FT
reaction and for producing a variety of fuels and valuable
chemicals. FT fuel is environmentally friendlier than fuel
obtained from crude oil [1–4]. In FTS, syngas (mixture of
H2 + CO) is converted to a wide range of products such
as alkanes, alkenes, alcohols, aldehydes and ketones [1,
5, 6]. Since this is a catalytic process, for optimization,
the properties of the catalysts such as physio-chemical
and mechanical, as well as particle size and shape must
be taken into consideration [7]. In order to control FTS
towards desired products, many studies have been carried
out. This action can be performed by choosing proper reac-
tor, catalyst composition and operating conditions [8–11].
In general, FT commercial reactors can be divided into
three basic classes: FT slurry reactor, fluid bed and fixed
bed [12]. The most common catalysts are used for the
FTS of group 8 metals such as iron, cobalt, ruthenium
and nickel [11, 13, 14]. Among the catalysts for com-
mercial applications, are those based on iron and cobalt
[9, 15, 16]. The effects of Ce as promoter over Co and
Fe-based catalysts on activity and selectivity were stud-
ied [17–22]. Fatima Pardo-Tarifa et al. [21] reported that
the presence of Ce causes decrease in Co-based catalyst,
reduction in temperature and improvement of the catalyst
activity. Wen-Ping Ma et al. [19] also reported that Ce pro-
moter improved the Co-based catalyst activity with high
methane selectivity. In the study of Sun et al. [22], it was
reported that addition of Ce to Fe-based catalyst resulted
in improved activity and selectivity to C+5 hydrocarbons.
Sergio-Rojas et al. [17] in their studies on Fe-based cata-
lysts, it was shown that Fe catalyst modified with Ce as
a promoter activated under CO continuous flow, subse-
quently, sped up the stabilization active intermediate of
carbon compounds. Generally, to achieve the thermal and
mechanical stability, high surface area, suitable disper-
sion of metals, etc. the presence of support is necessary
[23, 24]. Traditional porous supports are Al2O3, SiO2 or
TiO2 [25, 26]. Although zeolite has been used as a sup-
port in FT catalysts for fabrication of long chain liquid
fuels [27, 28], it seems reasonable that by controlling
the operation condition (for example low pressure), the
selectivity of light olefins is considered as a goal. The
aim of this study was to investigate catalyst composition
(metals and support loading) and process conditions effect
on the catalyst activity. The structural investigation was
done by applying Simultaneous Thermal Analysis (STA),
Brunauer–Emmett–Teller (BET) surface area measure-
ments, X-ray diffraction (XRD) and scanning electron
microscopy (SEM).
2 Methods and Materials
2.1 Catalyst Synthesis
In this work, an impregnation method was used for sup-
ported Fe–Co–Ce/Zeolite A-3 catalyst synthesis. To obtain
the desired catalyst composition, the required values of
Cobalt (II) nitrate hexahydrate (purity of 97% Merck), Iron
(III) nitrate nonahydrate (purity of 98% Merck) and Cerium
(III) nitrate hexahydrate (purity of 98.5% Merck) were mixed
in minimum amount of distilled water. The meshed (No.
150) Zeolite A-3 support was calcined for 6 h at 600 °C in
air before impregnation. The solution was prepared from the
metals that were directly impregnated onto calcined Zeolite
A-3 and it was dried for 6 h in an oven at 120 °C. Finally,
in order to perform the calcination, the dried catalyst was
placed in the furnace at 600 °C for 6 h.
2.2 Characterization of Catalysts
2.2.1 X-ray Diffraction (XRD)
X-ray diffraction (XRD) technique was used to identify
the catalyst phases. In order to perform this analysis, an
ADVANCE D8 device (made in Germany) and Cu–Ka radi-
ation were used. Catalyst phases were identified by matching
the experimental patterns with standard samples diffraction
patterns.
2.2.2 BET Measurements
The BET method was used to specify the surface area of the
catalyst samples. BET measurements were taken by using
Quantachrome Nova 2000 (made in USA). Prior to the anal-
ysis of BET, all the catalyst samples (precursor, calcined
before and after the reaction) were placed under nitrogen
atmosphere for 4 h at a temperature of 300 °C.
2.2.3 Simultaneous Thermal Analysis (STA)
To determine catalyst precursor mass changes caused by
temperature changes, STA technique was used. This tech-
nique includes two simultaneous tests: Thermal Gravimet-
ric Analysis (TGA) and Differential Scanning Calorimetry
(DSC) tests. The TGA was used to indicate sample weight
variations and DSC for specify endothermic or exothermic
change in sample’s mass. In this work, Rheometric Scientific
STA 1500+ was used for simultaneous TGA and DSC.
1 3

Synthesis of Fe–Co–Ce/Zeolite A-3 Catalysts and their Selectivity to Light Olefins for Fischer–…
29
2.2.4 Scanning Electron Microscopy (SEM)
28.5
The morphology of catalyst was observed using A Cam-
28
bridge S-360 scanning electron microscope device.
27.5
27
2.3 Catalytic Tests
26.5
A fixed bed reactor was used to perform catalytic tests.
26
Details and setup of the reactor system are provided in
25.5
previous work [29]. Briefly, in each experiment, one gram
25
of calcined catalyst was put into the reactor. Prior to the
5.5
7
8.5
10
11.5
13
13.5
15
start of reaction, the fresh catalyst was reduced in pure
t (h)
hydrogen (H2) stream for 14 h (T = 400 °C, P = 2 bar,
flow rate = 37.5 ml/min). These reduced conditions were
Fig. 1 Effects of contact time on CO conversion by Fe–Co–Ce/Zeo-
used for all samples in the next sections. The FT catalytic
tests were carried out under the following operating con-
ditions: T = 300–360 °C, P = 2–10 bar, gas hourly space
velocity (GHSV) = 1800–4200 h−1 and molar ratio of H2/
CO=1/1–4/1. To calculate CO conversion and selectivity,
Eqs. 1 and 2 were used, respectively.
lite A-3 catalyst (catalyst 45%Fe–45%Co–10%Ce/70%Zeolite A-3,
P=2 bar, T=320 °C, molar ratio of H2/CO=2/1,GHSV=3600 h−1
)
Catalyst 45%Fe–45%Co–10%Ce/70%Zeolite A-3 (No.
2), catalyst 42.5%Fe–42.5%Co–15%Ce/70%Zeolite A-3
(No. 3) and catalyst 40%Fe–40%Co–20%Ce–70%Zeolite
A-3 (No. 4) under the same reduction and operating con-
ditions (P = 2 bar, T = 320 °C, molar ratio of H2/CO = 2/1
and GHSV = 3600 h−1) were exposed to FT reaction. The
optimum amount of cerium was chosen taking into account,
increased selectivity of light olefins or olefin to paraffin ratio
and suppression of methane formation. The selectivity of
catalysts (No. 1–No. 4) towards light olefins was 24.86,
29.93, 27.10, 24.98%, and methane was calculated as 47.11,
43.46, 51.08 and 49.48%, respectively. Catalyst number 2
had the highest selectivity for light olefins (29.93%) and the
least selectivity was for methane (43.46%). Therefore, the
molar ratio of 10% of cerium was selected as the optimal
ratio. The findings of BET for all the calcined catalysts are
represented in Table 1. As indicated in the table, the BET
measurements’ results such as pore volume and surface area
for catalysts with various Ce loading were different. It is
clear that these features are affected by increase in the Ce
molar ratio. High activity of No. 2 catalyst can be explained
by surface areas data. Between catalysts, No. 2 catalyst had
more surface area; that was why 10% of cerium was chosen
as an optimum molar ratio.
Mol of COin − Mol of COout
CO Conversion (% ) =
× 100
Mol of COin
(1)
niCi
Selectivity (% ) =
× 100
∑
(2)
niCi
ni is carbon number of hydrocarbon i.
3 Results and Discussion
At the beginning of the reaction, all the catalyst’s active sites
were available, so, the catalyst had its highest activity. In this
state, the catalyst’s behavior was not stable and shifted to a
near constant conversion [30–39]. Therefore, the data could
not be provided. To find the best time to record data after the
beginning of the reaction, CO convertion was calculated at
different times. As shown in Fig. 1, after 10 h, the catalyst
reached a steady state and the CO conversion was almost
constant. 10 h was considered as the optimum time and all
the data in this work were recorded 10 h after the beginning
of the reaction.
3.1.2 Effect of Fe/Co Molar Ratio on the Catalyst
Performance
3.1 Catalyst Composition Impact on the Light
Olefins Selectivity
Like the previous step, Fe–Co ratio was changed in the
optimal amount of cerium obtained in the previous section.
Catalysts’ performances are indicated in Fig. 3. In the pre-
vious section, optimum molar ratio of cerium (10%) was
chosen. Now, the optimum amount of Fe–Co is sought.
FT reaction was carried out under similar process condi-
tions (P = 2 bar, T = 320 °C, molar ratio of H2/CO = 2/1
3.1.1 Effect of Ce Content
In this step, Ce ratio in catalysts preparation was
changed. The selectivity plot of catalysts with different
amounts of cerium is presented in Fig. 2. First, the cata-
lyst 47.5%Fe–47.5%Co–5%Ce/70%Zeolite A-3 (No. 1),
1 3

R. Nohtani et al.
CH4
C2H4
C3H6
C4+
(No. 1) 47.5%Fe-47.5%Co-5%Ce/70%Zeolite A-3
C2H6
(No. 2) 45%Fe-45%Co-10%Ce/70%Zeolite A-3
(No. 3) 42.5%Fe-42.5%Co-15%Ce/70%Zeolite A-3
(No. 4) 40%Fe-40%Co-20%Ce-70%Zeolite A-3
C3H8
CO converꢀon
60
50
40
30
20
10
0
N O . 1
N O . 2
N O . 3
N O . 4
Fig. 2 Effects of Ce content on the catalytic performance (P=2 bar, T=320 °C, molar ratio of H2/CO=2/1, GHSV=3600 h−1
)
Table 1 BET surface area measurements for catalysts with different
loading. The actual detected phases for them were Fe3O4
(cubic-orthorhombic), Ce2O3 (hexagonal), Fe2O3 (rhombo-
hedral), Co3O4 (cubic) and Ce3O4 (hexagonal). Although,
the phases observed were similar, the relative intensity of
the peaks was different.
Ce contents
Calcined catalysts
Surface
areas
Pore
volume
(m2/g)
(cm3/g)
47.5%Fe–47.5%Co–5%Ce/70%Zeolite A-3
No. 1
45%Fe–45%Co–10%Ce/70%Zeolite A-3
No. 2
42.5%Fe–42.5%Co–15%Ce/70%Zeolite A-3
No. 3
40%Fe–40%Co–20%Ce–70%Zeolite A-3
No. 4
3.1.3 Effect of Support Loading
50.14
67.26
60.73
58.42
0.017
0.025
0.020
0.020
In addition to metals, support properties such as surface area
and distribution, volume and diameter of pores can influence
the catalyst activity and selectivity [40, 41]. In this work,
Zeolite A-3 was used as catalyst support. After optimizing
the molar ratio of metals, the support loading impact on
the performance of 45%Fe–45%Co–10%Ce/xZeolite A-3
catalyst in the synthesis of FT for the production of light
olefin was investigated. Different amounts of Zeolite A-3
(50 [No. 10], 60 [No. 11], 70 [No. 12], 80 [No. 13] and
90% [No. 14]) were loaded in catalyst preparation. All cata-
lysts were reduced and FT reaction was started under the
following conditions (P=2 bar, T=320 °C, molar ratio of
H2/CO = 2/1 and GHSV = 3600 h−1). The results are dis-
played in Fig. 5. The obtained results show that decrease
in the amount of support from 90 to 50% did not signifi-
cantly change CO conversion (about 25.2–27.3%), while
selectivity for light olefins increased from 9.82 to 28.95%.
Among the amounts of support loading, 60% support had
maximum selectivity to light olefins. Finally, catalyst with
optimum amounts of metals and support were determined.
The 45%Fe–45%Co–10%Ce/60%Zeolite A-3 catalyst
and GHSV = 3600 h−1) on the catalysts with different Fe/
Co ratios. Catalysts: 15%Fe–75%Co–10%Ce/70%Zeolite
A-3 (No. 5), 30%Fe–60%Co–10%Ce/70%Zeolite A-3
(No. 6), 45%Fe–45%Co–10%Ce/70%Zeolite A-3 (No.
7), 60%Fe–30%Co–10%Ce/70%Zeolite A-3 (No. 8) and
75%Fe–15%Co–10%Ce/70% Zeolite A-3 (No. 9) were
shown to have different behaviors. The results are shown
in Fig. 3 increase in molar ratio of iron from 15 to 75%
increased methane production from 37.78 to 46.43% and
C4+ production decreased from 13.88 to 2%. Finally, the ratio
of 45%Fe–45%Co was selected as Fe–Co optimum ratio.
This selection is due to the maximum selectivity of light
olefins and minimum of C4+ production. Figure 4 shows XRD
patterns for calcined catalysts, including different Fe–Co
1 3

Synthesis of Fe–Co–Ce/Zeolite A-3 Catalysts and their Selectivity to Light Olefins for Fischer–…
Fig. 3 Effect of Fe/Co molar
ratio on the catalytic perfor-
mance (P=2 bar, T=320 °C,
molar ratio of H2/CO=2/1 and
CH4
C2H4
C3H6
C4+
15%Fe-75%Co-10%Ce/70%Zeolite A-3 (No.5)
30%Fe-60%Co-10%Ce/70%Zeolite A-3 (No.6)
45%Fe-45%Co-10%Ce/70%Zeolite A-3 (No.7)
60%Fe-30%Co-10%Ce/70%Zeolite A-3 (No.8)
75%Fe-15%Co-10%Ce/70% Zeolite A-3 (No.9)
C2H6
C3H8
GHSV=3600 h−1
)
CO conversion
50
45
40
35
30
25
20
15
10
5
0
N O . 5
N O . 6
N O . 7
N O . 8
N O . 9
was considered as the optimum catalyst for the following
sections.
the undesired product are high, and also, olefins production
is low. In addition to this, at high temperature, the catalyst
can be deactivated by sintering, coke formation and car-
bon deposition [46, 47]. T=340 °C was chosen as optimum
temperature because of high selectivity to light olefins.
The catalysts were characterized after reaction by SEM.
SEM images of used catalysts at T=340, 350, 360 °C are
presented in Fig. 6. The images show that the catalyst at
T=340 °C (Fig. 6a), T=350 °C (Fig. 6b) and T=360 °C
(Fig. 6c) have different morphologies. These images result
show that at T=350 and 360 °C, there are some agglomera-
tions in used catalysts which may be due to sintering during
the FT reaction, and coke formation.
3.2 Impact of Operational Conditions
In addition to catalyst composition, reaction conditions play
an important role on activity and performance of catalyst in
FT reaction. In order to find the optimum operating condi-
tion, reaction conditions such as T, P, GHSV and feed ratio
were varied in FT reaction on the optimum catalyst.
3.2.1 The Temperature Impact
Generally, increase in temperature enhances the catalyst per-
formance but too high or low temperatures are not appro-
priate for the reaction [42]. To study the effects of reac-
tion temperature, the optimum catalyst was investigated at
300–360 °C under same situations (P=2 bar, molar ratio of
H2/CO=2/1 and GHSV=3600 h−1). The results in Table 2
show that, by increasing the temperature from 300 to 360 °C,
CO conversion significantly increased from 24.57 to 38.34%.
This result is in agreement with [43–45]. At T = 300 °C,
methane formation (35.77%) was low and both CO conver-
sion (24.57%) and light olefins production (17.3%) were low.
This can be because of absence of adequate energy at low
temperatures to activate the reactant on the catalyst [29].
In other words, it can be said that a significant amount of
the reactants did not react with each other. At temperatures
340 and 350 °C, CO conversion and methane formation as
3.2.2 Effect of Gas Hourly Space Velocity (GHSV)
The residence time has an inverse relationship with GHSV,
as the residence time decreases with increase in GHSV.
The GHSV impact on the efficacy of optimum catalyst was
assessed under the same conditions (P=2 bar, T=340 °C,
molar ratio of H2/CO = 2/1) and variable GHSV = 1800,
2400, 3000, 3600 and 4200 h−1. According to the data
presented in Table 3, with a GHSV increase of 1800 to
4200 h−1, production of heavy products decreased, due to
reduction of the residence time. This result is in agreement
with that of [48, 49]. In other words, decrease in GHSV
causes a significant increase in the number of carbon in
the products. On the other hand, light olefins production
increased from 26.56 to 32.91%. It also increased methane
1 3

R. Nohtani et al.
production from 39.19 to 54.69%. This behavior is expected
due to less residence time for the feed and products over the
catalyst. At low GHSV, feed (H2 +CO) had enough time for
adsorption and reaction on the catalyst. Very low GHSV
is not suitable because of the presence of external mass
transfer limitation [50]. GHSV = 3600 h−1 was chosen as
the optimum GHSV due to high production of light olefins
(32.17%).
3.2.3 Effect of H2/CO Molar Ratio
To evaluate H2/CO molar ratio impact on the optimum cata-
lyst selectivity and CO conversion, FT reaction was per-
formed at T=340 °C, P=2 bar, GHSV=3600 h−1 and the
variable H2/CO molar ratio from 1 to 4. Table 4 shows the
effect of change in H2/CO molar ratio on optimum catalyst
selectivity and CO conversion. These results demonstrated
that by changing H2/CO molar proportion of 1–4, different
selectivity towards any products and CO conversion were
obtained. Table 4 shows that increase in the ratio of H2/
CO increased CO conversion, while selectivity for light ole-
fins decreased. This behavior is in agreement with [51–53].
However, H2/CO=1 was considered as optimum molar ratio
because of maximum selectivity to light olefin and minimum
selectivity to CH4.
3.2.4 Impact of Reaction Pressure on the Olefin to Paraffin
Ratio
It is well known that total pressure plays pivotal role in FT
reactions, and affects the catalyst performance directly.
Increment in pressure has a positive impact on the catalyst
selectivity [49]. In the current study, in order to investigate
Fig. 4 XRD patterns of calcined catalysts No. 5, No. 6, No. 7, No. 8,
No. 9
Fig. 5 Product selectivity and
CH4
C2H4
C3H8
C2H6
C4+
45%Fe-45%Co-10%Ce/50%Zeolite A-3 (No.10)
45%Fe-45%Co-10%Ce/60%Zeolite A-3 (No.11)
45%Fe-45%Co-10%Ce/70%Zeolite A-3 (No.12)
45%Fe-45%Co-10%Ce/80%Zeolite A-3 (No.13)
45%Fe-45%Co-10%Ce/90%Zeolite A-3 (No.14)
CO conversion of 45%Fe–
45%Co–10%Ce catalyst with
different support loadings
(P=2 bar, T=320 °C, molar
ratio of H2/CO=2/1 and
C3H6
CO conversion
45
40
35
30
25
20
15
10
5
GHSV=3600 h−1
)
0
N O . 1 0
N O . 1 1
N O . 1 2
N O . 1 3
N O . 1 4
1 3

Synthesis of Fe–Co–Ce/Zeolite A-3 Catalysts and their Selectivity to Light Olefins for Fischer–…
Table 2 Temperature
T (°C)
Selectivity (%)
(%) CO
Olefin/paraffin
Effect on the selectivity
of 45%Fe–45%Co–
conversion
C+
CH4
C2H4
C2H6
C3H6
C3H8
4
10%Ce/60%Zeolite A-3
catalyst (T=300–360 °C,
P=2 bar, H2/CO=2/1 and
300
310
320
330
340
350
360
35.77
36.8
38.09
38.91
40.04
41.64
42.98
2.95
3.59
4.94
6.97
8.89
9.25
9.27
20.83
19.47
18.55
17.56
15.58
14.89
14.7
14.35
13.39
15.41
17.42
19.56
19.17
18.9
14.12
13.29
10.62
8.75
7.08
6.45
11.98
13.46
12.39
10.39
8.85
8.6
7.92
24.57
25.33
26.27
27.81
30.45
33.43
38.34
0.21
0.20
0.25
0.32
0.40
0.40
0.39
GHSV=3600 h−1
)
6.23
Fig. 6 SEM images of 45%Fe–45%Co–10%Ce/60%Zeolite A-3 catalyst at different temperatures a 340 °C, b 350 °C, and c 360 °C
Table 3 Effect of different
GHSV on the 45%Fe–45%Co–
10%Ce/60%Zeolite A-3
GHSV (h−1
)
Selectivity (%)
(%) CO con- Olefin/paraffin
version
C+
CH4
C2H4
C2H6
C3H6
C3H8
4
catalytic performance
1800
2400
3000
3600
4200
39.19
39.49
40.23
42.45
54.69
7.52
9.6
10.72
11.92
13.05
15.07
13.41
12.56
11.77
9.59
19.04
20.01
20.21
20.25
19.86
6.84
5.52
4.77
4.28
2.51
12.34
11.97
11.51
9.33
32.3
0.36
0.42
0.45
0.47
0.49
(T=340 °C, P=2 bar, H2/
CO=2/1 and GHSV=1800–
21.58
16.64
15.53
13.64
4200 h−1
)
0.3
Table 4 H2/CO molar ratio
effect of the 45%Fe–45%Co–
10%Ce/60%Zeolite A-3
H2/CO ratio
Selectivity (%)
(%) CO
Olefin/paraffin
conversion
C+
CH4
C2H4
C2H6
C3H6
C3H8
4
catalytic performance on
1
2
3
4
39.58
41.47
44.54
48.12
10.83
7.64
7.05
6.51
12.23
15.35
15.69
16.57
20.45
17.93
14.63
13.02
5.24
7.61
7.81
8.06
11.67
10
10.28
7.72
21.08
26.81
29.29
30.02
0.46
0.34
0.28
0.24
distribution of products of FT
reaction (T=340 °C, P=2 bar
and GHSV=3600 h−1
)
the effect of total reaction pressure on the catalytic activity
and selectivity, FT reaction was carried out under similar
conditions (molar ratio of H2/CO=1, GHSV=3600 h−1 and
T=340 °C) and variable total pressure of reactants from 2
to 10 bar. The results are presented in Table 5. According
to the results of Table 5, it is clear that with increase in the
total pressure from 2 to 10 bar, significant increase in CO
conversion (from 17.14 to 39.63) was observed and selectiv-
ity towards C+ increased. Abbaslou et al. [54] explained that
by increasing4the pressure, the supercritical media indicates
a liquid-like density that can increase the extraction from
the catalyst pores. This phenomenon aids adsorption of car-
bon monoxide and hydrogen onto active sites, leading to
increase in conversion. It is also evident that with increase
in the total pressure in the range of 2–10 bar, the selectivity
to light olefins decreased from 31.41 to 12.04%. P=2 was
selected as the optimum pressure because of highest selec-
tivity to light olefins and olefin to paraffin ratio. XRD pat-
terns of the catalyst samples (precursor, fresh calcined and
used calcined) are displayed in Fig. 7. The actual phases that
1 3

R. Nohtani et al.
Table 5 Effect of pressure
on the 45%Fe–45%Co–
10%Ce/60%Zeolite A-3
catalytic performance
P (bar)
Selectivity (%)
(%) CO
Olefin/paraffin
conversion
C+
CH4
C2H4
C2H6
C3H6
C3H8
4
2
3
4
5
6
7
8
9
38.92
39
38.83
38.6
39.18
39.09
38.76
38.62
39.06
11
12.78
14.98
15.95
16.73
17.79
17.97
18.09
18.41
18.85
20.41
19.52
17.98
16.43
14.71
13.6
11.91
11.14
10.3
5.42
7.21
8.75
10.23
11.65
12.29
13.51
13.78
14.74
11.47
11.14
12.21
13.22
12.96
13.99
15.42
16.01
15.31
17.14
18.78
19.92
23.11
25.72
28.3
31.99
34.91
39.63
0.46
0.38
0.32
0.27
0.23
0.20
0.17
0.15
0.14
(T=340 °C, H2/CO=1/1 and
GHSV=3600 h−1
)
8.15
6.28
4.79
3.71
3.06
2.31
2.04
1.74
10
Fig. 7 XRD patterns of optimal
catalyst (45%Fe–45%Co–
10%Ce/60% Zeolite A-3) sam-
ples (precursor, fresh calcined
and used calcined catalyst)
1 3

Synthesis of Fe–Co–Ce/Zeolite A-3 Catalysts and their Selectivity to Light Olefins for Fischer–…
Table 6 BET measurements of optimum catalyst (45%Fe–45%Co–
TGA and DSC curves are presented in Fig. 8. As shown
in TGA curve, a weight loss of 8 percentage was observed
in the temperature from 80 until 360 °C. The initial weight
loss happened at a temperature range of 80–180 °C, which
can be attributed to the loss of physisorbed water and sec-
ond mass loss in the temperatures 180–360 °C can be due
to decomposition of the nitrate phase. The DSC spectrum
of this sample also confirms the TGA result. As shown in
Fig. 8, DSC spectra have two peaks. An endothermic peak
appears at low temperatures, and is referred to as the loss
of physisorbed water. Another endothermic peak appears at
higher temperatures, which is related to decomposition of
the nitrate phase.
10%Ce/60%wt) samples (precursor, calcined and used calcined)
Precursor
Surface areas (m2/g)
63.45
Fresh calcined
68.15
Used calcined
34.06
were detected in the precursor were FeOH (orthorhombic),
Ce(OH)3 (hexagonal), Fe(OH)3 (cubic) and Co(OH)2 (hex-
agonal). The phases that were detected in the calcined cata-
lyst before the reaction were Fe3O4 (cubic-orthorhombic),
Ce2O3 (hexagonal), Fe2O3 (rhombohedral), Co3O4 (cubic)
and Ce3O4 (hexagonal) which were different from the pre-
cursor. These results show that different phases detected in
the precursor were transferred to oxidic phases. The calcined
catalyst after the test had different phases [Co3C (orthorhom-
bic), Fe3C (orthorhombic), CeC2 (tetragonal), FeO (cubic),
CeCo2 (cubic), Ce2O2C2 (hexagonal) and Fe3O4 (cubic)]
compared to the calcined catalyst before the reaction. The
calcined catalyst after the reaction had oxidic and carbidic
phases, these are active phases for the FT reaction [55, 56].
BET measurements for both calcined catalyst (before and
after reaction) and precursor are presented in Table 6. The
data showed that the calcined catalyst before the start of
the reaction had the greatest surface area. Used catalyst had
lower surface area than calcined catalyst before the reac-
tion, which might be due to sintering and coke formation.
It is worth mentioning that one of the application of ZSM5
is in the Mobil/Badger process since 1980 [57]. One of the
faced problems in this process is coking, and the different
conditions required for regeneration. However, there are
numerous studies that have illustrated a high resistance to
coking of ZSM-5 [58–60]. Also, the results in this research
demonstrated that under Fischer–Tropsch operating condi-
tions (P=2 bar, T=320 °C, molar ratio of H2/CO=2/1 and
GHSV = 3600 h−1), the catalyst has a good stability. It is
worth mentioning that there is a probability for coking at
high temperatures (T=360 °C, Sect. 3.2.1). Similar result
was obtained for Co–Ce/zeolite in our previous study [61].
An optimal catalyst precursor was investigated by evalu-
ating thermal behavior using TGA and DSC technique.
4 Conclusion
The effect of operational conditions and catalyst com-
position on the selectivity over Fe–Co–Ce/Zeolite A-3
supported catalysts was studied. BET surface areas of
catalysts showed that different amounts of cerium causes
changes in surface areas of catalysts and catalyst with
higher surface area has maximum olefin to paraffin ratio.
Amounts of metals and support were variable finally,
45%Fe–45%Co–10Ce%/60%Zeolite A-3 catalyst was found
to be the optimum catalyst because of its high selectivity
towards light olefin and methane suppressed. Catalyst per-
formance was also impacted by reaction condition such as
pressure, temperature, GHSV and feed molar ratio. The SEM
images showed that the catalysts were deactivated at a tem-
perature higher than 340 °C by formation of coke and sin-
tering, as well. The results of XRD showed that calcination
caused the transformation of different phases of the hydrox-
ide in the optimum precursor catalyst to oxidic phases. The
TGA curve of optimum precursor catalyst showed two steps
of mass loss at temperatures of 80–180 °C and 180–360 °C,
corresponding to evaporation of physisorbed water, decom-
position of the nitrate phase and stable oxide production,
respectively. Optimal reaction conditions were found to be
T = 340 °C, P = 2 bar, GHSV = 3600 h−1 and feed molar
ratio=1/1.
1 3

R. Nohtani et al.
Fig. 8 TGA and DSC curves for
(45%Fe–45%Co–10%Ce/60%
Zeolite A-3) precursor
11. Van Der Laan GP, Beenackers A (1999) Catal Rev 41:255
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