Fixed Co2 System Diagram

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Studies in Surface Science and Catalysis 153 S.-E. Park, J.-S. Chang and K.-W. Lee (Editors)

© 2004 Elsevier B.V. All rights reserved. 67

CAMERE Process for methanol synthesis from C02 hydrogénation

Oh-Shim Joo*, Kwang-Deog Jung and Yonsoo, Jung

Eco-Nano Research Center, Korea Institute of Science and Technology (KIST), P.O. Box 131, Cheongryang, Seoul 130-650, Korea.

Tel: +822-958-5215, FAX: +822-958-5219, E-mail: [email protected]

A pilot plant to produce methanol by CO2 hydrogénation has been constructed with the government research fund in participating with POSCO and Korea Electric Power Research Institute (KEPRI). The pilot plant is consisted of a reverse water gas shift reactor and a methanol reactor. Two reactors are serially connected to remove water in the first reactor and then, synthesize methanol in the second reactor. The production capacity of the plant is 100 kg methanol per day. The methanol production yield in CAMERE Process is higher than twice as compared with the yield in the direct hydrogénation of CO2 into methanol without reverse water gas shift reaction. The pilot plant for methanol synthesis from CO2 was combined with the pilot plant for separation of CO2 discharged from a power plant.

1. introduction

The CAMERE process (Carbon dioxide hydrogénation to form methanol via a reverse-water gas shift reaction) was developed to fix CO2 into methanol. The reverse water gas shift reactor and the methanol synthesis reactor are serially aligned to form methanol by CO2 hydrogénation. Carbon dioxide is firstly converted to CO and H20 via the reverse water gas shift reaction (RWGSR) and then, the water is removed from the reactant gas before injection into the methanol reactors. The higher the conversion of CO2 to CO is in the RWGSR, the higher the methanol productivity can be increased, because methanol yield is dependant on the CO concentration in the CO2/CO/H2 mixture gas. Therefore, the volume of the recycle gas in the methanol synthesis reactors can be minimized by increasing the conversion of CO2 to CO in RWGSR as compared with the direct CO2 hydrogénation into methanol.1

The RWGSR should be carried out at higher temperature than 600 °C to obtain CO2 conversion over 60 % in the thermodynamic point of view.2 Therefore, development of an active and stable catalyst for the RWGSR at higher temperature than 600 °C was a critical requirement for the CAMERE process.

The water-gas-shift reaction has been studied intensively for the last several decades in order to adjust H2/CO ratio in the synthesis gas. On the contrary, a reverse-water-gas-shift reaction has attracted little attention because of little demand. The Fe203/Cr203 catalyst is a well-known commercial catalyst for the water-gas-shift reaction. The commercial catalyst, Fe203/Cr203, was not a good candidate for the RWGSR of the CAMERE process because of the severe deactivation.3 Deactivation of the catalyst was attributed to the reduction of Fe203 to the Fe metal. On the other hand, a new type of catalyst, ZnAl204, for the RWGSR was developed, which showed good activity and stability without coke formation.4

In this paper, ZnAl204 catalyst was optimized for RWGSR and the pilot plant was operated using the ZnAl204 and Cu/Zn0/Al203 catalysts.

2. experimental

The ZnAl204 catalyst (Zn:Al=l:2 in molar ratio) was prepared by a co-precipitation of the corresponding metal nitrates.5 The catalyst was calcined at 850 °C before the RWGSR. The activity of the ZnAl204 catalyst was studied in a test reactor of 3/8 "o.d. for the RWGSR at the temperature range of 400-750 °C with different GHSV (ml/gca,.h). The RWGSR was carried out at atmospheric pressure in the mixed gas with H2/CO2 ratio of 3 in the reactant gas. The concentration of the exit gas was measured with a gas chromatograph (Shimadzu 8A instrument with a TCD, Porapak Q column l/8"o.d., 3m long in program temperature mode for analysis of CO, C02, H2, CH4 and H20.

In the pilot plant, the RWGSR is operated on Z11AI2O4 in the temperature range of 600-700°C and at atmospheric pressure in the pilot plant and methanol is synthesized on Cu0/Zn0/Al203 under the reaction conditions of 250-300°C and 50-70atm. The feed gas of CO2/H2 (1/3) mixture gas is preheated before RWGSR. The RWGSR is endothermic reaction and three electric heaters are used for the reaction temperature control. The reactor size is 2"(I.D.) x 120cm (length). The methanol synthesis from C02/C0/H2 mixture gas is exothermic reaction and the reaction temperature is controlled to minimize the hot spot by the exothermic reaction by steam. Four fixed bed reactors for multiple purposes are used for methanol production. A reactor size is 2"(I.D.) x 100cm (length). A two-stage compressor (Diaphragm compressor, Sera MV 4714-IIK) is used for compressing the C02/C0/H2 mixture gas before injection into the methanol synthesis reactor. The concentrations of gas stream at four points of the pilot plant were monitored with two gas chromatographs and water content in final products was analyzed using karl-fisher (658KF).

3. results and discussion

Z11AI2O4 catalysts were prepared by a coprecipitation and the activity of the prepared catalysts for RWGSR was dependent on the pH of the solution (Figure 1). The Z11AI2O4 catalyst prepared at pH=7 shows the highest activity as compared with catalysts prepared at acidic condition. Especially, Z11AI2O4 catalyst prepared at pH =5.4 shows lower activity and broaden diffraction pattern. We obtained the ZÎ1AI2O4 catalyst with the highest activity at pH=7.

Figure 1. CO2 conversion and the X-ray diffractogram of ZnAl204 catalyst depending on the preparation pH. (a) pH=5; (b) pH=6.0; (c) pH=6.4; (d) pH=7.0

Figure 2 shows CO2 conversion with respect to GHSV over ZnAl204 catalyst prepared at pH=7. The dashed line is the equilibrium conversion for RWGSR. When the reaction temperature is increased, CO2 conversion over ZnAl2C>4 catalyst approaches to the equilibrium conversion. Most oxide catalysts show high catalytic activity for RWGSR at atmosphere pressure, but are rapidly deactivated because the RWGSR condition is very reductive above 400°C and the reactant ratio of H2/C02=3/l.2'3 Moreover, it should be operated above 600°C to obtain higher C02 conversion than 60%.

Co2 Conversion

Figure 1. CO2 conversion and the X-ray diffractogram of ZnAl204 catalyst depending on the preparation pH. (a) pH=5; (b) pH=6.0; (c) pH=6.4; (d) pH=7.0

Methanol Yield Co2 Conversion

400 500 600 700

Temperature (°C) Figure 2. C02 conversion over ZnAl204 with GHSV

400 500 600 700

Temperature (°C) Figure 2. C02 conversion over ZnAl204 with GHSV

Co2 Conversion

And the mixed gas from the RWGSR with part of recycled gas is fed into methanol reactor of R-201. Part of recycled gas is fired in the FN-101 to regulate the reaction conditions. The temperature of the four fixed bed reactors was well controlled by steam in the temperature range of 250-300°C for methanol production.

The pilot plant has been operated to obtain the optimum reaction conditions and the data for the evaluation of the methanol production cost. To evaluate the effect of the RWGSR on the methanol production yield, the pilot plant was operated in the RWGSR (On) or RWGSR(Off). RWGSR (On) means that the RWGSR was operated at the temperature range of 600-700°C. On the contrary, RWGSR (Off) means that the RWGSR was not operated during methanol production so H3/C03=3/l was just injected into the methanol reactor. Table 1 distinctively shows the effect of the reverse water-gas shift reaction on the methanol production yield. The methanol yield in the RWGSR (On) becomes more than twice in comparison with the yield in the RWGSR (Off) at the same reaction conditions. Moreover, the CO2 conversion over the Z11AI2O4 catalyst was about 35% in the RWGSR (On). It is worth noting that the reverse water-gas shift reaction shows a significant effect on the increasing of the methanol yield.

RWGSR(On)

RWGSR(Off)

C02(kgmol/day)

3.34

3.35

3.35

3.34

3.35

3.35

H2(kgmol/day)

10.04

10.09

9.94

10.18

10.15

10.15

CH3OH(kg/day) /100% Yield

107.1

107.3

107.3

107.1

107.3

107.3

CH3OH(kg/day)/ Pilot plant

71.67

72.58

75.46

35.95

34.87

39.82

CH3OH Yield

66.9

67.63

70.46

33.57

32.49

37.11

Pressure(atm)

51.0

51.0

60.7

61.1

61.3

70.9

Table 1. CH3OH Yield for CAMERE process based on RWGSR(On) or (Off)

Table 1. CH3OH Yield for CAMERE process based on RWGSR(On) or (Off)

CAMERE Process was simulated using the simulation program of Aspen Plus and the state equation of UNIFAC based on the 2000 tons methanol production in a year. The construction cost for the plant was evaluated based on the Guthrie's Modular Method. The methanol production cost was calculated depending on the hydrogen cost and methanol production capacity as shown in the figure 5. The methanol production cost proportionally increases with hydrogen cost in market and dramatically decreases up to 50,000 tons of methanol production capacity and then, becomes stable with the production capacity. It indicates that methanol can be produced with 300US$/ton from CO2 hydrogénation through CAMERE process if a commercial plant of 100,000ton/year is constructed. It also means we cannot economically produce methanol from CO2 hydrogénation because the methanol is sold at 100-150 US$/ton in the recent market. To become an economical process, the target material of the CAMERE process should be changed into another one having a value added, for example, DME (Dimethyl ether).

Figure 5. Methanol production cost via CAMERE Proccss depending on hydrogen cost and methanol production capacity,

4. CONCLUSIONS

Methanol yield of 70 % was obtained from the pilot plant for CAMERE process. Methanol of 75kg was produced in a day from the pilot plant for which about the 100kg of CO2 was consumed. Based on the results, we estimated the methanol production cost depending on the hydrogen cost and methanol production capacity. Operating cost of about 300USS was requested for Iton methanol production through CAMERE process.

REFERENCES

1. Joo, O.S, et al, Ind. Eng. Chem. Res., 38(5), 1808,1999.

2. Park, S.W., Joo, O.S., Jung, K.D., Han, S.H, Appl. Catal. A:general, 211, 81,2003.

3. Park, S.W., Joo, O.S., Jung, K.D., Kim, H., and Han, S.H, ICor. J. of Chem. Eng, 17(6), 719,2000.

4. Joo, O.S., and Jung, K.D., Bull. Korean Chem. Soc., 24(1), 86, 2003.

5. Joo, O. S., Jung, K. D., Han, S. H„ Uhm, S. J., Lee, D. K. and Ihm, S. K., Appl. Catal, A: General 135,273, 1996.

Methanol Production

Figure 5. Methanol production cost via CAMERE Proccss depending on hydrogen cost and methanol production capacity,

600 000 1000 1200 1400 1600 1800 2000 Hydrogen Cost (USS/tun)

600 000 1000 1200 1400 1600 1800 2000 Hydrogen Cost (USS/tun)

Methanol Production capacity —2,000 ton/yr

On the other hand, the methanol production cost of 300US$/ton calculated here would be a standard value for carbon dioxide sequestration process. Jn addition, whenever the carbon tax starts to work, CAMERE process to sequestrate carbon dioxide should be evaluated based on the real situation.

Studies in Surface Science and Catalysis 153 S.-E. Park, J.-S. Chang and K.-W. Lee (Editors)

© 2004 Elsevier B.V. All rights reserved. 73

Product Distribution Analysis for Catalytic Reduction of C02 in a Bench Scale Fixed Bed Reactor

Sang-Bong Lee,a Jun-Sik Kim,a Won-Young Lee,a Kyu-Wan Lee,b and Myoung-Jae Choia'*

a Advanced Chemical Technology Division, KRICT, PO Box 107, Taejon 305-600, Korea b Yanbian University of Science & Technology (YUST), Yanji, Jilin, China *E-mail: [email protected]

Hydrogenation of C02 was carried out over Fe-K/Al203 catalyst in a bench scale fixed bed recycle reactor with an aim to get higher molecular weight hydrocarbons including mainly olefinic fractions. Conversion of C02 reached 88.2% at the recycle ratio of 5 where reaction temperature, space velocity, and H2/C02 ratio were 300°C, 2,000ml/gcat.hr and 3, respectively. On increasing recycle ratio, selectivity to C2-C4 olefins tends to decrease and selectivity to higher branched hydrocarbons was increased via C2-C4 olefins oligomerization. Overall hydrogenation mechanism of C02 was proposed on the basis of the product distribution of oily and aqueous phases.

1. intoduction

Direct hydrogenation of C02 which is called here as modified Fisher-Tropsch (MFT) reaction over Fe based promoted catalysts have been highlighted because of no sulfur and nitrogen containing liquid fuels and chemicals. Similar product distribution of Fisher-Tropsch reaction could be obtained via one-step reaction [1-5]. In our previous work, screening of catalyst, parameter optimization for increasing C3+ alpha olefins, and de/reactivation of catalysts have been investigated from a process development point of view. Meanwhile, it has been a big problem to separate H2, C02, CO, and light hydrocarbons, especially CH4, due to its difficulty and high cost. In our process development study for technology of oil production for liquid fuel or chemicals, more simple and economically feasible process is needed. Our suggestion to solve this problem is to skip separation of H2, C02, CO, and light hydrocarbons but to recycle light gaseous fractions lighter than C5 fraction aiming at liquid fuels or chemicals without sulfur content.

In this paper, C02 hydrogenation in a bench scale fixed bed reactor with recycling of gaseous product and product oriented overall mechanism of MFT reaction are described.

2. experimental

2.1. Catalyst preparation

The catalyst Fe-K/Al203 was prepared in a large scale by previously reported method [4] except palletizing. Regular 2x5mm size of the catalyst was prepared by extrusion using 1 wt% of PVA (poly vinyl alcohol) as a binder.

2.2. Hydrogenation of C02 in a bench scale fixed bed system (Fig 1)

252 g of catalyst was loaded in a fixed bed reactor and was reduced with sufficient H2 for

1. Muss flow controller

2. Fixed bed P' reactor

3. Fixed bed 2"d reactor

4. Electric heater

5. Gas-liquid separator

6. Condenser

7. Heat exchanger

8. Back pressure regulator

9. Gas compressor

10. Buffer tank

11. Pressure regulator

12. Wet gas meter

Fig. 1. Flow Diagram of bench scale fixed bed system for CO2 Hydrogenation.

24 hrs at 450°C. CO2 and H2 were passed through the reactor using MFC at elevated temperatures. Pressure was controlled with the aid of BPR. Liquid and oil were trapped in a 3L tank at 5°C and uncondensed gases and hydrocarbons were compressed and kept in a 20L tank at 20 atm and recycled to the reactor according to the flow rate proposed using MFC.

Conversion of CO2 (Xco2) and selectivity to CO and CH4 were determined by a gas chromatography on Carbosphere column and TCD by analyzing vent gas. N2 was used as internal standard gas. Light hydrocarbons were analyzed on Poraplot-Q column and FID using the analyzed CH4 as an internal standard gas. Liquid oils and water-soluble organics were analyzed on DB-1 and HP-5 column, respectively.

2.3. Separation of oil product on column chromatograph

5 g of oil product was separated in a silica gel column chromatograph 5 cm-ID x 60 cm-H according to a typical method using hexane and ethylacetate as effluent. Compositions of the separated 4 kinds of fractions with different polarity are analyzed by GC/MS and IR. Carboxylic acids were separated as Na salts followed by acidification with HC1.

3. results and discussion

3.1. Preparation of catalyst and activity test

As shown in Table 1, the composition of PVA-bound catalyst was in the proportion of Fe:K:Al203=20:7:100 by weight. The BET surface area was measured to be 92 m2/g. The C02 and H2 uptake were found to be 514 and 20|imol/g of catalyst, respectively. These results were similar to that of reported conventional pressurized catalyst and were expected to show similar activity. 1 wt% of organic binder (PVA) was enough to make uniform extrudate and to give proper strength.

Activity test of the new PVA-bound catalyst was conducted under the same conditions as reported elsewhere [6]. Almost the same conversion of CO2 (38%) was obtained, and therefore the catalyst prepared in large scale using PVA binder was confirmed to be usable as like pressure-formed catalyst.

Fixed Co2 System Diagram

Fig. 1. Flow Diagram of bench scale fixed bed system for CO2 Hydrogenation.

Spectra Cr203

Fig. 2. GC of each fraction separated from Fig. 3. FT-JR spectrum of each fraction oily and aqueous phase products, from oiiy and aqueous phase products.

Fig. 2. GC of each fraction separated from Fig. 3. FT-JR spectrum of each fraction oily and aqueous phase products, from oiiy and aqueous phase products.

■ Fe CH; f -Fe —C—C-0 •—» » » Fe+3-CH2-CH2

Fe4"2 + ch2=ch2

Cat.

Fig. 4. Plausible overall reaction mechanism of CO2 hydrogénation.

Cat.

Major route Medium route Minor route

Fig. 4. Plausible overall reaction mechanism of CO2 hydrogénation.

From these results overall plausible mechanism of CO2 hydrogenation is summarized in Figure 4. In the beginning C02 is reduced by iron (II) followed by H radical abstraction on the catalyst surface. When the residual H radical attacked carbonyl C and OH, formic acid and CO are formed, respectively. Aldehydes and alcohols can be produced in the next stage. By the same manner Fe-CH2 radical is formed and is regarded as carbon-carbon propagation species. In our reaction system chain propagation is considered as a major route because higher hydrocarbon is major product. Higher a-olefin selectivity to paraffin is attributable to less H2 uptake and no excess H2 in this reaction system.

3.3. CO2 hydrogenation in the fixed bed recycle reactor.

As mentioned in the introduction, recycle of gaseous fraction lower than C\ was investigated in order to avoid difficult separation process. As shown in Fig. 5, conversion of CO2 was tended to increase on increasing recycle ratio. Selectivity to CO was tended to decrease probably due to rather high concentration and reactivity of CO than C02 in the recycling gas as compared to fresh gas. Conversion might be reached above 90% if recycle can be increased above 6, however, continuous feeding was difficult due to the lack of the recycle gases above recycle ratio of 3 where conversion was higher than 70%. Life-time test of the catalyst was conducted at recycle ratio of 3 and SV of 4,000 due to the reason described above. The catalyst activity maintained for 1,000 hrs (Fig. 6).

In order to predict the effect of recycle ratio more precisely, product distribution comparison was summarized in Table 2. Conversion of C02 in a single reactor was slightly increased to 40.8% at reduced SV 1,000, i.e. low SV is one of a tool to increase conversion. When a series reactor was used having the similar concept of recycle reactor, the CO2 conversion drastically increased to 68.5%. In this case, however, pressure drop was to big to continue the reaction at constant pressure probably due to the additional catalytic layer and

Ozone Depletion Layer DiagramScientific Liquid Conversion Chart

Studies in Surface Science and Catalysis 153 S.-E. Park, J.-S. Chang and K.-W. Lee (Editors) © 2004 Elsevier B.V. All rights reserved.

Process Evaluation of Biomass to Liquid Fuel Production System with Gasification and Liquid Fuel Synthesis

Tomoaki Minowa*, Toshiaki Hanaoka, and Shin-ya Yokoyama

Biomass Technology Research Lab, National Institute of Advanced Industrial Science and Technology (AIST),

AIST Chugoku, Suehiro, Hiro, Kure, Hiroshima 737-0197 Japan

We have proposed biomass to liquid fuel production process with gasification and liquid fuel synthesis (BTL process). The process was experimentally studied using rice straw as a biomass, steam gasification, and FT and Oxo synthesis, supported by New Energy and Industrial Technology Development Organization (NEDO), Japan. Overall mass and energy balance for the process was calculated. Liquid products of 8 wt% on the dry biomass basis was estimated, and its energy yield was around 25% on HHV basis. The process could be operated without energy from outside, when obtained off gas was used for the process.

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