Design of a plant producing 500,000 tones/annum synthetic oil products from natural gas, using Fischer-Tropsch technology

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CPD NR

3296

Conceptual Process Design

Basic of Design

Process Systems Engineering DelftChemTech - Faculty of Applied Sciences

Delft University of Technology

Subject

Final Report:

Design of a plant producing 500,000 tones/annum

synthetic oil products from natural gas, using

Fischer-Tropsch technology

Authors

(Study nr.)

Telephone

Binbin Bai 1134345 0641763172

Junying Hu 1160842 0641763830

Nan Liu 1132016 0641439516

Yan Jiao 1160915 0624560836

Zhiyong Wang 1129767 0618241976

Keywords

Fischer-Tropsch synthesis, Hydrocracking, Syngas

production, Combined autothermal reforming, Natural

gas.

Assignment issued

:

Sep. 22, 2003

Report issued

:

Dec. 15, 2003

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Group Conceptual Process Design Project CPD_3296

Summary

-ii-Summary

The conceptual process design is an important course for chemical engineering student. According to the project, the mythology of design technology is used and increases the creative and economic thinking.

In this project, natural gas is used as feedstock to produce 500,000 ton/year syngas through Fischer-Tropsch synthesis process technology, which is going to be converted into synthetic oil products. Among them, the target products are diesel (C15-C20), Kerosene (C10-C14), Naphtha (C5-C9) and LPG (C2-C4) is accepted as by-products.

According to the requirements, the chosen process consists of four operation units that are syngas production unit, Fischer-Tropsch synthesis process,

hydrocracking unit, and Separation unit. Combined autothermal reforming (CAR) reactor is applied to convert natural gas into syngas, which is the feedstock of Fischer-Tropsch synthesis process. In order to improve product quality and quantity, hydrocracking is placed after Fischer-Tropsch synthesis. Finally, diesel, kerosene, LPG and Naphtha will be separated respectively by distillation column. The reactor selection and design is based on literature and also include our creative design. Each unit has several options and the total process has alternative too.

The process is simulated in ASPEN and all the product specifications satisfy the requirement of the client. The annual production is 144452.4tone naphtha, 187726.8tone kerosene and 207612tone diesel. The process yield of each product is 27%, 35% and 38% (defined as t/t products).

The total investment cost is 34.68 [million $/year]. The income is 63 [million $/year]. The production cost is 133.2 [million $/year]. Then the net cash flow is -70.198 [million $/year], which means our margin is negative.

There are some wastes generated in our plant and we only consider the treatment of the indirect wastes that are CO2, coke, oxygenates, wax and

nitrogen oxide. The emission of the waste satisfies the emission standard of the Europe Commission (EC).

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-iii-Table of Content

Summary ……… II

1. Introduction ……… 1

1.1 Conceptual process design ……… 1

1.2 Project CPD_3296 ……… 1

1.3 Fischer-Tropsch synthesis ……… 1

1.4 Brief process description ……… 2

1.5 Environment ……… 2

2. Process Options and Selection ……… 3

2.1 Syngas unit ……… 3

2.1.1 Oxygen supply ……… 3

2.1.2 Energy recovery method ……… 3

2.1.3 Carbon dioxide recycle ……… 3

2.1.4 Raw syngas purification operation sequence ……… 4

2.1.5 Pure hydrogen separation route ……… 4

2.2 Fischer-Tropsch synthesis unit ……… 5

2.2.1 The conversion in Fischer-Tropsch synthesis ……… 5

2.2.2 Catalyst and wax separation of FT synthesis ……… 5

2.2.3 Basic block scheme of FTS process ……… 5

2.3 Process options of Hydrocracking unit ……… 6

3. Basis of Design ……… 10

3.1 Description of the Design ……… 11

3.2 Process Definition ……… 11

3.2.1 Process concepts chosen ……… 11

3.2.2 Block schemes ……… 15

3.2.3 Thermodynamic properties ……… 17

3.2.4 Pure component properties ……… 18

3.3 Basic Assumptions ……… 19

3.3.1 Plant capacity ……… 19

3.3.2 Plant location ……… 19

3.3.3 Battery limit ……… 20

3.3.4 Definition In- and Outgoing streams ……… 20

3.4 Economic Margin ……… 22

3.4.1 Calculation of economic margin ……… 22

3.4.2 Calculation of maximum allowable investment ……… 22

4. Thermodynamic Properties and Reaction Kinetics ……… 23

4.1 Operating windows ……… 23

4.1.1 Syngas production unit ……… 23

4.1.2 Fischer-Tropsch unit ……… 26

4.1.3 Hydrocracking operation unit ……… 27

4.1.4 Brief summary of operating windows ……… 29

4.2 Heat data ……… 29

4.3 Models for vapor/liquid equilibrium ……… 30

4.4 Reaction kinetics ……… 30

5. Process Structure and Description ……… 31

5.1 Criteria and Selections ……… 31

5.1.1 Syngas production unit ……… 31

5.1.2 FT synthesis unit ……… 33

5.1.3 Hydrocracking unit ……… 36

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Table of Content

-iv-5.2 Process Flow Scheme (PFS) ……… 38

5.3 Process Stream Summary ……… 39

5.4 Utilities ……… 40

5.4.1 Utility introduction ……… 40

5.4.2 Pinch and heat exchanger network ……… 40

5.5 Process yields ……… 41

6. Process Control ……… 43

6.1 Syngas production unit (U100) ……… 43

6.2 Fischer-Tropsch synthesis unit (U200) ……… 44

6.3 Hydrocracking unit (U300) ……… 46

6.4 Separation unit (U400) ……… 47

7. Mass and Heat Balances ……… 48

8. Process and Equipment Design ……… 49

8.1 Integration by process simulation ……… 49

8.2 Equipment selection and design ……… 49

8.2.1 Syngas reactor design ……… 49

8.2.2 Reactor design of Fischer-Tropsch synthesis ……… 50

8.2.3 Hydrocracking design ……… ……… 51

8.2.4 Separation unit design ……… 52

8.2.5 Shell and tube exchanger design ……… 54

8.2.6 Single flash column design ……… 56

8.2.7 Pump and compressor design ……… 57

8.2.8 H2 embrane separation (S101) ……… 57

8.2.9 CO2 removal separator (S102) ……… 58

8.3 Equipment data sheets ……… 58

9. Waste ……… 59

9.1 Introduction ……… 59

9.2 Waste treatment ……… 59

9.3 Emission limit values ……… 60

9.3.1 Air emission limit value ……… 61

9.3.2 Water emission limit value ……… 61

10. Process Safety ……… 62

10.1 Hazard and operability studies (HAZOP) ……… 62

10.1.1 Introduction of HAZOP ……… 62

10.1.2 HAZOP Analysis ……… ……… 63

10.2 Dow Fire and Explosion Index (F&EI) method ……… 63

10.3 Conclusion ……… 65

11. Economy ……… 66

11.1 Investment ……… 66

11.2 Cash flow ……… 67

11.3 Economic evaluation of the project ……… 69

11.4 Cost review ……… 70

11.5 Sensitivities ……… 70

11.6 Negative cash flows ……… 70

12. Process Safety ……… 72

12.1 Group relation diagram ……… 72

12.2 Group creativity evaluation ……… 72

12.3 Creativity implication in CPD ……… 74

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-v-13. Conclusions and Recommendations ……… 77

13.1 Conclusions ……… 77

13.2 Recommendation ……… 78

List of abbreviation ……… 78

List of symbols ……… 80

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Table of Content

-vi-Table of Content for Appendix

Appendix 1. ……… A1 1.1 Overall process scheme ……… A1 Appendix 2. ……… A2

2.1 Oxygen supply evaluation.……… A2

2.2 Conversion route ……… A3

2.3 Catalyst and wax separation in FTS process ……… A5

Appendix 3. ……… A8

3.1 Feedstock specifications .……… A8 3.2 Product specifications ..……… A9 3.3 List of prices for feedstock and product . ……… A10 3.4 Physical Properties of Pure components…. ……… A10

Appendix 4. ……… A17 4.1 The process for choosing a property method .……… A17 4.2 Models for vapor/liquid equilibrium ……….……… A19 4.3 Thermodynamic properties in Aspen ..……… A22 Appendix 5. ……… A27

5.1 Syngas production unit . ……… A27

5.2 Syngas ratio adjustment . ……… A32

5.3 CO2 removal technology .. ……… A38

5.4 Fischer-Tropsch synthesis design ….……… A42 5.5 Process flow scheme (PFS) …..….……… A61 5.6 Process Stream Summary ………..……… A62 5.7 Available utility conditions and costs ..……… A80 5.8Pinch Technology.. ……….……… A81 5.9 Utility summary……….……….. A88

Appendix 6. ……… A89

6 Process control………..……… A89

Appendix 7. ……… A90 7Heat and Mass Balance………..……… A90 Appendix 8. ……….……… A94

8.1 Oxygen supply evaluation …….……… A94 8.2 Description of ASPEN simulation…….……… A96 8.3 Process simulation scheme in ASPEN ….………..……… A98 8.4 The kinetics of combined autothermal reforming (CAR) ……… A100 8.5 Fisher-Tropsch reactor design………..……… A105 8.6 Hydrocraking catalyst and kinetics and hydrocracker design….………… A114 8.7 Design procedure of distillation column………..……… A123 8.8 Calculation for hydrocracker sizing………..……… A128 8.9 Equipment summary………..……….……… A131

8.10 Equipment specification………..……….…… A141

Appendix 9. ……… A169

9.1 The HAZOP analysis……… A169 9.2 Dow Fire and Explosion Index analysis….……… A182

Appendix 10. ……… A169

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-1-1. Introduction

1.1 Conceptual process design

The course Conceptual Process Design (CPD, CE3811) is part of the 4th_year’s curriculum for students studying Chemical Process Technology (CPT), Bio Process Technology (BPT) and Master of Science International Programme (MSc) at the DelftChemTech (DCT) Department of the Faculty of Applied Sciences (TNW) at Delft University of Technology. The CPD is coordinated by the section Process Systems Engineering (PSE) from the DelftChemTech Department. With this course, students are expected to produce an innovative, integrated, consistent and sound process design, and course time is limited within 12 weeks.

1.2 Project CPD_3296

The objective of this conceptual process design (CPD_3296), performed as part of course CPD by a group of five people, is to design a plant producing 500,000 tonnes/annum synthetic oil products from natural gas, using Fischer-Tropsch technology. The principal/client is Ir. Pieter Swinkels and Austine Ajah, and Cristhian Almeida Rivera is responsible for creativity and group process coaching. This CPD project (CPD_3296) focuses on the production of diesel and kerosene from natural gas, and LPG and naphtha can be concerned as by-products. For the 500,000 t/a capacity diesel, kerosene and naphtha are acceptable. Regarding the detailed product specification, please see Appendix 1. Moreover, the design project is quite special, and will be used for comparison to an alternative design made in the past. Therefore, price level related to 1999 is used, and literature information from 1998 and before is used.

1.3 Fischer-Tropsch synthesis

Main process involved in the design is the well-known Fischer-Tropsch (FT) synthesis operation. In 1923, Dr.Franz Fischer and Dr.Hans Tropsch developed the so-called Fischer-Tropsch synthesis process at the Kaiser Wilhelm Institute in Mullheim. In the FT process, synthesis gas, a mixture of predominantly CO and H2, obtained from natural gas, is converted to a multicomponent mixture of hydrocarbons. Currently, a promising topic in the energy industry is the conversion of remote natural gas to environmentally clean fuels, specialty chemicals and waxes. Fuels produced with the FT process are of high quality due to a very low aromaticity and absence of sulfur.

At present, there are several plants using this technology all over the world, such as Sasol’s Slurry Phase Distillate Process in South Africa, Shell’s Middle Distillation Synthesis (SMDS) Process in Malaysia. However, now (1998) this technology cannot still compete with the production of middle distillates derived from crude oil. That is because the natural gas price is not cheap enough; therefore, it does not make a big price difference between product and

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Final report

-2-feedstock. Moreover, a high capital investment and operating cost needed, due to the considerable amount of energy consumption. That also agreed with our negative economic margin. However, it can be believed that this promising technology will become economically feasible in the near future.

1.4 Brief Process Description

From natural gas to produce synthetic oil, the main process is quite straightforward; we do not have too many choices on that. A block scheme is shown in Appendix 1. Firstly natural gas is converted into syngas by so-called Combined Autothermal Reforming (CAR), which can be used as Fischer-Tropsch Synthesis (FTS) feedstock. The product of FTS process is quite broad, including unconverted syngas, LPG, Naphtha, Kerosene, Diesel, wax and so on. In order to improve product quality and increase product quantity, hydrocracking unit operation is placed behind FTS process. Finally, product will be fractionated within separation unit operation, in terms of requirements from client. Regarding procedure of fractionation and hydrocracking unit operations, this is the place where it is most likely to make an alternative. That is to say, we may place separation unit operation behind hydrocracking unit directly, and then wax will be fed to hydrocracker. However, the shortcoming of this treatment is leading to products with lower quality, due to lacking of alkalization of olefins, and a bigger volume of distillation column. Regarding the detailed comparison and description, we will come to that later in this report.

1.5 Environment

Although, the plant will be located in Brunei, South-East Asia, European emission rules are used. The main wastes are wastewater and carbon dioxide. They are going to be treated outside the plant, and carbon dioxide could be thought as a by-product, which is sold to food industry as utility. Of course, some solid waste will be produced during the plant operation, such as uncrackable wax, useless catalyst, carbon dioxide, etc. After certain treatment, the solid waste will be transported to landfill or discharged. Therefore, everything complied with European emissions standard.

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-3-2. Process Options & Selection:

2.1 Syngas unit

2.1.1 Oxygen supply

Option 1:Buy pure oxygen from suppliers, the cost should be taken into

account.

Option 2:Oxygen can be separated from air by several technologies (See Appendix 2.1). The cost of building air separation plant should be considered.

On the other hand, Nitrogen can also be sold, if we separate oxygen from air.

Option3: We can use air instead of oxygen in our autothermal reactor, which is

more safe, cheap and easy to control the reactor temperature. But we cannot recycle other gases; otherwise the inert gas such as nitrogen will be accumulated in our system. This will lead to the increasing of the raw material consumption, deactivate the catalyst in the reactor, and consequently affect the economy. Moreover, a reactor with rather big volume has to be applied in this process operation, if air is used as feedstock, instead of oxygen.

Conclusion:

Since we have chosen to use pure oxygen as feedstock, we have two choices. One is to purchase pure oxygen from some producer; the other is to build an oxygen plant. The criteria on whether to build oxygen plant or not, depend on our oxygen supply rate. If it is larger than 20 tons/day, to build oxygen plant is economically feasible (See Appendix 2.1). After rough calculation, our oxygen supply rate is 6.879 kg/s, which means 594 tons/day. Therefore, it’s more economical to build oxygen plant for our process.

The supplier will build an oxygen separation factory near our design project, so the cheaper oxygen than market is supplied and the construction of this factory excluded in our investment.

2.1.2 Energy recovery method

In our design case, we try to use heat-exchanging network to recover the energy, but this can only make up for part of the needed energy input, we still need hot steam and cooling water at the same time. In this plant design, much of hot steam is required. Part of this hot steam is used as reactant for syngas production unit, and the rest is going to be used to heat up the stream. At the same time, much of fuel gas is produced, which is difficult to send out of the factory because the location of factory is in remote area. We are planning to burn the fuel gas in the factory to generate hot steam, which is going to be applied in our process system. And therefore, we need not buy steam.

2.1.3 Carbon dioxide recycle

In order to avoid that carbon dioxide dilutes the reactant concentration of the FTS, it has to be removed from the raw syngas. At the same time, the separated carbon dioxide can be recycled to the CAR, which will promote the carbon dioxide reforming reaction and increase the selectivity of methane converting to carbon monoxide. Consequently, methane can be used more effective. But carbon

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Final report

-4-dioxide reforming will increase the possibility of hot pot formation. All in all, we have the following three choices:

1. No carbon dioxide recycles while the consumption of methane will increase; 2. Carbon dioxide recycles to the primary reforming zone with coke formation; 3. Carbon dioxide recycles to the secondary reforming zone.

From the safety point of view, we mix the recycled carbon dioxide not with natural gas but with the oxygen, and send them to the secondary reforming zone to avoid carbon dioxide reforming occurrence in steam reforming zone, and lower the possibility of the coke formation. Moreover, in the primary zone, three times steam as the stoichiometric amount is sent to the reactor to inhibit coke formation. All in all, we have chosen option 3 in this project design, which is CO2 recycled to the secondary reforming zone.

2.1.4 raw syngas purification operation sequence

To purify the syngas and adjust the hydrogen/carbon monoxide ratio, three separators should be arranged after the CAR, which are partial pure hydrogen separation unit, carbon dioxide removal unit, and water removal unit.

The sequence of the above three unit operations is arranged in the direction of decreasing the operating temperature, namely, first hydrogen separation (750K), water removal unit (200K) and carbon dioxide removal unit (70K). This idea follows the logic thinking of saving energy input. Although there is much impurity in raw syngas, the high selective penetration of hydrogen to the Pd membrane is available. And the purity of separated hydrogen still can be kept more than 99.75%. And water was removed before the carbon dioxide removal, which will avoid that the MEDA mixed with water erodes the pipeline.

2.1.5 pure hydrogen separation route

There are two hydrogen recovery routes, which can be chosen in this project design. And scheme is shown in the below figures.

Option1: Total syngas hydrogen recovery

All the syngas will flow to the hydrogen recovery flash, and the separation will be controlled that only the amount of the pure hydrogen we need is produced.

Option 2: Partial syngas hydrogen recovery

Syngas

Hydrogen

Syngas

Syngas Syngas

Hydrogen

Figure 2.1 Total syngas hydrogen

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-5-Syngas will be split to two branches, one flow is to the hydrogen recovery pipeline, and the other is by-pass. According to the flow ratio control, we can separate almost all of the hydrogen, the residue gas is in mixture with the un-recovery syngas, and the H2/CO ratio of the mixture is 2:1. It is assumed that the two gas streams will mix again in the pipe, so it need not place another mixer afterwards. Another advantage is smaller volume of operating vessel needed in the latter case, since part of flow is by pass.

Conclusion:

According to the requirement of hydrocracking unit, only small amount of pure hydrogen is needed, and in the meantime, the separation efficiency is so high that we need not send all the raw syngas to the separator. In a word, partial syngas recovery system has been chosen in this process.

2.2 Fisher-Tropsch synthesis unit

2.2.1 conversion in Fischer-Tropsch synthesis unit

According to the unsatisfied conversion rate of carbon monoxide in one-stage slurry reactors, where the conversion is just about 80%. And therefore, it is necessary to have a better use of unconverted syngas. In order to improve the reaction conversion more effectively, there are two alternatives to serve this purpose theoretically. One is to apply a syngas recycle; the other is to add more reactors. To select a better route, some comparison was conducted, and process and results are shown in Appendix 2.2.

In overall speaking, applying syngas recycle can save capital investment of reactor, but it is not economically feasible if we take operating cost into account, due to huge energy consuming in distillation separation. Therefore, Option2 to add one more reactor afterwards is taken in our plant design.

2.2.2 Catalyst and wax separation of FT synthesis

In order to achieve the separation of catalyst from wax in FTS process, at present, to our knowledge there are two options available: Gravity

sedimentation and Extraction. By overall consideration and specific

comparison, extraction process is chosen to achieve this separation goal in our case. After separation from wax, the catalyst can be reused in FTS reactor, until the end of catalyst life. Regarding the specific process of consideration and comparison, please see Appendix 2.3.

2.2.3 Basic block scheme of FTS process

Now the basic block scheme for FTS process is already fixed. Altogether there are three slurry reactors involved in this process unit operation. Of course, some auxiliary operations, such single flash distillation column, and catalyst recovery

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Final report

-6-system are needed within this unit. The specific conversion route and block scheme of FTS process is shown the table below.

2.3 Process options of Hydrocracking Unit1

How to arrange separation of the FTS product mixture is the key step of the whole process, which affects not only process equipment size and energy consumption but also the final product quality.

Option 1: Apply a single flash after FTS reaction unit, in order to separate light

components (C7-) with FTS wax. Then, send wax to hydrocracking reactor.

1_{Julius Scherzer, A.J.Gruia ,Hydrocracking Science and Technology, 1996} Uncovered Syngas Syngas First stage F-T Reactor First stage F-T Reactor Syngas Second stage F-T Reactor Flash Liquid Hydrocarbon Catalyst removal To separation Catalyst Clean Wax

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-7-Diesel Waxrecycles C7-C7+ HC reactor Fuel gas Naphth Kerosene LPG F-T reaction Unit

Figure 2.4 Product separation unit option1

Option2: All the products of FTS unit will be sent to the distillation column

firstly, only the separated wax flows to the HC reactor. All the products of hydrocracker will flow to a flash and split to cracked hydrocarbon, which will flow to the distillation column, and the unconverted wax back to the hydrocracker again.

Figure 2.5: Product separation unit option2

Option 3: Put two distillation units separately: one is after FTS reaction unit, the

other is after Hydrocracking reaction unit. The advantage of this option is that no cracked hydrocarbon back mixes with heavy feed. On the other hand, two separate distillation units will spend client a lot of money.

H2recovery F-T reaction Unit _HC reactor Unconverted wax Cracked hydrocarbon Wax Fuel gas Napht Kerosen Diesel LPG

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Group Conceptual Process Design Project CPD_3296 Final report -8-H2 recovery Wax F-T Reaction Unit HC reactor Wax Fuel gas Naphth Kerosene Diesel LPG Fuel gas Naphth Kerosene Diesel LPG

Table 2.1 Comparisons of different options

Criteria Option 1 Option 2 Option 3

Operability + + + Capital cost + + - Innovative design - + - Product quality + - - Total number ++ ++ - Note:

+ Positive for certain criteria - Negative for certain criteria

After the rough comparison, we can say that the third option is not a wise choice. For option 1 and 2, it is hard to make a decision. Therefore, we will come to ASPEN to simulate two processes individually. (See ASPEN file HCoption_1.BKP and HCoption_2.BKP).

The main function of hydrocracker is the heavy paraffin cracking. Some other reactions also happened, such as hydrogenation of olefins and removal of the small amounts of oxygenates.

Comparing option 1 and option 2, the main difference is whether the C7-C20 FTS products go to hydrocracker or not. The advantage of option 1 is that nearly all the alkenes and small oxygenates can be removed, which improves the quality of oil products. On the other hand, all the light FTS waxes, which have large volume, will go through hydrocracker. It will end up with a huge reactor and for option2, vice versa. So the question is whether the oil products from option 2 can satisfy our required quality or not. Therefore, we come to ASPEN to find the result.

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-9-The results from ASPEN simulation are shown in table 2.2. We can see that the only thing we worried has been solved and all the requirements are satisfied for option 2. Thus, we can say that Option 2 is a very good choice for us. Unfortunately, we have done nearly all the calculation based on the traditional process, which is option1. We are not going to change that. But we designed both systems and made economic evaluation, which can be found in Chapter 8.

Table 2.2 Comparison of different options in ASPEN

Product HCoption_1 HCoption_2

Total Production

[ton/yr] 500000 492969.6 50209372.8

Main Product quality Kerosene 5% ASTM D86 [°C] 95% ASTM D86 [°C] Diesel 5% ASTM D86 [°C] 95% ASTM D86 [°C] Product Distribution [wt%] Naphtha Kerosene Diesel 185 290 240 350 184.9 289.9 240.1 350.1 26.8 34.6 38.5 185.1 289.9 240.0 350.3 21.9 37.7 40.4

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Final report

-10-3. Basis of Design (BOD)

Summary

A basis of design has been made for this conceptual process design project. In this project, natural gas is used as feedstock to produce syngas, which is going to be converted into synthetic oil products. Among them, the target products are diesel (C15-C20), and kerosene (C10-C14). Naphtha (C5-C9) and LPG (C2-C4) will be accepted as by-products.

According to the requirements, the chosen process consists of syngas production, Fischer-Tropsch synthesis process, hydrocracking, and Separation, altogether four unit operations. Combined autothermal reforming is applied to convert natural gas into syngas, which is the feedstock of Fischer-Tropsch synthesis process. In order to improve product quality and quantity, hydrocracking is placed after Fischer-Tropsch synthesis. Finally, diesel, kerosene, LPG and Naphtha will be separated respectively by separation operation.

The Anderson Flory Schulz distribution is used to calculate the amounts of hydrocarbon formed from the syngas feed, the CO: H2 ratio is around 1:2.

Economic evaluation results are summarized as follows:

Table 3.1 Economic evaluation results summary

Product/Feedstock Price ($/ton) Amount (ton/a) Profit ($/a)

1 LPG 154.80 0.00 0

2 Naphtha 130.00 1.32E+05 1.71E+07

3 Kerosene 135.00 1.71E+05 2.31E+07

4 Diesel 120.00 1.90E+05 2.28E+07

5 Natural gas -92.50 6.61E+05 -6.11E+07

6 Steam -18.55 1.86E+01 -3.44E+02

7 Oxygen -27.00 4.61E+05 -1.24E+07

Total (Economic Margin) -1.05E+07

It can be seen that economic margin is negative, which directly shows that this process is not profitable. From this table, a brief conclusion can be drawn that under current conditions, the production of transportation fuel from natural gas using Fischer-Tropsch synthesis technology is still not economically feasible. However, this is just based on this simple calculation, and as the crude oil price increases, it could be a promising technology for the future transportation fuel production. Moreover, optimization and creative design will be made in the following part.

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-11-3.1 Description of the Design

The objective of this conceptual process design project is to design a plant producing 500,000 tons/annum synthetic oil products out of natural gas, using Fischer-Tropsch technology. The plant will be located in a remote area: Brunei, South-East Asia. The target products are diesel (C15-C20), and kerosene (C10-C14). Naphtha (C5-C9) and LPG (C2-C4) will be accepted as by-products. For the 500,000 t/a capacity diesel, kerosene and naphtha are included. And the natural gas will be transported to the site from the well by pipeline.

This design will be used for the comparison to an alternative design made in the past. Therefore, price levels related to 1999 should be used. Also no information from previous T.U. Delft design efforts on Fischer-Tropsch plants should be used. In addition only literature information regarding conversion technologies from 1998 and before should be used. Any information regarding technical developments after 1998 should be discarded. This will allow a fair comparison between this and the alternative design.

Project principal is Ir. Pieter Swinkels, from Process Systems Engineering, DelftChemTech, T.U.Delft, and his assistant, Austine Ajah. And Cristhian Almeida Rivera is response for creativity and group process coaching.

3.2 Process Definition

3.2.1 Process concepts chosen

In chapter 2, we discussed all the possible process options for every unit operations. The results are shown in table 3.2.

Table 3.2 Summary of process options and selections

Unit Issue Conclusions

Oxygen supply Build an oxygen plant to supply pure oxygen CO2 recycle CO2 recycled to the secondary reforming zone Syngas purification

sequence H2 separation ! water removal ! CO2 removal

U100

Pure hydrogen

separation route Partial syngas hydrogen recovery Conversion in FTS

unit Two-stage slurry reactors without syngas recycle

U200

Catalyst and wax

separation Extraction process

U300 and U400

Sequence of U300

and U400 U300 (Hydrocracking unit) followed by U400 (separation unit)

In chapter 8, we explained all the detailed information to size the reactors, which are including reaction stoichiometry, kinetics and catalysts chosen. The summary of those in formations per unit operation is shown as followed.

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Final report

-12-" U100 Syngas production unit Catalyst chosen in U100 is Ni/Al2O3

Table 3.3 Reaction stoichometry

Reactions Reaction stoichiometry

Combustion

( )

4 2 2 2 2 2 CH g + O g ! CO g + H O g CO2 Reforming

_{( )}

4 2 2 2 2 CH g +CO g ! CO g + H g Steam Reforming

( )

4 2 3 2 CH g +H O g ! CO g + H g Water gas shift

_{( )}

2 2 2

CO g +H O g ! CO g +H g Table 3.4 Reaction kinetics

Reaction Reaction rate Rate constant [m6_kg cat-1mol-1] Combustio n 2 2 CH4 O2 cb cb p p T R k r = ref ₌₀_.₀₈ cb k CO2 Reforming



_









−

=

2 4 2 2 4 2 2 2 2 2 2

1

CO CH cr H CO CO CH cr cr

p

K

T

R

p

T

R

k

r

ref =0.051 cr k Steam Reforming



_









−

=

O H CH sr H CO O H CH sr sr

p

K

T

R

p

T

R

k

r

2 4 2 2 4 2 2 3 2 2

1

ksrref =0.128 Water Gas Shift



_









−

=

O H CO gw H CO O H CO gw gw

_K

_p

p

T

R

k

r

2 2 2 2

1

2 2 kgwref =0.073 Where:

Kcr - Carbon dioxide reforming equilibrium constant [Pa2]

Kgw - Gas water shift equilibrium constant [Pa2]

kj - Reaction rate constant of reaction j [m6kgcat-1mols-1]

Ksr - Steam reforming equilibrium constant [-]

pi - Partial pressure of component i [Pa]

R - Gas constant [Jmol-1_°C-1_]

rcr - Carbon dioxide reforming reaction rate [molkgcat-1s-1]

rgw - Water gas shift reaction rate [molkgcat-1s-1]

rj - Reaction rate of reaction j [molkgcat-1s-1]

rsr - Steam reforming reaction rate [molkgcat-1s-1]

T - Temperature [°C]

The influence of temperature on the rate constant will be described by the Arrhenius equation.         − −

=

ref i A T T R E ref i i

k

e

k

1 1 , (8.14)

(20)

-13-Where,

Ea - Activation energy [Jmole-1_]

kj - Reaction rate constant of reaction j [m6kgcat-1mols-1]

R - Gas constant [Jmol-_1°C-1_]

The equilibrium constants as function of temperature for the steam reforming, carbon dioxide reforming and the water gas shift reaction are given in table 3.5 [Twigg, 1989].

Table 3.5 Equilibrium constant 2

CO2 reforming: gw sr cr

_K

K

=

(8.15) Steam reforming:

(

)

( ) ( ) ( ) ( 0.2513 0.3665 0.58101 27.1337 3.2770) 2 5

10 01325

.

1 ⋅

−ΨΨΨ Ψ− − + −

=

e

K

_sr (8.16) Water gas shift: ( ) ( ) (ΨΨ 0.63508−0.29353Ψ+4.1778+0.31688) =e K_gw (8.17) Where, Ψ is given by: 1 15 . 273 1000 ₋ + = Ψ T (8.18) T - Temperature [°C]

" U200 Fischer-Tropsch Synthesis unit Catalyst chosen in U200 is Co/MgO/SiO2

Table 3.6 Main reactions in Fischer-Tropsch synthesis unit

Reaction Stoichiometry 0

(

)

,298 / ∆Hr kJ mol Paraffin _{nCO g}

( ) (

₂_n ₁

) ( )

_{H g}₂ _{C H}_n ₂_n ₂

( )

_{g l}_/ _{nH O g}₂

( )

+ + + → + -165 Olefins _{nCO g}

_{( )}

₊₂_{nH g}₂

_{( )}

_→_{C H}_n ₂_n

_{( )}

_{g l}_/ ₊_{nH O g}₂

_{( )}

-165

The kinetic equation in Fischer-Tropsch unit is first order in hydrogen concentration: H RT E H

A

m

e

C

r

A −

=

(8.22) Where

A - Pre-exponential factor [m3 m3catalyst.s-1] CH - Liquid phase hydrogen concentration [mol m 3]

Ea - Activation energy [J/mol] m - Hydrogen distribution coefficient [mol mol-1] R - Gas constant [Jmol-1°C-1]

RH - Reaction rate with respect to hydrogen [mol m-3catalyst.s-1]

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Final report

-14-" U300 Hydrocracking unit

Catalyst used in this unit is Pt/Y-Zeolite.

Reactions happened in hydrocracker are summarized in table 3.7

Table 3.7 Reactions in Hydrocracking Unit 3

Reaction Stoichiometry Paraffin hydrocracking Hydro-isomerization Hydrogenation of olefins 2 2 2 2n + → n n+ nH H C H C Reduction of oxygenates _C _H _O _H _C _H _H _O n n n n 2 +2 + 2 → 2 +2 + 2

The kinetics in hydrocracking unit can be separated into three parts: cracking of paraffins, Isomerisation of paraffin and Conversion of FTS by-products.

1. Cracking of paraffins

∑

+

−

=

N j j ij J i i i

k

C

k

P

C

R

(8.23) Where:

Ci - Molar concentration [moli/m-3L]

ki - Reaction rate constant [m3Lkgc-1s-1]

Pij - Probability of ith component formation from the jth component [-]

Ri - Rate of reaction [molkgc-1s-1]

The reaction rate constant is described as:

RT Ea

e

k

=

0 − (8.24) Where:

k - First order reaction rate constant [molalkane_feedkgcat-1s-1]

k0 - Pre-exponential reaction rate term [molalkane_feedkg-1cats-1]

Ea - Activation energy [Jmol-1]

R - Gas constant [Jmol-1_°C-1_]

(

)

5 0 1.12 10 c 6 k = ⋅ ⋅ N − (8.26) Where Nc - carbon number 2. Isomerisation of paraffins

Since no data are available on product isomer distributions, we assume a 50 % branched to normal ratio for the product mixture, where the branched fraction consists of 2-methyl alkanes only22_.

3. Conversion of FTS by-products

For both hydrogenation of olefins and reduction of oxygenates, complete conversion can be assumed.

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-15-3.2.2 Block Schemes

As stated earlier, syngas production unit, Fischer-Tropsch synthesis unit, hydrocracking unit, and product separation unit, four unit operations are involved in this design project. And please see details in the following figure. The data on stream specification of each stream is from Aspen simulation. Moreover, the dash line indicates the battery limit of this plant design.

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G ro u p C onc ept u al P ro ce ss D es ign P roj ec t CP D_ 3 2 96 F ina l r epo rt Fi g u re 3. 1 B lo ck S che me o f t h e p ro ce ss P ro je ct ID Nu mb er : C P D _3 296 C o m p le ti o n D ate : D ec. 1 2, 2 00 3 T ot al Inpu t: 1, 613 ,491 .2 t/ a (3 .27) T ot al Ou tp ut : 1, 613 ,491 .2 t/ a (3 .27) 604, 281 .6 t/ a (1 .23) F-T Sy nt he si s 523 K 30 b ar Se par at io n 427-720 K 1-2 ba r Hy dr o-C rack ing 610~ 710 K 100~ 150 ba r Sy ngas gener at io n 1473 K 20 b ar W ax <4 08 > 106, 934 .4 t/ a ( 0. 22) Di es el < 407> 189 ,6 19 .2 t/ a (0. 038 ) Ke ro se ne < 406> 17 1, 41 7. 6 t/a (0. 35 ) W as te W at er < 413> 967, 708 .8 t/ a ( 1. 96) W ax R ecy cl e < 412> 101, 606 .4 t/ a ( 0. 21) H 2 O <121> 331, 718 .4 t/ a (0 .67) W as te W at er < 229> 627, 148 .8 t/ a ( 1. 27) W as te W at er < 405> 8, 84 1. 6 t /a ( 0. 02 ) W ax pu rg e < 410> 5, 35 6. 8t /a ( 0. 01) Na pht ha < 40 4> 1 31, 904 t/ a (0. 27 ) St ea m < 10 2> 4 83 ,5 80 .8 t/ a (0 .9 8) O xy gen <103> 4 60, 425 .6 t/ a ( 0. 93) N at ur al ga s < 10 1> 66 0, 816 t/ a ( 1. 34) Sy ng as <124> H C < 233 > L igh t H C < 228 > 24, 537 .6 t/ a (0 .0 5) Fuel g as < 414> 132 ,393 .6 t/ a (0 .2 7) H 2 <118> 18 ,374 .4 t/ a (0 .03) Fuel g as < 232> 98, 38 0. 8 t /a (0 .20) CO 2 Pur ge < 125> 15, 091 .2 t/ a ( 0. 03) 1, 23 9, 609 .6 t/ a (2 .5 1) 48, 9542 .4 t/a ( 0. 99 ) Batte ry L im it H2 purge < 307> 5, 241 .5 t/ a (0. 01) H2 r ec yc le < 301> 47, 17 4. 4 t /a (0 .10) C O2 re cycl e <126> 2 85, 696 t/a (0 .5 8) Fu el ga s < 403 > 4, 204 .8 t/ a (0. 01) H C < 303 > 591, 120 t/ a (1. 20 ) H C pr oduc t < 311 > St eam < 401> 8 ,668 .8 t/ a ( 0. 02) H2 r ecycl e < 30 2> 65, 57 7. 6 t /a (0 .1 3)

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-17-3.2.3 Thermodynamic properties

The detail thermodynamic properties and reaction kinetics is in chapter 4 and 8. $ Operating windows

Combined U100, U200 and U300, the summary of the total process is shown in the below table.

Table 3.8 Operating windows for the whole process

Unit Catalyst Reactor Temperature (K) Pressure (Bar)

Ni/Al2O3 Multi-tube 1000 20

1 Syngas

production _{- Fixed}_bed_1473-1573₂₀

2 FT synthesis Co/MgO/SiO2 Slurry 493-523 30 3 Hydrocracking Pt/Y-Zeolite Multi-fixed

bed 623-913 30-70

$ Property model methods in ASPEN

When we did the simulation in ASPEN, we divided the whole process into four parts, which are syngas production unit, FT synthesis unit, hydrocracking unit and separation column. Each unit set up its own property method that is summarized in the following table, and the total simulation property is Peng-Robinson method.

Table 3.9 Property method of each unit

Unit Unit Name Property Method

U100 CAR reactor PR-BM

U200 FT reactor PRMHV2

U300 Hydrocracking PENG-ROB

U400 Distillation PRMHV2

$ Reaction kinetics

The reaction kinetics is divided into three parts, which are CAR reactor, FT reactor and hydrocracking reactor. The detail is in chapter 8.

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Final report

-18-3.2.4 Pure component properties

For this CPD design project and manual calculations, the properties of pure components involved in this designing project are very useful. This section presents the most representative components’ properties, such as technological data (Boiling point, Melting point, etc), safety and health data (Explosion limits, Maximum allowable concentration, etc.). All the properties can be found in

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-19-3.3 Basic Assumptions

3.3.1 Plant capacity

The objective of this conceptual process design project is to design a plant producing 500,000 tones/annum synthetic oil products using Fischer-Tropsch synthesis technology.

Figure 3.2 Feedstock and products sketch map

As shown in the above figure, we have three material feedstock streams, four product streams, and three side product streams. The specific flow rate of each stream is given in the following table.

Table 3.10 Stream summaries of feedstock and products

Product/Feedstock Amount (ton/a) Profit ($/a)

S8_LPG-sep. 0.00 0

S9_Naphtha-sep. 1.32E+05 1.71E+07

S10_Kerosene-sep. 1.71E+05 2.31E+07

Product

S11_Diesel-sep. 1.90E+05 2.28E+07

S1_Natural gas feed 6.61E+05 -6.11E+07

S2_Steam feed 1.86E+01 -3.44E+02

Feedstock

S3_Oxygen feed 4.61E+05 -1.24E+07

S17_Wastewater 967,715,00 -

S19_CO2 removal - -

Side product (waste)

S20_Purge - -

Total - - -1.05E+07

Regarding economical plant life, this plant will be operated for 15 years and has 2-year construction time, as agreed by our client, due to the big investment consideration.

3.3.2 Plant location

The location of this plant is set in a remote area: Brunei, South-East Asia. Brunei is the fourth-largest producer of LNG in the world and the third-largest natural gas producer in Southeast Asia. Because of convenient geography location and high quality in natural gas, it is profitable to have a trading between Malaysia,

(27)

Final report

-20-China, Japan and other countries in Asia according to growing demands for transport oil. The client has provided the feedstock composition. The natural gas contains a high percentage of methane, which is good for the syngas production, and it has been desulphurised at the well, so we don’t need to add a unit for the desulphurization of natural gas.

3.3.3 Battery limit

As shown in Figure 3.1, there are four main units applied in this plant, and the battery limit (dash line) has defined an imaginary fence around this plant. What inside and outside this battery limit, are described as follows.

$ Inside battery limit

Syngas production unit: it consisted of a Gas Heated reforming, which

contains two reactors, primary reformer (steam reforming) and secondary reformer (autothermal reforming), hydrogen separator and carbon dioxide remover.

Fischer-Tropsch synthesis unit: there are four reactors to covert the

syngas to hydrocarbon. The syngas will be split to two same reactors firstly, all the products of first two reactors will be separated in a single flash, the light unconverted syngas part will be transferred to the third reactor. And analogously, we have the fourth reactor, in order to achieve higher conversion. The overall conversion of syngas to hydrocarbon is 94.1%. Simultaneously, water will be removed from the gas mixture.

Hydrocracking unit: the wax will be cracked in a fixed bed reactor, and

the products will be separated into two parts, light one is the cracked hydrocarbon, which will be separated again in the separation unit. The heavy one is unconverted wax and will be recycled to crack again.

Seperation unit: the gas mixture from FT and hydrocracking unit will be

separated to six parts according to their different relative volatilities, hydrogen, fuel gas, LPG, naphtha, diesel, kerosene and wax. The wax and hydrogen will be sent to the hydrocracking unit. Regarding the produced fuel gas, it will be used in our factory, due to the fact that we need large amount of heat to increase the temperature of some reactors and streams. And the LPG, naphtha, diesel, and kerosene will be sold as products or by products.

$ Outside battery limit

The facilities outside the battery limit: In our factory the following four facilities will be needed: electricity, oxygen, steam and cooling water.

3.3.4 Definition In- and Outgoing streams

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-21-Regarding specific feedstock specification, the detailed data are available in

Appendix 3, which is provided by our client. The amount needed as feedstock is

also summarized as below:

Table3.11 the flow rate of and price of the feedstock

Steam No. S1 S2 S3

Steam name Natural gas feed Steam feedOxygen feed Flow rate kg/s 2.29E+01 1.68E+01 1.60E+01 Flow rate ton/a 6.60E+04 4.84E+04 4.61E+04 Price $/ton 9.25E+01 1.86E+01 2.7E+01 $ Product:

We have two main products Kerosene and diesel, and two by-products LPG and Naphtha. With respect to product composition, please find them Appendix 3.2.

Table3.12 the flow rate of and price of the products

Steam No. S8 S9 S10 S11

Steam name LPG Naphtha Kerosene Diesel Flow rate kg/s 3.25E+01 4.98E+00 7.37E+00 4.70E+00 Flow rate ton/a 9.54E+05 1.46E+05 2.16E+05 1.38E+05 Price $/ton 1.55E+02 1.30E+02 1.35E+02 1.20E+02 $ Wastes:

All of the waste of our factory should satisfy the Europe emission standard.

1. Wastewater 2.CO2 purges

$ Utilities:

1. Steam 2.Electricity 3. Cooling water $ Catalysts:

There are four kinds of catalysts used in the whole process, and the location is list in the table 3.7:

Table 3.13 List of Catalyst and the relative applied unit

Name Ni/ Al2O3 Ni/ Al2O3 Co/ Al2O3 Pt/ Zeolite

Apply unit Syngas production F-T

synthesis Hydrocracking Reaction

name reforming Steam Auto thermal reforming synthesis F-T Hydrocracking Shape 4-hole

cylinder cylinder with 4-hole domed ends

- Zeolite Bulk density 1100 kg/m3 1000 kg/m3 0.27 g/cc -

(29)

Final report

-22-3.4 Economic Margin

3.4.1 Calculation of economic margin

The Economic Margin can be calculated from the following equation:

Margin = Total Value (Products, Waste OUT)

- Total Value (Feedstock's, Process Chemicals, IN)

_._(3.1)

Within our Process, there are three feed streams, which are S1-Natual gas feed, S2-steam feed and S3-Oxygen feed; four products streams, which are S8-LPG, S9-Naphtha, S10-Kerosene and S11-Diesel. Considering each stream, and substituting the corresponding values into equation, we can get the margin. In this stage, utilities, capital cost and labor cost, etc. have not been taken into account.

Table 3.14 Economic margin breakdowns

Feedstock Products

Steam No. S1 S2 S3 S8 S9 S10 S11

Steam name Natural gas Steam Oxygen LPG Naphtha Kerosene Diesel flow rate kg/s 2.30E+01 1.71E+01 1.60E+01 0.00E+00 4.58E+00 5.95E+00 6.58E+00 flow rate ton/a 6.61E+05 1.86E+01 4.60E+05 0.00E+00 1.32E+05 1.71E+05 1.90E+05 Price $/ton 92.500 18.550 27.000 154.800 130.000 135.000 120.000 Cost $/a 6.11E+07 3.44E+02 1.24E+07 0.00E+00 1.71E+07 2.31E+07 2.28E+07

Total 7.36E+07 6.30E+07

Margin $/a -1.05E+07

From the results shown above, we can see that our calculated margin is equal to –10.5 million $/yr. This negative value indicates that basically we cannot earn money by rough evaluation.

3.4.2 Calculation of maximum allowable investment

DCFROR is the economic criteria to judge if a project can be economically feasible during lifetime; the definition of DCFROR is shown below:

1

0 (1

)

(

,

17)

= =

=

+

=

∑

n t n n

NFV

DCFROR

t the life of the project in our case

(3.2)

The NFV is the Net Future value, which is equal to the Margin as we just calculated: -1.05E+07 $/yr. Because the NFV is negative, there is not any possibility to earn money. From pure economic opinion, this factory should not be constructed. So there is no meaning to calculate the DCFROR. From list of feedstock and products, it also can be seen that this process is quite hard to earn money, due to considerable small price difference between feedstock and the desired products. However, this technology from natural gas to transportation oil is still promising, as the price of crude oil increases.

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-23-4. Thermodynamic Properties and Reaction Kinetics

4.1 Operating Windows

4.1.1 Syngas production unit

As mentioned before, in syngas production unit operation, we have chosen the combined autothermal-reforming reactor (CAR) and there are two main reactions that are primary reforming (steam methane reforming reaction) and second reforming (autothermal reforming reaction). The main and side reactions are listed in table4.1 about steam reforming and partial oxidation.

Table 4.1 Reactions in syngas production from methane:

No. Reaction stochimometry 0

(

)

,298 / r H kJ mol ∆ 4 4.1 CH g4

( )

+H O g2

( )

! CO g

( )

+3H g2

( )

206 4.2 CO g

( )

+H O g2

( )

! CO g2

( )

+H g2

( )

-41 4.3 CH g4

( )

+0.5O g2

( )

! CO g

( )

+2H g2

( )

-36 4.4 CH g4

( )

+CO g2

( )

! 2CO g

( )

+2H g2

( )

247 4.5 CH g4

( )

+2O g2

( )

! CO g2

( )

+2H O g2

( )

-803 4.6 2CO g

( )

→CO g2

( ) ( )

+C g -173 4.7 CO g

( )

+0.5O g2

( )

! CO g2

( )

-284 4.8 H g2

( )

+0.5O g2

( )

! H O g2

( )

-242 4.9

CH g

₄

( )

!

C

+

2 H g

₂

( )

75 " Primary reforming (SMR)

Reaction 4.1 and 4.2 are steam-reforming reaction (SMR). All the components are calculated in equilibrium. In ASPEN the property method is PR-BM that is recommended in Aspen Plus 11.1 user guide5_{. The conditions are P=1 bar and} the mole fraction of CH4/H2O=1 mole/mole. The result is shown in figure 4.1, which is the equilibrium gas composition of the reaction of methane with steam as a function of temperature.

4_{Jacob A. Moulijn, Chemical process technology, 2001, p133}

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Group Conceptual Process Design Project CPD_3296 Final report -24-0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 500 1000 1500 2000 Temperature (K) mola r f ra c tion H2_CO CO2 H2O CH4

Figure 4.1 Equilibrium gas composition at 1 bar as a function of temperature

(CH4/H2O=1 mole/mole)

Figure 4.1 shows that the reaction is highly endothermic and should be carried out at high temperature (>1000K). This is obvious that at high temperature only H2 and CO is present and the ratio of H2/CO is 3.

To look at the effect of pressure on the equilibrium gas composition in steam reforming of methane, we compare the reaction at 1bar and 20bar with H2O/CH4=1 mole/mole. The result is shown in figure 4.2.

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 500 800 1100 1400 1700 2000 Temperature (K) m o lar f ract io n H2 CO CO2 H2O CH4 H2, 20bar CO,20bar CO2,20bar H2O, 20bar CH4, 20bar

Figure 4.2 Effect of temperature and pressure on equilibrium gas composition in

steam reforming reaction with H2O/CH4=1 mole/mole.

H2 CO CH4~H2O CO2 H2 CO CH4~H2O CO2

(32)

-25-Figure 4.2 shows that when increasing pressure steam reforming reaction is hindered. At 20bar, the equilibrium conversion to H2 and CO is only complete at a temperature of over 1400K. However in steam reforming zone, we don’t need that kind of high conversion and only 28% of methane attends the reaction, so we can choose lower temperature (1000K) as the reaction temperature.

" Second reforming (ATR)

For autothermal reforming, the reactions 4.3, 4.7 and 4.8 are considered to realize this kind of reactions. All the mole fractions of all components are calculated in equilibrium. The property method for thermodynamics in ASPEN is still PR-BM because we considering SMR and ATR are in the same reactor. The conditions are P=1bar and O2/CH4=0.756 mole/mole. Because this is a combined autothermal reforming, the result from steam reforming reaction is the reactants for the autothermal reaction. So the mole fraction of the components should be calculated using mass balance. The mole fractions are listed in table 4.2.

Table 4.2 Mole fractions of all the components in autothermal reaction.

Components H2 CO CO2 CH4 O2 H2O Total

Mole fraction 31.3126 10.4257 0.4977 32.5876 24.6427 0.5219 100

Ratio of H2/CO 3.00

The result is shown in figure 4.3 that is the equilibrium gas composition of the autothermal reaction as a function of temperature at P=1bar.

0 0.1 0.2 0.3 0.4 0.5 0.6 700 1000 1300 1600 1900 Temperature (K) mo le f ra ct io n O2 H2 CO H2O CH4 CO2

Figure 4.3 Equilibrium gas composition at 1 bar as a function of temperature for

autothermal reaction with O2/CH4=0.756 mole/mole.

From figure 4.3 we can see that the reaction is also at high temperature (>1300K) and methane is almost completely converted. To compare the effect of

H2

CO

H2O

CH4

(33)

Final report

-26-pressure on the reaction, we select the -26-pressure at 1bar and 20bar and the comparison is shown on figure 4.4.

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 700 900 1100 1300 1500 1700 1900 Temperature (K) m o le fr a c ti o n O2 H2 CO H2O CH4 O2,20BAR H2,20BAR CO,20BAR CO2,20BAR H2O,20BAR CH4,20BAR CO2

Figure 4.4 Effect of temperature and pressure on equilibrium gas composition in

autothermal reaction with O2/CH4=0.756 mole/mole.

Obviously at higher pressure, the reaction temperature is higher too.

4.1.2 Fischer-Tropsch Unit

Within this process unit operation, two stages slurry reactors are applied to convert syngas into hydrocarbon. The main reactions are summarized in the table4.3.

Table 4.3 Main reactions in Fischer-Tropsch synthesis unit

Reaction Stoichiometry 0

(

)

,298 / ∆Hr kJ mol Paraffin

_{( ) (}

_{) ( )}

_{( )}

2 2 2 2 2 1 _n _n / nCO g + n+ H g →C H ₊ g l +nH O g -165 Olefins

_{( )}

2 2 2 2 _n _n / nCO g + nH g →C H g l +nH O g -165 Among literatures on the kinetics and selectivity of the Fischer-Tropsch synthesis, most studies aim at catalyst improvement and postulate empirical power law kinetics and assume a simple polymerization reaction following an Anderson-Schulz-Flory (ASF) distribution for the total hydrocarbon product yield. ASF distribution formula is expressed as:

2 n 1 n n n

w

(1

)

m

(1

)

n

;

−

− α

= − α α

=

α

(4.10)

Where the growth probability factor α is independent of n, and m is the mole _n fraction of a hydrocarbon with the chain length n. The range of α is dependent

H2 CO H2O CO2 CH4 At 1bar At 20bar

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-27-on the reacti-27-on c-27-onditi-27-ons and catalyst type. In our design, cobalt was chosen as catalyst in Fischer-Tropsch reaction. Regarding cobalt as catalyst, the range of α

is defined between 0.70~0.95 for operating condition, T=523K and P=30bar. And high pressure or low temperature can shift products composition to heavy product, which means we will have a higher value for α accordingly. Since our designed FT system will work on 523 K and 30 bars, α is given 0.92. From figure 4.5, it also can be proven that 0.92 is a good choice. It can meet our requirement that the heavy product can be obtained as much as possible.

0 10 20 30 40 50 60 70 80 90 100 0.0 0.2 0.4 0.6 0.8 1.0

Chain growth probability, a

M a ss rat io, wt (%) _C1 C2~C4 C5~C9 C10~C14 C15~C20 C21~C45 C46~C100

Figure 4.5 Hydrocarbon selectivity as function of the chain growth probability

factor

Regarding the specific operating window, the following data can be given by literature6_:

Table 4.4 Operating windows for Fischer-Tropsch synthesis

Catalyst Reactor

Type Temperature (°C) Pressure (MPa) Hfeed ratio 2/CO

Co/MgO/SiO2 Slurry 220~250 1.5~3.5 1.5~3.5

4.1.3 Hydrocracking operation unit

In Hydrocracking unit, multi-fixed bed reactor has been applied. The typical reactions show in table 4.5.

6_{Kinetics, selectivity and scale up of the Fischer-Tropsch synthesis, chapter 2, P63, 1999 (note:} the reactor conditions designed here is basing on the experiment data before 1998.)

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Final report

-28-Table 4.5 Reactions in Hydrocracking Unit7

Reaction Stochiometry

(

)

0 ,298 / ∆Hr kJ mol Alkanes hydrocracking -44 Hydro-isomerization -4

Due to too many reactions happen in this unit, we choose some data from literature to find out the operating window.

Table 4.6 hydrocracking yield response to reactor temperature8

Feed stock: Gravity, °API Nitrogen, wt% Sulfur, wt% Aniline point,°C 343-350°C Gas oil 23.6 1250 2.0 85 Unicracker reactor avg. temp. ,°C 376 367 360

Product objective: PC naphtha Turbine fuel Diesel

Yield, vol% feedstock C4 C5-60°C 60°C+ Naphtha 149°C+ Distillate Total C4+ C6-C8 C6-C9 C6 19.6 21.7 87.0 -- 128.3 29.7 37.9 41.0 8.9 11.3 45.7 54.1 120.0 18.7 23.1 23.04 4.2 6.3 28.5 75.2 114.2 13.2 16.5 16.6

From table 4.6, we can see that the target products can shift from naphtha to diesel with a decrease of reactor temperature. The total reaction is exothermic, and low temperature is favorable. On the other hand, in order to maintain the conversion constant, the operating temperature is gradually increased to make up for the loss in acidity.

Another important factor in process condition is the hydrogen partial pressure. Hydrogen partial pressure has a dual effect on catalytic cracking and isomerization. On one hand, an increase in pressure has a favorable effect due to enhanced hydrogenation of coke and cleaning of the catalyst surface. On the

7_{Julius Scherzer, A.J. Gruia, Hydrocracking Science and Technology, chapter6, p74, 1996} 8_{Julius Scherzer, A.J. Gruia, Hydrocracking Science and Technology, chapter11, p205, 1996}

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-29-other hand, the rate of cracking and isomerization reactions decreases with the increasing of hydrogen partial pressure.

Our target products are Kerosene and Diesel, which determines that our process condition should not be too severe, and belongs to mild hydrocracking. So, the typical process condition of mild hydrocracking is summarized in table 4.7.

Table 4.7 Typical process and operating condition for mild hycrocracking

Process One stage

Operating conditions Conversion wt% 20-70 Temperature °C 350-440 H2 pressure bar 30-70 LHSV h-1 _0.3-1.5 H2/oil Nm3/m3 300-1000

Thus, our operating window of hydrocracking unit is determined by this means.

4.1.4 Brief summary of operating windows

Combine the thermodynamic data of individual unit, the valid operating conditions are shown in table 4.8 for the total process.

Table 4.8 Operating windows summary for the whole process

Unit Catalyst Reactor Temperature (K) Pressure (Bar) Ni/Al2O3 Multi-tube 1000 20-40 1 Syngas

production _{- Fixed}_bed_1473-1573_20-40

2 FT synthesis Co/MgO/SiO2 Slurry 493-523 30 3 Hydrocracking Pt/Y-Zeolite Multi-fixed

bed

623-913 30-70

4.2 Heat data

The thermodynamic properties of components in ASPEN are shown in Appendix

4.3 where include the vapor, liquid and solid phase properties at constant

pressure (P= 1bar) and temperature (T= 273 K). The thermo properties are Cp, G, H, S, RHO, PL, viscosity (MU). Table 4.9 gives the vapor enthalpy and the heat capacity of the main feedstock and products.

Table 4.9 Vapor enthalpy and Heat capacity from ASPEN database

tb ∆vapH (tb) Cp Phase

Component Formula o_{C KJ/mol}_J/(mol_K)

Hydrogen H2 -252.87 0.9 Gas

Carbon monoxide CO -191.5 6.04 29.1 Gas

Carbon dioxide CO2 37.1 Gas