CPD NR 3296 Conceptual Process Design

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 FischerTropsch technology
Authors
Binbin Bai
Junying Hu
Nan Liu
Yan Jiao
Zhiyong Wang
(Study nr.)
1134345
1160842
1132016
1160915
1129767
Telephone
0641763172
0641763830
0641439516
0624560836
0618241976
Keywords
Fischer-Tropsch synthesis, Hydrocracking, Syngas
production, Combined autothermal reforming, Natural
gas.
Assignment issued
Report issued
Appraisal
:
:
:
Sep. 22, 2003
Dec. 15, 2003
Jan. 23, 2004
Group Conceptual Process Design Project
CPD_3296
Preface
Preface
-i-
Group Conceptual Process Design Project
CPD_3296
Summary
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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Group Conceptual Process Design Project
CPD_3296
Table of Content
Table of Content
Summary
………………………………………………………………………………………………………
II
1. Introduction ……………………………………………………………………………………………
1.1 Conceptual process design ………………………………………………………………
1.2 Project CPD_3296
…………………………………………………………………………
1.3 Fischer-Tropsch synthesis
……………………………………………………………
1.4 Brief process description
………………………………………………………………
1.5 Environment
…………………………………………………………………………………
1
1
1
1
2
2
2. Process Options and Selection …………………………………………………………
2.1 Syngas unit ……………………………………………………………………………………
2.1.1 Oxygen supply ………………………………………………………………………
2.1.2 Energy recovery method ………………………………………………………
2.1.3 Carbon dioxide recycle
………………………………………………………
2.1.4 Raw syngas purification operation sequence ………………………
2.1.5 Pure hydrogen separation route ……………………………………………
2.2 Fischer-Tropsch synthesis unit ………………………………………………………
2.2.1 The conversion in Fischer-Tropsch synthesis ………………………
2.2.2 Catalyst and wax separation of FT synthesis ………………………
2.2.3 Basic block scheme of FTS process
……………………………………
2.3 Process options of Hydrocracking unit ……………………………………………
3
3
3
3
3
4
4
5
5
5
5
6
3. Basis of Design
……………………………………………………………………………………
3.1 Description of the Design ………………………………………………………………
3.2 Process Definition ……………………………………………………………………………
3.2.1 Process concepts chosen ………………………………………………………
3.2.2 Block schemes …………………………………………………………………………
3.2.3 Thermodynamic properties
…………………………………………………
3.2.4 Pure component properties
…………………………………………………
3.3 Basic Assumptions …………………………………………………………………………
3.3.1 Plant capacity …………………………………………………………………………
3.3.2 Plant location …………………………………………………………………………
3.3.3 Battery limit …………………………………………………………………………
3.3.4 Definition In- and Outgoing streams
…………………………………
3.4 Economic Margin ………………………………………………………………………………
3.4.1 Calculation of economic margin ……………………………………………
3.4.2 Calculation of maximum allowable investment ………………………
10
11
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15
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20
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22
4. Thermodynamic Properties and Reaction Kinetics ………………………
4.1 Operating windows ………………………………………………………………………
4.1.1 Syngas production unit …………………………………………………………
4.1.2 Fischer-Tropsch unit ……………………………………………………………
4.1.3 Hydrocracking operation unit ………………………………………………
4.1.4 Brief summary of operating windows ……………………………………
4.2 Heat data …………………………………………………………………………………………
4.3 Models for vapor/liquid equilibrium
……………………………………………
4.4 Reaction kinetics
…………………………………………………………………………
23
23
23
26
27
29
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30
30
5. Process Structure and Description …………………………………………………
5.1 Criteria and Selections …………………………………………………………………
5.1.1 Syngas production unit …………………………………………………………
5.1.2 FT synthesis unit …………………………………………………………………
5.1.3 Hydrocracking unit
……………………………………………………………
5.1.4 Separation unit ……………………………………………………………………
31
31
31
33
36
37
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Table of Content
5.2 Process Flow Scheme (PFS) ………………………………………………………………
5.3 Process Stream Summary ………………………………………………………………
5.4 Utilities ………………………………………………………………………………………………
5.4.1 Utility introduction
…………………………………………………………………
5.4.2 Pinch and heat exchanger network ……………………………………………
5.5 Process yields
…………………………………………………………………………………
38
39
40
40
40
41
6. Process Control ………………………………………………………………………………………
6.1 Syngas production unit (U100) ………………………………………………………
6.2 Fischer-Tropsch synthesis unit (U200) ……………………………………………
6.3 Hydrocracking unit (U300)
……………………………………………………………
6.4 Separation unit (U400) ………………………………………………………………………
43
43
44
46
47
7. Mass and Heat Balances
………………………………………………………………………
48
8. Process and Equipment Design …………………………………………………………
8.1 Integration by process simulation ……………………………………………………
8.2 Equipment selection and design ………………………………………………………
8.2.1 Syngas reactor design ……………………………………………………………
8.2.2 Reactor design of Fischer-Tropsch synthesis ………………………
8.2.3 Hydrocracking design ………………………… ……………………………………
8.2.4 Separation unit design
…………………………………………………………
8.2.5 Shell and tube exchanger design …………………………………………
8.2.6 Single flash column design ………………………………………………………
8.2.7 Pump and compressor design ………………………………………………
8.2.8 H2 embrane separation (S101) ………………………………………………
8.2.9 CO2 removal separator (S102) ………………………………………………
8.3 Equipment data sheets ………………………………………………………………………
49
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51
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56
57
57
58
58
9. Waste …………………………………………………………………………………………………………
9.1 Introduction ………………………………………………………………………………………
9.2 Waste treatment …………………………………………………………………………………
9.3 Emission limit values …………………………………………………………………………
9.3.1 Air emission limit value ………………………………………………………………
9.3.2 Water emission limit value ………………………………………………………
59
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60
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61
10. Process Safety
……………………………………………………………………………………
10.1 Hazard and operability studies (HAZOP) …………………………………………
10.1.1 Introduction of HAZOP
…………………………………………………………
10.1.2 HAZOP Analysis ……… ……………………………………………………………
10.2 Dow Fire and Explosion Index (F&EI) method ………………………………
10.3 Conclusion ………………………………………………………………………………………
62
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11. Economy
………………………………………………………………………………………………
11.1 Investment ………………………………………………………………………………………
11.2 Cash flow
………………………………………………………………………………………
11.3 Economic evaluation of the project …………………………………………………
11.4 Cost review ………………………………………………………………………………………
11.5 Sensitivities ………………………………………………………………………………………
11.6 Negative cash flows …………………………………………………………………………
66
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70
12. Process Safety
……………………………………………………………………………………
12.1 Group relation diagram ……………………………………………………………………
12.2 Group creativity evaluation ………………………………………………………………
12.3 Creativity implication in CPD ……………………………………………………………
12.4 Group process tools evaluation ………………………………………………………
72
72
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74
74
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Table of Content
13. Conclusions and Recommendations …………………………………………………
13.1 Conclusions ………………………………………………………………………………………
13.2 Recommendation ………………………………………………………………………………
77
77
78
List of abbreviation …………………………………………………………………………………………
List of symbols
…………………………………………………………………………………………
Literature …………………………………………………………………………………………………………
78
80
82
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Table of Content
Table of Content for Appendix
Appendix 1. ……………………………………………………………………………………………………
1.1 Overall process scheme
…………………………………………………………………
A1
A1
Appendix 2. ……………………………………………………………………………………………………
2.1 Oxygen supply evaluation.…………………………………………………………………
2.2 Conversion route
…………………………………………………………………………
2.3 Catalyst and wax separation in FTS process ……………………………………
A2
A2
A3
A5
Appendix 3. ……………………………………………………………………………………………………
3.1 Feedstock specifications .……………………………………………………………………
3.2 Product specifications
..……………………………………………………………………
3.3 List of prices for feedstock and product . …………………………………………
3.4 Physical Properties of Pure components…. …………………………………………
A8
A8
A9
A10
A10
Appendix 4. ……………………………………………………………………………………………………
4.1 The process for choosing a property method .…………………………………
4.2 Models for vapor/liquid equilibrium ……….…………………………………………
4.3 Thermodynamic properties in Aspen ..………………………………………………
A17
A17
A19
A22
Appendix 5. ……………………………………………………………………………………………………
5.1 Syngas production unit . …………………………………………………………………
5.2 Syngas ratio adjustment . …………………………………………………………………
5.3 CO2 removal technology .. …………………………………………………………………
5.4 Fischer-Tropsch synthesis design ….……………………………………………………
5.5 Process flow scheme (PFS) …..….…………………………………………………………
5.6 Process Stream Summary …………………..………………………………………………
5.7 Available utility conditions and costs ..………………………………………………
5.8Pinch Technology.. ………………………….……………………………………………………
5.9 Utility summary…………….……………………………………………………………………..
A27
A27
A32
A38
A42
A61
A62
A80
A81
A88
Appendix 6. ……………………………………………………………………………………………………
6 Process control……………………..…………………………………………………………………
A89
A89
Appendix 7. ……………………………………………………………………………………………………
7Heat and Mass Balance………..…………………………………………………………………
A90
A90
Appendix 8. ……………………………….…………………………………………………………………
8.1 Oxygen supply evaluation …….……………………………………………………………
8.2 Description of ASPEN simulation…….……………………………………………………
8.3 Process simulation scheme in ASPEN ….………..……………………………………
8.4 The kinetics of combined autothermal reforming (CAR) ……………………
8.5 Fisher-Tropsch reactor design………………………………..……………………………
8.6 Hydrocraking catalyst and kinetics and hydrocracker design….…………
8.7 Design procedure of distillation column……………………..………………………
8.8 Calculation for hydrocracker sizing……………………………..………………………
8.9 Equipment summary……………………………..……………………………….……………
8.10 Equipment specification……………………………..……………………………….……
A94
A94
A96
A98
A100
A105
A114
A123
A128
A131
A141
Appendix 9. ……………………………………………………………………………………………………
9.1 The HAZOP analysis………………………………………………………………………………
9.2 Dow Fire and Explosion Index analysis….……………………………………………
A169
A169
A182
Appendix 10. ……………………………………………………………………………………………………
10. The price estimate of the purchased equipment………………………………
A169
A187
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Group Conceptual Process Design Project
CPD_3296
Final report
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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Group Conceptual Process Design Project
CPD_3296
Final report
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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CPD_3296
Final report
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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CPD_3296
Final report
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.
Hydrogen
Hydrogen
Syngas
Syngas
Syngas
Figure 2.1 Total syngas hydrogen
recovery
Option1:
Syngas
Figure 2.2 Partial syngas hydrogen recovery
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
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Final report
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 unrecovery 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
system are needed within this unit. The specific conversion route and block
scheme of FTS process is shown the table below.
Uncovered
Syngas
Flash
To separation
Second
stage F-T
Reactor
First stage
F-T
Reactor
Liquid
Syngas
Clean Wax
First stage
F-T
Reactor
Hydrocarbon
Catalyst
removal
Syngas
Catalyst
Table 2.3 Block scheme of FTS process
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
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Final report
C7-
Fuel gas
F-T reaction
Unit
LPG
HC
reactor
Naphth
Kerosene
C7+
Diesel
Wax recycles
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.
Cracked hydrocarbon
Fuel gas
F-T reaction
Unit
H2 recovery
LPG
Napht
Kerosen
HC
reactor
Diesel
Wax
Unconverted wax
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.
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Fuel gas
H2 recovery
Fuel gas
LPG
F-T
Reaction
Unit
Naphth
LPG
Kerosene
Naphth
HC
reactor
Diesel
Kerosene
Diesel
Wax
Wax
Figure 2.6 Product separation unit option3
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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Group Conceptual Process Design Project
CPD_3296
Final report
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
Total Production
[ton/yr]
Main Product quality
Kerosene
5% ASTM D86 [°C]
95% ASTM D86 [°C]
Diesel
5% ASTM D86 [°C]
95% ASTM D86 [°C]
HCoption_1
HCoption_2
500000
492969.6
50209372.8
185
290
184.9
289.9
185.1
289.9
240
350
240.1
350.1
240.0
350.3
26.8
34.6
38.5
21.9
37.7
40.4
Product Distribution
[wt%]
Naphtha
Kerosene
Diesel
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Group Conceptual Process Design Project
CPD_3296
Final report
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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Group Conceptual Process Design Project
CPD_3296
Final report
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 (C10C14). 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
U100 Oxygen supply
Build an oxygen plant to supply pure oxygen
CO2 recycle
CO2 recycled to the secondary reforming zone
H2 separation ! water removal ! CO2 removal
Syngas purification
sequence
Pure hydrogen
Partial syngas hydrogen recovery
separation route
U200 Conversion in FTS
Two-stage slurry reactors without syngas recycle
unit
Catalyst and wax
Extraction process
separation
U300 Sequence of U300
U300 (Hydrocracking unit) followed by U400
and
and U400
(separation unit)
U400
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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Group Conceptual Process Design Project
CPD_3296
Final report
" U100 Syngas production unit
Catalyst chosen in U100 is Ni/Al2O3
Table 3.3 Reaction stoichometry
Reactions
Reaction stoichiometry
Combustion
CH 4 ( g ) + 2O2 ( g ) ! CO2 ( g ) + 2 H 2O ( g )
CO2 Reforming
CH 4 ( g ) + CO2 ( g ) ! 2CO ( g ) + 2 H 2 ( g )
Steam Reforming
CH 4 ( g ) + H 2 O ( g ) ! CO ( g ) + 3H 2 ( g )
Water gas shift
CO ( g ) + H 2O ( g ) ! CO2 ( g ) + H 2 ( g )
Table 3.4 Reaction kinetics
Reaction
Reaction rate
Combustio
n
rcb =
k cb
R 2T 2
Rate constant
[m6kgcat-1mol-1]
p CH 4 p O2
ref
k cb
= 0.08
CO2
Reforming
2

p CO
p H2 2
k cr

rcr = 2 2 p CH 4 p CO2 1 − 2 2

R T
R T K cr p CH 4 p CO2

Steam
Reforming

p CO p H3 2
k sr
rsr = 2 2 p CH 4 p H 2O 1 − 2 2

R T
R T K sr p CH 4 p H 2O

Water Gas
Shift
rgw =
Where:
Kcr
Kgw
kj
Ksr
pi
R
rcr
rgw
rj
rsr
T
-
k gw
R 2T 2

p CO2 p H 2
p CO p H 2O 1 −

K gw p CO p H 2 O













ref
k cr
= 0.051
ref
k sr
= 0.128
ref
k gw
= 0.073
Carbon dioxide reforming equilibrium constant
[Pa2]
Gas water shift equilibrium constant
[Pa2]
6
-1
Reaction rate constant of reaction j
[m kgcat mols-1]
Steam reforming equilibrium constant
[-]
Partial pressure of component i
[Pa]
Gas constant
[Jmol-1°C-1]
Carbon dioxide reforming reaction rate
[molkgcat-1s-1]
Water gas shift reaction rate
[molkgcat-1s-1]
Reaction rate of reaction j
[molkgcat-1s-1]
Steam reforming reaction rate
[molkgcat-1s-1]
Temperature
[°C]
The influence of temperature on the rate constant will be described by the
Arrhenius equation.
ki = k
ref
i
−
e
E A, i  1 1
 −
R  T Tref




(8.14)
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Group Conceptual Process Design Project
CPD_3296
Final report
Where,
Ea
kj
R
T
-
Activation energy
Reaction rate constant of reaction j
Gas constant
Temperature
[Jmole-1]
[m kgcat-1mols-1]
[Jmol-1°C-1]
[°C]
6
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
K cr =
CO2
reforming:
2
K sr
K gw
(8.15)
(
)
Steam
reforming:
K sr = 1.01325 ⋅ 10 5 e − ( Ψ ( Ψ ( Ψ (0.2513Ψ −0.3665 )−0.58101)+ 27.1337 )−3.2770 ) (8.16)
Water gas
shift:
K gw = e ( Ψ (Ψ (0.63508− 0.29353 Ψ )+ 4.1778 )+ 0.31688 )
2
(8.17)
Where,
Ψ is given by:
T
-
Ψ=
Temperature [°C]
1000
−1
T + 273.15
(8.18)
" 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
Paraffin
nCO ( g ) + ( 2n + 1) H 2 ( g ) → Cn H 2 n + 2 ( g / l ) + nH 2 O ( g )
Olefins
nCO ( g ) + 2nH 2 ( g ) → Cn H 2 n ( g / l ) + nH 2 O ( g )
∆H r0,298 ( kJ / mol )
-165
-165
The kinetic equation in Fischer-Tropsch unit is first order in hydrogen
concentration:
rH = A m e
Where
A
CH
Ea
m
R
RH
2
-
−
EA
RT
CH
(8.22)
Pre-exponential factor
[m3 m3catalyst.s-1]
Liquid phase hydrogen concentration
[mol m 3]
Activation energy
[J/mol]
Hydrogen distribution coefficient
[mol mol-1]
Gas constant
[Jmol-1°C-1]
Reaction rate with respect to hydrogen
[mol m-3catalyst.s-1]
Twigg M.V. 1989, Catalyst Handbook, Second edition, Wolfe Publishing Ltd. London 1989
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Group Conceptual Process Design Project
CPD_3296
Final report
" 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
Reaction
Paraffin hydrocracking
3
Stoichiometry
Hydro-isomerization
Hydrogenation of olefins
C n H 2n + H 2 
→ C n H 2 n+ 2
Reduction of oxygenates
C n H 2n+2 O + H 2 
→ C n H 2n + 2 + H 2 O
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
Ri = − k i C i + ∑ k j Pij C J
(8.23)
j
Where:
Ci
ki
Pij
Ri
-
Molar concentration
[moli/m-3L]
Reaction rate constant
[m3Lkgc-1s-1]
th
th
Probability of i component formation from the j component [-]
Rate of reaction
[molkgc-1s-1]
The reaction rate constant is described as:
k = k0e
−
Where:
k
k0
Ea
R
T
Ea
RT
(8.24)
-
First order reaction rate constant
Pre-exponential reaction rate term
Activation energy
Gas constant
Temperature
k0 = 1.12 ⋅105 ⋅ ( N c − 6 )
Where
Nc
[molalkane_feedkgcat-1s-1]
[molalkane_feedkg-1cats-1]
[Jmol-1]
[Jmol-1°C-1]
[°C]
(8.26)
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.
3
Julius Scherzer, A.J. Gruia, Hydrocracking Science and Technology, chapter6, p74, 1996
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CPD_3296
Final report
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.
-15-
Steam
Total Input:
1,613,491.2 t/a (3.27)
Steam <401> 8,668.8 t/a (0.02)
<102> 483,580.8 t/a (0.98)
Natural gas <101> 660,816 t/a (1.34)
Oxygen <103> 460,425.6 t/a (0.93)
591,120 t/a
(1.20)
48,9542.4
t/a (0.99)
HydroCracking
610~710 K
100~150 bar
Wax Recycle <412>
101,606.4 t/a (0.21)
HC <303>
H2 recycle <302>
65,577.6 t/a (0.13)
HC <233>
Waste Water <229>
627,148.8 t/a (1.27)
F-T
Synthesis
523 K
30 bar
Light HC <228>
24,537.6 t/a
(0.05)
H2 recycle <301>
47,174.4 t/a
(0.10)
-16-
Figure 3.1 Block Scheme of the process
1,239,609.6 t/a
(2.51)
Syngas<124>
Fuel gas <232>
98,380.8 t/a
(0.20)
H2O<121>
331,718.4 t/a
(0.67)
Syngas
generation
1473 K
20 bar
CO2 recycle
<126> 285,696
t/a (0.58)
H2 <118> 18,374.4 t/a
(0.03)
Battery Limit
Group Conceptual Process Design Project
CPD_3296
Final report
Waste Water <413>
967,708.8 t/a (1.96)
Waste Water <405>
8,841.6 t/a (0.02)
Wax purge <410>
5,356.8t/a (0.01)
Diesel <407> 189,619.2 t/a
(0.038)
Kerosene <406> 171,417.6 t/a
(0.35)
Total Output:
1,613,491.2 t/a (3.27)
Separation
427-720 K
1-2 bar
Naphtha <404> 131,904 t/a
(0.27)
Fuel gas <403>
4,204.8 t/a
(0.01)
Project ID Number: CPD_3296
Completion Date: Dec. 12, 2003
Wax <408>
106,934.4 t/a (0.22)
604,281.6 t/a
(1.23)
HC product
<311>
H2 purge <307>
5,241.5 t/a
(0.01)
Fuel gas <414> 132,393.6 t/a (0.27)
CO2 Purge <125> 15,091.2 t/a (0.03)
Group Conceptual Process Design Project
CPD_3296
Final report
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
1
Unit
Syngas
production
Catalyst
Ni/Al2O3
Reactor
Multi-tube
Temperature (K)
1000
Pressure (Bar)
20
-
2
3
FT synthesis
Hydrocracking
Co/MgO/SiO2
Fixed bed
Slurry
Multi-fixed
bed
1473-1573
493-523
623-913
20
30
30-70
Pt/Y-Zeolite
$ 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 PengRobinson method.
Table 3.9 Property method of each unit
Unit
Unit Name
U100
CAR reactor
U200
FT reactor
U300
Hydrocracking
U400
Distillation
Property Method
PR-BM
PRMHV2
PENG-ROB
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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Group Conceptual Process Design Project
CPD_3296
Final report
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
Appendix 3.4.
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Group Conceptual Process Design Project
CPD_3296
Final report
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
Side product (waste)
Total
Product/Feedstock
S8_LPG-sep.
S9_Naphtha-sep.
S10_Kerosene-sep.
S11_Diesel-sep.
S1_Natural gas feed
S2_Steam feed
S3_Oxygen feed
S17_Wastewater
S19_CO2 removal
S20_Purge
-
Amount (ton/a)
0.00
1.32E+05
1.71E+05
1.90E+05
6.61E+05
1.86E+01
4.61E+05
967,715,00
-
Profit ($/a)
0
1.71E+07
2.31E+07
2.28E+07
-6.11E+07
-3.44E+02
-1.24E+07
-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,
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Group Conceptual Process Design Project
CPD_3296
Final report
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
$
Feedstock:
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Group Conceptual Process Design Project
CPD_3296
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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 feed Oxygen 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
Flow rate ton/a
Price $/ton
3.25E+01 4.98E+00 7.37E+00 4.70E+00
9.54E+05 1.46E+05 2.16E+05 1.38E+05
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.
2.CO2 purges
1. Wastewater
$
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
Apply unit
Ni/ Al2O3
Ni/ Al2O3
Syngas production
Reaction
name
Shape
Steam
reforming
4-hole
cylinder
Bulk density
1100 kg/m3
Auto thermal
reforming
4-hole
cylinder with
domed ends
1000 kg/m3
-21-
Co/ Al2O3
F-T
synthesis
F-T
synthesis
-
Pt/ Zeolite
Hydrocracking
0.27 g/cc
-
Hydrocracking
Zeolite
Group Conceptual Process Design Project
CPD_3296
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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:
n =t
NFV
∑ (1 + DCFROR)
n =1
n
=0
(3.2)
(t = the life of the project , in our case = 17)
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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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:
∆H r0,298 ( kJ / mol )
No.
Reaction stochimometry
4.1
CH 4 ( g ) + H 2 O ( g ) ! CO ( g ) + 3H 2 ( g )
4.2
4.3
4.4
4.5
CO ( g ) + H 2O ( g ) ! CO2 ( g ) + H 2 ( g )
CH 4 ( g ) + 0.5O2 ( g ) ! CO ( g ) + 2 H 2 ( g )
CH 4 ( g ) + CO2 ( g ) ! 2CO ( g ) + 2 H 2 ( g )
CH 4 ( g ) + 2O2 ( g ) ! CO2 ( g ) + 2 H 2O ( g )
206
-41
-36
247
-803
4.6
2CO ( g ) → CO2 ( g ) + C ( g )
-173
4.7
CO ( g ) + 0.5O2 ( g ) ! CO2 ( g )
-284
4.8
H 2 ( g ) + 0.5O2 ( g ) ! H 2 O ( g )
4.9
CH 4 ( g ) ! C + 2 H 2 ( g )
4
-242
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
5
Jacob A. Moulijn, Chemical process technology, 2001, p133
http://www.eng.auburn.edu/users/halljoh/ASPEN_Manuals/APLUS%20111%20User%20Guide.pdf
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0.8
H2
0.7
molar fraction
0.6
0.5
H2
CO
CO2
H2O
CH4
CH4~H2O
0.4
0.3
CO
0.2
0.1
CO2
0
500
1000
1500
2000
Temperature (K)
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.
H2
0.8
H2
CO
CO2
H2O
CH4
H2, 20bar
CO,20bar
CO2,20bar
H2O, 20bar
CH4, 20bar
molar fraction
0.7
0.6
0.5
CH4~H2O
0.4
0.3
CO
0.2
0.1
0
500
CO2 800
1100
1400
1700
2000
Temperature (K)
Figure 4.2 Effect of temperature and pressure on equilibrium gas composition in
steam reforming reaction with H2O/CH4=1 mole/mole.
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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
Mole fraction
Ratio of H2/CO
H2
31.3126
CO
CO2
CH4
O2
H2O
10.4257 0.4977 32.5876 24.6427 0.5219
3.00
Total
100
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.
H2
0.6
mole fraction
0.5
O2
H2
CO
H2O
CH4
CO2
0.4
CO
0.3
CH4
0.2
H2O
0.1
CO2
0
700
1000
1300
1600
1900
Temperature (K)
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
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pressure on the reaction, we select the pressure at 1bar and 20bar and the
comparison is shown on figure 4.4.
At 1bar
0.7
mole fraction
At 20bar
H2
0.6
O2
H2
0.5
CO
H2O
0.4
0.3
CH4
CH4
O2,20BAR
CO
H2,20BAR
CO,20BAR
0.2
CO2,20BAR
H2O
0.1
H2O,20BAR
CH4,20BAR
CO2
0
700
900
1100
1300
1500
1700
CO2
1900
Temperature (K)
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
Paraffin
nCO ( g ) + ( 2n + 1) H 2 ( g ) → Cn H 2 n + 2 ( g / l ) + nH 2 O ( g )
Olefins
nCO ( g ) + 2nH 2 ( g ) → Cn H 2 n ( g / l ) + nH 2 O ( g )
∆H r0,298 ( kJ / mol )
-165
-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:
m n = (1 − α)α ;
n −1
w n (1 − α) 2 n
=
α
n
α
(4.10)
Where the growth probability factor α is independent of n, and m n is the mole
fraction of a hydrocarbon with the chain length n. The range of α is dependent
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Mass ratio, wt(%)
on the reaction conditions 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.
100
90
80
70
60
50
40
30
20
10
0
0.0
C1
C2~C4
C5~C9
C10~C14
C15~C20
C21~C45
C46~C100
0.2
0.4
0.6
0.8
1.0
Chain growth probability, a
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 Temperature
Pressure
H2/CO
feed ratio
Type
(°C)
(MPa)
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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Table 4.5 Reactions in Hydrocracking Unit 7
Reaction
Stochiometry
∆H r0,298
( kJ / mol )
Alkanes
hydrocracking
Hydroisomerization
-44
-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:
343-350°C Gas oil
Gravity, °API
23.6
Nitrogen, wt%
1250
Sulfur, wt%
2.0
Aniline point,°C
85
Unicracker reactor
376
367
avg. temp. ,°C
Product objective:
PC naphtha
Turbine fuel
Yield, vol%
feedstock
8.9
19.6
C4
11.3
21.7
C5-60°C
45.7
87.0
60°C+ Naphtha
54.1
-149°C+ Distillate
120.0
128.3
18.7
29.7
Total C4+
23.1
37.9
C6-C8
23.04
41.0
C6-C9
C6
360
Diesel
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
8
Julius Scherzer, A.J. Gruia, Hydrocracking Science and Technology, chapter6, p74, 1996
Julius Scherzer, A.J. Gruia, Hydrocracking Science and Technology, chapter11, p205, 1996
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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
-1
LHSV h
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
1
Unit
Syngas
production
Catalyst
Ni/Al2O3
Reactor
Multi-tube
Temperature (K)
1000
Pressure (Bar)
20-40
-
2
3
FT synthesis
Hydrocracking
Co/MgO/SiO2
Fixed bed
Slurry
Multi-fixed
bed
1473-1573
493-523
623-913
20-40
30
30-70
Pt/Y-Zeolite
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
Component
Formula
o
C
∆vapH
(tb)
Cp
KJ/mol
J/(mol K)
Hydrogen
H2
-252.87
0.9
Carbon monoxide
CO
-191.5
6.04
Carbon dioxide
CO2
Water
H2O
100
-29-
40.65
Phase
Gas
29.1
Gas
37.1
Gas
Liquid
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Oxygen
O2
-182.95
6.82
Gas
Methane
CH4
-161.48
8.19
35.7
Gas
Ethane
C2H6
78.29
38.56
52.5
Gas
Propane
C3H8
-42.1
19.04
73.6
Gas
Butane
C4H10
-1.9
22.9
140.9 Liquid
Pentane
C5H12
36.06
25.79
167.2 Liquid
Isopentane
C5H12
27.88
24.69
Liquid
Nonane
2,2,4,4-tetramethylpentane
C9H20
150.82
37.18
284.4 Liquid
C9H20
122.29
32.51
Liquid
Decane
C10H22
174.15
39.58
300.8 Liquid
Tetradecane
C14H30
253.58
48.16
Liquid
Pentadecane
C15H32
270.6
50.08
Liquid
Eicosane
C20H42
343
58.49
Solid
Heneicosane
C21H44
Solid
Nonacosane
C29H60
Solid
4.3 Models for vapor/liquid equilibrium
In Appendix 4.1, there are the steps for choosing a suitable property method in
ASPEN Plus. According to these guidelines, we can find the models for syngas
production unit, Fischer-Tropsch unit, Hydrocracking unit and Distillation column.
The details for choosing the property methods in the process are shown in
Appendix 4.2.
In table4.10 the property method of each unit is summarized and Peng-Robinson
method is the total simulation property we used in ASPEN simulation.
Table 4.10 summary of Property method
Unit
Unit Name
U100
CAR reactor
U200
FT reactor
U300
Hydrocracking
U400
Distillation
Property Method
PR-BM
PRMHV2
PENG-ROB
PRMHV2
The applied data we use from ASPEN simulation with the tolerance of 0.0001,
which can make sure of our design accuracy.
4.4 Reaction kinetics
The reaction kinetics can be described as three parts, which are CAR reactor, FT
reactor and hydrocracking reactor. Because each reactor uses typical catalyst
and different kinds of reactors, the kinetics is quite complex and is related to the
reactor design. The detail description of the kinetics can be seen in chapter 8,
8.2 equipment selection and design.
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5. Process Structure and Description
5.1 Criteria and Selections
The whole process is divided into four parts according to their functions, e.g.
syngas production unit, FT synthesis unit, hydrocracking unit, separation unit,
which will be described in details below. The criteria are used in different units
and the choice of the design criteria is explained also.
5.1.1 Syngas production unit
The main object of the syngas production unit is to convert the natural gas into
syngas. This unit is the beginning of the whole process, which supply enough
syngas for F-T synthesis unit with the ratio of H2 / CO equal to 2.0 and pure
hydrogen for hydrocracking unit. The feedstock of this unit is natural gas and
pure oxygen, and steam. All the raw materials enter the unit are vapor at room
temperature.
This unit can be divided as two parts, syngas generation and syngas purification.
The main process is shown in figure 5.1.
Pure Hydrogen
Syngas
750K
Pure Syngas
Carbon dioxide
Wastewater
Syngas
Generation
Hydrogen
Water
Separation Removal
Carbon dioxide
Removal
Figure5.1 The flowsheet of syngas production unit.
There are many methods to manufacture syngas, CAR reactor is chosen finally
because of the energy saving and suitable H2/CO ratio. The reason is descript in
Appendix 5.1. The operating pressure of CAR is 20-40 bar, we prefer the lower
one when considering the process safety. In a CAR reactor, there are primary
reformer (SMR) and secondary reformer (ATR) zones. About 24% of methane
react with steam in the primary zone with nickel catalyst at the condition of
P=20bar and T=1000K, and the remained methane is partial oxidized by pure
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oxygen in the secondary zone at the condition of P=20bar and higher
temperature T=1473K, the amount of oxygen is limited to avoid further
oxidization. The thermodynamic conditions are defined in chapter 4. The total
conversion of methane is 99.1% in CAR reactor.
Methane and steam are first mixed and heated to reaction temperature 623K and
transfer to the primary reforming zone; oxygen and recycled carbon dioxide is
heated to 523K and transfer to secondary reforming zone. The temperature of
oxygen is lower than the reaction temperature because autothermal reforming is
an exothermal reaction. The produced hot gas from the autothermal reforming
zone can supply heat for steam reforming zone, in this way the generated energy
is used and the whole reactor need no extra heating or furnace. In ASPEN we try
to simulate this kind of reactor by using two stoichiometric reactors, a heat
exchanger and a cooler. The graph of this kind of function is shown in figure 5.2.
The function of the heat exchanger is transferring the heat from the secondary
reformer to the primary reformer. CO2 is split from CO2 removal separator.
Natural gas
and steam
Syngas
Primary
reformer
(SMR)
O2, CO2
Secondary
reformer
(ATR)
Figure 5.2 The CAR reactor work in ASPEN
The products of the CAR reactor are hydrogen, carbon monoxide, carbon dioxide,
steam, unconverted methane and the temperature is as high as 750K. There are
three separation units applied to separate carbon dioxide, excess steam, and
hydrogen. Three equipments that are determined by their respective operating
condition requirement fulfill the separation of these components. And we choose
the temperature from high to low to avoid unnecessary heat loss. The operating
conditions of the separation units are listed in table5.1.
Table5.1 Operating conditions of separation units
Unit
Function
Operating condition
S101 H2 membrane separation
750K 35bar
E105 Steam removal
300K 20bar
300K 20bar
S102 CO2 removal
In CAR reactor the H2/CO ratio of this raw syngas product is controlled to be a
little higher than 2.0. The excess H2 can be compressed to hydrocracking unit,
so there is a H2 membrane separator needed after syngas production. (The detail
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can see Appendix 5.2 syngas ratio adjustment) And only part of the raw syngas
product will be deal in the separator, because of the high efficiency of hydrogen
separation, this route is more flexible and has already shown in chapter 2
process option.
After the cooler E104, the temperature of the syngas product is decreased to
750K, which is the best condition for hydrogen separation operation. Composite
Pt based membrane supported on stainless steel separator is used. The
advantage of this kind of membrane separator is obviously; such as allow high
temperature, sustainable, low investment and operation costing. The purity of
the separated hydrogen is 98.7%.
After the cooler E105, the syngas is further cooled to 300K, now the steam is
condensed to water and separated from raw syngas product.
300K is also the operating temperature of the last purification to remove carbon
dioxide, which dilute the reactant of FTS. Chemical and physical Absorption with
MDEA solvent is employed because of its low operating cost, high efficiency and
reliability. Carbon dioxide is absorbed in aMDEA solution, and pure syngas leave
the column. A low-pressure flash column regenerates the laden solution, and the
separated carbon dioxide will be recycle to CAR reactor and the purged carbon
dioxide amount is very small and pure, which can be sold to the food industry. In
Appendix 5.3 Carbon dioxide removal is described in detail.
5.1.2 FT synthesis unit
Fischer-Tropsch unit is the core of this design, in which syngas can be converted
into light and heavy hydrocarbons. The produced hydrocarbons will be sent to
the following part, hydrocracking unit, in order to have the desired products. As
to obtain a higher conversion rate of syngas, two stages of slurry reactors are
applied in FT unit. And this unit can be divided into three parts, which are
hydrocarbon production, slurry / catalyst recovery and hydrocarbon separation.
The simplified scheme is shown in Figure 5.3.
Syngas is feed with the H2/CO ratio of 2.0, which is near to the H2/CO ratio
basing on stoichiometry of Fischer-Tropsch reactions. Temperature and pressure
are very important in Fischer-Tropsch reactions, so the input syngas is firstly
compressed (K201) and heated (E201) to the desired process conditions.
For Fischer-Tropsch synthesis process, several types of reactors can be applied,
but according to the highly exothermal reactions involved and the preference of
producing more heavy wax in this process, bobble column slurry reactor can
satisfy the requirements. The comparison and selection of reactors check
Appendix 5.4. According to research of slurry reactors with Cobalt catalyst, the
conversion rate of syngas can reach 80% (one-through); the remaining 20%
would be discarded as fuel gas and economically not feasible. As a result, it is
necessary to raise the conversion rate of syngas; in this process, two-stage
reactors are applied. The unconverted syngas from the first stage reactor can
feed to the second stage slurry reactor; the total syngas conversion rate can
reach 96%.
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333K
30 bar
Fuel gas
333K
30 bar
Wastewater
Wastewater
523K
523K
30 bar
373K
30 bar
30 bar
Hydrocarbon
Slurry
Syngas
Heavy HC
Hydrocarbon
Solvent
Recycle Solvent
1st stage
reactor
Extractor
Wastewater
2nd stage
reactor
Slurry
recovery
Solvent
recovery
Flash
Figure 5.3 Preliminary flowsheet of Fischer-Tropsch unit
The slurry reactors operate at T=523 K and P=30 bars, and the cooling coils
inside the reactors remove the large amount heat generated by FTS reactions.
There are two streams leaving slurry reactor, one is vapor and the other is liquid
stream. The vapor stream contains unconverted syngas, light hydrocarbons and
steam, while the liquid stream mainly contains heavy hydrocarbons (wax) and
slurry, so a simple three-phase single flash (S203) should be applied to the
vapor stream. After cooled by E202 and flash separation, most of the water is
decanted and the unconverted syngas is sending to the second stage reactor,
another oil stream contains mostly light hydrocarbon can be gotten from the
flash too.
The liquid phase leaving slurry reactors contains wax and slurry, because the
catalyst is dispersed in the slurry, the recycle of slurry should also be taken into
account. Hydrocarbon can be extracted from slurry and following by solvent
recycle, at the end of this recycle part, clean wax can be obtained. And the slurry
can be settled and separated from wastewater, and then slurry that containing
catalyst will be recycled back to reactors.
In Aspen simulation, we combined a stoichiometry reactor and a single flash in
order to simulate a slurry reactor that has two outlet streams (vapor and liquid),
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these two units should have the same process condition (temperature &
pressure), and the following figure illuminates this simulation:
Syngas
523K, 30bar
Vapor
Rstoic.
Flash
Liquid
523K, 30bar
Figure 5.4 Simulation of slurry reactor in Aspen
Considering the sizing of reactors, if the feed flow rate is too high, there is no
doubt to have reactors with very large volume. So the first-stage includes two
parallel reactors that have the same size. Because each reactor has the same
conversion rate of 80%, which means the volume of the second stage reactor
should be much smaller than the first stage.
Before entering the second stage reactor, the unconverted syngas should be
preheated (E205) to the suitable temperature in order to achieve a better
conversion. And the reactions as well as operations in second stage slurry reactor
are the same as the first stage. Details see the explanation above. After coming
out of the second stage Fischer-Tropsch reactor, the vapor is sent to a simple
flash separator (S207), while the heavy hydrocarbon is also sent to the same
flash, in order to separate C7+, which will be sent to hydrocracking unit.
The operating conditions including temperature and pressure of equipments
mentioned above are shown in Table 5.2.
Table 5.2 Equipment operating
Equipment Type
R210
Slurry reactor
R220
Slurry reactor
R230
Slurry reactor
S203
3-phase flash
S204
3-phase flash
S206
Single flash
S207
Single flash
condition summary
Temperature
Pressure
523 K
30 bar
523 K
30 bar
523 K
30 bar
333 K
30 bar
333 K
30 bar
368 K
9 bar
373 K
30 bar
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Catalyst
Co/MgO/SiO2
Co/MgO/SiO2
Co/MgO/SiO2
N.A.
N.A.
N.A.
N.A.
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5.1.3 Hydrocracking unit
The F-T products have a carbon number ranging from C1 to around C77. In order
to get our target produces, which are Kerosene and Diesel, we have to reduce
the carbon number of heavy F-T wax to middle distillate range, e.g. C10 –C20.
On the other hand, some other reactions are happened in this unit, such as
hydrogenation of olefins; removal of the small amounts of oxygenates, mainly
primary alcohols; hydrocracking of the n-paraffins to isoparaffins of the desired
length/boiling range. Therefore, a hydrocracking unit is necessary to increase
both the production and quality of aimed products.
For hydrocracking section, we use Single-stage, single-catalyst, recycle
hydrocracking (figure 5.5). Because comparing to other hydrocracking process,
there are several advantages. Firstly, many of the units designed to maximize
diesel product utilize this configuration; Secondly, the input for hydrocracking in
our case contain low Sulphur and low Nitrogen, which fit for this configuration;
Finally, this is the simplest and cheapest one.
Makeup
H2
Fresh
feed
RecycleH2
Product gas
Light Naphtha
Heavy
Naphtha
Kerosene
Diesel
Singlestage
Product
Fuel Oil
FCC feed
Ethylene
feed
Lube Oil
base
Recycle Bottoms
Single-stage
reactor
Separators
Fractionator
Figure 5.5 Simplified flow diagram of single-stage, single-catalyst hydrocracking
process9
The overall reaction in hydrocracking unit is exothermic. The feedstock of this
unit is mainly the F-T product, ranging from C7 to C77. Some light fuel gas has
been separated before F-T effluent goes to hydrocracking unit, in order to reduce
the hydrocracking reactor size. Excess hydrogen feed into reactor. Hydrogen can
9
Julius Scherzer, A.J.Gruia, Hydrocracking science and Technology, chapter 10, p176
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Final report
be recycled back as a quench feed; the route can reduce the reactor temperature
generated by exothermal reaction.
The operating temperature in hydrocracking reactor is very important, because
target product can shift from naphtha to diesel with a decrease of reactor
temperature.10 The total reaction is exothermic, so low temperature is favorable.
On the other hand, in order to maintain 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
other hand, the rate of cracking and isomerization reactions decreases with
hydrogen partial pressure increase.
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 5.3.
Table5.3 Typical process and operating condition for mild hycrocracking11
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 condition of hydrocracking unit is determined by this means,
which are P=40bar and T=350°C.
5.1.4 Separation unit
Distillation column is used to separate the products and by-products, the detail is
in chapter 8, 8.2.4.
10
11
Julius Scherzer, A.J.Gruia, Hydrocracking science and Technology, chapter 14, page 244,1996
Julius Scherzer, A.J.Gruia, Hydrocracking science and Technology, chapter 12, page 216,1996
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5.2 Process Flow Scheme (PFS)
The process flow scheme (PFS) is composed as U100, U200, U300 and U400,
and combined together. The total process flow scheme is shown in Appendix
5.5.
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Final report
5.3 Process Stream Summary
The total streams specification in ASPEN is summarized in Appendix5.6.
According to it, we can do the overall component mass balance and stream heat
balance.
Table 5.4 the overall component mass balance & stream heat balance
Overall Component Mass Balance & Stream Heat Balance
STREAM Nr.
NAME
:
:
101
IN
Natural gas
kg/s
Total
kmol/s
22.95
Pressur
Temp.
Bara
K
Enthalpy kW
STREAM Nr.
:
NAME
:
Enthalpy kW
STREAM Nr.
NAME
:
:
kmol/s
kg/s
kmol/s
0.932 15.99
401
IN
Steam
kg/s
kmol/s
0.499 0.301
20
298
2
500
-105789.62
-213802.7612
-98.85
-3929.72
125
OUT
Carbon Dioxide
kmol/s
414
OUT
Fuel gas
kg/s
0.012 4.597
kmol/s
0.203
413
kg/s
-17145.86
-522291.16
404
OUT
Naphata
406
OUT
Kerosene
Enthalpy kW
-10985.54
kg/s
kmol/s
-323620.95
1.855
-4691.47
kmol/s
407
OUT
Diesel
kg/s
kmol/s
410
OUT
Wax purge
kg/s
kmol/s
Total out
kg/s
0.042 5.952 0.033 6.584 0.028 0.186 0.001 56.024
5
1.2
1.5
2
273.15
486.6
563.6
715
-9455.05
-8542.11
-132.58
kmol/s
2.174
-573243.79
0.001
OUT-IN:
2.794
kmol/s
33.6
5
361
4.58
kmol/s
OUT
5
354.3
kg/s
kg/s
Waste water
20
300.7
Total
Pressure Bara
Temp.
K
Total IN
0.017 56.025
40
683
0.524
Pressure Bara
Temp.
K
kg/s
1.346 16.79
103
IN
Oxygen
20
298
kg/s
Total
102
IN
Steam
-0.62
-249622.84
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5.4 Utilities
5.4.1 utility introduction
The utility applied in the system is hot steam, cooling water and electricity, which
are also described in Appendix 5.7. Hot steam is applied to supply the heat
needed in the system; because the whole system temperature is so high that
only the high pressure (HP) and middle pressure (MP) steam is used. Now the
steam is bought from suppliers, we may generate it within the factory also, from
where many fuel gas produced as by-product, but it is outside the battery limit of
this design. The location of the factory is in the remote area, the availability of
the utility is very important. Cooling water, steam and electricity are the most
normal, cheap utility we can find, so they are applied here.
For the operating temperatures of different vessels are dramatically different,
cooling water is used as coolant to remove the heat energy. Electricity is
necessary for modern factory; it is the driving force of the most equipments. The
summary of the utility is shown in Appendix 5.9.
5.4.2 Pinch and heat exchanger network
There are 12 heat exchangers in the ASPEN simulation, including heaters and
coolers. To recovery the energy as much as possible, we are planning to make a
heat exchanger network analysis.
First we calculate the pinch point of the all steams, and then arrange the hot
streams and the cold streams to exchange heat. To balance the heat in the
process and use of energy efficiently, some streams will be split, which means
more than 12 heat exchangers will be used in practice. The whole calculation
procedure is shown in Appendix 5.8. Actually 26 heat exchangers should be
applied in plant design. The arrangement of the heat exchanger is also drawn in
Appendix 5.8.
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5.5 Process Yields
To show the performance of the design, process yield is listed in table 5.5, which
represents the consumption per ton of main product (Naphtha, Kerosene, and
Diesel).
Table 5.5. Process yields summary
Name
Feed
Natural gas
Oxygen
Steam
Products Naphata
Kerosene
Disel
Total
By products LPG
Fuel gas
Wax
Wastes
Carbon dioxide
Waste Water
Total
Process Stream
Ref.
kg/s
t/h
t/t products
Stream IN
OUT IN
OUT
IN
OUT
101 22.95
82.63
1.34
103 15.99
57.56
0.93
102+401 17.09
61.53
1.00
404
4.58
16.49
0.27
406
5.95
21.43
0.35
407
6.58
23.70
0.38
17.12
61.62
1.00
0.00
0.00
0.00
414
4.60
16.55
0.27
410
0.19
0.67
0.01
125
0.52
1.89
0.03
413
33.60
120.96
1.96
56.03 56.02 201.72 201.69 3.27 3.27
Utilities
Name
MP Steam
HP Steam
C.W.
Electricity
Ref.
Stream kg/s
kW
t/h
kWh/h t/t
0.346
1.2456
61.378
220.96
63.222
227.6
6919
6919.3
The block scheme is shown in figure 5.6.
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kWh/t
products products
0.02022
3.58599
3.69375
112.2936
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CPD_3296
Final report
CO2 1.89t/h
Natural
Gas
Oxygen
Syngas
production
1500K 20Bar
FT
synthesis
1500K 20Bar
Stream
Hydrocracking
1500K
20Bar
Fuel Gas 16.55 t/h
(0.27)
Products 61.62t/h
Distillation
1500K
20Bar
Wax 0.67t/h
(0.01)
Wastewater
MP Steam
HP Steam
1.25 t/h
(0.02 t/t
products)
220.96 t/h
(3.59 t/t
products)
Products
Total Process
Feed
201.72 t/h
(3.27 t/t
products)
By-products
6919.3 t/h
(112 t/t
products)
227.6 t/h
(3.70 t/t
products)
Cooling
Water
Electricity
Figure 5.6 Block Scheme
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61.62 t/h
(1.00 t/t
products)
0.67 t/h
(0.01 t/t
products)
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6. Process control
In this Chapter, the control of the whole process will be described unit by unit.
Only the basic control is given in this CPD design project and is not going to be
simulated in any program. All the controllers mentioned below are shown in
Appendix 6. The purpose of this chapter is to clarify why certain combinations
of control loops were chosen and positioned from a process point of view and
how these choices have influenced the design process.
6.1 Syngas production unit (U100)
Feed Streams (101, 102 and 103): There are three raw materials attending
the reaction of syngas production, which are natural gas, steam and oxygen, and
the flow rate of these feedstock affect the final products quantity and quality
directly, so flow controller should be added to make sure the flow rate in a
correct value.
To keep the reactants amount in stoichiometry and to make sure the correct
ratio feeding to the reactor, we use two ratio controllers. One is to control the
flow rate ratio of natural gas (101) and steam (102); the other is to control the
flow rate ratio of natural gas (101) and oxygen (103).
Heaters (E101 and E102): Two heaters E101 and E102 before the CAR reactor
R100 are used to heat up the reactants to the reaction temperature. We control
these two temperatures by measuring the temperatures of the outlet streams
(106 and 108). If there is any fluctuation in the feed temperature, the valve will
control the flow rate of the hot steam, and then the feed temperatures in syngas
production reactor will be kept as constant.
CAR Reactor (R100): It is very important to keep the reactor in right operating
conditions. Therefore pressure controller is applied to execute this function. The
ratio controller of the feedstock between natural gas and steam is the set point.
If the pressure is higher than 20 bar, the valve will be closed and make the set
point of natural gas flow rate decrease. At the same time the flow rate of the
steam and oxygen will decrease with the certain ratio, the input decrease will
lower the reactor pressure. Also, a temperature controller is employed to ensure
the reactor in good temperature range, which decides the set point of the heat
exchanger E101.
Coolers (E104 and E105): These two heat exchangers are placed, in order to
cool down the raw syngas to the operating temperature of hydrogen membrane
separator and carbon dioxide removal separator respectively. The separators
temperature should be constant, but the change in a small range is inevitable,
these two temperature controllers can ensure that syngas enters the separator at
suitable temperature by adjusting the cooling water flow rate.
Membrane Separator (S101): For only part of syngas entering the hydrogen
membrane separator, the fraction of the by pass flow rate/ whole syngas flow
rate is dependent on the amount of the purified hydrogen. In order to control the
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Final report
hydrogen flow rate at the set point, a three-way valve and a flow controller are
used. The set point is given by hydrocracking unit, which means the hydrogen
amount is decided by requirements from hydrocracking unit operation.
Compressor (K101): The purified hydrogen pressure should be controlled in 40
bars; otherwise the hydrocracking reactor pressure will be influenced. A pressure
controller is applied to adjust the inlet flow rate of compressor, which can ensure
the outlet pressure at good posistion.
Controller
type
Flow
Ratio
Ratio
Temperature
Temperature
Pressure
Temperature
Temperature
Flow rate
Pressure
Temperature
Flow
Pressure
Temperature
Sensor
location
101
101
101
105
108
R100
R100
113
116
118
122
115
117
121
Process
variable
Flow rate
Flow rate
Flow rate
Temperature
Temperature
Pressure
Temperature
Temperature
Flow rate
Pressure
Temperature
Flow rate
Pressure
Temperature
Manipulate
variable
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Temperature
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Manipulate
location
101
102
103
E101
E102
111
106
E104
K101
E105
112
K101
E105
6.2 Fischer-Tropsch synthesis unit (U200)
Compressor (K201): The function of the compressor is to increase the pressure
of the syngas to ensure that FT reactions carry out in the setting conditions. The
pressure control is also a necessary tool to fulfill this purpose, which can shun
the influence of changing the syngas inlet flow rate.
Heat exchanger (E201, E205): Temperature is important to the FT synthesis,
the temperature controller measures the temperature of the inlet of the FT
reactors while comparing to the set point, then adjusts the hot stream flow rate
of the heat exchanger, so the FT temperature will not deviate the required value.
FT reactor (R210, R220, R230): There are three slurry reactors in this unit,
and the process control applied here is similar to each other. Since the FischerTropsch reactions are highly exothermal, the generated heat will be removed by
coolant in the cooling coils, and the coolant flow rate is dependent on the heat
removal rate; the flow rate should be controlled with a set point, which is
determined by the real-time reactor temperature.
Changing the vapor outlet flow rate performs the pressure control of the
reactors. The outlet flow rate of the vapor phase should always be surveyed and
compared to the set point, while the reactor pressure designs the set point.
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Furthermore lever controllers are used here in order to avoid flooding, by
changing the flow rate of the liquid phase outlet.
With these three types of controllers, the Fischer-Tropsch reactions can be
processed in the desired condition.
Heat Exchanger (E202, E203): The heat exchangers are placed before the
flash units; the temperature of the reactor vapor outlet is firstly cooled down to
the desired value. This can be achieved by setting the coolant flow rate while
measuring the outlet temperature of the heat exchangers, compared to the set
point.
Flash (S203, S204, S207): The flash pressure can affect the separation result
greatly. While determining the pressure drops in the flash units and comparing to
the set point, the vapor flow rate can be adjusted by the valves, which are
controlled by the pressure controllers.
Extractor: The solvent flow rate should be in a certain ratio to the flow rate of
heavy hydrocarbon. By determining the flow rate of the inlet heavy HC, a ratio
controller can be set and control the flow rate of solvent, hence the desired
product can be separated in this extractor in the optimal conditions.
Settler: In settler, slurry and wastewater is settled and separated. After
separation slurry will be recycled to slurry reactors and wastewater will be
purged out of the system. A level controller should be added to settler in order to
avoid the split out of liquid. While the liquid level inside settler extends the
acceptable limits, level controller can open the valve of the purge flow to justify
the conditions inside.
Pump (P201): The pump is served to increase the pressure drop of the heavy
hydrocarbon, which will be sent to the following unit. The outlet pressure of the
pump is detected while comparing to the set point, which can adjust the valve on
the bypass of this pump, in order to obtain the satisfactory pressure.
Heat exchanger (E204): In the wastewater treating part of the FischerTropsch unit, the stream should be heated before going into the flash that can
separate wastewater and fuel gas. The temperature controller measures the
temperature of the inlet of the flash while comparing to the set point, and then it
can adjust the hot stream flow rate of the heat exchanger.
Flash (S206): A pressure controller is employed here to ensure the separation
is performed under an optimized pressure, so the residue hydrocarbon and other
gases can be separated completely from wastewater. To keep the flash in the
desired pressure, by determining the flash vapor phase flow rate and comparing
it to the set point, the valve is adjusted and hence pressure drop in the flash is
controlled.
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Controller
type
Pressure
Temperature
Temperature
Temperature
Temperature
Temperature
Pressure
Pressure
Pressure
Level
Level
Level
Level
Temperature
Temperature
Pressure
Pressure
Pressure
Ratio
Pressure
Temperature
Pressure
Sensor
location
201
202
216
R210
R220
R230
R210
R220
R230
R210
R220
R230
Settler
212
220
S203
S204
S207
234
226
227
S206
Process
variable
Pressure
Temperature
Temperature
Temperature
Temperature
Temperature
Pressure
Pressure
Pressure
Level
Level
Level
Level
Temperature
Temperature
Pressure
Pressure
Pressure
Flow rate
Pressure
Temperature
Pressure
Manipulate
variable
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Manipulate
location
K201
E201
E205
Water
Water
Water
207
209
223
208
210
222
Water Purge
E202
E203
213
221
232
Solvent
P201
E204
228
6.3 Hydrocracking unit (U300)
Hydrocracking reactor (R301): To keep the hydrocracking at the desired
temperature, cooled hydrogen is used to quench the column, temperature sensor
is added in every reaction zone, and sends the information to the temperature
controller, which will adjust the flow rate of the cooled hydrogen. In this way, the
temperature of different parts of reactor will be kept in similar level. The ratio of
the two reactants of hydrocracking, hydrogen (302) and wax (303) are kept in
the design value by the flow rate ratio controller. Pressure of the reactor is
controlled by the flow rate of the outlet (305).
Flash (S301): The pressure and temperature is key element of the separation
efficiency of the flash and should be controlled in the favorite value. In this
column, the flow rate of 306 and 308 is adjusted to keep the column pressure
and temperature constant. The vapor from flash 301 is hydrogen, which will
partly be recycled to hydrocracking and the rest will be purged. The amount of
purge gas is depended on the flow rate of the total hydrogen to ensure the
recycle hydrogen amount is in the design value. So the flow rate of 306 can
affect the flow rate of 307 with a ratio controller.
Heat exchanger (E301): Obviously, the inlet temperature will affect the
separating efficiency of distillation column. So a temperature controller to ensure
the correct cooling degree of heat exchanger E301 is necessary, which will adjust
the cooling water flow rate to avoid the influence of the change of temperature.
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Heat exchanger (E302): The hydrocarbon feed to hydrocracking reactor should
satisfy its temperature and pressure requirements, so heater E302 is added in
order to preheat the hydrocarbon stream entering HC unit. The temperature
sensor detects temperature in stream 301. After comparison to the desired
temperature, the flow rate of heating steams will be adjusted.
Controller
type
Ratio
Temperature
Pressure
Pressure
Temperature
Ratio
Flow
Temperature
Temperature
Sensor
location
303
R310
R310
S301
S301
S301
307
311
310
Process
variable
Flow rate
Temperature
Pressure
Pressure
Temperature
Flow rate
Flow rate
Temperature
Temperature
Manipulate
variable
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Manipulate
location
302
302(part)
305
306
308
306
307
E301
E302
6.4 Separation unit (U400)
Distillation column (D401): It is the very important to operate the distillation
in the appropriate temperature, it is influenced by the flow rate of the steam
(401), which heats up the whole column. So a flow rate controller of the steam
inlet is necessary. The pressure is affected by the reflux rate of the condenser,
which is used here to keep the pressure in desired value. To avoid the flooding of
the column, a lever controller is applied, which adjusts the flow out rate of the
liquid fraction (408) to keep correct amount of the wax inside the column.
Flash (S401): Analogously, in the single flash distillation column, pressure
control is also necessary. There is one flow rate controller, which controls the
outlet stream in this flash to keep the pressure constant.
Pump (P401): the wax stream after the distillation column will partly be
recycled to the hydrocracking again, a pressure controller is added here to
ensure the input for the hydrocarcking in desired value.
Coolers (E401): A heat exchanger is bestowed in the process to cool down
recycled wax. The cooling water flow rate is modulated according to the
temperature of the exchanger outlet, so the temperature or flow rate of the
recycled fractions of wax will not change the final product inlet temperature too
much. Thus, ensure the hydrocracking reactor safe enough.
Controller
type
Temperature
Pressure
Level
Pressure
Pressure
Temperature
Sensor
location
D401
D401
D401
S401
412
411
Process
variable
Temperature
Pressure
Temperature
Pressure
Pressure
Temperature
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Manipulate
variable
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Flow rate
Manipulate
location
401
402
408
P401
E401
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7. Mass and heat balance
Here the mass and heat balance will be performed unit by unit, and all the
calculation procedure and results are available in Appendix7. The following data
in the table just give a brief description about that. Data in the last column
indicates the difference between enthalpies of IN- and OUT-going steams, which
is almost equal to the difference between heats IN and OUT.
Table7.1
The heat balance
Stream
U100
Equipment
Stream
U200
Equipment
Stream
U300
Equipment
Stream
U400
Equipment
Stream
Process
Equipment
IN
-3.20E+05
-7.74E+04
-1.51E+05
-2.42E+05
-3.86E+04
1.26E+04
-5.63E+05
-4.99E+03
-1.07E+06
-3.11E+05
OUT
-3.38E+05
-5.92E+04
-3.86E+05
-6.52E+03
-2.60E+04
0.00E+00
-5.72E+05
0.00E+00
-1.32E+06
-6.57E+04
OUT-IN
-1.83E+04
1.83E+04
-2.35E+05
2.35E+05
1.26E+04
-1.26E+04
-8.95E+03
4.99E+03
-2.50E+05
2.46E+05
From the tables in Appendix 7, it can be seen that there are still some deficits
in mass and heat balance. For mass balance, we have difference of 0.001 kg/s
between inlet and outlet flow rate. However, it might be due to calculation
accuracy and could be omitted, when compared to the amount of feed flow rate
(56.025 kg/s). Form heat balance, the difference is –0.04e5kW, which is caused
by the work and heat transfer, but it is still in the acceptable range, compare to
–2.50e5 kW heat feed.
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8. Process and Equipment Design
8.1 Integration by process simulation
Tools used in our process are mainly ASPEN and Matlab. ASPEN does the whole
process simulation; Matlab is used to model the reaction rate of Syngas
production unit. The m-file can be found in Appendix8.1. ASPEN is used as an
important design tool, the description of the ASPEN simulation is in Appendix
8.2 and the process flow sheet is shown in Appendix8.3.
The major problems encountered in using ASPEN is described as follows:
•
•
•
•
In Syngas production unit, the chosen reactor is a combined autothermal
reforming (CAR) reactor. However, there is no such reactor model in
ASPEN at all. Therefore, two stoichiometric reactors are used together with
one heat exchanger, which has been applied to simulate the CAR reactor.
In FTS unit, the chosen reactor is the slurry reactor, where wax and light
hydrocarbons are separated automatically. In ASPEN, we chose a
stoichiometric reactor combined with a single flash to simulate this slurry
reactor.
In hydrocracking unit, the multi-fixed bed reactor has been chosen. Due to
too many reactions happened here, it is difficult to exactly simulate all
reactions. Since we know the predicted product distribution, a Ryield
reactor model has been applied here in ASPEN.
In separation unit, we use a model called Petrofrac to simulate and size
the real distillation column. However, an SCfrac model has been applied
first in ASPEN to obtain the necessary data. Then, the corresponding data
is filled in PetroFrac model to achieve the final design for this separation
unit.
The results simulated by ASPEN, namely the stream summary can be found in
Appendix5.6.
8.2 Equipment selection and design
All the equipments in this system are made of carbon steel, because there are no
strongly corrosive materials involved in the process.
8.2.1 syngas reactor design
The reactor design is based on the reaction kinetics. The reaction kinetics and
the reaction rate calculation is described in Appendix8.4; the result is listed in
the below table:
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Table 8.1 Producing rate of syngas and the required catalyst amount
Reaction rate [mol/kg catatlyst/s]
SMR
ATR
rcb
rcr
rsr
rgw
rH2
rCO
0
100.8823
20.3323
31.3567
3.22E+03
202.2267
-92.4656
9.66E+03
5.77E+02
3.22E+03
3.57E+02
Productivity [mol/s]
FH2
SMR
ATR
9.33E+02
1.89E+03
FCO
3.11E+02
9.50E+02
Weight [kg]
9.67E-02
3.28E+00
Catalyst
Density [kg/m3]
1970
Volume [m3]
4.91E-05
1.66E-03
An autothermal reforming reactor with Ni catalysts supported is at the contact
resident time of normally 0.1 second 12 [P.M. en and X. Chu, 1994]. A typical
steam reformer operates at 850~900°C with a Ni/Al2O3 catalyst and the
superficial contact time is 0.5~1.5 second 13 [S.S. Bharadwaj and Schmidt,
1994]. Because the residence time of autothermal reaction is very small, only
accounting for one tenth of steam reforming reaction, we neglect the contact
time of autothermal reaction, and estimate the volume of autothermal zone is in
one tenth of the steam-reforming zone.
According to Catalyst handbook [1996], the overall length of reformer tubes is
usually in the range 7.5~12.0m; tube diameter usually lies between 7 and
13cm.14 we have chosen the length of 7.5m with 8cm diameter. In table 8.2 the
parameters of SMR and ATR are given; the volume of CAR reactor and the tube
numbers can be calculated. The result shows that we need 553 tubes and the
reactor volume is 21m3.
Table 8.2 CAR reactor volume and tube number
Volume Flow [m3/s]
SMR
Resident time
[s]
Volume
[m3]
ATR
10% of SMR [m3]
Total Volume
[m3]
18.954
1.000
18.954
1.895
Tube parameter
H
[m]
D
[m]
Volume/tuber [m3]
20.850 Tube number
7.5
0.08
0.0377
553
8.2.2. Reactor design of Fischer-Tropsch process
The reactor design of FTs process is written in Appendix 8.5 and results are
listed in the flowing table. The design profiles for the1st stage reactor is shown as
following:
12
13
14
P.M. Torniainen, X. Chu, and L.D. Schmidt, Journal of Catalysis, 1994, 146, p1-10
S.S. Bharadwaj and L.D. Schmidt, Journal of Catalysis, 1994, 146, p11-21
Twigg M.V. 1989, Catalyst Handbook, Second edition, Wolfe Publishing Ltd. London 1996, p265
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Table 8.3 Design profiles for 1st stage reactor
Reactor Parameters
Value
Temperature K
523
Pressure bar
30
Diameter DT [m]
2.7
Reactor Height H [m]
20.1
Un-gassed Slurry Height [m]
13.4
Height of suspension [m]
16.1
Diameter of cooling tubes [m]
50E-03
Height of vertical cooling tubes [m] 16.1
Temperature of coolant [K]
293
Superficial velocity [m/s]
0.4
Total gas holdup ε [-]
0.167
0.35
Slurry concentration ε S [-]
Catalyst amount [kg/reactor]
1.74E+04
The design profiles for the 2nd stage reactor are shown in Table 8.4.
Table 8.4 Design profiles for 2nd stage reactor
Reactor Parameters
Value
Temperature K
523
Pressure bar
30
Diameter DT [m]
1.7
Reactor Height H [m]
20.6
Un-gassed Slurry Height [m]
13.9
Height of suspension [m]
16.5
Diameter of cooling tubes [m]
50E-03
Height of vertical cooling tubes [m] 16.5
Temperature of coolant [K]
293
Superficial velocity [m/s]
0.3
Total gas holdup ε [-]
0.155
0.35
Slurry concentration ε S [-]
Catalyst amount [kg/reactor]
7.16E+03
Regarding the detailed calculation sheet, please see FT reactor calculation.xls.
8.2.3. Hydrocracking design
The reactor used in hydrocracking unit is fownflow, fixed-bed catalytic reactor.
The detailed description and calculation are shown in Appendix 8.6. The
summary of hydrocraker design data is shown in table 8.5.
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Table 8.5 Hydrocracker design results summary
Design specification
Symbol
Option1
100
000
Catalyst amount [kg]
Wcat
Cat. Density [kg/m3]
Option2
100 000
ρbed
500
500
Vcat
D
200
200
4
4
Lcat
N
L per cat bed
16
16
4
4
4
4
Resident time [hr]
Volume of hydrocracker [m3]
τ
0.6
1566.0
0.6
1047.6
Hydrocracker Diameter [m]
Dreactor
H
6
6
55
37
3
Cat. Bed volume [m ]
Hydrocraker inner diameter
[m]
Length of total cat. Bed [m]
Number of cat. Bed
Length per cat. Bed [m]
Height of Hydrocracker [m]
Vreactor
There are two options in hydrockacking reactor design, the detail description and
calculation are shown in Appendix 8.6, and the results are in table 8.6
Table 8.6 Summary of hydrocracker volume calculation
Option 1
Option 2
3
0.725
0.485
m / s
φ
v
V


 m3 
1566.0
Cost [Million $] 0.152
1047.6
0.129
From Table 8.6, we can see that option2, which is our innovation design, cost
less money. But all the other design is based on option1. The saved money in
hydrocracking unit does not take into account of the final cost estimation.
8.2.4 Separation unit design
Separation is the last one of four unit operations within plant design. The output
from Hydrocracking unit operation will be fed to this area, in order to separate
LPG, Naphtha, Kerosene, Diesel and Wax, respectively.
a). Boundaries
Separation unit operation consists of a distillation column, a single flash
separator, and several pumps, compressors and heat exchangers. The boundary
is from Hydrocracking effluent to product output.
b). Performance specifications for distillation
During the separation simulation in Aspen plus, the following performance
specifications are set:
- Top product mass flow: 1.25 kg/s;
- 5% ASTM D86 of Kerosene: 185 °C;
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95% ASTM D86 of Kerosene: 290 °C;
5% ASTM D86 of Diesel: 240 °C;
95% ASTM D86 of Diesel: 350 °C;
Mole purity of wax: 0.96.
c). Design variables
During this distillation column design, the following variables will be valuated:
- Number of theoretical stages;
- Number of actual trays;
- Tray efficiency;
- Feed tray;
- Draw off trays;
- Type of trays/packing;
- Tray diameter;
- Weir length;
- Active area;
- Number of bubble caps/valves/sieve holes
- Heating/cooling duties.
d). Configuration
A distillation column with several side streams is chosen as main operation unit
in this area, since we have to separate several products from byproducts. In
Aspen simulation, SCFrac column model is used first to simulate this complex
distillation process in whole system simulation by Aspen plus, named as
CPD_3296final12.07.bkp. However, SCFrac cannot offer all the data for design
variables, and therefore, Petrofrac column model (named as Petrofrac
Design.bkp) in Aspen plus is used, when doing the specific design for separation
unit operation. The procedure of the specific column design is shown in
Appendix 8.7. The results are listed below.
Figure 8.1 Column internals
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Table 8.7 Aspen calculates the size of each tray, and results are given below.
Section
Stage range
Dc [m]
Ad ratio
Vsd [m/s]
1
2~4
3.27
0.10
0.054
2
5~8
3.08
0.10
0.045
3
9~11
3.00
0.10
0.044
4
11~25
2.46
0.10
0.013
5
26~28
0.61
0.24
0.113
Stripper Stage range
Dc [m]
Ad ratio
Vsd [m/s]
1
1~4
0.82
0.19
0.119
2
1~4
0.93
0.19
0.115
Note:
- Dc: Column diameter, m;
- Ad ratio: Downcomer area / Column area;
- Vsd: Side downcomer velocity, m/s;
- Lsw: Side weir length, m.
Lsw [m]
2.38
2.24
2.18
1.78
0.55
Lsw [m]
0.71
0.80
650
600
550
500
450
400
350
300
250
200
1.7
1.5
1.3
1.1
0.9
T
0.7
P
Pressure, bar
Temperature,K
Temperature and Pressure profile in Distillation
Column
0.5
0
5
10
15
20
25
30
Number of stage
Figure 8.2 Temperatures and Pressure Profile in Distillation Column
8.2.5 Shell and tube exchangers design
The transfer of heat to and from process fluid is an essential part of the chemical
process. The most commonly used type of heat-transfer equipment is shell and
tube heat exchanger and we choose one shell two tube pass exchanger.
The general equation for heat transfer across a surface is:
Q = U ⋅ A ⋅ ∆Tm → A =
Q
U ⋅ ∆Tm
(8.1)
Q - Heat transferred per unit time, W
U - The overall heat transfer coefficient, W/m2 °C
A - Heat transfer area, m2
∆Tm - Logarithmic mean temperature difference, °C
Logarithmic mean temp difference:
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∆Tlm =
(T1 − t2 ) − (T2 − t1 )
(T − t )
ln 1 2
(T2 − t1 )
(8.2)
∆Tm = Ft ∆Tlm
(8.3)
∆Tlm - Logarithmic mean temp difference
T1 – inlet shell-side fluid temperature
T2 – outlet shell-side fluid temperature
t1 – inlet tube-side temperature
t2 – outlet tube-side temperature
The value of Ft depends upon the exact arrangement of the streams within the
exchangers, the number of exchangers in series, and two parameters defined in
terms of the terminal temperatures of the two streams:
R=
T1 − T2 range of shell fluid
=
t2 − t1
range of tube fluid
S=
t2 − t1
range of tube fluid
=
T1 − t1 max temperature difference
The overall heat-transfer coefficient can be estimated according to the typical
overall coefficient for shell and tube exchangers in Chemical Engineering Volume
6. Table 8.8 only gives the required coefficient in our design and we choose the
average value as our approximation.
Table 8.8. Typical overall coefficient15
Hot fluid
Steam
Steam
Steam
Gases
Heavy oils
Light oils
Shell and tube exchangers
Cold fluid
Heaters
Water
Gases
Heavy oils
Coolers
Water
Water
Water
o
U [W/m2 C]
2750
165
255
160
180
625
Table 8.9 is the summary of the coolers and heaters in our design and the detail
description is in equipment data specification.
15
Coulson & Richardson’s Chemical engineering ,volume 6,1993, p580-581
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Table 8.9 Coolers and heaters specification summary
Type
Unit
St No.
In
HeaterE101
HeaterE102
CoolerE105
HeaterE201
CoolerE202
CoolerE203
HeaterE204
HeaterE205
HeaterE301
HeaterE302
CoolerE401
105
108
120
201
211
218
226
213
309
233
409
Flow rate
Out
[kg/s]
106
109
122
202
212
220
227
216
311
310
411
39.735
25.907
65.004
43.042
35.453
9.139
22.628
10.541
20.982
16.998
3.528
T [K]
In
Out
dT
Q
[oC]
[kW]
440.6
623 70.7601 2.10E+04
297.2
523 191.166 5.70E+03
750
300 -104.01 -9.62E+04
357.6
523 162.835 1.94E+04
515 323.1 92.8931 -5.83E+04
523 323.1 46.2397 -1.15E+04
336.4 368.1 245.334 4.72E+03
323.1
523 179.361 5.10E+03
274.5 573.1 170.038 1.92E+04
373.1 623.1 98.2757 1.49E+04
715 623.1 -364.87 -1.13E+03
U
W/m2
o
C
A
m2
165 1.79E+03
165 1.81E+02
160 5.78E+03
165 7.22E+02
160-3.92E+03
160-1.55E+03
2750 6.99E+00
165 1.72E+02
255 4.42E+02
255 5.95E+02
180 1.72E+01
8.2.6 Single flash column design
A single flash distillation column is used as the separation of liquid droplets from
vapor or vapor and liquid mixing streams. The design scheme of vertical
separator is shown below:
Vapor
0.4m
DV =
Dv
4VV
πµ s
µ s = µ t , if with demister
µ t = 0.07[(ρL − ρV ) / ρV ]1/ 2
Dv/2
DV : minimum vessel diameter, m;
VV : vapor volumtric flowrate, m3 /s;
µ s : settling velocity, m/s;
Hl
ρL : liquid density, kg/m3 ;
ρV : vapor density,kg/m3 ;
Liquid
Figure 8.3 Design scheme for single flash distillation column
Firstly, settling velocity is calculated by the above formula. µ s can be equal to
the settling velocity, since there is a demister placed on the top of single flash
column. In terms of empirical correlation, the vapor height can be considered as
1.5 minimum vessel diameters. Then, residence time and liquid height are all
available.
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8.2.7 Pump and Compressor design
1 Compressors Design
The horsepower can be read from ASPEN. But if unknown, it can also be
calculated by the following equation 8.416.
 P γ 
 3.03 ⋅ 10−5 
out
hp = 
 − 1
 ⋅ Pin ⋅ Qin ⋅ 
γ
P
 in 



With: γ
P
Q
(8.4)
constant (0.23)
pressure (lbf/ft2)
feed (ft3/min)
-
-
K101 is installed to increase the pressure and decrease the temperature of
hydrogen, which will be fed into hydrocracking unit. K201 is installed between
Syngas production unit and F-T unit, in order to increase the feed Syngas to a
desired pressure.
2 Pumps
The power requirement if not calculated by Aspen, is calculated using equation
8.5.
Power =
∆P ⋅ Q p
With: ∆P
ηp
⋅ 100
Qp
Pressure difference (N/m2)
Flow rate (m3/s)
ηp
Pump efficiency (0.72)
(8.5)
P401 is installed between hydrocracking unit and separation unit, in order to
increase the pressure of recycled wax, which can be separated from the final
separation unit.
8.2.8 H2 membrane separation (S101)
The purpose of H2 membrane separator is to adjust the H2/CO ratio and provide
H2 for hydrocracking reaction. The design is based on the data from ASPEN and
literature.
Table 8.10 Membrane separator permeating data
H2 Permeation Flow
H2 Permeation Flux
H2 Permeation Area
cm3/min
cm3/cm2 min
m2
3.408E+07 (ASPEN)
19.000 (Literature)17
179.349
16
James M. Douglas, Conceptual Design of Chemical Processes. Page 490
M.Konno, M.Shindo, S.Sugawara and S.Saito; A composite palladium and porous aluminum
oxide membrane for hydrogen gas separation, Journal of membrane science, 37, 1988, p193-197.
17
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The diameter and the length of the tube are approximated by experience. Then
we can calculated the volume of one tube roughly and how many tubes we need
in the separator.
Table 8.11 Membrane separate sizing
Diameter of the tube
Hight of the vessel
Membrane area of one tube
Volume of vessel
Tube numbers
m
m
m2
m3
-
0.080
7.000
1.759
5.978
101.944
8.2.9 CO2 removal separator (S102)
We use physical and chemical absorption of CO2 in a MEDA solvent. From ASPEN
data the inlet gas flow (122) is 53.49kg/s (5.13 m3/s), the absorbed outlet flow
(123) is 10.44kg/s (0.263m3/s). The loading ratio of CO2/MDEA is 1mol/mol 18
and the density is 1.038kg/l, which means that the flow of MEDA solvent is
28.28kg/s (0.027m3/s). We approximate the ratio of the length and the diameter
of the vessel is L/D=3.0. The liquid holdup time is 10min normally. The
calculation of the vapor, liquid volume and the total volume of the separator are
shown in table 8.12.
Table 8.12 the approximation and calculation of vessel volume
Approximation
L/D ratio
Diameter of the vessel
(D)
Length of the vessel
(L)
Height of the liquid/ Diameter (Hl/D)
Height of the vapor/ Diameter (Hv/D)
Calculation
Vapor cross section area
Vapor velocity
Vapor residence time
Vapor volume
Liquid holdup time
Liquid volume
Total vessel volume
m
m
3
1.5
4.5
0.5
0.5
m2
m/s
s
m3
s
m3
m3
0.88
5.81
0.78
3.98
600.00
3.98
7.95
8.3 Equipment data sheets
All the equipment specification is shown in the Appendix 8.9 and 8.10
18
http://chemfinder.cambridgesoft.com/result.asp
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9. Waste
9.1 Introduction
As environment was getting worse, people started to realize importance of
environmental protection. Therefore, nowadays some regulations have to be
complied with by people, especially by company. In our design case, although
the plant is planning to be placed in Brunei, South-East Asia, and European
emissions standard has to be followed.
There will be some wastes produced in our designed plant, though we have tried
our best to avoid it. As far as we know, wastes can be classified into direct and
indirect wastes. Indirect wastes include all pollution as a result of product usage,
e.g., fuels for heating or traction. This category is not included in our CPD design
project19. The direct wastes of this plant are listed here:
$
$
$
$
$
$
Wastewater from synthesis gas production and FTS process;
CO2: from synthesis gas production, FTS process and burning fuel gas;
Carbon as coke or soot: from syngas production, FTS process and
hydrocracking reaction;
Oxygenates: alcohols from FTS process (assuming no acid produced)
Wax: unconverted hydrocarbon from hydrocracking
Nitrogen or oxides: NOx, NH3 or Nitrogen
9.2 Waste treatment
In fact, the waste does not exist in the process separately, but in the following
forms. In order to satisfy the emission standard of the Europe Commission, all
the wastes will be treated first before emission to environment.
1. Waste water from synthesis gas production and FTS process
There are two sources for production of wastewater within this plant, one is
produced by FTS process, which contains some alcohols or certain amount of
hydrocarbon inside; the other is from excess steam of syngas production unit,
which contains some NOx or NH3, due to nitrogen presence in the feedstock.
In order to meet the European emissions standard, the wastewater has to be
treated. The wastewater can be dealt in a simple flash to separate the alcohol,
which might be sold in market. Due to low concentration of alcohols, the income
of alcohol cannot make up for the energy cost of the flash operation. Therefore,
another treatment method for the wastewater is taken to meet the
environmental requirements in Europe, which is so called biological method.
After treatment, the water can be discharged according to European emissions
standard.
2. Separated CO2 after syngas production unit
In the syngas production unit, we use physical and chemical absorption method
to absorb CO2; the purity of separated CO2 is so high that it might be sold to food
19
J.Grievink, C.P. Luteijn, P.Swinkels; Instruction manual of conceptual design, p26-27; July, 2002
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industry. For this purpose, several compressors are needed to increase the CO2
pressure, however, which is beyond the design limit. Therefore, we are going
into details here.
3.Waste gas from syngas productions or burning the fuel gas
The side reaction of syngas production unit will generate some coke, which exists
as soot in the emission gas and will flow to fuel gas branch at last. Excess steam
is fed to the first unit (R100) to inhibit coke formation. Therefore, we may
assume no soot will exist at the end. Even in the abnormal situation such as
shutdown or startup, the content of soot in emission gas can suit the European
standard. There are also some NOx, NH3 or N2 in the waste gas, which originates
from the feedstock of natural gas and oxygen. These things will flow to the fuel
gas in the process and be emitted to the atmosphere, because the content of
nitrogen in natural gas or oxygen is very low, the waste gas can be emitted
directly.
4. Coke formation in F-T process and hydrocracking reaction
Coke formation on catalyst in FTS unit is inevitable. Because the catalyst in
slurry reactor is regenerated during the process, coke can be moved away with
the used catalyst. In hydrocracking unit, the coke will attach to the catalyst and
poison the catalyst. The poisoned catalyst is regenerated by combustion in a
steam of diluted oxygen or air. Upon combustion at temperature 400-500ºC, the
coke is converted to CO2 and H2O20.
5. Wax
To avoid the unconverted wax accumulated in the system, a small amount of
wax is purged, which can be dealt by certain treatment. Then, it can be
transported to landfill.
6. Used catalyst
Unrecyclable catalyst has to be thrown away after its lifetime. The used catalyst
can be regarded as special waste. After some treatment, it can be sent to landfill.
9.3 Emission Limit Values
According to the task description, the plant emissions should follow the EU
standard. The corresponding standards to plant emissions gas, water and solid
waste are chosen by handbook of Industrial Pollution Control and Risk
Management.21
The definition of emission standard is the maximum amount of discharge legally
allowed from a single source, mobile or stationary.22
20
J. Scherzer, A. Gruia, Hydrocracking science and technology; P123, Marcel Dekker, Inc., 1996
http://europa.eu.int/comm/environment/enlarg/handbook/pollution.pdf
22
ETC/CDS. General Environmental Multilingual Thesaurus (GEMET 2000)
http://glossary.eea.eu.int/EEAGlossary/E/emission_standard
21
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9.3.1 Air Emission Limit Values
The purpose of the incineration plants established and operated in accordance
with the Directive, which is a consequence of the fifth Environment Action
Programs, is to reduce the pollution-related risk of waste through a process of
thermal treatment. The aim of this Directive is to prevent negative effects on the
environment, in particular the pollution of air, soil, surface water and
groundwater, from the incineration and co-incineration of waste and, to that end,
to set up and maintain appropriate operating conditions and emission limit
values.
The EU standards 88/609/EC “The limitation of emissions of certain pollutants
into the air from large combustion plants” 23 will be applied here. Although this
standard has been amended by 94/66/EC24 in 1994, it will not affect this plant.
Because only a new limitation of sulfur is made, the 88/609/EC still can be used.
Table 9.1 Daily average values of air emission limit
Total dust
Gaseous and vaporous organic substances, expressed
as total organic carbon
Nitrogen monoxide (NO) and nitrogen dioxide (NO2),
Expressed as NO2 for existing incineration plants with
a capacity of 3 tonnes per hour or less
10
mg/m3
10
mg/m3
400 mg/m3
9.3.2 Water emission limited value
There is no any heavy metals in the wastewater from the plant, so no attention
need to be paid to these kind of standard. The surface fresh water standard is
used here. The standard “Council Directive 80/778/EEC of 15 July 1980 relating
to the quality of water intended for human consumption25” is applied here.
23 http://europa.eu.int/smartapi/cgi/sga_doc?smartapi!celexplus!prod!CELEXnumdoc&lg=en&numdoc=31988L0609
24
http://europa.eu.int/smartapi/cgi/sga_doc?smartapi!celexplus!prod!DocNumber&lg=en&type_doc
=Directive&an_doc=1994&nu_doc=66
25
http://europa.eu.int/smartapi/cgi/sga_doc?smartapi!celexplus!prod!CELEXnumdoc&lg=en&numd
oc=31980L0778
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10. Process Safety
Nowadays the requirements for occupational health and safety in plants have
increased substantially. As a designer, we must try to reduce the risks when we
are doing the design. In order to quantify the potential hazards, two tools will be
used, i.e. Hazard and operability study (HAZOP) and Dow Fire and Explosion
Index (FEI) assessment. And the specific analysis is shown below.
10.1 Hazard and operability studies (HAZOP)
10.1.1 Introduction of HAZOP
A hazard and operability study is a procedure for the systematic, critical,
examination of the operability of a process. When applied to a process design or
an operating plant, it indicates potential hazards that may arise from deviations
from the intended design conditions. This type of study is usually referred to as a
HAZOP study.
Even experienced and competent designers make mistakes and omissions during
the design process, so HAZOP is a necessary assistant design tool. Though the
accidents will still happen even when HAZOP is carefully applied, the number of
accidents should be smaller and their consequences should be less severe.
HAZOP will produce a significant number of design changes, such as additional
plant control or equipment protection, which will prevent accidents and improve
safety, then save money at last.
In view of that our process deals with flammable gases and operates in high
pressure (20-40bar) conditions, HAZOP seems to be extremely important to the
plant safety. The HAZOP is applied here according to Coulson & Richardson’s
CHEMICAL ENGINEERING volume 6 Chemical Engineering Design. With HAZOP
the design is systematic studied vessel-by-vessel, and line-by-line, using “guide
words” to help generate the thought about the way deviations from the intended
operation conditions can cause hazardous situations. Common used guidewords
are given in table 1. The process parameters can be chosen from flow, pressure,
temperature, mixing control, level, reaction, start/stop, separation, operate,
maintain. The meaningful combination of the guidewords and parameters are
used. 26
Table10. 1 Standard guidewords and their generic meanings27
Guide word
Meaning
No (not, none)
None of the design intent is achieved
More (more of, higher)
Quantitative increase in a parameter
Less (less of, lower)
Quantitative decrease in a parameter
As well as (more than)
An additional activity occurs
Part of
Only some of the design intention is achieved
Reverse
Logical opposite pf the design intention occurs
Other than (other)
Complete substitution – another activity takes
place
26
R.K.Sinnott; Coulson & Richardson’s Chemical Engineering Volume 6, P339-P347; 1993
Prof. Dr. Ir. H.J. Pasman, Dr.Ir.S.M.Lemkowitz, Chemical risk management, 2003, p289,
table5.9
27
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Except the recommended guidewords shown above, the following words are also
used in a special way, and have the precise meanings given bellows:
$
$
$
$
$
Intention: the intention defines how the particular part of the process
was intended to operate; the intention of the designer.
Deviations: these are departures from the designer’s intention, which are
detected by the systematic application of the guidewords.
Causes: reasons why, and how, the deviations could occur. Only if a
deviation can be shown to have a realistic cause is it treated as
meaningful.
Consequences: the results that follow from the occurrence of a
meaningful deviation.
Necessary action: the actions that need to be done to avoid this kind of
hazards happen.
10.1.2 HAZOP Analysis
The HAZOP analysis will proceed on the basis of unit, and the sequence will
follow Coulson & Richardson’s CHEMICAL ENGINEERING volume 6 Chemical
Engineering Design. And therefore, the corresponding result summary is shown
as Appendix 9.1.
10.2 Dow Fire and Explosion Index (F&EI) method
The hazard of the plant is ranked by the method of the Dow Fir and Explosion
Index, which is based on amounts of substances, material properties of those
substances, process conditions, and preventive a protective measures. The
higher value means the highest risk.
To estimate the system hazard potential, the material factor of the feedstock,
products and some intermediates are calculated here. The MF is decided by the
National Fire Protection Agency (NFPA) values for flammability (Nf), Reactivity
(Nr), and Health (Nh), the values of all the substance is show in the following
table10.2 28:
Table10. 2 The safety property of the related material
298
301
1038
Hc
[BTU/lb.*10-3]
21.5
51.6
4.3
Flash
point [F]
Gas
Gas
Gas
683
-626
18.7
18.0
19.9
100/130
28/85
Gas/-40
Common name
Td [K]
Natural gas (methane)
Hydrogen
Carbon monoxide
Kerosene29
Diesel
Naphtha
LPG (Butane/Pentane)
28
Nh
Nf
Nr
MF
1
0
3
0
0
1
1
4
4
3
2
2
3
4
0
0
1
0
0
0
0
21
21
16
10
10
16
21
Dow’s fire& explosion index hazard classification guide; American institute of chemical
engineers; 1981
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The general process hazards (GPH) and Special process hazards (SPH) providing
weighing factors as penalties and credits in positive or negative sense
respectively added to the value 1.00. The resulting value after the additions is
multiplied with the MF-value of the substance involved. In formula:
Index value = MF * GPH * SPH 30 . There are always several substance involved in
one unit, the highest MF will be employed.
The property of the reaction in unit and the plant process conditions are taken
into account in the GPH calculation. The whole process is in an open area. So the
item D of GPH is zero of all units. In the plant the access for the emergency
equipment to approach an operation area from at least two sides, so the item F
also equals to zero. About F, there are adequate designs of drainage, which can
directly spill away the spill from the process unit.
In SPH, the operating condition combined with the reactant property is involved.
A: Process temperature:
Syngas production: process temperatures are above the boiling point of the
material, so apply a penalty of 0.6.
Other units: process temperature is above the flash point but below the boiling
point of the material, so apply a penalty of 0.3.
B: Low pressure:
There is no any low pressure applied.
C: Operation in or near flammable range:
The operation of the whole process is below or near the flammable range. Only
in case of instrument of equipment upsets or purge failed, the penalty is 0.3.
D: Dust explosion:
No danger of dust explosion because the absence of the dust in the whole
system.
E: Relief pressure:
The whole process is operated above atmospheric, the hazard results from the
potentially large release of liquid and gas can occur. Use the safety relief valve
setting or rupture disk rating to determine the penalty point, the results show in
the table10.3.
Table10. 3 The pressure penalty of different unit
Unit
Pressure
Pressure
Set pressure
[bar]
[psig]
[Psig]
Syngas
FT
Hydrocracking
Seperation
20
30
40
2
275
421
566
15
305
450
595
44
Penalty
0.54
0.65
0.71
0.20
F: Low temperature:
There is not any unit running below the transition temperature.
G: Quantity of Flammable material:
29
Material safety data sheet master list (MSDS master sheet list);
http://bcdhscwebs.tambcd.edu/bcdfacility/msds_main.html; Safety data sheet:
http://www.elgas.com.au/safety/msds_LPG.pdf
30
H.Pasman, S.Lemkowitz; Chemical risk management, P279-290; TU Delft, 2003
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The flash point of all material involved in the process are less than 140ºF, and
belong to the category 1 “liquids or gases in process”, and the penalties
determined by enthalpy according to Chemical Risk Management31.
Table10. 4 The flammable penalty of different units
Unit
Material
Amount
BTU*
(Pound)
10-3/lb
Syngas
Natural Gas
666
21.5
FT
Hydrogen
563
51.6
Carbon monoxide
3915
4.3
Hydrocracking Hydrogen
122
51.6
Separation
Naphtha
529
1.80
Kerosene
368
1.87
Diesel
478
1.87
None
BTU *10-9
Penalty
0.014
0.046
0.1
0.1
0.151
0.25
0.025
0.1
H: Corrosion and erosion:
J: Leakage- Joints and packing:
Penalty equals to 0.1, because of the pumps and gland seal leakage.
The calculation of GPH, SPH and Index Value of four units in the plant are show
in the Appendix10.2 and results are show in the table10.5:
Table10. 5 the F&EI value of four units:
Syngas production FT synthesis
Hydrocracking
126.5
120.8
119.4
Distillation
84
The final index is 126.5, which indicates that the degree of hazard for F&EI of our
process is intermediate according to the table10.6
Table10. 6 The hazard degree standard
Index range
1-61
62-96
97-127
Degree of hazard Light
Moderate Intermediate
128-158
Heavy
159-up
Severe
10.3 Conclusions
Because the materials applied in this design are flammable, the F&EI value is a
little bit too high. Usually if the resulting index value is above 100, in many
companies the degree of hazard is judged to be too high, and risk-reducing
measures are required. Therefore, more attention should be paid to the safety
here.
31
H.Pasman, S.Lemkowitz; Chemical risk management, P20, figure 3; TU Delft, 2003
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11. Economy32
11.1 Investment
The detailed purchased equipment cost calculation is introduced in the Appendix
10; the result is summarized in the table11.1. Abbreviation of PEC indicates the
purchased equipment costs.
Table 11.1 Equipment costs and PEC
Equipment
Cost [million £]
Reactor
0.323
0.061
Column
Heat exchanger
1.280
Compressors and pump
0.628
Mixer
0.004
PEC
2.296
And the physical plant cost (PPC) is estimated by the following equation:
PPC=PEC*Lang’s factor
The process type of the whole FT synthesis is fluid one, according to the
33
reference , the Lang’s factors of this process is 3.4. Accordingly then estimate
the PPC, indirect capital costs and fixed capital costs. The above calculation is
based on the price in 1992, in order to estimate the costs in 1998, the profit 7%
per year is used. And we transfer UK Pond to US dollar by the ROE=1.633, and
Nfl to US dollar by ROE=0.505, which is the data of 199834. All results are listed
in Table 11.2:
Table 11.2 The cost of different part and year
Lang's factor
Cost [million £]
PCE
2.296
PPC
3.4
7.806
Indirect Capital Costs
0.45
3.513
Fixed Capital Costs @1992
1.45
11.318
Fixed Capital Costs @1998
16.986
From the fixed capital costs, and the percentage of fixed costs to total
investment, we can deduce the total investment cost and other costs.
Table 11.3 The fix license, working and total cost
Percentage to
Costs
Costs
total costs
[Million £] [Million $]
Fix costs
80
16.986
License costs
14
297.255
Working costs
6
127.395
Total investment costs
100
21.232
34.679
32
33
34
Coulson& Richadson’s Chemical Engineering, Vol.6, 2nd edition, 1993, chapter
Coulson& Richadson’s Chemical Engineering, Vol.6, 2nd edition, 1993, chapter
http://fx.sauder.ubc.ca/data.html
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11.2 Cash flow
Operating cost
There are two kinds of operating costs. One is the fixed cost, which is
independent on the produced quantity; the other is the variable production cost,
which is dependent on the amount of production and the process conditions.
Firstly, we estimate the variable production, which consists of
• Raw material
• Miscellaneous operation materials
• Utilities
• Shipping and packaging, which is assumed as zero
All the data needed to estimate the raw material costs and unity costs are taken
from the Chapter 5. It is assumed that the catalyst lasts for five years, this
results in catalyst cost shown in table 11.4. The raw materials and utility costs
per year are summarized in table 11.5 and table 11.6.
Table 11. 4 Catalyst cost of the whole process:
Type
Amount
Life
Consumption
Price
Cost
[kg]
[y]
[kg/year]
[$/kg]
[k $/y]
Unit
100 Ni/Al2O3
200 Co/MgO/SiO2
300 Pt/Zeolite
Total
3.50E+02
4.20E+04
1.00E+05
5
5
5
1.42E+05
7.00E+01
8.39E+03
2.00E+04
17
10
25
2.84E+04
1.190
83.920
500.000
585.110
Table11.5 The raw material cost per year
Raw material
Natural gas
Oxygen
Catalyst
Consumption [ton/year]
6.61E+05
4.61E+05
Price [$/ton]*
92.5
27
Cost [m $/y]
61.139
12.434
0.585
Total
74.158
* The price is gotten from the client
Table11.6 The utility cost per year
Consumption
Cost per unit **
[Unit/y]
[Nfl/unit]
Utility
Electricity
55,354,400kwh
Cost
[k Nfl/y]
Cost
[k $/y]
0.13per kwh
7196.072
3634.292
Steam
1,777,648t
30.00per Ton
53329.440
26933.411
Cooling water
Total
1,820,800t
0.05per Ton
91.040
60616.552
45.979
30613.682
**The price is gotten from manual
Secondly, the fixed costs consist of:
• Maintenance
• Operating labor
• Laboratory costing
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Final report
•
•
•
•
•
•
Supervision
Plant overheads
Capital charges
Insurance
Local taxes
Royalties
Operator cost:
One operator for every unit and a supervisor of the whole process will be
arranged every shift, three shifts are normal, and 5 shifts are set for space.
5 operators/shift * 5shift*100,000Dfl/operator/year
=2.5million Nfl/year
(11.1)
=1.263 million US dollar/year
The other fixed cost is calculated according to the description of the table 11.7
and all the results are displayed in the below table also.
Table11. 7 The layout of all the costs35:
Cost type
Description
Cost
Percent
[M $/a]
%
1Raw material
See table 11.5
2Miscellaneous materials 10% of the maintenance
3Utilities
See table 11.6
4Shipping and packaging Usually negligible
Sub-total A
74.158
0.208
30.614
55.66%
0.16%
22.98%
Zero
0.00%
104.978
78.79%
1.56%
0.95%
0.19%
0.19%
Fixed costs
8Supervision
20% of the operating labor
2.081
1.263
0.253
0.253
9Plant overheads
50% of the operating labor
0.631
0.47%
10Capital chargers
15% of the Fixed capital
11Insurance
1% of the Fixed capital
12Local taxes
2% of the Fixed capital
13Royalties
1% of the Fixed capital
0.189
0.347
0.694
0.347
0.14%
0.26%
0.52%
0.26%
6.056
4.55%
111.035
83.33%
22.207
16.67%
22.207
16.67%
5Maintenance
6% of the Fixed capital
6Operating labor
7Laboratory costs
20% of the operating labor
Sub-total B
Direct production costs
14Sales expense
15General overheads
20% of the direct production cost
16Research and development
Sub-total C
Annual production cost=A+B+C
133.241
Production costs per mass unit: ($/Ton)
266.483
(Annual production costs/Annual production rate)
35
J.Grievink, Supplementary lecture notes on process systems design, 2003
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Final report
The annual production costs is 133.241million US dollar, and to produce one ton
product, 266 US dollar will be cost.
Income
The income of this factory comes from three main products, i.e. naphtha,
kerosene and diesel and one by-product LPG, other by-product such as wax or
pure carbon dioxide and fuel gas can be omitted.
Table 11.8 The income from the product per year
Product Price [$/t]*** Productivity [t/y]
Income [million $/y]
LPG
154.8
0.000
0.00E+00
Naphtha
130
1.32E+05
17.148
Kerosene
135
1.71E+05
23.141
Diesel
120
1.90E+05
22.754
Total
4.93E+05
63.043
*** The price is get from client
So, the sales income should be 63.043 million $ per year
The net cash flow
NCF=sales income-production costing
=63.043-133.241
=-70.198 [million dollar/year]
(11.2)
11.3 Economic evaluation of the project
Rate of return (ROR) and pay out time (POT)
Cumulative net cash flow at end of project
*100%
Life of project * original investment
NCF (before tax )
*100%
=
original investment
ROR =
(11.3)
Pay back time=100/ROR
Table 11.9 The economic criteria of the project
Value [m $]
34.679219
63.043
133.24145
-70.198
-2.024
-0.494
Investment
Income
Cost
Net Cash flow
ROR
POT
To calculate the DCFRR, we just try to find a value that satisfy the equation
n =t
NFV
∑ (1 + DCFRR)
n =1
n
= 0 (t = the life of the project , here = 17)
(11.4)
We cannot calculate the DCFRR of this system because the margin is negative,
which means the cumulative cash flow will get more negative with time, and
there is no chance to earn the investment back.
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11.4 Cost review
From the table 11.7, we can see that, in the operation cost, the most important
items are feedstock and utility. Because the conversion of related reaction is high
enough, there is small space for the feedstock amount deceasing. What we hope
is the price of the feedstock decrease, perhaps in remote area, which can be
realized. It means the location of the factory is extremely important. For the
utility, much steam is used, which is bought from the outside. By the way, there
are several fuel gas sources involved in our designed plant. However, fuel gas is
quite difficult to be transported, therefore, we are not going to sell it. Instead the
fuel gas is burnt to generate hot steam, which is going to be as utilities in our
plant.
11.5 Sensitivities
Sensitivity of economic criteria is analyzed here with respect to investment,
operation costs (two biggest items are selected, feedstock and utility) and
product price (for there is no any LPG produced in this case, the LPG price
change is excluded). First each item will increase 10% and get the new value,
estimate the NCF, ROP and POT of the project again, and then compare with old
one, the sensitivity of the related item price to the economic criteria is expressed
as changing percentage. The result is shown in the following table.
Table 11.10 the sensitivity of the investment operation costs and product price
to the economic criteria
NCF
ROP POT
New
Value
New
NCF
Variances (+10%)
New
Change
ROP
degree
New
POT
Change
Degree
Invest
34.679
38.147
-70.198
-1.840
9.09%
-0.543 10.00%
Feedstock
Operation
cost
Utility
74.157
81.572
-77.614
-2.238
-10.56%
-0.447 -9.55%
33.675
-73.260
-2.112
-4.36%
-0.473 -4.18%
18.862
-68.484
-1.975
2.44%
-0.506
2.50%
Product
price
30.614
2.024 0.494
Naphtha
17.148
Kerosene
23.141
25.456
-67.884
-1.957
3.30%
-0.511
3.41%
Diesel
22.754
25.030
-67.923
-1.959
3.24%
-0.511
3.35%
Feedstock cost is most sensitive item to the economic criteria.
11.6 Net cash flows
Just as mentioned before, the net cash flow of this case is negative, -70.198. For
requirement of the earning power 12%, 5.397 million US dollars must be earned
every year. There is two ways to arrive this object; one is the price of all
products increase 220%, in table 11.11, the original and new situation are
compared, and it is obvious that the NCF rise from -70.198 to 5.454, which a
little more than 5.397 million dollar. Another way is decrease the operation cost,
in which raw material and utility is the major item. So once raw material and
utility cost decrease 59.6%, and because the item 14- 16 is partly based on
them, the total operation cost will decrease dramatically, and also the NCF will
bigger than5.397 million dollar.
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Final report
Table11.11 The effect of the products price to the NCF
Stream
Unit
LPG
Naphtha
Kerosene
Diesel
Total Income
NCF
Productivity
t/y
0.00E+00
1.32E+05
1.71E+05
1.90E+05
4.93E+05
Price
$/ton
154.8
130
135
120
Income New price New Income
Million $/y $/ton
Million $/y
0.000
17.148
23.141
22.754
63.043
-70.198
340.56
286
297
264
0.000
37.725
50.911
50.059
138.695
5.454
Table 11.12 The effect of the raw material and utility price to the NCF
Cost type
Raw material
Utilities
Others
Total operating cost
NCF
Cost [M$]
74.157
30.614
22.207
133.24
-70.198
New Cost[M$]
29.959
12.368
8.972
57.563
5.480
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12. Creativity and Group process tools
12.1. Group relation diagram
Client
Pieter
Swinkels
Assistant
Augustine Ajah
Creativity Coach
Cristhian Almeida
Rivera
Group
(5 person)
Tools
Design tools
Creativity
methods
DDM/ Notebook
Reference
books
Computer
Soft wares
Planning
tools
AAA
PIQUAR
Time
planning
Figure 12.1 Group Relation Diagram
12.2 Group Creativity evaluation
During the CPD project, several creativity tools have been applied to improve our
innovation design and creativity thinking. There are sixteen creativity methods in
the "Process conditions for using creativity in design work”, which we has applied
in our daily job of CPD project.
" Brainstorming
When problems arise, all of us will sit together and write down what any new
idea, one or two hours later all new ideas will be collected and discuss together
to find out the best choice.
" +/- Evaluations
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There are so many aspects in the whole design; evaluation is the option to
decide the steps of the final choice and give us a correct answer. Usually we will
list all the alternatives in the paper, and then write down all the advantages and
disadvantages of them, then eliminate some alternatives according to the quality
evaluation factor (PIQUAR) such as safety, economy, reliability, availability and
so on. And also use ASPEN Plus that is a very useful assistance.
" Diary with ideas, associations, new solutions
We work eight to ten hours every day (sometimes including Saturdays), but we
always take our small notebooks with us even at rest time. Conception of a new
idea often occurs in an intuitive flash of insight, in which the more or less
complete idea is revealed. So we will write it down when afflatus strike you and
spend some time to think over it when you are free or discuss with others the
possibility.
" Discussions about contradictory elements
Some of our requirements are contradict to each other. In this situation, we will
discuss the contradictory elements together and try to balance them or ever
create our own method to satisfy the entire requirement. It’s a hard work
" Methods for improved group work
We always work together and it’s convenient for our communication. We try to
give everyone enough thinking time and help each other when he or she has
problem and try to solve it together if need. We think creativity should be
cultivated in a good environment, and grow up with enough attention, patience
and freedom. So we do it together.
" Exploring alternative solutions, and values of different approaches
More alternatives mean more possible creativity. We try to find more alternative
solutions to the main process unit and then use the method 2 to make a final
choice.
" Open discussion on improvements of mistakes, on more direct
communication on productivity and on participation
Because all of us have the similar academic background and origin, we have the
same way of thinking, which means we will make same mistakes. So once we
find the mistake, we will inform all and discuss this kind of mistake in our group
meeting, then try to find similar mistake in other’s job and pay more attention to
this aspect in order to avoid them in our future job.
" Use of more outside- or inside –information; and of more experience in
practice
When someone has difficulty, he or she can get information from other group
members; if fail, we will try to get help from the professor. Although we are lack
of relative experience, we still can get the kind of information from our creativity
coach and client.
" Exploration of solutions applied in different, but comparable situations
Sometimes it is not possible to find an exactly method or technology in literature
which has the same situations as us, we just try to look for a similar one and do
some adjustment according to our situations.
" Reporting how mistakes have been found and improved
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All the mistake that have been found by anyone will be report to all the members
of the group, write it down, and discussion will be made to improve it.
" Explanation why quotes, statements and formulas were used
Our BOD design is base on the literatures. It is required Reporting how mistakes
have been found and improved in our group that everyone should explain his/her
quote, statements and formulas that will use in the design. We try to avoid any
misunderstanding of the literature.
An overview of main creativity methods used in CPD project is shown in table
below:
Table 12.1 Overview of creativity method
Time
Creativity method used
Week 1
Brainstorming
Week 2
Brainstorming
Week 3
+/-Evaluation
Week 4
+/-Evaluation
Week 5
Discussions about contradictory elements
Week 6
Open discussion
Week 7
Exploration of solutions applied in comparable situations
Week 8
Exploring alternative solutions
Week 9
Open discussion
Week 10
Explanation why quotes, statements and formulas were used
Week 11
Explanation how mistakes have been found and improved
Week 12
Evaluation
Week 1~12
+/-Evaluation
Week 1~12
Diary with ideas
12.3 Creativity Implication in CPD Project
During CPD project, there are some important innovation designs. For example,
in week6, at one routine group meeting, some of our team members illustrated
the flow diagram of basis of design. Then, all the team members started to
question the unclear demonstrations. After some arguments and discussions, we
found several process alternatives. After several weeks, some of those process
alternatives have been denied because of our enriching knowledge. But there is
one alternative we cannot determine from reference book or by common sense.
That is whether to apply distillation column before or after hydrocracking unit.
Because each has their own advantage and disadvantage, like we illustrated in
process options. Then, we come to ASPEN to simulate two options respectively.
To combine both hydrocracking unit and distillation unit is not an easy work,
because the existing two recycled loop. We have to optimize both units, and then
try to connect them. Option 1 is our first choice, however option 2 is very difficult
to tune. Then, we come to our conclusion that the desired quality of oil products
is difficult to reach for option2. Then, we choose option 1 in our final process.
After one week, some group members come up with a better way to optimize
distillation column. Then, we try to simulate option 2 again and got very good
results, which saved one third of hydrocracker size. This is one sample of our
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creative activities. We applied those creativity methods mentioned above through
our CPD project. We all think that those methods stimulate our way of thinking
and make our work more efficient and innovative.
12.4 Group Process Tools Evaluation
" DDM
We attempted to use Delft Design Matrix at our basis of design phase. But we
found that it took us too much time, and it’s impossible to apply DDM for our
project, which just last 12 weeks. In the main CPD design phase, we didn’t take
DDM as a main tool.
" AAA
The Advanced Activity Assistant has been used throughout our design activity to
recorder each of our activity per day and per person. This is a very used tool not
only for knowing what team members have done, but also providing a good
timetable to plan our further activity. Moreover, after comparing the time spent
on the same activity, we can have a good estimation on our time in further
activities, which make our work more efficient. For example, after knowing how
long do we need to prepare the kickoff meeting presentation, we left about the
same time for preparing BOD presentation, then, we can have enough time to
make our report better. After two times presentation, we have experience for
preparing presentation. I believe that we can also estimate how long do we need
to prepare our final review presentation. We must thank the tool AAA. Because
otherwise, nobody will pay attention to the time spent on each activity, and we
have to roughly estimate every time.
" PIQUAR
PIQUAR was found useful to help us identify the progress have been made during
design project. A graph of our PIQUAR numbers during the design spaces is
shown in Figure 12.1
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CPD_3296
Final report
1.00
0.90
0.80
PIQUAR numbers
0.70
0.60
wk4
wk7
wk8
wk9
wk11
wk12
wk3
0.50
wk2
0.40
0.30
wk5
wk6
wk10
wk1
0.20
0.10
0.00
Time [week]
Figure 12.1 PIQUAR numbers development during the design process.
Graph above shows that at the basis of design phase, we worked really hard and
found a lot of useful information. After BOD review, which is in week 5, we
became less tense and didn’t do much work. After that, we made a further plan
and realize that so much work is waiting for us. Then, we increase our workload.
Our progress is obviously from the graph. The deviation of the graph shows that,
some people work more and some less because the large deviation. After week
5, all the team-members work together, and everybody nearly have the same
workload and we always helped each other.
All in all, PIQUAR has helped us quantify our feeling of the design. It showed a
valuable tool in the development of a chemical process.
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Group Conceptual Process Design Project
CPD_3296
Final report
13 Conclusions and Recommendations
13.1 Conclusion
We fulfill the object, which is to produce 500,000 ton/year synthetic oil fuel,
naphtha, kerosene and diesel from natural gas by applying the Fisher-Tropsch
technology. To satisfy the demand of nearby market, the product distribution of
FT synthesis is adjusted to heavy products. Our plant produces almost no LPG.
As by-products, purged wax and fuel gas can be sold or reused inside the plant.
The process is consisted of four major units. In U100, 99% of natural gas is
reformed and combustion to raw syngas in combined autothermal reforming
(CAR) reactor. The purification and ratio adjustment of raw syngas can be
divided into three steps, which are membrane separation of hydrogen step,
water removal and carbon dioxide removal steps. In hydrogen purification step,
the separated hydrogen is compressed to hydrocracking unit U300; in CO2
removal step, H2/CO ratio is adjusted to about 2.0 and sends to FTS reactor
U200 and most of the separated carbon dioxide will recycle to CAR reactor
second reformer zone.
Because the low conversion rate is 80% in FTS, two stages slurry bubble column
reactor (SBCR) is arranged in U200. R210 and R220 are in parallel and are in
series with R230. Totally 96% of syngas converted to hydrocarbon and a large
amount of wastewater is produced. Then F-T wax is cracked at U300 by
hydrogen. All the hydrocracking products are transferred to the distillation
column U400 in turn, and the final products are separated according to the
products standard. Some unconverted wax is recycled back to U300.
In ASPEN simulation, all units are designed according to the products
requirement of the client. And after the simulation, the results can satisfy the
products specifications.
The most major wastes in our process are wastewater and carbon dioxide. The
amount of wastewater is large comparing with the products, but the quality is
not bad, because only a small amount of alcohol and acid are dissolved in
wastewater. Anyway it should be treated before pipe out of to the environment,
the treatment is outside of our battery limit.
Carbon dioxide comes from not only the syngas production unit but also the FTS
unit, which is mixed with fuel gas. Physical and chemical absorption in MDEA
solvent technology is employed in U100 to removal CO2, the purity of the
separated carbon dioxide is so high that can be sold to food industry. The rest of
carbon dioxide from other units can be purged to the atmosphere.
Concerning the safety aspect, the behaviour of this design is just fair. But for the
nature of the related reactions in this factory, it is acceptable.
Based on our design, the most serious flaw is in economic aspect. The net cash
flow is negative (–70.198 million dollar per year). The result is not good, but it’s
inevitable, because the price of the feedstock is expensive and the main products
are so cheap, the price difference is not much. We can consider the plant is only
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Final report
for make experience on FTS area and it maybe make money on a long-term
investment.
13.2 Recommendation
The efficiency of the feedback usage is high enough for the technology in the
year of 1998. Perhaps in syngas production unit, membrane reactor can be
applied in the future, and then oxygen will be replaced by air, which can save the
oxygen separation cost. Although, the nitrogen purification in the feedstock or in
the reactor will remove some energy, it still can be recovery by the heat
exchange network, furthermore, it is safer than the exist technology.
Though the heat exchange network here recovers heat, there is still some
unrecovered heat waste and removed by cooling water. To recovery this energy,
one suggestion is to generate steam inside factory, which does not be applied in
our design because of the safety reason, but if the furnace can be build a little bit
far from the factory, it’s still an attractive option. Another one is to construct an
energy required factory nearby where can also make use of the fuel gas.
There are two choices on the sequence of hydrocracking unit and separation unit.
We arrange hydrocracking unit before the separation unit. Actually option 2 is
better which is separation is before the hydrocracking (the details see chapter 8).
Unfortunately, we got the simulation too late and have no time to reconnect and
calculation everything. We suggest using this option in the future work.
According to the price in 1998, it is LPG production, not liquid fuel, more
profitable.
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Group Conceptual Process Design Project
CPD_3296
List of Abbreviation and Symbols
List of Abbreviation
AAA
AFS
ASTM
ATR
BLEVE
BOD
BPT
CPD
CPO
CPT
CAR
DCFROR
DDM
Dfl
EOS models
EP
F&EI
FTS
GPH
GTL
HAZOP
HC
HC
HP steam
HT
HTFT
I/O
LP steam
LNG
LPG
LTFT
MD
MF
MP
MP steam
MPDO
MPPD
MUSD
NCF
NFPA
NFV
NG
NPSH
PCE
PENG-ROB
PFR
Advanced Activity Assistant
Anderson-Flory-Schulz (distribution)
American Society for Testing and Materials
Auto Thermal Reforming
Boiling Liquid Expanding Vapour Explosion
Basis of Design
Bio Process Technology
Conceptual Process Design
Catalytic Partial Oxidation
Chemical process technology
Combined Autothermal Reforming
Discounted Cash Flow Rate of Return
Delft Design Matrix
Dutch florin
Equation Of State models
European Patent
Fire and Explosion Index
Fischer-Tropsch Synthesis
General process hazards
Gas To Liquid
Hazard and operability study
Hydrocracking
Hydrocrack(er)(ing)
High Pressure steam (40 bar)
Hydrotreating
High Temperature Fischer-Tropsch
Input Output
Low Pressure steam (3 bar)
Liquefied natural gas
Liquid Petrol Gas
Low Temperature Fischer-Tropsch
Middle Distillates
Material Factors
Medium Pressure
Medium Pressure steam (10 bar)
Maximum Probable Days Out
Maximum Probable Property Damage
Million U.S. Dollars
Net Cash Flow
National Fire Protection Association
Net future value
Natural Gas
Net Positive Suction Head
Purchased equipment cost
Peng Robinson
Plug Flow Reactor
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Group Conceptual Process Design Project
CPD_3296
List of Abbreviation and Symbols
PIQUAR
POT
POX
PR
PPC
PRMVH2
PSE
PSRK
RKS-BM
Plant design Improvement by QUAlity Review
Pay out time
Partial Oxidation
Peng Robinson
Physical plant cost
Peng Robinson of state with modified Huron-Vidal mixing
Process Systems Engineering
Predictive Redlich-Kwong-Soave equation of state
Redlich-Kwong-Soave equation of state with Boston-Mathias
modifications
ROR
Rate of return
SBCR
SMDS
SMR
SPH
SR-POLAR
Slurry Bubble Column Reactor
Shell’s Middle Distillation Synthesis
Steam Methane Reforming
Special process hazards
Schwartentruber-Renon equation of statefor highly non-ideal
systems
Synthesis gas
True Boiling Point
Tubular Fixed Bed Reactor
The Faculty of Applied Sciences
Delft, University of Technology
Utilities and auxiliaries
Unified Activity Coefficients model
UNIFAQ modified by Dortmund
UNIFAQ modified by Lungby
UNIFAC for liquid-liquid systems with Redlich-Kwong equation of
state and Henry’s law
United States of America Dollar
Syngas
TBP
TFBR
TNW
TU Delft
U&A
UNIFAC
UNIF-DMD
UNIF-LBY
UNIF-LL
USD
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Group Conceptual Process Design Project
CPD_3296
List of Abbreviation and Symbols
Symbol List
Symbol
Description
a
Gas-liquid interfacial area
a,b
reaction constant
A
heat transfer area
BoC
Bodenstein number for
catalyst particles
BoG
Bodenstein number for
gas phase
c
concentration
CG
Concentration in gas
phase
Ci
Molar concentration
SI Units
cm-1
m2
CH
mol/m3
cp,S
d
dp
D
DG
DL
E
Ea
G
h
H
Hydrogen concentration
in liquid phase
specific heat capacity at
const. pressure,
diameter
particle diameter
diameter
Gas-phase dispersion
coefficient
Gas-phase dispersion
coefficient
energy
activation energy
free enthalpy
heat transfer coefficient to
the cooling wall
specific enthalpy
He
H
H0
Hs
∆Hr
∆ Hv
k
k
Henry coefficient
Height of Hydrocracker
Height of un-gassed
Height of suspension
enthalpy of reaction
enthalpy of evaporation
reaction rate constant
First order reaction rate
constant
k0
Pre-exponential reaction
rate term
equilibrium constant
coefficient
Carbon dioxide reforming
equilibrium constant
mol/m3
mol/m3
moli/m-3L
J/(kg. K)
m
m
m
m2/s
m2/s
J
kJ/mol
kJ/mol
J/(cm2 s
K)
kJ/mol
cm3
kPa/mol
m
m
m
kJ/mol
KJ/mol
KJ/mol
molalkane_fee
dkgcat-1s-1
Symbol
Description
Kgw
Gas water shift equilibrium
constant
kj
Reaction rate constant of
reaction j
ksr
Steam reforming
equilibrium constant
Lcat
Length of total cat. Bed
m.p
melting point
Mw
molecular mass
MU
Viscosity
n
number of moles
n
reaction rank
Nc
Carbon number
p
pressure (absolute, total)
pi
Partial pressure of
component i
Pij
Probability of ith
component formation from
the jth component
PL
Vapour Pressure
Q
amount of heat
r
reaction rate (production)
rcr
Carbon dioxide reforming
reaction rate
rgw
Water gas shift reaction
rate
rj
Reaction rate of reaction j
SI Units
Pa2
rsr
molkgcat-1s-1
R
R
RHO
R
S
StH
StG
StL
K
Kcr
molalkane_fee
dkgcat-1s-1
-
t
t
tb
Pa2
-80-
Steam reforming reaction
rate
reaction rate
Gas constant
Density
gas constant
entropy
Stanton number for heat
transfer
Stanton number for gas
phase
Stanton number for liquid
phase
time
temperature
boiling point
m6kgcat1mols-1
m
oC
g/mol
cP
Pa, bar
Pa, bar
barg
J
kg/(m3 s)
molkgcat-1s-1
molkgcat-1s-1
molkgcat-1s-1
Jmol-1°C-1
mol/(m3 s)
mol/l
J/(mol K)
s
C
oC
o
Group Conceptual Process Design Project
CPD_3296
List of Abbreviation and Symbols
T
TW
∆Tm
U
U
Udf
Vsmall
V
Wcat
x
x
X
XH2
Thermodynamic
temperature
cooling wall temperature
Logarithmic mean
temperature difference
Gas superficial velocity
Overall heat transfer
coefficient
Gas velocity through the
dense phase
Rise velocity of small
bubble
volume
concentration of catalyst
in suspension %wt
mole fraction
axial coordinate
conversion
hydrogen conversion
K
K
oC
Greek
α
α
α*
ε
m/s
W/m2 °C
ε
εb
m/s
εdf
m/s
εs
m3
-
η
η
λS
θ
θ
ν
ν
ρ
ρ
ρ SK
ϕ
Φ
µ
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Description
relative volatility
contraction factor
modified contraction factor
porosity
total gas hold up
gas holdup in the dilute
phase
gas holdup in the dense
phase
catalyst concentration
(vol cat/ vol slurry)
dynamic viscosity
efficiency
thermal conductivity
SI unit
m3/m3
surface coverage (catalysis)
dimensionless temperature
stoichiometric coefficient
kinematic viscosity
T/Tw
m2/s
density
density of suspension
skeleton density
volume fraction
flow rate (mass, volume,
etc.)
viscosity of suspension
kg/m3
kg/m3
kg/m3
kg/s, m3/s
Pa s
W/m/K
kg/m s
Group Conceptual Process Design Project
CPD_3296
Final report
References:
1.
2.
3.
4.
5.
Julius Scherzer, A.J.Gruia ,Hydrocracking Science and Technology, 1996
Twigg M.V. 1989, Catalyst Handbook, Second edition, Wolfe Publishing Ltd. London 1989
Julius Scherzer, A.J. Gruia, Hydrocracking Science and Technology, chapter6, p74, 1996
Jacob A. Moulijn, Chemical process technology, 2001, p133
http://www.eng.auburn.edu/users/halljoh/ASPEN_Manuals/APLUS%20111%20User%20Gui
de.pdf
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.)
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
9. Julius Scherzer, A.J.Gruia, Hydrocracking science and Technology, chapter 10, p176
10. Julius Scherzer, A.J.Gruia, Hydrocracking science and Technology, chapter 14, page
244,1996
11. Julius Scherzer, A.J.Gruia, Hydrocracking science and Technology, chapter 12, page
216,1996
12. P.M. Torniainen, X. Chu, and L.D. Schmidt, Journal of Catalysis, 1994, 146, p1-10
13. S.S. Bharadwaj and L.D. Schmidt, Journal of Catalysis, 1994, 146, p11-21
14. Twigg M.V. 1989, Catalyst Handbook, Second edition, Wolfe Publishing Ltd. London 1996,
p265
15. Coulson & Richardson’s Chemical engineering ,volume 6,1993, p580-581
16. James M. Douglas, Conceptual Design of Chemical Processes. Page 490
17. M.Konno, M.Shindo, S.Sugawara and S.Saito; A composite palladium and porous aluminum
oxide membrane for hydrogen gas separation, Journal of membrane science, 37, 1988,
p193-197.
18. http://chemfinder.cambridgesoft.com/result.asp
19. J.Grievink, C.P. Luteijn, P.Swinkels; Instruction manual of conceptual design, p26-27; July,
2002
20. J. Scherzer, A. Gruia, Hydrocracking science and technology; P123, Marcel Dekker, Inc.,
1996
21. http://europa.eu.int/comm/environment/enlarg/handbook/pollution.pdf
22. ETC/CDS. General Environmental Multilingual Thesaurus (GEMET 2000)
http://glossary.eea.eu.int/EEAGlossary/E/emission_standard
23. http://europa.eu.int/smartapi/cgi/sga_doc?smartapi!celexplus!prod!CELEXnumdoc&lg=en&
numdoc=31988L0609
24. http://europa.eu.int/smartapi/cgi/sga_doc?smartapi!celexplus!prod!DocNumber&lg=en&ty
pe_doc=Directive&an_doc=1994&nu_doc=66
25. http://europa.eu.int/smartapi/cgi/sga_doc?smartapi!celexplus!prod!CELEXnumdoc&lg=en&
numdoc=31980L0778
26. R.K.Sinnott; Coulson & Richardson’s Chemical Engineering Volume 6, P339-P347; 1993
27. Prof. Dr. Ir. H.J. Pasman, Dr.Ir.S.M.Lemkowitz, Chemical risk management, 2003, p289,
table5.9
28. Dow’s fire& explosion index hazard classification guide; American institute of chemical
engineers; 1981
29. Material safety data sheet master list (MSDS master sheet list);
http://bcdhscwebs.tambcd.edu/bcdfacility/msds_main.html; Safety data sheet:
http://www.elgas.com.au/safety/msds_LPG.pdf
30. H.Pasman, S.Lemkowitz; Chemical risk management, P279-290; TU Delft, 2003
31. H.Pasman, S.Lemkowitz; Chemical risk management, P20, figure 3; TU Delft, 2003
32. Coulson& Richadson’s Chemical Engineering, Vol.6, 2nd edition, 1993, chapter
33. Coulson& Richadson’s Chemical Engineering, Vol.6, 2nd edition, 1993, chapter
34. http://fx.sauder.ubc.ca/data.html
35. J.Grievink, Supplementary lecture notes on process systems design, 2003
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