Off-shore Gas Utilization

FVO No. .

1

3161

1

Preliminary Design

Department of Chemical Process Technology

L

,filti

T

U Delft

Technische Universiteit Delft

Subject

Off-shore Gas Utilization

Authors

Arjan den Bouw

Gijsbert Korevaar

Pieter Lusse

Karin Remery

Keywords

Phone

015 2136281

015 2147266

015 2132562

015 2621830

Autothermal Reforming, Fischer-Tropsch, FPSO, Natural

Gas Utilization, Off-shore Technology, Synfuels, Syngas

Date of Assigment

Date of Report

November 29 1995

March 8 1996

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Abstract

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~o· This paper presents the study performed on oft-shore gas utilization. During oil production associated gas obtained which until now is flared. A study has been made of a process to convert the associated natural gas into synfuels bya Fischer-Tropsch synthesis route on a 'floating production, storage and offloading (FPSO)' system: Most of the process eqUlpment has been modeled and designed. Furthermore the process control has been designed and an economic study as weil as a safety, health and environment study have been performed.

The process has two main steps: conversion of the natural gas into syngas in an autothermal reformer followed by the conversion of this syngas into synfuels in a Fischer-::j. Tropsch slurry reactor. The designed process has a production capacity 3892 barrels per

~ day at a maximum feed rate of 29 million standard cubic feet per day, at the beginning of a field exploitation, and a production capacity of 590 barrels per day at a minimum feed "t'I'

i

(\

rate of 4 million standard cubic feet per day, obtained towards the end of a field life. The

\l '

produced synfuels are within the specification for tanker storage (Reid vapor pressure <

Q-~V

v 10 ps ia at 100 OF) so that they can be commingled with the crude oil. The designed

I. installation can deal with the large flow turndown ratio of 7.25. It can withstand up to 800

t"0 h ()~

vJ" ppm hydrogensulfide in the feed, any carbon dioxide concentration, but no mercury.

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Compositional variations of the feed of a maximum of 5 mole % methane replaced by

-..)0-.'< \ (, jleavies can be handled. The external utilities required for this unit are 6,000 m3/h of sea

water for cooling purposes, 8 MW electricity, a waste water tr.§.9..tme~capacity of 38,000

o

m3/h, and a permanent flare availability. The process rurts 360 days 'per year during 5

ft

i

years consecutively with a minimum of maintenance. The msla1~ation has been designed

to operate tor 20 years , i.e. exploit 4 fields, each during 5 years. The total installation fits in an area of 35 by 65 meters.

From the hazard and operability study it follows th at beside a normal control system some complementary safetyactions are required because of high explosion risks and the toxicity of components such as carbon monoxide use? in tbi~ process.

"",Cl.r~ l ~

7

The economic study shows that the process is profitable. The total investment costs are 131.106 US dollars and the productions costs are 15.106 US dollars per field. Two cases have been investigated. In the first case the oil price of crude oil (16 US dollars per barrel) have been used and in the second case, where the oil stored separately from the crude oil, the price of Fischer-Tropsch oil (26 US dollars per barrel) has been used. If the product is mixed with the crude oil the net income before taxes is 25.106 US dollars per field. In the second case the net income is 67.106 US dollars per field. The investments costs are high because the process equipment has to operate at sea and therefore the materials of construction have to be resistant against corrosion.

Table of Contents

1.

Introduction

2.

Objectives and Process Choice

2.1. Objectives of the Design 2.2. Possible Process Routes 2.3. Process Choice

3.

Assumptions

3.1. Input-Output Structure 3.2. Major Reactions 3.3. Plant Location

3.4. External Specifications and Conditions 3.5. The Feed

3.6. Used Chemicals 3.7. Availability of the Plant 3.8. List of Assumptions

4.

Process Structure and Flowsheet

4.1. Feed Pretreatment 4.2. Oxygen Production

4.3. From Natural Gas to Synthesis Gas 4.3.1. The Autothermal Reformer 4.3.2. Reformer Catalyst Selection 4.3.3. Operating Conditions

4.4. Syngas Cleaning: The Selexol® Unit 4.5. From Synthesis Gas to Synthetic Fuels

4.5.1. Choice of the Reactor

4.5.2. Description of the Slurry Reactor 4.5.3. Fischer-Tropsch Catalyst Selection

4.5.3.1.Catalytic Material : Iron or Cobalt 4.5.3.2.Choice of Mean Chain Length

4.5.3.3.The Catalyst ParticIe Size and Concentration 4.6. Heat Integration

4.6.1. The Steam Turbine

4.6.2. The Cooling of the Slurry Reactor 4.7. The Process Flowsheet

4.7.1. The Main Stream

5.

5.1. 5.2. 5.3. 5.4. 5.5. 5.6. 5.7. 5.8.

4.7.2. The Power Generation Process Simulation

=

~Q~::;') The Autothermal Reformer

The Selexol® Unit

The Fischer-Tropsch Reactor The Knock-Out Drums Pumps

Compressors

Expander and Steam Turbine Results

FVO 3161: Off-share Gas Utilization

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Table of Cantents

1

t(

2 2

~

2 3

4

4 4 5 5 6 7 7 7

9

9 9 9 10

_Vtl(

11 12 12 -

-P

o

13 13 13 14

A+~

14 14 14 15 15 15 15 15 16

?

17 ..-17 -

V"{

r:::-17

-

'

f-\

tG-

·

18 18

₍

18

_,

_{Ak b .}

18 18

)

18

6.

Equipment Design 19

6.1. Design of the Reactors 19

6.1.1. The Autothermal Reformer 19

6.1.1 .1. The Bed Sizing 19

_7~

6.1 .1.2. The Construction Materials 19

6.1.1.3. The Burner 20

6.1.2. The Fischer-Tropsch Reactor 20

+\

f:

G-6.1.2.1. Modeling Slurry Reactor 20

6.2. Other Pieces of Equipment 21

6.2.1. Knock-Out Drums 21

Af' .

6.2.2. Pumps 21 I

6.2.3. Compressors 22

6.2.4. Expanders 22(

6.2.5. Heat Exchangers and Condensers 22

6.3. Plant Layout 22

6.4. Results 22

7.

Mass and Heat 8alances

23

l

7.1. The Feed Streams 23

7.1.1. Natural Gas 23

~

7.1.2. Oxygen 23

7.2. The Product Streams 23

7.2.1. Synfuels 23

7.2.2. Purge 23

7.2.3. Carbon Dioxide 23

7.2.4. Water 23

7.3. The Cooling Streams 24

7.3.1. Closed Loops 24· 7.3.2. Sea Water 24

8.

Feed Variations

25

1

8.1. Flow Turndown 2S Y

+~

8.2. Compositional Variations 25 8.3. Feed Impurities 26 8.3.1. Heavy Metals 26 8.3.2. H2S 26 9. Process Control

27

!

9.1. Autothermal Reformer 27

9.2. Fischer-Tropsch Reactor and Recycle 27

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t

Y-9.3. Other Process Units 28

9.3.1. Knock-Out Drums 28

9.3.2. Compressors 28

9.3.3. Pumps 28

9.3.4. Heat Exchangers and Condensers 28

9.3.5. Stream Splitters 28

9.4. The Flare 29

10. Safety, Health, and Environment 30

\,

10.1. Safety 30

) A

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10.1.1. Hazard Identification 30

10.1.2. HAZOP Study 30

10.2. Health 31

10.3. Environment 32

Table of Contents

11.

Economics

33

[

11.1. Investment Costs 33

11.1.1. The Taylor Method 33

11 .1.2. The Lang Method 34

G.

11.1.3 The Zevnic-Buchanan Method 34

11.2 The Costs of the Process 35

11.3 The Net Income 35 I

11.4 Return on Investment 35

J

11.5 Internal Rate of Return 36

12.

Conclusions

37

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13.

Recommendations

38

_'-f

Acknowledgements

39

List of Symbols

40

References

42

Appendices A-G

List of Figures

Figure 2.1: Figure 3.1: Figure 4.1: Figure 4.2: Figure 4.3 : Figure 4.4: Figure 4.5: Figure 4.6: Figure 6.1: Figure 11.1:

Conceptual Process Design.

Input-Output Structure of the Process.

Computational Fluid Dynamic Calculation of the Flow Pattern in the Combustion Zone of an Autothermal Reformer.

A Schematic Representation of the Autothermal Reformer. A schematic representation of the ICI 54-8 catalyst.

Computational Fluid Dynamic Calculation of the Temperature Profile in the Combustion Zone of an Autothermal Reformer.

Schematic Representation of the Fischer-Tropsch Slurry Reacor. The Anderson-Schulz-Flory Distribution with CL

=

0.86.

The Schematic Plant Layout.

Process Flowsheet for the Taylor Co st Calculation.

List of Tables

Table 3.1. Table 3.2. Table 3.3. Table 3.4. Table 4.1. Table 6.1. Table 10.1. Table 10.2.

Specifications of Raw Materials and Products Composition and Conditions of Gas Feed Streams Gas Flow Turndown

Physical Properties of Used Chemicals ICI 54-8 catalyst data.

Final design parameters of the Fischer Tropsch reactor Results of the DOW Fire and Explosion index

MAC-values and volatility of chemical compounds

FVO 3161: Off-shore Gas Uti/ization

List of Appendices

A Conceptual Design According to Douglas A-1

ç

B Process Flowsheet 8-1

fi-c

Calculations and Results of the Simulation C-1

1

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C1 ChemCad Flowsheet C-2

C2 Stream Compositions C-3

C3 Outlet Composition of Autothermal Reformer C-13

)

J.\ G

C4 Stoichiometrics Fischer-Tropsch Reactor C-14

D Equipment Design and Specifications D-1

01 Effect of Compositional Variation of the Feed 0-2

02 Design Calculations of Knock-Out Drums 0-3

?-+~

03 Design Calculations of Pumps, Compressors,

and Expanders 0-6

04 Design Calculations of Heat Exchangers and Condensers 0-9

A

05 Equipment Specification Sheets 0-14

- Pump P26 0-14

-Compressor C4 0-15

- Heat Exchanger H33 0-16

06 Equipment List 0-17

- Reactors, Columns, and Vessels 0-17

- Pumps, Compressors, Expanders, and Filters 0-21

- Heat Exchangers and Condensers 0-25

List of Symbols used in Appendix 0 0-27

E Mass and Heat Balance E-1

F Safety Analysis

F-1

1

~

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F1 Hazop Studies F-2 - Autothermal Reformer F-2 - Fischer-Tropsch Reactors F-4

(

F2 Dow Fire & Explosion Index Hazard Classification F-5

- Autothermal Reformer - Fischer-Tropsch Reactors - Knock-Out Drums

G

Economics G-1

~

G1 Taylor Investment Calculation G-2

G2 Zevnic-Buchanan (Jansen) Investment Calculation G-3

\

Û

G3 Lang Investment Calculation G-4

G4 Total Costs and Revenues G-5

G5 Return on Investment G-6

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G6 Internal Rate of Return G-7

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Chapter 1: Introduction

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Introduction

During oil production associated gas is always obtained in varying quantities. Until now,

when there is no gas pipeline or infrastructure~ this gas is being flared or reinjected into

the weil. In this project, a preliminary design of a process to convert associated natural gas into synthetic fuels has been made.

The goal of reinjection is to increase the pressure in order to favor the extraction of oil from the weil. But reinjection of associated gas is not mandatory; water can be reinjected as weil to achieve this. Reinjection can however be justified because the gas could be recovered at a later time. This is not the case for flaring. It serves no purpose, produces emissions, and the energetic resource is lost. At this time, as it has become c1ear th at the oil resources are not endless and the energy needs are still growing, attention is given to other, often harder to recover, energy sources. In this perspective, it will probably become

unacceptable to flare gas and profitable to process it, even in remote locations.

Single Buoy Moorings (SBM), a company specialized in oft-shore oil production systems, is therefore interested in an alternative for flaring. It was our assignment to study the possibilities.

The challenge is to get the associated gas from remote off-shore locations to an on-shore ,

site with infrastructure. Pipelines are excluded because of the remote character of the

location. Transporting the gas in tankers requires liquefying. The required LNG tankers, in I

addition to the oil tankers, makes this solution economically not feasible. The idea of the process studied in this project is to convert the associated gas into an additional oil

product stream which can be commingled with the main oil product stream. The

specifications for the synfuels are those of stabilized oil suitable for tanker storage, On-shore, it will be further processed in a refinery.

From an economic point of view, it is worthwhile to mention that the natural gas feed stream is available without additional costs. The other two main required feed streams are

oxygen and steam. These will have to be produced yl~shore and will bring major costs.

The synfuel prodUct stream will increase the overall Y)'~f of the weil.

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2.

Objectives and Process Choice

In this chapter the objectives of the design will be presented. Af ter a brief description of the available technology, the process route on which the plant design will be based will be described.

2.1.

Objectives of the Design

The goal of the assignment is to design an installation, th at meets the following requirements:

• The unit converts associated gas into synfuels that meet the specifications of

stabilized oil suitable for tanker storage (RVP < 10 psia at 100 OF).

• The unit will be installed on a floating production, storage and offloading (FPSO)

system and thus must be able to operate with the utilities available on board and meet the requirements of an offshore installation.

• The unit has to operate with a maximum gas feed of 29 MMSCFD and a gas feed

minimum of 4 MMSCFD. ~ O\l~ r <;

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- .

.

. .

• The unit has to require a minimum of maintenance and have a lifetime of 20 years ..

2.2.

Possible Process Routes

There are two main pathways to convert natural gas into synthetic fuels ("synfuels" for short).

The first one consists of two steps. In the first step, natural gas is converted into synthesis gas, a mixture of hydrogen and carbon monoxide, also called syngas. This syngas is then in turn converted into liquid hydrocarbons, i.e. synfuels. This process originated in the beginning of this century with the Germans F. Fischer and H. Tropsch. They discovered that syngas made out of coal could be converted into a range of products such as methane, methanol and hydrocarbons, depending on the process conditions and the type of catalyst[1],[28]. This process was used by the Germans during World War II to produce gasoline out of coal, but has been neglected shortly after the war with the rise of oil based technologies. However, it has become interesting again in the last decades because of the declining oil resources.

Today, several commercial installations based on this technolo~~ are operating, such as

Sasol's process to convert coal into synfuels in South Africa 8] and the Shell Middle

Distillate Synthesis (SM DS)[40],[43],[52] process in Malaysia using natural gas to make synfuels.

Another route to convert natural gas into synfuels is via methanol. In this process the natural gas is converted into syngas, which is converted into methanol. The methanol can in turn be converted into synfuels[28],[57]. This is done by Mobil in its Methanol to Gasoline (MTG) process in New Zealand.

Air

Simplified flowsheet

N, and other components

Water

CO,to Flare

Waste water

Figure 2.1: Conceptual Process Design.

Water Recycle gas Purge Product M = Mixer 0= Divider

2.3.

Process Choice

For a first conceptual design of a process, basic choices can be made by following a hierarchy of decisions. This hierarchy is presented by Douglas[601 and developed for the present case in appendix A.

Moreover, a literature study has revealed the existence of experience with these kind of processes described above. This can be taken into account when choosing the process design. It is obvious that the process via methanol has an additional step compared to the first process. This makes this process less attractive for the present case because of the restrictions in available oft-shore area and weight. This is why the two step process, in which natural gas is converted to syngas and subsequently directly to synfuels has been chosen.

The process is presented schematically in figure 2.1.

Purge Oxygen Natural gas Carbondioxide Nitrogen Light Hydrocarbons Carbondioxide Carbonmonoxide Hydrogen Water

Liquid Hydrocarbons /Wax

Waste Water

3.

Assumptions

3.1.

Input-Output Structure

From the process choice made in the previous chapter, an "input-output" structure of the process can be made. It is represented in figure 3.1. This diagram shows the inlet and outlet streams of the process during stationary operation.

3.2. Major Reactions

First Step

In the first step, methane and other hydrocarbons are converted into syngas via two reactions:

• Combustion:

co

+ llrH298

=

-519 kJ/mole (3.1 )

• Reforming:

co

+ llrH298

=

206 kJ/mole

(3.2)

At their exit, reformers operate near the equilibrium of reaction

3.2.

A measure of the

conversion of methane is the approach to equilibrium given as a temperature. The equilibrium constant is linked to the exit gas composition and temperature according to equations 3.3 and 3.4. -!:l

G

In

K

=

r T eq RT

(3

.3

)

(3.4)

ICI Katalco has calculated that an approach to equilibrium of 5 °C is reasonable under the

given circumstances. This means that the actual exit gas composition is the equilibrium

composition at 5 °C below the actual exit gas temperature.

Second Step

After the combustion and reforming of the natural gas, the synthesis of Iinear hydrocarbons

takes place according to the Fischer-Tropsch (FT) mechanism. The FT synthesis can be

reduced to two fundamental reactions:

llrH298

=

-165 kJ/mole

llrH298

=

-41 kJ/mole

1) The hydrogenation of carbon monoxide, equation 3.5, in which -CH2- represents a

segment of a straight paraffin chain.

FVO 3161: Off-share Gas Utilizatian

(3.5) (3.6)

Chapter 3: Assumptians

2)

The water gas shift reaction, equation

3.6,

which takes places most easily on iron

catalysts as a competitive reaction.

From kinetic experiments presented in the Iiteraturel52J.153J it followed th at the rate of conversion of CO can be described by a rate equation which is first order in the hydrogen gas-phase concentration and zero order in the carbon monoxide gas-phase concentration, as described in equation 3.7.

(3.7) The FT reaction occurring over iron, ruthenium or cobalt catalysts is known to lead to products which are highly linear. The distribution of the molecular length can be described as a function of one parameter, the probability of chain growth, a. Since atoms of the intermediate alkyl chain will hardly influence the reactions at the terminal growth center, it is

conceivable that 0: will be independent of the length of the alkyl chain, except when this

chain is very short.

The forminp of the linear hydrocarbons can be assumed as a polymer mechanism. Andersonl29 modified the Schulz-Flory distribution, which describes the forming of polymers, to make clear the synthesis of hydrocarbons of very different length by the FT. The

probability, a, is the single parameter in the Anderson-Schulz-Flory (ASF) distribution. This

mass distribution function is given in equation 3.8.

m"

=

(1 - a)

_a

PaP

3.3.

Plant Location

(3.8)

As mentioned earlier, the unit will be installed on a FPSO. A typical North Sea field was

selected by SBM as a representative example, and this will be the location for the

designed installation.

The area available on the FPSO for the gas conversion unit is 80 x 35 meters. As the unit

will be installed on a floating system, it will have special requirements. Deck load capacity

will be taken into account in the choice of the technology and the equipment, but na detailed calculations will be performed on it. Transportation of the oil product to on-shore

locations is done with tankers. Further transportation of materials from and to this remote

location is difficult and should be kept at a minimum.

3.4.

External Specifications and Conditions

The specifications of the raw materials used and the product made are fixed by respectively the suppliers, the client and the surroundings. These specifications are given

below in table 3.1.

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Purge Gas C02 Water Coo

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3.5.

The Feed

802C,1 bar TVP < 14 psia at 100 2F RVP < 10 psia at 100 2F . to Flare

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to Flare --+

Nu ;

V1er

_{h tc. {'tClre}

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to waste water treatment on the site

1-2 MW available on the site

1

steam available to generate extra power

r

tT

.

Sea Water: Tin

=

15 maximum

=

40 QC

IÎ

P

0 W ('r

c

J

C

lr<..A-cvh

'

-cyc.Lz .

The feed stream, natural gas, is not an individual chemical species but a mixture of several components, with a composition as given in table 3.2.

osition and Conditions of Gas Feed Streams /

/

The feed stream is assumed to be free of impurities such as H2S and heavy metals. Gas

phase impurities and compositional variations of the feed stream will be discussed but not modeled in this design. A typical compositional variation which can be expected in the feed is given by SBM:

±

5 mole% methane,

±

1 mole% propane, and

±

0.5 mole% hexane's.

Likewise, the gas feed is not a constant stream. Due to a flow turndown of the weil, the feed stream to the installation designed in this project will decline as presented in table

3.3.

3.6.

Used Chemicals

5.0 5.0 4.3 2.3 1.0 0.4 {Uv-\l0-..J-

J.)

In this paragraph a list of all chemica~ used in this process

properties will be used in the calculations.

Nitrogen _N2 28.02 1.229 -195.79

Carbondioxide CO2 44.01 0.720 (25) -78.5 (subi)

Methane CH4 16.04 0.4228 -161.5 Ethane _C2H6 30.07 0.5446 -88.6 Propane C3Hs 44.09 0.493 (25) -42.1 n-Butane C4HlO 58.12 0.573 (25) -0.5 n-Pentane CSH'2 72.15 0.6262 (20) 36 n-Hexane C6H,4 86.17 0.6548 (20) 68.7

Carbon monoxide CO 28.01 0.968 (A) -191.5

Hydrogen H2 2 0.06948 (A) -252.87 Oxygen O₂ 32 1.1053 (A) -182.95 Heptane C7H'6 100.2 680 98 Octane CsH,s 114.2 700 126 Paraffin C'2- C,S 170-254 800 < 240 Water 18 1000 100

• (A) related to air; ( ) related to °C.

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3.7.

Availability of the Plant

Chapter 3: Assumptions 29.0 29.0 25.1 13.1 5.6 2.5

is given. Their physical

-210 -56.57 -182.4 -182.8 -187.6 -138.2 -129.7 -95.3 -205 -259.34 -218.79 -91 -57 ±8 0 5000 600 600 100 50 400 300

Since the unit will be operating at sea where no elaborate servicing facilities are at hand, the target is to have a maximum availability of the plant. This is realized by reserving extra

funds to purchase the most durable equipment and the right technology. The ATR

catalyst has a lifetime of 6 years and hence need not be replaced during the exploitation

of one field. The FT atalyst can be re laced on-line so no sh ~ necessary. Major

servicing and maintenance wi e perfo e wtien t e FPSO is moved to the next field.

This leads to the conclusion that the plant should be able to operate for at least 20)years

with an average production time of 360 days per year during this period. ) / l -.:

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3.8.

List of Assumptions

In th is process design, the following assumptions have been made:

• The oxygen plant required for this process is designed and built by an external

contractor. A unit producing 1000 ton oxygen per year costs 30 MMUS$. It has a

power requirement of 20 MW, and dimensions of 26x32m by 30 meters high.

• The oxygen produced consists of 93 % oxygen and 7 % nitrogen and is delivered at

47

oe

and 20 bar.

• Sea water is used as a cooling liquid with an inlet temperature of 15°C. It is assumed that the water can be returned to the sea at a temperature more of 40°C.

• All calculations were performed for steady state operation of the installation.

• Pressure drops and heat losses in pipes are neglected. In the equipment, pressure drops and heat losses have been estimated or calculated where possible. There are no heat losses to the surroundings.

• The feed is free of H2S and metal impurities.

• The composition of the outlet of the autothermal reformer is as given by ICI Katalco. The approach to equilibrium is 5 °C.

• The Selexol® unit removes 90 mole % of the CO2 from the gas stream and is

purchased from Union Carbide. The scrubbing liquid is fully recycied. The unit costs 10 MMUS$.

• The composition of Fischer-Tropsch reactor product has an An.derson-Schulz-Flory distribution.

In the Fischer-Tropsch unit, the filter placed in the liquid recycle stream has not been designed. For its operation, it is assumed that 10% of the inlet passes through the filter. The rest, 90%, is retentate.

• A waste water treatment is present on the FPSO, to bring the water on the required specifications before it is returned into the sea or reinjected in the weil.

• Extra power can be generated from the available steam. The utilities to perform this fall outside the battery limits of this design.

• Utilities for storage and transportation of fuels are present. Storage and transportation of the fresh and spe nt catalysts can be implemented in this system.

• The reformer catalyst is changed during the shut down between the exploitation of two fields.

• The installation can operate during 20 years, 8640 hours per year (360 days/y). During this period of time, 4 identical North Sea fields are exploited, each during 5 years.

Chapter 4: Process Structure and Flawsheet

4.

Process Structure and Flowsheet

In this chapter, a motivation of the choice of the reactors and other unit operations is given. It is ~ased on a Iite~ature study for the tech~i51ues and reaction kinetics for natural gas reformlng and the Fischer-Tropsch process[][ 81• The resulting flowsheet is then described in detail.

4.1.

Feed Pretreatment

Before entering the reforming reactor both natural gas feed streams, coming from the second and first separator respectively, are subjected to a series of pretreatment steps to bring them up to the reformer feed specification. Both feed streams pass through a knockout drum to remove any liquid present. These knockout drums also act as safeguards should Iiquid slugs arrive. Subsequently the streams are compressed to the reformer feed pressure of 30 bar. No interstage cooling is applied because the temperature rise caused by the compression is used as the feed pre heat. The feed streams are mixed and steam is added for dilution of the hydrocarbon stream as to prevent coking in the reformer.

4.2.

Oxygen Production

Oxygen can be separated from air in two different ways. Firstly by cryogenic separation.

This method is suitable for large quantities and is the least energy intensive. However it requires that the feed streams to the main freezer are absolutely hydrocarbon free because any trace of hydrocarbons in a Iiquid oxygen stream makes it highly explosive. In our case it would mean that an extensive air c1eaning operation would be necessary negating all previously mentioned advantages.

The other option is pressure swing adsorption. Here two columns filled with an adsorbent, either a molsieve or activated carbon, are used. Since the adsorption constant differs for the main substances, nitrogen and oxygen, and varies with the operating pressure, separation can be achieved. This process however requires substantial amounts of electricit to com ress the as streams. Still this remains the favorite option because the

éxp osion dangers are limited. It is assumed that this unit will be bought from a specialist

contractor and no calculations will be performed on it (see chapter 3)

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4.3.

From Natural Gas to Synthesis Gas

The conventional way to convert natural gas into syngas is steam reforming, where reactions 4.1 and 4.2 take place. However a steam reformer is a huge device which can not be installed off-shore. The presence of open flames is deemed to great a risk in view of explosion dangers.

co

+ .0.rH298

=

206 kj / mole (4.1 )

co

+ .0.rH298

=

-41 kj / mole (4.2)

An alternative is direct oxidation[111,[181. With dry oxidation, methane can be converted directly into syngas. Several problems occur with the application of this technique. Firstly water acts as an autocatalytic species, which means that slight traces of water in either the gas or oxygen feed streams will promote the formation of more water.

Catalytic Bed

Figure 4.1: Computational Fluid Dynamic Calculation of the Flow Pattern in the

Chapter 4: Process Structure and Flowsheet

The desired reaction 4.3 will not be the only reaction to take place and reactions 4.4 and 4.5 wil! occur as weil.

CH 4

+

1/2 O2 -7 CO

₊

_{2 H2 ~rH298}

=

_{-36 kJ/mole (4.3)}

G H4

+

%02 -7 CO

₊

_{2 H20 ~rH298}

=

_{-519 kJ/mole}

---

_(4.4)

- - -

-CH 4

+

2 O2 -7 CO2

₊

_{2 H2}0 ~rH298

=

-802 kJ/mole (4.5) The result will be that the exit gas stream will have a H2 to CO ratio lower than 2 (approx.

1.3) which is unsuitable for the FT process where a H2/CO ratio of 2.05 is required.

Furthermore any variation in the feed stream composition towards heavier components or any increase in the feed stream will necessitate a higher oxygen feed, which means that surplus capacity has to be provided within the oxygen plant increasing the co st of this unit. Finally, to prevent a high methane slip a higher oxygen partial pressure is required. This promotes the unwanted side reaction 4.5, decreasing the yield of the reactor.

To avoid the oxygen plant capacity problem the use of air instead of oxygen can be considered. The major problem in this situation is the formation of nitrous oxides (NOx).

+ 2 NO ~rH298

=

91 kJ/mole (4.6)

~rH298

=

34 kJ/mole (4.7) At temperatures higher than 900°C the equilibrium of reactions 4.6 and 4.7 lies towards the formation of NOx. Besides being an environmentally unfriendly and toxic chemical when flared these reactions consume part of the oxygen making a higher air feed rate necessary. The use of air also increases the size of all equipment used in the following steps of the process or requires a higher operating pressure, making it more expensive. A third alternative is a combination of these two techniques in one reactor. In an autothermal reformer both combustion and reforming reactions take place. This reactor has the advantage of a good heat integration as weil as areasonabie size[2j,[3j,[12j,[17j,[24j. 4.3.1. The Autothermal Reformer

\

The goal of an autothermal reformer is to convert natural gas into syngas. This is done in two consecutive steps within one reactor. At the top of the reactor natural gas, mixed with steam, enters sideways. This ensures a turbulent flow pattern in this gas stream which in turn promotes good mixing further down in the reactor. The oxygen feed stream, mixed~

with steam, is injected in the burner head. Since the burner head is downward facing and the hot gas will intend to rise upwards, recirculation zones will be established thereby ensuring a homogenous gas mixture where all oxygen will be consumed (see figure 4.1). The design of the burner head is of great importance for a good combustion. Especially bypassing of gas or plugging of the burner head by coke formation should be avoided. Since the enthalpy of reaction increases with a higher carbon number, the heavier natural gas components will be consumed first according to reaction 4.8.

The remainder of the oxygen will be consumed by the methane in the feed stream according to reaction 4.4. Since partial oxygen pressure is low, reaction 4.5 can be neglected. Should the feed stream composition change towards heavier components less

FuelFeed Refraetory Lined wa 11 Inert Beads (filling) Oxygen Briek dome - - - 1 1 7 . (eatalyst support) 6 lil M M

6

M 6 lil Iè Briek dome

(to proteet eatalyst)

Catalytie Bed ICI 54-8 Catalyst

;;;3!

...

_ _ r - -~ Raw Syngas

Chapter 4: Process Structure and Flowsheet

methane will be consumed by the combustion and will enter the catalytic bed. By slightly oversizing the catalytic bed it can be ensured th at the methane slip is as small as possible and any feed stream changes can be processed without a greater oxygen demand.

After the combustion section the gas passes through a brick wall, installed to protect the catalyst from the flames, into the catalytic bed. Here the classic steam reforming reactions take place (reactions 4.1 and 4.2).

The necessary heat for reaction 4.1 is provided by the earlier combustion as opposed to conventional steam reforming where this heat is supplied by burners outside the tube walls. Af ter passing through the catalytic bed and the catalytic support the raw syngas

stream exits the reactor.

A schematic representation is given is figure 4.2. 4.3.2. Reformer Catalyst Selection

All reformer catalysts are basically the same. They use nickel as the active metal on

which the actual reaction takes place[251.[261. As a support different forms of silica or

alumina are used. The shape of the catalyst is determined by the conditions at which the

process takes place.

Conventional steam reforming catalysts, operating at approximately 500 °C, are ring

shaped. This to reduce the pressure drop over the catalyst bed and to make a more

active use of the catalyst since the process is severely diffusion limited. This means that

the rate of diffusion inside the catalyst particIe is much lower than the rate of reaction, so

before the reactants reach the center of the particIe they will have reacted. Therefore only a small part (i.e. the outer layer) of the catalyst particIe is used. The classic way around this problem is to reduce the particle size of the catalyst with the major disadvantage of a

massive increase in pressure drop, which, in most cases, cannot be allowed.

In an autothermal reformer the catalyst will have to be more rugged because of the higher

operating temperatures (1300 °C - 1100 0C). This means that the catalyst particles wil!

have to be denser. The chosen catalyst is the high throughput ICI 54-8 catalyst[261. A

schematic representation is given in figure 4.3. This catalyst consists of a calcium

aluminate support with nickel oxide dispersed on it. Further information is given is table

4.1.

T bi 4 1

a

e

..

ICI 54 8

-

ca alys a at I t d t .

Shape Optimized 4 - Hole

Outer Diameter (mm ) 14

Lenqth (mm) 17

Hole Diameter (mm) 4

Charged Bulk Density

1

kq / m;j ) 850

The advantages of this catalyst are that the effective surface area is increased over the

surface area of the normal solid pellet catalyst by the four holes while retaining its original

strength. These holes also decrease the total pressure drop over the catalytic bed, in

comparison with solid pellets. The sulfur tolerance of this catalyst is approximately 800 ppm.

<1200

oe

1.200

oe

1300

oe

1400

oe

1500

oe

1600 - 2100

oe

Figure 4.4: Computational Fluid Dynamic Calculation of the Temperature Profile in the

Chapter 4: Process Structure and Flowsheet

4.3.3. Operating Conditions

Typical operating pressures for autothermal reformers lie between 20 and 40 bar. The actual operating pressure is a compromise between compressor costs and equipment size. In our case it is mostly determined by the pressure required in the FT process, which is 21 bar. Taking pressure drops in the equipment between the reformer and the FT unit into account and to provide some leeway for special circumstances an operating pressure of 30 bar is chosen. The overall pressure drop in the reformer is estimated at 3 bar.

I

(j...V'

(/VI.

v/

(,tV I j'V\o

The H2/CO ratio is mainly ~ntrolled by the steam to carbon ratio. To obtain a H2/CO ratio

of 2.05 a steam to carbon ratio of 0.9 is required, according to the calculations performed by ICI Katalco. The oxygen to carbon molar ratio is set at 0.6 in accordance with the Iiterature[21•

The temperature profile in the reformer is steep. The hydrocarbon feed stream temperature is not critical but should remain between 200°C and 300°C. The oxygen feed stream temperature is again not critical but should be approximately within 50°C of the hydrocarbon feed stream. Both these temperatures can be regulated by the temperature of the steam which is mixed with the feed streams. The combustion reaction releases a tremendous amount of heat causing a temperature rise of roughly 1000 °c (see figure 4.4). At this temperature the gas enters the catalytic bed where heat will be consumed by the reforming reaction. In this section the temperature drops approximately 200°C bringing the exit gas temperature to 1000 °C.

4.4.

Syngas Cleaning: The Selexol® Unit

The side reaction 4.5 and the water-gas-shift reaction 4.2 can both produce some carbon dioxide. Initially this presents no problems because of the low concentration but at higher concentrations (> 5 mole%), due to accumulation in the FT recycle, the competition between CO2 and CO for active surface sites on the FT catalyst will become noticeable and catalyst activity will decrease. To prevent this from happening the CO2 needs to be removed. Several processes can be applied.

The chemisorption processes using amines are too complicated (2 towers, one for adsorption one for regeneration with steam an.d several other pieces of equipment) to be used on the FPSO. Adsorption systems using molsieves or membranes are not specific enough towards CO2 so some CO will also be absorbed, disturbing the H2/CO ratio.

The least complicated remaining absorption system which is highly specific towards acid gas es (C02 and H2S) is the Selexol® process licensed by Union Carbide. This process uses one column where the acid gas is absorbed by the Selexol® solvent (a mixture of polyethyleneglycol (PEG) and other unspecified components). This mixture is subjected to a high pressure flash where the small amount of hydrogen, which is co-absorbed, is recycled to the feed stream. Subsequently the remaining stream is subjected to a low pressure flash where the CO2 is desorbed. The clean solvent is repressurized and fed to the top of the absorption column. The CO2 is lead to the flare. Since this is a physisorption process it will perform better with decreasing temperature and increasing partial pressure of the component to be removed. This would indicate that the unit should be placed af ter the FT reactors since the CO2 partial pressure will be the highest at this

point. However to remove the low levels of H2S which might pass through the reformer and might destroy the FT catalyst, placement before the FT reactors is the wisest.

The Selexol® unit can be bought from Union Carbide. Since no information, concerning the exact solvent composition or the liquid stream required to remove most of the CO_{2 ,} has arrived until the completion of this project, neither the design nor the calculation of the stream compositions can be performed.

4.5.

From Synthesis Gas to Synthetic Fuels

4.5.1. Choice of the Reactor

There are three reactor systems which are suitable for synthesis of heavy paraffin's. The reactor systems are: a fixed bed reactor, an ebulliating bed reactor and a slurry reactor. The main technical problem of the FT synthesis is the rapid removal of the heat of reaction. This to avoid local overheating of the catalyst, which favors methane formation. Another problem is the uniform distribution of the synthesis gas over the catalyst.

In a fixed bed reactor the removal of heat is very difficult. Regenerating the catalyst is not possible on-line, hence a shut down of the plant is necessary for a month per year. Besides this the capacity of a fixed bed is low, therefore the feed stream in our case wil I lead to the use of more than two reactors, which is not very attractive.

To avoid these problems slurry reactors have been chosen. The most important advantages are:

1) Uniform temperature in the reactor 2) High catalyst and reactor productivity

3) A catalyst activity of about 1, due to the very small particIe size 4) Good heat transfer

5) On-line catalyst refreshing 6) Simple construction 7) Low investment costs

4.5.2. Description of the Slurry Reactor

The reactor consists of a shell with cooling coils in which steam is generated. Syngas is distributed at the bottom by an gas sparger and rises through the slurry that consists of liquid reaction products, predominantly wax, with catalyst particles suspended in it. The reactants diffuse from the gas bubbles through the liquid phase to the suspended catalyst particles where the reaction takes place to produce hydrocarbons and water. The heavy liquid hydrocarbons form part of the slurry phase whereas the lighter gaseous products and water diffuse through the liquid to the gas bubbles. The gaseous products and the unreacted syngas pass through the freeboard and demister above the bed, to the gas outlet.

The gas sparger design needs very much attention, because of the influence of the primary bubble size on the mass-transfer and the possibility of bubble coalescence. Stopping the gas flow, without draining the reactor, can be made possible by installing bubble caps on the sparger to prevent the liquid flowing down.

The catalyst has to be separated from the liquid product. Complete separation is not necessary because the catalyst can be recycled to the reactor, suspended in the liquid phase. For the removal a filter is used. According to the Iiteraturel32] the design of th is filter is

the most complicated part of this reactor. In thjs design it is assumed th at 10 vol% of the recycled liquid stream is the permeate stream. A sm all part of the retentate stream is removed to make it possible to feed fresh catalyst. The removed catalyst cannot be regenerated and af ter removing as much of the hydrocarbons as possible the remaining

4.5 m ...

.

•

_•

•

9' Steam " .Q' iI' ~ --19.5 m

Boiler Feed Water

Product gas out

Demister

Slurry bed

Wax and catalyst to filter

Pump

Filter

~--

-Removal Cat and Wax /

/

, /

Liqu id Prod uct

... _---,-_ __ _ _ Syngas in Gas Sparger

residue has to be burnt to cokes. The choice of temperature and pressure for the reactor are strongly dependent on the selected catalyst and the required product distribution. A schematic representation of the slurry reactor is given in figure 4.5.

Based on values for conventional FT reactors a syngas conversion of 80% is assumed301,[4l1,[491. It is to be expected that this conversion will be higher at lower syngas flowrates because of the longer residence time of the bubbles in the reactor.

4.5.3. Fischer-Tropsch Catalyst Selection 4.5.3.1.Catalytic Material : Iron or Co balt

The water gas shift reaction can be avoided by use of a cobalt catalyst. The conversion of carbon monoxide and the production of hydrogen

by

the water gas shift reaction can be neglected under the following conditions: temperature 473-513 K, pressure 21 bar. Other advantages of the co balt catalyst, besides the reduction of the water gas shift are the lower production costs (according to Jager321 about 30%), a long er catalyst lifetime and higher reactivi~ at lower temperatures. The catalyst usage is about a third of that of a fixed bed reactor 21.

For cobalt catalysts it is weil known that they are more resistant to oxidation than iron based catalysts. Sulfur is the main poison for FT reactions. Shell recommend th at for a cobalt catalyst the feed should essentially be sulfur free. In our case it can be supposed that all sulfur is removed by the Selexol® step.

The use of the cobalt catalyst also determines the required H2/CO ratio, which in th is case is 2.05[421,[501,[531. Because of the neglection of the water gas shift reaction, the H2/CO ratio in

the feed has to be equal to the HiCO ratio in the alkanes produced.

4.5.3.2.Choice of Mean Chain Length

The product composition of the FT reactor is governed by one catalyst parameter : the probability of chain growth, a. Here a

=

0.86 is chosen, because it gives the greatest fraction useful products, like gas oil and an acceptable quantity of waxes. This a is typical for a Cobalt catalyst. Calculation of the ASF-distribution with a

=

0.86 gives these values : Cl -C6

=

21%

C7 - C20

=

57%

C2l - C50

=

22%

An overview of the product range is given in figure 4.6.

4.5.3.3.The Catalyst ParticIe Size and Concentration

The solid particles in the slurry reactor have to be kept in suspension. This defines their maximum size. The influence of small amounts of catalyst particles on the viscosity of the continuous liquid phase is very important. Sep·arability of catalyst particles from the liquid product is also an important factor which has to be taken into account in the choice of particIe size. An optimum between these boundary conditions is found with an average particIe size of 1 00 ~m[4l1,[491. Because of the small si ze of the catalyst particles the effectiveness factor approaches unity and the size of the slur&i' catalyst particIe is no longer a matter of concern from the viewpoint of diffusion limitations[ 1.

In this project a life time of the cobalt catalyst of 1.5 years is assumed and a weight percentage of 15% is chosen[4l 1.

0.4

0.3

a

=

0.86 ~

0.2

0.1

0.0

+---1--r--+----r----.-=:::::=:~==-_

o

10

20

30

P

40

50

60

Figure 4.6: The Anderson-Schulz-Flory Distribution with ex = 0.86.

4.6.

Heat Integration

Power is required for the compressors, pumps and oxygen generation. This power is not available and therefore generated by using the heat produced in the processes. Both the reactors in the process are at exotherm ic conditions and can thus deliver heat used for heating the process streams and the generation of power. An expander and a steam turbine are integrated in the plant. One is combined with the cooling coil in the Fischer-Tropsch reactor. The other uses the heat removed from the syngas stream.

4.6.1. The Steam Turbine

The syngas stream exiting from the autothermal reformer is approximately at 1030 QC. Cooling this stream by means of a heat exchanger gives high pressure steam. This

steam, af ter having been used to heat the streams from the Selexol® unit and the FT

recycle, is led through a steam turbine. The power generated in this way is 8 MW.

A steam turbine is chosen because the main goal is to generate as much power as

possible. Since the ower consuming equipment 0 erates at a s ecific volta e e s

9

turbines have 0 operate at a specl IC telJlperpt!:!.re anJ! Rressme. The stream entering the

s eam fü76l~nè-iSät 40b"ö"C whiêh according to Iiterature[6Tl should be accompanied by a

pressure of 27.6 bar.

4.6.2. The Cooling of the Slurry Reactor

Expanding the steam made in the cooling coil of the slurry reactor produces 8 MW of

power. The stream coming from the expander has to be cooled down further to the

temperature needed in the cooling coil. This takes place in a sea water heat exchanger.

For this stream the first goal is to cool the water stream. It is not opportune to try to

achieve the necessary temperature using solely water heat exchangers. Using a turbo

expander gives a fair cooling and also produces power. This is the reason that an

expander is chosen.

4.7. The Process Flowsheet

The process flowsheet can be found in appendix B.

4.7.1. The Main Stream

The first feed stream, coming from the second separator at 100°C and 1.5 bar, is led to a

vertical knock-out drum, V1, to make sure that no liquid flows into the system. After the knock-out drum the gas stream is led through a compressor, C2, where it is compressed to a pressure of 15 bar, the pressure of the second feed stream, accompanied with temperature rise to 261°C. Next, the second feed stream, coming from the first separator

at 60 °C and 15 bar, is mixed with this stream and led to a horizontal knock-out drum, V3,

to remove any liquids present. After the horizontal knock-out drum the stream is

compressed to 30 bar and 174 °C and again led to a knock-out drum to remove the

possible liquid formed during the compression stage.

The stream is led to the autothermal reformer where it will be converted with oxygen and

steam to syngas and by-products like water and CO2. The product leaves the autothermal

reformer at 27 bar and at a temperature of 1023 °C and is led through a heat exchanger,

HS, and a condenser, H9, where the stream is cooled to 47 °C. The water formed in the

Chapter 4: Process Structure and Flowsheet

condenser will be removed in a knock-out drum V12. After the knock-out drum the gas stream is led to a Selexol® unit T17, where the gas will be cleaned. The water stream from V12 is recycled to autothermal reformer. It is pressurized to 30 bar and turned into steam of 327°C by leading it through two heat exchangers, H11 and H8.

The cleaned gas is led through heat exchanger H18, where it is heated to 220°C and led

0

to the FT section. The syngas stream is mixed with the recycle stream and at a

temperature of 214°C led to the FT reactors R22 and R27. The product with suspended

catalyst is taken from the slurry reactors and with a pump, P20/P28, led through a filtration unit where the catalyst and the product are separated. The final product is cooled with a heat exchanger, H11, to the specification temperature of 80°C. The gas stream from the reactors R22 and R27, consisting of unconverted syngas, light gasses and water, is cooled to 47°C and led to aflasher, V31, where the water is separated from the

gases and the gases are recycled to the FT reactors.

The two FT reactors are cooled with a single cooling system. The cooling water stream from the FT reactors is led to an expander E24, where the stream is expanded from 11 bar and 212°C to 2 bar and 112°C. After the expander the stream is led through a condenser, H25, in which the water is cooled to 77°C and with pump, P26, pumped back to the reactors R22 and R27.

Sea water is pumped up with pump, P35, at 15°C to supply the condensers H33, H25 and H30 with cooling water.

4.7.2. The power generation

The power generation cycle starts with stream 64. This stream is led through condenser H9, where the stream is used as a coolant and heated to a temperature of 638°C. After the condenser the stream is led through two heat exchangers, H18 and H19, where the

stream is cooled to a temperature of 404°C. Af ter this the steam goes through the steam

turbine E24, where power is generated. The stream leaves the expander at 164°C and 2 bar and is led through condenser, H33, where it is condensed and cooled to 35°C. With

pump, P32, it is repressurized to 28.6 bar.

5.

Process Simulation

In this chapter, it is explained how the process has been simulated. This simulation gives

the composition and properties of every stream, which can be found in Appendix C2.

The process has been simulated with ChemCad 3.20 in which heat- and mass flows as weil as temperatures and pressures are simulated. A representation of the flowsheet as inputted in ChemCad can be found in Appendix C1. The thermodynamic model used for the calculations is SRK (Soave-Redlich-Kwong). This model is designed to deal with hydrocarbon streams at moderate to high temperatures and pressures.

The autothermal reformer, the FT reactors, and the Selexol® unit, which are complicated

and specific units, could not be simulated properly in ChemCad. The outlet streams of

these units, calculated outside ChemCad, have been imported in the simulation program, as described below.

Beside performing a process simulation, ChemCad is also capable of calculating the power of pumps, compressors, and expanders.

5.1.

The Autothermal Reformer

The autothermal reformer (ATR) has two reaction sections. In the first one, combustion

reactions take place which are weil known. The second one however, the catalytic

section, could not be simulated in ChemCad. It has therefore been chosen to calculate

the outlet composition of the autothermal reformer outside ChemCad. In ChemCad, the

disappearance of the reactants is simulated with an ideal component separator, 'CSEP' .

The reactants are represented by a product stream leaving the 'CSEP'. In the same way,

the formed products are represented by a feed stream, which is mixed up with the unreacted stream in a mixer unit, 'MIXE'.

As ICI Katalco is a producer of reformer catalysts, we asked this company for help about the simulation of the ATR. ICI could not provide us with a model, which is considered confidential information, but agreed to calculate our case with this model. The composition of the outlet stream of the ATR was thus calculated by the ICI Katalco

material balance program. This was done for two cases, giving a H2/CO ratio of

respectively 1.81 and 2. For our case, the desired H2/CO ratio was set at 2.05, which is

the overall stoichiometric ratio in which H2 and CO react in the Fischer-Tropsch unit. The

parameters for this H2/CO ratio were extrapolated from the two other runs.

It has been assumed that a 5 °C approach to methane-steam equilibrium can be reached,

which corresponds to a methane conversion of 98.8%. The exit temperature is 1023 °C.

The pressure drop has been estimated at 3 bar over the whole reactor height. The outlet composition is calculated by solving the material balance, with the methane conversion of 98.8% mentioned before. The results of this spread sheet calculation can be found in Appendix C3.

5.2.

The Selexol® Unit

As Selexol® is a trade product of Union Carbide, no detailed information is available on

this unit. It is known, however, that a proper Selexol® unit is capable of removing 90% of

the CO₂ present in the process stream. This is simulated with a component separator,

'CSEP', in ChemCad.

Chapter 5: Process Simulation

5.3.

The Fischer-Tropsch Reactor

The Fischer-Tropsch unit, consisting of two parallel slurry reactors, is simulated as one single reactor in ChemCad.

The FT reactor is simulated with a stoichiometrie reactor, 'STOE', in ChemCad. The reactions given in this unit are defined in order to obtain an Anderson- Schulz-Flory (ASF)

distribution with the desired ex of 0.86 at the outlet of the reactor. According to Zhao[50J the

PLoduction of 1 mole hydrocarbons, with no water-gas-shift reaction can be given as :

~

1~

-

I \

19.757 CO

+

40.~14

H2 ->

~

M20

+

-(CH2)", I Cj

~

~ ~~

tl -:. •

~.

!

'>

1

(5.1)---.1

The quantities of every single product can be calculated with the ASF-distribution. The resulting stoichiometrie coefficients are given in Appendix C4. The product leaves the reactor in two separate streams, namely a gas and a liquid stream. In ChemCad, this gas/liquid separation is simulated in a separate unit, placed after the FT reactor.

The cooling of the FT reactor is simulated as a separate loop. Heat exchanger 'HTX'R' nr.

114 in ChemCad represents the cooling coil of the FT reactor. lts heat stream equals the heat of reaction of the Fischer-Tropsch reaction.

5.4.

The Knock-Out Drums

The multiple knock-out drums in the flowsheet have been simulated with 'FLASH' units,

operating at inlet temperature and pressure.

5.5.

Pumps

The pumps have been simulated with 'PUMP' units, in which the outlet pressure was given as an input. The efficiency was set at 75%, which is the default value of ChemCad, in all cases.

5.6.

Compressors

The compressors have been simulated with 'COMP' units, operating polytropically. This has been chosen because it is most realistic for gas mixtures. This is the most

conservative operation mode, i.e. it gives the largest energy requirement. The efficiency

was set at 78% in all cases[66J.

5.7.

Expander and Steam Turbine

Both the expander and the steam turbine have been simulated in ChemCad with expander 'EXP' units. They have been defined as operating adiabatically, as it was

assumed that there is no h:at exchange with the environ.men~. The ~fficienc(s1 was set at

its default value of 75%, whlch corresponds to the value glven In the literature J.

5.8.

Results

The result of the simulation is a report in which the temperature, pressure, size, and

composition of each stream is given. This report can be found in Appendix C2.

6.

Equipment Design

In this chapter, the calculations performed to design the various pieces of equipment are given. Detailed calculations are described in Appendix Di to 04 and the results are given in Appendix 05 and 06.

6.1.

Design of the Reactors

6.1.1. The Autothermal Reformer

It

was attempted to simulate the ATR with Froment's model for reforming[9J_.[19).[20J_{.[22).[23J•}

This was done in a computer program, Matlab. This attempt however failed because of the great non-linearity of the set of differential equations. The probable reason of failure is that the error introduced when making the equations discrete was too large. This resulted in conversions greater than 100% and oscillating parameters. The simulation was therefore given up.

6.1.1.1. The Bed Sizing

The alternative for the sizing of the ATR came from ICI Katalco. ICI's simulation program calculated a required catalyst volume of 11.7 m3 and a bed heightldiameter ratio of 3.9. The installation has to be insensitive for limited compositional variations of the feed. The goal is to convert the feed with a different composition at a constant oxygen feed stream. The variations one can expect are given in chapter 3. The worst case is when the methane mole percentage is 5% lower than in the standard composition. This case has been analyzed below.

Those 5% methane are replaced by heavier components, in proportional quantities. As the heavies have a higher C-mole/mole ratio, this compositional variation results in a higher C-mole feed: around 10% higher. According to the steam/C-mole ratio of 0.9, a 9% higher steam feed is required. These two parameters will contribute to the increase of the required catalytic bed volume, as described below.

The heavies are combusted according to reaction 4.8. The higher C-mole/mole ratio they have, the more oxygen they use. The heavies react first, using a lot of oxygen. The result is more unreacted methane that has to be converted to syngas in the bed. It has been calculated that 10% more unreacted methane is present after the combustion zone.

Therefore the catalytic bed has to be oversized by 10 % to account for this effect, see Appendix Di.

The second point is that a larger total flow goes through the bed due to the higher steam feed. To keep the residence time of the hydrocarbons in the bed constant, a larger bed is required. The total flow is 4 mole% higher, leading to a 4% extra oversizing. A total oversizing of 14% is estimated, resulting in a bed size of 13.3 m3.

6.1.1.2. The Construction Materials According to Coulson and Richardson[59J

, considering the pressure in the vessel, the minimum wall thickness of the ATR should be 3.6 cm, calculated with equation 6.1.