The production of middle distillates from CO2 rich natural gas with the "Shell Middle Distillates Synthesis Process"

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FVONr.

Preliminary Plant Design

Dept. of Chemical Process Technology

Subject

The produetion of middle distillates from CO

₂

rieh

natural gas with the "Shell Middle Distillates

Synthesis Proeess"

Authors

J .A. (Jantien) Baeker

A.I. (Aldo) de Jong

K. (Kavitha) Neelakantan

A.C.J. (Natasja) Romijn

Keywords

TeZephone

0152122756

0152140687

071 5233483

0152133362

Natural gas, Steam reforming, Syngas, Fisher-Tropsch, Heavy paraffins, Hydrocracking, Fractionation, Naphtha, Kerosene and Gas oil.

Date

of

assignment :

13

th

_{of February 1997}

Date

of

report

9

th

of June 1997

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FV03201 Abstract

Abstract

The objective ofthis project was to design a plant to produce 917 900 ton/year of

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cl

7

naphtha, kerosene and diesel from C02 rich natura! gas based on the Shell Middle 0 W Wu,..L . Distillate Synthesis (SMDS) process. The original SMDS process uses natura! gas

with anorma! C02 content. The process comprises of the following steps: 1. Production of synthesis gas

2. Remova! of the excess C02

3. Heavy paraffms synthesis (HPS) by the Fisher-Tropsch reaction 4. Heavy paraffms conversion (HPC) by hydrocracking

5. Product separation by fractionation

When the maxima! use of CO2 is desired it was deduced that steam reforming is the best option for the production of synthesis gas. After desulphurization and partia! CO2 removal the feedgas is fed to a prereformer where light hydrocarbons present in the feed are converted into methane. This feed is mixed with steam to be fed to a steam reforming furnace where production of syngas with a H2/CO ratio of 2.11 takes place at 26 bar and 800°C by a combination of the steam reforming and reverse water-gas shift reactions. The maximum consumption of C02 is one third of the consumption of C14. Therefore just a part ofthe C02 present in the feed is eventually converted into liquid hydrocarbons, while a major part has to be removed. The process can be started up without the need of importing fuel because the natural gas from the weIl is used as fuel itself after desulphurization and partial CO2 removal. Subsequently the syngas is converted into linear hydrocarbons ranging from Cl to C200 in the heavy paraffm .

synthesis. The heavy paraffms (C20+1, are converted into naphtha, kerosene and diesel by hydrocracking and hydrogenation in the HPC where hydrogen is added. The required hydrogen is produced by water-gas shift followed by pressure swing

adsorption. The hydrogenated and hydrocracked hydrocarbons are then fractionated in a distillation column. The bottoms and tops (mainly hydrogen) are recycled to the HPC. The products (naphtha, kerosene and diesel) are of excellent quality with respect to the same produets derived from crude distillation.

They are free of sulphur, aromatic and nitrogen, and consist of mainly linear alkanes. The process has a minimum environmental impact due to the chosen recycle structure

and treatrnent facilities. ~ er

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In this project the process structure is established, the total process is simulated,

:J

energy is reduced by heat integration, the basic design calculations for the equipment are made, the basic process control is designed, a safety, health and environment analysis is done and an economical evaluation is carried out. .

From the economical evaluation it is concluded that in this form the process is not

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economically feasible. Recommendations are made to make the process more

profitabie but despite these recommendations it is not sure whether the process will become feasible. But as oil prices will increase in the future the prospeets for this process are brighter.

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~"'-FV03201 Pre/ace

Preface

The preliminary plant design is an obligatory part ofthe curriculwn ofthe Faculty of Chemical Engineering study at the Delft University ofTechnology. It is a

comprehensive project which should be carried out in the last phase ofthe study by a group of students.

We were glad to receive an assignment from the Shell Research and Technology Centre Amsterdam (SRTCA) in February 1997.

The assignment consists of the design of a plant to produce liquid hydrocarbons from C02 rich natural gas based on the Shell Middle Distillate Synthesis (SMDS) concept.

In this report the main-features ofthe design and the choices ofthe process structure are discussed. FVO 3201, Jantien Backer Aldo de Jong Kavitha Neelakantan Natasja Romijn

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FV03201 Conlents

4.7.1 Heat exchangers 4.7.2 Furnace 4.8 Pumps 4.9 Compressors 5. Design results 5.1 Production 5.2 Utilities 5.3 Product quality 6. Process control 6.1 Norma/operation 6.1.1 Reformer furnace 6.1.2 HPS reactors 6.1.3 HPC reactor 6.104 Flashes 6.1.5 Pumps 6.1.6 Compressors 6.1.7 Heat exchangers 6.1.8 Columns 6.1.9 Purges 6.2 Start-up procedure

7. Safety, health and environment

7.1 HAZOP 7.1.1 Syngas section 7.1.2 Absorber/washer/stripper section 7.1.3 HPS section 7.104 HPC section 7.1.5 Fractionator section 7.2 Complete shutdown

7.3 Environmental aspects ofthe process

7.3.1 General 7.3.2 Surroundings 7.3.3 Products Conlenis 45 46 47 48 49 49 49 50 51 53 53 53 54 54 55 56 56 58 59 60 61 61 62 63 65 65 65 65 66 66 66 66 66 67 67 68 69 69 70 72 74 76 77 77 78 78 78 78

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FV03201

7.3.4 7.3.5 7.3.6

Discharges to the atrnosphere Waste water

Heat and energy integration 8. Economics

8.1 Cost assumptions and estimations

8.2 Production volume dependant operating costs 8.3 Direct operating labour costs

8.4 Investment costs

8.4.1 Zevnik-Buchanan 8.4.2 Wils on

8.4.3 Taylor

8.4.4 Investment costs: summary & conclusion

8.5 Total costs 8.6 Sales income

Total income

8.7 Net profit

8.8 Evaluation ofthe process economics 8.9 Profitability analysis

8.9.1 Return on investment (ROl) 8.9.2 Intemal Rate of Return ORR)

8.9.3 Conclusions drawn from profitability analysis 9. Conclusions and Recommendations

9.1 Conc/usions 9.2 Recommendations 10. Literature 11. Acknowledgements List of symbols Contents 78 79 79 80 80 81 83 83 84 84 84 85 85 86 86 86 87 88 88 88 89 90 90 90 92 94

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FVOfol

Appendices

A Process flowsheet B Mass and Heat Balance

C Stream and Component Balance

Dl Carbon limit

D2 Product distribution

D3 Calculation of conversion coefficients RPS and RPC E Equipment calculations EISteam reformer E2 AbsorberlWasher/Stripper E3 RPS E4 RPC ES Fractionator E6 Flashes

E7 Heat transfer units

E8 Pumps

E9 Compressors F Equipment lists

Fl Equipment list for reactors, columns and vessels F2 Equipment list for heat exchangers and fumaces F3 Equipment list for pumps, blowers and compressors G Specification forms

G 1 Specification form for fractionator

G2a Speci:{ication form for heat exchanger El 09 G2b Specification form for heat exchanger E231 G3 Specification form for pump P 113

G4 Specification form for compressor Kl 04

H Economics

Hl Direct operating labour costs H2 Investment costs

H3 Profitability analysis I Original assignment

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FV03201 Introduction

1. Introduction

Over the years, there has been substantial and sustained growth in proven natural gas reserves. Moreover, the depletion rate of oil is twice that of gas. If this trend

continues, proven gas reserves may soon exceed those of oil. This implies that gas will be of increasing importance as an energy provider in the near future. As many gas fields are in remote areas, the gas needs to be converted into liquefied natural gas (LNG) to be transporled economically to these markets. Another way ofusing the gas is by converting into products such as fertilizers, methanol or liquid hydrocarbons. The Shell Middle Distillate Synthesis (SMDS) Pro ce ss has been developed to convert (remote) natural gas to easily transportabie liquid hydrocarbons. The world's first commercial SMDS plant came on stream at Sarawak, Malaysia in 1993. It produces 12000 barrels/day ofmiddle distillates. The SMDS process comprises the conversion of natura! gas into synthesis gas, subsequently Fischer-Tropsch synthesis yields heavy paraffinicproducts, which are partly cracked into middle distillates. The heavy paraffms can also be converted into high valuewaxes. In the origina! SMDS process the syngas is produced by partial oxidation of methane with pure oxygen.

~

Wh~

~e market prices of oil increasing rapidly, even gas fields of less quality are considered as they may become profitable in the future. A good example of such a remote and fO~ unattractive gas field is the Natuna gas field in Indonesia, which contains up to OO/C of C02. The goal ofthis preliminary plant design is to investigate how to make é lcient use of this low-energy gas and to determine whether this is economically feasible. This implies that certain alterations to the original SMDS-concept are made.

In this project a process is designed to pro duce 917 900 tons per year of naphtha, kerosene and diesel (20 000 barrels/day) from CO2 rich natural gas. The process is simulated in the flowsheeting program AspenPlus.

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-FV0320J Basis of design

2. Basis of design

In this chapter a hierarchical approach described by Douglas (1988) is used to define the assignment and to explain the process choices that were made.

2.1 Definition of the assignmentlhierarchical approach

The assignment for this preliminary plant design is to design a process to convert natural gas from the Natuna field in Indonesia containing high amounts of carbon dioxide into middle distillates. The products are naphtha, kerosene and diesel. The gas composition from the Natuna field (Shell) is given in table 2.1.1. The gas contains also small amounts of hydrogen sulfide, nitrogen and light hydrocarbons which have to be accounted for. A block diagram of this process is given in figure 2.1.1 .

.... C02

,

fuel naohtha proces kerosene C02 rieh ,.

natural gas diesel water

Figure 2. J. J Total process represented by one blo ck

component vol % CHt 26.46 C2H6 0.44 C3Hg 0.14 C4

+

0.13 N2 0.49 C02 71.80 H2S 0.54

Table 2.1.1 Natuna gas composition

Like most processes the process consists of areaction and a separation section. In the reaction section natural gas is converted into middle distillates. Middle distillates and byproducts are separated in the separation section (see figure 2.1.2).

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FV03201 Basis of design

'"

C02

,.

fuel

naDhtha

reaction sep. kp.m~p.nI!

C02rich

natural gas diesel

water

Figure 2.1.2 Process divided into reaction and separation section

For the reaction section several routes from natural gas to middle distillates are possible. Direct conversion methods, like oxidative coupling are not commonly used. In practice two stage processes are found, starting with the production of synthesis gas followed by a synthesis step. This configuration is chosen based on the original SMDS concept.

""

,

C02

I

filel naDhtha

syngas synthesis sep. kp.m~p.nI!

C02 rich

natural gas _diesel

water

Figure 2.1.3 Process divided into basic steps

To pro duce hydrocarbons form synthesis gas Fischer-Tropsch synthesis is the next step in many processes like Sasol and the SMDS process. The byproduct in the FT-reaction is water. It was agreed in the assignment that this synthesis section of the process wiU be left the same as in the original SMDS process. It consists of the so called heavy paraffms synthesis (RPS) and heavy paraffms conversion (HPC), where the heavy paraffins produced will all be converted into rniddle distillates. The HPC is a hydrocracking process and therefore needs hydrogen. When taking a closer look at the HPC it appears that next to the middle distiUates, lighter and heavier hydrocarbons leave the HPC and will have to be separated in the separation section. They leave the process as tops and bottoms (see also next section, figure 2.2.1). To limit the

assignment, the plant wiU not pro duce any waxes nor specialty products. With these constraints the only unit for which the different fundamental routes wiU be evaluated is the syngas reactor. The choice for the way syngas wiIl be produced is discussed in the next section.

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FV0320J Basis of design

2.2 Process alternatives and choices

For the production ófthe syngas the following premises were set:

1. the ratio of H2/CO should be 2, this is required for the Fischer-Tropsch synthesis 2. there should be maximum use of C02

The second premise is set because ofthe high C02 content ofthe feed. We dedicated ourselves to make maximum use of this component rather than just removing it upstream and copying the SMDS process.

Synthesis gas can be produced by partial oxidation, partial combustion, steam reforming and the (reverse) water-gas shift reaction. Combinations of these routes, which are commonly used in practice are discussed bel ow.

• Partial oxidation

Llli = -41 kJ/mol [2.2.1a] This reaction yields immediately the required H2/CO ratio, but none of the carbon dioxide is used. Air could be used to supply the oxygen required for this reaction, but this will yield a large inert strearn of nitrogen which is hardly separabie form other gases. Therefore an air separation unit is needed to pro duce pure oxygen; ho wever a fumace is not needed as the reaction is exothermic.

• Autothermal reforming

With autothermal reforming the heat required for stearn reforming is supplied by the partial combustion reaction. The partial combustion reaction is as follows:

Llli=-519 kj/mol [2.2.1 b]

Part ofthe methane is steam reformed:

Llli=206 kJ/mol [2.2.1 c]

When reaction [2.2.1 c] is carried out with a ratio of 2 to reaction [2.2.1 b] the desired H2/CO ratio is reached. Again an air separation unit is needed and the exothermic reaction eliminates the need of a fumace. An extra difficulty is that the process should be controlled very weIl to ensure· that the reactions occur in the desired ratio.

The overall heat effect of autothermal reforming would be -107 kj/mol. Theoretically this heat could be used to convert CO2 by the endothermic reverse water-gas shift (equation [2.2.1 d]). But, to maintain the desired H2/CO ratio, every CO2 molecule converted must be compensated with 3 times the stearn reforming reaction [2.2.1 c], which means that the C02 consumption will be very low.

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• Steam reforming

To yield the right H2/CO ratio steam reforming [2.2.1c] can be combined with the reverse water-gas shift in which CO2 is consumed and the surplus ofhydrogen is converted.

The reverse water-gas shift reaction is:

MI=44 kJ/mol [2.2.ld]

When reaction [2.2.1 c] and [2.2.1 d] take place in a ratio of 1 to 0.33, the following overall reaction is obtained:

CH. + 0.67 H20 + 0.33 CO2 --t _{1.33 CO + 2.67 H2 llH=221 kj/mol} [2.2.1e]

In

this way the highest achievable amount of carbon dioxide is used.

In

this case an air

separation unit is not required. However the combination of two endothermic reactions demands a furnace to supply sufticient heat for the reactions.

• Mathematical approach

A more mathematical approach to ca1culate the ideal mix is the use of the C,H and 0 balanee is:

C: w+x= 1 H: 4w+2y=4 O:2x+y+2z=1

w>O (methane must be converted) z>=0 (oxygen can't be produced) maxUll1se x

o

.

k

.

Eventually the residual equation to maximise x is: 4x+2z=1 This yields the following values:

w=3/4 x=I/4 y=1/2

z=O

[2.2.lf]

y7-

,_c-x-

1

₁

'

I

'il:.

-+

X'"

'=t \ -t A -+:2.:-~ ., ?, 'j.

'i

X. +.;( '1

:-

-1

This means that oxygen should not be used to maximise the use of CO2. It yields the

same overall equation as [2.2.le).

To match the premises best, steam reforming in combination with the reversed water-gas shift is chosen. The process scheme with the four sections without recycle loops is now defined as shown in figure 2.2.1.

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...

,

H2 C02 fuel ,~ naohtha

steam HPS HPC sep. kerosene

C02rich refonner

natural gas diesel

bottoms ... _,

steam water

Figure 2.2.1 Process divided into functiona/ sections

The next step is to integrate certain streams and to define the recycle structure between the sections. The water produced in the HPS section can be boiled up

partially as steam for the syngas section. The bottoms are recycled to the HPC section. Because the partial pressure of H2 must be high in the HPC section, a large recycle stream of H2 should be applied. The light gases from the HPS and the H2/tops purge will be used as fue! for the refonner furnace. The alternative to recycle the tops to the prereformer is rejected as it would be a loss of energy to convert the tops to methane, subsequently to syngas and then to hydrocarbons which are cracked and eventually partly end up as tops again. Also the build-up ofinerts can be kept lower. These choices result in the following block diagram (figure 2.2.2)

.-C02 H2 fuel recyde H2 purge H2Itops te fuel steam HPS HPC sep. C02 rich reformer

natural gas '--.,---' diesel

ter eam

water

Figure 2.2.2 Process divided into functiona/ sections with recycle structure

Now we will de fine of what units each section is composed. The syngas section consists of a prereformer, a reformer, a flash and an absorber.

In the prereformer C2+ hydrocarbons are converted into methane to prevent carbon deposition in the steam reformer. Another measure to prevent this is to set a minimum ratio of C02 to C~ of 1 to operate at the right side of the carbon limit (Dibbern,

1986). This leads to a steam to methane ratio of 2.6 on mole basis due to equilibrium considerations. Part of the C02 originally present in the feed needs to be separated

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because the original ratio of CO2 to C~ is 3rather than 1. Also the small amount of H2S in the feed should be removed before entering the catalysed processes because it poisons the catalysts. The desulphurizer/C02 removal section is agreed to be

considered outside the boundary limit.

The excess of steam fed to the refonner is separated as water in the flash after the reformer. Part of it is boiled up to steam to be recycled to the refonner. It is recycled to the reformer instead of to prerefonner because in the prereformer not such high amounts of steam are necessary. Another part leaves the process as waste water. This stream is of the same magnitude of the water which is net produced in the process. Because C02 and C!4 react in a ratio of 0.3 to 1, the excess CO2 has to be separated as well. This is do ne by an absorber in which MBA (monoethanol amine) is used to absorb the C02. The CO2 is subsequently stripped offrom the MEA in a stripper. The HPS section consists of reactors in which mainly linear hydrocarbons (69% alkanes, 30% alkenes and 1 % oxygenates) and water is produced. Because the conversion is not complete, the unreacted syngas, light hydrocarbons and water are flashed offfrom the heavy hydrocarbons (C20+) in a fITst flash after the reactor. The reason why first the heavy hydrocarbons are separated from the rest is that they would congeal if they would be cooled further down. III a second three-phase flash the light hydrocarbons are separated from the water and the syngas. The syngas is recycled to the reactor; to avoid build-up of tops (C

_s-)

in this recycle, part ofthe gas is purged to use as fuel gas. The water leaving the three-phase flash is led to the prereformer to fulfill the steam requirement. The light hydrocarbons (C20-) bypass the HPC to the fractionator because they don't need to be converted into lighter products. An

important assumption which justifies this choice is that the alkenes don't influence the quality of the end products too much. This assumption is founded on the fact that in the original SMDS process, according to our information, the same is done.

In the next step, the HPC, the heavy hydrocarbons are hydrogenated and hydrocracked into lighter alkanes by use ofhydrogen. This hydrogen is produced from syngas by shift reactors in which the water-gas shift reaction takes place and pressure swing absorption is used to separate the hydrogen from the rest of the gases. The section which should carry out this conversion and separation is agreed to be considered off-limit too.

To obtain a satisfactory conversion, hydrogen should be in large excess to the amount which would be needed stoichiometrically. This excess ofhydrogen could be

separated from the products by means of a flash. A disadvantage of such operation is that a large part of the hydrocarbons which should leave the process as naphtha is built up and eventually purged as fuel unless cryogenic separation is used. Therefore it is decided to let the fractionator, which is designated to yield a better separation, separate the hydrogen to be recycled to the HPC.

In the fractionator the light hydrocarbons from the HPS and the products from the HPC are fractionated in the following products: H2/tops, water, naphtha, kerosene, diesel and bottoms. The water mainly originates from the stripping steam applied in this unit. The H2/tops are recycled to the HPC and partly purged as fuel.

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The amount of fuel from the syngas purge and the H2/tops purge is not enough to meet the heat requirements ofthe reformer fumace. Therefore an extra flow ofthe feed gas is mixed with these streams to obtain enough fuel gas. This feed gas is split from the main stream after the desuiphuriser/C02 removal section. This choice is made for two reasons. Firstly, the original natural gas has a heating value which is insufficient to reach the high temperature in the fumace. Secondly the upstream removal of H2S means a cleaner flue gas and less flue-gas treating.

All these choices and considerations result in the simplified flow sheet as dr~wn in figure 2.2.3. ...

·

. ·

.

·

.

· -

.

,----~~!'SA 1---7-+--,

Figure 2.2.3 Simplifiedjlow sheet with allfunctional units and recycle structure

2.3 Reactions

In this section the most important reactions that occur in the different reactors are glven.

2.3.1 Prereformer

In the prereformer the light hydrocarbons (C2+) are converted into methane and

carbon dioxide:

CnHm + 1/4 (4n-m) H20 ~ 1/8 (4n+m) CRt + 1/8 (4n-m) CO2 i1H>O [2.3.la]

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2.3.2 Steam reformer

As mentioned before steam reforming of methane occurs in a combination with the reversed water-gas shift:

C~ + H20 -+ CO + 3 H2 C02 + H2 -+ CO + H20 C~ + C02 -+ 2 CO + 2 H2 &1=206 kJ/mol &1=44 kJ/mol &1=250 kJ/mol [2.3.2a] [2.3.2b] [2.3.2c] The C02 reforming equation [2.3.2c] is commonly mentioned in the literature as a separate reaction, whereas it in fact made up ofthe steam reforming reaction [2.3.2a] and the reversed water-gas shift equation [2.3.2b].

2.3.3 Heavy paraffin synthesis (HPS)

_PI

T

· _.I6.'··lR'

' J e )

1

In the HPS syngas is converted into alkanes, olefins, water, oxygenates and carbon dioxide according to the following reactions:

(2n+ 1) H2

+

n CO -+ CnH2n+2 (alkanes)

+

n H20 68% [2.3.3a]

30% [2.3.3b]

1% [2.3.3c]

(n+ 1) H2

+

2n CO -+ CnH2n+2 (alkanes)

+

n C02 1% [2.3.3d] The percentages refer to what extent the reactions occur and were obtained from Shell. The reactions are all exothermic. Other reactions which may occur are agreed to be neglected because they involve similar products but occur to an even smaller extent then the above mentioned reactions.

2.3.4 Heavy paraffin conversion (HPC)

In the HPC the olefins and oxygenates are hydrogenated to alkanes and hydrocracking takes place according to the following reactions:

[2.3.4a] [2.3.4b] [2.3.4c] Hydroisomerisation occurs as weIl, but can be neglected as the selectivity 0 . -paraffins is very low at high operating temperature .. The reactions are al exothennic.

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2.4 Battery limit

For the SMDS process for C02 rich natural gas the following equipment is taken into account:

Prereformer Steam refonner

C02 absorber (inci. washer) MEA stripper

HPS (Fischer-Tropsch reactor) HPC (Hydrocracker)

Fractionator

Heat transfer equipment

Pumps, turbines and compressors Flash vessels

The equipment considered to be off-limit is: Desulphurization and CO2 removal section Waste water treatment

ShiftlPSA section Flue gas treating Utility equipment Storage facilities

Equipment necessary for start-up/calamities Engineering and maintenance workshops Auxiliary buildings (offices, canteens etc.)

Transportldistribution areas (railroads, truck load-off areas etc.)

With these considerations taken into account the conditions and compositions in weight percentage ofthe streams fed to the process are as given in table 2.4.1.

Water MEA-sol. Natural gas Hydrogen steam

(101) (102) (103)

*

(207) (326) P (bar) I I 5 40 3 T (OC) 25 25 25 250 190 Cf4 25.5 % C2H6 0.8% C3Hg 0.4 % C4HIO 0.4 % N2 0.8% CO2 72.1 % H2S 0 H2 98% CO 2% H20 100% 98.9% 100% MEA 1.1 %

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Referring to table 2.4.1 on the previous page:

*

This composition is different from the composition defrned by the assignment because the desulphurisation/C02 removal section is considered off-limit

All product streams leave the pro ce ss at an elevated pressure (> 1.4 bar) to be stored in tanks. The temperature of the product streams should be lower than 40°C. The required boiling ranges and carbon number cut points ofthe products (Scherzer, 1990) are given in table 2.4.2.

Naphtha Kerosene Diesel

low cut point (OC) 100 140 290

high cut point (OC) 140 290 370

low cut point (C no.) Cs C9 C16

high cut point (C no.) Cs CIS C23

Table 2.4.2 Required boiling range and carbon number cut points products

Properties ofkerosene and diesellike cetane-number, pour point, freezing point and smoke point should be investigated to establish a view ofthe product quality.

2.5 Utilities

Concerning the utilities required for operation, the following assumptions are made with respect to the availability.

Standard high, medium and low pressure steam produced by steam production units are:

Steam HP MP LP

TCOC) 410 220 190

P (bar) 40 10 3

Tsat (0C) 250 180 133.5

Table 2.5.1 Steam utility

Cooling water is available at a temperature of 20°C. The pressure is elevated to 3 bar for transportation.

Cooling and combustion air is available at a temperature of25 °C and at an elevated pressure of2 bar.

Electricity is available at a voltage of380 V; it probably needs to be generated by turbines, as the plant is situated at a remote gas field. This is however considered to be off-limit and the standard prices are taken for electricity.

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FV0320J

2.6 Life time

In the plant four units use catalysts: Ni-cat in prerefonner

NilMgAh04 in the steam refonner

Co-based Shell proprietary catalyst in HPS Dual functional Shell proprietary catalyst in HPC

Basis of design

The average life time of all three catalysts is assumed to be 5 years. The economical plant life is assumed to be 15 years.

2.7 Location

According to the assignment the plant is located near the Natuna gas field. It is assumed to be situated on one ofthe relative remote Natuna isiands in the South Chinese sea (figure 2.7.1). It is assumed that there will be no facilities already available at all and there is no infrastructure or other industry.

Figure 2.7. I Location of Natuna gas fields

2.8 Market demand

Due to the rapid economic growth and a large population the energy consumption in Asia is predicted to increase by an average of3.5%/year between 1995 and 2015 (Tijm, 1995). The consumption of naphtha is forecasted to rise. By the year 2000 huge naphtha deficits are expected (Asia pacific fuel oil, 1997). Asian nations that are short of methane find naphtha a useful feedstock for the production of urea and ammonia. Naphtha is also a feed for the production of ethylene. East Asia's proportion ofworld-wide ethyl ene demand is predicted to increase by 2000. The future demand rate for

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FV0320J Basis of design

kerosene is predicted at 4.5%/year (Tijm, 1995). The world market for diesel is 370 million tonnes per year.

In the real situation (speaking 1997) plans are being made in Indonesia for the Natuna LNG project (Greenhouse issues, Jan 1996). Construction activities are scheduled to start in 1998/1999 and the LNG will come on stream by the year 2004/2005. The gas will be produced and treated offshore and subsequently liquefied onshore. Treatment would involve separation of the CO2 and other gases from the methane. The potential buyers ofthe LNG are identified to be Japan, South Korea and Taiwan. This implies that there is a ready market for the methane itself.

The above forecasts indicate that there is an increasing demand for the middle distillates in Asia. The SMDS products could be distributed internally in Indonesia and exported to other countries. Due to the excellent properties it can be sold for a higher than normal price to be blended with lower grade diesel fuels to meet new strict environmental regulations. But it is important to point out that there is already a good market for the methane itself. Incr,ease of the oil price and more rigorous

environmental and fuel quality demands will make the products more profitabie.

,

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.-

~

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Name Struc.Formula molwt b.p. m.p. IIq denslty MAC MAC Expl IImlts

g/mol °C °C rel (water) ppm mg/m3 vol %

methane CH4 16 -162 -182 - nv nv 4.4-16 ethane C2H6 30.1 -89 -183 0.4 nv nv 2.7-12.5 propane C3H8 44 -42 -187 0.5 nv nv 1.7-9.5 butane C4H10 58.1 -0.5 -13.8 0.58 600 1430 1.3-8.5 nitrogen N2 28 -196 -210

-

nv nv

-hydrogen sulfide H2S 34.1 -60 -68 0.8 10 15 4.0-46 carbon dioxide C02 44 -79 - 0.8 500 9000

-water H20 18 100 0.997 1 nv nv

-hydrogen H2 2 -253

-

nv nv apr-76

carbon monoxide CO 28 -191 -205

-

25 29 nov-75

pentane C5H12 72.2 36 -130 0.6 600 1800 1.4-8 1-pentene C5H10 70.2 30 -138 0.6 1-pentanol C5H110H 88.2 137 -79 0.8 hexane C6H14 86.2 69 -95 0.7 25 90 1.1-7.5 1-hexene C6H12 84.2. 63 -140 0.7

--

1-hexanol C6H130H 102.2 158 -47 0.8 ~ heptane C7H16 100.2 98 -91 0.7 300 1200 1.0-7.0 1-heptene C7H14 98.2 94 -119 0.7 1-heptanol C7H150H 116.2 176 -34 0.8 octane C8H18 114.2 126 -57 0.7 300 1450 0.8-6.5 1-octene C8H16 112.2 121 -102 0.7 1-octanol C8H170H 130.2 194 -17 0.8 kerosene (C9-C 15) C11H24 156.2 196 -25.6 0.7 nv nv 0.7-6.5 "-ene C11H22 154.2 193 -49.2 0.8 "-ol C11H230H 172.2 243 19 0.8 diesel (C16-C23) C18H38 254.2 316 28.2 0.8 nv nv 0.6-6.5 "-ene C18H36 252.2 313 17.5 0.8 "-ol C18H370H 270.2 210 59 0.8 bot!oms (C24+) C30H62 422.2 450 65.8 0.8 "-ene C30H60 420.2 448 55 0.8 "-ol C30H610H 438.2

-

88 0.8 MEA H2NC2H40H 61.08 170 10.3 1.0 ammonia NH3 17.2 -33 -78 0.8 25 I Ign. vap. energy denslty mJ rel. (air) 0.28 0.6 0.24 1.04 0.25 1.6 0.25 2.01

-

0.97 -0.07 1.2

-

1.5

---

-0.01 0.07 -0.1 0.97 0.22 2.5 0.24 3 0.24 3.45

-

3.9 5

-

7 autolgn. temp °C 537 515 470 365 260 605 260 1 223 204 210 220 >220 N

\0

~

...

fIl f0t-O

,...,

n o

e

"'Cl

o

=

~

=

~ 'l')

è5

I.., I-..) <:::>

....

ll:l

a

t;j. ~

~

OQ' ::s

(22)

FV0320J Process structure

3. Process structure

3.1 Process description

In this section the total process is described. The process flowsheet (appendix A)

together with this section should give a clear view ofthe process

• Prereformer

The natural gas from the desulphurisation/absorption is preheated to 500°C in the convective section of the steam reforming furnace. After preheat C2+ hydrocarbons are converted with steam into methane in the prereformer to prevent carbon deposition in the steam reformer. A very tiny part «1%) ofthe nitrogen, which is for less than one percent present in the feed, is converted into ammonia. This is no problem for the rest ofthe process because it leaves the process with the water at the fust flash or with the washing water in the washer. The nitrogen leaves the process with the syngas

purge and should therefore be no problem too .

• ' Steam reforming furnace

The stream is preheated again in the convective section ofthe furnace to 500°C

before entering the reformer furnace. In this reactor, which consists of multiple

catalyst tubes inserted in a furnace, CH4 and CO2 are converted to synthesis gas (H2 and CO) by steam. The furnace is fired with a mix of feed-gas, and the syngas and H2/tops purge streams. The temperature is almost linear increased along the tubes. The stream leaving the furnace at 800°C is first led to a thermosyphon reboiler to generate steam for the prereformer. After that it is further cooled by generating steam for the reformer in the next thermosyphon reboiler. Still enough heat is left to warm up the boiling feed water for the HPS section. After this last heat exchanger the water is separated from the raw synthesis gas by a flash at 75°C and 25.5 bar. This water is for 83 % fed to the above mentioned reboiler, the rest is purged to the waste water

treatment.

• Absorber/washer and stripper

The raw syngas is contacted with a MEA-solution in an absorber where the CO2 is absorbed. To remove traces ofMEA in the syngas, a washer section is placed on the

absorber, where washing water is flown in countercurrent flow with the gas. The CO2

is subsequently stripped of from the MEA in a stripper and led to an off-gas treating

unit. After this section 2% ofthe syngas is split offto the shiftlPSA section where H₂

is produced. Assumed that in the shift reactor the molar ratio of H2/CO is elevated at

approximately 9, that in the pressure swing adsorber the hydrogen is recovered for 90% with a purity of 99.9 mol%, the amount of pure hydrogen produced is enough to

compensate for the use in the HPC and the 10ss through the H2/tops purge.

• HPS

After some heat exchangers to preheat the syngas it is fed to the first HPS (heavy paraffin synthesis) reactor. In this reactor mainly linear hydrocarbons (69% alkanes,

(23)

FV03201 Process structure

remove the heat produced boiling water is flowing countercurrently with the stream in the reactor tubes. Because the conversion is not complete, the unreacted syngas, light hydrocarbons and water are flashed offfrom the heavy hydrocarbons (C20+) in a flTst flash at 150°C and 22.4 bar after heat exchange with the feed of the RPs. The heavy hydrocarbons are led to the HPC. The gas stream is further cooled to 30°C. In a second three-phase flash the light hYdrocarbons are separated from the water and the syngas. The unreacted syngas form the flTst reactor is fed to the second HPS reactor where at 250

oe

and 22.3 bar the same reactions occur. After this reactor the same heat exchangers and flashes are instalied is after the flTst. The syngas is recycled to the flTst reactor, the water to the prerefonner and the light hydrocarbons (C20-) bypass the HPC to the fractionator. To avoid build-up of tops (C

_s-)

in the syngas recycle, 10% of the gas is purged to use as fuel gas .

• HPC

In the next step, the HPC, the heavy hydrocarbons are hydrogenated and hydrocracked into lighter alkanes by using hydrogen in a hydrogen quenched multibed reactor at 250 °C, 40 bar. To obtain a satisfactory conversion, hydrogen is in large excess to the amount which would be needed stoichiometrically. The stream leaving the HPC has a pressure of37 bar. After a knock-out drum to separate the gas stream from the liquid the pressure of both streams is relieved to 1.8 bar in two turbines where electricity is generated before entering the fractionator. The vapour stream is preheated before it is mixed with the evaporated light hydrocarbons from the HPS section, at 320°C and 1.8 bar. The liquid stream is without further heating (299°C) introduced in the fractionator.

• Fractionator

In the fractionator the light hydrocarbons from the RPS and the produets from the HPC are fractionated in the following products: H2/tops to fuel, water, naphtha, kerosene, diesel and bottoms. The water mainly originates from the stripping steam applied in this unit. The pressure ranges from 1.8 bar at the bottom to 1.2 bar at the top, while the temperature is 305°C at the bottom and 40°C at the top. A three-phase flash at the top separates the condensor water, H2/tops and naphtha. The kerosene and dieselleave the fractionator via the bottom stream of two sidestreamstrippers. The gas streams containing lighter hydrocarbons leave these strippers at the top and are fed back to the fractionator. The water stream is sent to the waste water treatment. The bottoms and the H2/tops are recycled to the HPC after respectively a pump and a compressor has brought the pressure back to 40 bar.

(24)

FV0320/ Process structure

3.2 Unit operations and process conditions

3.2.1 Syngas reactors

In the prereformer the ethane, propane and butane, present in small percentages in the feedgas, are reformed with steam to methane and carbon dioxide, according to

reaction equation [2.3.1a] . The reactor is an adiabatic fixed bed reactor. The feed is preheated to 500°C and leaves the reactor at 463°C. The inlet pressure is chosen to be 26.6 bar to achieve a pressure of 25 bar in the HPS reactor downstream.

In the main reformer steam reforming takes place at temperatures ranging from 500 to 800°C. In this reactor the reactions [2.3.2a], [2.3.2b] and [2.3.2c] take place catalysed by a NilMgAh04 catalyst. The reactor is a gas fired furnace in which the tubes filled with catalyst particles are inserted. The pressure at the inlet is 26.1 bar. This pressure is chosen to achieve the desired pressure of 25 bar in the HPS reactor downstream. With steam reforming it is important to prevent carbon formation. Therefore the CO2 to CHt and B20 to CHt ratios should be high enough to operate at the right side of the carbon limit (Dibbem, 1986, appendix D2). With a higher pressure, the chance of carbon formation is even lower (Rostrup-Nielsen, 1984). An artist's impression of a reforming furnace is given in figure 3.2.1.

Figure 3.2.1.1 ReJarmer Jurnace. Tube and burner arrangement. Taps@e design. (Rastrup-Nielsen /984)

The steam reforming is simulated using a Simulink programme to design the reactor (outlined in section 4.1). The results ofthis prograrnme are implemented in the Aspen simulation by use of a stoichiometric reactor.

(25)

3.2.2 AbsorberlWasher and Stripper

In the absorber, as much C02 as possible is removed from the gas stream to minimize equipment volumes and to prevent accumulation of C02 in the process. In spite of the fact that the RPS is a CO2 producing unit, the absorber is placed upstream ofthe RPS section. As the HPS produces only a small amount of C02, the advantages of the reduced reactorvolume are considered to have more impact than the small increase of C02 accumulation in the recycle.

• Process description

CO2-rich gas enters the bottom ofthe trayed absorber. A lean amine solution is introduced at the top tray of the absorber section and moves down the column. Contact between the gas and the amine liquid on the trays results in C02 being

absorbed into the amine. As amines poison the downstream HPC catalyst, the lean gas is led to a washer-section to remove any entrained amine before leaving the top. Here, the lean gas from the top absorber tray is washed with water to ensure that no MEA enters the process. C02-rich amine leaves the bottom ofthe absorber and is led to a stripping column. The stripper is reboiled with saturated steam; no extra water is added but the water in the C02-rich amine stream is evaporated (thermosyphon reboiler) and used as stripping stearn. The CO2 is stripped off and the lean amine is cooled and sent back to the absorber.

• Monoethanol amine

There are several amine-based solvents in commercial use for the removal of COz. All these solvents depend upon their amino nitrogen group to react with the acidic COz in performing their absorption. Most commonly used are monoethanol amine (MEA) and diethanol amine (DEA), which have as main difference their degree of reactivity. MEA is the most reactive one and can remove nearly all the C02 in the gas stream. The process is weIl proven in refmery operations.

MEA is considered to be a chemically stable compound, but does oxidize when exposed to air. Storage and surge tanks must therefore be provided with inert blanket gases as N2 or 'sweet' natural gas to avoid this degradation. Also, high temperatures promote amine degradation, therefore saturated stearn instead of superheated steam is used as stripping medium.

A disadvantage of MEA compared to DEA is that it has a higher vapour pressure (30 times that ofDEA) which causes high MEA losses.

absorber

According to the literature found (lones, 1991), a MEAlCOz ratio varying from 2.4 to 3 is used in order to remove all CO2 from the feed gas. For the recirculation solution, the weight% of MEA in water normally is between 15 and 20%.

washer

The literature (Jones, 1991) reveals that the water wash rate in the washer-section should be about 25% ofthe amine circulation rate. The effluent water, which contains the absorbed MEA, should be steam stripped befor disposal to the effluent system.

(26)

stripper

The CO2-rich amine stream contains a lot water, which can be evaporated and used to strip the CO2 from the MEA. As this is enough water to strip off all the CO2, no extra steam is needed. A thennosyphon reboiler is used to evaporate the water. Saturated steam is used to avoid excessive temperatures which lead to amine deterioration. As MEA has quite a high vapour pressure, MEA loss due to MEA leaving the top of the stripper is unevitable. This loss is compensated with a make-up stream ofMEA solution.

• Simulation

In Aspen, the absorber and the washer are simulated as two separate columns. In

reality, the washer section is placed directly on top ofthe absorber section. CO2-free gas leaving the topsection of the absorber is directly led to the bottom of the washer-section.

The ratio kmol MEAlkmol C02 was set at 2.88; a solution of 16.7 weight% MEA in

water was used. In the absorber as weIl as in the washer and stripper, the gas feed

enters the column at the bottom tray and leaves the column at the top tray. The liquid feed on the other hand enters at the top tray and leaves at the bottom tray.

For the simulation in Aspen, the following thennodynamic models were used:

option set K-value method motivation

absorber ammes Kent-Eisenber amines used for system C02 in MEA

model solution

Table 3.2.2.1 Aspen option set

option liquid phase activity vapor phase fugacity motivation

set coeff method coeff method

washer elecnrtl electrolyte NRTL Redlich-K wong used for low conc.

. stripper amines and acid gas

stripping from ammes

Table 3.2.2.2 Aspen option sets

• Results and conclusion simulation A WS

From the gas feed, 99.0% ofthe CO2 is removed. A spore of MEA (0.002%) leaves

the absorber at the top and is fully removed in the washer section. In the stripper,

99.6% ofthe MEA is recycled to the absorber. The weight% MEA in this stream is 17.7%, this is compensated by the MEA solution make-up stream.

The process is not yet optimalised:

- the ratio MEAlC02 in the absorber can be decreased to a minimum of2.4 - a condensor can be placed to prevent MEA leaving over the top ofthe stripper

(27)

3.2.3 Heavy paraffin synthesis (HPS)

The syngas produced in the steam refonner wiIl be synthesized via the Fischer-Tropsch reaction. This production step to heavy linear paraffms wiIl be based on the existing reactors at Bintulu, i.e. the same catalyst will be used under the same process conditions yielding the same product distribution. Therefore the most important objective of the simulation is to imitate the original process as realistically as possible. This means the reaction wiIl take place in multitubular fixed bed reactors at a pressure of25 bar and a temperature of250 °C catalyzed by the Shell proprietary catalyst. An overall conversion of 40 % is assumed, based on Anderson (1984).

The catalyst used is cobalt based on a silica carrier. The hydrogen sulfide in the feed gas is removed upstream, because of the poisonous effect on this type of catalyst: 1.124 Co + H2S ~ 1.124 COSO.89 + H2 [3.2.3a] with

K

= 2.8*106 at 500

K

(Anderson, 1984).

The ammonia produced in the prerefonner is on the contrary non-poisonous to Co-catalysts, as the equilibrium [3.2.3b] is far on the left side:

[3.2.3b] with K = 1.2*10,5 at 500 K (Anderson, 1984); this effect is thereföre assumed to be

negligible. • Reactions

In the synthesis step a number of catalyzed reactions take place simultaneously. The most important which wiIl be simulated, are:

(2n + 1) H2 + n CO ~ CnH2n+2 (al kanes) + n H20 [3.2.3c] [3.2.3d] [3.2.3e] (n + 1) H2 + 2n CO ~ CnH2n+2 (alkanes) + n CO2 [3.2.3f] Side reactions which wiIl not be taken into account, are:

[3.2.3g] [3.2.3h] [3.2.3i]

(28)

FV03201 Process struclure

By these reactions a wide range ofheavy linear paraffins is produced. To avoid

simu1ation of all these different reactions~ groups of carbon numbers are 1umped and

the median on weight bases is chosen as representative(see Table 3.2.3.1). The tops and naphtha groups are not lumped to be able to simulate the downstream flash vessels realistically.

Group Range representative

Tops CI- C4 Cl C2 C3 C4 Naphtha Cs - Cs Cs C6 C7 Cs

Kerosene C9 - CI5 CII

Diesel _{CI6 - C23} CIS

Bottoms C24 - C200 C30

Table 3.2.3.1 Lumping of product range

The selectivities ofreactions [3.2.3c-f] on CO, as obtained from Shell, are

respectively 68%, 30%, 1 % and 1 %, which are simulated for the carbon numbers above 5. This yields the overall reaction of:

(1.99n

+

0.69) H₂

+

1.01n CO ~

0.69 _CnH_2n+₂

+

0.30 C_nB_2n

+

0.01 _CnH_2n+_IO

+

O.Oln C02

+

(0.99n - 0.01) B20

forn=5,6, 7,8,11,18,30 [3.1.1j]

For the carbon numbers 1 to 4 only the production of al kanes (reaction [3.2.3c]) is

taken into account:

for n = 1, 2, 3, 4 [3.1.1k]

• Productdbtribution

The product distribution ofthe Bintulu process, shown in figure 3.2.3.1 (Eiiers, 1990), is taken as premises for the simulation ofthe reaction. From this plot a product

(29)

.. : FV03201 10 8 6 PRODUCT CQMPOSITION. %w IIIr.II··SEVenITY IIYllnOCnAGKmr. MEflIUM··SF.VEnll Y IIYDROCnACKING .-" FISCIIE 11 .. 1 norS!:II ~/rn()Il11CT ~

..

...

..

_...

..~

•..

_....

--

...

---o LLC.J.LI..-'-''-'-.L...L....L-I-L-L-J...l. .. ::....<..:,;· L=.' .L.L..1 . .I o 10 20 30 ~O CIInOON NUI.mrn

. Û1,onn·dL<I,ihulinn nra FL<ehc,.Trup<eh pruducl ~dll'c :oud :o/ll'l<ckcli,·c h)'dro<"lackill~ .

Figure 3.2,3.1 Product distribution (Ei/ers, 1990)

Process structure

However, this figure only shows the distribution of Cs and up. The amount oflighter organics produced can be derived as follows. The Schulz-Flory constant for a Fischer-Tropsch catalyst is a measure for the chance on chain growth; the larger this alpha-value is, the heavier the product distribution will be (see Table 3.2.3.2, (Eiiers, 1990». From the rough product distributions of catalysts with different alpha-values (see Table 3.2.3.2), the alpha-value ofthe Bintulu-distribution is determined by comparing the ratios of weight percentages of heavy paraffms (>C20) on the weight percentages of middle distillates _{(C IO-C20).}This results in an alpha-value of 0.90 for the used catalyst. Now the weight percentage ofthe carbon numbers bel ow 4 can be determined as 17 wt%.

Product distribution (wt%) Ratio

r>C20 lIrCIO-20]

Growth chance a < CIO CIO - C20 > C20

0.80 62.4 31.8 5.8 0.18 0.85 45.6 38.9 15.5 0.40 0.90 26.4 37.1 36.5 0.98 0.95 8.6 19.8 71.7 3.62 0.98 1.6 4.9 93.5 19.08 0.99 0.4 1.4 98.2 70.14 Bintulu 11.3 47.2 41.5 0.88 Bintulu _{CI - C4} Cs - C9 CIO - C20 > C20 a= 0.90 17.0 9.4 39.2 34.4 0.88

Tab/e 3.2.3.2 Product distributions with different a-va/u es (Ei/ers, 1990)

According to Shell data, the arnount of gaseous hydrocarbons (Cl -C4) consists for

(30)

• Reactor placement

Based on the four reactors used for the Fischer-Tropsch synthesis at Bintulu to

produce 12000 barrrels/day, six reactors will be used in this design to pro duce 20 000 barrels/day. These six reactors can be placed either all parallel, as two sections of three reactors (sections in series, reactors in one section parallel), as three sections of two reactors or all in series. To choose the best option of these, the following is considered:

When reactors are placed parallel, the main inlet stream needs to be divided in equal substreams, which will obviously be very difficult in the case all reactors are parallel placed.

2 Reactors placed in series allow for intermediate product and water removal,

which re duces the inlet stream and raises the reactant concentration for the next reactor. However, extra removal units will increase the capital and operational costs considerably, which mIes out the option to place all reactors m senes.

Comparing the two remaining options, - three sections of two reactors or two sections ofthree reactors -, it is decided to couple the six reactors as twosections ofthree reactors, as the larger division head is not outweighed by an extra removal unit.

The units RPS, FlashPR and Flash3 in the overall flowsheet will in practice look as

shown in figure 3.2.3.2 r---"e~'E---________________ ,

-~

---'~ HS

4-}[U

Jr

'S

~

I T [l

(31)

• Conversion

Knowing the product distribution, the overall conversion of 40% and the number of reactor sections (2), the conversion with respect to CO for reactions [3.1.1j] and [3.1.1 k] are calculated by assuming an equal conversion in both reactor sections of 22.5% (see appendix D4 and Table 3.2.3.3).

Carbon number Conversion

Cl 0.017119 C2 0.006087 Cl 0.006225 C4 0.006297 Cs 0.001114 C6 0.001679 C7 0.002621 Cs 0.006382 ClI 0.065255 CIS 0.048713 C30 0.063507

Table 3.2.3.3 Conversion per compoundfor an overall reactor convers ion of22.5%

With these coefficients the reaction can be simulated using a simpie stoichiometric reactor.

3.2.4 Heavy paraffin conversion (HPC)

The hydrocracking reactions occur in a three phase trickle flow reactor where the gas and liquid phases flow cocurrently downwards through a fixed bed of catalyst. The liquid phase consists of the heavy paraffins produced in the HPS reactors and the bottoms recycle from the fractionator. The gas phase reactants consist of the hydrogen feed from the shiftlPSA unit and the H2/tops recycle from the fractionator (containing about 85 mol% hydrogen). The solid phase consists of a dual functional catalyst (Shell proprietary). The catalysed hydroconversion reactions are:

[3.2.4a] [3.2.4b] [3.2.4c] According to Minderhoud (1986) the process conditions for hydrocracking are: - partial pressure ofhydrogen: 25 - 150 bar

- reactor temperature: 250 - 375°C - hydrogen to feed ratio: 250 - 2500 NI/kg

lt is decided to operate the reactor at a hydrogen partial pressure of 33 bar, whiie the total pressure is 40 bar. The temperature at the inlet is 300°C and the hydrogen to feed ratio is 1400 NI/kg (an excess hydrogen of approximately 1300%).

(32)

The HPC reactions are simulated by use of a stoichiometric reactor. First alkenes and

oxygenates are hydrogenated according to reactions [3.2.4a] and [3.2.4b]. To accommodate the simulation ofthe hydrocracking it is assumed that the

hydrogenation has a selectivity of 100%; all alkenes and oxygenates entering the reactor are converted into alkanes. Subsequently the alkanes are cracked via reaction [3.2.4c]. The product distribution is again taken from Eilers (figure 3.2.3.1, mild hydrocracking). Comparing the Fischer-Tropsch and the hydrocracking distribution (see appendix D3) the overall reaction is:

0.190443 H2 + 0.017931 CsH12 + 0.022518 C6HI4 + 0.030124 C7HI6 + 0.064174

CsHls + 0.477239 CllH24 + 0.217713 CIsH3S + 0.170301 C30H62 ~

0.022299 C4HlO + 0.029938 CsH12 + 0.045116 C6H14+ 0.058199 C7H16+ 0.088868 CsHIs+ 0.630076 C IIH24+ 0.234222 CIsH3S + 0.081726 C30H62

[3.2.4d]

This yields the following netto reaction:

0.190443 H2 + 0.088576 C30H62 ~ 0.022299 C4HlO + 0.012007 CsH12 + 0.022598

C6H14+ 0.028075 C7H I6 + 0.024694 CsHIs+ 0.152837 CllH24+ 0.016509 CIsH38

[3.2.4e]

3.2.5 Fractionator

The gas feed to the fractionator is at a temperature of 326°C and the liquid feed at

299

oe.

Both streams have a pressure of 1.8 bar. Therefore no preheat furnace is

needed. At atmospheric pressure the condensor temperature should be 20°C which is

difficult to achieve, especially in a tropical climate. Therefore the condenser is operated at an elevated pressure of 1.2 bar and 40°C. The temperature of the

condensor is 40°C to condense the naphtha and water but leave the H2/tops gaseous. Naphtha is withdrawn from the three-phase flash as a liquid after the partial vapour-liquid condens er. The H2/tops stream leaves the flash at the top, while water is withdrawn from the same flash at the bottom. To ensure vapour-liquid equilibrium in the bottom stages ofthe fractionator, LP steam (190 °C, 3 bar) is injected in the bottom of the column. The kerosene and diesel fractions are withdrawn from two sidestreamstrippers in which steam is injected to strip out the lighter components. These streams are fed back to the fractionator. To ensure vapour-liquid flow on all stages, two pumparounds are used.

To simulate the fractionator an assay of pseudocomponents is generated from the boiling curve ofthe feed by Aspen. These pseudocomponents are defined with 'virtual' physical properties. This boling curve is constructed from the product distribution as mentioned in appendix D3. To avoid the use of pseudocomponents in the whole flowsheet, the fractionator is replaced by an ideal separator and simulated

separately. It is operated in such a way that the product massflows are the same as in

the overall flowsheet. Then the product composition ofthe fractionator simulated by the assay of pseudocomponents is a realistic approach of the real product composition. The thermodynamic model used is Grayson.

(33)

The fractionator is designed with 29 real trays (17 theoretical). The feed is introduced at the 25th tray. The fractionator resembles an atmospheric crude fractionator (25 theoretical stages) in various aspects. This is not unexpected because the feed consists of a synthetic crude.

3.2.6 Flashes

As outlined in the previous paragraphs the process comprises 7 flash vessels: 1 after the steam reformer, 4 in the HPS section, 1 after the HPC and 1 at the top ofthe fractionator. The process conditions and the desired separation are suminarized below:

Equipment no. P (bar) T(°C) top bottom

V112 25.5 75 raw syngas water

V210 22.4 150 syngas, water, light heavy hydrocarbons hydrocarbons

V214 22.3 30 syngas water / light

hydrocarbons V222 20.2 150 syngas, water, light heavy hydrocarbons

hydrocarbons

V225 20.1 30 syngas water / light

hydrocarbons V303 37.0 300 hydrogen, light heavy hydrocarbons

hydrocarbons

V315 1.08 23.3 hydrogen, light water / naphtha hydrocarbons

Table 3.2.6.1 Process conditionsjlash vessels

Before entering the flash vessel, the stream is first cooled or heated to the flash temperature, thus ensuring the flash is operated without heat duty. The pressure drop over each vessel is estimated as 0.1 bar and the separation is simulated using the Peng-Robinson thermomodel.

3.2.7 Heat transfer equipment

The heat transfer equipment ofthis process comprises several types ofheat exchangers:

fired heater

thermosyphon reboiler condensors

exchanger (involving two process streams) cooler (using cooling water as coolant) heater (using steam as heating medium) air cooler (using air as coolant)

(34)

The equipment is simulated by use ofthe Peng-Robinson thermomodel and pressure

drop over the units is not ~en into account.

3.2.8 Pumps

Pumps are simply simulated by only defining the oudet pressure. The simulation programme Aspen Plus calculates the theoretical power and the oudet temperature using Peng-Robinson as thermomodel.

3.2.9 Compressors

The three compressors in the process are defmed as polytropic compressors, whereas the gas turbine is simulated isentropically. The theoretical power and the outlet temperature is calculated by use ofthe Peng-Robinson thermomodel.

3.3 Heat integration

In any plant design a considerable number of streams needs either to be heated or to

be cooled. The heat removed from the hot streams can in theory be used to heat the cold. However, one should consider whether these exchanges are physically possible, i.e. at any place in the exchanger the hot stream must have a higher temperature than the cold. To examine this integration and to fmd the best heat exchanger network, the pinch technology as described by Douglas (1988) is used.

All heat exchanges are taken into consideration, except for the reboiler duty ofthe stripper and the condenser and cooling duties ofthe fractionater. Thereboiler duty cannot satisfactory be evaluated by this method and the fractionator will certainly be cooled by cooling water to insure an accurate control ofthe fractionater. However, the preheating of the boiling feed water (BFW) for the HPS reactors is taken into account,

as this requires a considerable amount of energy. An overview ofthe streams to be

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Stream Hot! Tin Tout M Qav FCp

number Coid (OC) (OC) (kg/sec) (MW) (kW/oC)

113 Hot Par. con. 800 226 290.331 362.77 632 gas

226 225 140.70 140700 condens

225 75 117.75 785 I!guid

124 Hot 227 30 2126.28 1658.55 8419

208 Hot Par. con. 247 150 357.152 89.59 924

215 Hot Par. con. 150 30 345.106 178.27 1486

225 Hot Par. con. 240 150 304.959 69.75 775

230 Hot Par. con. 150 30 295.829 144.82 1207

309 Hot 422 250 12.200 9.71 56

316 Hot 142 30 20.654 5.36 48

318 Hot 241 30 10.901 5.71 27

118 Coid Evap. 79 225 62.911 -39.42 270 Iiquid

225 226 -115.10 115100 condens

226 500 -36.99 135 I gas

206 Coid Evap. 30 225 54.942 -51.87 266 Iiquid

225 226 -100.50 100500 condens

226 500 -36.44 133 [&as

214 Coid 25 250 119.415 -191.44 930

220 Coid 30 250 304.959 -170.12 773

238 Coid Par. evap. 29 320 16.327 -16.20 56

239 Coid 151 300 21.177 -8.70 58

302 Coid Par. evap. 171 330 25.866 -14.92 94

BFWto HPS Cold 20 224 219.1 -186.85 916

Table 3.3.1 Overview of streams to be heated or cooled

In

this tabie the temperature ofthe Iisted streams is given as Tin, the desired

temperature is given as Tout. The enthaipy change involved in this temperature change is the heat avaiIabie (Qav). The Iisted Fcp- values are calculated by:

Qav = FCp.(Tin-Tout) [3.3a]

In

three cases (streams 113, 118 and 206) the heat exchange involves a considerabie phase change which implies the FCp-vaIues will not be constant over the whoie temperature range. We can incorporate phase changes that take place at constant temperature into this forrnalism simpIy by assuming a 1 °C temperature change at the temperature ofthe phase change and then caiculating a fictitious FCp-vaIue that gives the same heat duty as the phase change; i.e.

[3.3b] As the phase changes of above mentioned streams all involve water/steam at

approximately 25 bar, the phase changes take place between 225°C and 226°C and the heat of evaporation or condensation is 1830 kJ/kg. The other partial condensers and evaporators all involve mixtures with large boiling ranges. For these streams the FCp-value will be taken as a constant.

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FV03201 _{Process structure}

The minimwn temperature difference to exchange heat from one (hot) stream to another (cold) is chosen to be 10

oe.

This is incorporated by shifting the temperatures ofthe cold streams 10

oe

upwards, obtaining the following temperature diagram (figure 3.3.1). 632

192~

I

17;5

I

Iill !ill

~

@!]

~

140700 ...."...,,-::-::--\: \~7!85:"""--J,-8_41_9... 1486

[IW7J

[!!)

~ ~ ~ 8oor--. _____________________________ ï _ _ _ _ _ _ _ _ _ _ _ _ _ _ _ _ _ _ _ _ _ _ _ _ _ _ _ _ _ _ ~790Istr.nol 600 610 500~-r---~---~510 400~~---~---_+---~_+----~---~410 300~~---_+---+__+----~---_,~--~~--~310 200~--_+--+_~--~~~---_+--+_--_1----~--~~--~ H---,~_l2 10 100~--~--+_----+_----+_----+__+---~--_1----~--~~~~----~ 0L-______________________________ ~ ____________________________ ~

Figure 3.3.1 Temperature diagram

The heat transfer in each temperature interval is calculated by:

Q. _I

=

["(Fe ) -"(Fe)

_L.J _{P bot}_,_i _L.J _{P cold,i}

]~T

_I [3.3c] The heat available in a certain interval can be transferred to the interval just below, If we evaluate all temperature intervals from the top down, adding heat by hot utility and removing heat by a cold utility where needed, the cascade diagram in figure 3.3.2 is obtained.

llO

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FV0320J Process structure 800 0 790 510 183.28 500 500 186.92 490 422 215.31 412 340 249.78 330 330 253.05 320 310 258.47 300 300 260.60 290 260 269.12 250 C 250 H 254.22 0 240 0 L 247 T 249.58 D 237 241 U 245.85 U 231 T T 240 I 245.26 I 230 L L 236 I 245.97 I 226 T T 235 Y 30.82 Y 225 234

~

30.73 224 227

~

23.71 217 226

~

31. 59 216 225

~

179.55 215 181 13t641 533.19 171 161 1 1 \2.621 695.80 151 150

~

785.88 140 142

cp

859.35 132 89

~

1348.59 79 85

~

1386.59 75 75

r:::pTI

1481.60 65 87.16 65 1568.76 55 50 40 40 30 39 29 35 25 30 1737.14 20

Figure 3.3.2 Cascade diagram

One ob serves that hot utility is not needed as sufficient heat is available over the whole of the temperature range. This means this system has no pinch temperature and the temperature difference is always larger than the minimum temperature difference. This is also shown in the temperature-enthalpy diagram (figure 3.3.3), where the temperature is plotted against the cumulative enthalpies ofthe hot and cold streams seperately. The minimum cooling duty is used as starting point for the cold streams and the heating duty is obviously zero, as both curves end at the same enthalpy.

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FV0320J 900 800 700 600 - 500 2-... 400 300 200 100 0 0 500 1000 1500 H(MW)

Figure 3.3.3 Temperature-enthalpy diagram

Process structure

2000 2500 3000

The enthalpy difference at each temperature is shown in the grand composite curve in figure 3.3.4. It is obvious that this system has no pinch as the enthalpy difference is larger than zero at any temperature (ex cept for 800°C due to the absence of the cold curve). 900 800 700 600 - soo 2-... 400 300 200 100 0 0 500

Figure 3.3.4 Grand composite curve 1000 H(MW)

1500 2000

In theory a heat exchanger network can now be designed without the need of hot utility. However, as this would probably be only achieved with a very large number of heat exchangers, this is not the most economical solution. In designing the network without pinch limiations (such as the evaluation of the FCp-values just above and just below the pinch), only the avoidance oftemperature crosses should be taken into account. Considering the process controllability, the streams exchanging heat should preferably be in the same section. The proposed network and the considerations which led to this network, are shortly described below.

The gas phase of stream 113 (leaving the steam reformer) is cooled to provide the heat of evaporation for stream 118 and 206, using thermosyphon reboilers. The remaining heat is removed in a condens er and transferred to preheat the BFW used in the HPS reactors. Doing this, there will still be an excess of heat to be removed from stream