Appendix 1: Project description, Conceptual Process Design

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Appendix 1: Project description, Conceptual Process Design

TITLE OF THE PROJECT:

Recovery of Carbon Monoxide and Hydrogen from Low Joule Gas followed by reaction with

Ethylene.

PROJECT NUMBER:

CPD-3262

COURSE INSTRUCTION:

Process Systems Engineering, DelftChemTech, DelftUT

Julianalaan 136, 2628 BL Delft

Ir. Pieter Swinkels,

P.L.J.Swinkels@tnw.tudelft.nl

Prof. ir. Johan Grievink,

j.grievink@tnw.tudelft.nl

COACHING:

Process Systems Engineering, DelftChemTech, TUDelft

Prof. ir. G.J. (Jan) Harmsen Btw,

G.J.Harmsen@tnw.tudelft.nl

PROJECT PRINCIPAL:

European Planning Center/ Rotterdam Refinery, ExxonMobil

Postbus Rotterdam

Jaap de Glopper, 0(10)4874578,

jaap.deglopper@exxonmobil.com

Jaap de Glopper will be absent in the period of 17 September to 8 October. During this

period he will be replaced on the project by Aad Rooijmans (Flexicoker Technical Console

Leader), telephone: 0(10)4874357.

TEAM MEMBERS:

Johan Breugem,

a.j.breugem@tnw.tudelft.nl

Lisette Gerritsma,

l.j.gerritsma@tnw.tudelft.nl

Ralph Krul,

r.a.krul@tnw.tudelft.nl

Maarten Over,

m.j.j.over@tnw.tudelft.nl

PROJECT DESCRIPTION PREPARED BY:

Ir.drs. Giljam Bierman, 22nd August 2001

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Designers: A.J. Breugem R.A. Krul L.J. Gerritsma M.J.J. Over Project ID-Number : CPD3262 Date : December 4th 2001

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DESCRIPTION:

The assignment on Conceptual Process Design (CPD) concerns the recovery of Carbon

Monoxide and Hydrogen from Low Joule Gas (LJG). LJG is produced in the Flexicoker

operated by the Esso refinery in Rotterdam and is mainly used as fuel gas in the refinery

furnaces. (see reference 1). Accordingly, the value of LJG is comparable to the value of

natural gas (Fuel value). To make a more economic use of the LJG, Esso aims at extracting

those components from the LJG, which can be used as feedstock for the production of

chemicals.

Low Joule Gas is formed by the gasification of cokes with air and composes of Carbon

Monoxide, Hydrogen, Nitrogen, Carbon Dioxide, Methane and small amounts of Hydrogen

Sulphide, Ammonia. Hydrogen Sulphide and Ammonia are removed from the LJG by wet

scrubbing. For the composition of the cleaned LJG reference is made to Table 1.

The main valuable components in LJG are Carbon Monoxide and Hydrogen. Depending on

the relative amounts at which they are made available, they can be used as feedstock for a

number of processes. Two options to be mentioned are:

-

97% Carbon Monoxide for the production of Ethylene copolymer

- 60-70%

H

2

/CO mixture for the production of Propion Aldehyde by hydroformylation of

ethylene using the ExxonMobil Rhodium Technology. (see reference 2,3)

Next to the Carbon Monoxide and Hydrogen, Ethylene is a feedstock for the reactions

mentioned. Economically it will be favourable to use a locally available source of Ethylene.

At the Esso Refinery this will be Unsaturated High Joule Gas (UHJG) which contains

Ethylene at concentration of about 8mol%.

OBJECTIVE:

The objective of the project is to make a conceptual process design for the recovery and use

of Carbon Monoxide and Hydrogen from Low Joule Gas originating from the Flexicoker at

the Esso refinery in Rotterdam to produce valuable products. For this Unsaturated High

Joule Gas (UHJG) containing Ethylene (8mol%) is available at the Esso refinery. For

detailed composition and properties reference is made to Table 2 and 3.

The process composes of a primary separation section in which Carbon Monoxide and/or

Hydrogen is removed from the LJG, a reaction section in which the Carbon Monoxide and

Hydrogen are reacted with Ethylene and a second separation section in which the final

products are separated from the unconverted feedstock and by-products.

The ultimate challenge within the project is to find a economic viable separation process for

the difficult separation of Carbon Monoxide and Hydrogen from the LJG. In this a chemical

separation driven step is most likely preferred over a physical separation step. One possibly

design alternative is to avoid the first separation step and directly carry out the reactions with

low concentrated reactants. This means that the relatively difficult separation step between

Carbon Monoxide and Hydrogen and the LJG will be replaced by a relatively easy

separation step between the reaction products (Ethylene Copolymer or Propion Aldehyde)

and the product stream. Drawbacks for this design alternative are the high flows through the

reactor and possible formation of by-products.

In the first part of the project (first 3 weeks), a choice has to be made which design

alternative will be most likely to be favourable. This will be reported as the Base of Design.

The chosen alternative will be worked in detail during the second part of the project (9

weeks).

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ORGANISATION AND SCHEDULE:

The team is responsible for communication between the team and the Principle, i.e.

appointments for meetings, writing of minutes, distribution of the report, inviting

representatives for the Assessment Meeting.

Project Schedule (TENTATIVE)

Date

Time

Location

Issue of the project:

04-09-01 14h00 PSE-Conference Room, Delft

Kick-off meeting with the Principle:

12-09-01 !!

Esso Rotterdam Refinery

Delivery of the Base of Design Report:

23-09-01 -

-

Review of the Base of Design:

27-09-01 15h00 PSE-Conference Room, Delft

Delivery of the Final Report:

27-11-01 -

-

Assessment Meeting:

04-12-01 -

PSE-Conference Room, Delft

Presentation of the work at Principle:

?

Esso Rotterdam Refinery

!! The Kick-off meeting with the Principle has to be confirmed with the Principle (deGlopper)

on Wednesday afternoon 5

th

of September by telephone. Preferably the course instructor

(Pieter Swinkels) will also attend this meeting.

Within the framework of the long-term education-project on creativity in design,

DelftChemTech provides the team with a mentor (Jan Harmsen). The mentor will coach the

team during the project. For this, meetings will be scheduled on a regular basis.

LITERATURE REFERENCES:

1 http://www.essobenelux.com/eaff/essobenelux/nl747/page12.html#rotterdam

2 Gabor Kiss et al., Hydroformylation Process, Exxon Chemical Patents Inc., Houston Tex.,

US6,049,011, April. 11, 2000.

3 Gabor Kiss et al., Direct Hydroformylation of a multi-componant synthesis gas containing

Carbon Monoxide, Hydrogen, Ethylene and Acethylene, Exxon Research & Engineering

Company, Florham Park, NJ, US5,675,041, Oct. 7.1997.

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A4

TABLES:

Table 1. Molar composition of Low Joule Gas (After Scrubbing)

Component

Concentration (mol%), dry

CO 22

H

2

16 CO2 8

N2 52.5

CH4 1.5

H2S 300

vppm

COS 100-150

vppm

Hydrocarbon

~10 wt ppm diesel like poly nuclear aromatics

Flexsorb-SE

saturated at 1.3 barg and 40°C, some liquid entrainment

Coke

~ 1 ppm range sub micron particles (we don't exactly know, but

we find the dirt everywhere)

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Table 2. Composition of Unsaturated High Joule Gas (UHJG) at the Esso Refinery

Rotterdam

Component Concentration

(mol%)

Typical

Range

H

2

16.3 16.0

-

19.0 C1 46.8 40.5

-

47.0 C2 20.4 20.2

-

20.7 C2 =

8.1 8.0 - 9.1

C3 0.9

0.4 -

1.0 C3 =

0.4 0.2 - 0.5

iC4= 0.092 0.06

-

0.12 C4 =

0.004 0.002 - 0.006

iC4 0.4

0.1 -

0.5 nC4 0.074 0.05

-

0.1 C5+ 0.02

0.01 -

0.03 CO 0.8

0.5 -

3.5 N2 5.4

4.4 -

5.5 CO2 10

vppm

5 -

15 H2O 0.33

Dew

Point

COS 0.7

vppm

0.5 -

1.5 RSH 6.3

vppm

4 -

8 H2S 14.2

vppm

10 -

21 Table 3. Normal operating conditions of Unsaturated High Joule Gas (UHJG) at the Esso

Refinery Rotterdam, (Feed to sales gas cleanup and compression unit, balance to

RHJG system)

Property

Value

Temperature,

o

C

40 Pressure, bar(g)

2.4 ANSI Flange Class

-

Molecular Weight

18.5-19.2

Heating Value (LHV), MJ/kg

44.5-47.0

Heating Value (HHV), MJ/kg

48.8-51.6

Sulfur Content, kg/FOET

0.05

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Designers: A.J. Breugem R.A. Krul L.J. Gerritsma M.J.J. Over

Project ID-Number : CPD3262

Date : December 4th 2001

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Appendix 3: Additional information on propion aldehyde

Physical properties

Propion aldehyde is completely miscible with organic solvents such as alcohols, ether, and

benzene. Solubility in water decreases at elevated temperature. For example, propion

aldehyde is miscible with water in all proportions below 15 °C, but becomes less soluble and

separates from solution in aqueous mixtures containing ca. 20 – 70 wt % propion aldehyde

as the temperature increases. Some physical properties of propion aldehyde follow.

Sources: [39-43].

TABLE A3.1: FLASH POINTS, AUTO IGNITION TEMPERATURE AND SOLUBILITY OF PROPION ALDEHYDE

1)

Property

Value

Note

Flash point

–30.0 °C

Closed cup

–7.2 – 9.4 °C

Open cup

Auto ignition temperature

206.85 °C

Solubility of propion aldehyde in water (20 °C)

35.6 wt %

Solubility of water in propion aldehyde (20 °C)

21.1 wt %

1)

More properties are given in appendix 2: Pure components list.

Storage and transportation

For storage and transportation of propion aldehyde, containers of stainless steel are

normally used. Vessels lined with polyethylene or other coatings are also suitable. For

aldehydes that enter the market as solutions aluminium vessels or containers of standard

steel should not be used, because the acids formed by auto-oxidation are corrosive, and the

corrosion products can cause discolouration of the aldehyde. Condensation of propion

aldehyde can be caused by light and oxygen, which is particularly troublesome in high-purity

propion aldehyde. Therefore, aldehydes are normally stored under a nitrogen atmosphere.

Propion aldehyde is highly reactive and forms explosive peroxides on exposure to air.

Addition of water also reduces peroxide formation and provides stabilization against

metal-catalysed condensations. Also other antioxidants and stabilizers can be added to prevent

auto-oxidation [24]. Caution is advisable with emptied containers containing residual

vapours, which may explode on ignition. Trimerization tendency of aldehydes occasionally

increases at lower temperatures, so that some compounds cannot be stored over long

periods at temperatures under 20 °C without stabilization [44].

Environmental release consequences

TABLE A3.2: ENVIRONMENTAL INFLUENCE OF PROPION ALDEHYDE

Release

Degradation

Influence

Into the soil

 Biodegrades to a moderate extent

Unknown

Into water

 Biodegrades readily

 Evaporates quickly

Half-life between 1 and 10

days

Into the air

 Degrades readily by reaction with

photo chemically produced

hydroxyl radicals

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A8

Appendix 4: Market information of propion aldehyde

TABLE A4.1: THE PRODUCTION OF PROPION ALDEHYDE DERIVATIVES IN 1988 (kt/a)

Compound

United States

Western Europe

1-Propanol 88.5

11.0 Propionic acid

47.2 -Trimethylolethane 2.7

-Others 1.4

2.0 Source: [45]

TABLE A4.2: OXO CAPACITIES BY REGION IN 1984 AND 1994 (KT/A)

Compound

1984

1994

Western Europe

2.5

2.4 United States

1.4

2.4 Japan 0.5

0.8 Others 0.6

1.6 Total 5.0

7.2 Source: [46]

TABLE A4.3: CONSUMPTION OF OXO CHEMICALS BY REGION IN 1993 AND 1998 (KT/A)

Compound

Western

Europe

United States

Japan

Totals

1993 1998

1993

1998

1993

1998

1993 1998

Propion aldehyde

11

12

161

183

1

173

196 n-Butyraldehyde 1,224 1,274

1,055

1,178

572

622 2,851 3,074

Isobutyraldehyde 133 128

234

263

65

72 432 463

Valeraldehydes

11 12

32

35 43 47

Totals

1,379 1,426

1,482

1,659

638

695 3,499 3,780

Source: [46]

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Appendix 5: Propionic acid usage

TABLE A5.1: WORLDWIDE PRODUCTION CAPACITIES OF PROPIONIC ACID IN 1989

Producer

Country

Capacity

(kt/a)

Notes

BASF FRG

60 Union Carbide

USA

68 Used for production of

various carboxylic acids.

Eastman – Kodak

USA

25 Used for production of

various carboxylic acids.

BP Chemicals

UK

30 Hoechst – Celanese

USA

7 Daicel

Japan

2 Total 192

The areas of use of propionic acid and its distribution

Animal feed and corn preservatives 42%

Herbicides 22% Calcium and sodium salts

17%

Cellulose acetate propionate 17%

Miscellaneous 2%

FIGURE 5.1: THE AREAS OF USE OF PROPIONIC ACID AND ITS DISTRIBUTION

In the medium-term an increase in its use as a preservative for animal feeds is predicted. In

contrast, the use in herbicides will decrease because of the increasing use of optically active

substances [6].

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A10

Appendix 6: Product options and selection

The following options are considered:

1. Recovery of pure hydrogen from the LJG streams

2. Recovery of pure carbon monoxide from the LJG streams

3. Recovery of pure ethylene from the UHJG streams

4. Production

of

methanol

5. Production

of

neoacids

6. Water/gas shift reaction to produce hydrogen

7. Copolymerisation of ethylene and carbon monoxide to polyketons

8. Hydroformylation of ethylene to propion aldehyde

These options are reviewed by checking the criteria mentioned in chapter 2.

Ad 1. Although hydrogen has a good economic potential (about 830$/t, [9]), the mass

percentage of hydrogen in both the LJG stream and the UHJG stream is very low (about

1 w%). At ideal separation the yield will still be low. Furthermore, the heat of combustion

of the hydrogen-free stream would be lowered significantly, resulting in relative low value

of this stream. In case of hydrogen recovery, cryogenic processing and pressure swing

adsorption are economical for this operating scale. Water, however, may cause

solidification.

TABLE A6.1: ECONOMIC POTENTIAL OF THE RECOVERY OF HYDROGEN

Item

Price / unit (k$/kt)

Units (kt/a)

Value (M$/a)

Feed LJG

19

(1)

77

1.4 Product H

2

830

(2)

1

0.8 Remaining 12

(1)

76

0.9 Margin

0.3 Total Capital

Investments

1 M$

(1)

Value is determined by comparing the heat of combustion with natural gas

(2)

[9]

Since relatively expensive compressors or refrigeration apparatuses are required to obtain

purified hydrogen, this option is rejected.

Ad 2. Carbon monoxide is a valuable chemical. The mass percentage in the LJG stream is

about a 25%. For the evaluation of the economic potential of the process options, a base

case is taken. The maximum amount of UHJG is used a stoichiometrically mixed with

LJG. At this feed composition the margin is determined.

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The results are depicted below.

TABLE A6.2: ECONOMIC POTENTIAL OF THE RECOVERY OF CARBONMONOXIDE

Item

Price / unit (k$/kt)

Units (kt/a)

Value (M$/a)

Feed LJG

(xCO =

22mol%)

19

(1)

77

1.4 Product CO

140

(2)

18

2.5 Remaining 12

(1)

58

0.7 Margin

1.8 Total Capital

Investments

6 M$

(1)

Value is determined by comparing the heat of combustion with natural gas

(2)

[47]

The recovery of carbon monoxide can be achieved by means of cryogenic processes;

liquid adsorption processes, new complexing compounds and pressure swing adsorption

[48]. An evaluation:

– Cryogenic processes are economically favoured at large scale processing and can

lead to efficient separation. Therefore, there are numerous plants operating

successfully. Energetically, however, it is very consuming. Furthermore, the LJG

stream contains water and nitrogen. To prevent plugging and solidification the water

needs to be removed. Nitrogen is a bigger problem. The boiling points of carbon

monoxide and nitrogen differ only a couple of degrees, leading to a difficult

separation and expensive apparatuses.

– A liquid adsorption process like the COSORB process is an apparent less costly

alternative. The selective and reversible chemical adsorption uses a complex CuAlCl4

dissolved in toluene, which is inert to H2, CO2, CH4 and N2. This carbon monoxide is

bound to the complex and separated from the other gases. After separation the CO is

reversibly freed from the complex by heating. The yield and purity are about 99% and

the process is in case of large nitrogen contents in the feed an economical choice. A

drawback, however, is the rapid degradation of the adsorbent due to the sulphur and

water content of the stream. Therefore stringent pre-treatment is necessary.

Water removal at 10 bara (pressure for COSORB process) needs equipment costs of

about 0.1 M$ for processing 300,000 Nm

3

/h with silica or alumina triethylene glycol

adsorption.

Tenneco Chemicals ltd. operates this process and a few plants have been built since

the 1970s. In the beginning it seemed to be a very promising process, during

processing however some technological problems occurred; The recovery of toluene

appeared to be quite difficult and the process is very sensible to oxygen

contamination [49]. Referring to several publications the COSORB process is

competing with cryogenic distillation [50].

The costs mentioned by [51] can offer an estimate for the costs of equipment and

operating. A simple calculation shows the investment costs would exceed the

maximum allowed 6 M$.

– New complexing compounds are economically competitive with pressure swing

adsorption. Very high recoveries and purities can be obtained, but the current

knowledge is limited to bench-scale experiments.

– Pressure swing adsorption is another process. At 98% recovery, the purity can be

about 98%. The amount of catalysts is very limited (Na-type mordenite, active carbon

supported copper and activated carbon). The catalyst is poisoned by CH

, which is

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A12

It is clear the allowed total capital investments of 6 M$ will not be sufficient to recover CO

from the LJG.

Ad 3. The determination of the economic potential for the recovery of ethylene is shown in

the following table.

TABLE A6.3: ECONOMIC POTENTIAL OF THE RECOVERY OF ETHYLENE

Item

Price / unit (k$/kt)

Units (kt/a)

Value (M$/a)

Feed UHJG

(x

C2H4

=

22mol%)

194

(1)

170

33.0 Product C

2

H

4

555

(2)

20

11.3 Remaining 194

(1)

149

29.0 Margin

7.3 Total Capital

Investments

25 M$

(1)

Value is determined by comparing the heat of combustion with natural gas

(2)

[52]

The price of ethylene is 0.25$/lb, which is 555$/t. At the Esso Refinery Rotterdam, there

is a ethylene stream available at 600$/t. From this point of view it would be interesting to

recover the ethylene. Some options for the recovery of ethylene are membranes;

cryogenic distillation and pressure swing adsorption. The most promising method for this

scale of processing is to make use of membranes. Compression to about 50 bar and two

membranes are needed. Unless the surmountable investment costs, the process is not

efficient due to the relatively low recovery.

Ad 4. For the production of methanol a purified feedstock is needed. Purification is also

necessary for other alternatives, as discussed below. On the case of production of

methanol especially, the market is competitive and the margins are relatively low,

depicted in the following table. This leads to the discarding of this process.

TABLE A6.4: ECONOMIC POTENTIAL OF THE PRODUCTION OF METHANOL

Item

Price / unit (k$/kt)

Units (kt/a)

Value (M$/a)

Feed LJG

19

(1)

77

1.4 Product Methanol

157

(2)

8

1.2 Remaining

9

(1)

69

0.6 Margin

0.5 Total Capital

Investments

2 M$

(1)

Value is determined by comparing the heat of combustion with natural gas

(2)

[52]

Ad 5. Neo-acids can be produced from olefins, carbon monoxide and water. These kinds of

products are specialty chemicals, implying high margins. Using the LJG and UHJG

streams, however, cancels this option, since it involves complex and sophisticated

processing. A clear drawback for the production of neo-acids is the lack of proven

technological knowledge.

Ad 6. The water/gas shift reaction is producing H

2

out of CO and H

2

O. The apparently best

option is to react the impure CO with water (45 bara, 700 K, [53]), followed by the earlier

discussed recovery of hydrogen.

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Therefore this option is comparable to the recovery of H

2

. The water/gas shift reaction

case roughly differs from the solely hydrogen recovery on:

- An extra pre-treatment section

- A compressor

- An extra reaction section

- A higher production of hydrogen and a consequently more efficient recovery

section (economy of scale)

TABLE A6.5: ECONOMIC POTENTIAL OF THE WATER/GAS SHIFT REACTION TO HYDROGEN

Item

Price / unit (k$/kt)

Units (kt/a)

Value (M$/a)

Feed LJG

18

(1)

77

1.4

2

12

0.0 Product H2 833

(2)

2

1.6 Remaining

6

(1)

63

0.4 Margin

0.8 Total Capital

Investments

2.7 M$

(1)

Value is determined by comparing the heat of combustion with natural gas

(2)

[9]

The higher TCI is cancelled out by the higher necessary investments. At last an

important insight is that the reactant CO is more valuable and needed within Exxon

chemicals than the product H

2

. Considering above-mentioned insights, the water/gas

shift reaction does not appear to be a very economical option.

Ad 7. The price of polyketones is relatively high, about 6 DM/kg [54] which is 2875$/t. The

economic margin is therefore high, which is depicted in the following table.

TABLE A6.6: ECONOMIC POTENTIAL OF THE COPOLYMERISATION OF ETHYLENE TO POLYKETONS

Item Price / unit (k$/kt)

Units (kt/a)

Value (M$/a)

Feed Ethylene

600

(1)

19

12 CO

1000

(1)

18

18 Product Polyketone

2875

(2)

41

117.1 Margin

87.1 Total Capital

Investments

293 M$

(1)

For a fair comparison, the value of the purified compounds is used.

(2)

[54], kDM/T]

Considering the margin, the copolymerisation of ethylene and carbon monoxide appears

to be a good option. The other criteria leave different results. The market, with a

relatively small amount of players, is growing 80% annually. The polyketones are

specialty chemicals, which require intensive capital usage. Furthermore the product

needs to be developed, needs to have a broad range in product specification possibilities

and needs to have a purpose, such as replacing traditional metal gasoline tanks. The

product specifications might be a problem, due to the quality of LJG. The higher olefins

might be built in, the catalyst may be poisoned and the coke fines will remain in the

product, yielding black polyketones. Technological drawbacks are the required minimum

97% purity of carbon monoxide, the purity of ethylene, the pressure of about 45 bara

(doubled pressure if compared with the other process options) and the three phase

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A14

Hydroformylation to propion aldehyde is the final reviewed option. The bulk price for propion

aldehyde is about 950$/t [7]. The margin in the base case scenario is shown below.

TABLE A6.7: ECONOMIC POTENTIAL OF THE HYDROFORMYLATION OF ETHYLENE TO PROPION ALDEHYDE

Item

Price / unit (k$/kt) Units (kt/a)

Value (M$/a)

Feed UHJG

194

(1)

160

31.1 LJG

18

(1)

77

1.4 Product Propion

aldehyde

942

(2)

32

29.8 Remaining

140

(1)

204

28.5 Margin

25.9 Total Capital

Investments

87 M$

(1)

Value is determined by comparing the heat of combustion with natural gas

(2)

[7].

As the principal suggested this option, it is assumed that there is an internal market. The

process conditions are relatively mild; the operating pressure is about 15 bara, the

temperature about 370 K [55]. Another advantage of this process is the presence of

Exxon Mobil patents on this process [20,12] though it should be said these patents are

concerning the production of propion aldehyde from a feed that is of higher quality than

the LJG and UHJG streams. A major drawback is the price of the rhodium catalyst, being

10 k$/kg, and its sensitivity towards poisoning.

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A16

Appendix 7.2: ASPEN PLUS Flow Sheet

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A18

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Appendix 10: Definition of feed and product streams

In the following tables the composition of all in and outgoing streams of the design are given.

Feed

TABLE A10.1: COMPOSITION OF THE FEED STREAMS

Stream 1:

Feed LJG

Stream 4:

Feed UHJG

Temperature (K)

313.15

313.15 Pressure (Bara)

2.3

25 Phase (V/L/S)

V

Component kt/a

kt/a

Hydrogen 0.9668

2.7793

Methane 0.7213

63.5045

Ethane 0

51.8845

Ethylene 0

19.2201

Propane 0

3.3568

Propylene 0

1.4237

iso-Butylene 0

0.4366

1-Butylene 0

0.0190

iso-Butane 0

1.9665

n-Butane 0

0.3638

2-Methylbutane 0

0.1221

Carbonmonoxide 18.4703

1.8953

Nitrogen 44.0816

12.7950

Carbondioxide 10.5529

0.0037

Water 1.3688

0.5028

Carbonylsulphide 0.0270

0.0004

Methylmercaptan 0

0.0026

Hydrogensulphide 0.0306

0.0041

Propionaldehyde 0

0 n-Butyraldehyde 0

0 iso-Butyraldehyde 0

0 3-Methylbutyraldehyde 0

0 Dimethylpropionaldehyde 0

0 Valeraldehdye 0

0 2-Methylbutyraldehyde 0

0 Methanol 0

0 Tetraglyme 0

0 Total 76.2193

160.2807

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A24

Products

TABLE A10.2: COMPOSITION OF THE PRODUCT STREAMS

Stream 3:

Condensed

Water

Stream 15:

Off-gas

Stream 19:

Product

Stream 20:

By-product

Temperature (K)

308.15

307.8 317.4 350.8

Pressure (Bara)

7.6

22 1 1.2

Phase (V/L/S)

L

V

L

Component kt/a

kt/a

Hydrogen 0

2.4123

0

0 Methane 0

64.2258

0

0 Ethane 0

51.8845

0

0 Ethylene 0

1.7298

0

0 Propane 0

3.3546

0.0022

0 Propylene 0

0.1281

0

0 iso-Butylene 0

0.0326

0.0067

0 1-Butylene 0

0.0014

0.0003

0 iso-Butane 0

1.6997

0.2668

0 n-Butane 0

0.2708

0.0930

0 2-Methylbutane 0

0.0549

0.0671

0 Carbonmonoxide 0

1.8330

0

0 Nitrogen 0

56.8766

0

0 Carbondioxide 0

10.5766

0

0 Water 1.0417

0.1630

0.6853

0 Carbonylsulphide 0

0

0 Methylmercaptan 0

0

0 Hydrogensulphide 0

0

0 Propionaldehyde 0

12.1177

23.8532

0.2395

n-Butyraldehyde 0

0.2668

0.0118

1.8749

iso-Butyraldehyde 0

0.0091

0.0121

0.0454

3-Methylbutyraldehyde 0

0.0259

0 0.5657

Dimethylpropionaldehyde 0

0.0017

0.0002

0.0164

Valeraldehyde 0

0.0005

0 0.0252

2-Methylbutyraldehyde 0

0

0 0.0008

Methanol 0

0.0002

0.0013

0.0002

Tetraglyme 0

0

0 Total 1.0417

207.6654

25.0002

2.7682

(25)

Appendix 11: BOD Margin

TABLE A11.1: MAXIMUM INVESTMENT DETERMINATION AT DCFROR 60%

NET PRESENT- @ FUTURE VALUES

NET FUTURE VALUES (1)

NET PRESENT VALUES

No Discount

Discounted, Accumulated

END

CAPIT. COSTS

CASH FLOW

DISC.

CAPIT.

CASH

YEAR

ANN.

ACCUM.

ANN.

ACCUM.

NFV

FACT.

COSTS

FLOW

NPV

NO.

@

ACCUM. ACCUM.

DCFROR

M$

60%

M$

1

9.3

9.3 -9.3

1.000

9.3 -9.3

2

9.3

18.6 -18.6

0.625

15.1 -15.1

3

14.5

14.5 -4.0

0.391

5.7 -9.4

4

14.5

29.1

10.5

0.244

9.2 -5.9

5

14.5

43.6

25.1

0.153

11.5 -3.6

6

14.5

58.2

39.6

0.095

12.8 -2.3

7

14.5

72.7

54.2

0.060

13.7 -1.4

8

14.5

87.3

68.7

0.037

14.3 -0.9

9

14.5

101.8

83.2

0.023

14.6 -0.5

10

14.5

116.4

97.8

0.015

14.8 -0.3

11

14.5

130.9

112.3

0.009

14.9 -0.2

12

14.5

145.5

126.9

0.006

15.0 -0.1

13

14.5

160.0

141.4

0.004

15.1

0.0

14

14.5

174.6

156.0

0.002

15.1

0.0

15 ( 3 ) :

0.9

175.5

156.9

0.001

15.1

0.0 ACCUM.

18.6

175.5

156.9

2.664

15.1

0.0

1.0

0.0 N.B. :

1. Cash-Flows "Including Taxes and Depreciation".

2. Earning Power =

Interest, for which [Cash Flow - Capital]@Disc. = 0

Disc. Factor =

1/(1 + r)^n with r = interest fraction

3. Rest Value =

5.0%

of Capital Investment

RATIO :

[Cash Flow / Capital] @ Disc.

(26)

A26

Appendix 12: Azeotropes

Figure A12.1: x/y-diagram for the mixture of propion aldehyde and water at 1 atm. and 22

bara.

TABLE A12.1: AZEOTROPIC DATA OF SOME ALDEHYDES WITH WATER [56]

Aldehyde

b

p,aldehyde

b

p,azeotrope

H

2

O in azeotrope

°C

w%

Propanal

47.9

47.5

2.0 Butanal 74.8

68.0

9.7 2-Methylpropanal 63.3

60.1

9.6 Pentanal 103.3

83.0

19.0 3-Methylbutanal 92.5

77.0

12.0

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

0

0.2

0.4

0.6

0.8

1 x

N-PRO-01

y

N-PRO -01

1 atm

22 bar

(27)

(28)

(29)

(30)

A30

Appendix 14: Mass and heat balances

(31)

Appendix 15: Heat integration

Results heat integration of reactor feed with reactor outlet (HE,1) and crude aldehyde stream

(HE,2). For these calculations a heat transfer coefficient of 30 W/(m

2

*K) is taken. T

cold,in

was

assumed to be 340 K.

TABLE A16.1: OVERVIEW OF AREAS FOR INTEGRATED HEAT EXCHANGE

Temperature cold

stream from HE 1

A

HE,1

A

HE,2

86 890

60 85.5 832

93 85 779

114 84.5 730

178 84 690

Error:

T

incold

> T

out hot

using

counter current HE

This shows an unrealistic large heat exchanger is needed to transfer heat from the gaseous

reactor outlet to the gaseous reactor inlet.

(32)

A32

(33)

(34)

A34

(35)

Appendix 17.2: Equipment calculation of the desulphurisation unit

General assumptions

The total sulphur content to has to be reduced to 50 wppb to ethylene. This is done in a

two-stage reactor, first COS is converted to H

2

S and secondly the H

2

S and CH

3

SH are converted

to H2O and CH3OH respectively. The calculations were performed for a period of 4 years of

continuous processing. Also to reduce pressure drop in the bed, three parallel beds are used

to convert the sulphuric compounds.

COS conversion

The COS conversion is performed according to [18, 33, 34]

in a zinc promoted -alumina

bed. This bed is of porous material with a specific area of 300 m

2

/g. [18, 33,34] stated the

following relation between the rate of conversion and the concentration of reactants, COS

and H

2

O. 







1 2 2

[COS]

Rate

k

1 k

[H O]

[17.2.1]

Data for the kinetics were present at 60

o

C and initial concentration of 150 ppm COS these

were used in the calculations. Using these data and the ingoing flow composition the

minimal bed volume is calculated, assuming a reduction of COS from 39.5 ppm to 10 wppb

to ethylene.

Reaction conditions

Temperature 60

o

C

Pressure 25

Bar

Feed stream per unit

480.72 kmol/hr

Initial COS stream

0.02 kmol/hr

Initial H

2

O stream

4.44 kmol/hr

Initial H

2

S stream

0.04 kmol/hr

Initial CO

2

stream

9.99 kmol/hr

Kinetics data

k

1

3.359655

-

k

2

0.001021

-

[H

2

O]

0

9229

ppm

[COS]

0

39.5 ppm

Basic design values of the reactor

Porosity 0.6

-

Density of -alumina

3970 kg/m

3

Diameter 2.75

m

Taken from calculation for H2S/CH3SH converter

Output concentration of COS

10 wppb

To ethylene

Bed volume of one unit

0.011 m

3 (1)and according to eq.8.21(1) Because of this small volume a bed height of 0.30 m is chosen.

Conversion 0.999993

-

Heat of reaction



0

H

f,COS

-142,000 kJ/kmol



0

H

f,H2O

-241,800 kJ/kmol



0

H

f,CO2

-393,500 kJ/kmol



0

H

f,H2S

-20,600 kJ/kmol

c

p,mixture

37.83 kJ/(kmolK)



T

H

f,COS

-140,510 kJ/kmol



T

H

f,H2O

-240,621 kJ/kmol

(36)

A36

The output results are given in table A17.2.1.

H

2

S and CH

3

SH conversion

A zinc oxide bed is used to convert both H

2

S as CH

3

SH. To calculate the bed volume the

kinetics at 60

o

C are used, which are given by [19] An assumption is made that CH

3

SH reacts

at the same rate as H

2

S as follows:



2



2 2 3 H O H O p H S CH SH

P

K

P

[17.2.1]

Using the partial pressures and equations given in chapter 8 the bed volume can be

calculated.

Reaction conditions

Temperature 60

o

C

Pressure 24.99

bar

Feed stream per unit

480.72 kmol/hr

Initial H

2

O stream

4.42 kmol/hr

Initial H

2

S stream

0.061 kmol/hr

Initial CH

3

SH stream

0.0022

kmol/hr

Initial CH

3

OH stream

0 kmol/hr

Initial partial pressure H

2

O 0.23

bar

Initial partial pressure H

2

S 0.032

bar

Initial partial pressure CH

3

SH 0.00012

bar

Kinetics data

K

p

(333K)

1.5810

11

bar/bar

Basic design values of the reactor

Porosity 0.4

-

Maximal sulphur loading

0.4 kg/kg

Density of zinc oxide

5606

kg/m

3

Bed volume of one unit

48.6 m

3

Conversion >0.9999

-

H/D-ratio 3

-

Heat of reaction



0

H

f,H2S

-20,600 kJ/kmol



0

H

f,ZnO

-350,500 kJ/kmol



0

H

f,H2O

-241,800 kJ/kmol



0

H

f,ZnS

-192,600 kJ/kmol



0

H

f,CH3SH

-22,300 kJ/kmol



0

H

f, CH3OH

-201,500 kJ/kmol

c

p,mixture

37.84 kJ/(kmolK)



T

H

f,H2S

-19,310 kJ/kmol



T

H

f,ZnO

-349,087 kJ/kmol



T

H

f,H2O

-240,619 kJ/kmol



T

H

f,ZnS

-190,987 kJ/kmol



T

H

f,CH3SH

-20,498 kJ/kmol



T

H

f,CH3OH

-199,914 kJ/kmol

H

Reaction,H2S

-63,210 kJ/kmol

H

Reaction,CH3SH

-21,317 kJ/kmol

Temperature increase

0.22

o

C

Pressure drop

p

0.80 bar

Using

Ergun

equation

(37)

TABLE A17.2.1: SUMMARY OF INGOING AND OUTGOING FLOWS OF THE COS CONVERTER

INPUT

<5>

Mole Flow

kmol/hr

INPUT per

unit

<5a>/3

Mole Flow

kmol/hr

OUTPUT

<6>

Mole Flow

kmol/hr

H

2

232.28 H

2

77.43 H

2

232.28 CH

4

500.43 CH

4

166.81 CH

4

500.43 C

2

H

6

215.68 C

2

H

6

71.89 C

2

H

6

215.68 C

2

H

4

85.64 C

2

H

4

28.55 C

2

H

4

85.64 C

3

H

8

9.52 C

3

H

8

3.17 C

3

H

8

9.52 C

3

H

6

4.23 C

3

H

6

1.41 C

3

H

6

4.23 i-C

4

H

8

0.97 i-C

4

H

8

0.32 i-C

4

H

8

0.97 n-C

4

H

8

0.04 n-C

4

H

8

0.01 n-C

4

H

8

0.04 i-C

4

H

10

4.23 i-C

4

H

10

1.41 i-C

4

H

10

4.23 n-C

4

H

10

0.78 n-C

4

H

10

0.26 n-C

4

H

10

0.78 2-C

5

H

12

0.21 2-C

5

H

12

0.07 2-C

5

H

12

0.21 CO

90.88 CO

30.29 CO

90.88 N

2

253.79 N

2

84.60 N

2

253.79 CO

2

29.98 CO

2

9.99 CO

2

30.04 H

2

O 13.31

H

2

O 4.44

H

2

O 13.25

COS

0.06 COS

0.02 COS

0.00 CH

3

SH 0.01

CH

3

SH 0.00

CH

3

SH 0.01

H

2

S 0.13

H

2

S 0.04

H

2

S 0.18

C

3

H

6

O 0.00

C

3

H

6

O 0.00

C

3

H

6

O 0.00

n-C

4

H

8

O 0.00

n-C

4

H

8

O 0.00

n-C

4

H

8

O 0.00

i-C

4

H

8

O 0.00

i-C

4

H

8

O 0.00

i-C

4

H

8

O 0.00

3-C

5

H

10

O 0.00

3-C

5

H

10

O 0.00

3-C

5

H

10

O 0.00

2,2-C

5

H

10

O 0.00

2,2-C

5

H

10

O 0.00

2,2-C

5

H

10

O 0.00

n-C

5

H

10

O 0.00

n-C

5

H

10

O 0.00

n-C

5

H

10

O 0.00

2-C

5

H

10

O 0.00

2-C

5

H

10

O 0.00

2-C

5

H

10

O 0.00

CH

3

OH 0.00

CH

3

OH 0.00

CH

3

OH 0.00

Total

1442.16

480.72 1442.16

(38)

A38

Appendix 17.3: Equipment calculation of the hydroformylation

reactor

(39)

(40)

A40

(41)

(42)

A42

(43)

(44)

(45)

Appendix 18: Equipment summary specifications sheet

EQUIPMENT NR. : NAME :

C301

Off gas stripper

C302 Distillation column R101 COS and H2S converter R201 Reactor Pressure [bara] : 22.00 / 22.02 1.0 / 1.2 24.2/25.0 24.2/24.0 Temp. [oC] : 35.0 / 172.3 46.2 / 81.8 60.0 100.0 Volume [m3] : Diameter [m] : L or H [m] : 1.6 0.5 8.0 0.3 0.2 20.5 50.0 2.75 8.52 39.4 2.56 7.67 Internals - Tray Type : - Tray Number : - Fixed Packing Type : Shape : - Catalys Type : Shape : Sieve trays 5 n.a. n.a. n.a. n.a. Sieve trays 30 n.a. n.a. n.a. n.a. n.a. n.a. ZnO palet -alumina spherical n.a. n.a. n.a. n.a. Rh-TPP Sol. complex Number - Series : - Parallel : 1 1 3 1 Materials of Construction(1)_: _{CS CS CS CS} Remarks: (1) CS = Carbon Steel EQUIPMENT NR. : NAME : V201A Flash R201 V201B Demister R201 V301 Reflux accumulator C301 V302 Reflux accumulator C302 Pressure [bara] : 24.1 24.1 22.0 1.01 Temp. [oC] : 100 100 34.8 44.4 Volume [m3] : Diameter [m] : L or H [m] : 1.18 0.79 0.94 1.02 0.81 0.81 6.20 1.58 3.16 3.17 1.26 2.53 Internals - Tray Type : - Tray Number : - Fixed Packing Type : Shape : - Catalyst Type : Shape : n.a. n.a. CS plate n.a. n.a. Number - Series : - Parallel : 1 1 1 1 Materials of Construction(1): CS CS CS CS Remarks: (1) CS = Carbon Steel

(46)

A46

EQUIPMENT NR. : NAME : Stage 1 Compressor Stage 2 Compressor

Jacket reactor Coil reactor Substance

- Tubes :

- Shell :

Low Joule Gas cooling water

Low Joule Gas cooling water Tetraglyme cooling water Cooling water tetraglyme Duty [kW] : 658 193 900 2096 Heat Exchange area [m2] : 109 10 61 62 Number - Series : - Parallel : 1 1 1 1 Pressure [bara] - Tubes : - Shell : 7.6 4 25 4 4 24 24 4 Temperature In / Out [o_C] - Tubes : - Shell : 200/35 30/54 207/141.5 30/54 30/54 100/100 100/100 30/54 Materials of Construction(1): CS CS CS CS Other : Remarks: (1) CS = Carbon Steel. EQUIPMENT NR. : NAME : E201 Pre-heater R201 E301 Effluent cooler R201 E302 Condenser C301 E303 Reboiler C301 Substance - Tubes : - Shell : Reactor feed steam hp Effluent reactor cooling water Off-gas cooling water Aldeh./water steam hp Duty [kW] : 527 626 810 464 Heat Exchange area [m2] : 21 116 155 5.5 Number - Series : - Parallel : 1 1 1 1 Pressure [bara] - Tubes : - Shell : 24.2 41 24 4 22 4 22 41 Temperature In / Out [oC] - Tubes : - Shell : 67.6/100 252/252 99.3/60 30/54 34.8/34.8 27/30 171/171 41/41 Materials of Construction(1): CS CS CS CS Other : Remarks: (1) CS = Carbon Steel

(47)

EQUIPMENT NR. : NAME : E304 Bottom cooler E305 Condenser C302 E306 Reboiler C302 Substance - Tubes : - Shell : Bottom stripper cooling water Propion aldehyde cooling water Higher aldehyde steam hp Duty [kW] : 346 883 876 Heat Exchange area [m2] : 11 110 5 Number - Series : - Parallel : 1 1 1 Pressure [bara] - Tubes : - Shell : 22 4 1 atm. 4 1 atm. 41 Temperature In / Out [o_C] - Tubes : - Shell : 171/50 30/54 44.4/44.4 27/30 171/171 252/252 Materials of Construction(1): CS CS CS Other : Remarks: (1) CS = Carbon Steel EQUIPMENT NR. : NAME : K101 LJG 2-stage compressor P201 Solvent recycle R201 P301 Reflux C301 P302 Reflux C302 Type : Number : Reciprocating 1 Centrifugal 1 Centrifugal 1 Centrifugal 1 Medium

transferred : Low Joule Gas Solvent Reflux C301 Reflux C302

Capacity [kg/s] : [m3/s] 2.65 1.21 0.01 1.01E-5 0.36 0.02 0.78 0.001 Density [kg/m3] : 2.19 738.70 18.44 799.92 Pressure [bara] Suct. / Disch. : 2.3 / 25.0 24.1 / 24.2 22.0 / 22.1 1.013 / 1.019 Temperature In / Out [oC] : 40 / 35 (1) 101.1 / 101.1 34.9 / 34.9 44.4 / 47.1 Power [kW] -Theory : -Actual : 934 1,075 0.00034 0.20 0.36 0.00094 0.0019 Number -Theory : -Actual : 1 2 1 2 1 2 1 2 Special materials of Construction (2) : C.S. C.S. C.S. C.S. Other : Remarks:

(1) With interstage cooling (2) C.S. = Carbon steel

(48)

A48

EQUIPMENT NR. : NAME : PV301 Depressurisation for C301 PV302 Depressurisation for C302 Type : Number : Centrifugal 1 Centrifugal 1

Medium transferred : Reactor effluent Bottom stripper

Capacity [kg/s] : [m3/s] : 8.17 0.45 0.964 0.002 Density [kg/m3_{] : 18.37}_628.86 Pressure [bara]

Suct. / Disch. : 24.0 / 22.0 22.0 / 1 atm.

Temperature In / Out [oC] : 100.0 / 99.3 50.0 / 50.0 Power [kW] -Theory : -Actual : - - - - Number -Theory : -Actual : 1 1 1 1 Special Materials of Construction (1) : C.S. C.S. Other : Remarks: (1) C.S. = Carbon steel

(49)

EQUIPMENT NUMBER : C301

NAME

: Off gas stripper

General Data

Service :distillation / extraction / absorption

Column Type :packed / tray / spray

Tray Type :cap / sieve / valve

Tray Number (1)

 Theoretical :8

 Actual :8

 Feed (actual) :1

Tray Distance (HETP) [m] :0.5 Tray Material : AISI 410 S (2) Column Diameter [m] :0.25 Column Material : C.S. (2) Column Height [m] :9.00

Heating (3) :none / open steam / reboiler Process Conditions

Stream Details Feed Top Bottom Reflux / Absorbent Temp. [oC] Pressure [bara] Density [kg/m3_] Mass Flow [kg/s] 60 22.0 19 8.2 35 22.0 18 7.2 171 22.0 598 1.0 35 22.0 598 0.9

Composition mol% wt% mol% wt% mol% wt% mol% wt%

Hydrogen 11.8 1.0 12.4 1.2 0.0 0.0 12.4 1.2 Methane 39.4 27.3 41.4 30.9 0.0 0.0 41.4 30.9 Ethane 17.0 22.0 17.9 25.0 0.0 0.0 17.9 25.0 Ethylene 0.6 0.7 0.6 0.8 0.0 0.0 0.6 0.8 Propane 0.7 1.4 0.8 1.6 0.0 0.0 0.8 1.6 Propylene 0.0 0.1 0.0 0.1 0.0 0.0 0.0 0.1 Iso-butylene 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 N-butylene 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 Iso-butane 0.3 0.8 0.3 0.8 0.9 1.0 0.3 0.8 N-butane 0.1 0.2 0.0 0.1 0.3 0.3 0.0 0.1 Methyl-butane 0.0 0.1 0.0 0.0 0.2 0.2 0.0 0.0 Carbon monoxide 0.6 0.8 0.7 0.9 0.0 0.0 0.7 0.9 Nitrogen 20.0 24.2 21.0 27.4 0.0 0.0 21.0 27.4 Carbon dioxide 2.4 4.5 2.5 5.1 0.0 0.0 2.5 5.1 Water 0.5 0.4 0.1 0.1 7.7 2.5 0.1 0.1 Carbonyl sulfide 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 Methyl mercaptan 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 Hydrogen sulfide 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 Propion aldehyde 6.1 15.4 2.2 5.8 83.9 86.8 2.2 5.8 N-butyraldehyde 0.3 0.9 0.0 0.1 5.3 6.8 0.0 0.1 Iso-butyraldehyde 0.0 0.0 0.0 0.0 0.2 0.2 0.0 0.0 3-meth.but.ald. 0.1 0.3 0.0 0.0 1.3 2.0 0.0 0.0 Dimeth.prop.ald. 0.0 0.0 0.0 0.0 0.0 0.1 0.0 0.0 Valeraldehyde 0.0 0.0 0.0 0.0 0.1 0.1 0.0 0.0 2-meth.but.ald. 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 Methanol 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 Column Internals Trays (4)

Number of caps / sieve holes/ :143

Active Tray Area [cm2] : 9.5

Weir Length [mm] : 90

Diameter of

Chute/ pipe /hole [mm] : 2.9

Packing N.A. Type : Material : Volume [m3_{] :} Length [m] :

Width [m] : Height [m] :

Remarks: (1) Tray numbering from top to bottom. (3) Reboiler is E303; operates with HP steam (40 barg).

(50)

A50

EQUIPMENT NUMBER : C302

NAME : Distillation column

General Data

Service :distillation extraction absorption

Column Type :packed tray spray

Tray Type :cap sieve valve

Tray Number (1)

 Theoretical :28  Actual :28  Feed (actual) :16

Tray Distance (HETP) [m] :0.5 Tray Material : AISI 410 S (2)

Column Diameter [m] :0.23 Column Material : C.S. (2) Column Height [m] :20.00

Heating (3) :none open steam reboiler Process Conditions

Stream Details Feed Top Bottom Reflux / Absorbent Temp. [oC] Pressure [bara] Density [kg/m3_] Mass Flow [kg/s] 49 1.1 181 1.0 44 1.0 624 0.9 78 1.2 676 0.1 44 1.0 676 0.8

Composition mol% wt% mol% wt% mol% wt% mol% wt%

Hydrogen 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 Methane 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 Ethane 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 Ethylene 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 Propane 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 Propylene 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 Iso-butylene 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 N-butylene 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 Iso-butane 0.9 1.0 1.0 1.1 0.0 0.0 1.0 1.1 N-butane 0.3 0.3 0.4 0.4 0.0 0.0 0.4 0.4 Methyl-butane 0.2 0.2 0.2 0.3 0.0 0.0 0.2 0.3 Carbon monoxide 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 Nitrogen 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 Carbon dioxide 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 Water 7.7 2.5 8.3 2.7 0.0 0.0 8.3 2.7 Carbonyl sulfide 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 Methyl mercaptan 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 Hydrogen sulfide 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 Propion aldehyde 83.9 86.8 90.0 95.4 10.9 8.7 90.0 95.4 N-butyraldehyde 5.3 6.8 0.0 0.0 68.7 67.7 0.0 0.0 Iso-butyraldehyde 0.2 0.2 0.0 0.0 1.7 1.6 0.0 0.0 3-meth.but.ald. 1.3 2.0 0.0 0.0 17.4 20.4 0.0 0.0 Dimeth.prop.ald. 0.0 0.1 0.0 0.0 0.5 0.6 0.0 0.0 Valeraldehyde 0.1 0.1 0.0 0.0 0.8 0.9 0.0 0.0 2-meth.but.ald. 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 Methanol 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 Column Internals Trays (4) Number of

Caps / sieve holes : 115 Active Tray Area [cm2_{] :16.0}

Weir Length [mm ] :90

Diameter of

chute pipe / hole [mm] :4.2

Packing N.A. Type : Material : Volume [m3_{] :} Length [m] : Width [m] : Height [m] : Remarks:

(1) Tray numbering from top to bottom.

(2) CS = Carbon Steel. AISI 410 S = Type of stainless steel (3) Reboiler is E306; operates with 40 barg steam.

(51)

EQUIPMENT NUMBER : R101

In Series :1

NAME

: COS / H

2

S reactor

In Parallel :3

General Data

Service :- Buffer / Storage / Separation / Reaction

Type :- Packed Bed

Position :- Horizontal - Vertical

Internals :- Demister / Plate / Coil

Heating/Cooling medium :- none / Open / Closed / External Hxgr - Type :n.a.

- Quantity [kg/s] :n.a. - Press./Temp.’s [bara/o_{C] :n.a.}

Vessel Diameter (ID) [m] :2.75

Vessel Height [m] :8.52 Vessel Tot. Volume [m3] :50.0

Vessel Material :Carbon steel

Other :

Process Conditions

Stream Data Feed Top Bottom

Temperature [oC] Pressure [bara] Density [kg/m3] Mass Flow [kg/s] 60.0 25.0 18.9 8.18 67.6 24.2 20.5 8.17 n.a. n.a. n.a. n.a.

Composition mol% wt% mol% wt% mol% wt%

(1) COS H2S CH3SH H2O CH3OH 0.0040 0.0089 0.0005 0.4014 0.0000 0.0116 0.0148 0.0011 0.3525 0.0000 2.7E-8 1.2E-10 4.6e-12 0.4103 0.0005 8.1E-8 2.1E-10 1.1E-11 0.3603 0.0007 Remarks:

(52)

A52

EQUIPMENT NUMBER : R201

In Series

:1

NAME

: Hydroformylation reactor

In Parallel

:1

General Data

Service :- Buffer / Storage / Separation / Reaction

Type :- Continuous/stirred tank reactor/gas/liquid phases

Position :- Horizontal-/Vertical

Internals :- Demister / Plate / Coil / Jacket

Heating/Cooling medium :- None / Open / Closed / External Hxgr

- Type :Cooling water jacket and coil

- Quantity [kg/s] :0.0299

- Press./Temp.’s [bara/o_C] _{:4 bara, 27-54}o_C

Vessel Diameter (ID) [m] :2.55

Vessel Height [m] :7.65

Vessel Tot. Volume [m3] :40

Vessel Material :Carbon Steel

Other :

Process Conditions

Stream Data Feed Top Bottom

Temperature [oC] Pressure [bara] Density [kg/m] Mass Flow [kg/s] 100.0 24.2 16.17 8.17 100.0 24.0 18.47 8.17 n.a. n.a. n.a. n.a.

Composition mol% wt% mol% wt% mol% wt%

(1) H2 CO C2H4 C3H6 propion aldehyde n-butanal i-butanal 11.78 0.64 0.61 0.03 6.14 0.29 0.01 1.02 0.77 0.73 0.05 15.37 0.91 0.03 16.19 6.34 5.97 0.29 0.00 0.00 0.00 1.59 8.65 8.16 1.43 0.00 0.00 0.00 Remarks: