cpd nr 3334 conceptual process design process systems

CPD NR 3334 Conceptual Process Design

Process Systems Engineering

DelftChemTech - Faculty of Applied Sciences Delft University of Technology

Subject

Design of a process to manufacture ethylene from ethane by means of a shock wave reactor

Authors (Study nr.) Telephone

Jurrian van der Dussen 1195166 06-41030220 Alan Farrelly 1184881 06-28235728 Gerold Kort 1115545 06-17626279 Vincent Twigt 1184628 06-18726719 Hao Weng 1158856 06-28188702

Keywords

Ethane, Ethylene, Shock wave, Pyrolysis

Assignment issued : 21/03/2006

Report issued : 02/06/2006 Appraisal :

Final Report Shock Wave Reactor CPD 3334

Table of Contents

Acknowledgement v

Summary vi List of abbreviations vii Quantities and their dimensions viii 1 Introduction 1

1.1 Background 1 1.2 Thermal cracking 2 1.3 Shock Wave Reactor 2 1.4 Comparison 4 1.5 Requirements 5 1.6 Approach 5

2 Criteria and assumptions 6

2.1 Criteria 6 2.1.1 Product quality 6 2.1.2 Location 7 2.1.3 By-products 7 2.1.4 Legislation 8

2.2 Assumptions 9 2.2.1 Feed quality 9 2.2.2 Reactions 10 2.2.3 Kinetics 12 2.2.4 Product destination 13

i


3 Overall mass balance 14

3.1 Ethylene production 14 3.2 By-products 14 3.3 In- and Outgoing streams 15

4 Process Scheme 16

4.1 I/O-diagram 16 4.2 Recycle diagram 17 4.3 Separation diagram 17

5 Reactor 19

5.1 Acceleration section 19 5.2 Mixing section 21 5.3 Pyrolysis section 24

5.3.1 Ideal gas? 24 5.3.2 Computational work 24

5.4 Assumptions 30 5.4.1 Residence time 30 5.4.2 Widening angle 30 5.4.3 Pre-shock temperature 30 5.4.4 Pre-shock pressure 31 5.4.5 Pre-shock velocity 32 5.4.6 Summary 33 5.4.7 Results 34

5.5 Reactor Dimensions 37

ii


6 Separation 38

6.1 Separation technology 38 6.1.1 Membrane 38 6.1.2 Distillation 39 6.1.3 Cryodistillation 39 6.1.4 Absorbers 39

6.2 Order of separation 40 6.2.1 Components 40 6.2.2 Separation sequencing 41

6.3 Simulation of the process 41 6.3.1 Water separation 42 6.3.2 Benzene\Water separation 43 6.3.3 H2S, CO and CO2 removal 43 6.3.4 Dryer 44 6.3.5 Hydrogenation of acetylene 45 6.3.6 Demethanizer 46 6.3.7 Hydrogen/Methane separation 46 6.3.8 Product separation 48 6.3.9 Deethanization 48

7 Heat & Power Integration 49

7.1 Heat 49 7.2 Power 50

8 Economics 51

8.1 Purchased equipment cost 51 8.2 Cost estimation for raw materials 52 8.3 Determining the cost of utilities 53 8.4 Labour cost for the SWR-plant 53 8.5 Capital cost of the SWR-plant 54 8.6 Economic evaluation of the SWR-plant 55

iii


9 Safety 57

9.1 Fire & Explosion index 57 9.1.1 Boundary 57 9.1.2 Material Factor 57 9.1.3 General process hazards 58 9.1.4 Special process hazards 59

9.2 Fire protection and prevention 61 9.2.1 Leak prevention 61 9.2.2 Leak detection 61 9.2.3 Leak dispersion, containment 62 9.2.4 Miscellaneous 62

10 Controllability 63

10.1 Inlet streams 64 10.2 Shock wave position 65 10.3 Emergency control of the SWR reactor 66 10.4 Separation of water 67 10.5 Water discharge 68 10.6 Distillation columns 69

10.6.1 Control in the top of the column 70 10.6.2 Control in the bottom of the column 70

10.7 Fixed bed reactor 71 10.8 Membrane 72 10.9 Ethane Purge 73

11 Conclusions and recommendations 74

11.1 Conclusions 74 11.2 Recommendations 75

Literature I Appendix Index III

iv


Acknowledgement As a group we would like to express our thanks to certain people, who helped us during

the project. At first we would like to thank the project supervisors from the TU Delft.

Prof. J. Grievink (Technical Supervisor)

Ir. M.W.M. van Goethem (Project Principal)

Ir. J. Nijenhuis (Creativity and Group Process Coach)

We also would like to thank the following people for there contribution during this project:

Dr. Ir. C.S. Bildea (TU Delft)

Prof. Dr. F. Kapteijn (TU Delft)

Dr. Ir. M. Makkee (TU Delft)

Dr. Ir. S.M. Lemkowitz (TU Delft)

Ir. A.van Miltenburg (TU Delft)

Ir. C. Dell’Era (Helsinki University of Technology)

P.D.De Carvalho Falcao (Msc-student TU Delft)

R. Bosma (Msc student TU Delft)

v


Summary The aim of this project was to evaluate the possibility of building an economically viable

SWR-plant while conforming to predetermined constraints and criteria.

Globally, 117-million t/a ethylene is produced. The plant designed produces 1 Mt/a

ethylene. The feedstock available is provided from neighbouring ethane-producing

facilities. Information about SWR technology was provided through a patent issued by

the project supervisor. This technology is relatively new and there are no known

operating petrochemical plants utilizing this.

The SWR plant designed has an annual runtime of 8400 hours. The total investment is

775 million US dollars and has an economical lifespan of 10 years, after which an

estimated profit of 520 million dollars is made.

Compared to current thermal cracking processes, SWR technology achieves a higher

conversion and selectivity. Also, the model used for designing the SWR is adaptable to

different kinds of chemical reactions affiliated with ethane pyrolysis.

vi


List of abbreviations /a per annum (year)

DACE Dutch Association of Cost Engineers

HEN Heat Exchanger Network

F&EI Fire and Explosion Index

M$ Million dollar

MEA Monoethanolamine

MF Material Factor

Mt Million tonnes

NCF Net Cash Flow

ODE Ordinary Differential Equation

PFS Process Flow Scheme

ppb parts per billion

ppm parts per million

SWR Shock Wave Reactor

wt% Weight percentage

vii


Quantities and their dimensions α Angle [°]

A Area [m2]

Ar Arrhenius constant [1/s]

ci Concentration component I [mol/m3]

cj Concentration of component in reaction j [mol/m3]

Cp Specific heat (cst P) [J/mol K]

Cv Specific heat (cst V) [J/mol K]

dn Nozzle spacing [m]

D Diameter [m]

∆H Heat of formation [kJ/kg]

Ea Activation energy [J/kg]

f Friction factor [-]

Fi Flow rate [m3/s]

k Rate constant [1/s]

κ Ratio [-]

Mw Molecular Weight [kg/mol]

n Number of moles [mol]

Nf Flammability [-]

Nn Number of nozzles [-]

Nr Reactivity [-]

η Dynamic viscosity [Pa s]

Pc Critical pressure [Pa]

Pi Pressure [Pa]

Poc Pressure carrier fluid [Pa]

Pof Pressure feedstock [Pa]

ri Reaction rate [mol/m3 s]

R Gas constant [J/mol K]

Re Reynolds number [-]

ρ Density [kg/m3]

T Temperature [K]

Tc Critical temperature [K]

ushock Shock velocity [m/s]

X Mixing distance [m]

V Volume [m3]

Zc Critical compressibility factor [-]

viii


1 Introduction For this project ethylene is to be produced from ethane feedstock, by means of the

relatively new shock wave reactor technology1. The production rate needs to be 1 million

tonnes per annum (Mt/a). This comes down to a production of 33 kg/s, assuming an

annual production time of 8400 hours. The economical potential for this type of process

is 520 million dollar (M$) after 10 years.

1.1 Background With an annual world production of over 117 million tonnes, ethylene is, in volume, the

largest organic chemical product2. Thermal cracking units produce the most significant

part of this ethylene. However, this cracking process is an energy consuming and

capital-intensive process.

Ethylene is an intermediate and is used to produce a final product, e.g. polyethylene,

before made into a consumer product. In Figure 1-1 the position of the reactor in the

total supply chain is given.

Consumer

product

Final Product Ethylene

Ethane

Shock Wave

Reactor

C2/C3

separator Separation

Distillation

Refinery

Operations

Ethane/

propane

Propane

Natural gas

Crude oil

Figure 1-1: Chain of supply

1 Hertzberg, A., et al, “Method for initiating pyrolysis using a shock wave”, US Patent 5,300,216,

1994 2 Grievink, J., “Project Objectives & Description”, TU Delft, 2006

1


1.2 Thermal cracking Conventional thermal cracking units are highly energy intensive. The energy required

during the pyrolysis in the reactor is supplied by preheating the steam-ethane mixture.

Thermal energy created in the furnace is converted into thermal energy in the mixture.

However, the temperature of the mixture is above the pyrolysis temperature, of

approximately 1100 K. Because of this some reactions already occur prior to the reactor,

causing coke formation. Increasing the steam/ethane ratio can reduce this to a certain

extent.

1.3 Shock Wave Reactor In 1993 Hertzberg et al.1, proposed to use gas dynamics to supply the energy to the

reactor, which is less energy intensive, compared to the thermal cracking unit. This is

done in a so-called shock wave reactor (SWR), shown in Figure 1-2.

Figure 1-2: Shock wave reactor

Using gas dynamics, the available kinetic energy is converted into thermal energy. This

is initiated by expanding the cross-sectional area of the reactor, decelerating the

supersonic steam-ethane mixture present.

2


The steam-ethane mixture, with supersonic velocity, ‘bumps’ into the decelerated steam-

ethane mixture converting the kinetic energy into thermal energy. This increase of

thermal energy is expressed by an increase of temperature, raising it above the pyrolysis

temperature, inducing thermal cracking.

Acceleration of the super heated steam to supersonic velocity, which is needed as a

carrier fluid, is done with the use of a jet tube. It increases the velocity of the super

heated steam from Mach 0.9 to almost Mach 3. Knowing that the Mach speed is

approximately 330 m/s it can be said that the velocity is about 1000 m/s.

The mixing that occurs in the reactor is done below the reaction (pyrolysis) temperature

of approximately 1100 K, fully mixing the steam and ethane prior to the reaction section.

Because this mixing occurs at a temperature below the pyrolysis temperature, coke

formation is reduced, compared to the thermal cracking reactor. However, compared to

the thermal cracking reactor an extra amount of steam is needed. This extra steam acts

as a buffer for the increase in temperature during pyrolysis and is needed to achieve

perfect mixing.

The reactor conditions, which are going to be used during the modelling of the reactor,

are taken from the patent1, mentioned before. It must be noted however that these

conditions are only an indication and will only be used in order to check the validity of the

results obtained.

3


1.4 Comparison As stated before the goal is to produce ethylene from ethane using shock wave

technology. This technology is chosen because important parameters, shown in Table

1-1, are better than that of the conventional thermal cracking process.

Table 1-1: Thermal cracking process versus SWR-technology2

Thermal Cracking SWR-Parameter Furnace Technology

Selectivity (%) 85 90Ethane conversion (%) 65 70Yield Ethylene (%) 55 63Energy requirement (kJ/kg ethylene) 57500 26300 The energy requirement for the thermal cracking furnace is taken from ECN3. The SWR

energy requirement is taken from this project.

Because of the reduced coke formation the SWR can stay on-stream longer than the

thermal cracking reactor. However, because of the novelty of the SWR, it is difficult to

say how long it will take for the reactor to achieve stable operation.

The position of the shock wave is controlled by means of an expander. Increasing, or

reducing, the speed, at with which the gas is expanded, makes this possible. This helps

to reduce the time to reach stable operation.

Varying the amount of steam entering the reactor, controls the temperature inside the

reactor. A slight change in temperature changes the product distribution, as will be

shown in paragraph 5.4. Therefore caution is needed when altering this variable.

Because this variable is easy to control:

• The stability of the process is increased

• The time needed to reach stable operation is reduced.

All in all it is hard to say how long the reactor can stay on-stream continuously and how

long it needs to run at stable operation. However the above-mentioned factors surely

affect the time, on-stream and at which stable operation is achieved, in a positive way.

3 Gielen, D.J., Vos, D., van Dril, A.W.N., “The petrochemical industry and its energy use

prospects for the Dutch energy intensive industry“, ECN-C—96-029, 1996

4


1.5 Requirements To achieve this goal certain requirements and boundaries are set by the supervisor and

principal, which are2:

• The process must be economically viable

• Product purity must be 99.9 wt%

• Annual ethylene production of 1 million tonnes

• No major changes in the local ecosystems are allowed

• Process materials should be recovered and recycled to maximum extent

• Energy consumption must be minimised

• Process must be safe and controllable

• Process has to be accepted by the US society and local communities

1.6 Approach To meet the requirements, the following steps will be taken:

• Generate an overall mass balance

• Set the production requirements in order to dimension the reactor

• Globally design the separation section

• Calculate the energy consumption of all the plant units

• Calculate an efficient recycle of water and ethane

• Integrate the power and heat of different units

• Design a control scheme

• Test for economic viability

This project is a conceptual design. Therefore it is decided not to design the SWR, the

piping and building structures in detail.

5


2 Criteria and assumptions This chapter states and discusses the basic assumptions and criteria that are set for this

project. In order to design the plant, some criteria have to be met, which are:

• Product quality

• Plant location

• By-product concentration

• Legislation

Based on these criteria, assumptions have to be made. Assumptions have to be made

on:

• Feed quality

• Occurring reactions

• Product take off

2.1 Criteria

2.1.1 Product quality

As stated before the product purity, in this case that of ethylene, must be 99,9 wt%.

Therefore the stream leaving the battery limit can be stated. The pressure and

temperature of the stream are chosen. These parameters, as for the stream quality, are

shown in

Table 2-1. They are set in such a way that the ethylene, meets the market demand set

by the costumers.

Table 2-1: Ethylene product stream

Stream Name : EthyleneComp. Units Specification Additional Information

Available Design Notes (also ref. note numbers)Ethylene wt% 99.9 99.9By-Products wt% 0.1 0.1Total 100.0

Process Conditions and PriceTemp. K 303Press. Bara 10Phase V/L/S VPrice $/tonne 650

6


2.1.2 Location

The SWR-plant must be situated in the southern part of the United States of America,

near the Mexican Gulf2. The Mexican Gulf area is one of the largest oil producing areas

in the world.

It is chosen to ‘build’ the plant near the city of Houston, Texas. This city is a large

intersection in the American oil industry. It has a large harbour, which makes it possible

to transport the ethane and ethylene by water.

Figure 2-1: Geographical position of Houston, Texas

The plant can receive its ethane from neighbouring plants, oil platforms, rigs in the

Mexican Gulf and from the large oil fields in the northern part of Texas. Consequently

ethylene production can be maintained, at all times.

2.1.3 By-products

During the pyrolysis 0.5 wt% of benzene must be formed. This weight percentage is

based on the total weight leaving the reactor, excluding water. The amount is specified

in agreement with the project principal.

Other by-products that are present in the reactor are:

• CO \ CO2

• H2S

7


These other by-products are not to be modelled in the reactor section, but must be taken

into account using amounts specified in agreement with the project principal. The

specified amounts are shown in Table 2-2. Note that the concentration specified is that

of the stream leaving the reactor, excluding water.

Table 2-2: By-product specification

ComponentTotal

concentration DimensionCO / CO2 0.5 wt%H2S 50 ppm

2.1.4 Legislation

Texas legislation4 states that the concentration of benzene in water, that is going to be

discharged, may not exceed 0.05 mg/l (0.05 ppm or 50 ppb). This criterion has to be met

before the water can be discharged5.

4 http://www.capitol.state.tx.us/statutes/wa.toc.htm 5 http://www.texas.gov

8


2.2 Assumptions

2.2.1 Feed quality

During operation two feeds will enter the reactor, namely:

• Ethane

• Steam (water)

These streams have a certain purity and quality. Table 2-3 and Table 2-4 state the

quality of the streams entering the battery limits of the plant.

Table 2-3: Ethane inlet stream

Stream Name : EthaneComp. Units Specification Additional Information

Available Design Notes (also ref. note numbers)Ethane wt% 94-97 95 (1) (1) Values taken in consultation.Propane wt% 1-3 2.5 (1) with Principal.Methane wt% 1-2 2 (2)CO/CO2 wt% 0-0.5 0.5 (2) (2) As 'worst case' scenario,Sulphur ppm wt 50 50 (3)

(3) Contaminants not harmful forthe process. Compounds not

Total 100.0 included in mass balance.Process Conditions and Price

Temp. K 298Press. Bara 10Phase V/L/S VPrice $/tonne 150

Table 2-4: Water inlet stream

Stream Name : Water InletComp. Units Specification Additional Information

Available Design Notes (also ref. note numbers)Water wt% 100 100.0 (1) Contaminants not harmful forImpurities ppm wt 80 80.0 (1) the process. Compounds not

included in mass balance.Total 100.0

Process Conditions and PriceTemp. K 300Press. Bara 1Phase V/L/S LPrice $/tonne 4.8

9


2.2.2 Reactions

From literature6 it can be found that the reactions occurring during thermal cracking

follow the reaction mechanism of radical reactions. These radical reactions inhibit 3

steps, namely:

• Initiation

• Propagation

• Termination

Initiation is the cleavage of a C-C bond, leading to two radicals. In case of ethane they

lead to two methyl radicals.

After initiation the propagation occurs. Here the radical ‘attacks’ another molecule after

which a different molecule and a primary radical are created. This primary radical then

decomposes to its most stable form while rejecting a hydrogen radical.

The final reaction that occurs is the termination of the reactions due to the combining of

two radicals. Forming either one saturated molecule or one unsaturated and one

saturated molecule.

This sequence is shown in Figure 2-2.

Figure 2-2: Radical reaction stages, taken from Chemical Process Technology7

6 Sundaram, K.M, Froment, G.F, “Modelling of thermal cracking kinetics - I”, Chem. Eng. Sc.,

1977, Vol 32, pp 601-608

10


From Figure 2-2 it can be seen that the reactions will lead to large products if the

residence time is large. If all the radical reactions, found in Hidaka, et al.8, are taken into

account the list of reactions would be extremely large.

This leads to a difficult modelling of the SWR and the rest of the plant. Therefore another

approach for the reactions is necessary. In Sundaram et. al.6 the reactions for the

pyrolysis of ethane are seen as equilibrium reactions between each component. This

approach helps to model the SWR to such a level that it would approach the normal

cracking situation.

During pyrolysis, the ethane-cracking and other side reactions take place. All reactions,

except the last reaction (9), the formation of benzene, are taken from Sundaram et.al.6

The components in the reaction are all in the gas phase. The reactions are:

2 6 2 4 2

2 6 3 8 4

3 8 3 6 2

3 8 2 4 4

3 6 2 2 4

C H C H + H2 C H C H + CH

C H C H + HC H C H + CHC H C H + CH

→→→

(1)(2)(3)(4)(5)

(6)(7)(8)(9)

2 2 2 4 4 6

2 6 2 4 4

2 4 2 6 3 6 4

4 6 2 2 6 6 2

C H + C H C H2 C H C H + 2 CH

C H + C H C H + CHC H + C H C H + H

→→→

All components in gas phase

The values for the kinetic parameters of the benzene reaction are assumed. Because

the total reactor outflow of benzene must be 0.5 wt%, excluding water, the parameters

can be found by trail and error, while running the Matlab-script.

7 Moulijn, Jacob A., Makkee, Michiel, van Diepen, Annelies, “Chemical process technology”, John

Wiley & Sons Ltd, 2001 8 Hidaka, Y. et al, “Shock-tube and modeling study of ethane pyrolysis and oxidation”, Comb. and

Flame, 120, page 245-264, 2000

11


2.2.3 Kinetics

In this paragraph the rates of each reaction are defined. The rate expressions for each

individual reaction are6:

he number below the rate expression (r), correspond with the reaction stated. The

The rate constant (k), for each reaction, is calculated using the Arrhenius equation.

27 2 t

6 6 2t

t17 7

t2

t1 28 8 2

t2 2

7 8 t 3 9 t9 9 -92 2

t t

F F Pr =kF TR

PFr =kF TR

PFFr =kF TR

F F P F F Pr =k - kF TR F TR

2t 2 3 t1

1 1 -1 2t t

t22 2

t

3 t3 3

t

t44 4

t2

5 t 7 4 t5 5 -5 2

t t

P F F PFr =k -kF TR F TR

PFr =kF TRF Pr =kF TR

PFr =kF TR

F P F F Pr =k -kF TR F TR

T

number stated below the flow rates (F) corresponds with the following component:

1 Ethane 6 Propane2 Ethylene 7 Acetylene3 Hydrogen 8 Butadiene4 Methane 9 Benzene5 Propylene

R*Tn rk = A *e

a-E

he Ft stated in the rate expressions is the total flow of all components leaving the

T

reactor:

t 1 2 3 4 5F =F +F 6 7 8 steam+F +F +F +F +F +F +F

12


2.2.4 Product destination

After separation, all the (by-) products have a different destination. The ethylene, for

instance, purified to 99.9 wt%, will be sold to a refinery nearby, which produces the final

product.

The un-reacted ethane will be recycled back to the reactor in order to reduce raw-

material costs.

Hydrogen is an economically interesting by-product considering its high sales price (±

2700 $/tonne). It is therefore decided to sell the formed hydrogen.

Acids, produced during pyrolysis, will be removed. After separation these acids will be

sent to special treatment plants. Acids comprise the following:

• CO

• CO2

• H2S

The benzene will be treated as a waste stream. The amount produced is not

economically interesting to sell. Therefore costs will be taken into account to dispose of

this waste correctly.

All other by-products will be used as fuel, preheating the steam. This is done in order to

reduce the costs for the amount of gas needed. The following components are

considered by-products:

• Methane

• Acetylene

• Propane

• Propylene

• Butadiene

13


3 Overall mass balance Now that the criteria and initial assumptions are stated, mass balances are created and

calculated. These are calculated in order to have a global idea on the amount of

components entering and leaving the plant.

3.1 Ethylene production The main reaction taking place in the reactor, is that of ethane to ethylene, which can be

denoted as:

2 6 2 4 2C H C H + H

This is an equilibrium reaction so a conversion of 100% will never be achieved.

3.2 By-products Using the reaction stated, it can be assumed that the amount of hydrogen, molar based,

formed is equal to the amount of ethylene formed. This is an approximation because

hydrogen and ethylene also react with the by-products formed (see Paragraph 2.2.2).

Benzene is calculated using the specified amount of 0.5 wt%.

The exact amount of all by-products is calculated during reactor modelling in chapter 5.

Therefore an approximated total amount of by-products is stated.

14


3.3 In- and Outgoing streams Using the selectivity, conversion and data from the patent, Table 3-1 is obtained. The

calculations are enclosed in Appendix B. Note that due to round off, it may look like

mass is not conserved.

Table 3-1: Mass balance on in- and outgoing streams

IN OUT Name Mt/a kg/s Name Mt/a kg/s

Ethane 1.70 56.2 Ethylene 1 33.0Water 11.34 375.0 Ethane 0.51 16.9

Hydrogen 0.07 2.4Benzene 0.0085 0.3By-products 0.11 3.7Water 11.34 375.0

Total 13.0 431.2 Total 13.0 431.2

Recycles streams are not taken into account for now, because all the values are an

indication. The real reactor output is calculated and stated in Chapter 5.

Product losses are expected during component separation, lowering the recyclable

amount.

15


4 Process Scheme In this chapter a simplified process scheme is created. A process scheme helps to

understand the process. It clarifies the task of the whole process or the selected unit.

The next paragraphs discuss how the simplified process scheme is created, starting

from a ‘black-box’ model9.

4.1 I/O-diagram Initially, not much is known about the process. From the previous chapters it is known

what the process should do, so therefore a ‘black-box’ model can be made. This is

known as an Input/Output-diagram, I/O-diagram for short, which is shown in Figure 4-1.

This shows the in- and outgoing streams of the process. The product destination is

stated in paragraph 2.2.4.

H2, CO, CO2, H2S, CH4

Benzene

Process Water

By-products

Ethylene

Water

Ethane

Figure 4-1: I/O-Diagram

16

9 Douglas, J.M., ”Conceptual Design of Chemical Processes”, McGraw-Hill, New York, 1988


4.2 Recycle diagram The I/O diagram can be expanded to a recycle diagram. Un-reacted ethane is recycled.

Recycling some of the water could prove to be economically viable. The recycle diagram

is shown in Figure 4-2.

H2, CO, CO2, H2S, CH4 Ethane-recycle

Benzene

Water-recycle

Separation

System

Water

By-products

Ethylene

Water

Ethane

Reactor

System

Figure 4-2: Recycle diagram

4.3 Separation diagram The separation system stated in Figure 4-2 is very simplified. For the real separation

system consider Figure 4-3. The separation sequence presented, is based on following

reasoning:

1. Water is the most abundant component.

2. Acids exhibit a negative affect on separation equipment, due to corrosion.

3. Acetylene is converted into ethylene to avoid a large distillation column, due to

close boiling points.

4. Hydrogen and methane are the lightest components in the system.

5. Ethylene is separated because it’s the desired product.

6. Ethane is recycled and its purity as that of the fresh ethane feed.

7. The benzene concentration has to be less than 0.05 ppm in order to legally

discharge wastewater.

8. Hydrogen is separated from methane because of its economical value.

17


18

H2S, CO, CO2

Ethane

H2

CH4

Hydrogen

removal

Acetylene

conversion

Benzene Benzene

removal

Ethane-recycle

De-

Ethanizer

Product

Separation

De-

Methanizer

Acid

removal

Water-recycle

Water/

Benzene

removal

Water

By-products

Ethylene

Water

Reactor

System

Figure 4-3: Separation diagram


5 Reactor The SWR-reactor incorporates three different segments, namely:

• Acceleration section

• Mixing section

• Pyrolysis section

From an engineering point of view this makes it difficult to model this reactor as a whole,

therefore all segments are modelled individually.

The segments will be dealt in chronological order, thus from the entrance (acceleration

section) of the reactor to the mixing section and then finally the pyrolysis section.

In order to calculate the acceleration and the mixing section first the pyrolysis section

was modelled. Some parameters used in the first or second paragraph will be explained

in more detail in the paragraphs following.

5.1 Acceleration section To increase the velocity of the steam entering the mixing section a jet tube is used. It

increases the steam velocity from Mach 0.9 (± 300 m/s) to about Mach 3 (± 1000 m/s).

To calculate the decreased diameter of the tube, venturi tube calculations are used. This

is due to the fact that the jet tube inside the reactor acts as a venturi tube10. Therefore it

may also be assumed that energy dissipation of the jet tube acts as a polytropic

compressor11.

10 van Kimmenaede, Ir. A.J.M.,”Warmteleer voor technici”, 8e druk, Wolters-Noordhoff,

Groningen, 2001 11 Bos, Ir. G.A.,”Stromingsmachines”, 1e druk., Stenfert Kroese, Houten, 1997

19


During these calculations two positions (see Figure 5-1) were defined, namely:

• Before the narrowing of the tube (Position 1)

• In the middle of the narrowed tube (Position 2)

Figure 5-1: Positions inside venturi tube

The calculations for the jet tube are enclosed in Appendix. The results are stated in

Table 5-1. Those with an * are values taken from the patent.

Table 5-1: Acquired parameters for the jet tube

Parameter Value DimensionDensity 5.946 kg/m3

Diameter 1 0.52 mDiameter 2 0.286 mFlowrate 63.1 m3/sPressure drop* 2594 kPaVelocity 1* 0.9 MachVelocity 2* 2.97 Mach The angle going from diameter 1 to diameter 2 must be smaller than 25°. The angle

going from diameter 2 to the final diameter of the mixing section (0.98 m) must be

smaller than 8°12.

12 van den Akker, H.E.A., Mudde R.F.,”Fysische Transportverschijnselen I”, Tweede druk, DUP

Blue Print, Delft, 2003

20


Because of energy dissipation, occurring when using a jet tube, calculations are

performed in order to reach the desired steam temperature, needed in the mixing section

(710 K). The origin of this temperature will be discussed in the mixing section.

As stated before the jet tube acts as polytropic compressor. Therefore the following

formula is used to calculate the energy dissipation: κ-1κ

2 2

1 1

T P = T P

(Equation 1)

This results in an entering steam temperature of 1290 K. Calculations are enclosed in

Appendix D.

5.2 Mixing section Mixing at high-speed velocities requires a specific approach. At high velocities

experience is lacking and common sense is sometimes not adequate. Calculations used

for normal velocities are not accurate enough, and do not describe the situation

correctly.

Knowlen et. al13. experimented with shock tubes measuring the length it would take,

compared to nozzle spacing and differing pressure, to reach a perfectly mixed system.

This data is used for predicting the mixing length of the reactor. In order to check

whether this experimental data is useable, it is recommended that experiments are

conducted for this specific reactor.

13 Knowlen, C. et.al., “Petrochemical pyrolysis with shockwaves”, AIAA., 1995, 95-0402.

21


The experiments were conducted in a similar reactor setup as used in the SWR. The

venturi tube, used to accelerate the steam to supersonic speed in the SWR, is replaced

by a supply unit, which creates the velocity of the carrier gas. During the experiments

measurements were done to check where perfect mixing occurred. Carrier gas pressure

and nozzle spacing were varied to see the influence on the mixing length. These

influences are shown in Figure 5-2.

Figure 5-2: Mixing distance over nozzle spacing as a function of pressure ratio

The pressure ratio in Figure 5-2 is defined as the pressure of the feedstock (ethane)

over the pressure of the carrier fluid (steam). The pressure ratio during this project is set

according to the patent. Because steam and ethane both enter the mixing section at

1.02 bar the ratio is 1.

of ethane

oc steam

P P = = 1P P

(Equation 2)

Using this ratio, the mixing distance over the nozzle spacing can be determined. From

Figure 5-2, it can be seen that this corresponds with 26.6. When the nozzle spacing is

set the total mixing length can be determined, using:

nX = d × 26.6 (Equation 3)

22


Knowing that the diameter of the mixing section is 0.98 m (D0) the number of nozzles

can be calculated using:

0n

n

2 × π × DN = d

(Equation 4)

Results are shown Table 5-2.

Table 5-2: Mixing distance resulting from chosen nozzle spacing

dn X Nn(m) (m)

0.20 5.32 30.80.21 5.59 29.30.22 5.85 28.00.23 6.12 26.80.24 6.38 25.70.25 6.65 24.60.26 6.92 23.70.27 7.18 22.80.28 7.45 22.00.29 7.71 21.20.30 7.98 20.5

There will be 25 nozzles used in the nozzle block. This results in a nozzle spacing of

0.246 m and a total mixing distance of 6.40 m.

The temperature at the end of the mixing section is set to 710 K. This is done in order to

achieve an ethane conversion of 70% and a selectivity of 90% towards ethylene.

Mixing the 2 gases entering the mixing section need to accomplish this final

temperature. Therefore, calculations are carried out in order to estimate the

temperatures at which both gases enter the mixing section, see Table 5-3. These

calculations are shown in Appendix E.

Table 5-3 : Temperature components entering the mixing section

Component T (K) T (°C)Ethane 788 515Steam 703 430

23


5.3 Pyrolysis section As a basis for the reactor design, the model made by R. Bosma14 is used. This model

describes the pyrolysis section of the SWR and assumes an ideal gas system.

5.3.1 Ideal gas?

To check whether the ideal gas assumption is correct, the van der Waals equation of

state15 (Equation 5) is compared to the ideal gas law (Equation 6):

2

R×T aP = - V-b V

(Equation 5)

P × V = n × R × T (Equation 6)

The two variables, a and b, in Equation 5 are defined as: 2 2

c c

c c

27×R ×T R×Ta = b = 64×P 8×P

Tc and Pc are the corresponding critical temperature and pressure, respectively for each

component. Filling in both equations resulted in the same pressure for each component

in the system. Therefore the use of the ideal gas law is justified.

5.3.2 Computational work

This paragraph covers the computational work that has been done to describe the

pyrolysis section of the reactor. Using the ideal gas law correlations and known reactions

(Paragraph 2.2.2) the dimensions of the SWR and stream composition of the reactor

effluent are calculated.

The final Matlab-file is enclosed in Appendix F. This file will be explained below

according to the sequence of the calculations. All the titles used here will also be used in

Matlab.

14 Bosma, R.,”Ethane cracking by means of a shock wave reactor”, TU-Delft, Delft, 2005 15 Smith, J.M., Van Ness, H.C.,”Introduction to chemical engineering

thermodynamics”, 4th ed., McGraw-Hill, New York, 1987

24


5.3.2.1 Heat capacity

First the heat capacity is calculated. This is done to get a value for kappa (κ), which

represents the ratio between the specific heat at constant pressure and volume

(Equation 10). This κ is needed to calculate the pressure, temperature and velocity at

the initial shock position.

For an ideal gas system the specific heat capacity can be calculated, using16: 2

p,i pa pb pc pdC = C +C ×T + C ×T + C ×T3 (Equation 7)

For the ease of use in Matlab, an average heat capacity is calculated. If separate Cp-

values would be used, all components would have a different velocity and temperature.

This of course is not the case because all components have the same velocity and the

temperature is uniform.

ip

i

F × CC =

F∑p,i

v

(Equation 8)

The universal gas constant relates the specific heat at constant volume to the specific

heat at constant pressure:15

pR= C - C (Equation 9)

Knowing all parameters the ratio κ can be calculated using:

p

v

Cκ =

C (Equation 10)

5.3.2.2 Initial shock values

To estimate the conditions at the occurrence of the shock wave a mean molecular

weight is calculated by using the fractions of the components entering the pyrolysis

section.

w i wM = γ × M ,i

(Equation 11)

16 Sinnott, R.K. , "Coulson & Richardsons's Chemical Engineering Vol. 6," 3th ed., Butterworth-

Heinemann, Oxford, 1999

25


Knowing κ (Equation 10) and the mean molecular weight (Equation 11) at the mixing

conditions, the shock velocity at the start of the pyrolysis section is calculated.

shockw

κ × R × Tu = M

(Equation 12)

Using an iteration script the velocity, pressure and temperature at the start of the

pyrolysis section are determined.

0 0 0 shock 0Mach number P T κ u u→ → → → →

During the iteration the κ0 mentioned, is checked with the κ from Equation 10. If the

relative difference is larger than the designated tolerance the script will continue to run.

5.3.2.3 Initial guess

In order to calculate all pyrolysis section variables, initial guesses have to be made for all

different parameters. Therefore the following calculations are done. As a basis the

following parameters are taken from the iteration script in the previous paragraph:

• Mach number

• u0

• T0

• P0

• Mw

• κ0

Using these parameters the density of the gas is calculated, using:

w

0

M × Pρ = R × T

0 (Equation 13)

Using the calculated density the starting diameter of the pyrolysis section is calculated.

This diameter is also chosen to be the diameter of the mixing section.

i w0

0

4 × F × MD =

ρ × u × π,i (Equation 14)

26


In order to get the starting initial guesses a distance of z has to be defined. For the start

of the pyrolysis section, z is defined as 0.

The temperature and pressure are from the initial shock values calculation. With the help

of D0 the cross-sectional area (A) can be calculated:

20

1A = × π × D4

(Equation 15)

From the molar flows, the concentration is acquired, which is then used in the rate

expression for the specific reaction.

ii

i

Fc = × F R × T∑

P (Equation 16)

a,jER × T

j jRate = k × e × c

j (Equation 17)

The P and T in Equation 16 at z=0 are P0 and T0 respectively. Note that ci and cj are not

the same.

• ci is the specific concentration of the component i,

• cj is the concentration of components used for the specific reaction j.

• kj is the rate constant of reaction j

• Ea,j is the activation energy of reaction j

The velocity of the gas is calculated using:

iR × Tu = × FP × A ∑ (Equation 18)

After this the specific heat capacity for each component is calculated using Equation 7 at

the current temperature. From this the heat of formation is calculated, using: 2 2 3 3 4 4

ref ref reff,i ref,i pa,i ref pb,i pc,i pd,i

(T -T ) (T -T ) (T -T )∆H = ∆H + C × (T-T ) + C × + C × + C × 2 3 4

(Equation 19)

27


Using the stoichiometric coefficients the heat of reaction is computed, using:

r,j f,i f,i∆H = ∆H (products) - ∆H (reactants)∑ ∑ (Equation 20)

The viscosity per component is calculated in order to calculate the Reynolds number

later on. 24/5 - 2/33

c,i c,i-7 3i w,i1/6

c,i c,i

Z P1.9 × Tη = - 0.29 × 10 × × 10 × M × T T 1.0134

` (Equation 21) 17

Because an ideal gas mixture is assumed, the mean viscosity can be used:

i

i

F × ηη = F∑

i (Equation 22)

To check whether a turbulent flow (Re>100000) can be assumed, the Reynolds number

is calculated. This is of interest due to the fact that a turbulent flow can reach ideal

mixing in a shorter distance and time in comparison to laminar flow.

ρ × u × DRe = η

(Equation 23)

Because of the occurring reactions, the density of the mixture changes. As a result the

pressure in the vessel will change as well. The new pressure is best described by the

following equation: 2

xP = P + ρ × u (Equation 24)

With the new temperature and molar flows the corresponding heat capacity (Equation 7)

and the resulting κ (Equation 10) are calculated. Knowing all these initial guessed

parameters, the real values for the remainder of the reactor are calculated.

17 Jossi, J.A., Stiel, L.I., Thodos, G.,”The viscosity of pure substances in dense gaseous and

liquid phases”, AlChe Vol 8 Issue 1 pp 59-63

28


5.3.2.4 Pyrolysis section

For the computation of the remainder of the pyrolysis section an Ordinary Differential

Equation-solver (ODE-solver) is used. This ODE-solver uses two known correlations

inside the SWR. The first correlation is that of the temperature and pressure dependency

of the length of the reactor. And the second correlation is the interrelation of the

reactions.

Most of the formulas stated in the previous paragraphs are used for solving the

mathematical relations for the pyrolysis section. However some equations are different.

The added and different equations will be stated below.

The diameter and cross-sectional area of the pyrolysis section increase along the

distance. The order in which this area increases is dependent on the chosen rise angle

(α). The diameter correlation, Equation 14, changes to:

02 × π × αD = D + 2 × z × tan

360

(Equation 25)

During the pyrolysis the temperature and pressure change. To calculate this, the

following relations are used:

j r,j

i p,i

-rate × ∆H × AdT = u × dt F × C

(Equation 26)

2dP 2 × π × α ρ × u = u × - 2 × f - 4 × tan ×

dt 360 D

(Equation 27)

Where f is the friction factor: -0.2f = 0.046 × Re (Equation 28)

29


5.4 Assumptions In order to obtain final results, some assumptions are made. These assumptions

concern:

• The residence time

• The widening angle of the pyrolysis section

• The pre-shock temperature

• The pre-shock pressure

• The pre-shock velocity

5.4.1 Residence time

It is stated in the patent that the residence time must be between 5 and 50 ms

(milliseconds). It is chosen to set the residence time to 50 ms, because this results in the

stated conversion and selectivity, in accordance with the assumptions following.

5.4.2 Widening angle

The angle of reactor tube widening is set to 5°. Altering this parameter does not interfere

with the actual result, but only influenced the final diameter and the length of the reactor.

5.4.3 Pre-shock temperature

Figure 5-3 shows the influence off the temperature on the conversion and selectivity. It

can be seen that an increase in temperature increases the conversion of ethane, but

slightly decreases the selectivity towards the ethylene.

Temperature Influence

50%

60%

70%

80%

90%

100%

660 670 680 690 700 710 720 730 740 750 760Temperature (K)

%

Conversion Ethane

Selectivity

Figure 5-3: Influence of temperature

30


It is chosen to increase the temperature in the mixing zone by 10 K, to 710 K, than

stated in the patent. With this temperature the stated conversion and selectivity

(paragraph 1.4) are obtained, in accordance with the other assumptions.

5.4.4 Pre-shock pressure

Besides the temperature, the pressure also influences the conversion and selectivity of

the process as is seen in Figure 5-4.

Pressure Influence

50%

60%

70%

80%

90%

100%

0 1 2 3 4 5 6P ressure (B ar)

Conversion Ethane

Selectivity

Figure 5-4: Influence of pressure

Increasing the pressure, prior to the pyrolysis section, slightly increases the conversion

of the ethane, but drastically decreases the selectivity towards ethylene. Therefore a low

pressure is desired. The pressure used will be the same as stated in the patent1, 1.02

bar.

31


5.4.5 Pre-shock velocity

The final parameter that can be adjusted, when the reactor is on-stream, is the velocity

of the gas prior to the pyrolysis section. The gas has supersonic velocity, which means it

is above Mach 1 (330 m/s).

The Mach nr used in Figure 5-5 is defined as:

VelocityMach nr = Speed of sound

Velocity Influence

30%

40%

50%

60%

70%

80%

90%

100%

2.4 2.5 2.6 2.7 2.8 2.9 3 3.1Mach Nr

%

Conversion Ethane

Selectivity

Figure 5-5: Influence of velocity

From Figure 5-5 it can be seen that an increase in velocity results in a higher conversion

of ethane, but a lower selectivity towards ethylene. It is chosen to use the Mach number

as stated in the patent, 2.8.

32


5.4.6 Summary

It showed that choosing the parameter values as stated in the previous paragraphs that

a conversion of 70% and a selectivity of 90% were obtained. It seems that minor

improvements can be obtained by adjusting these parameters. However the span in

which this can be done is rather small. Figure 5-5 shows an overview of the estimated

parameters. The variable parameters can be altered during operation the fixed

parameter cannot.

Table 5-4: Assumed parameters

Parameter Value Dimension

VariableVelocity 2.8 -Temperature 710 KPressure 1.02 barResidence 50 ms

FixedAngle 5 °

33


5.4.7 Results

This paragraph will discuss the results that are obtained from the MATLAB-file. At first

the flows of the three most abundant components are presented in Figure 5-6. It can be

seen that the flow of ethane decreases and the ethylene and hydrogen increase in time,

which is expected due to the reactions that occur.

Figure 5-6: Ethane, Ethylene and Hydrogen flow in reactor

34


As stated before not only ethane, ethylene and hydrogen are present, but also some by-

products. From Figure 5-7 it can be seen that methane (CH4) is the most abundant by-

product with a production of almost 180 mol/s.

Figure 5-7: By-product flow in reactor

35


From Figure 5-8 it can be concluded that the rate, at which ethane reacts, decreases in

Figure 5-8: Conversion of ethane

time. The total conversion of ethane is 70%.

must be noted that the results obtained and stated above are in line with the results,

It

which were expected at the beginning of the modelling. This can also be seen in Table

5-5, where the flow composition which is expected and which is obtained, just before the

quencher, is tabulated. Acids, which are also present, are not tabulated, but are taken

into account in the further process.

36


Table 5-5: Reactor outlet composition

Stream Name : Reactor outletComp. Units Specification Additional Information

Expected Design Notes (also ref. note numbers)Ethylene wt% 53 - 58 54.88 (1) For this mixture the priceEthane wt% 25-30 28.41 could not be calculatedMethane wt% 4 - 5 4.72Hydrogen wt% 4 - 4.5 4.16Propane wt% 3 - 4 3.40Butadiene wt% 0-0.5 3.09Propylene wt% 0.3 - 0.8 0.48Acetylene wt% 0 - 1 0.42Benzene wt% 0-0.5 0.42Total 100.0

Process Conditions and PriceTemp. K 1273 1290Press. Bara 10 9.8Phase V/L/S VPrice $/tonne - (1)

5.5 Reactor Dimensions To get a clear view on the reactor size the results from all different paragraphs are

tabulated in Table 5-6. From these dimension an artistic impression is made. This

impression is enclosed in Appendix G.

Table 5-6: Reactor dimensions

(m) (m) (m)Speed 0.51 0.286 0.32Mixing start 0.286 0.98 2.83Nozzleblock 0.98 0.98 0.25Mixing final 0.98 0.98 6.40Pyrolysis 0.98 2.07 6.17Total 15.97

Starting Diameter

Final Diameter LengthSection

From Table 5-6 it can be seen that the total reactor length is almost 16 meter. This

length excludes the quencher after the pyrolysis section.

37


6 Separation Now that the composition and the amount of effluent from the reactor are known the

separation section can be designed. There are several types of separation methods. In

this chapter the manner in which separation was approached is discussed. Detailed

information about the employed units is given in Appendix H.

Component properties play an important role determining which separation method is

adequate. The boiling point was the main property used for determining the separation

sequence, because it implicates volatility.

The block diagram stated in paragraph 4.3, gives a global insight in the separation

sequence. While the separation system was configured, innovative separation

technologies were also considered. In this search membrane technology formed the

main focus.

6.1 Separation technology

6.1.1 Membrane

Membrane technology applications in the petrochemical industry are an upcoming trend.

Conventional separation methods, such as distillation and cryodistillation are still used

worldwide, but membrane technology offers a whole range of advantages in comparison

with the former:

• It is an ideal solution for remote locations with limited utilities and sour gas. • Membrane units have no moving parts so maintenance costs are minimal. • For gas sweetening no additional hazardous materials, e.g. amines, are needed. • Low energy consumption, low pressure drop. • Most membrane units are lightweight and compact. • Additionally, they are easy to install and operate.

However, there are a couple of marginal notes:

1. Because it’s an upcoming technology, membrane units are still expensive. 2. Membranes are sensitive to fouling and other impurities present in the feed. 3. Membranes exhibit a short lifetime.

38


Weighing the advantages and disadvantages it must be said that no obvious reason can

be pointed out to make a choice for membrane technology. The only point that might

make the difference between membrane and distillation is the lower energy consumption

of the membrane, making it more economically attractive after a long period.

6.1.2 Distillation

Conventional distillation units constitute the most widespread means of separation in the

petrochemical industry. All distillation units have been programmed in the Aspen Plus

simulation software package.

6.1.3 Cryodistillation

Like distillation, cryodistillation facilities are widely used in the petrochemical industry. In

this process, demethanization, deethanization and ethylene removal require

cryogenically operated distillation towers. These separation steps use the most of the

plant net energy requirement. In the future it might be possible to replace these units by

highly pressurized membrane units, leading to lower energy costs.

6.1.4 Absorbers

Absorbers also constitute a widely used technology. The petrochemical industry mainly

uses them for gas sweetening. Gas sweetening is a process in which sour gases,

present in reactor effluent, are removed. Sour gas removal is of paramount importance

since its corrosive properties form a mayor liability for expensive separation equipment.

In this process an alkanolamine, specifically monoethanolamine (MEA), is used to

absorb H2S and CO2. Usually a mixture of monoethanolamine (15 – 20 wt%) or

diethanolamine (20 – 30 wt%) with water is used.

39


6.2 Order of separation As stated before, an ethylene production of 1 Mt/a must be achieved, with a purity of

99.9 wt%. In order to attain to the above criteria, an adequate separation system is

configured.

In the following sections, separation selection procedures are outlined in detail,

ultimately leading to an efficient separation system. Other separation options will also be

considered and dealt with in an appropriate manner.

6.2.1 Components

The separation system immediately follows the reactor section, after being cooled and

depressurized. The reactor effluent is the feed to the separation system. In Table 6-1,

the feed composition is given.

Table 6-1: Reactor stream outlet composition

Component Structure mol/s kg/s DestinationMethane CH4 176.84 2.84 FuelAcetylene C2H2 9.75 0.25 ConvertedEthylene C2H4 1174.94 32.96 SoldEthane C2H6 567.34 17.06 RecycledPropylene C3H6 6.90 0.29 FuelPropane C3H8 46.37 2.05 FuelButadiene C4H6 34.32 1.86 FuelHydrogen H2 1239.62 2.50 SoldBenzene C6H6 3.25 0.25 DischargedW ater H2O 20734.80 373.54 Discharged/

RecycledCarbonmonoxide* CO 12.80 0.36 UpgradingCarbondioxide* CO2 12.00 0.53 UpgradingHydrogensulfide* H2S 1.38*10-3 4.45*10-5 UpgradingTotal 24018.93 434.49 * = Sour gases

The last column indicates the destination of the separated products. Evidently, all

products with significant economical value will be sold.

40


From an economic point of view ethane and water are being recycled back into the

process. The recycle streams conform to specified conditions regarding purity, pressure

and temperature. To protect the environment, the sour gases are selected for further

processing before end of pipe discharge. The products, which are not sold, recycled or

selected for processing, function as a fuel source, hereby minimizing process fuel costs.

6.2.2 Separation sequencing

Feed composition and component characteristics are used as measures for determining

the separation sequence. The heuristics used, in descending priority, are:

• Most plentiful

• Corrosive components

• Lightest until the region in which the product is located is reached

• Product

• Residual components

6.3 Simulation of the process As stated before all separation units are modelled in the Aspen Plus 11.1 simulation

program. In the following paragraphs, every unit will be dealt with in detail. For an

overview of the separation system see Appendix R.

The tables stated in every section only state the relevant products. The rest is denoted

as by-products. A relevant product is defined as a product separated in that specific unit.

Ethylene is always stated, since this is the desired product. Recoveries and mole

fractions are stated in ranges in order to compensate for fluctuations.

41


6.3.1 Water separation

As can be seen from the heuristics the most abundant component must and is separated

first. At a temperature of 298 K, water and benzene are liquids. It is therefore decided to

use a flash drum to separate these.

Because the boiling points of water and benzene are relatively close to each other,

compared to the other components, they are separated together. Butadiene displays

some affinity towards the water-benzene mixture because 15% of the total butadiene

amount is entrained in the bottom product.

For this flash drum the Lee Kessler Plöcker thermodynamic model is used, because it is

applicable for non-polar or mildly polar compounds, which are present in the vapour

phase.

Table 6-2: Flash drum for water separation

Unit: Flash drum (Water separation) Thermodynamic Model: LK-PLOCKAspen model: Flash2 Outlet Temperature: 298 KFeed: 449 kg/s Column pressure: 2 barHeat duty: -33.11 MW

Recovery Top (%) Top stream fraction (molfrac)

Ethylene >99.9 0.33 - 0.37Benzene 0.01 – 0.05 0.03 - 0.04Water 0.1 – 0.3 0.010 - 0.012Byproducts 0.62 - 0.63

42


6.3.2 Benzene\Water separation

For the water-benzene system a conventional distillation column is used. Some water is

entrained with the top column vapour stream but is within acceptable limits. Considering

the non-polar nature of benzene, it is justified to use the Peng Robinson thermodynamic

model, applicable to mildly polar to non-polar compounds.

Table 6-3: Water-Benzene distillation colum

Unit: Distillation Column (Separation benzene-water)Aspen model: Radfrac Thermodynamic model: Peng- RobinsonNet Heat duty: 754.5 MW Column pressure: 1 barTop T 373.25 K Bottom T 374.65 K

Water 0.05 – 0.1 0.96 - 0.97Benzene >99.99 0.013 - 0.014


6.3.3 H2S, CO and CO2 removal

Sour gas, which is formed during pyrolysis, is to be removed as soon as possible,

because of its corrosive nature and the negative effects on the low temperature

distillation columns.

As has been stated in paragraph 6.1.4 an amine absorber will be used for gas

sweetening. In Aspen a mixture of MEA/water is used of which 15 wt% consists of MEA.

43


The MEA is regenerated using a regenerator, in which the absorbed sour gases are

desorbed and sent for further treatment (Claus process). After regeneration the MEA is

recycled to the absorber. During regeneration no significant amounts of MEA are lost.

Table 6-4: Acid removal unit

Unit: Amine absorber Thermodynamic model: AminesAspen model: Radfrac Temperature: 315 KAmine stream (mass fraction): 0.85 water & 0.15 MEA

H2S <0.01 -CO2 <0.01 -CO - 1.01E-07Ethylene >99.99 0.35 - 0.36Water 0.1 - 0.2 0.039 - 0.040Monoethanolamine

<<0.01 0.00005 - 0.00006

Byproducts 0.60 - 0.65

Top stream fraction (molfrac)

Recovery Top (%)

6.3.4 Dryer

The sweetened vapour, coming from the absorber, still contains a certain amount of

water. Before sending the stream into the demethanizer this amount is separated in a

dryer. In Aspen this step is modelled using a SEP2 block. Water must be removed to

prevent freezing and plugging under the cryogenic conditions.

44


6.3.5 Hydrogenation of acetylene

During pyrolysis an acetylene amount of approximately 0.26 kg/s is formed. Acetylene

has to be removed from the product stream because it poisons the catalysts used for

downstream ethylene processing. In addition, acetylene can form metal acetylides,

which are explosive contaminants.18

There are two ways for removing acetylene from an ethylene rich environment:

1. Selective hydrogenation of acetylene to ethylene.

2. Separation of acetylene from the mainstream.

The most common industrial method of eliminating acetylene is hydrogenation, as the

separation method is both expensive and dangerous19. Acetylene removal takes place

after the main product stream has been stripped of residual moisture.

Acetylene is to be removed by means of selective hydrogenation in a fixed bed reactor,

because of the deactivation of the catalyst. To catalyse, a Pd/Al2O3 catalyst is proposed,

consisting of 95% aluminium based support and 5% Pd.

Due to lack of time and importance of modelling this unit in detail, this reaction has been

simulated in Aspen whereby an acetylene conversion of 95% was aimed at.

18 http://www.che.lsu.edu/COURSES/4205/2000/McNeely/paper.htm 19 Mostoufi, N., Ghoorchian, A., Sotudeh-Gharebagh, R., “ Hydrogenation of acetylene: Kinetic

studies and reactor modeling””, Int. Journal of Chem. Reactor Eng., Vol 3. Article A14, 2005

45


6.3.6 Demethanizer

According to heuristics, the lightest components are separated after gas sweetening. As

the name implies, methane is separated as an overhead component from C2 and heavier

bottom components. However, since hydrogen is a lighter component than methane it is

entrained with the overhead methane stream. Except for hydrogen, CO is entrained

along with the overhead stream. Further down the separation line CO will be selected for

upgrading.

Table 6-5: Methane-Hydrogen distillation column

Unit: Demethanizer Thermodynamic model: Peng RobinsonAspen model: RadfracTop T 134.35 K Bottom T 251.15 KNet Heat duty: 15.55 MW

Methane 99.95 – 99.99 0.12 - 0.14Ethylene 0.40 – 0.10 0.0037 - 0.0039Hydrogen >99.99 0.87 - 0.90Byproducts 0.00002 - 0.00003


6.3.7 Hydrogen/Methane separation

The methane-hydrogen overhead stream from the demethanizer, is led through a

palladium-based membrane reactor in order to separate hydrogen from methane.

Lacking accurate data this unit is not modelled accurately, therefore it is chosen to

explain theoretically how this separation occurs.

In this process a hollow fibre membrane unit is used to separate hydrogen from

methane. The driving force is the partial pressure difference across the membrane for

hydrogen and methane. The pressure feed gas enters the membrane from the tube side

and hydrogen is collected at low pressure.

46


‘Fast gases’ such as H2 with a high permeation rate diffuse through the membrane into

the hollow interior and are channelled to the permeate stream. Table 6-620 indicates

which gases are categorised as fast or slow. ‘Slow gases’ flow around the hollow fibre,

making sure a fast gas, like hydrogen, is separated from the slower gas, in this case

methane. A schematic representation on how this membrane unit works is given in

Figure 6-120.

Table 6-6: Relative permeation rates through a membrane

Fast H2O He H2 NH3 CO2 H2S O2 Ar CO N2 CH4 C2H4 C3H6 SlowRelative permeation rates

Figure 6-1: Schematic representation of membrane unit

47

20 http://www.medal.airliquide.com/en/membranes/hydrogen/index.asp


6.3.8 Product separation

After demethanization, ethylene is separated as overhead component. Conform the

criteria this stream is 99.9 wt% pure. It was possible to model the column in Aspen while

delivering on-spec ethylene. As can be seen half of the total acetylene amount is

entrained with the overhead product. However, it must be noted that the stated

percentage denotes the recovery and not the actual amount.

Table 6-7: Ethylene distillation column

Unit: Fractionator Thermodynamic model: Peng RobinsonAspen model: Radfrac Column pressure: 8 barTop T 214.85 K Bottom T 235.55 KNet Heat Duty: 58.90 MW

Ethylene 99.50 – 99.90 0.9990487Byproducts 0.000931


6.3.9 Deethanization

In the last step, residual ethane is separated and recycled to the reactor. Acetylene and

ethylene are also separated as overhead components from C3+ bottom components.

Table 6-8: Ethane distillation column

Unit: Deethanizer Thermodynamic model: Peng RobinsonAspen model: Radfrac Column pressure: 5 barTop T 219.85 K Bottom T 287.35 KNet Heat Duty: 34.34 MW

Ethane >99.99 0.94 - 0.95Byproducts 0.050 - 0.060


48


7 Heat & Power Integration During reactor operation heat and power are needed in order to run the plant. Some

operations use heat and power, others supply this. It is, therefore, economically sensible

to exchange this heat and integrate the power to a maximum extent. For this integration

Douglas9 is used.

7.1 Heat Overall there are two streams that need to be cooled and three streams that need to be

heated. The streams that need to be cooled are:

• Leaving reactor (10)

• Before separation (12)

The streams that need to be heated are:

• Water to steam (6)

• Ethane to reactor (3)

• Ethylene product (49)

The numbers behind the stream corresponds with the number from the PFS enclosed in

Appendix R. All these streams have specific heat capacities and temperatures as shown

in Table 7-1.

Table 7-1: Stream data

Ident. Nr. Hot Cold MW /K Tin Tout DT MW attReac out 1 x 1.2 1128 873 255 306

After expander 2 x 1.05 648 298 350 367.5W ater 3 x 0.78 358 1293 -935 -729.3Ethane 4 x 0.09 298 788 -490 -44.1

Ethylene 5 x 0.048 231 303 -72 -3.5-103.4Total

DTmin=10°C

Stream Data

Stream Conditions F*Cp Temperatures (K) Q avail.

49


The F*Cp-values are taken from the Matlab file, except for the ethane and the stream

after the expander. These are estimated using the ideal gas Cp calculation (Equation 7)

and the total flow that needs to be cooled or heated.

Using the pinch technology it is found that the following equipment is needed:

• Three heat exchangers

• Three coolers

• One heater

The Heat Exchanger Network (HEN) built, is enclosed in Appendix I. Stream 1 is not

really split into 3 different streams but remains as a whole. This stream is quenched

immediately using the stream 3 and 4. Due to this the temperature drops to

approximately 970 K. Directly after the quencher a cooler is placed to cool the

temperature to 870 K to make sure the pyrolysis is stopped. If technically possible, the

coolant could also be introduced directly in the quencher, lowering the temperature

directly to 870 K.

7.2 Power In the process different units, e.g. compressors, require power to operate. Other units,

e.g. expanders, supply power. During the operation of this plant more power is

accumulated then needed in the process. Therefore all units requiring power are

supplied. The surplus of power is led to the electricity net. The income for this amount is

calculated in chapter 8.

50


8 Economics An economic evaluation for the SWR plant is made. For this, a method from Coulson

and Richardson16 is used. In order to use this method, certain parts need to be specified.

These parts are:

• Equipment cost

• Raw materials cost

• Utility prices

• Labour costs

• Sales Income

From these estimations an economical study can be proposed for the economical life

span of the SWR plant, which is stated to be 10 years, although the physical life span of

the SWR plant could easily be 15 years. The study will concentrate on the economical

life span of 10 years. If relevant, the 15 years study costs and benefits will be stated.

8.1 Purchased equipment cost The purchased equipment cost is composed of a list of components needed in the SWR

plant. This list and its calculations are stated in Appendix J.

The total equipment cost is $106 million. The cost for each individual unit is found with

the help of one of the three following references:

• Coulson & Richardson16

• DACE21,

• Peters and Timmerhaus22.

21 Dutch Association of Cost Engineers, “Prijzenboekje”, 22th ed., Elsevier, mei 2002 22 Peters, Max S., Timmerhaus, Klaus D., “Plant design and economics for chemical engineers”,

4th ed. McGraw-Hill, 1991

51


8.2 Cost estimation for raw materials Using the mass balances of the process, the required raw materials are calculated. For

the stated production of 1 million ton ethylene per annum the process needs the

following amounts of ethane and water.

Table 8-1: Raw materials for ethylene

Component Amount [Mt/a] Cost [$/t] Cost [M$]Ethane 1.25 150 188.58Water 2.47 0.675 1.67Total 190.25 The amounts in Table 8-1 are the amounts needed to make up the recycle streams to

the desired amounts for the production of ethylene. The cost of raw materials is reduced,

because of the use of recycle streams. The prices, used for the components stated, are

from the project description or from Platts23.

The amounts of components needed for the make up streams are calculated using the

mass balances and amounts recycled. The water recycle stream was set on 80% using

a common sense engineering point of view. This percentage could be changed if

needed, but it will affect the estimation of the cost of raw materials. The calculations are

shown in Appendix K

23 http://www.platts.com

52


8.3 Determining the cost of utilities The utilities used in the SWR-plant mainly concern electricity and gas, but also

monoethanolamine (MEA) is needed. MEA, which is introduced with water, is used in the

acid gas absorber to absorb H2S, CO and CO2 from the product stream. The mass

fraction of MEA in the stream is 15 wt%. This stream absorbs all the acid gasses in the

product stream. The MEA/water mixture is then led through a regenerator to release the

acid gasses and regenerate the MEA/water mixture in order to re-use it. This way the

MEA is only bought once per annum. This has positive effects on the economical

estimation because the price of MEA is very high. In Table 8-2 the total cost of utilities is

presented, the calculations of these costs are enclosed in Appendix L

Table 8-2: Cost estimation of utilities

Unit Amount Cost Total costm3/a $/m3 M$/a

Cooling in Process cooling water 8.34E+07 0.13 10.56

Heat in Process heat MW $/MMBTU M$/a852.72 5.40 125.38

Utility Monoethanolamine (MEA) 1088 267.86 0.29Water 6165 0.68 0.004

Total 136.24

8.4 Labour cost for the SWR-plant The SWR-plant has an operating span of 8400 hours per annum. During this operating

span, operating personnel is required. Because the plant will run continuously, and an

average working day is 8 hours, multiple shifts are needed. Dividing the 24 hours in a

day by 8 hours, results in 3 shifts per day. To make sure a safe amount of personnel is

working at the plant, 5 employees are needed per shift. With an average pay of 31 dollar

per hour this results in the following cost per annum (Table 8-3). The total calculation is

enclosed in Appendix M

Table 8-3: Labour cost for the SWR-plant

Number of personnel Shifts a day Hours per shift Hour wage Total cost$/hr M$/a

5 3 8 31 1.30

53


8.5 Capital cost of the SWR-plant The capital cost of the SWR-plant is made up from:

• Fixed capital cost

• Variable costs.

The fixed capital costs can be found from the required equipment cost of the SWR-plant.

The variable costs can be calculated from the raw materials, utilities and labour. First the

fixed capital cost is shown in Table 8-416. Using the fixed capital cost the variable and

total capital cost are calculated. These values are mentioned in Table 8-516. The

calculations are enclosed in Appendix N.

Table 8-4: Fixed capital cost of the SWR-plant

Item M$Purchased Equipment Cost 105.77Equipment erection 42.31Piping 74.04Instrumentation 21.15Electrical 10.58Buildings, process 15.86Utilities 52.88Storages 15.86Site development 5.29Ancillary buildings 15.86Sub-total physical plant costs 359.60

Design and Engineering 107.88Contractor's fee 17.98Contigency 35.96

Total fixed capital cost 521.42

Table 8-5: Annual production costs

M$Variable costs 364.52Fixed costs 153.68Direct production costs 518.20Extra costs 155.46Annual production costs 673.66 From Table 8-5, the production cost per kg ethylene can be calculated, using:

54


Annual production cost $=Annual production kg ethylene

This is about 0.67 $/kg ethylene.

8.6 Economic evaluation of the SWR-plant From all the estimation on costs, the total cost for the SWR-plant is known. To get a

clear view on the economical potential of the plant the gross income is calculated. Using

the gross income and the annual production cost, the Net Cash Flow (NCF) is found.

This NCF can be used to estimate the Rate of Return of the SWR-plant how much the

SWR-plant is worth. The method of calculating the NCF is found in Coulson16 and is

shown in Appendix O. In Table 8-6 a summary of the gross income is shown.

Table 8-6: Gross income of SWR-plant

Product Mt/a $/t M$/aEthylene 1 650 650.00Hydrogen 0.077 2700 208.85

Subtotal 858.85M$/a

Electricity 67.18

Subtotal 926.03Cost for wastewater disposal

m3/a $/m3 M$/a2550000 0.51 1.30

Total 924.72 The Net Cash Flow becomes:

Gross income - Annual production cost = Net Cash Flow924.72 - 673.66 = 251.06 M$/a

55


The Rate of Return16 and the value for the SWR-plant are enclosed in Appendix O. In

Figure 8-1 the Net Present Value is plotted against the life span of the plant. The Net

Present Value after 10 years is approximately 571 M$. After 15 years this comes down

to approximately 863 M$.

Feasible Net Present Value

-800

-600

-400

-200

0

200

400

600

800

0 1 2 3 4 5 6 7 8 9 10 11 12 13

Years

Cum

ulat

ive

cash

flow

[M$]

Figure 8-1: Net Present Value of the SWR-plant

The plant will be built in three years. The investments (fixed capital cost) are spread

amongst these three years. At the end of the economical life span of the plant (10 years)

a new investment is made. This investment is used as starting capital for a possible new

plant or revise of the old plant.

During the calculations it is assumed that 15% of all the catalyst, used in the acetylene

to ethylene reactor, is replaced after 6 months (The time after which the catalyst is

deactivated19). Therefore 2 reactors are built. While one of the reactors regenerates the

catalyst and fresh catalyst is added the other reactor is on-stream.

56


9 Safety During the cracking of ethane, many flammable hydrocarbons are present and formed.

In order to assess the potential threat of the SWR-plant a DOW Fire & Explosion Index

(F&EI) calculation is carried out.

9.1 Fire & Explosion index The calculation is done using the standard form for the F&EI, which is enclosed in

Appendix P. Notes on the decisions taken and the factors implemented are stated

below.

9.1.1 Boundary

The main attention of this project is producing ethylene by means of a SWR. Therefore

the F&EI-calculation is only an indication whereby only the plant as a whole is

considered excluding storage facilities. Note that if storage facilities are taken into

account, the huge amount of ethylene present would pose an additional risk.

9.1.2 Material Factor

The material factor (MF) is a measure for the intensity of energy release from a chemical

compound or a mixture of compounds or substances. It is the starting point for the

calculation of the F&EI. The MF is determined by using two potential hazards acquired

from the National Fire Protection Association (NFPA) classification. These hazards are:

• Flammability (Nf)

• Reactivity (Nr)

These factors combined lead to a MF, ranging from 0 to 40, where 0 means no hazard

and 40 means serious hazard24.

24 Lemkowitz, S.M., Pasman, H.J., “Chemical Risk Management”, TU Delft, 2002

57


During the F&EI-calculation the highest material factor, of a compound that is present in

significant quantities, must be taken into account. In this case, ethane and ethylene are

the two most plentiful components present. Both have a MF of 24. Acetylene, which is

also present, has a larger MF (40) but its concentration is too small to be considered the

dominant material. For an overview of all the material factors see Table 9-1.

Table 9-1: Material factors25

Nh Nf Nr

Acetylene C2H2 40 1 4 4Benzene C6H6 16 2 3 13-Butadiene C4H6 29 2 4 3

Carbon Monoxide CO 16 3 3 1Ethane C2H6 24 1 4 0Ethylene C2H4 24 1 4 2Hydrogen H2 21 0 4 0Methane CH4 21 1 4 0Propane C3H8 21 1 4 0Propylene C3H6 21 1 4 0

Compound Formula MFNFPA Classification

9.1.3 General process hazards

A & B Overall the occurring reactions are endothermic. Therefore a factor of

0.2 must be taken into account for B.

C As stated before no shipping and handling will be taken into account

during this F&EI-analyses. Thus no penalty is given.

D All plant units are stated to be outside, which means that no penalty

has to be taken into account for D.

E&F It is assumed that adequate measures are taken. No penalty.

25 Dow’s Fire Explosion Index hazard classification guide, AlChe, New York, 1981

58


9.1.4 Special process hazards

A Ethylene and ethane are not marked as toxic. No penalty.

B The pressure in the system never drops below 500 mmHg. No

penalty.

C The process always operates in the flammability range of the

components, so a factor of 0.8 is taken into account.

D Because no solids are present, dust explosions cannot occur.

Therefore no penalty is given

E The highest operation pressure, present in the reactor, is 10 bar, This

equals 130.5 psi of overpressure. The relief valve is 20% above

operating pressure, which is 157 psi. From the first Figure in Appendix

Q the penalty factor is taken (0.35)

F The ethylene product distillation column is made from steel. The

temperature is well below a temperature of 244 K, which implies a

factor of 0.30.

G The process is run continuously, with an ethane feed flow rate is

56.2kg/s (123.9lb/s). Because not all residence times are known or

can be determined, only the reactor feed is taken into account and not

the total amount of ethane in the plant. The heat of combustion for

ethane is 20.4 BTU/lb25, resulting in a possible energy release of

2.528 MBTU. Looking at the second figure in Appendix Q it can be

seen that no penalty needs to be registered.

H Because there is a small quantity of acid gas (CO, CO2, H2S) present

a factor of 0.1 is applied due to corrosion and erosion

59


I Welded joints are used therefore the minimum factor (0.1) is taken

into account. Full equipment details are not known.

J A carrier fluid heater is used. It is assumed that the distance between

the reactor, a possible leak source, and the heater is 15. According to

the third figure in Appendix Q a penalty of 0.89 needs to be applied

because the condition, in which the materials are used, is above

boiling point.

K Heat transfer systems using a combustible liquid as the heat

exchange media are not used in the process. Therefore no penalty is

applied.

L Large turbines and compressors are used, implying a penalty of 0.5

As can be seen in Appendix P the index works out 116, meaning that the degree of

hazard is classified as ‘intermediate’ as shown in Table 9-2. In many companies, when

the F&EI is above 100, the degree of hazard is judged to be too high. Risk reducing

measures are required.

Table 9-2: Degree of hazard for F&EI

Index range Degree of hazard

1~61 Light62~96 Moderate97~127 Intermediate128~158 Heavy159~up Severe

60


9.2 Fire protection and prevention Ethane and the major part of the products of the pyrolysis, alkanes and alkenes, are

flammable. Because steam cracking is a closed process, the primary fire potential is

from leaks. From these leaks, liquids, gases or vapors can reach an ignition source,

such as a heater. This contributes to a large penalty in the F&EI. Below recommended

measures for fire protection and prevention are given26 :

9.2.1 Leak prevention

In order to prevent leakage, good maintenance of the piping system is essential. Also

the use of correct and adequate construction materials is of the utmost importance.

• Painted equipment may have a better anti-corrosion effect than unpainted

equipment.

• Installation of double sealed pump.

• Plugs with safety chains can be installed on all hydrocarbon drains and vents to

prevent leak of inadvertent opening.

• Minimizing the use of flanges in flammable material service reduces this potential

hazard as well.

9.2.2 Leak detection

Early detection of a release is a key to rapid containment. Hydrocarbon leak detection

devices, e.g. gas detectors, should be installed. They should be installed at location with

a high hydrocarbon concentration present, such as the SWR reactor and distillation

columns. Also detectors should be located on elevated structures where potential

hydrocarbon release is high, such as compressor shelters and cold-service exchangers.

26 Olivo, J.“Loss prevention in a modern ethylene plant”, Loss Prev. Process Ind Vol. 7, No. 5,

1994.

61


9.2.3 Leak dispersion, containment

In case of a spill or leak, deluge systems, steam curtains and water curtains systems

can be activated to isolate the leak.

The use of curbs, dikes and trenches to confine or divert hazardous materials to safe

locations can also be considered during plant design. This way, the spill dispersion area

is limited resulting in risk reduction.

9.2.4 Miscellaneous

Moving the carrier fluid heater, a potential ignition source, further away reduces the risk

considerably. Thus it’s preferable to keep it as far away possible from the reactor

section.

Excess flow valves, fire-actuated valves and remote activated isolation valves can be

installed to limit the amount of hazardous materials released.

Building a firewater storage tank and installing firewater pumps reduces risk as well.

Using fixed water spray protection for storage of liquefied gasses is also recommended.

Fireproofing of selected structural supports, equipment and cable trays also reduces

potential risk.

62


10 Controllability In the SWR-plant a certain degree of control is required to make sure that three

objectives27 are met:

• Safe process operation

• Production rate is maintained

• Product quality is maintained

The first objective is making sure that the safety of the immediate and remote

environment is assured. The quality of the ethylene may not have the desired

specifications, but the plant will operate without damaging the environment.

The specified production rates are maintained to make sure that the process equipment

of the plant is not damaged and that the market demand is still fulfilled.

The last objective is set to make sure that the product can be sold on the market. This

product amount could be under the market demand, but it does not have to be discarded

like waste.

Using these objectives as guidance, a control scheme for the SWR-plant is created. This

is shown in Appendix R. Not all the control sections will be discussed. Only the most

important controls will be dealt with according to their location in the PFS.

27 Ogunnaike, Babatunde A., Harmon Ray, W.,”Process dynamics, modelling, and control”,

Oxford university press, 1994

63


10.1 Inlet streams The first control to make sure that the process is running correctly is that of a ratio

control between the water and ethane feed streams. Both streams have a recycle

stream as shown in Figure 10-1.

Ethane Feed

Ethane recycle

Water Recycle

P01

T01

FC

X

FC

V01

FT

FT2a

1

359

2

1

300

1a

10

289

1

10

300

59

10

262

Ethane

Water

Figure 10-1: Control on feed

The water recycle is measured, and this will control a valve to make sure that the total

amount of water entering the SWR to ensure safe operation.

The same is done with the ethane recycle and the fresh ethane feed. The only difference

is that the set point of the flow control on the fresh ethane stream is set by the total flow

of water entering the plant. The ratio between water and ethane is set to 6.67 kg water

on 1 kg of ethane. So if the total amount of water is know, the total amount of ethane

entering the system can be adjust so that the market demand is met.

64


10.2 Shock wave position The shock wave inside the SWR is controlled from outside the reactor. In Figure 10-2

the control on the shock wave is shown.

R01

��

��

T02

��

E02

TC

PC

CW

12

2

648

11

10

873

10

10

970

Ethane

Steam

Products

Figure 10-2: Shock wave controller

The pressure of the reactor outlet is measured. The signal from the pressure sensor is

send to the valve. The valve is operated to make sure that the shock wave front starts at

the widening of the pyrolysis section of the reactor. This is the desired point of the start

of the shock wave front.

If the shock wave front is before of the widening of the pyrolysis section, the mixing

quality of the ethane and steam could be less. This will lead to more ethylene production

as well as more by-products.

When the shock wave front starts after the widening of the pyrolysis section, the total

pyrolysis time is shortened leading to lower ethylene production. This control is placed to

meet the market demand and quality of the ethylene product.

65


10.3 Emergency control of the SWR reactor The emergency control is placed to make sure that the surrounding environment and

personnel are safe in case of a possible pressure built up in the SWR reactor. If the

pressure after the expander is too high for the remainder of the process, a signal is sent

to the purge valve. If this happens the total amount of stream inside the system is sent to

a burner. The schematic control is shown in Figure 10-3.

The unit after this purge will notice the pressure drop and reduces its liquid. This triggers

two controls in Figure 10-4 to close and shut down the rest of the plant.

From R01

V02

20

2

298

Bleed

��

��E04

E03��

TC

��

To V02

��

TCE05

E06

E07

PC

5051

CW

CW

19

2

298

18

2

374

17

2

378

16

2

648

15

2

298

14

2

51313

2

648Products

Steam

Products

21

-

-

Figure 10-3: Emergency control in SWR plant

66


10.4 Separation of water Water is separated from the gas products with a condenser. This condenser needs two

controls. The bottom stream is controlled using the liquid hold-up inside the vessel. If the

liquid drops too fast, the valve is closed. This makes sure that the water releases all the

gasses, which could be entrained in the stream.

The pressure in the top of the condenser is measured to maintain the pressure for the

rest of the separation section of the SWR-plant. If the pressure drops, the valve will

close to built up pressure. Otherwise, water could be sent into the other part of the

separation section. This will lead to problems in the absorber and other distillations. The

set-up of this control is shown in Figure 10-4

To C08(Benzene / water)

PC

LC

C01

24

2

298

232 298

Products with water

Products

Figure 10-4: Condenser and controls

For the absorber and the dryer the same control technique is used. Making sure no

fluids enter the separation section of the gas products. This way the ethylene will reach

the desired quality.

67


10.5 Water discharge To make sure the benzene concentration in the water discharge is not too high a

concentration controller is used. This controller sends a signal to the temperature

controller, where it is used as a set point for the reboiler duty. The control scheme is

shown in Figure 10-5.

E17

LCTC

WaterFC

Water recycle

Water/ Benzene

��

23

2

298

62

1

375

63

1

375

64

1

375

ST

CC

Figure 10-5: Benzene concentration control scheme

68


10.6 Distillation columns This short description will explain the controls on one of the distillation columns in the

SWR plant. This control scheme is applicable to all other distillation columns as well.

The distillation column shown in Figure 10-6 is used for the separation of methane and

hydrogen of the ethylene product.

It can be seen that there are three controls in the top and two controls in the bottom.

First the top part is discussed.

V04

��

E10

E11

P06

LCTC

PC

CH4/ H2

LC

TC

C05

39

44

20

251

40

41

42

CW

ST

43

20

134

ProductsWith CH4/ H2

Products

Figure 10-6: Controls in a distillation column

69


10.6.1 Control in the top of the column

The top of the distillation has three different measurements at different locations.

Starting in the top of the column, the temperature is measured to know how much

cooling is needed to condense the products, which are led back to the column.

In the condenser the pressure is measured to make sure that only the gas product is

taken out. The liquid level inside the condenser is measured to ensure a constant flow

rate. This makes sure that the distillation column will operate at the set conditions for the

top.

10.6.2 Control in the bottom of the column

The bottom of the column has two controls. A temperature controller is used to set the

heat duty of the reboiler. In case the temperature is too low, more duty is needed and

vice versa.

The level controller makes sure that the column operates safely. The level control gives

a signal to the valve of the stream leaving the bottom of the column. This way the

column will not dry up.

70


10.7 Fixed bed reactor In the reactor certain products are formed which are removed with the help of distillation

columns. However the acetylene formed, makes distillation very complex. To solve this

problem a fixed bed reactor is used to hydrogenate the acetylene with the hydrogen

present to form ethylene. The scheme of this reactor is shown in Figure 10-7.

R02

CC

37

2

360

ProductsProducts with

Acetylene

Figure 10-7: Acetylene fixed bed reactor

The gas stream is fed from the bottom of the reactor. Before the stream enters the

reactor the concentration of acetylene is measured. The signal from this measurement is

sent to a valve above the fixed bed reactor. The valve will be closed a bit, if the

concentration is to high and vice versa. This valve determines the residence time in the

fixed bed reactor, to assure that most of the acetylene is converted to ethylene. This

valve is a air-to-close valve, to make sure that the gas can leave at all time.

71


10.8 Membrane It was decided to use a membrane to separate methane and hydrogen. If a normal

distillation column would be used, this would be large and very costly. The membrane

has three controls as can be seen in Figure 10-8.

C09

��

H2

CC

PC

CH4

��

��

F02 T03

43

20

134

70

20

243

712 157

73

2

157

FuelTC

CH4/H2

Figure 10-8: Membrane separation and controls

As with all gas phase operations, the membrane pressure is measured in the top of the

separator. This way it can be determined how much methane can be sent off to maintain

the pressure inside the membrane. Methane from this section is used in the furnace to

heat up the water stream that enters the SWR-reactor.

Because the hydrogen is sold as a by-product, a certain quality is required. To make

sure this is obtained, a concentration measurement is performed after the membrane

section. If the quality is not within the desired specifications, the feed to the membrane

section is reduced. This way the quality of the hydrogen is increased.

To make sure the membrane is not damaged, the temperature needs to be in a certain

range. A temperature controller, which measures the temperature before the heater, is

used to set the heat duty of the furnace.

72


10.9 Ethane Purge It is possible that the ethane quality, meant for recycling is too low. A concentration

controller is used to check whether the concentration suffices. If not a valve closes the

recycle loop and sends the ethane to a burner. This is shown in Figure 10-9.

��

E14

P08

TC

PC

LC

��

C07V06

54

55

56

57

58 5220

CW

59

--

CC

PurgeEthane recycle

Figure 10-9: Ethane recycle control scheme

73


11 Conclusions and recommendations

11.1 Conclusions The project description was set to design a concept for a SWR plant, which produces 1

Mt/a (1 million tonnes per annum). Certain requirements are stated for the design to

achieve a realistic conceptual design.

Market demand and product quality constraints were the most essential. From the

modelling of the SWR, it can be concluded that both constraints are met. With this

knowledge the economic potential of the plant was evaluated. It can be said that the

plant will make profit after an economical plant life of 10 years. This is reached due to

recycling and recovering most of the process materials.

With the concept control scheme, the product will reach its quality and quantity. These

controls will also ensure safe plant operation, protecting both people and local

ecosystems.

Pinch technology evaluation resulted in lower energy consumption than in the currently

used thermal-cracking processes.

Conform Texas wastewater legislation, SWR-plant waste streams are processed in

order to minimise environmental impact upon effluent discharge. This is also done for

the off gasses (H2S, CO2) and the benzene content in wastewater.

Benchmarking SWR-plant performance against current day thermal cracking processes

shows that the former performs just as good as the latter. Therefore, no objections could

be raised against SWR-technology.

74


11.2 Recommendations In the development of the SWR plant certain points were encountered which required

extra attention. These points were simplified or approached in another way. Below a

summary is made of the major problems requiring attention in further development.

Determining the required mixing degree in the SWR was problematic. In this report the

mixing is correlated with the nozzle block. Additional testing is required to evaluate the

impact of the feed mixing degree on the reactor length. Knowledge that is obtained

about the mixing behaviour can be of great value because this determines the product

quality.

The membrane unit was another aspect that was simplified. To see if this unit is

applicable for the process, more information is needed. This data could be retrieved from

testing the set-up in a pilot plant.

Pilot plant testing of the SWR could provide more insight in the total separation section.

This way the separation section could be tuned in such a way that the need for extra

utilities is reduced to a minimum.

75


Literature 1 Hertzberg, A., et al, “Method for initiating pyrolysis using a shock wave”, US

Patent 5,300,216, 1994

2 Grievink, J., “Project Objectives & Description”, TU Delft, 2006

3 Gielen, D.J., Vos, D., van Dril, A.W.N., “The petrochemical industry and its

energy use prospects for the Dutch energy intensive industry“, ECN-C—96-029,

1996

4 http://www.capitol.state.tx.us/statutes/wa.toc.htm

5 http://www.texas.gov

6 Sundaram, K.M, Froment, G.F, “Modelling of thermal cracking kinetics - I”, Chem.

Eng. Sc., 1977, Vol 32, pp 601-608

7 Moulijn, Jacob A., Makkee, Michiel, van Diepen, Annelies, “Chemical process

technology”, John Wiley & Sons Ltd, 2001

8 Hidaka, Y. et al, “Shock-tube and modeling study of ethane pyrolysis and

oxidation”, Comb. and Flame, 120, page 245-264, 2000

9 Douglas, J.M., ”Conceptual Design of Chemical Processes”, McGraw-Hill, New

York, 1988

10 van Kimmenaede, Ir. A.J.M.,”Warmteleer voor technici”, 8e druk, Wolters-

Noordhoff, Groningen, 2001

11 Bos, Ir. G.A.,”Stromingsmachines”, 1e druk., Stenfert Kroese, Houten, 1997

12 van den Akker, H.E.A., Mudde R.F.,”Fysische Transportverschijnselen I”,

Tweede druk, DUP Blue Print, Delft, 2003

13 Knowlen, C. et.al., “Petrochemical pyrolysis with shockwaves”, AIAA., 1995, 95-

0402.

14 Bosma, R.,”Ethane cracking by means of a shock wave reactor”, TU-Delft, Delft,

2005

15 Smith, J.M., Van Ness, H.C.,”Introduction to chemical engineering

thermodynamics”, 4th ed., McGraw-Hill, New York, 1987

16 Sinnott, R.K. , "Coulson & Richardsons's Chemical Engineering Vol. 6," 3th ed.,

Butterworth-Heinemann, Oxford, 1999

17 Jossi, J.A., Stiel, L.I., Thodos, G.,”The viscosity of pure substances in dense

gaseous and liquid phases”, AlChe Vol 8 Issue 1 pp 59-63

18 http://www.che.lsu.edu/COURSES/4205/2000/McNeely/paper.htm

I


19 Mostoufi, N., Ghoorchian, A., Sotudeh-Gharebagh, R., “ Hydrogenation of

acetylene: Kinetic studies and reactor modeling””, Int. Journal of Chem. Reactor

Eng., Vol 3. Article A14, 2005

20 http://www.medal.airliquide.com/en/membranes/hydrogen/index.asp

21 Dutch Association of Cost Engineers, “Prijzenboekje”, 22th ed., Elsevier, mei

2002

22 Peters, Max S., Timmerhaus, Klaus D., “Plant design and economics for

chemical engineers”, 4th ed. McGraw-Hill, 1991

23 http://www.platts.com

24 Lemkowitz, S.M., Pasman, H.J., “Chemical Risk Management”, TU Delft, 2002

25 Dow’s Fire Explosion Index hazard classification guide, AlChe, New York, 1981

26 Olivo, J.“Loss prevention in a modern ethylene plant”, Loss Prev. Process Ind

Vol. 7, No. 5, 1994.

27 Ogunnaike, Babatunde A., Harmon Ray, W.,”Process dynamics, modelling, and

control”, Oxford university press, 1994

II


Appendix Index A Pure Component Properties IV B Mass Balances V C Venturi diameter calculation VII D Polytropic compression VIII E Mixing of gases IX F Matlab file X G Artistic impression SWR XXIV H Stream and Unit summary XXV I Heat Integration XXXIII J Equipment cost XXXV K Raw Materials cost XXXVIII L Utility cost XXXIX M Labour cost XL N Fixed capital cost XLI O Rate of return XLII P Fire and Explosion Index XLIV Q Safety figures XLV R Process flow scheme XLVII

III


Appendix A

PURE COMPONENT PROPERTIESComponent Name Technological Data Health &Safety dataDesign Systematic Formula Mol. Phase Boiling Melting Flash Liquid Vapour Auto-ignition Flammable Lower Upper LC 50 MAC LD50

Weight Point Point Point Density Density Temp. Limits Explosion Explosion In air/ Value Oral[1] [1] [1] [2] [3] [1] % by vol Limit Limit water [4]

g/moloC oC oC oC in air % % mg/m3 mg/m3

g

Hydrogen Hydrogen H2 2.0 G -253 -259 - 0.07 0.07 560 - 4 76 - - -Water Water H2O 18.0 G 100 0Hydrogensulfide Hydrogensulfide H2S 34.1 G -60 -86 - 0.9 1.2 260 - 4.3 46 - 15 -Methane Methane CH4 16.0 G -162 -182 - - 0.60 537 - 4.4 16 - - -Carbonmonoxide Carbonmonoxide CO 28.0 G -191 -205 - - 0.97 605 - 11 75 - 29 -Carbondioxide Carbondioxide CO2 44.0 G -79 n.b. - - 1.5 - - - - - 9000 -Acetylene Acetylene C2H2 26.0 G -83.8 -80.6 - 0.40 0.90 305 - 2.3 80 - - -Ethylene Ethylene C2H4 28.1 G -104 -169.0 -136 0.57 1.00 425 - 2.3 34 - 330 -Ethane Ethane C2H6 30.1 G -88.6 -183.0 - 0.54 1.00 515 - 2.7 12.5 - - -Propylene Propylene C3H6 42.1 G -48.0 -185.2 -108 0.50 1.50 497 - 2.0 11.1 - 900 -Propane Propane C3H8 44.1 G -42.2 -187.6 -104 0.50 1.56 470 - 1.7 9.5 - - -1-Butylene 1-Butylene C4H8 56.1 G -6 -185 - 0.60 1.90 384 - 1.6 10 - - -2-Butylene 2-Butylene C4H8 56.1 G 2 -120 - 0.60 1.90 324 - 1.6 10 - - -Iso-Butylene Iso-Butylene C4H8 56.1 G -7.0 -140.0 - 0.60 1.99 465 - 1.8 9.6 - - -n-Butane n-Butane C4H10 58.1 G -0.5 -138 - 0.58 2.01 365 - 1.3 8.5 - 1430 -Iso-Butane Iso-Butane C4H10 58.1 G -12 -160 0.60 2.10 460 - 1.8 8.4 - - -Benzene Benzene C6H6 78.1 L 80 6.0 -11 0.90 2.70 555 - 1.2 8 - 3.25 3800Monoethanolamine Monoethanolamine C2H7NO 61.1 L 171 10 85 1.02 2.1 410 2.5 23.5 - 2.5

Notes:[1] At 101.3 kPa[2] Relative (water =1)[3] Relative (air =1)[4] Oral ingestion in (g) for a male of 70kg weight. Benzene oral rat mg/kg

Project ID Number: CPD3334Completion Date: 17 April 2006

IV


Appendix B

ORIGIN 1:= year 8400 hr⋅:=

The following factors are known:

Conversion of ethane: Selectivity of ethane to ethylene reaction:

Conversion 0.7:= Selectivity 0.9:=

Moleculair Weights

Methane 30.1gmmol

⋅:= Methylene 28.1gmmol

⋅:= Mhydrogen 2gmmol

⋅:=

We can define the production of ethylene:

methylene 1 106⋅

tonneyear

:= methylene 33.069kgs

=

The annual molar production:

nethylenemethyleneMethylene

:= nethylene 1.177 103× katal=

The same amount of ethane is needed to produce this amount of ethylene, however there is a selectivity. Therefore the amount of ethane converted is:

nethaneconvnethyleneSelectivity

:= nethaneconv 1.308 103× katal=

methaneconv 39.358kgs

= methaneconv nethaneconv Methane⋅:=

Only 70% of the ethane is converted, so the total amount of ethane feed is:

nethanenethaneconvConversion

:= nethane 1.868 103× katal=

methane 56.226kgs

= methane nethane Methane⋅:=

The 30% that is unconverted is then calculated by:

methaneunconv 16.868kgs

= methaneunconv methane methaneconv−:=

Because it is know that during the reaction for one mol of ethylene formed also 1 mol of hydrogen is formed, the approximated amount can be calculated:

nhydrogen nethylene:=

mhydrogen nhydrogen Mhydrogen⋅:= mhydrogen 2.354

kgs

=

V


In consultation with the supervisor it has been chosen to produce approximately 0.5 wt % of benzene measured at the reactor outlet. Because of the law of mass conservation the mass entering the reactor equals the mass leaving the reactor. Therefore it can be said that:

mbenzenemethane

1000.5⋅:= mbenzene 0.281

kgs

=

Because the law of mass conservation must be met the by-products can be calculated. Note that this is the total amount of by-product. The final distribution must be calculated during the modeling of the reactor.

mbyproducts methane methaneunconv methylene+ mhydrogen+ mbenzene+( )−:=

mbyproducts 3.655kgs

=

Steam calculation From the patent the ratio of steam is known:

Ratioethane 6.67:= Ratioethyl 10.2:=

However the ratio of steam per kg ethylene is for a selectivity of 100%. The selectivity here is 90% thus the ratio becomes:

RatioethyleneRatioethylSelectivity

:= Ratioethylene 11.33=

The amount of steam needed during the reaction is:

mwater1 methane Ratioethane⋅:= mwater2 methylene Ratioethylene⋅:=

mwater1 375kgs

= mwater2 375kgs

=

VI


Appendix C Diameter estimations

Physical constants

R 8.314J

mol K⋅⋅:= Mach 330

ms

:= Mw.water 18.01510 3−⋅kgmol

⋅:= Mw.ethane 30.0710 3−⋅kgmol

⋅:=

Known variables

P1 26.7bar:= V1 0.9Mach:= Positon 1: inlet

P2 1.02bar:= V2 2.97Mach:= Position 2: throat

Fsteam 375kgs

:= T1 1290 K⋅:=

Initial Values

ρP1 Mw.water⋅

R T1⋅:= Q

Fsteamρ

:= ∆P P1 P2−:= D2 15in:= D1 20in:=

Calculations Given

2 ∆P⋅

ρ

16 Q2⋅

π2

1

D24

1

D14

−

QFsteam

ρ

D2 4Q

V2 π⋅⋅ D1 4

QV1 π⋅

⋅ ∆P P1 P2−

D2

D1

Q

∆P

ρ

Find D2 D1, Q, ∆P, ρ,( ):=

Q 63.066m3

s= ρ 5.946

kg

m3= V1 297

ms

= V2 980.1ms

=

∆P 2.594 106× Pa= D1 51.996cm= D2 28.623cm=

Distance of tube α1 20 deg⋅:= α2 7deg:= D3 98cm:=

L1 D1 D2−( ) 12 tan α1( )⋅

⋅:= L2 D3 D2−( ) 12 tan α2( )⋅

⋅:=

L1 32.109cm= L2 282.515cm=

VII


Appendix D Energyloss over jet tube

Assumptions Steam acts as an ideal gas

The behavior of the tube can be best described by a polytropic compression Calculations The Cp-value is taken from the Matlab file for a temperature of 1282 K.

Cp 44.7J

mol K⋅⋅:=

For an ideal gas the following relations may be assumed:

Cv 36.3861

KmolJ= Cv Cp R−:=

κCpCv

:= κ 1.228=

A polytropic compression can be described by:

T2T1

P2P1

κ 1−( )κ

With all but T1 known

Tsteam.inTsteam

P2P1

κ 1−( )

κ

:= Tsteam.in 1290K=

Because the final temperature calculated is only 8 K higher than that of the

original temperature from the Matlab file it is assumed that the Cp-value is the

same at both temperatures.

VIII


Appendix E Mixing of the 2 gases

Temperature estimation

Fethane 56.2kgs

⋅:= Fsteam 375kgs

=

Tassumed 710 K⋅:=

nethaneFethane

Mw.ethane:= nwater

FsteamMw.water

:=

ntotal 2.268 104×

mols

= ntotal nethane nwater+:=

Total 1.611 107×

Kmols

= Total ntotal Tassumed⋅:=

If one temperature is fixed for one of the streams the other stream temperature can be

calculated. However the temperature of the ethane may not exceed the pyrolysis temperature. This is 515°C, which is well below the pyrolysis temperature) Tethane 788 K⋅:=

TsteamTotal Tethane nethane⋅−

nwater:= Tsteam 702.997K=

IX


Appendix F function [t,y] = CPD(Mps,Tps,Pps,ethanefeed);

% Mps = pre-shock Mach number [M], Tps = pre-shock Temperature [K]

% Pps = pre-shock Pressure [Pa], SDF = Steam dilution factor [-]

% tend = time span of the reactor

% p.ethanefeed is taken as pure ethane feed needed to produce 1*10^6 t/a

% ethylene.

% Mps = 2.8, Tps = 710 K, Pps = 1.02*10^5 Pa, SDF = 11.1, tend = 0.05 s

% INPUT SECTION:

% The spatial domain for the integration over the pyrolysis section

tspan = [0 0.05];

% feed specs

p.SDF = 11.1; % The steam dilution factor taken from U.S. patent 5,300,216

p.ethanefeed = ethanefeed; % [mol/s]

% feedratio = [CH4 C2H2 C2H4 C2H6 C3H6 C3H8 C4H6 H2 C6H6 H2O] This ratio

% is calculated over p.ethanefeed.

p.feedratio = [0.039 0 0.01 1 0.008 0.019 0 0 0 p.SDF]; %C2H4,C2H6,C3H6 ratio from Froment 79 model

p.Fmoli0 = p.ethanefeed*p.feedratio';

X


% reactor geometry parameters

p.alpha = 5; % angle of reactor tube widening [°]

%=========================================================================

% Constants for calculations

%=========================================================================

% constants for isobaric heat capacity (Cp=CpA+CpB*T+CpC*T^2+CpD*T^3)

p.CpA = [19.250 26.820 3.806 5.409 3.710 -4.224 -1.687 27.140 -33.917 32.240]'; % for i=1:9

[J/mol/K]

p.CpB = [ 5.213e-2 7.578e-2 1.566e-1 1.781e-1 2.345e-1 3.063e-1 3.419e-1 9.274e-3 47.436e-2

1.924e-3]'; % [J/mol/K^2]

p.CpC = [ 1.197e-5 -5.007e-5 -8.348e-5 -6.938e-5 -1.160e-4 -1.586e-4 -2.340e-4 -1.381e-5 -3.017e-4

1.055e-5]'; % [J/mol/K^3]

p.CpD = [-1.132e-8 1.412e-8 1.755e-8 8.713e-9 2.205e-8 3.215e-8 6.335e-8 7.645e-9 71.301e-9 -

3.569e-9]'; % [J/mol/K^4]

p.Tref = 298.15;

% constants for standard heat of formation (at 298.2 K and 1 atm)

p.DfH0i = [-7.490e4 2.269e5 5.234e4 -8.474e4 2.043e4 -1.039e5 1.102e5 0 8.298e4 -

2.420e5]'; % J/mol

XI


% component properties for viscosity correlation

p.Mwi = 1e-3*[16.043 26.038 28.054 30.070 42.081 44.094 54.092 2.016 78.114 18.015]'; % kg/mol

p.Pci = [46.0 61.4 50.4 48.8 46.0 42.5 43.3 12.9 48.9 221.2]'; % bar

p.Zci = [0.288 0.270 0.280 0.285 0.274 0.281 0.270 0.303 0.268 0.235]'; % [-] Formula: Zci = Pci*Vci/R*Tci with Vci

= critical volume

p.Tci = [190.4 308.3 282.4 305.4 364.9 369.8 425.0 33.0 562.1 647.3]'; % K

% kinetic parameters

p.k0j = [4.65e13 7.88e5 3.85e11 7.08e10 9.81e8 5.87e1 1.03e9 1.8e5 3.0e3]'; % [1/s or m^3/mol/s]

p.Eaj = 1e3*[273.02 136.87 273.19 253.01 154.58 29.48 172.75 75.00 100.50]'; % [J/mol]

p.R = 8.314; % [J/(mol*K)]

% stoechiometric reacting component matrix - j reaction rows and i component columns

p.Mstoech=[ 0 0 1 -1 0 0 0 1 0 0

0 0 -1 1 0 0 0 -1 0 0

1 0 0 -2 0 1 0 0 0 0

1 0 -1 -1 1 0 0 0 0 0

1 1 0 0 -1 0 0 0 0 0

-1 -1 0 0 1 0 0 0 0 0

0 -1 -1 0 0 0 1 0 0 0

0 -1 0 0 0 0 -1 1 1 0

0 1 0 0 0 0 1 -1 -1 0];

XII


% Reactions are mentioned below

% ========================================================================

% Pyrolysis section

% ========================================================================

shock=Fshock(Mps,Tps,Pps,p);

p.M0 = shock(1);

p.u0 = shock(2);

p.T0 = shock(3);

p.P0 = shock(4);

p.meanMw0 = shock(7);

p.meankappa = shock(10);

rhom = p.meanMw0*p.P0/(p.R*p.T0);

p.D0 = sqrt(4*(p.Fmoli0'*p.Mwi)/(rhom*p.u0*pi));

% compute the initial guess

y0 = CompInitialGuess(p);

M = zeros(81);

M([34 63 64], [34 63 64])=eye(3); % eye matrix with [z,T,Px] from pyro

XIII


M(22:31,22:31)=eye(10); % eye matrix with [Fmoli] from pyro

options = odeset('Mass', M, 'RelTol', 1e-5, 'AbsTol', 1e-5, ...

'Vectorized','on');

[t,y] = ode15s(@(t,y) ff(t,y,p) ,tspan, y0,options);

conversion = (100-(y(:,25)./p.ethanefeed)*100);

figure(1);

plot(t,y(:,1),'r',t,y(:,34),'b'),title('reactor diameter and length vs. residence time'),ylabel('reactor diameter/ length [m]'),xlabel('t [s]')

legend1 = legend('Diameter','Length',-1);

figure(2);

plot(t,conversion),title('conversion of ethane’),ylabel('conversion [%]'), xlabel('t [s]')

figure(3);

plot(t,y(:,25),'r',t,y(:,24),'b',t,y(:,29),'g'),title('ethane,ethylene and hydrogen vs. residence time'),ylabel('amount [mol/s]'),xlabel('t [s]')

legend3 = legend('Ethane','Ethylene','Hydrogen',-1);

figure(4);

plot(t,y(:,22),'r',t,y(:,23),'b',t,y(:,26),'k',t,y(:,27),'g',t,y(:,28),'-.',t,y(:,30),':'),title('byproducts vs. residence time'),ylabel('amount

[mol/s]'),xlabel('t [s]')

legend4 = legend('Methane','Acetylene','Propylene','Propane','Butadiene','Benzene',-1);

figure(5);

hl1 = line(t,y(:,63),'Color','r');

XIV


ax5 = gca;

set(ax5,'XColor','k','YColor','k');

ax6 = axes('Position',get(ax5,'Position'),'XAxisLocation','top','YAxisLocation','right','Color','none','XColor','k','YColor','k');

hl2 = line(t,1e-5*y(:,65),'Color','b','Parent',ax6);

% Built in check for conversion and selectivity in reaction. Behind the

% ethylene value the needed mol/s ethylene is given. This 1.170e+3 mol/s is

% a total amount of 1 million ton/a ethylene.

q = size(y);

uncon = y(q(1),25);

conversion = (100 - (uncon/p.ethanefeed)*100);

ethylene = (y(q(1),24));

Selectivity = (y(q(1),24)/(p.ethanefeed-uncon))*100;

fprintf('\n');

fprintf('conversion selectivity ethylene');

fprintf('\n%5.1f %% %5.1f %% %5.2f - 1.170e+3\n', conversion, Selectivity, ethylene);

fprintf('\n');

function [pyro] = ff(t,y,p);

D0 = p.D0;

XV


alpha = p.alpha;

Steamfeed = p.Fmoli0(10);

% kinetic parameters

k0j = p.k0j;

Eaj = p.Eaj;

R = p.R;

Mstoech = p.Mstoech;

CpA = p.CpA;

CpB = p.CpB;

CpC = p.CpC;

CpD = p.CpD;

Tref = p.Tref;

DfH0i = p.DfH0i;

Mwi = p.Mwi;

Pci = p.Pci;

Zci = p.Zci;

Tci = p.Tci;

% ~~~~~~~~ calculation part ~~~~~~~~

XVI


pyro = [0,0];

ny = size(y);

pyro = zeros(81,ny(2));

for i=1:ny(2);

D = y(1,i);

A = y(2,i);

c = y(3:12,i);

ratej = y(13:21,i);

Fmoli = y(22:31,i);

F = y(32,i);

u = y(33,i);

z = y(34,i);

Cpi = y(35:44,i);

DfHi = y(45:53,i);

DrHj = y(54:62,i);

T = y(63,i);

Px = y(64,i);

P = y(65,i);

meanMw = y(66,i);

rhom = y(67,i);

f = y(68,i);

XVII


Re = y(69,i);

eta = y(70,i);

etai = y(71:80,i);

kappa = y(81,i);

pyro(:,i) = [D0+2*z*tan(2*pi*alpha/360)-D; % D #1

0.25*pi*D^2-A; % A #2

(Fmoli/F)*(P/(R*T))-c; % c(1:9) #3-12

(k0j.*exp(-Eaj./(R*T))).*[c(4) c(3)*c(8) c(4) c(3)*c(4) c(5) c(1)*c(2) c(2)*c(3) c(2)*c(7) c(8)*c(9)]'-ratej; % rate(1:7) #13-21

u*(A*[ratej'*Mstoech(:,1:9)]'); % dFmoli(1:9)/dt #22-30

Steamfeed-Fmoli(10); % Fmoli(10) #31

sum(Fmoli)-F; % F #32

F*((R*T)/(P*A))-u; % u #33

u; % dz/dt #34

CpA+CpB*T+CpC*T^2+CpD*T^3-Cpi; % Cpi #35-44

DfH0i(1:9)+CpA(1:9)*(T-Tref)+CpB(1:9)*(T^2-Tref^2)/2+CpC(1:9)*(T^3-Tref^3)/3+CpD(1:9)*(T^4-Tref^4)/4-DfHi;% DfHi =

#45-53

Mstoech(:,1:9)*DfHi-DrHj; % DrHj #54-62

u*(-ratej'*DrHj*A/(Fmoli'*Cpi)); % dT/dt #63

u*(-(2*f+4*tan(2*pi*alpha/360))*rhom*u^2/D); % dPx/dt #64

Px-rhom*u^2-P; % P #66

Fmoli'*Mwi/F-meanMw; % meanMw #67

meanMw*P/(R*T)-rhom; % rhom #68

XVIII


0.046*Re^-0.2-f; % f %% oftewel 4f=0.316*Re^-0.25 %%% if Re>1e5, state f = 0.046*Re^-0.2 #69

rhom*u*D/eta-Re; % Re #70

Fmoli'*etai/F-eta; % eta #71

((1.9*T./Tci-0.29).^(4/5))*1E-7.*Zci.^(-2/3)./Tci.^(1/6).*sqrt(1e3*Mwi).*(Pci/1.0134).^(2/3)-etai; % etai = #72-81

(corresponding states method, from Froment)

kappa-Fkappa(T,Fmoli,p)]; % kappa #82

end

% ========================================================================

% Heat Capacity

% ========================================================================

function kappa = Fkappa(T,Fmoli, p)

F=sum(Fmoli);

Cpi = p.CpA+p.CpB*T+p.CpC*T^2+p.CpD*T^3;

Cp = Fmoli'*Cpi/F;

kappa = Cp/(Cp-p.R);

% ========================================================================

% Initial Shock Values

% ========================================================================

function shock = Fshock (Mps, Tps, Pps, p)

XIX


p.normfeedratio = p.feedratio/sum(p.feedratio);

meanMw0 = p.normfeedratio*p.Mwi;

kappaps = Fkappa(Tps,p.Fmoli0,p);

kappa = kappaps;

aps = sqrt(kappa*p.R*Tps/meanMw0);

eps = 1;

reltol = 0.000005;

while eps>reltol;

M0 = sqrt((Mps^2+2/(kappa-1))/(2*kappa/(kappa-1)*Mps^2-1));

P0 = Pps*(1+kappa*Mps^2)/(1+kappa*M0^2);

T0 = Tps*(P0*M0/(Pps*Mps))^2;

kappa0 = Fkappa(T0,p.Fmoli0,p);

meankappa = (kappaps + kappa0)/2;

eps = abs((meankappa-kappa)/kappa);

kappa = meankappa;

end

a0 = sqrt(kappa0*p.R*T0/meanMw0);

u0 = a0*M0;

XX


shock = [M0; u0; T0; P0; a0; aps; meanMw0; kappaps; kappa0; meankappa];

% ========================================================================

% Initial guess of the SWR

% ========================================================================

function [y0] = CompInitialGuess(p);

z =0; % [m]

T = p.T0; % [K]

P = p.P0; % [Pa]

D = p.D0; % [m]

A = 0.25*pi*D^2; % [m^2]

Fmoli = p.Fmoli0; % [mol/s]

F = sum(Fmoli); % [mol/s]

c = (Fmoli/F)*(P/(p.R*T)); % c(1:10) [mol/m^3]

ratej = (p.k0j.*exp(-p.Eaj./(p.R*T))).*[c(4) c(3)*c(8) c(4)^2 c(3)*c(4) c(5) c(1)*c(2) c(2)*c(3) c(2)*c(7) c(8)*c(9)]'; % rate(1:9)

[mol/(m^3*s)]

u = F*((p.R*T)/(P*A)); % [m/s]

Cpi = p.CpA + p.CpB*T + p.CpC*T^2 + p.CpD*T^3; % [J/(mol*K)]

DfHi = p.DfH0i(1:9) + p.CpA(1:9)*(T-p.Tref) + p.CpB(1:9)*(T^2-p.Tref^2)/2 + p.CpC(1:9)*(T^3-p.Tref^3)/3 + p.CpD(1:9)*(T^4-

p.Tref^4)/4; % DfHi(1:9) [J/mol]

DrHj = p.Mstoech(:,1:9)*DfHi; % [J/mol]

XXI


etai = ((1.9*T./p.Tci-0.29).^(4/5))*1e-7.*p.Zci.^(-2/3)./p.Tci.^(1/6).*sqrt(1e3*p.Mwi).*(p.Pci/1.0314).^(2/3); % [kg/(m*s)] from AlChE

vol. 8 nr 1. page 59 (Jossi, Stiel, Thodos)

eta = Fmoli'*etai/F; % [kg/(m*s)]

meanMw = Fmoli'*p.Mwi/F; % [kg/mol]

rhom = meanMw*P/(p.R*T); % [kg/m^3]

Re = rhom*u*D/eta; % [-]

f = 0.046*Re^-0.2; % f = %% oftewel 4f=0.316*Re^-0.25 %%% if Re>1e5, state f = 0.046*Re^-0.2

Px = P+rhom*u^2; % [Pa] (normaal gevonden met diff. eq. maar nu uit relatie voor P en waarde van P)

kappa = Fkappa(T,Fmoli,p); % [-]

y0 = [D; A; c; ratej; Fmoli; F; u; z; Cpi; DfHi; DrHj; T; Px; P; meanMw; rhom; f; Re; eta; etai;kappa];

% y0 = [D=1; A=2; c=3-12; ratej=13-21; Fmoli=22-31; F=32; u=33; z=34;

% Cpi=35-44; DfHi=45-53; DrHj=54-62; T=63; Px=64; P=65; meanMw=66;

% rhom=67; f=68; Re=69; eta=70; etai=71-80; kappa=81]

% ~~~~~~~~~~~~~~~

% Some information to see what is what:

% comp # component

% 1 CH4 methane y(22)

% 2 C2H2 acetylene y(23)

% 3 C2H4 ethylene y(24)

% 4 C2H6 ethane y(25)

% 5 C3H6 propylene y(26)

XXII


% 6 C3H8 propane y(27)

% 7 C4H6 butadiene y(28)

% 8 H2 hydrogen y(29)

% 9 C6H6 benzene y(30)

% 10 H20 steam y(31)

% reaction# reaction

% 1 C2H6 -> C2H4 + H2

% 2 C2H4 + H4 -> C2H6

% 3 2C2H6 -> C3H8 + CH4

% 4 C2H6 + C2H4 -> C3H6 + CH4

% 5 C3H6 -> C2H2 + CH4

% 6 C2H2 + CH4 -> C3H6

% 7 C2H2 + C2H4 -> C4H6

% 8 C4H6 + C2H2 -> C6H6 + H2

% 9 C6H6 + H2 -> C4H6 + C2H2

% ~~~~~~~~~~~~~~~

XXIII


Appendix G

XXIV


Appendix H

COMP MW Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole FracHydrogen 2.02 - - - - - - trace trace - - - -Methane 16.04 - - - - - - trace trace - - - -W ater 18.02 - - - - 1.00 1.0000 1.00 1.0000 - - - -Acetylene 26.04 - - - - - - trace trace - - - -Carbonmonoxide 28.01 - - - - - - trace trace - - - -Ethylene 28.05 - - - - - - trace trace - - - -Ethane 30.07 1.00 1.0000 1.00 1.0000 - - trace trace 1.00 1.0000 1.00 1.0000Hydrogensulfide 34.08 - - - - - - trace trace - - - -Propylene 42.08 - - - - - - trace trace - - - -Carbondioxide 44.01 - - - - - - trace trace - - - -Propane 44.10 - - - - - - trace trace - - - -1,3-Butadiene 54.09 - - - - - - trace trace - - - -Monoethanolamine 61.08 - - - - - - trace trace - - - -Benzene 78.11 - - - - - - trace trace - - - -

Total 1.00 1.0000 1.00 1.0000 1.00 1.0000 1.00 1.0000 1.00 1.0000 1.00 1.0000PhasePressure barTemperature K

COMP MW Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole FracHydrogen 2.02 - - trace trace trace trace trace trace trace trace 0.01 0.0512Methane 16.04 - - trace trace trace trace trace trace trace trace trace 0.0004W ater 18.02 - - 1.00 1.0000 1.00 1.0000 1.00 1.0000 1.00 1.0000 0.87 0.8704Acetylene 26.04 - - trace trace trace trace trace trace trace trace trace 395 ppmCarbonmonoxide 28.01 - - trace trace trace trace trace trace trace trace trace 13 ppbEthylene 28.05 - - trace trace trace trace trace trace trace trace 0.08 0.0485Ethane 30.07 1.00 1.0000 trace trace trace trace trace trace trace trace 0.04 0.0228Hydrogensulfide 34.08 - - trace trace trace trace trace trace trace trace trace 3 ppbPropylene 42.08 - - trace trace trace trace trace trace trace trace trace 280 ppmCarbondioxide 44.01 - - trace trace trace trace trace trace trace trace 0.00 17 ppbPropane 44.10 - - trace trace trace trace trace trace trace trace 0.00 446 ppm1,3-Butadiene 54.09 - - trace trace trace trace trace trace trace trace 0.00 0.0014Monoethanolamine 61.08 - - trace trace trace trace trace trace trace trace - -Benzene 78.11 - - trace trace trace trace trace trace trace trace trace 140 ppm


7.6300 289 300 359 278 63810 10 1 1

W ater feed P01 Ethane feed E03

7.6

Ethane feed E01

V V L L V V

Name : Fresh ethane feed T01 Ethane feed T01 Fresh water feed V01

1290 970

Process Stream SummarySTREAM Nr. : 1 IN 1A = 1 + 61 2 IN 2A = 2 + 64 3 4

788 359 638 863

V7.6 48.5 48.5 48.5 26.7 10V L V

10Name : Ethane feed R01 W ater feed E04 W ater feed E01 W ater feed F01 W ater feed R01 R01 mix E02

8 9

V V

STREAM Nr. : 5 6 7

XXV


COMP MW Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole FracHydrogen 2.02 0.01 0.0512 0.01 0.0512 0.01 0.0512 0.01 0.0512 0.01 0.0512 0.01 0.0512Methane 16.04 trace 0.0004 trace 0.0004 trace 0.0004 trace 0.0004 trace 0.0004 trace 0.0004Water 18.02 0.87 0.8704 0.87 0.8704 0.87 0.8704 0.87 0.8704 0.87 0.8704 0.87 0.8704Acetylene 26.04 trace 395 ppm trace 395 ppm trace 395 ppm trace 395 ppm trace 395 ppm trace 395 ppmCarbonmonoxide 28.01 trace 13 ppb trace 13 ppb trace 13 ppb trace 13 ppb trace 13 ppb trace 13 ppbEthylene 28.05 0.08 0.0485 0.08 0.0485 0.08 0.0485 0.08 0.0485 0.08 0.0485 0.08 0.0485Ethane 30.07 0.04 0.0228 0.04 0.0228 0.04 0.0228 0.04 0.0228 0.04 0.0228 0.04 0.0228Hydrogensulfide 34.08 trace 3 ppb trace 3 ppb trace 3 ppb trace 3 ppb trace 3 ppb trace 3 ppbPropylene 42.08 trace 280 ppm trace 280 ppm trace 280 ppm trace 280 ppm trace 280 ppm trace 280 ppmCarbondioxide 44.01 0.00 17 ppb 0.00 17 ppb 0.00 17 ppb 0.00 17 ppb 0.00 17 ppb 0.00 17 ppbPropane 44.10 0.00 446 ppm 0.00 446 ppm 0.00 446 ppm 0.00 446 ppm 0.00 446 ppm 0.00 446 ppm1,3-Butadiene 54.09 0.00 0.0014 0.00 0.0014 0.00 0.0014 0.00 0.0014 0.00 0.0014 0.00 0.0014Monoethanolamine 61.08 - - - - - - - - - - - -Benzene 78.11 trace 140 ppm trace 140 ppm trace 140 ppm trace 140 ppm trace 140 ppm trace 140 ppm


COMP MW Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole FracHydrogen 2.02 0.01 0.0512 0.01 0.0512 0.01 0.0512 0.01 0.0512 - - 0.01 0.0512Methane 16.04 trace 0.0004 trace 0.0004 trace 0.0004 trace 0.0004 - - trace 0.0004Water 18.02 0.87 0.8704 0.87 0.8704 0.87 0.8704 0.87 0.8704 - - 0.87 0.8704Acetylene 26.04 trace 395 ppm trace 395 ppm trace 395 ppm trace 395 ppm - - trace 395 ppmCarbonmonoxide 28.01 trace 13 ppb trace 13 ppb trace 13 ppb trace 13 ppb - - trace 13 ppbEthylene 28.05 0.08 0.0485 0.08 0.0485 0.08 0.0485 0.08 0.0485 - - 0.08 0.0485Ethane 30.07 0.04 0.0228 0.04 0.0228 0.04 0.0228 0.04 0.0228 - - 0.04 0.0228Hydrogensulfide 34.08 trace 3 ppb trace 3 ppb trace 3 ppb trace 3 ppb - - trace 3 ppbPropylene 42.08 trace 280 ppm trace 280 ppm trace 280 ppm trace 280 ppm - - trace 280 ppmCarbondioxide 44.01 0.00 17 ppb 0.00 17 ppb 0.00 17 ppb 0.00 17 ppb - - 0.00 17 ppbPropane 44.10 0.00 446 ppm 0.00 446 ppm 0.00 446 ppm 0.00 446 ppm - - 0.00 446 ppm1,3-Butadiene 54.09 0.00 0.0014 0.00 0.0014 0.00 0.0014 0.00 0.0014 - - 0.00 0.0014Monoethanolamine 61.08 - - - - - - - - - - - -Benzene 78.11 trace 140 ppm trace 140 ppm trace 140 ppm trace 140 ppm - - trace 140 ppm


2873 648 648 513 298 64810 2 2 2

R01 mix E07 R01 mix

2

R01 mix E04

V V V V V/L V

Name : R01 mix T02 R01 mix R01 mix E03

298

Process Stream SummarySTREAM Nr. : 11 12 = 13 + 16 13 14 15 16

378 374 298 298

V/L2 2 2 2 2V V/L V/L

22Name : R01 mix E05 R01 mix E06 R01 mix R01 mix V02 Bleed V01 mix C01

20 = 15 + 19 21

V/L

STREAM Nr. : 17 18 19

XXVI


COMP MW Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole FracHydrogen 2.02 trace 454 ppb 0.04 0.3914 trace 29 pbb - - 0.07 0.5516 0.04 0.3804Methane 16.04 trace 6 ppb 0.03 0.0332 trace 1 ppb - - 0.02 0.0200 0.03 0.0323Water 18.02 1.00 0.9996 0.01 0.0122 0.85 0.9504 0.85 0.9505 - - 0.04 0.0398Acetylene 26.04 trace 399 ppb trace 0.0030 trace trace - - 0.02 0.0160 trace 0.0029Carbonmonoxide 28.01 trace trace trace 103 ppb trace trace - - trace 311 ppb trace 100 ppbEthylene 28.05 0.00 5 ppm 0.58 0.3706 trace 9 ppb - - 0.29 0.1728 0.56 0.3602Ethane 30.07 0.00 7 ppm 0.29 0.1745 trace 1 ppb - - 0.05 0.0273 0.28 0.1696Hydrogensulfide 34.08 trace trace trace 25 ppb trace 2 ppb - - 0.06 0.0313 - -Propylene 42.08 0.00 1 ppm trace 0.0021 trace trace - - trace 187 ppm trace 0.0021Carbondioxide 44.01 trace trace trace 131 ppb trace 9 ppb - - 0.42 0.1627 - -Propane 44.10 0.00 4 ppm 0.01 0.0034 trace trace - - trace 62 ppm 0.01 0.00331,3-Butadiene 54.09 0.00 226 ppm 0.03 0.0095 trace trace - - 0.03 0.0083 0.03 0.0092Monoethanolamine 61.08 - - - - 0.15 0.0496 0.15 0.0495 trace 50 ppmBenzene 78.11 trace 0.0002 trace 5 ppm trace trace - - 0.05 0.0098 trace 4 ppm


COMP MW Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole FracHydrogen 2.02 - - 0.04 0.3962 0.04 0.3945 0.04 0.3945 0.57 0.9173 trace traceMethane 16.04 - - 0.03 0.0336 0.03 0.0337 0.03 0.0337 0.39 0.0784 trace 819 ppbWater 18.02 1.00 1.0000 - - - - - - - - - -Acetylene 26.04 - - trace 153 ppm trace 153 ppm trace 153 ppm trace 600 ppb trace 268 ppmCarbonmonoxide 28.01 - - trace 104 ppb trace 104 ppb trace 104 ppb trace 242 ppb trace traceEthylene 28.05 - - 0.58 0.3752 0.59 0.3792 0.59 0.3792 0.04 0.0042 0.63 0.6620Ethane 30.07 - - 0.29 0.1767 0.29 0.1772 0.29 0.1772 trace 29 ppm 0.32 0.3108Hydrogensulfide 34.08 - - - - - - - - - - - -Propylene 42.08 - - 0.01 0.0022 trace 0.0022 trace 0.0022 trace trace trace 0.0038Carbondioxide 44.01 - - - - - - - - - - - -Propane 44.10 - - 0.01 0.0034 trace 0.0034 trace 0.0034 trace trace trace 0.00601,3-Butadiene 54.09 - - 0.03 0.0096 0.03 0.0096 0.03 0.0096 trace trace 0.03 0.0169Monoethanolamine 61.08 - - trace 52 ppm trace 52 ppb trace 52 ppb trace trace trace 91 ppmBenzene 78.11 - - trace 5 ppm trace 5 ppm trace 5 ppm trace trace trace 8 ppm


Process Stream Summary

STREAM Nr. :Name :

L2

298

V L2

2982 2

313 315 317

3602 20

3602

3172

317

L V V

L V V V

31522

30134

V L20

251

Bottom C02 to C03 Discharge P03 to C02 Discharge C03 Discharge C02 to C04

Bottom C04 Discharge C04 to R02 Discharge R02 Discharge P05 to C05 Discharge C05 to F02 Bottom C05 to C06

Bottom C01 to C08 Discharge C01 to C0233 34

35 36 43 44

23 24 25 28STREAM Nr. :Name :

37 38

XXVII


COMP MW Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole FracHydrogen 2.02 - - - - - - - - - - - -Methane 16.04 trace 1 ppm trace 1 ppm trace 1 ppm trace trace - - - -Water 18.02 - - - - - - - - - - - -Acetylene 26.04 trace 207 ppm trace 207 ppm trace 207 ppm trace 382 ppm trace 12 ppm trace 414 ppmCarbonmonoxide 28.01 - - - - - - - - - - - -Ethylene 28.05 1.00 0.9991 1.00 0.9991 1.00 0.9991 0.03 0.0361 trace 426 ppb 0.04 0.0393Ethane 30.07 trace 719 ppm trace 719 ppm trace 719 ppm 0.85 0.8868 0.02 0.0409 0.96 0.9603Hydrogensulfide 34.08 - - - - - - - - - - - -Propylene 42.08 trace trace trace trace trace trace 0.01 0.0110 0.19 0.1357 trace 25 ppbCarbondioxide 44.01 - - - - - - - - - - - -Propane 44.10 trace trace trace trace trace trace 0.02 0.0170 0.12 0.2153 trace 2 ppb1,3-Butadiene 54.09 trace trace trace trace trace trace 0.08 0.0484 0.66 0.6045 trace traceMonoethanolamine 61.08 trace trace trace trace trace trace trace 260 ppm 0.00 0.0033 - -Benzene 78.11 trace trace trace trace trace trace 0.00 23 ppm trace 291 ppm trace trace


COMP MW Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole FracHydrogen 2.02 - - - - - - trace trace trace trace trace traceMethane 16.04 - - - - - - trace trace trace trace trace traceWater 18.02 - - - - - - 1.00 1.0000 1.00 1.0000 1.00 1.0000Acetylene 26.04 - - trace 414 ppm trace 414 ppm trace trace trace trace trace traceCarbonmonoxide 28.01 - - trace trace trace trace trace traceEthylene 28.05 - - 0.04 0.0393 0.04 0.0393 trace trace trace trace trace traceEthane 30.07 - - 0.96 0.9603 0.96 0.9603 trace trace trace trace trace traceHydrogensulfide 34.08 - - - - - - trace trace trace trace trace tracePropylene 42.08 - - trace 25 ppb trace 25 ppb trace trace trace trace trace traceCarbondioxide 44.01 - - - - - - trace trace trace trace trace tracePropane 44.10 - - trace 2 ppb trace 2 ppb trace trace trace trace trace trace1,3-Butadiene 54.09 - - trace trace trace trace trace trace trace trace trace traceMonoethanolamine 61.08 - - - - - - - - - - - -Benzene 78.11 - - trace trace trace trace trace trace trace trace trace trace

Total 0.00 0.0000 1.00 1.0000 1.00 1.0000 1.00 1.0000 1.00 1.0000 1.00 1.0000PhasePressure barTemperature K 375 375

L L1 1

Name :63 OUT 64

Wast Water Water Recycle

Name :

Process Stream Summary

STREAM Nr. :

STREAM Nr. :

V8

215

V10

231

L5

287

V5

220

V5

220 262

L1

375

V10

V10

303

L8

235

Discharge C06 Discharge K01 Etylene Product Bottom C06 to C07 Bottom C07 to F01 Discharge C07 to K02

Purge Ethane Recycle Discharge K02 Bottom C0861 6259 60

53 5849 50 51 OUT 52

XXVIII


COMP MW Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole Frac Mass Frac Mole FracHydrogen 2.02 trace 45 ppm 0.57 0.9173 0.57 0.9173 0.03 0.1816 1.00 1.0000Methane 16.04 trace 626 ppb 0.39 0.0784 0.39 0.0784 0.89 0.7761 - -Water 18.02 0.87 0.9597 - - - - - - - -Acetylene 26.04 trace 40 ppm trace 600 ppb trace 600 ppb trace 6 ppm - -Carbonmonoxide 28.01 trace trace trace 242 ppb trace 242 ppb trace 2 ppm - -Ethylene 28.05 trace 473 ppm 0.04 0.0042 0.04 0.0042 0.08 0.0420 - -Ethane 30.07 trace 679 ppm trace 29 ppm trace 29 ppm trace 283 ppm - -Hydrogensulfide 34.08 trace 3 ppb - - - - - - - -Propylene 42.08 trace 131 ppm trace trace trace trace trace trace - -Carbondioxide 44.01 trace trace - - - - - -Propane 44.10 trace 356 ppm trace trace trace trace trace trace - -1,3-Butadiene 54.09 0.06 0.0226 trace trace trace trace - - - -Monoethanolamine 61.08 - - - - - - - - - -Benzene 78.11 0.06 0.0161 trace trace trace trace - - - -

Total 1.00 1.0000 1.00 1.0000 1.00 1.0000 1.00 1.0000 1.00 1.0000PhasePressure barTemperature K

Process Stream SummarySTREAM Nr. : 69 OUT 70 71 72 73 OUT

Name : Benzene Discharge F02 Discharge T03 to C09 Discharge C09 to F01 Hydrogen Product

1 20 2 2 2V V V

157 157

VV

157374 243

XXIX


EQUIPMENT NR.: C-01 C-01 C-03 C-04 C-08

NAME : Flasher AbsorberSour gas removal

StripperAmine regenaration

DryerSeparating residual

mositure

Distillation (Water Benzene)

Vertical Vertical Vertical Vertical

Pressure [bar] : 2 2 2 2 1

Temp. [K] : 298 315 316.85 316.65 373.23/375.23

Volume [m3] : 2441 444 44.18 1446 76.58

Reflux ratio 15.1 63

- Tray Number : - - 30

- Catalyst - -

Number- Series : 1 1 1 1

- Parallel : - -

Materials of C i

SS SS SS SS

XXX


EQUIPMENT NR.: R-01 R-02 C-05 C-06 C-07

NAME : ReactorPyrolysis

ReactorAcetylene

hydrogenation

Distillation (demethanizer)

Distillation (deethylenizer)

Distillation (deethanizer)

Horizontal Vertical Vertical Vertical Vertical

Pressure [bar] : 9.728 2 20 8 5

Temp. [K] : 1127 360 134.35/251.85 214.85/237.85 226.95/294.05

Volume [m3] : 32.82 930 4.52 75.1 55.42

Length [m] : 15.97

Reflux ratio 3.5 2.5 3.7

- Tray Number : 20 41 20

- Catalyst Pd/Al2O3

Number- Series : 1 1 1 1 1

- Parallel :Materials of C i

SS SS SS SS SS

Remarks: 1. In the distillation columns the two temperatures denote the top and bottom column temperature respectively

XXXI


Heaters/CoolersEquipment nr: F-01 E-02 E-06 E-07Equipment name Furnace Cooler Cooler CoolerDuty in MW 335.5 117 62 51.6Heat exchange areaPressureTemperature In - Out (K) 863.15 - 1293.15 970.65 - 873.15 374.15 - 298.15 513.15 - 288.15Material of Construction SS SS SS SS

HeatexchangersEquipment nr: E-01 E-03 E-04 E-06Equipment name: Quencher Xchanger Xchanger XchangerDuty 175.5+13.5 32.4 218 3.5Heat exchange coefficient

Steam: 900Ethane: 400 400 900 400

Heat Exchange Area [m2]

Steam: 655.52Ethane: 200.38 1136.52 16789.57 83.04

LMTD Steam: 297.74Ethane: 336.24 71.27 14.43 105.37

Tin - Tout (K) ethane: 638.15 - 788.15 ethane: 278.15 - 638.15 water: 358.15 - 638.15 reactout2: 378.15 - 374.15Tin - Tout (K) steam: 638.15 - 863.15 reactout1: 648.15 - 513.15 reactout2: 648.15 - 378.15 ethyleen: 231.15 - 303.15Material SS SS SS SS

XXXII


Appendix I

XXXIII

873

231

13.5 MW

32.4

303

367.5 MW

511 MW

648 K

218 MW

306 MW

FCp (MW/K)

1293

358

278

788

298

648 638 K

3.5 MW

1128

1.05 0.78 0.048 0.09 1.2

5 4 3 2 1


XXXIV

1 2 3 4 5

FCp (MW/K)

0.09 0.78 0.048 0.81 0.24 0.457 0.69 0.053

1293 1128 511

175.5 335.5 335.5

0

873 788 13.5

13.50

117117

0

175.5175.5

0

13.513.5

0

648 K 638 K 648

32.432.4

0

218218

0

358

298 303 283.5

218.065.5

3.562.062.0

0

84.032.451.651.6

0

3.53.5

0278

231


Appendix J Purchased Equipment Cost £ 1 = 1.8355 $ 12-5-2006

ethylene: 1.00E+06 t/a € 1 = 1.2885 $ 15-5-2006working hr 8400 3.31E+01 kg/s

nr Equipment1 Expander ethane kg/hr m3/h with electromotor € in 2006 $ $*1000

7015.714 kmol/hr 210962.52 31023.90 --> € 2002 667,854 875,421 1,127,980 1,1282 Pump water 3 pompen 1 pump

375.00 kg/s 5943.87225 gal/min 1981.29075 3620 10,860 10.863 Expander reactor outlet kg/hr m3/h with electromotor € in 2006 $ $*1000

89735.689 kmol/hr 1.62E+06 666260.55 --> € 2002 7,197,930 9,435,017 12,157,020 12,157

4 Compressor ethylene kg/hr m3/h with electromotor € in 2006 $ $*10004285 kmol/hr 1.20E+05 7029.89 --> € 2002 211,287 276,955 356,856 357

5 Compressor Ethane m3/h with electromotor € in 2006 $ $*1000recycle 61295.94 kg/hr 9014.11 --> € 2002 256,197 335,823 432,707 433

6 Expander methane/ m3/h with electromotor € in 2006 $ $*1000hydrogen 19009.64 kg/hr 5002.54 --> € 2002 162,305 212,748 274,126 274

T in T out Tlm U Q A [m2] € in 2002 € in 2006 $ $*10007 Quencher part 1 855 697.5 297.47 900 1.76E+08 655.52 172,904.53 226,642.57 292,029 292

reactor out - steam 365 5908 Quencher part 2 855 697.5 336.24 400 1.35E+07 100.38 39,948.03 52,363.72 67,471 67

reactor out - ethane 365 5159 Heat-exchanger 3 375 105 14.43 900 2.18E+08 16789.57 4,037,008.83 5,291,695.06 6,818,349 6,818

reactor out exp - H2O 85 36510 Heat-exchanger 4 375 240 71.27 400 3.24E+07 1136.52 288,104.58 377,646.33 486,597 487

reactor out exp - ethane 5 36511 Heat-exchanger 5 105 101 105.37 400 3.50E+06 83.04 35,796.61 46,922.05 60,459 60

reactor out exp - ethylene -42 30

MW kW 10^6*BTU/hr in 1990 $ $ $*100012 Heater 335.5 3.36E+05 1145.79 11,428,571.43 33,739,014 33,739

Water13 Heater 4.38 4.38E+03 14.96 500,000.00 1,476,082 1,476

Hydrogen/ Methane outlet14 Cooler 117 1.17E+05 399.58 8,000,000.00 23,617,310 23,617

Reactor outlet15 Cooler 62 6.20E+04 211.74 4,000,000.00 11,808,655 11,809

Reactor out exp. 116 Cooler 51.6 5.16E+04 176.22 3,500,000.00 10,332,573 10,333

Reactor out exp. 2

XXXV


Size unit S C n Ce [£ in 92] £ in 2006 $ $*1000

17 Flash drum 2441 m3 2441 1250 0.6 134,726 347,395 637,644 63818 Absorber 444 m3 444 1250 0.6 48,455 124,943 229,333 22919 Dryer 1446 m3 1446 1250 0.6 98,404 253,737 465,734 46620 Distillation 1 4.52 m3 5 1250 0.6 3,283 8,466 15,539 16

Demethanizer21 Distillation 2 75.10 m3 76 1250 0.6 16,803 43,328 79,529 80

Deethylene22 Distillation 3 44.18 m3 45 1250 0.6 12,270 31,638 58,072 58

Regeneration23 Distillation 4 55.42 m3 56 1250 0.6 13,990 36,074 66,214 66

Deethanizer24 Distillation 5 76.58 m3 77 1250 0.6 16,936 43,669 80,155 80

Benzene removal

€ in 2002 € in 2006 $ $*100025 Reactor 1 425,000.00 557,088.30 717,808 718

SWRSize unit S C n Ce [£ in 92] £ in 2006 $ $*1000

26 Reactor 2 930 m3 930 1250 0.6 75,509 194,703 357,377 357Acetylene removal

Total PCE 105,765,494 105,765

Item numbers 1,3,5,6,7,8,9, and 20 are found from Dutch Association of Cost Engineers.Item numbers 2, 10 and 11 are found with the help of Peters and Timmerhaus.The remaining item numbers are found from Coulson.

Dutch Association of Cost Engineers, "Prijzenboekje", 22th ed., Elsevier, may 2002Peters, Max S., Timmerhaus, Klaus D., "Plant Design and Economics for Chemical Engineers", 4th ed., McGraw-Hill, 1991Sinnott, R.K. , "Coulson & Richardsons's Chemical Engineering Vol. 6," 3th ed., Butterworth-Heinemann, Oxford, 1999

On the next page all the equipment is listed with their corresponding cost. This will show the total cost for purchased equipment somewhat easier than the calculation above.

XXXVI


nr Equipment Cost$*1000

1 Ethane Expander 1,1282 Water pump 113 Reactor outlet expander 12,1574 Ethylene compressor 3575 Ethane recycle compressor 4336 Hydrogen expander 2747 Quencher 3598 Heat-exchanger 3 6,8189 Heat-exchanger 4 487

10 Heat-exchanger 5 6011 Water heater 33,73912 Hydrogen/methane heat 1,47613 Reactor outlet cooler 23,61714 Reactor out exp. 1 11,80915 Reactor out exp. 2 10,33316 Flash drum 63817 Absorber 22918 Dryer 46619 Demethanizer 1620 Ethylene removal 8021 Regenerator 5822 Deethanizer 6623 Benzene removal 8024 SWR reactor 71825 Acetylene removal 357

Total 105,765

XXXVII


Appendix K CALCULATION RAW MATERIALS COST

It must be stated that this are the needed streams with the recycle implemented. +/- 70% of new ethane and +/- 20 % of new water needed

kg/hr kg/a t/a $/aWater for steam 294273.668 2471898811 2471898.811 1,669,976Ethane 149666.576 1257199238 1257199.238 188,579,886Pd/Al2O3 cat 2 reactors used 1 reactor needs: 309.86 m3 cat

m3 cat kg/m3 cat kg/a t/a $/a619.73 4371.5 2709147.514 2709.147514 32,812,924

Total 223,062,785First year total amount of catalyst, rest 15% fresh for 9 years

XXXVIII


Appendix L CALCULATION UTILITIES COST

MW kJ/s kg/s m3/a $/aIn Process cooling 230.60 2.31E+05 2.76E+03 83,413,206 10,566,122

MW $/MMBTU $/aprocess heat 852.72 5.40 125,379,708

MW kW kWh $/aOut Total electricity needed 233.25 2.33E+05 8.40E+08 67,177,042

from to € mean € 2002 mean € 06 mean $ 06gas price 0.13 0.27 0.20 0.26 0.34 m3

31.65 MJ/m3electricity 0.08 $/kWhwater 4.18 kJ/kg°C

Sub-total 135,945,830 invest67,177,042 back

once boughtkmol/hr needed 2x kg t $

MEA 8904.83 17809.66 1088170.226 1088.170226 291,477 lifetime MEAWater with MEA 171095.169 342190.338 6164558.939 6164.558939 4,165 1

295,642 roughly M$/aTotal 136,241,472 0.30

Mol.W.MEA 61.1 kg/kmolwater 18.015 kg/kmol

XXXIX


Appendix M CALCULATION LABOUR COST

5 man per shift3 shifts a day8 hrs per shift

6.538 £/hr

cost a day 784.56 £/day5 men* 3 shifts* 8 hr/shift* wage

days operating 350cost/annum 274596 £/a

708,055 £/a1,299,635 $/a

XL


Appendix N ESTIMATION OF PROJECT FIXED CAPITAL COST

Table 6.1 from Coulson and Richardson $1. Major equipment, total purchase cost PCE 105,765,494f1 Equipment erection 0.40 42,306,198f2 Piping 0.70 74,035,846f3 Instrumentation 0.20 21,153,099f4 Electrical 0.10 10,576,549f5 Buildings, process 0.15 15,864,824f6 Utilities 0.50 52,882,747f7 Storages 0.15 15,864,824f8 Site development 0.05 5,288,275f9 Ancillary buildings 0.15 15,864,824

2. Total physical plant cost PCC 359,602,680f10 Design and Engineering 0.30 107,880,804f11 Contractor's fee 0.05 17,980,134f12 Contingency 0.10 35,960,268

Fixed capital 521,423,886 521.42 M$

Table 6.6 from Coulson and RichardsonVariable costs $/a

1 Raw materials 223,062,7852 Miscellaneous materials 5,214,2393 Utilities 136,241,4724 Shipping and packaging 0

Sub-total A 364,518,496Fixed costs

5 Maintenance 52,142,3896 Operating labour 1,299,6357 Laboratory costs 259,9278 Supervision 259,9279 Plant overheads 649,818

10 Capital charges 78,213,58311 Insurance 5,214,23912 Local taxes 10,428,47813 Royalties 5,214,239

Sub-total B 153,682,234Direct production costs A+B 518,200,730

14 Sales expense 51,820,073 0.115 General overheads 51,820,073 0.116 Research and development 51,820,073 0.1

Sub-total C 155,460,219

Annual production cost A+B+C 673,660,949 673.66 M$

Production cost $/kg 0.67

XLI


Appendix O CALCULATION DCFROR AND OTHERS Chapter 6.10 from Coulson and Richardson

Production cost 0.67 $/kgannual cost 673,660,949 $/a 674 M$/a

Gross incomeamount t/a $/a M$/a $/ton ethylene

Ethylene 1.00E+09 kg/a 1.00E+06 6.50E+08 650.00 650.00kmol/hr kg/hr t/a $/a M$/a

Hydrogen 4567.75 9208.58 77352.07 208,850,590 208.85 208.85$/a M$/a

Electricity 6.72E+07 67.18Fee for disposal kg/hr m3/hr m3/a gal/a gal $/2e+5gal $/a M$/a $/ton ethyleneof water 2.78E+05 303.63 2550518.45 673775815 200000 387.29 1,304,733.18 1.30 1.30density 9.16E+02 kg/m31 m3 = 264.1721 gallon

M$/a income 924.72Net Cash Flow 251.06 NCF

$ M$Fixed Capital Cost 673,660,949.36 673.66Working Capital 101,049,142.40 101.05Total investment 774,710,091.76 774.71

XLII


0.1 Lifespan plantyears investments NCF Cum.NCF NFV DCFROR NPV

0 0 0.00 0 0.00 0.00 amounts in M$1 224.55 0 -224.55 -224.55 -204.14 -204.142 224.55 0 -224.55 -449.11 -185.58 -389.723 224.55 0 -224.55 -673.66 -168.71 -558.434 0 251.06 251.06 -422.60 171.48 -386.95 15 0 251.06 251.06 -171.54 155.89 -231.06 26 0 251.06 251.06 79.52 141.72 -89.35 37 0 251.06 251.06 330.59 128.83 39.49 48 0 251.06 251.06 581.65 117.12 156.61 59 0 251.06 251.06 832.71 106.47 263.09 610 0 251.06 251.06 1,083.77 96.80 359.88 711 0 251.06 251.06 1,334.83 88.00 447.88 812 0 251.06 251.06 1,585.90 80.00 527.87 913 101.05 251.06 150.01 1,735.91 43.45 571.33 10

XLIII


Appendix P

XLIV


Appendix Q

XLV


XLVI


Appendix R

XLVII

cpd nr 3334 conceptual process design process systems

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