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
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
3
Overall mass balance
14
3.1 Ethylene production 14
3.2 By-products 14
3.3 In- and Outgoing streams 15
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 416.3 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
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
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)
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.
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
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]
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
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
The 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.
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 90 Ethane conversion (%) 65 70 Yield Ethylene (%) 55 63 Energy 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
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
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 : Ethylene
Comp. Units Specification Additional Information Available Design Notes (also ref. note numbers) Ethylene wt% 99.9 99.9
By-Products wt% 0.1 0.1
Total 100.0
Process Conditions and Price
Temp. K 303
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
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
Component Total concentration Dimension CO / 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.
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 : Ethane
Comp. 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 for the process. Compounds not Total 100.0 included in mass balance.
Process Conditions and Price
Temp. K 298
Press. Bara 10 Phase V/L/S V Price $/tonne 150
Table 2-4: Water inlet stream
Stream Name : Water Inlet
Comp. Units Specification Additional Information Available Design Notes (also ref. note numbers) Water wt% 100 100.0 (1) Contaminants not harmful for Impurities ppm wt 80 80.0 (1) the process. Compounds not
included in mass balance.
Total 100.0
Process Conditions and Price
Temp. K 300
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.,
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 + H
2 C H
C H + CH
C H
C H + H
C H
C H + CH
C H
C H + CH
→
→
→
U
U
(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 2C H + C H
C H
2 C H
C H + 2 CH
C H + C H
C H + CH
C H + C H
C H + H
→
→
→
U
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
8Hidaka, Y. et al, “Shock-tube and modeling study of ethane pyrolysis and oxidation”, Comb. and
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. 2 7 2 t 6 6 2 t t 1 7 7 t 2 t 1 2 8 8 2 t 2 2 7 8 t 3 9 t 9 9 2 -9 2 t t
F F
P
r =k
F
TR
P
F
r =k
F TR
P
F F
r =k
F
TR
F F
P
F F
P
r =k
- k
F
TR
F
TR
2 t 2 3 t 1 1 1 -1 2 t t t 2 2 2 t 3 t 3 3 t t 4 4 4 t 2 5 t 7 4 t 5 5 -5 2 t tP
F F
P
F
r =k
-k
F TR
F
TR
P
F
r =k
F TR
F
P
r =k
F TR
P
F
r =k
F TR
F
P
F F
P
r =k
-k
F TR
F
TR
Tnumber stated below the flow rates (F) corresponds with the following component:
1 Ethane 6 Propane 2 Ethylene 7 Acetylene 3 Hydrogen 8 Butadiene 4 Methane 9 Benzene 5 Propylene R*T n r
k = 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 5
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.
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 2
C H
U
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%.
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 33.0 W ater 11.34 375.0 Ethane 0.51 16.9 Hydrogen 0.07 2.4 Benzene 0.0085 0.3 By-products 0.11 3.7 W ater 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.
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
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.
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
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
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 Dimension
Density 5.946 kg/m3
Diameter 1 0.52 m
Diameter 2 0.286 m
Flowrate 63.1 m3/s
Pressure drop* 2594 kPa
Velocity 1* 0.9 Mach
Velocity 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
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-1 κ 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.
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
=
= 1
P
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:
n
Knowing that the diameter of the mixing section is 0.98 m (D0) the number of nozzles can be calculated using:
0 n n
2 × π × D
N =
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 30.8 0.21 5.59 29.3 0.22 5.85 28.0 0.23 6.12 26.8 0.24 6.38 25.7 0.25 6.65 24.6 0.26 6.92 23.7 0.27 7.18 22.8 0.28 7.45 22.0 0.29 7.71 21.2 0.30 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 788 515
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
a
P =
-
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×T
a =
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
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 pd
C = C +C ×T + C ×T + C ×T
3 (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. i p i
F × C
C =
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
p
R= C - C
(Equation 9)Knowing all parameters the ratio κ can be calculated using: p
v
C
κ =
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 w
M = γ × M
,i
(Equation 11)
16 Sinnott, R.K. , "Coulson & Richardsons's Chemical Engineering Vol. 6," 3th ed.,
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.
shock
w
κ × R × T
u
=
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 0
Mach 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 × P
ρ =
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.
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:
2 0
1
A = × π × D
4
(Equation 15)From the molar flows, the concentration is acquired, which is then used in the rate expression for the specific reaction.
i i i
F
c =
×
F
R × T
∑
P
(Equation 16) a,j E R × 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:
i
R × T
u =
×
F
P × 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 ref
f,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
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. 2 4/5 - 2/3 3 c,i c,i -7 3 i 1/6 w,i c,i c,iZ
P
1.9 × T
η =
- 0.29
× 10 ×
× 10 × M ×
T
T
1.0134
` (Equation 21) 17Because an ideal gas mixture is assumed, the mean viscosity can be used:
i i
F × η
η =
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 × D
Re =
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 x
P = 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
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:
0
2 × π × α
D = 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 × A
dT
= 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: -0.2
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 700Temperature (K)710 720 730 740 750 760 % Conversion Ethane Selectivity
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 6 P re ssu re (B a r) C onversion Ethane S electivity
Figure 5-4: Influence of pressure
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:
Velocity
Mach 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 3.1 Mach Nr % Conversion Ethane SelectivityFigure 5-5: Influence of velocity
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
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.
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 by-production of almost 180 mol/s.
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
Table 5-5: Reactor outlet composition
Stream Name : Reactor outlet
Comp. Units Specification Additional Information Expected Design Notes (also ref. note numbers) Ethylene wt% 53 - 58 54.88 (1) For this mixture the price Ethane wt% 25-30 28.41 could not be calculated Methane wt% 4 - 5 4.72 Hydrogen wt% 4 - 4.5 4.16 Propane wt% 3 - 4 3.40 Butadiene wt% 0-0.5 3.09 Propylene wt% 0.3 - 0.8 0.48 Acetylene wt% 0 - 1 0.42 Benzene wt% 0-0.5 0.42 Total 100.0
Process Conditions and Price
Temp. K 1273 1290 Press. Bara 10 9.8 Phase V/L/S V Price $/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 0.32 Mixing start 0.286 0.98 2.83 Nozzleblock 0.98 0.98 0.25 Mixing final 0.98 0.98 6.40 Pyrolysis 0.98 2.07 6.17 Total 15.97 Starting Diameter Final Diameter Length Section
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:
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
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 Destination Methane CH4 176.84 2.84 Fuel Acetylene C2H2 9.75 0.25 Converted Ethylene C2H4 1174.94 32.96 Sold Ethane C2H6 567.34 17.06 Recycled Propylene C3H6 6.90 0.29 Fuel Propane C3H8 46.37 2.05 Fuel Butadiene C4H6 34.32 1.86 Fuel Hydrogen H2 1239.62 2.50 Sold Benzene C6H6 3.25 0.25 Discharged W ater H2O 20734.80 373.54 Discharged/ Recycled Carbonmonoxide* CO 12.80 0.36 Upgrading Carbondioxide* CO2 12.00 0.53 Upgrading Hydrogensulfide* H2S 1.38*10-3 4.45*10-5Upgrading Total 24018.93 434.49 * = Sour gases
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.
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-PLOCK Aspen model: Flash2 Outlet Temperature: 298 K Feed: 449 kg/s Column pressure: 2 bar Heat duty: -33.11 MW
Recovery Top (%) Top stream fraction (molfrac)
Ethylene >99.9 0.33 - 0.37
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- Robinson Net Heat duty: 754.5 MW Column pressure: 1 bar
Top T 373.25 K Bottom T 374.65 K
Water 0.05 – 0.1 0.96 - 0.97 Benzene >99.99 0.013 - 0.014
Recovery Top (%) Top stream fraction (molfrac)
6.3.3 H
2S, CO and CO
2removal
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.
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: Amines Aspen model: Radfrac Temperature: 315 K Amine stream (mass fraction): 0.85 water & 0.15 MEA
H2S <0.01 -CO2 <0.01 -CO - 1.01E-07 Ethylene >99.99 0.35 - 0.36 Water 0.1 - 0.2 0.039 - 0.040 Monoethanolam ine <<0.01 0.00005 - 0.00006 Byproducts 0.60 - 0.65
Top stream fraction (molfrac) Recovery Top (%)
6.3.4 Dryer
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
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 Robinson Aspen model: Radfrac
Top T 134.35 K Bottom T 251.15 K Net Heat duty: 15.55 MW
Methane 99.95 – 99.99 0.12 - 0.14 Ethylene 0.40 – 0.10 0.0037 - 0.0039 Hydrogen >99.99 0.87 - 0.90
Byproducts 0.00002 - 0.00003 Recovery Top (%) Top stream fraction
(molfrac)
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.
‘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 Slow
Relative permeation rates
Figure 6-1: Schematic representation of membrane unit
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 Robinson Aspen model: Radfrac Column pressure: 8 bar
Top T 214.85 K Bottom T 235.55 K Net Heat Duty: 58.90 MW
Ethylene 99.50 – 99.90 0.9990487 Byproducts 0.000931
Recovery Top (%) Top stream fraction (molfrac)
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 Robinson Aspen model: Radfrac Column pressure: 5 bar
Top T 219.85 K Bottom T 287.35 K Net Heat Duty: 34.34 MW
Ethane >99.99 0.94 - 0.95 Byproducts 0.050 - 0.060
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 att
Reac out 1 x 1.2 1128 873 255 306 After expander 2 x 1.05 648 298 350 367.5 W ater 3 x 0.78 358 1293 -935 -729.3 Ethane 4 x 0.09 298 788 -490 -44.1 Ethylene 5 x 0.048 231 303 -72 -3.5 -103.4 Total DTmin=10°C Stream Data
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
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 $106 million. The cost for each individual unit is found with the help of one of the three following references:
• Coulson & Richardson16 • DACE21,
• Peters and Timmerhaus22.
21 Dutch Association of Cost Engineers, “Prijzenboekje”, 22th ed., Elsevier, mei 2002
22 Peters, Max S., Timmerhaus, Klaus D., “Plant design and economics for chemical engineers”,
8.2 Cost estimation for raw materials
Using the mass balances of the process, the required raw materials are calculated. For the stated production of 1 million ton ethylene per annum the process needs the following amounts of ethane and water.
Table 8-1: Raw materials for ethylene
Component Amount [Mt/a] Cost [$/t] Cost [M$] Ethane 1.25 150 188.58 Water 2.47 0.675 1.67
Total 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
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 cost
m3/a $/m3 M$/a Cooling in Process cooling water 8.34E+07 0.13 10.56
Heat in Process heat MW $/MMBTU M$/a 852.72 5.40 125.38 Utility Monoethanolamine (MEA) 1088 267.86 0.29 Water 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
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 105.77 Equipment erection 42.31 Piping 74.04 Instrumentation 21.15 Electrical 10.58 Buildings, process 15.86 Utilities 52.88 Storages 15.86 Site development 5.29 Ancillary buildings 15.86 Sub-total physical plant costs 359.60 Design and Engineering 107.88 Contractor's fee 17.98 Contigency 35.96
Total fixed capital cost 521.42
Table 8-5: Annual production costs M$ Variable costs 364.52 Fixed costs 153.68 Direct production costs 518.20 Extra costs 155.46 Annual production costs 673.66
