The document describes a styrene production plant that uses a two-reactor process to produce styrene via the dehydrogenation of ethylbenzene. A series of distillation columns are then used to separate and purify the styrene, benzene, and toluene products. An economic analysis found that the total capital investment was $27.8 million and the plant would need to sell styrene for $1.003/lb to achieve a 15% rate of return, which is not competitive with larger styrene facilities. Further increasing the scale of the plant by 10 times could potentially make it more competitive.

1. Sty-Guys
ChE 0613
Group 17
Authors:
Benjamin Longstreet
Chris Merz
Michael Thompson
Ari Ufondu

2. Chris Merz Benjamin Longstreet Michael Thompson Ari Ufondu
Table of Contents X X X X
Executive Summary X X X
Process Background X X X X
General Plant Description and
Assumptions
X X X X
Detailed Process Description X X X X
Reactor Design X X X X
Column Sequencing and Column
Design
X X X X
Other Unit Operations X X X X
Pinch and Exergy Analysis X X X X
PDF and Stream Table of the Plant
with the Traditional Heat Exchange
Network
X X X X
Process Equipment List X X X X
Process Economics X X X X
Conclusions X X X X
Nomenclature X X X X
References X X X X
Appendix A X
Appendix B X
Appendix C X
Appendix D X

3. Table of Contents
Executive Summary.....................................................................................................................1
1. Process Background..................................................................................................................3
2. General Plant Description and Assumptions ...........................................................................10
3. Detailed Process Description...................................................................................................11
Material Balance .....................................................................................................................11
Detailed Description................................................................................................................13
4. Reactor Design........................................................................................................................23
5. Column Sequencing.................................................................................................................34
6. Other Unit Operations.............................................................................................................43
Heat Transfer Equipment........................................................................................................43
Separation and Storage Tanks.................................................................................................50
Pumps......................................................................................................................................56
7. Pinch and Exergy Analysis ......................................................................................................69
Pinch Analysis .........................................................................................................................69
“Traditional” Plant and Tabular HX Network........................................................................72
Exergy Analysis.......................................................................................................................84
8. PDF and Stream Table of the plant with Traditional HX Network .........................................88
9. Process Equipment List......................................................................................................... 110
10. Process Economics............................................................................................................... 112
Conclusion: Design and Simulation of the Plant................................................................................ 115
Nomenclature.................................................................................................................................118
References .....................................................................................................................................119
Appendix A: Membrane Reactors....................................................................................................120
Appendix B: Primary Column Redesign For 99.9% .......................................................................... 122
Appendix C: Separating heavies (naphthalene/ethyl naphthalene) from high purity styrene in a single
column........................................................................................................................................... 125
Appendix D: Water Remediation .....................................................................................................127

4. 1
Executive Summary
The styrene plant produces styrene from the dehydrogenation of ethyl-benzene in a series of high
temperature adiabatic packed bed reactors using a potassium promoted iron oxide catalyst. The gaseous reactor
products are condensed and separated to yield high purity styrene for commercial use.
The styrene plant is in production for 300 days a year, 24 hours a day. The plant produces 215 million
pounds of styrene/yr and 90% of the total styrene is considered high purity and is 99.8% pure, while the remaining
10% of styrene produced is considered low purity and is 98% pure. Other products sold for profit from the plant
include benzene and toluene.
In the plant simulation a Gibbs reactor is used in Aspen to model the adiabatic packed bed reactor operating
conditions. The stream going into the first reactor must be 1175oF; this temperature is reached by bringing in ethyl-
benzene at 965oF and superheated steam at 1425oF. The ratio of steamto EB for the first reactor is about 6:1 to reach
the correct reactor temperature. In addition to bringing up the temperature of the reactant, steamacts as a diluent that
helps increase the conversion of EB to styrene. The second reactor is fed from a mixture of the effluent stream of
reactor one and more superheated steam, with a ratio of 8:1 steam to EB, to bring the operating temperature back up
to 1125oF. The reaction is endothermic so the products lose heat in reactor one and must be reheated by the
superheated steam to reach the reactor temperature. The steam and ethyl-benzene feeds have an inlet pressure of 26
psi to meet the operating conditions of both reactors 1 and 2. Potassium promoted iron oxide catalyst is used in both
reactors to bring the reactors up to the correct conversion values. Reactor one holds 28% of the total catalyst and the
conversion of EB to styrene is 21%. The remaining catalyst is in the second reactor, and the conversion of the
remaining EB to styrene is 33%.
A traditional plant employs the simplest heat exchanger network which simply integrates the largest
temperature hot streams with the largest temperature cold streams. A traditional heat exchange network was used in
the styrene plant simulation, hot gas was used to heat the water feed stream through the water heat exchanger
(WHX) and cold utility water was used to cool the product stream through the three phase heat exchanger (3PHX).
The target heating utility is the amount of heat required to bring the water feed into the plant up to 1425oF through
the WHX heat exchanger. The heating duty was reached by manipulating the flow rate of methane to be mixed with
the hot flue gas to reach the correct steam temperature. The cooling utility was reached by manipulating the flow
rate of cold utility water at 70oF through the 3PHX heat exchanger until the correct outlet temperature of 85oF was
achieved. A pinch analysis was calculated on the plant to determine what the actual minimum heating and cooling
utilities were, and to compare these values to the traditional heat exchange network. When the two networks are
compared, the pinched network was more efficient than the traditional plant.
A system of five distillation columns is used to separate the products and bring the styrene, benzene, and
toluene to their appropriate purity levels. Three of the distillation columns are vacuum pressurized to 30 torr
preventing styrene from polymerizing. The vacuum columns are more expensive to run than the other columns and
therefore took more time during the separation process. Styrene and ethyl-benzene were separated first within the
column sequencing, ultimately reducing the number of vacuumcolumns needed reducing the costs of plants.

5. 2
There are some health and safety concerns for the chemicals produced in the styrene plant. Some of the
chemicals can be dangerous to one’s health if exposed to them. Benzene, styrene and toluene are the most hazardous
chemicals at the facility as they all have low flash points and storing a large quantity can raise some concern. These
chemicals also have the hazardous lower explosive limits (LEL). Gas detectors and proper ventilation should be
used within in the storage tanks to monitor the levels of the air to benzene, toluene, and styrene ratio to ensure that
an explosion can be prevented or contained.
An economic analysis of the styrene plant was determined using Aspen Process Economic Analyzer
software. The total capital investment for the facility is about $27.8 million. To get a rate of return of 15%, the price
of styrene would need to be $1.003/lb. The net present value of the plant after 15 years of operation compared to an
interest rate of 10% is $23.9 million. This plant is not competitive to other styrene facilities when trying to achieve a
15% rate of return and a lower rate of return is not desirable for a potential investor.
Further study of this particular styrene plant would be unnecessary as the plant is not currently attractive to
a typical investor. The competitive styrene producing facilities are much larger, and can sell the styrene at a higher
profit because there is only a small increase in overhead equipment costs compared to the production rates. Scaling
the styrene plant up by about ten times its current size would give the plant a competitive edge. Some of the
concerns with taking a model and scaling it up are that some of the factors do not scale linearly such as reactor
kinetics, pressure changes, and thermodynamics. Health and safety risks also increase when larger quantities of
potentially dangerous chemicals are stored, an explosion or leak would be considerably more catastrophic from a
larger plant.

6. 3
1. Process Background
Styrene (chemical formula C6H5CH=CH2) is an organic compound that is derived from ethylbenzene in the
presence of a catalyst and heat. Styrene is a colorless unsaturated liquid hydrocarbon that is stable and easily
handled When exposed to air quickly evaporates often causing a sweet aroma to the surrounding area, however, at
high volumes or when mixed with other chemicals, styrene is often described as producing an unpleasant odor1.
Styrene, also known as phenylethylene, vinylbenzene, and styrol and is an important industrial unsaturated aromatic
monomer. The development process for the commercial manufacturing of styrene is based on two different methods.
One is a method known as SM/PO (styrene monomer/propylene oxide). This process occurs when ethyl-benzene is
treated with oxygen to form ethyl-benzene hydro-peroxide. The hydro-peroxide is used to oxidize propylene into
propylene oxide, and the product 2-phenylethanol is dehydrated to form styrene. The method investigated in this
report is the dehydrogenation of ethylbenzene which was achieved in the 1930’s and was widely expanded during
World War II as the need for cheap plastics and rubbers increased greatly1. This market for a high purity monomer
that could be stabilized into cheap materials expanded rapidly after the war’s end in 1946 and is still used in high
volumes today as the least expensive thermoplastic. Styrene is a miscible compound when mixed with most organic
solvents, and is also a good solvent for synthetic rubber, polystyrene, and other non-cross linked high polymers1.
The most important engineering properties of styrene, listed in Table 1, are used to determine whether or not styrene
is a capable and stable fit for polymerization with other monomers and functional groups as well as the health and
safety of the plant.
# Property Value
1 Molecular weight 104.153 g/mol
2 Boiling point 145.15° C
3 Freezing point -30.6° C
4 Critical density 0.297 g/mL
5 Critical pressure 3.83 MPa
6 Critical temperature 362.1° C
7 Critical volume 3.37mL/g
8 Flammable limits in air 1.1 – 6.1 vol%
9 Flash point 31.1° C
10 Auto-ignition point 490° C
11 Heat of combustion at constant P -4.263 MJ/mol
12 Density 0.9050 g/mL
13 Heat of Formation ∆HF Gas (25°C) 147.4 kJ/mol

7. 4
Heat of Formation ∆HF Liquid(145°C) 103.4 kJ/mol
14
Heat of Vaporation ∆HV Gas (25°C) 421.7 J/g
Heat of Vaporation ∆HV Liquid (145°C) 356.7 J/g
15 Volume Shrinkage on Polymerization 17.0%
16 Refractive Index, nD (20°C) 1.54682
17 Heat of Polymerization ∆Hp (25°C) -69.8 kJ/mol
18 Viscosity (20°C) 0.762 centipoise
19 Specific Heat (20°C) 1.73 kJ/kg°C
20 Thermal Conductivity 0.16 W/m°C
21 Solubility in Water <1%
22 NFPA 704
Table 1.1: 20 Most Important Engineering Properties of Styrene
Styrene is easily and widely used in the creation of plastic, rubber, and resin components because of the high activity
of the vinyl group. These components are made when styrene monomers are linked together to form an aromatic
polymer, polystyrene1 (chemical formula (C8H8)n). This polymerization reaction is used to form many different
products listed in table 2.
Acrylonitrile butadiene styrene (ABS) 1 Appliances, automotive parts, pipe, business machines,
telephone components 2
Styrene/butadiene co-polymer (SBR) 1 Automobile tires, latex adhesives, wire insulation,
footwear 4
Polystyrene (PS) 7 Packaging, domestic appliances, consumer electronics,
construction insulation, medical tissue culture trays 6
Styrene-acrylonitrile (SAN) 1 Cosmetics, household containers and mixing bowls,
thermally insulated jugs 5
Expanded polystyrene (EPS) 7 Household trays and plates, molded sheets for building
insulation, packing peanuts 6
Table 1.2: Specific Uses of Styrene

8. 5
Two styrene products that society utilizes daily are packaging material and the rubber based component in
automobile tires. Polystyrene foams are good thermoplastic insulators and are widely used in the food industry when
perishable foods need transported between facilities. Packaging material is an example of a polystyrene foam. In
atactic polystyrene production2, which is the commercially produced polymer, the phenyl groups are randomly
positioned on all sides of the polymer chain and therefore are inhibiting the polystyrene to crystalize giving the
compound the ability to be molded into different materials as seen below in figure 15.
Figure 1.1: Polymerization of Styrene
Automobile tires are an example of styrene-butadiene rubber or SBR as seen in figure 2. The butadiene component,
1,3-butadiene, provides the rubber composition due to the unsaturated carbon-carbon bond in the polymer while the
styrene provides the molecular stability of the tire4. SBR has replaced natural rubber in the manufacturing of
automobile tires based on its cheap cost and the strength that it provides to the tire4.
Figure 1.2: Styrene-Butadiene Rubber
The majority of styrene is made from the dehydrogenation of ethylbenzene. Dehydrogenation is the removal of
hydrogen molecule(s) and reallocation of electrons in the form of stronger bonds. In the formation of styrene, a
single hydrogen molecule is removed from ethylbenzene and a double bond is formed. The major reversible reaction
involved in the production of styrene by dehydrogenation is seen in figure 3.

9. 6
Figure 1.3: Formation of Styrene by Dehydrogenation
This highly endothermic reaction is carried out over a catalyst bed, typically iron-oxide. Potassium
carbonate/hydroxide or chromium-oxides can also be added to the catalyst to increase the reaction selectivity. The
most dominant composition of the catalyst used in the market today is 84.3% iron oxide, 2.4% chromium-oxide and
13.3% potassium carbonate1. During this reaction there are several side reactions that also take place. There are
competing thermal reactions that degrade ethylbenzene to benzene, ethylene and carbon. Carbon is a poison to the
catalyst which is why potassium is a large component of the catalyst as it assists in converting the carbon into
carbon dioxide and removed as a vented gas. Ethylbenzene can also have a side reaction with hydrogen to produce
toluene and methane. It is important to note that this reaction operates under low pressure and high temperature
conditions. The low pressure is due to Le Chatelier’s principle. As the reaction is carried forward, there will be more
products than reactants. To counter act the reverse reactions, low pressure is required. To pressure cannot be taken
too low because the reaction will not occur at all. The high temperature is required because the reaction is
endothermic and thus requires thermal energy to be carried out. However, if the temperature is too high, the
reactants could burn and the reaction would not occur. Also, too high of temperatures could cost the facility money.
In industry there are two dominant process methods in which styrene can be mass produced; (1) isothermal
dehydrogenation and (2) adiabatic dehydrogenation of ethylbenzene. More than 75% of all styrene plants utilize
adiabatic dehydrogenation of ethylbenzene to produce styrene which can be seen in figure 4. The industrial process
uses two adiabatic reactors in order to achieve maximum yield of pure styrene. The process begins with a fresh
stream of ethylbenzene being mixed with a recycled stream of unreacted ethylbenzene and fed to a heat exchanger.
The ethylbenzene is then converted into a vapor and heated to between 600°C-640°C. The ethylbenzene vapor is
mixed with steam and fed to the first reactor. In figure 4, the steam is mixed at the beginning of the process with the
ethylbenzene and recycled ethylbenzene. The placement of the steamaddition is only necessary prior to the reactor,
thus the exact location is arbitrary—the steam is used to prevent the ethylbenzene from forming coke and heating
the ethylbenzene stream to reactor temperatures. The ratio of steam to ethylbenzene optimizes the yield of styrene
while minimizing utility costs. These ratios can vary depending on the size of the process and input parameters.
Reactor one usually has approximately a 35% conversion of ethylbenzene to styrene. Since the reactors are
adiabatic, the product stream from reactor one will need to be reheated to reactor temperatures due to the heat lost
from the endothermic reaction and is achieved using an additional heat exchanger. It is important to note that in
some cases, the steam can be utilized to reheat the product stream from reactor one to reduce utilities. Steam is

10. 7
mixed with this product stream and the mixture is fed to reactor two resulting in around a 65% con version of the
ethylbenzene to styrene. Each of the reactors used in the adiabatic plant operate at near vacuumconditions where the
pressures can vary depending on the total conversion desired and consequently reactor temperatures. The product
stream is then condensed and separated into vent gases, crude styrene, and ethylbenzene by means of distillation
columns2.
Figure 1.4: Adaibatic Dehydrogenation of Ethylbenzene
Another major process for dehydrogenating ethylbenzene is by means of isothermal reactors. The reactor is built
similarly to a shell and tube heat exchanger with a catalyst bed but unlike the adiabatic process, the isothermal
process requires only one reactor undervacuum pressure conditions as seen in figure 5. The beginning of the process
is similar to that of the adiabatic process—fresh ethyl benzene is mixed with recycled ethylbenzene and steamwhere
the ratio of steam to ethylbenzene is between 0.6-0.9 by weight. The mixture is heated up to the reaction temperature
of approximately 600°C and then fed to the isothermal reactor. The heat of the endothermic reaction is supplied by a
hot flue gas on the shell side of the reactor. An example of the heating medium used for the reactor is molten
sodium, lithium, and potassium carbonates3. The product stream from the reactor is then separated using distillation
columns similarly used in the adiabatic process. The running costs of isothermal plants compared to adiabatic plants
are generally cheaper because they require less heating utilities while producing a higher conversion of styrene from
ethylbenzene. However, the major disadvantage to using an isothermal reactor instead of an adiabatic reactor is the
practical size limitation of the reactor and the expense of constructing multi-tubular reactors. Larger plants using
isothermal reactors will have to invest more capital into the styrene plant2.

11. 8
Figure 1.5: Isothermal Dehydrogenation of Ethylbenzene
Global styrene production, and subsequent consumption, is driven by polystyrene, expandable polystyrene (EPS),
acrylonitrile butadiene styrene (ABS), and styrene-acrylonitrile (SAN) resins demand (mainly used within the
automobile industry). To provide feedstock for these products, global styrene production capacity has had to
increase from 52.8 billion pounds in 2012 to an estimated 59.4 billion pounds in 20158. Asia controls a majority of
the market, holding about half of the world’s styrene capacity and production. Europe, North America, The Middle
East, and Latin America account for the other half of the world’s styrene capacity as seen in figure 68. When
examined on a country-by-country basis, China, Japan, and the United States are the world’s largest styrene
producers, responsible for about 44% of global production, while Taiwan and South Korea follow closely behind
with about 16% of the world’s styrene production as seen in figure 78.
Figure 1.6: Global Styrene Capacity by Region

12. 9
Figure 1.7: Global Styrene Capacity by Country
Due to its large end user sector, global demand for styrene has recorded stable growth in recent years, with Asia
having the largest growth in demand. This has been attributed to its huge emerging middle class and its economies
growing steadily. Interestingly, Asia is the world’s largest importer and exporter of styrene9. The production
capacity of styrene plants worldwide are highly variable typically producing anywhere from 350 million pounds per
year to approximately 2.5 billion pounds per year9. Such high capacities are needed to keep up with the ever surging
demand for styrene and its applications. Large publicly owned petrochemical companies typically own and operate
the larger styrene plants, while smaller companies make do with the smaller capacity plants 9. Figure 8, seen below,
depicts several companies in which styrene is produced. The figure displays the producer, the location and the
amount of styrene produced on an annual basis.
Producer Plant location Output (millions of pounds/year)
Amoco Texas City, TX 840
Arco Channelview, TX 2,525
Chevron St. James, LA 1,500
Cos-Mar Carville, LA 1,900
Dow Freeport, TX 1,500
Huntsman Chemical Bayport, TX 1,250
Sterling Chemicals Texas City, TX 1,600
West Lake Charles, LA 350
Rexena Odessa, TX 320
Figure 1.8: Styrene individual plant production rates in 19939

13. 10
2. General Plant Description and Assumptions
The main purpose of a styrene plant is to produce large quantities of high purity styrene. The process
needed to produce a high purity styrene involves the catalytic dehydrogenation of ethyl benzene at high
temperatures. The traditional plant minimizes the amount of heating and cooling utilities required to maintain
reaction temperatures for each stream by use of heat exchangers. This designed styrene plant is intended to operate
for 300 days, 24 hours a day, producing an estimated total of 215 million pounds of styrene.
A mixture of fresh and recycled liquid ethylbenzene is heated by superheated steam and is fed to an
Adiabatic Plug Flow Reactor (PFR). This steam is created by using a heat exchanger (WHX) to superheat the water
inlet feed (WATFEED) which sequentially heats the ethyl benzene feed stream (EBFEED) with the ethyl benzene
heat exchanger (EBHX) to its corresponding reaction temperature. It is important that the heated ethyl benzene
stream reaches the required reaction temperature of 1175oF from its entering temperature of 965 oF. This ethyl
benzene stream is heated by superheated steam measured at 1425oF. After exiting reactor 1, 21% of the ethyl
benzene entering reactor 1 was converted to styrene. To reach a higher yield of styrene production, a mixture of
styrene and the remaining ethylbenzene (R1PROD) blends with more superheated steam and passes through a
second adiabatic PFR (R2) converting 33% of the remaining ethyl benzene. The superheated steamused throughout
the styrene plant uses a high amount of utilities in order to continuously heat the necessary streams.To save costs on
heating utilities and lower the energy needed, each of the steam sources are split (WATSPLIT) fromthe same water
stream. Highly pure styrene is formed when the procedures are followed and the reaction temperatures are reached,
but this styrene is not the only product made from both reactors. Carbon dioxide (CO2), Hydrogen (H2), and other
organic compounds (benzene, ethylene, ethyl-naphthalene, methane, naphthalene, and toluene) are created due to
side reactions involving ethyl benzene and steam.
The products generated from R1 & R2, now in the gaseous phase, are then fed to a condenser, followed by
a 3 Phase Separator to remove oil and water from this production mixture. In this unit, the oil phase, water phase,
and vapor phase components are separated into three separate streams. The oil and water phases are separated by
their specific gravities. Water is sent to a stripping column for purification and the oil is sent to a system of
distillation columns. The gaseous products from the condenser comprising the vapor phase, are pulled off fromthe
top of the separation vessel and sent to the FIRE unit. The size of the separator is determined by its need to keep
methane, carbon dioxide, and hydrogen gas in the vapor phase. Utilizing a large enough separator will allow these
gases to not condense and mixwith the other two liquid phases.
The oil mixture from the separator is then sent through a series of distillation columns. All of the organic
products are separated by the distillation columns into their proper storage containers except for ethhylbenzene
which is recycled back to the ethylbenzene feed. Monomers have a tendency to polymerize in the re-boiler portion
of the columns thus fouling and causing damage to the distillation column. To reduce the rate of polymerization,
vacuum columns will be utilized when the feed stream contains any styrene. Columns PRIMARY, STHCOL, and
STLCOL, which are labeled in Figure 6, would encounter problems with fouling if they were not vacuumed.

14. 11
Compared to the other columns in this distillation system the initial column in the sequence, PRIMARY, is larger
than the others since it is the column that is fed all of the organic compounds at a high volumetric flow rate.
Flow rates were obtained for each of the products streams of the distillation columns. A flow rate of
benzene at 17.83 lbmol/hr, toluene at 30.59 lbmol/hr, high purity styrene at 254.34 lbmol/hr, low purity styrene at
28.22 lbmol/hr, naphthalene and ethyl-naphthalene both at 4.77 lbmol/hr were obtained. The difference between
high purity and low purity styrene is that high purity styrene is extracted fromnaphthalene and ethyl-naphthalene in
the STHCOL while low purity needs to be extracted by an additional distillation column, STLCOL. Ethylbenzene
collected from the EBRECOL is recycled and fed back to the EB feed at the beginning of the process.
3. Detailed Process Description
Material Balance
In order to determine the amount of ethylbenzene, recycled ethylbenzene and water needed to achieve a
production rate of 286.74 lbmol/hr of styrene within an operational plant, a material balance was completed to
govern the necessary inputs required in order to minimize running costs and maximize profits. All of the material
balance calculations were performed using the dehydrogenation of ethylbenzene reaction.
The styrene plant consists of two reactors connected in series. Reactor one is assumed to have a 21%
conversion of ethylbenzene to styrene, while reactor two is assumed to have a 33% conversion of ethylbenzene to
styrene. In reactor two, several side reactions take place which also consume 12.5% of the ethylbenzene feed. Each
reactor can form a carbon deposit, coke, on the catalyst bed which can potentially make the catalysts beds inert. To
avoid this from happening, a large quantity of steam is supplied to the reactors. The steam not only reacts with the
carbon deposits to protect the catalyst beds, but also cuts down on utility expenses. The heat energy supplied by the
steam is used to bring the two reactor feeds to optimal reaction temperatures. Based on a 21% and 33% conversion
of the two reactors, the production of styrene from ethylbenzene is given by equation 1:
𝐸𝐵1 ∗ 0.21 + 𝐸𝐵2 ∗ 0.33 = 286.74
𝑙𝑏𝑚𝑜𝑙 𝑆𝑡𝑦𝑟𝑒𝑛𝑒
ℎ𝑟
Eqn 3.1
Where EB1 and EB2 are the flowrates of ethylbenzene feeds (lbmol/hr) going to reactors one and two respectively.
𝐸𝐵1 ∗ 0.21 + (𝐸𝐵1 − ( 𝐸 𝐵1 ∗ 0.21)) ∗ 0.33 = 286.74
𝑙𝑏𝑚𝑜𝑙 𝑆𝑡𝑦𝑟𝑒𝑛𝑒
ℎ𝑟
Eqn 3.2.
𝐸𝐵1 = 609.18
𝑙𝑏𝑚𝑜𝑙 𝐸𝑡ℎ𝑦𝑙 𝐵𝑒𝑛𝑧𝑒𝑛𝑒
ℎ𝑟
Reactor one, as seen in figure 1, describes the material balance and its respective components in the formation of
styrene. Selectivity of the dehydrogenation of ethylbenzene producing styrene is given by equation 3:
Eqn 3.3.
𝑆𝑒𝑙𝑒𝑐𝑡𝑖𝑣𝑖𝑡𝑦 =
𝐶𝑜𝑛𝑣𝑒𝑟𝑠𝑖𝑜𝑛 𝑜𝑓 𝐸𝑡ℎ𝑦𝑙𝑏𝑒𝑛𝑧𝑒𝑛𝑒 𝑡𝑜 𝑆𝑡𝑦𝑟𝑒𝑛𝑒
𝑇𝑜𝑡𝑎𝑙 𝐶𝑜𝑛𝑣𝑒𝑟𝑠𝑖𝑜𝑛
= 0.826
Based on the selectivity of ethylbenzene, 609.18 lbmol/hr is required to produce 286.74 lbmol/hr. Prior to entering
reactor one, the ethylbenzene stream is mixed in a 6:1 ratio of steamto ethylbenzene. The 3655.08 lbmol/hr

15. 12
of steam and 609.18 lbmol/hr of ethylbenzene are then directed to reactor one. From here, styrene is then formed
consuming 21% of the entering ethylbenzene resulting in a 3.0% exiting stream composition of styrene product at
127.92 lbmol/hr. The unconverted ethylbenzene and the inert cooled steam make up the rest of the exiting stream
composition at 11.29% (481.25 lbmol/hr) and 85.71% (3655.08 lbmol/hr).
In addition to the product stream from reactor one, fresh steam is required for reactor two, as seen in figure 2. This
fresh steam makes up a 47.45% composition at an 8:1 ratio of steam to ethylbenzene at a flowrate of 3850 lbmol/hr.
The fresh steam will be applied prior to entering reactor two, to heat up the feed stream to optimal reactor
temperatures for the dehydrogenation of ethylbenzene. Reactor two will convert 33% of the incoming unreacted
ethylbenzene into styrene, producing a final styrene product of 286.74 lbmol/hr. Ethylbenzene and styrene are also
subjected to side reactions which is expected to consume 12.5% of the ethylbenzene entering reactor two accounting
for 60.15 lbmol/hr of the product stream. There is 262.28 lbmol/hr of unreacted ethylbenzene remaining that will be
used as the recycle stream for reactor one cutting costs by consuming less fresh ethylbenzene. The inert steamthat is
exiting accounts for the overall amount of steam that was used throughout the process which totals 7505.08
lbmol/hr.

16. 13
In conclusion,in order for this process to be successfulin forming the final styrene product,inputs of
346.89 lbmol/hr of fresh ethylbenzene, 262.28 lbmol/hr of recycled ethylbenzene and 7505.07 lbmol/hr of steamare
required. The water stream, after being converted into steam using a heat exchanger, is split where 48.7% of the
composition going to reactor one and the remainder (51.3%) going to reactor two. There is more steambeing
charged into reactor two because the bulk mass of reactor two’s feed is greater and thus requires more thermal
energy. Ultimately, these inputs will produce a styrene product of 286.74 lbmol/hr. It is important to note that if the
styrene plant was not running, the inputs of reactor one would change significantly at start-up. The recycle streamof
ethylbenzene entering reactor one would be unused,resulting in a larger quantity of fresh ethylbenzene being
required.
Detailed Description
The styrene plant being described in this report presumes a traditional heat exchanger network. The amount
of styrene this plant will produce annually is 215 million pounds. Utilizing this information, a material balance can
be used to determine the inputs of the facility, namely the liquid ethyl benzene and water feeds. The feeds for the
fired heater, methane and air, are also accounted for. These feeds are found in table 1:

17. 14
Fresh Feed Streams
Molar Flow Rate
(lbmol/hr)
Temperature
(°F)
Pressure
(psia)
Vapor
Fraction
(% )
Composition
Pure Ethyl Benzene
(EBFEED)
346.01 70 14.7
-- 1
Pure Water
(WATFEED)
6540.02 70 14.7
-- 1
Pure Methane Feed
(CH4FEED)
445 70 14.7
-- 1
Air Feed
(AIRFEED)
7272 70 14.7
-- YN2 = 0.79
YO2 = 0.21
Table 3.1. Fresh Feeds
In order to achieve a dehydrogenation reaction of ethyl benzene 346.01 lbmol/hr of ethyl benzene and
6540.02 lbmol/hr of water is required. The temperature and pressures of each stream are at ambient conditions, i.e.
they are not specified but drawn from a source tank kept at atmospheric conditions. Since, the dehydrogenation
reaction does not fully consume the ethyl benzene, a recycle streamof ethyl benzene is added to the fresh streamof
ethyl benzene. The recycled stream was estimated to contain 262.28 lbmol/hr of ethyl benzene. Table 2 depicts the
actual recycled stream (EBRECYCLE):
Substream
Molar Flow Rate
(lbmol/hr)
Fraction
Vapor Fraction
(% )
Ethyl Benzene 257.463 0.8912616 0
Styrene 31.3814 0.1086331 --
Benzene 3.6003E-10 1.2463E-12 --
Toluene 0.030391 0.000105205 --
Naphthalene 5.92163E-06 2.0499E-08 --
Ethyl Naphthalene 5.9231E-06 2.05041E-08 --
Totals 288.875 1 --
Table 3.2. Recycled Stream of Ethyl Benzene
The actual recycled amount of ethyl benzene within the styrene plant is 257.463 lbmol/hr. This stream is
sent to EBMIX and combined with the fresh ethyl benzene stream EBFEED. Once mixed, this combined stream,
EBTOTAL, is sent through EBPUMP to bring the mixture to the specified reactor pressure of 26 psia given in the
problem statement. Once the streams are pressurized, they are sent into heat exchangers in which they are heated to
reactor temperatures necessary for specified reactor conversions. Table 3 shows the specs for both the ethyl benzene
heat exchanger (EBHX) and the water heat exchanger (WHX).

18. 15
Specifications
Ethyl Benzene Heat Exchanger
(EBHX)
Water Heat Exchanger (WHX)
Hot Stream R2OUT HOTGAS
Cold Stream EBHIP WATHIP
Temperature (°F) 965 1425
Valid Phases Vapor-Liquid Vapor-Liquid
Table 3.3: Input Heat Exchanger Specifications
EBHX utilized the product stream of reactor two (R2OUT) as the heat source for the mixed feed of ethyl
benzene (EBHIP). This heat exchanger is only required to ensure the ethyl benzene is heated to 965°F. The water
heat exchanger (WHX) utilized the hot stack gases (HOTGAS), which are incinerated after being separated fromthe
reactor two product stream, as the hot stream and the water feed (WATHIP) as the cold stream. WHX was specified
at a temperature of 1425°F. This elevated temperature was specified since it will be utilized in future streams to
further heat the ethyl benzene stream to reactor specifications. The allowed phases for both cold and hot streams is
vapor-liquid since at the low pressures of the streams both liquid and vapor will be present. It is important to note
that this plant was designed traditionally. This means that the heat exchanger network was designed easily while
maintaining thermodynamic laws, i.e. the hot streams were matched with the cold streams ensuring the hot streams
were actually at a higher temperature than the cold streams.
After the ethyl benzene and water streams have been heated, the ethyl benzene stream(EBHIPT) is sent to
a mixer (R1MIX) where it is mixed with the superheated steam. Prior to being mixed with the higher pressure and
temperature ethyl benzene, the steam is sent to a splitter (WATSPLIT) where the water will be split at an inexact
specified molar ratio of 6:1. This ratio was utilized in order to achieve the desired reaction temperature of 1175°F in
both reactor one and two. Once the controllers, TEMPR1 and TEMPR2 were run the new splitting came to 0.576 or
3767.05 lbmol/hr of the WATHIPT stream to reactor one and 0.424 or 2772.97 lbmol/hr of the stream to reactor
two. The controllers utilized during this process will be discussed later in Table 8.
The reactors were specified at 26 psia and 1175°F to have explicit conversions of ethyl benzene to styrene
as well as side reactions. Each of the reactors used the Peng-Robinson package of thermodynamics to determine the
transport properties. Reactor one (R1) was assumed to contain no side reactions, thus the only reaction that occurs in
the dehydrogenation of ethyl benzene. The first reactor was specified to have a conversion of 21% of the ethyl
benzene to styrene. The second reactor (R2) contained all of the side reactions where it was specified that 33% of
the ethyl benzene converted to styrene, 4% to toluene and methane, 0.5% to benzene and ethylene, 2.5% to
toluene/hydrogen/carbon dioxide, 3.5% of benzene/hydrogen/carbon dioxide, 1% to naphthalene and ethylene, a nd
1% to ethyl-naphthalene and ethylene. The balanced reactions and stoichiometry are listed below.
Reactor1
[1] 𝐶8 𝐻10(𝑒𝑡ℎ𝑦𝑙 𝑏𝑒𝑛𝑧𝑒𝑛𝑒) = 𝐶8 𝐻8(𝑠𝑡𝑦𝑟𝑒𝑛𝑒) + 𝐻2(ℎ𝑦𝑑𝑟𝑜𝑔𝑒𝑛) 21% conversion

19. 16
As the product stream leaves reactor one (R1), the stream exits at a temperature of 1090.89°F which is not
at the reactors specified temperature of 1175°F. To achieve this temperature R1PROD is sent to R2MIX where it is
combined with the 2772.97lbmol/hr of steamfrom stream WATR2.
Reactor2
[1] 𝐶8 𝐻10(𝑒𝑡ℎ𝑦𝑙 𝑏𝑒𝑛𝑧𝑒𝑛𝑒) = 𝐶8 𝐻8(𝑠𝑡𝑦𝑟𝑒𝑛𝑒) + 𝐻2(ℎ𝑦𝑑𝑟𝑜𝑔𝑒𝑛) 33% conversion
[2] 𝐶8 𝐻10(𝑒𝑡ℎ𝑦𝑙 𝑏𝑒𝑛𝑧𝑒𝑛𝑒) + 𝐻2(ℎ𝑦𝑑𝑟𝑜𝑔𝑒𝑛) = 𝐶7 𝐻8(𝑡𝑜𝑙𝑢𝑒𝑛𝑒) + 𝐶𝐻4(𝑚𝑒𝑡ℎ𝑎𝑛𝑒) 4%
[3] 𝐶8 𝐻10(𝑒𝑡ℎ𝑦𝑙 𝑏𝑒𝑛𝑧𝑛𝑒) = 𝐶6 𝐻6(𝑏𝑒𝑛𝑧𝑒𝑛𝑒) + 𝐶2 𝐻4(𝑒𝑡ℎ𝑦𝑙𝑒𝑛𝑒) 0.5%
[4] 𝐶8 𝐻10(𝑒𝑡ℎ𝑦𝑙 𝑏𝑒𝑛𝑧𝑒𝑛𝑒) + 2𝐻2 𝑂(𝑤𝑎𝑡𝑒𝑟) = 𝐶7 𝐻8(𝑡𝑜𝑙𝑢𝑒𝑛𝑒) + 3𝐻2(ℎ𝑦𝑑𝑟𝑜𝑔𝑒𝑛) + 𝐶𝑂2(𝑐𝑎𝑟𝑏𝑜𝑛 𝑑𝑖𝑜𝑥𝑖𝑑𝑒) 2.5%
[5]𝐶8 𝐻10(𝑒𝑡ℎ𝑦𝑙𝑏𝑒𝑛𝑧𝑒𝑛𝑒) + 4𝐻2 𝑂(𝑤𝑎𝑡𝑒𝑟) = 𝐶6 𝐻6(𝑏𝑒𝑛𝑧𝑒𝑛𝑒) + 6𝐻2(ℎ𝑦𝑑𝑟𝑜𝑔𝑒𝑛) + 2𝐶𝑂2 (𝑐𝑎𝑟𝑏𝑜𝑛𝑑𝑖𝑜𝑥𝑖𝑑𝑒) 3.5%
[6] 2𝐶8 𝐻10(𝑒𝑡ℎ𝑦𝑙 𝑏𝑒𝑛𝑧𝑒𝑛𝑒) = 𝐶10 𝐻8(𝑛𝑎𝑝ℎ𝑡ℎ𝑎𝑙𝑒𝑛𝑒) + 3𝐶2 𝐻4(𝑒𝑡ℎ𝑦𝑙𝑒𝑛𝑒) 1%
[7] 2𝐶8 𝐻10(𝑒𝑡ℎ𝑦𝑙 𝑏𝑒𝑛𝑧𝑒𝑛𝑒) = 𝐶12 𝐻12(𝑒𝑡ℎ𝑦𝑙 𝑛𝑎𝑝ℎ𝑡ℎ𝑎𝑙𝑒𝑛𝑒) + 2𝐶2 𝐻4(𝑒𝑡ℎ𝑦𝑙𝑒𝑛𝑒) 1%
The heat exchanger EBHX heats the EBTOT stream utilizing the R2OUT stream. This results in a cooled
R2PROD (reactor 2 products). R2PROD must be cooled in order to properly separate within the distillation columns
downstream. For this reason, cooling water (CLDUTIN) is utilized within heat exchanger 3PHX to cool down this
stream to 85 oF before it is sent to the three phase separator(3PSEP). 3PHX is referred to as the three phase heat
exchanger because its product stream contains vapor,liquid and free water. The reactor product 3-Phase Heat
Exchanger specifications can be seen in table 4. Figure 3.3 also depicts a 3-dimensional view of the reactors (R1 and
R2), heat exchangers (EBHX and 3PHX), as well as the fired heater (FIRE) in the styrene plant.
Stream COLDUTIN
Cold Side CLDUTOUT
Hot Side 3PH
Mass Flow (lbmol-water/hr) 190,268
Temp (ºF) In 70
Valid Phases Vapor-Liquid
Table 3.4: 3 Phase Heat Exchanger Specifications

20. 17
Figure 3.3: Styrene Plant Reactors and Heat Exchangers
Once the products stream (3PH) exits the 3PHX, it is sent to the 3PSEP where this input streamis separated
into vapor, liquid, and oil. The vapor stream leaving 3PSEP, referred to as 3PVAP, leaves at 85oF which flows
thorough a valve (H2VALVE) lowering the pressure to from 26 psia to 14.7 psia. This stream passing through the
valve, H2LOWP, is sent to a mixer (FIREMIX) where it is combined with 7272 lbmol/hr of air (AIRFEED) and 445
lbmol/hr of methane (CH4FEED) operating at 70 oF and 14.7 psia. Two of the three feeds that are fed to the
FIREMIX utilize a feedback controller measuring temperatures upstream increasing and decreasing the flowrates as
necessary. The AIRFEED utilizes controller XSO2 and CH4FEED utilizes controller CH4TRAD. The FIREFEED is
the product stream from the FIREMIX flowing at 8202.36 lbmol/hr and enters the FIRE reactor at 70.91 oF and 14.7
psia. This reactor is utilized to heat the product stream from the FIREMIX to extreme temperatures (3222.14 oF) and
is utilized to heat the input water for the WHX heat exchanger to a reaction temperature of 1425oF. The reactions
occurring within the Fire Heater Reactor is seen below:
Fire Heater Reactor
[1] 2𝐻2
(ℎ𝑦𝑑𝑟𝑜𝑔𝑒𝑛) + 𝑂2 (𝑜𝑥𝑦𝑔𝑒𝑛) = 𝐻2 𝑂(𝑤𝑎𝑡𝑒𝑟)
[2] 𝐶2 𝐻4(𝑒𝑡ℎ𝑎𝑛𝑒) + 3𝑂2 (𝑜𝑥𝑦𝑔𝑒𝑛) = 2𝐶𝑂2(𝑐𝑎𝑟𝑏𝑜𝑛 𝑑𝑖𝑜𝑥𝑖𝑑𝑒) + 𝐻2 𝑂(𝑤𝑎𝑡𝑒𝑟)
[3] 𝐶𝐻4
( 𝑚𝑒𝑡ℎ𝑎𝑛𝑒) + 3𝑂2(𝑜𝑥𝑦𝑔𝑒𝑛) = 𝐶𝑂2(𝑐𝑎𝑟𝑏𝑜𝑛 𝑑𝑖𝑜𝑥𝑖𝑑𝑒) + 𝐻2 𝑂(𝑤𝑎𝑡𝑒𝑟)
[4] 2𝐶6 𝐻6(𝑏𝑒𝑛𝑧𝑒𝑛𝑒) + 15𝑂2(𝑜𝑥𝑦𝑔𝑒𝑛) = 12𝐶𝑂2 (𝑐𝑎𝑟𝑏𝑜𝑛 𝑑𝑖𝑜𝑥𝑖𝑑𝑒) + 6𝐻2 𝑂(𝑤𝑎𝑡𝑒𝑟)
[5] 𝐶7 𝐻8(𝑡𝑜𝑙𝑢𝑒𝑛𝑒) + 9𝑂2 (𝑜𝑥𝑦𝑔𝑒𝑛) = 7𝐶𝑂2 (𝑐𝑎𝑟𝑏𝑜𝑛 𝑑𝑖𝑜𝑥𝑖𝑑𝑒) + 4𝐻2 𝑂(𝑤𝑎𝑡𝑒𝑟)
[6] 𝐶8 𝐻8(𝑠𝑡𝑦𝑟𝑒𝑛𝑒) + 10𝑂2(𝑜𝑥𝑦𝑔𝑒𝑛) = 8𝐶𝑂2(𝑐𝑎𝑟𝑏𝑜𝑛 𝑑𝑖𝑜𝑥𝑖𝑑𝑒) + 4𝐻2 𝑂(𝑤𝑎𝑡𝑒𝑟)
FIRE
R1
R2
EBHX
3PHX

21. 18
[7]2𝐶8 𝐻10(𝑒𝑡ℎ𝑦𝑙𝑏𝑒𝑛𝑧𝑒𝑛𝑒) + 21𝑂2
( 𝑜𝑥𝑦𝑔𝑒𝑛) = 16𝐶𝑂2(𝑐𝑎𝑟𝑏𝑜𝑛 𝑑𝑖𝑜𝑥𝑖𝑑𝑒) + 10𝐻2 𝑂(𝑤𝑎𝑡𝑒𝑟)
[8] 𝐶10 𝐻8
( 𝑛𝑎𝑝ℎ𝑡ℎ𝑎𝑙𝑒𝑛𝑒) + 12(𝑜𝑥𝑦𝑔𝑒𝑛) = 10𝐶𝑂2(𝑐𝑎𝑟𝑏𝑜𝑛 𝑑𝑖𝑜𝑥𝑖𝑑𝑒) + 4𝐻2 𝑂(𝑤𝑎𝑡𝑒𝑟)
[9] 𝐶12 𝐻12
( 𝑒𝑡ℎ𝑦𝑙 𝑛𝑎𝑝ℎ𝑡ℎ𝑎𝑙𝑒𝑛𝑒) + 15𝑂2(𝑜𝑥𝑦𝑔𝑒𝑛) = 12𝐶𝑂2 (𝑐𝑎𝑟𝑏𝑜𝑛 𝑑𝑖𝑜𝑥𝑖𝑑𝑒) + 6𝐻2 𝑂(𝑤𝑎𝑡𝑒𝑟)
The oil stream (3POIL) that is separated within the 3PSEP leaves the separator at 636.48 lbmol/hr, 85 oF
and 26 psia and is sent to a GIBBS reactor referred to as MAGSEP. This separator is utilized to separate water,
hydrogen, and methane from the oil stream before it is sent to the distillation columns. The MAGSEP vents the
gases into the atmosphere while it sends the oils to the columns through stream OILS2COL which are fed to
PRIMARY to begin the distillation process.
The column sequence was determined using the logic displayed in progress report 7 to produce the most
effective product separations. The input values for each distillation column is displayed in Table 5. Each of the
design columns that are displayed in table 5 are estimated to have a reflux ratio of -1.35. This value of -1.35 is
actually the equation of 1.35 * Minimum Reflux Ratio that is determined by ASPEN. See figure 3.3 for 3-
dimensional view of columns.
Column Input PRIMARY EBRECOL STHCOL BTCOL STLCOL
Light Key Ethylbenzene Toluene Styrene Benzene Styrene
Light Key Recovery 0.9995 0.999 0.9 0.9995 0.999
Heavy Key Styrene Ethylbenzene Naphthalene Toluene Naphthalene
Heavy Key Recovery 0.1 0.001 0.01 0.001 0.001
Condenser Pressure
(psia) 0.5801 14.7 0.5801 14.7 0.5801
Reboiler Pressure
(psia) 0.5801 14.7 0.5801 14.7 0.5801
Reflux Ratio -1.3 -1.35 -1.35 -1.35 -1.35
Table 3.5: Distillation Column Input Variables

22. 19
Figure 3.4: Styrene Plant Columns 3-D View
The OILS2COL stream enters the distillation column PRIMARY, which operates at vacuum conditions,
and two streams are produced, EBBENTOL and SH. The distillate, EBBBENTOL, emerges from the top of the
column. The bottoms product, SH, emerges from the bottom. EBBENTOL then must enter a pump to bring it to
14.7 psia, as EBRECOL operates at atmospheric conditions. The product streams of EBBENTOL are distillate,
BENTOL and bottoms, EBRECYCL. This stream is then fed to EBMIX and combined with fresh feed EBFEED
and sent to a pump to bring the EBTOTAL streamto reactor conditions of 26 psia as stated previously in this report.
The distillate, BENTOL, is fed to BTCOL which produces distillate product BENPROD and bottoms product
TOLPROD. These two streams are product quality benzene and toluene which will be stored in tanks and later sold.
The bottoms product of PRIMARY, SH, is fed to STHCOL and separated to STHPROD and SH2. As noted in
Table 5, STHCOL operates at vacuum conditions. This stream is the high purity styrene product desired in the
problem statement. The bottoms product is SH2. This stream is fed to STLCOL which also operates at vacuum
conditions. The products are, distillate, STLPROD (low purity styrene), and bottoms, HVYPROD(naphthalene and
ethyl-naphthalene).

23. 20
Radfrac columns were created using specs summarized in Table 6 in order to determine the correct column
sizing and rating.
Column Input RPRIMARY REBRECOL RBTCOL RSTHCOL RSTLCOL
Stages 49 39 30 17 23
Feed Tray Location 13 20 15 13 12
Pressure (psia) 0.5801 14.7 14.7 0.5801 0.5801
Column Diameter (ft) 21.286 4.651 1.503 8.736 2.936
Tray Spacing (ft) 2 2 2 2 2
Type of Tray Sieve Sieve Sieve Sieve Sieve
Tray Efficiency (% ) 100 100 100 100 100
Weir Height (in) 2 2 2 2 2
Reflux Ratio (mole) 4.14 5.707 2.5019 0.086244 0.143078
Distillate to Feed Ratio
(mole)
0.54 0.1436 0.36816 0.885179 0.855459
Table 3.6: Radfrac Distillation Column Input Values
The stream specifications used in the Radfrac columns were taken directly from the plant simulation.
Those streaminput values are summarized in Table 7.
Radfrac Stream Inputs ROILS2COL REBHIPT RBENZTOL RSHIN RSH2IN
Temperature (F) 85 99.003 206.681 128.559 133.706
Pressure (psia) 26 14.7 14.7 0.5801 0.5801
Total Flow Rate (lbmol/hr) 624.611 337.29 48.424 287.321 32.991
Ethylbenzene Flow Rate (lbmol/hr) 257.842 257.713 0.258 0.129 0.00542
Styrene Flow Rate (lbmol/hr) 313.806 31.381 0.000155 282.425 28.243
Benzene Flow Rate (lbmol/hr) 17.806 17.806 17.806 4.30E-26 0
Toluene Flow Rate (lbmol/hr) 30.39 30.39 30.36 7.22E-13 3.46E-15
Naphthalene Flow Rate (lbmol/hr) 2.383 5.92E-06 3.54E-11 2.383 2.359
Ethylnaphthalene Flow Rate
(lbmol/hr) 2.384 5.92E-06 3.54E-11 2.384 2.383
Table 3.7: Radfrac Distillation Feed Stream Input Values
The ROILS2COL stream has the exact specifications as OILS2COL in order to simulate conditions identical to the
plant conditions. This stream is fed to RPRIMARY which provides the proper column sizing and rating. The
distillate product stream is REBBENTOL and the bottoms product stream is RSH. REBHIPT, whose properties are
identical to that of EBTHIP, is fed to REBRECOL to produce RBENTOL as the distillate and REBRECYC as the
bottoms. RBENZTOL, copied from BENTOL, is fed to RBTCOL. The products of this column are RBENPROD
as the distillate and RTOLPROD as the bottoms. The stream RSHIN, copied from SH, is fed to RSTHCOL. The
distillate product is RSTHIPROD and the bottoms product is RSH2. Stream RSH2IN is copied from SH2 is fed to
RSTLCOL where the distillate product is RSTLPROD and the bottoms product is RHVYPROD.
Table 8 describes the five controllers utilized in our APSEN simulation.

24. 21
Specs TEMPR1 TEMPR2 COLDUT XSO2 CH4TRAD
Type Stream-Var Stream-Var Stream-Var Mole-Flow Stream-Var
Variable Temperature Temperature Temperature Excess Oxygen
Mole Fraction
Temperature
Reference
Stream
R1FEED R2FEED CLDUTOUT FIREFEED STACK
SPEC TR1 TR2 TC O2OUT/(O2IN-
O2OUT)
TSTACK
Target 1175°F 1175°F 120°F 0.25 250°F
Tolerance 0.01 0.01 0.01 0.01 0.01
Manipulated
Variables
WATFEED WATSPLIT CLDUTIN AIRFEED CH4FEED
Convergence 6540.02 lbmol-
water/hr
Fractions:
0.576 to WATR1
0.434 to WATR2
190268 lbmol-
water/hr
7272 lbmol-
air/hr
445 lbmol-
CH4/hr
Table 3.8: Controller Specs
TEMPR1 was the first established controller which allowed us to adjust the amount of water necessary to
heat the 965°F reactor one ethyl benzene/water mixture feed (R1FEED) up to reactor temperatures of 1175°F. Note
that the water splitter, WATSPLIT, was set at an inexact setting of a 6:1 (R1:R2) ratio. This allowed our input of
water to become an exact value. The total amount of water necessary to achieve the target temperature of 1175°F
was 6540.02 lbmol/hr.
Next, controller TEMPR2 was implemented. This controller was designed such that the correct amount of
water was split between reactor one (R1) and reactor two (R2). Referencing R2FEED, the reactor one products are
mixed with the water from the splitter (WATR2) such that the input stream of reactor two was at the target
temperature of 1175°F. The correct splitting of the WATHIPT streamwas 0.576 to WATR1 and 0.434 to WATR2.
The next controller, COLDUT, was used to control the heat exchanger, 3PHX. This heat exchanger was
designed such that it will cool down the products ofreactor two even further. Before the products can continue to the
distillation columns, they must be reduced to a temperature of 120°F. This was achieved using water. The controller
specified that 190,268 lbmol/hr of water was necessary to achieve this cooling duty.
The next controller, XSO2 was designed so that the fired heater (FIRE) could burn away all of the excess
vapors.After the three-phase separator, vapors are sent to FIRE in order to burn away unwanted product. In order to
achieve this, the excess oxygen of the fire heater was specified to a fractional amount of 0.25. This was set by using
O2OUT, the oxygen leaving the fired heater, and O2IN the oxygen entering the fired heater. This was done utilizing
a spec equation of
𝑂2𝑂𝑈𝑇
𝑂2𝐼𝑁−𝑂2𝑂𝑈𝑇
. By adjusting the AIRFEED stream, this was achieved with the controller. The final
stream flow was specified to 7272 lbmol/hr of air.
The final controller designed for the traditional styrene plant was the CH4TRAD. This controller was
utilized in order to achieve a target STACK gas stream temperature of 250°F. In order to achieve this, the controller

25. 22
manipulated the CH4FEED. After the controller was run, the final value of the CH4FEED was found to be 445
lbmol/hr of methane.
The following figures depict a 3-dimensional representation of what the styrene plant looks like from a
bird’s eye view. Each major component is labeled with the equivalent name given to it in aspen.
Figure 3.5: Styrene Plant Birds-Eye View
EBTank STHITank
TOLTANK
BENTANK
3PSEP
PRIMARY
R1 & R2
FIRE
STHPROD
BTCOL
STLPROD
EBRECOL

26. 23
4. Reactor Design
In industry, more than 75% of all styrene plants utilize adiabatic reactors to complete the dehydrogenation
of ethyl benzene. Our styrene plant being designed requires two adiabatic plug flow reactors with axial flow to
maximize conversion and minimize utility costs. The two reactors operate at a pressure of 26 psia. The temperature
of both reactors are determined by the input streams of steam and ethyl benzene which are measured at 1175°F. The
reason steam is added to the input stream is to reduce coking and to also raise both reactor to their respective
temperatures of 1175oF. The given specifics for the reactors were that the first reactor only had the process of
dehydrogenation of ethyl benzene and hence will only contain ethyl benzene, styrene, hydrogen and water. The
second reactor will include several side reactions and will include benzene, ethyl benzene, ethylene, ethyl-
napthalene, hydrogen, methane, naphthalene, styrene and toluene. The reactor is to include four feet of mixing space
and two separate bead volumes. These beads are to ensure the catalyst remains in the same position throughout the
reaction and are designed so that each beads height in the reactor is one foot. Each reactor is designed in such a way
that the optimal reactor will be produced. The following tables, figures and equations depict how the reactor will be
designed in terms of the design material of the reactors, the volume of the catalyst bed, the volume of the reactors,
the insulting material and sizing.
There are a variety of choices for the material the PFR reactors can be designed with. Our styrene plant will
utilize carbon steel as the reactor material which is the most commonly used material in processing plants. The
reasoning for this decision is because it is relatively cheap, strong, ductile, can be forged and welded, and has good
corrosion rates. Utilizing the Ergun equation, it was found that the pressure drop was 0.24 psia and 0.22 psia for
reactors one and two respectively. Since these are extremely low, the minimum thickness of reactors one and two
will be utilized at 1/8 inch (0.125”) and 1/4 inch (0.25”).
The catalyst used in this plant is the potassium promoted iron oxide catalyst. The composition of this
catalyst will be 77% Fe2O3 and 23% K2O, where the catalyst is the iron oxide, and potassium oxide acts as the
promoter having a total life span of 2 - 6 years. The catalyst is spherical in shape with a diameter of 1/8 inch for
both reactors used within the styrene plant.
The dehydrogenation surface reaction of ethyl benzene to form styrene has an intermediate step as shown:
Figure 4.1: Dehydrogenation of Ethyl Benzene

27. 24
The α-hydrogen and β-hydrogen create a partial bond with each of the two irons catalyst while
simultaneously the partial bonding between the carbons are occurring to stabilize the molecule forming the
intermediate molecule. In the final stage, the carbons form a double bond, breaking the bond between the two
styrene generating hydrogen molecules.
Reactions occurring in Reactor 1
EB
Conversion
21%
Reactions occurring in Reactor 2
EB
Conversion
33%
4%
0.5%
2.5%

28. 25
Table 4.1 Occurring reactions and their conversions in Reactors 1 and 2
In order to find the amount of catalyst required for each of the reactors, several calculations must be
performed. From aspen, utilizing a temperature sensitive curve for a perfect Gibbs Reactor, we were able to
determine the equilibrium constant Keb to be 0.08654
𝑘𝑚𝑜𝑙 𝐸𝐵
𝑘𝑔 𝑐𝑎𝑡 ∗ℎ𝑟∗𝑏𝑎𝑟
. Utilizing given values and equation (1) we were
able to determine the forward reaction constant:
𝑘𝑓 = exp ( 𝐴1 − (
𝐸1
𝑅𝑇
)) Eq. 4.1
Where A1 is the frequency factor of the reaction from collision theory, E1 is the activation energy (kJ/kmol) of the
reaction on the catalyst surface, T is the temperature of the reactor, and R is the gas constant.
𝑘𝑓 = exp (0.851 − (
90891 (
𝐽
𝑚𝑜𝑙
)
8.314 (
𝐽
𝑚𝑜𝑙 ∗ 𝐾
) ∗ 828 𝐾
)) = 4.32006 ∗ 10−6
𝑘𝑚𝑜𝑙 𝐸𝐵
𝑘𝑔𝑐𝑎𝑡 ∗ 𝑠 ∗ 𝑏𝑎𝑟
𝑘𝑓 = 4.32006 ∗ 10−6 𝑘𝑚𝑜𝑙 𝐸𝐵
𝑘𝑔 𝑐𝑎𝑡 ∗𝑠∗𝑏𝑎𝑟
∗ 3600 (
𝑠
ℎ𝑟
) = 0.015552232
𝑘𝑚𝑜𝑙 𝐸𝐵
𝑘𝑔 𝑐𝑎𝑡 ∗ℎ𝑟∗𝑏𝑎𝑟
.
Next we must determine the partial pressures of each reactant and product in the reactor. In order to do this we must
first find the mole expansion of the volume (𝜖).
𝜖 = 𝑦𝑒𝑏𝑜 ∗ 𝛿 Eq. 4.2
3.5%
1%
1%

29. 26
𝛿 (which is a relative term) represents the increase in the total number of moles per mole of ethyl benzene reacted
using stoichmetric coefficients. 𝑦𝑒𝑏𝑜 is the initial mole ratio of ethyl benzene to total moles:
𝛿 = # 𝐻20 + # 𝐻2 + # 𝑆𝑇 − # 𝐻20 − # 𝐸𝐵 = 8 𝐻20 + 1 𝐻2 + 1 𝑆𝑇 − 8 𝐻20 − 1 𝐸𝐵 = 1 Eq. 4.3
𝑦𝑒𝑏𝑜 =
# 𝑜𝑓 𝑒𝑡ℎ𝑦𝑙 𝑏𝑒𝑛𝑧𝑒𝑛𝑒 𝑚𝑜𝑙𝑒𝑠 𝑖𝑛𝑖𝑡𝑖𝑎𝑙𝑙𝑦
# 𝑜𝑓 𝑡𝑜𝑡𝑎𝑙 𝑚𝑜𝑙𝑒𝑠
=
1
8+1
=
1
9
Eq. 4.4
Equations 5 through 8 depict the partial pressures of ethyl benzene, styrene and hydrogen as seen below where xis
the conversion of ethyl benzene:
𝑃𝐸𝐵𝑜 = 𝜖 ∗ 𝑃𝑇𝑂𝑇 Eq. 4.5
𝑃𝐸𝐵 = 𝑃𝐸𝐵𝑜 ∗
1−𝑥
1+𝜖𝑥
Eq. 4.6
𝑃𝑆𝑇 =
𝑃 𝐸𝐵𝑜
1+𝜖𝑥
∗ 𝑥 Eq. 4.7
𝑃𝐻2 =
𝑃 𝐸𝐵𝑜
1+𝜖𝑥
∗ 𝑥 Eq. 4.8
Finally, the rate of the reaction can be calculated using:
𝑟 = 𝑘𝑓 ∗ (𝑃𝐸𝐵 −
𝑃 𝑆𝑇 ∗𝑃 𝐻2
𝐾 𝐸𝐵
) Eq. 4.9
This rate equation was calculated using conversion values in increments of 0.01. After this calculation was
performed, the area under the curve was determined by utilizing the trapezoidal rule and plotting r-1 vs conversion
(x) as seen in figure 2.

30. 27
Figure 4.2. Levenspiel Plot
Figure 4.3: Bulk Catalyst Volume
0
500
1000
1500
2000
2500
3000
3500
4000
0 0.05 0.1 0.15 0.2 0.25 0.3 0.35 0.4 0.45 0.5
1/r
x Conversion Fraction
1/r vs x
Reactor 2
Reactor 1

31. 28
The molar feed ratio of water to ethyl-benzene which favors the ethyl-benzene causing the selectivity to be
high, a 82.66% selectivity was achieved while the yield was found to be 47.1% and the conversion was 56.9% (this
conversion is accounting for the side-reactions that occur within the second reactor).
Next, the volume of the catalyst can be calculated. The specifics of each of our reactors were that the first
reactor reached a conversion of 22% and the second reached a conversion of 44%. The volume of the catalyst is
given by:
𝑉𝑐𝑎𝑡 = 𝐹𝐴0 ∗
∫ 𝑑𝑥
𝑟
𝜌 𝑝𝑎𝑟𝑡𝑖𝑐𝑙𝑒 (1−𝜙)
Eq. 4.10
Where FAo is the intial flow rate of ethyl benzene, 𝜌𝑝𝑎𝑟𝑡𝑖𝑐𝑙𝑒 is the density of the catalyst particle, and 𝜙 is
the porosity of the particle. In figure 3, found below, is the plot of volume of catalyst versus conversion.
In order to optimize the reactor size, we were specified to make the reactor have a height that is equal to the
diameter (i.e. d/h = 1). We were also specified to include a 4 foot mixing space with 1 foot of beads on the top and
bottom of the catalyst. In order to determine the diameter of the tank, we can utilize the catalyst’s volume since the
catalyst’s diameter will have the same diameter as the reactor.
The volume of the catalyst is given by equation 13 assuming the catalyst is spherical:
𝑉𝑐𝑎𝑡 = 𝜋 ∗
𝑑2
4
∗ (ℎ 𝑐𝑎𝑡) Eq. 4.11
Where hcat is the height of the catalyst and can be found by adding the specified mixing space and beads together
and subtracting off the height of the reactor. Thus the volume of the catalyst becomes :
𝑉𝑐𝑎𝑡 = 𝜋 ∗
𝑑2
4
∗ (ℎ − 6) Eq. 4.12
Since h = d for an optimal reactor size this equation is written as:
𝑉𝑐𝑎𝑡 = 𝜋 ∗
𝑑2
4
∗ ( 𝑑 − 6) Eq. 4.13
The volume of the reactors are given by:
𝑉𝑟𝑒𝑎𝑐𝑡𝑜𝑟1 = 𝜋 ∗ 𝑟1
2
∗ 𝐿1 = 𝜋 ∗
𝑑1
2
4
∗ 𝐿 = 𝜋 ∗
𝑑1
3
4
Eq. 4.14
𝑉𝑟𝑒𝑎𝑐𝑡𝑜𝑟2 = 𝜋 ∗ 𝑟2
2
∗ 𝐿2 = 𝜋 ∗
𝑑2
2
4
∗ 𝐿 = 𝜋 ∗
𝑑2
3
4
Eq. 4.15
Specification Reactor 1 Reactor 2
Catalyst Volume(ft3) 296.735 716.241
Diameter(ft) 9.877 12.164

32. 29
Table 4.2. Reactor Specs
Table 2 above, has values for the calculated catalyst volume, diameter of the reactor, and the inside reactor
volumes. It is important to note that the volumes of the reactors calculated here do not account for the insulating
material taking up portion of the volume of the reactor. Also, there are several assumptions that were made in order
to perform these calculations. First, the second reactor, though it contains side reactions, was assumed to only have
the dehydrogenation of ethyl benzene reaction. Next, the 𝛿 value that was calculated was for the first reactor. The
second reactors 𝛿 will be different from the first reactors because there are more moles associated with its input
stream. This value was assumed to remain constant from reactor to reactor. The ratios of catalysts were first
assumed to be 1/3 to 2/3 between reactors one and two. The actual catalyst split values (0.2929 and 0.7071) are
slightly different from the assumption values. Utilizing this information, the heat loss to the environment can be
calculated. The dehydrogenation of ethyl benzene is an endothermic reaction which is why the temperature of the
inlet stream changes from 1175°F to 1090°F. For our reactors, we will accept 10°F of temperature loss in our
reactors with a 50°F of accepted error. This equates to an exiting temperature of 1030°F. In order to achieve this
minor temperature loss, the reactors will use Skamol POROS 500 insulator10. The conductivity value of this
firebrick insulation is 1.18
𝐵𝑇𝑈∗𝑖𝑛
ℎ𝑟∗𝑓𝑡∗°𝐹
. The reason this insulator was chosen is because it had one of the lowest thermal
conductivities of the insulators researched that are used in processing plants. Inside each of the reactors, insulating
material is required to reduce temperature drop of the exiting stream of a perfectly adiabatic reactor. A perfectly
insulated adiabatic reactor would have an exit stream temperature of 1090°F. The heat loss from the reactors are
798,000 BTU/hr and 1.058*106 BTU/hr for reactors one and two respectively. In order to find the length of
insulation, a thermal transport equation can be applied. The following equations are for thermal resistivity[1]:
1
𝑅
=
1
𝑅1
+
1
𝑅2
Eq. 4.16
𝑅 𝑟𝑒𝑎𝑐1 =
𝐿1
𝑘∗𝐴1
+
𝐿2
𝑘2 ∗𝐴2
=
1
𝐴
∗ (
𝐿1
𝑘1
+
𝐿2
𝑘2
) =
1
𝐴
∗ (
0.125
252
+
𝐿2
1.18
) Eq. 4.17
Where k1 is the thermal conductivity of the insulator, k2 in the thermal conductivity of the carbon steel, L1
is the thickness of the insulator, L2 is the thickness of the carbon steel, A1 is the cross sectional area of the insulator
and A2 is the cross sectional area of the carbon steel. The thermal conductivity of carbon steel is 252
(BTU*in)/(hr*ft2*°F) and the thermal conductivity for insulating material is 1.18 (BTU*in)/(hr*ft2*°F). The total
heat transfer is given by:
𝑄
𝐴
=
Δ𝑇
𝑅
= (1175 − 70) ∗
𝐴
0.125
252
+
𝐿2
1.18
= 98,612
𝐵𝑇𝑈
ℎ𝑟
Eq. 4.18
𝐿 𝑟𝑒𝑎𝑐𝑡1𝑖𝑛𝑠𝑢𝑙𝑎𝑡𝑖𝑜𝑛
=.001 ft
𝐿 𝑟𝑒𝑎𝑐2𝑖𝑛𝑠𝑢𝑙𝑡𝑎𝑖𝑜𝑛
= 0.0093
Reactor Volume w/ Insulation (ft3) 756.685 1413.461
Catalyst Split 0.2929 0.7071

33. 30
The specification of our reactors in our styrene plant were such that the minimum amount of insulation
required is 6” meaning that both reactor one and reactor two will require 6” of insulation. Reinserting these
thicknesses into the equations, the true temperature drop can be calculated. Table 3 depicts all of the thermal
information calculated.
Reactor 1 Reactor 2
Heat Loss to environment (BTU/hr) 98,612 152,654.20
Temperature Drop (°F) 1.54 2.28
Insulation Thickness (inches) 6.0 6.0
Temperature of Steel Interface (oF) 70.12 70.17
Temperature Drop of Reactor (oF) 1.22 1.44
Table 4.3. Thermal Calculations
Using these values, the true value of the volume of the reactors can now be calculated. The new diameter of
the reactor one is 10.877 feet and reactor two at 13.164 feet. The cylindrical portion of the reactors are 1010.69 feet
and 1791.65feet for reactors one and two respectively. These volumes are for the volume of the carbon steel portion
of the reactors without insulation. There are also tori-spherical heads on both ends of the reactors. These volumes
are (both heads) 378.8 ft3 for the first reactor and the second reactor 707.86 ft3. Thus the true volume of the entire
reactors are 1389.49 ft3 and 2499.51 ft3 respectively.
Figure 4 and 5 shows the overall dimensions of the cylindrical portions of the reactors one and two
respectively (NOTE: Not to scale):
Figure 4.4: Reactor One Cylindrical Dimensions

34. 31
Figure 4.5: Reactor Two Cylindrical Dimensions
Using the Ergun equation ∆𝑃 = 150𝜇
1−𝜀
𝐷𝐺
+
7
4
(
1−𝜀
𝜀3
) ∗ (
𝐿
𝐷
)(
𝐺2
𝜌
) to calculate the pressure drop, it was found
that R1 had a pressure drop of 0.24 psia and R2 had a pressure drop of 0.22 psia. These low values of pressure helps
the reaction proceed in favor of the product and move from left to right. Since within the styrene plant we are
assuming reactors 1 and 2 are operating under ideal conditions at low pressures and high temperatures, the low
pressures ideally favor the products.
Using Gibb’s reactor to calculate the effect of temperature, pressure and water dilution on K equilibrium
and equilibrium conversion shows that equilibrium constant K is strongly dependent on temperature. As shown
below in figure 6, Keq is directly proportional to Temperature, as the reactor temperature increases, the equilibrium
constant increases resulting in an increase in the equilibrium conversion of the reaction. At higher temperatures the
reaction has a higher probability to proceed in favor of producing more styrene thus conversion and Keq is
prominently dependent on temperature.

35. 32
Figure 4.6. Effect of Keq with increasing temperatures
Figure 4.7. Effect of Keq with increasing pressure
Figure 7 shows a constant value of Keq as the pressure increases. As the chemical reaction of ethyl benzene
proceeds styrene and hydrogen, increases the total number of moles within the reactor. The steam that is added
throughout the process cancels the partial pressure interaction that the products (styrene and the side products) have
on the equilibrium constant. The Aspen simulation analysis displays that the conversion of styrene is inversely
proportional to pressure.
0
20
40
60
80
100
120
450 650 850 1050 1250 1450 1650 1850 2050
Keq
Temperature oF
Keq vs Temperature (oF)
0.4
0.42
0.44
0.46
0.48
0.5
0 2 4 6 8 10 12 14 16
Keq(Bar)
Pressure (Psi)
Keq vs Pressure

36. 33
Figure 4.8. Keq versus Moles of Water
Figure 4.8 shows that the equilibrium constant does not depend on the moles of water.
Figure 4.9: Conversion effects of Temperature, Pressure, and Dilution
Figure 4.9 shows the effect on conversion on temperature, water dilution, and pressure. As temperature
increases, the amount of conversion also increases. It is important to note that this value reaches a maximum and
then the temperature does not affect the conversion anymore (this is when conversion is equal to 1). As the moles of
water increases, the conversion also increases. Finally, as pressure increases the conversion decreases. This is
because of Le Chatelier’s principle. There are more products than reactants because of all of the side products that as
pressure increases, the equilibrium will drive towards the reactants resulting in less conversion of ethyl benzene to
styrene.
0.4
0.45
0.5
0.55
0.6
0 5 10 15 20 25
KEQBar
Water (Moles)
Keq vs Water mole
0
0.2
0.4
0.6
0.8
1
0 0.2 0.4 0.6 0.8 1
Conversion
% of final values
Effect on Conversion
Conversion for
Temperature
Conversion for
dilution
Conversion for
Pressure

37. 34
5. Column Sequencing
A sequence within a plant is a specific order in which distillation columns are strategically placed in order
to separate the components efficiently. For a two component separation, there is only a single accepted plant
combination for separation available. For a three component separation there are two design combinations, five
components there are 14 design combinations, and for 10 components there are approximately 4,800 design
combinations. A heuristic is a guideline utilized to increase the process of properly setting up a plants distillation
column ultimately cutting costs during the designing stage of the plant. Many heuristics include the removal of
corrosive materials and reactive monomers or other reactive component which will lessen the potential damage that
may be caused to the column; Removing products as distillates resulting in a more pure final product; Removing the
recycle streams as distillates; Removing the lightest products first, and favor equal molar splits. Although these
guidelines seem easy to comprehend and potentially fulfill, it is often unlikely that many of the sequences
potentially used are able to satisfy each of the heuristics described.
Heuristic Reason
Remove corrosive materials as soon as possible Corrosive resistant materials are expensive
Remove reactive components or monomers as soon as
possible
Reactive components may foul the re-boilers in the
distillation columns
Remove products and recycle streams as distillate in a
packed column
To minimize contamination from heavies such as rust
Remove the most plentiful compound first To reduce distillation column size and also the amount
of materials being dealt with
Do the most complicated separations last That way there are no other columns that must maintain
a complicated design to accommodate for the complex
separation.
Favor equi-molar splits in the column Different molar rates will cause the boil up to dry out
and to prevent liquid entrainment in the distillation trays.
Obtain recycle early To reduce contamination to the recycle stream
Minimize the number of vacuum columns Vacuum columns are more expensive than ordinary
columns
Do separations of the highest purity last To maintain the purity at the highest at the last and not
be contaminated from other columns
Table 5.1: Heuristics for Column Sequencing

38. 35
The best sequence for which a distillation column process is designed is seen in figure 1:
Figure 5.1. Best Column Sequence
Rules Followed Rules Disobeyed
Product is distillate
The most difficult separation is done first
Most plentiful exits the column first
Obtained the recycle stream early in the process
Ethyl Benzene recycle stream is removed as a bottom
product stream
Used a minimum amount of vacuum columns
Subsequent columns are cheaper than a traditional
column
Table 5.2: Rules and guidelines followed and disobeyed from the best column sequencing

39. 36
Figure 5.2: Second Best Column Sequence
Rules Followed Rules Disobeyed
Recovered products as distillate
The bottomproducts of the first two columns and
much greater than the distillate (columns >> distillate)
Recovered the recycle stream as distillates
Monomer makes its way through three separate
columns
Subsequent columns are more expensive
Table 5.3: Rules and guidelines followed and disobeyed from the second best column sequencing

40. 37
Product Streams Mass Flow Rates (lb/hr)
Streams
Low-Purity
Styrene
High-Purity
Styrene
Benzene Toluene
Component
Total
H2O 0 0 0 0 0
Ethyl Benzene 0.574998 13.1121 0 27.3607 41.0478
Styrene 2938.56 26473.5 0 0.0161151 29412.0761
H2 0 0 0 0 0
Benzene 0 0 1390.22 0.695457 1390.9155
Toluene 0 6.624*10-11 2.79737 2794.57 2797.3674
Methane 0 0 0 0 0
C2H4 0 0 0 0 0
CO2 0 0 0 0 0
N2 0 0 0 0 0
O2 0 0 0 0 0
Naphthalene 0.302384 3.05438 0 4.53548*10-09 3.3568
Ethyl
Naphthalene
0.0019148 0.0209647 0 5.52955*10-09 0.02288
Stream Total 2939.4347 26489.6874 1393.0174 2822.6423 ---
Needed wt% and Volume % vs. Actual wt% and Volume %
Low-Purity Styrene
High-Purity
Styrene
Benzene Toluene
Required wt% > 98 > 99.8% ≥ 99.0% ≥ 99.0%
Actual wt% 99.97% 99.93% 99.8 99.0%
Required Vol. % ≤ 10% ≥ 90% --- ---
Actual Vol. % 9.99% 90% --- ---
Table 5.4: Mass flow rates (lb/hr) of Product streams in from distillation columns
The table above displays the mass flow rates in lb/hr from the product stream results within the Aspen
simulation. From the specifications needed, the amount of high-purity styrene needed at 99.8 wt% purity is at least
90% of the final product, leaving at most, 10% of the 97 wt% purity low-purity styrene. Subsequently for the
benzene and toluene, each of these products must be at least 99 wt% purity. Equations 1 – 6 below display the
purities and the percentage amount from each of the product streams in table 3:
𝐻𝑖𝑔ℎ 𝑃𝑢𝑟𝑖𝑡𝑦 𝑆𝑡𝑦𝑟𝑒𝑛𝑒 ∶
26473.5 𝑙𝑏/ℎ𝑟
26489.6874 𝑙𝑏/ℎ𝑟
= 99.93 𝑤𝑡% > the required 99.8 wt% Eq. 5.1

41. 38
𝐻𝑖𝑔ℎ 𝑃𝑢𝑟𝑖𝑡𝑦 𝑆𝑡𝑦𝑟𝑒𝑛𝑒 𝑉𝑜𝑙. % ∶
26473 .5 𝑙𝑏/ℎ𝑟
29412.0761 𝑙𝑏/ℎ𝑟
= 90 % of the total styrene volume Eq. 5.2
𝐿𝑜𝑤 𝑃𝑢𝑟𝑖𝑡𝑦 𝑆𝑡𝑦𝑟𝑒𝑛𝑒 ∶
2938.56 𝑙𝑏/ℎ𝑟
2939.4347 𝑙𝑏 /ℎ𝑟
= 99.97 𝑤𝑡% > the required 97 wt% Eq. 5.3
𝐿𝑜𝑤 𝑃𝑢𝑟𝑖𝑡𝑦 𝑆𝑡𝑦𝑟𝑒𝑛𝑒 𝑉𝑜𝑙. % ∶
2938.56 𝑙𝑏/ℎ𝑟
29412 .0716 𝑙𝑏/ℎ𝑟
= 9.99 % of the total styrene volume Eq. 5.4
As seen in equations two and four, the total styrene volume does not equal 100%, this is because the
toluene product streamproduces a minimal amount of styrene accounting for ~ 0.01% of the total styrene.
𝐵𝑒𝑛𝑧𝑒𝑛𝑒 ∶
1390 .22 𝑙𝑏/ℎ𝑟
1393.017 𝑙𝑏/ℎ𝑟
= 99.8 𝑤𝑡% > the required 99 wt% Eq. 5.5
𝑇𝑜𝑙𝑢𝑒𝑛𝑒 ∶
2794.57 𝑙𝑏/ℎ𝑟
2822 .6423 𝑙𝑏/ℎ𝑟
= 99.0 𝑤𝑡% ≥ the required 99.0 wt% Eq. 5.6
As seen from equations 1-6 or the values displayed at the bottom of table 3, the requirements needed to
operate the designed styrene plant have been met for each of the low and high purity styrene, benzene and toluene
from the use of the design column values.
RPRIMARY REBRECOL RBTCOL RSTHCOL RSTLCOL
# Trays 49 39 30 17 23
Feed Tray 13 20 15 13 12
Condenser P 30 torr 14.7 psia 14.7 psia 30 torr 30 torr
Pressure Drop (psi) 3.27348 3.84422 3.32234 1.20447 1.31586
Diameter (feet) 21.286 4.646 1.503 8.736 17.844
Spacing (feet) 2 2 2 2 2
Type Sieve Sieve Sieve Sieve Sieve
Thickness (gauge) 10 10 10 10 10
Hole Diameter (ft) 0.0417 0.0417 0.0417 0.0417 0.0417
Weir Height (in) 2 2 2 2 2
Weir Length (ft) 15.467 3.376 1.092 6.348 2.134
Condenser Duty
(BTU/hr)
-30420590.6 -4670852.4 -817530.647 -4893174.21 -571379.955
Condenser Type Water Water Water Water Water

42. 39
RPRIMARY REBRECOL RBTCOL RSTHCOL RSTLCOL
Condenser Temp
(oF)
98.937 207.411 176.398 127.937 297.235
Cooling Water T
(oF)
70 70 70 70 70
Re-boiler Type
Low
Pressure-
Steam
Low
Pressure-
Steam
Low
Pressure-
Steam
Low
Pressure-
Steam
Low
Pressure-
Steam
Saturated Steam T 327.74 327.74 327.74 327.74 327.74
Re-boiler Duty
(BTU/hr)
32215256.8 7570372.04 834086.525 4889416.16 596393.264
Re-boiler Temp 214.15 294.61 234.05 133.73 249.68
Material C Steel - C C Steel - C C Steel - C C Steel - C C Steel - C
Wall Thickness (in) 0.011055 0.002413 0.000781 0.004537 0.009267
Actual Thickness
(in)
0.125 0.125 0.125 0.125 0.125
Tray Efficiency 100 100 100 100 100
Table 5.5: Styrene Plant Distillation Column Specifications
When designing the styrene plant distillation columns, there are several parameters, assumptions and
correlations that were applied to the Aspen simulation to simplify the process. The simulation in Aspen simulates
two different types of column designs, the design column and the radfrac column. The design column is based on
the desired bottom stream products and applies four different equations, the Fenske, Gilliland, Kirkbride and
Underwood. These equations are applied to solve for each columns respective number of equilibrium stages,
optimum feed stage, minimum reflux ratio and number of minimum theoretical plates as seen in equations 7 – 10.
Unlike the design column, the radfrac column does not specifically use any of the listed equations but instead
utilizes the answers as inputs to solve for column pressure drop and the column diameter. Applying these numbers
from equations 7 – 10 within the Aspen simulation, Aspen is able to estimate the pressure drop needed for gas to
pass through the tray and ultimately through the liquid layer.
The Fenske Equation was used to solve for the minimum number of theoretical plates required within a
fractional column when operating at total reflux:
𝑁, 𝑚𝑖𝑛 =
𝐿𝑜𝑔[(
𝑥 𝑏
𝑥 𝑡
)(
𝑥 𝑡
𝑥 𝑏
)]
log(𝛼 𝑎𝑣𝑔)
Eq. 5.7
 Nmin is the minimum number of theoretical
plates required at total reflux
 Xt is the mole fraction of more volatile
component in the overhead distillate

43. 40
 Xb is the mole fraction of more volatile
component in the bottoms
 is the average relative volatility of the
more volatile component
The Gilliland correlation, shown in equation 3, was used to find the number of equilibrium stages for a
given reflux ratio, in this case using the minimum reflux ratio and the minimum number of stages calculated:
𝑁−𝑁 𝑚𝑖𝑛
𝑁+1
= 0.75 − 0.75 ∗ (𝑋0.5668) Eq. 5.8
 RD is the actual reflux ratio
 RDmin is the minimum reflux ratio calculated
 N is the number of equilibrium stages
 Nmin is the minimum number of stages
calculated using the Fenske Equation
(Equation 7)
 Where  𝑋 =
𝑅 𝐷−𝑅 𝐷𝑚𝑖𝑛
𝑅 𝐷+1
The Kirkbride equation is used to determine the optimum stage to set the feed:
𝐿𝑜𝑔
𝑁𝑒𝑛𝑟𝑖𝑐ℎ𝑖𝑛𝑔
𝑁𝑠𝑡𝑟𝑖𝑝𝑝𝑖𝑛𝑔
= (0.206) log[(
𝑥 𝑡𝑜𝑙𝑢𝑒𝑛𝑒,𝑓
𝑥 𝑏𝑒𝑛𝑧𝑒𝑛𝑒,𝑓
)
𝑊
𝐷
(
𝑥 𝑏𝑒𝑛𝑧𝑒𝑛𝑒,𝑏𝑜𝑡𝑡𝑜𝑚
𝑥 𝑡𝑜𝑙𝑢𝑒𝑛𝑒,𝑑𝑖𝑠𝑡𝑖𝑙𝑎𝑡𝑒
)
2
] Eq.5.9
 N,enriching. is the number of trays above the feed
stage
 N,stripping is the number of trays below the feed
stage
 B is the molar flow rate of the bottoms stream
 D is the molar flow rate of the distillates
 D is the molar flow rate of the distillates
In order to find the minimum reflux ratio, the Underwood equation is utilized:
(
𝐿
𝐷
)
𝑚𝑖𝑛
+ 1 = ∑
𝑥 𝐷
(𝛼−𝜃)/𝛼
𝑛
1 Eq. 5.10
In addition to the designs created in Aspen, several parameters required calculating based on known
physical parameters. Using AP-135B carbon steel, the thickness of the column shell was calculated using equation 6.
𝑡 =
𝑃𝑅
𝑆𝐸−0.6𝑃
Eq. 5.11

44. 41
 t is thickness
 P the pressure drop across the column wall,
assuming ΔP = 15 psia
 R is the column radius
 S is the strength of the selected material
 E is the elasticity of the selected material
The primary columns’ function is to successfully separate the final product styrene and the remaining unreacted
ethyl benzene. Because the boiling points of each of these materials are close to one another (BP,styrene = 293oF ,
BP,ethyl-benzene = 276.8oF) making this separation a high degree of difficulty to complete. For this reason, this
distillation column is tall because of the large number of stages to ensure the completion of the separation. The feed
rate into the column contains all of the plant products which therefore requires the column to have a larger diameter
than the remaining columns. The top stream from the primary column contains ethylbenzene, benzene and toluene,
which becomes the feed stream to the ethyl benzene recycle column. This column is also tall because of the high
volume of its feed due to the nature of separation occurring. With the boiling point of benzene at 176.2°F and
toluene at 231.1°F, this particular separation process isn’t difficult but the ethyl benzene but this product must be
high quality because it will be recycled back into the initial feed stream within the styrene plant. The Benzene-
Toluene column is fed by the top of the ethyl benzene recycle column, and designed to separate benzene and
toluene. This column has the smallest diameter in the plant due to the small volume of the feed stream. The bottom
from the primary column feeds the high-purity styrene column. With styrene and the heavies within the high-purity
styrene column, this is a simple distillation process with a large volume of feed which makes this column needing a
large column diameter. The final column, low-purity styrene column is fed with the tops from the high-purity
styrene column and is designed to produce the purest form of styrene and separate it fromthe lower quality styrene
making this column require a significant amount of trays but is narrow due to the small volume flow rate.

45. 42
The primary columns’ function is to successfully separate the final product styrene and the remaining unreacted
ethyl benzene. Because the boiling points of each of these materials are close to one another (BP,styrene = 293oF ,
BP,ethyl-benzene = 276.8oF) making this separation a high degree of difficulty to complete. For this reason, this
distillation column is tall because of the large number of stages to ensure the completion of the separation. The feed
rate into the column contains all of the plant products which therefore requires the column to have a larger diameter
than the remaining columns. The top stream from the primary column contains ethylbenzene, benzene and toluene,
which becomes the feed stream to the ethyl benzene recycle column. This column is also tall because of the high
volume of its feed due to the nature of separation occurring. With the boiling point of benzene at 176.2°F and
toluene at 231.1°F, this particular separation process isn’t difficult but the ethyl benzene but this product must be
high quality because it will be recycled back into the initial feed stream within the styrene plant. The Benzene-
Toluene column is fed by the top of the ethyl benzene recycle column, and designed to separate benzene and
toluene. This column has the smallest diameter in the plant due to the small volume of the feed stream. The bottom
from the primary column feeds the high-purity styrene column. With styrene and the heavies within the high-purity
styrene column, this is a simple distillation process with a large volume of feed which makes this column needing a
large column diameter. The final column, low-purity styrene column is fed with the tops from the high-purity
styrene column and is designed to produce the purest form of styrene and separate it fromthe lower quality styrene
making this column require a significant amount of trays but is narrow due to the small volume flow rate.

46. 43
6. Other Unit Operations
Heat Transfer Equipment
Heat exchangers defined as pieces of equipment that are used to exchange thermal energy between two
fluids. The most common type of heat exchanger is the tube and shell heat exchanger, which is used in refineries
petro-chemical and power processes. Shell and tube heat exchanges can operate under a wide range of temperatures
and pressures and withstand large heat duties. The materials of construction can be diverse without affecting a
technician’s ability to repair or maintain the device. As its name suggests, the two main components are a shell and
tubes, but most also have baffles. Baffles are used to support tubes and divert the flow on the shell side and increase
the heat transfer coefficient with the exchanger. Heat is exchanged through the tube walls either fromor to the shell
side fluid. Having lots of tubes causes the velocity of the fluid to increase as well as increasing the heat transfer
coefficient.
The Heat exchanger designed in this report are the heat exchangers for the pinched network. For this heat
exchange network the Heat loss is approximately zero. None of the singular heat exchanges in this network are done
in parallel but instead flow in a countercurrent manner. This is to achieve the best heat transfer between the two
fluids which makes countercurrent heat exchange common within styrene plants.The tube section of this heat
exchanger is typically utilized for the fluid that has the most buildup for it is much easier to clean than the shell
section.
Figure 6.1. Countercurrent Flow Shell and Tube Heat Exchanger
In the analysis, two heat exchangers were designed, the ethyl benzene heat exchanger (EBHX) and the
three phase heat exchanger (3PHX). The EBHX heat exchangers purpose was to increase the temperature of the
ethyl benzene from 172oF to 965oF which was achieved by using the hot stream R2PROD from reactor 2 which is
cooled from 1065.49 oF to 695.18 oF.
ffasdfffasdf
ffasdf

47. 44
Figure 6.2. Ethyl Benzene Heat Exchange Network (EBHX)
Figure 6.3. Ethyl Benzene Heat Exchange Network Inner Tube Layout

48. 45
Table 6.1 TEMA Sheet for EBHX
As shown in figure 2, 3 and figure 1 above, the cost effective carbon steel EBHX is 157.2 inches long and
46 inches wide consisting of 486 tubes at 96 inches in length and 0.75 inches in diameter, arranged in a 30-

49. 46
Triangular pattern single segmental baffle type weighing 5283 lbs when empty. This shell and tube heat exchanger
after labor, tube and material costs comes to an estimated price of $109,194.
The second heat exchanger modeled was the 3PHX and was utilized to cool the hot stream R2PROD
leaving the EBHX from 703oF to 85oF. This was achieved by using the cold water stream CLDUTIN ultimately
being heated from 70oF to 120oF.
Figure 6.4. Three Phase Heat Exchanger (3PHX) Network
Figure 6.5. Three Phase Heat Exchanger (3PHX) Network Inner Tube Layout

50. 47
Table 6.2: TEMA Sheet for EBHX
As shown in figure 4 and 5 above, the cost effective carbon steel EBHX is 309.9 inches long and 38 inches
wide consisting of 1101 tubes at 120 inches in length and 0.75 inches in diameter, arranged in a 60-Rotated-
Triangular pattern double segmental baffle type weighing 19,343.1 lbs when empty. This shell and tube heat
exchanger after labor, tube and material costs comes to an estimated price of $154,492.

51. 48
Heat
Exchanger
Labor Cost Tube Costs
Material
Costs
Number of
Shells
Cost per
Shell
Total Cost
EBHX $85,969 $10,173 $13,050 1 $36,398 $109,194
3PHX $100,033 $34,568 $19,891 1 $77,246 $154,492
Table 6.3. Heat Exchanger Cost Summary
By the use of ASPEN simulation we were able to identify the acceptable estimated pressure drop associated with
each heat exchanger:
Heat Exchanger Estimated Pressure Drop (psia)
EBHX 2
3PHX 3
Table 6.4: Estimated Pressure Drops
In order to establish a heat exchanger that simulates the WHX, water heat exchanger, a simple method of
designing a fired heater was used.This fired heater will be a cylindrical heater with vertical, annular tubes in the
radiant zone. Note that this method is very approximate, simplistic and crude. The following tables equations are
how the dimensions of the fired heater were determine.
Stream Temperature (°F) Enthalpy (BTU/lbm)
Input 70 38.8
Output 1425 1762.2
Table 6.5: Enthalpies of WHX Streams
The above table utilized steam tables in order to determine the enthalpies of the input and output streams of
the water heat exchanger. Using these values,the total duty used for WHX to convert the water at 70°F to steamat
1425°F can be calculated utilizing thermodynamics
𝐷𝑢𝑡𝑦 (
𝐵𝑇𝑈
ℎ𝑟
) = 𝑚̇ ∗ (𝐻 𝑜𝑢𝑡 − 𝐻𝑖𝑛) Eq. 6.1
𝐷𝑢𝑡𝑦 = 117820
𝑙𝑏
ℎ𝑟
∗ (1762.2
𝐵𝑇𝑈
𝑙𝑏𝑚
− 38.8
𝐵𝑇𝑈
𝑙𝑏𝑚
) = 2.03 ∗ 108 𝐵𝑇𝑈
ℎ𝑟
Eq. 6.2
where 𝑚̇ is the mass flow rate, Hout and Hin are the enthalpies of the output and input streams respectively.
Next, the volume of the fired heater can be calculated. In order to calculate the volume, the approximation of radiant
heat transfer release in combustion chamber must be used,76,800 BTU/(hr*ft3). The equation for volume is given
by
𝑉𝑜𝑙𝑢𝑚𝑒 ( 𝑓𝑡3) =
𝐷𝑢𝑡𝑦
𝐵𝑇𝑈
ℎ𝑟
76,800
𝐵𝑇𝑈
ℎ𝑟∗𝑓𝑡3
=
2.03∗108 (
𝐵𝑇𝑈
ℎ𝑟
)
76,800
𝐵𝑇𝑈
ℎ𝑟∗𝑓𝑡3
= 2644.844 𝑓𝑡3
Eq. 6.3
Now that the volume has been calculated, the height and diameter of the fired heater can be calculated.
Anotherequation for volume using the fired heater’s dimensions is assuming that diameter is equal to height
𝑉𝑜𝑙𝑢𝑚𝑒( 𝑓𝑡3) =
(𝐷𝑐𝑧
2
∗𝜋∗𝐻)
4
= 0.785 ∗ 𝐷𝑐𝑧
2
∗ 𝐷 Eq. 6.4

52. 49
Where Dcz is the combustion zone diameter of the fired heater. In order to determine, Dcz and D, there are
several equations that can be used.First the area of the tubes needs to be calculated
𝐴𝑟𝑒 𝑎 𝑇𝑢𝑏𝑒𝑠
( 𝑓𝑡2) =
𝐷𝑢𝑡𝑦 (
𝐵𝑇𝑈
ℎ𝑟
)
𝐻𝑒𝑎𝑡 𝑡𝑟𝑎𝑛𝑠𝑓𝑒𝑟 𝑟𝑎𝑡𝑒
=
2.03∗108 𝐵𝑇𝑈
ℎ𝑟
14,000
𝐵𝑇𝑈
𝑓𝑡2∗ℎ𝑟∗°𝐹
= 27709.88 𝑓𝑡2
Eq. 6.5
After the area of the tubes has been calculated, anotherequation representing the area of the tubes using the
diameter of the fired heater and the number of tubes can be used.
𝐴𝑟𝑒 𝑎 𝑇𝑢𝑏𝑒
( 𝑓𝑡2) = 𝜋 ∗ (
1
6
)𝐷 ∗ 𝑁𝑡 = 13854.94 𝑓𝑡2
Eq. 6.6
𝑁𝑡 =
27709.88
𝐷
Eq. 6.7
The cross section of the tube sheet equation is given by:
0.087 ∗ 𝑁𝑡 = 0.785 ∗ (𝐷2
− 𝐷𝑐𝑧
2
) Eq. 6.8
Solving these equations, Dcz, D, and Nt can be calculated. The answers to these are given in the following table.
Variable Value
Dcz (ft) 13.457
D = H (ft) 18.61
Nt (tubes) 1489
D with insulation (ft) 20.61
Table 6.6: Calculated Dimensions

53. 50
Separation and Storage Tanks
In the chemical industry,storage tanks are typically used to hold liquids and vapors.These materials can be
reagents,products or even waste material. There are a few things that need to be considered in tank design.These
include material of construction,geometry, roof type, as well as other features such as heating, cooling or mixing
requirements. It is imperative that the material of a tank must be able to withstand the temperature and pressures of
the fluid being stored,otherwise a rupture is likely to happen. The type of roof is also important for safety reasons.
A floating roof is well suited for highly volatile materials; the possible flammable/explosive vapors are easily vented
from the gaps between the floating roof and the tank wall. Alternatively, a fixed roof with an inert gas blanket may
also be used to prevent fires and explosions. Mixers can be implemented if the contents require a desired consistency
and heat exchangers for contents that require heating and/or cooling. An inexpensive design is one of cylindrical
shape; however tanks that hold high pressure material are usually spherical to minimize the surface area to volume
ratio. The vapor pressures ofthe products and feed tanks are below 11 psia so cylindrical tanks will be utilized in the
styrene plant. The following figure depicts the basic shape of each storage tank.
Figure 6.6: Tank Diagram
The following storage tanks are designed to store product and feed materials for seven days of continuous
production. The tanks are cylindrical in geometry with equivalent height and diameter and a hard cone roof. The
wall thickness is determined by ASTM equations based on maximumpressure drop across the tank wall. A pressure
of 15 psia was accounted for the wall thickness in case the tanks became pressurized. Another important factor to
consider when constructing the tanks is an issue of safety on how full the tank should be. For safety reasons, a 10%
increase in the tanks volume will be added to account for vapor which may accumulate. In order to maximize the

54. 51
safety of the tanks, carbon steel was chosen as the material for the storage tanks due to its compatibility, strength
and temperature tolerance.
In the styrene plant facility being designed there are in total 6 storage tanks. The first tank is a feed storage
tank utilized for ethyl benzene which will be the largest tank in the facility. The other 5 tanks are used to store
products; benzene, toluene, high purity styrene, low purity styrene, and heavies. It is important to note that both the
ethyl benzene and high purity styrene tanks will have a larger thickness than all of the other tanks due to the high
volume of product inside each of these tanks. The larger volume tanks will induce large pressures inside the tank
from vapor. Extreme temperatures are considered when creating the storage tanks of 100°F and -20°F. Benzene is
stored well within its freezing temperatures and thus a heater will be required to ensure benzene does not solidify
and ensure it is pump-able liquids. The heavies, naphthalene and ethyl-naphthalene, are solid at roomtemperatures
and because of this a heater was also installed on the heavies storage tank this is so that the removal of the heavies is
easy and efficient. The following table depicts all of the chemical properties of the products and tank sizes for all of
the required storage tanks.
Ethyl
Benzene
Benzene Toluene
Hi-Purity
Styrene
Lo-Purity
Styrene
Heavies
CHEMICAL PROPERTIES
T Boiling Pt. (Fº) 277.16 176.18 232.00 293.00 293.00 424.00
Ethyl
Benzene
Benzene Toluene
Hi-Purity
Styrene
Lo-Purity
Styrene
Heavies
T Freeze Pt. (Fº) -139.00 41.90 -139 -23.8 -23.8 176.00
T Flash Pt. (Fº) 72 11.07 43.00 88.00 88.00 174.00
Comp. Vol (ft3) 116,515.59 4,445.90 9,851.76 84,467.74 9,374.04 1,991.43
Auto Ignition(°F) 914
TANK PROPERTIES
Flow Rate (lb/hr) 36,733.92 1,393.016 2,822.641 26,486.7 2,939.44 677.382
Volume (safety) ft3-
week
128,167.15 4,879.49 10,836.94 92,914.51 10,311.44 2190.57
Material (ft)
Carbon
Steel
Carbon
Steel
Carbon
Steel
Carbon Steel Carbon Steel
Carbon
Steel
Height (ft) 54.65 18.387 23.99 49.09 23.59 13.079
Diameter (Ft) 54.64 18.387 23.99 49.09 23.59 13.079
Thickness (in) 1.000 0.250 .375 0.875 0.375 0.250
Vapor Pres. (psia) 0.19 1.45 0.087 0.087 0.087 0.0
Roof Type Hard Cone
Hard
Cone
Hard Cone Hard Cone Hard Cone
Hard
Cone
Blank Gas N2 N2 N2
O2-enriched
N2
O2-enriched
N2
N2
Concentration
(ppm)
--- --- --- 19.25 19.23 ---
Heating No Yes Yes No No Yes
Cooling Duty
(BTU/hr)
NA NA NA 288,737.7 67,464.02 NA
Mixer Type None Sidewall Sidewall Top Mount Top Mount None

55. 52
Ethyl
Benzene
Benzene Toluene
Hi-Purity
Styrene
Lo-Purity
Styrene
Heavies
Exterior Color White White White White White White
Insulation None None None None None None
Table 6.7: Chemical and Storage Tank Properties
The main purpose of our facility is to produce styrene and store it in a manor such that the quality of the
product is maintained. In the plant being specified there are two forms of styrene,high and low purity thus two tanks
must be made to store these products.Assuming that the maximum ambient temperature will be 100°F, a cooling
unit must be installed for safety purposes.A Freon refrigerant will be installed at the top of the tank to ensure the
tank maintains a temperature of 60°F. It is of utmost important that the cooling duty be performed because styrene at
high temperatures can begin to polymerize which can reduce the quality of the styrene.The cooling was through
heat loss from the ambient air as well as the ground. The total amount of heating duty for the high purity and low
purity styrene tanks was found to be 288,737.7 BTU/hr and 67,464.02 BTU/hr respectively. In addition to the
cooling, two polymerization inhibitors were also placed within the styrene tanks. The inhibitors used at our facility
are tertiary-butyl-catechol (TBC) and di-nitro-cresol (DNC). TBC requires dissolved oxygen for the inhibitor to
perform property so 10-20ppm of dissolved O2 was added to the tank. The O2 cannot be added prior to the column
separators which is why DNC was also included as an inhibitor as well. DNC does not require any O2 to perform
optimally. Oxygen creates a problem because it will increase the flammability of the styrene thus O2 enriched with
N2 was added to the tank. It was determined that 2.5% oxygen concentration was required with the nitrogen in order
to maintain 15ppm of oxygen at 60°F. During the column separation process,some vapors still remain within the
system. The vapors are removed off of the column and the vacuumis pulled via heat exchangers. The vapors are
then compressed and the liquid is then taken to the storage tank. Throughout this removal process,the column is
purged with an inert gas and then the vacuumis pulled. This is done in contrary to pulling the vacuum first because
it is cheaper. Inside the storage tank, when the styrene is being cooled, a pump is pumping styrene into the tank at an
angle, which induces agitation thus a mixer is not required for this storage tank. The tanks were also lined with zinc
primer to ensure no polymerization occurred on sites inside the vessel.The outside of the tanks were painted white
to reduce thermal radiation absorption.
The three phase separator is designed to separate the incoming streaminto vapor, organics, and free water.
This separator is useful in refining a material streamthat contains different phases such as was implicated within the
traditional plant design. This designed three phase separator was designed to separate the vapors, liquids, and heavy
liquids (oils) from each other before being sent to their corresponding distillation columns for further refining.
Separators can feature different designs such as the orientations (vertical or horizontal), a weir, or a boot
design.The design of the separatordepends sole upon the streamthat is in need of separation. Vertical separators are
used when a significant amount of the inlet stream contains vapor ranging from 10-20% by weight and thus is not
needed within the traditional plant for the vapor stream is <10% by weight. The amount of heavy liquids within the
stream is considered when choosing between the weir or boot design. Because of the amount of weight that the

56. 53
heavy liquids account for within the inlet stream, a weir design was utilized because of its ability to handle the high
volume and weight of the styrene and ethyl benzene. Table 2 describes the parameters of the three phase separator:
Three Phase Separator Parameters
Hold-up Time 5 min
Surge Time 3 min
Terminal Velocity of Oil in Water 2.529 ft/sec
Terminal Velocity of Water in Oil 1.897 ft/sec
Diameter 12.047 ft
Total Length 24.095 ft
Length of Left of Weir 23.614 ft
Length of Right of Weir 2.844 ft
Weir Height 9.626 ft
Water Oil Interface Position 4.892 ft
Normal Oil Height to Right of Weir 4.813 ft
Cross Sectional Area Light Liquid 42.575 ft2
Cross Sectional Area Heavy Liquid 57.580 ft2
Table 6.8: Parameters of the Three Phase Separator
After calculating these parameters, understanding the terminal velocity of how fast the fluids travel into the
separator is important. The Three Phase Separator was designed to withstand a 3 minute surge time and a 5 minute
hold-up. A surge time is defined as when the outlet stream(s) clog and can no longer drain either the water or oils.
This can cause a rise in the oil and water interface ultimately causing the oil to overflow and making the right of the
weir reach the top of the separator. A hold-up unlike a surge, decreases in height while keeping the height of the oil
the same. The oil to the right of the weir decreases to within a foot of the bottomoil drain.
The designed three phase separator tank is a cylinder with two separate compartments, right of the weir and
left of the weir. The cylinder has a diameter of 12.047 ft with a total top of cylinder area of 115.178 ft2. The total
length of the weir is 24.095 ft and the length to the right and left of the weir was calculated to be 23.614 ft and 2.844
ft. The weir has a height of 9.626 ft with a water and oil interface position of 4.892 ft. The normal oil height of the
weir was calculated to be 4.813 ft almost totaling the water and oil interface position. As seen below in figures 2 and
3 describe the dimensions of the three phase separator and the weir within: