Sty-Guys

ChE 0613

Group 17

Authors:

Benjamin Longstreet

Chris Merz

Michael Thompson

Ari Ufondu

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

Table of ContentsExecutive 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.................................................................................................................................22

5. Column Sequencing.........................................................................................................................32

6. Other Unit Operations....................................................................................................................41

Heat Transfer Equipment...............................................................................................................41

Separation and Storage Tanks........................................................................................................48

Pumps...............................................................................................................................................54

7. Pinch and Exergy Analysis..............................................................................................................67

Pinch Analysis..................................................................................................................................67

“Traditional” Plant and Tabular HX Network.............................................................................72

Exergy Analysis...............................................................................................................................83

8. PDF and Stream Table of the plant with Traditional HX Network.............................................88

9. Process Equipment List.................................................................................................................109

10. Process Economics.......................................................................................................................111

Conclusion: Design and Simulation of the Plant.....................................................................................114

Nomenclature..........................................................................................................................................117

References...............................................................................................................................................119

Appendix A: Membrane Reactors...........................................................................................................121

Appendix B: Primary Column Redesign For 99.9%................................................................................123

Appendix C: Separating heavies (naphthalene/ethyl naphthalene) from high purity styrene in a single column.....................................................................................................................................................126

Appendix D: Water Remediation............................................................................................................128

Executive SummaryThe 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 steam to 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, steam acts 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 1425 oF 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 vacuum columns needed reducing the costs of plants.

1

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.

2

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 odor 1.

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 polymers 1.

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

3

Heat of Formation ∆HF Liquid(145°C) 103.4 kJ/mol

14Heat 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

4

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.

5

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 steam addition 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

6

mixed with this product stream and the mixture is fed to reactor two resulting in around a 65% conversion of the

ethylbenzene to styrene. Each of the reactors used in the adiabatic plant operate at near vacuum conditions 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 under vacuum 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 steam where

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.

7

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

8

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 styrene 9. 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

9

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 steam used 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) from the 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 (CO 2), 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 from the

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 mix with 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.

10

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 from naphthalene 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:

E B1∗0.21+E B2∗0.33=286.74 lbmol Styrenehr

Eqn

3.1

Where EB1 and EB2 are the flowrates of ethylbenzene feeds (lbmol/hr) going to reactors one and two respectively.

E B1∗0.21+(E B1−( E B1∗0.21 ))∗0.33=286.74 lbmol Styrenehr

Eqn

3.2.

E B1=609.18 lbmol Ethyl Benzenehr

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.

11

Selectivity=Conversion of Ethylbenzene¿Styrene ¿TotalConversion

=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 steam to ethylbenzene. The 3655.08 lbmol/hr

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 steam that is

exiting accounts for the overall amount of steam that was used throughout the process which totals 7505.08

lbmol/hr.

12

In conclusion, in order for this process to be successful in 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 steam are

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 steam being

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 stream of

ethylbenzene entering reactor one would be unused, resulting in a larger quantity of fresh ethylbenzene being

required.

Detailed DescriptionThe 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:

Fresh Feed Streams Molar Flow Rate Temperature Pressure Vapor Composition

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(lbmol/hr) (°F) (psia)Fraction

(%)

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 stream of ethyl benzene is added to the fresh stream of

ethyl benzene. The recycled stream was estimated to contain 262.28 lbmol/hr of ethyl benzene. Table 2 depicts the

actual recycled stream (EBRECYCLE):

SubstreamMolar 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).

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SpecificationsEthyl 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 from the

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, and

1% to ethyl-naphthalene and ethylene. The balanced reactions and stoichiometry are listed below.

Reactor 1

[1]C8 H10(ethyl benzene)=C8 H 8(styrene)+H 2(hydrogen) 21% conversion

15

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 steam from stream WATR2.

Reactor 2

[1]C8 H 10(ethyl benzene)=C8 H 8(styrene)+H 2(hydrogen) 33% conversion

[2] C8 H 10(ethyl benzene)+H 2(hydrogen)=C7 H 8( toluene)+C H 4(methane) 4%

[3]C8 H 10(ethyl benzne)=C6 H 6(benzene )+C2 H 4(ethylene ) 0.5%

[4]

C8 H 10(ethyl benzene)+2 H 2O(water )=C7 H 8(toluene)+3 H 2(hydrogen)+C O2(carbondioxide )

2.5%

[5]

C8 H10(ethylbenzene )+4 H 2O(water )=C6 H 6(benzene )+6 H2(hy drogen)+2C O2(carbondioxide) 3.5%

[6]2C8 H 10(ethyl benzene )=C10 H 8(naphthalene)+3 C2 H 4(ethylene) 1%

[7]2C8 H 10(ethylbenzene)=C12 H12(ethyl naphthalene)+2 C2 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

16

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 stream is separated

into vapor, liquid, and oil. The vapor stream leaving 3PSEP, referred to as 3PVAP, leaves at 85 oF 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 1425 oF. The reactions

occurring within the Fire Heater Reactor is seen below:

Fire Heater Reactor

[1]2 H2 (hydrogen )+O2(oxygen)=H2O(water )

[2] C2 H 4 (ethane)+3O2(oxygen)=2CO2(carbon dioxide)+ H 2O(water)

[3]CH 4 (methane )+3 O2(oxygen)=CO2(carbondioxide)+H2O(water )

[4]2 C6 H 6(benzene )+15O2(oxygen)=12 CO2(carbon dioxide)+6 H 2O(water)

[5]C7 H 8(toluene )+9O2(oxygen)=7CO2(carbondioxide)+4 H 2O(water )

17

FIRE

R1

R2

EBHX

3PHX

[6]C8 H 8(styrene)+10O2(oxygen)=8 CO2(carbondioxide)+4 H 2O(water)

[7]2C8 H 10(ethylbenzene)+21O2 (oxygen)=16CO2(carbon dioxide)+10 H2 O(water )

[8]C10 H 8 (naphthalene )+12(oxygen)=10 CO2(carbon dioxide)+4 H2 O(water )

[9]C12 H 12 (ethyl naphthalene )+15 O2(oxygen)=12 CO2(carbondioxide)+6 H 2O(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 Benzen

e Styrene

Light Key Recovery 0.9995 0.999 0.9 0.9995 0.999

Heavy Key StyreneEthylbenze

neNaphthale

ne TolueneNaphthale

neHeavy Key Recovery 0.1 0.001 0.01 0.001 0.001Condenser Pressure

(psia) 0.5801 14.7 0.5801 14.7 0.5801Reboiler Pressure

(psia) 0.5801 14.7 0.5801 14.7 0.5801Reflux Ratio -1.3 -1.35 -1.35 -1.35 -1.35

Table 3.5: Distillation Column Input Variables

18

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 stream to 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).

19

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 RSTLCOLStages 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 0.143078Distillate 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 stream input values are summarized in Table 7.

Radfrac Stream Inputs ROILS2COL REBHIPT RBENZTOL RSHIN RSH2INTemperature (F) 85 99.003 206.681 128.559 133.706Pressure (psia) 26 14.7 14.7 0.5801 0.5801Total Flow Rate (lbmol/hr) 624.611 337.29 48.424 287.321 32.991Ethylbenzene Flow Rate (lbmol/hr) 257.842 257.713 0.258 0.129 0.00542Styrene Flow Rate (lbmol/hr) 313.806 31.381 0.000155 282.425 28.243Benzene Flow Rate (lbmol/hr) 17.806 17.806 17.806 4.30E-26 0Toluene Flow Rate (lbmol/hr) 30.39 30.39 30.36 7.22E-13 3.46E-15Naphthalene Flow Rate (lbmol/hr) 2.383 5.92E-06 3.54E-11 2.383 2.359Ethylnaphthalene 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.

20

Specs TEMPR1 TEMPR2 COLDUT XSO2 CH4TRADType Stream-Var Stream-Var Stream-Var Mole-Flow Stream-VarVariable Temperature Temperature Temperature Excess Oxygen

Mole FractionTemperature

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°FTolerance 0.01 0.01 0.01 0.01 0.01Manipulated Variables

WATFEED WATSPLIT CLDUTIN AIRFEED CH4FEED

Convergence 6540.02 lbmol-water/hr

Fractions:0.576 to WATR10.434 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 stream was 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 of reactor 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 ofO2 OUT

O2∈−O 2OUT. By adjusting the AIRFEED stream, this was achieved with the controller.

The final stream flow was specified to 7272 lbmol/hr of air.

21

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

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

22

EBTank STHITankTOLTANK

BENTANK

3PSEP

PRIMARY

R1 & R2

FIRE

STHPROD

BTCOL

STLPRODEBRECOL

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:

23

Figure 4.1: Dehydrogenation of Ethyl Benzene

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%

24

3.5%

1%

1%

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 0.08654kmolEB

kgcat∗hr∗¿¿. Utilizing given values and equation (1) we

were able to determine the forward reaction constant:

k f=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.

k f=exp(0.851−( 90891( Jmol )

8.314( Jmol∗K )∗828 K ))=4.32006∗10−6 kmolEB

kgcat∗s∗¿¿

25

k f=4.32006∗10−6 kmolEB

kgcat∗s∗¿∗3600( shr )=0.015552232

kmolEB

kgcat∗hr∗¿ ¿¿

.

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

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=¿H 20+¿H2+¿ST−¿H 20−¿EB=8H 20+1H 2+1ST−8H 20−1EB=1 Eq. 4.3

yebo=¿of ethyl benzenemoles 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 x is

the conversion of ethyl benzene:

PEBo=ϵ∗PTOT Eq. 4.5

PEB=PEBo∗1−x

1+ϵx Eq.

4.6

PST=PEBo

1+ϵx∗x Eq.

4.7

PH2=PEBo

1+ϵx∗x Eq.

4.8

Finally, the rate of the reaction can be calculated using:

r=kf∗(PEB−PST∗PH 2

KEB) Eq.

4.9

26

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.

Figure 4.2. Levenspiel Plot

Figure 4.3: Bulk Catalyst Volume

27

0 0.05 0.1 0.15 0.2 0.25 0.3 0.35 0.4 0.45 0.50

500

1000

1500

2000

2500

3000

3500

4000

1/r vs x

x Conversion Fraction

1/r

Reactor 2

Reactor 1

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:

V cat=FA0∗∫ dx

rρparticle(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:

V cat=π∗d2

4∗(hcat) 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:

V cat=π∗d2

4∗(h−6 ) Eq. 4.12

Since h = d for an optimal reactor size this equation is written as:

V cat=π∗d2

4∗(d−6 ) Eq. 4.13

The volume of the reactors are given by:

V reactor 1=π∗r12∗L1=

π∗d12

4∗L=

π∗d13

4 Eq. 4.14

V reactor 2=π∗r22∗L2=

π∗d22

4∗L=

π∗d23

4 Eq. 4.15

28

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∗¿

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]:

1R

= 1R 1

+ 1R 2 Eq.

4.16

Rreac 1=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

29

Specification Reactor 1 Reactor 2

Catalyst Volume(ft3) 296.735 716.241

Diameter(ft) 9.877 12.164

Reactor Volume w/ Insulation (ft3) 756.685 1413.461

Catalyst Split 0.2929 0.7071

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:

QA

= ΔTR

=(1175−70 )∗A0.125252

+L2

1.18

=98,612 BTUhr Eq. 4.18

Lreact 1insulation=¿.001 ft

Lreac 2insultaion=0.0093

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):

30

Figure 4.4: Reactor One Cylindrical Dimensions

Figure 4.5: Reactor Two Cylindrical Dimensions

Using the Ergun equation ∆ P=150 μ 1−εDG

+ 74 ( 1−ε

ε3 )∗( LD )( 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

31

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.

450 650 850 1050 1250 1450 1650 1850 20500

20

40

60

80

100

120Keq vs Temperature (oF)

Temperature oF

Keq

Figure 4.6. Effect of Keq with increasing temperatures

0 2 4 6 8 10 12 14 160.4

0.41

0.42

0.43

0.44

0.45

0.46

0.47

0.48

0.49

0.5 Keq vs Pressure

Pressure (Psi)

Ke

q (

Bar

)

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

32

on the equilibrium constant. The Aspen simulation analysis displays that the conversion of styrene is inversely

proportional to pressure.

0 5 10 15 20 250.4

0.420.440.460.48

0.50.520.540.560.58

0.6

Keq vs Water mole

Water (Moles)

KEQ

Bar

Figure 4.8. Keq versus Moles of Water

Figure 4.8 shows that the equilibrium constant does not depend on the moles of water.

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 10

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1Effect on Conversion

Conversion for Temperature

Conversion for di-lution

Conversion for Pressure

% of final values

Conv

ersi

on

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.

33

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

34

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 distillateThe 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 streamUsed 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

35

Figure 5.2: Second Best Column Sequence

Rules Followed Rules Disobeyed

Recovered products as distillateThe bottom products 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

36

Product Streams Mass Flow Rates (lb/hr)

Streams Low-Purity Styrene

High-Purity Styrene Benzene Toluene Component

TotalH2O 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 3.3568Ethyl

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 Pur ity Styrene : 26473.5 lb /hr26489.6874 lb /hr

=99.93 wt % > the required 99.8 wt%

Eq. 5.1

37

High Purity Styrene Vol . %: 26473.5lb /hr29412.0761lb /hr

= 90 % of the total styrene volume

Eq. 5.2

Low Purity Styrene : 2938.56lb /hr2939.4347 lb /hr

=99.97 wt % > the required 97 wt%

Eq. 5.3

Low Purity StyreneVol . %: 2938.56 lb /hr29412.0716 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 stream produces a minimal amount of styrene accounting for ~ 0.01% of the total styrene.

Benzene : 1390.22lb /hr1393.017 lb /hr

=99.8 wt % > the required 99 wt%

Eq. 5.5

Toluene : 2794.57 lb /hr2822.6423 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 12Condenser P 30 torr 14.7 psia 14.7 psia 30 torr 30 torr

38

Pressure Drop (psi) 3.27348 3.84422 3.32234 1.20447 1.31586Diameter (feet) 21.286 4.646 1.503 8.736 17.844Spacing (feet) 2 2 2 2 2

Type Sieve Sieve Sieve Sieve SieveThickness (gauge) 10 10 10 10 10Hole Diameter (ft) 0.0417 0.0417 0.0417 0.0417 0.0417Weir Height (in) 2 2 2 2 2Weir Length (ft) 15.467 3.376 1.092 6.348 2.134Condenser Duty

(BTU/hr) -30420590.6 -4670852.4 -817530.647 -4893174.21 -571379.955

Condenser Type Water Water Water Water WaterRPRIMARY 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 TypeLow

Pressure-Steam

Low Pressure-

Steam

Low Pressure-

Steam

Low Pressure-

Steam

Low Pressure-

SteamSaturated 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 214.15 294.61 234.05 133.73 249.68Material 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 0.009267Actual Thickness

(in) 0.125 0.125 0.125 0.125 0.125

Tray Efficiency 100 100 100 100 100Table 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.

39

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 [( xb

x t )(x t

xb )]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

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−Nmin

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=RD−RDmin

RD+1

The Kirkbride equation is used to determine the optimum stage to set the feed:

logN enriching

N stripping=(0.206) log [( x toluene, f

xbenzene , f )WD ( xbenzene , 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

40

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:

( LD )

min+1=∑

1

n xD

(α−θ) /α 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= PRSE−0.6 P Eq.

5.11

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

41

with the tops from the high-purity styrene column and is designed to produce the purest form of styrene and

separate it from the lower quality styrene making this column require a significant amount of trays but is

narrow due to the small volume flow rate.

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 = 293 oF ,

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 from the lower quality styrene

making this column require a significant amount of trays but is narrow due to the small volume flow rate.

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 from or 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.

44

ffasdf

ffasdf ffasdf

Figure 6.2. Ethyl Benzene Heat Exchange Network (EBHX)

Figure 6.3. Ethyl Benzene Heat Exchange Network Inner Tube Layout

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-

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

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.

48

Heat

ExchangerLabor Cost Tube Costs

Material

Costs

Number of

Shells

Cost per

ShellTotal 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 23PHX 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 1762.2Table 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 steam at 1425°F can be calculated utilizing thermodynamics

Duty ( BTUhr )=m∗(Hout−H ¿) Eq.

6.1

Duty=117820 lbhr

∗(1762.2 BTUlbm

−38.8 BTUlbm )=2.03∗108 BTU

hr

Eq. 6.2

where m 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 ( f t3 )=Duty BTU

hr

76,800 BTUhr∗f t3

=2.03∗108( BTU

hr )76,800 BTU

hr∗f t3

=2644.844 f t 3

Eq. 6.3

49

Now that the volume has been calculated, the height and diameter of the fired heater can be calculated. Another equation for volume using the fired heater’s dimensions is assuming that diameter is equal to height

Volume ( f t3 )=(D¿¿cz❑¿¿2∗π∗H)4

=0.785∗Dcz2 ∗D ¿¿

Eq. 6.4

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

AreaTubes ( f t 2 )=Duty ( BTU

hr )Heat transfer rate

=2.03∗108 BTU

hr

14,000 BTUf t 2∗hr∗° F

=27709.88 f t2

Eq. 6.5

After the area of the tubes has been calculated, another equation representing the area of the tubes using the diameter of the fired heater and the number of tubes can be used.

AreaTube ( f t2 )=π∗(16)D∗N t=13854.94 f t 2 Eq.

6.6

N t=27709.88

D Eq.

6.7

The cross section of the tube sheet equation is given by:

0.087∗N t=0.785∗(D2−D cz2 ) Eq.

6.8

Solving these equations, Dcz, D, and Nt can be calculated. The answers to these are given in the following table.

Variable ValueDcz (ft) 13.457

D = H (ft) 18.61Nt (tubes) 1489

D with insulation (ft) 20.61Table 6.6: Calculated Dimensions

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 of the 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 maximum pressure 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

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 room temperatures

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

StyreneLo-Purity Styrene Heavies

CHEMICAL PROPERTIEST Boiling Pt. (Fº) 277.16 176.18 232.00 293.00 293.00 424.00

Ethyl Benzene Benzene Toluene Hi-Purity

StyreneLo-Purity Styrene Heavies

T Freeze Pt. (Fº) -139.00 41.90 -139 -23.8 -23.8 176.00T Flash Pt. (Fº) 72 11.07 43.00 88.00 88.00 174.00Comp. Vol (ft3) 116,515.59 4,445.90 9,851.76 84,467.74 9,374.04 1,991.43

Auto Ignition(°F) 914TANK PROPERTIES

Flow Rate (lb/hr) 36,733.92 1,393.016 2,822.641 26,486.7 2,939.44 677.382Volume (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

SteelHeight (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 13.079Thickness (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

ConeBlank Gas N2 N2 N2 O2-enriched N2 O2-enriched N2 N2

Concentration (ppm) --- --- --- 19.25 19.23 ---

Heating No Yes Yes No No YesCooling Duty

(BTU/hr) NA NA NA 288,737.7 67,464.02 NA

Mixer Type None Sidewall Sidewall Top Mount Top Mount NoneEthyl Benzene Toluene Hi-Purity Lo-Purity Heavies

52

Benzene Styrene StyreneExterior Color White White White White White White

Insulation None None None None None NoneTable 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 vacuum is 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 vacuum is 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 stream into vapor, organics, and free water.

This separator is useful in refining a material stream that 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 separator depends sole upon the stream that 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

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:

53

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 bottom oil 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 ft 2. 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:

54

Figure 6.7-1: Normal Operations within the Three Phase Separator

Figure 6.7-2: Normal Operations within the Three Phase Separator (Top Angle)

Figure 6.8-1: Surge Operation within the Three Phase Separator

55

Figure 6.8-2: Surge Operation within the Three Phase Separator (Top Angle)

56

Pumps

Pumps are essential piece of equipment necessary for plants to operate. The main purpose of a pump is to

move and control fluids within a piping network. Pumps are needed in two different situations, when there is a

height that the fluid must overcome to reach its destination and when the fluid needs pressurized. When the fluid

experiences a change in height, the fluid experiences difficulties from both frictional forces and head losses.

Controlling a fluid consists of pressurizing the fluid in terms of what is necessary for the plant specifications. The

way that pressurization is handled is through mechanical energy. The fluid going into the pump is known as the

suction side and the fluid going out is known as the discharge side. Using pistons or impellers, the suction side is

pressurized and released to the discharge side.

There are many different classifications of pumps that can be used in the styrene plant facility. These range

from dynamic, special effect pumps, and centrifugal. Dynamic pumps use displacement action in order to get the

fluid to its proper pressure. Special effect pumps use a multitude of different actions including jet, gas lift, hydraulic

ram, and electromagnetic. The centrifugal pumps use a rotating impeller to pressurize the fluids and force them

outward by. The centrifugal pumps are the most optimal choice for our pumping necessities in the styrene plant

because they are simple, cheap, consistent, uniform in pressure, have a relatively low head, can tolerate pressures up

to 1000°F, and discharge lines can be equipped with valves to shut off if necessary. There are two different types of

centrifugal pumps, volute and diffuser. The volute type is the one that will be utilized in the traditional styrene plant.

The volute draws in fluid into the suction eye where impellers rotate the fluid. The fluid gains centripetal velocity

and then is released through the discharge side where the centripetal velocity is transformed into a pressure head.

The volute centrifugal pump can be seen in figure 1.

Figure 6.9: Volute Centrifugal Pump

57

Using aspen, the traditional styrene plant only required two pumps for simulation. This is an

inaccurate of what the actual facility will require because the dimensions and locations of the storage tanks, utilities,

separation columns, etc. were not incorporated. In reality, a fully functional styrene plant will have many more

pumps. These pumps include the ethyl benzene feed pump, the water feed pump, the methane feed pump, the utility

water pump, the air feed pump, product pumps (high purity styrene, low purity styrene, toluene, benzene, and

heavies), ethyl benzene recycle pump, and many others. It is important to note that each separation column also has

a pump associated with the reboiler. Also, if a fluid is elevated high enough above its destination and the hydro

static pressure head is sufficient, a pump will not be required. All of these pumps will not be investigated in this

report, however, four pumps that will be investigated are the high purity styrene pump (HIPURSTYPUMP), the

toluene pump (TOLPRODPUMP), the heavies pump (HEAVYPUMP), and the ethyl benzene-benzene-toluene

pump (EBTPMP). The HIPURSTYPUMP is designed to pump high purity styrene from the condenser unit of

STHICOL distillation column into the high purity styrene storage tank. The toluene pump is designed to pump

liquid toluene bottoms product from the BTCOL distillation column and into the toluene storage tank. The

HEAVYPUMP is designed to pump all of the bottoms (both naphthalene and ethyl-naphthalene) from the

STYLOCOL column and into the heavies’ tank for disposal. The fourth pump, EBTPUMP, is located immediately

after the primary separation column and condenser. This one is necessary because the stream leaving the primary

column (PRIMARY) is under vacuum conditions and needs to be pressurized. The stream is then sent to another

column, EBRECOL, where ethyl benzene and benzene/toluene will be separated.

The toluene product pump, TOLPRODPUMP, is responsible for taking the toluene from the BTCOL

column as a bottom and pumping it to the toluene storage tank. The piping network of the suction side consists of 1

globe valve and 2 fittings. The suction network is located 0.083’ above the pump level and the length of the piping is

10.33’. The discharge piping network consists of 1 globe valve and 7 fittings. The discharge network is 20’ above

the pump and is 130.25’ in length. Note that the piping network cartoon does not depict a 3-D view of the pipes thus

some of the fittings are depicted on top of one another. The primary assumption made was that the toluene storage

tank was considered full. Whenever the stream reaches the tank, it is also assumed that the outside temperatures will

bring the tank temperature to ambient temperatures.

58

Figure 6.10: Toluene Piping and Pump Network (AVEVA)

Toluene Pump Parameters Units Values

Suction Pressure Psia 14.7

Discharge Pressure Psia 23.599

Manufacture Goulds Pumps Inc.

Material of Construction Steel 100-938 54153

RPM Revolutions per minute 1750

Impeller Diameter Inches 4

Efficiency 0.45

Volumetric Flow Rate Ft3/hr 58.64

HP Requirement Hp 0.09278

Units Values

59

BTCOL

Toluene Tank

TOLPRODPUMP

NPSH Required Ft 1.5

NPSH Actual Ft 19.46

Cavitation No

Pipe Diameter Inches 0.95

Pipe Length Ft 140.583

Friction on Suction Side 0.623

Friction on Discharge Side 5.743

Suction Temp Fº 234.162

Suction Mass Flow Lb/hr 2822.64

Discharge Temp Fº 234.162

Discharge Mass Flow Lb/hr 2822.64

Table 6.9: Toluene Pump Parameters

Figure 6.11: Overhead View of Toluene Piping and Pump Network

60

The next pump designed was the heavies pump which was responsible for pumping the heavies liquid from

the bottom of the STYCOL up and into the heavies storage tank, 18.167 feet above the pump level. The bottom feed

allows the assumption of a liquid stream but is important to note that the column operates under vacuum, which

needed to be accounted for in the design of the pump. The pressure assumed was calculated to total 22.35 psi

assuming that the storage tank was considered full.

Figure 6.12: Heavies piping and pump diagram (AVEVA)

61

Heavy Tank

HVYPUMP

Figure 6.13: Overhead View of Heavies Piping and Pump Network

Heavies Pump Parameters Units Values

Suction Pressure Psia 14.7

Discharge Pressure Psia 22.35

Manufacture Goulds Pumps Inc.

Material of Construction Steel 100-938 54153

RPM Revolutions per minute 1750

Impeller Diameter Inches 5

Efficiency 0.55

Volumetric Flow Rate Ft3/hr 11.85

HP Requirement Hp 0.0132

NPSH Required Ft 1.5

NPSH Actual Ft 17.892

Cavitation No

Pipe Diameter Inches 0.5

62

Units Values

Pipe Length Ft 48.667

Friction on Suction Side 0.2746

Friction on Discharge Side 1.7058

Suction Temp Fº 234.162

Suction Mass Flow Lb/hr 2822.64

Discharge Temp Fº 234.162

Discharge Mass Flow Lb/hr 2822.64

Table 6.10: Heavies Pump Parameters

The EBTPMP was responsible for taking the distillate stream of PRIMARY, which exits the column at 30

torr, to atmospheric conditions, which the EBRECOL operates. Several assumptions were made in the design of this

pump and piping network. It was assumed that the pressure at the outlet of the condensate storage tank of

PRIMARY to be 30 torr, the stream to be a liquid composition at 70 °F, to simulate an uninsulated piping network

and the condensate storage tank to be nearly empty. The EBTPMP is also responsible for pumping this stream to the

EBRECOL feed tray height. This is 33 ft above the outlet for the condensate storage tank. The suction side of the

network is positioned 6 inches above the pump, is 29 ft in length and consists of one globe valve and two 90 o

elbows. The discharge side is 33.5 ft above the pump and consists of one globe valve and four 90o elbows.

63

Figure 6.14: EBENTOL piping and pump diagram (AVEVA)

64

Inlet Parameters:

Pressure: 0.5801 psia

Temperature: 70 oF

Mass Flow: 34820.2 lb/hr

Outlet Parameters:

Pressure: 14.7 psia

Temperature: 70 oF

Mass Flow: 34820.2 lb/hr

PRIMARY

EBRECOL

EBTPUMP

Figure 6.15: Overhead View of EBENTOL Piping and Pump Network

EBTPMP Parameters Units Values

Suction Pressure Psi 0.5801

Discharge Pressure Psi 27.525

Manufacture Goulds Pumps Inc

Material of Construction Steel 100-938 54153

65

RPM Rounds per minute 3500

Impeller Diameter Inches 4

Efficiency 0.509

Volumetric Flow Rate Ft3/hr 654.941

HP Requirement Hp 2.52

NPSH Required Ft 0.5

NPSH Actual Ft 3.842

Cavitation No

Pipe Diameter Inches 3

Pipe length Ft 94

Friction on Suction Side Psi 0.192

Friction on Discharge Side Psi 0.454

Suction Temp Fº 70

Suction Mass Flow Lb/hr 34820.2

Discharge Temp Fº 70

Discharge Mass Flow Lb/hr 34820.2

Table 6.11: EBTPMP Pump Parameters

As mentioned before, the high purity styrene pump is responsible for pumping the high purity styrene from

the condenser unit STHICOL to the high purity styrene storage tank. The suction side consists of 1 globe valve and

6 fittings. The STHICOL condenser is elevated to a total height of 24.25 ft. above the pump and 54 ft. in length.

However, the discharge side consists of 1 globe valve and four fittings. It is also 47 ft. in height and 13.17 ft. in

length. Again, it is important to note that the schematics are given from a bird’s eye view. That is to say that the

vertical sections of the piping network are not explicitly shown. A table showing a summary of the notable

parameters of the pumping network is shown below.

66

Figure 6.16: High purity styrene piping and pump diagram (AVEVA)

STHIPMP Pump Parameters Units Values

Suction Pressure Psia 7.36

Discharge Pressure Psia 17.53

Manufacture Goulds Pumps Inc.

Material of Construction Steel 100-938 54153

RPM Revolutions per minute 3500

Impeller Diameter Inches 4

Efficiency 0.5

Units Values

Volumetric Flow Rate Ft3/hr 502.83

HP Requirement Hp 0.09278

NPSH Required Ft 3.2

NPSH Actual Ft 43.01

67

STHIPUMP

STHI Tank

Cavitation No

Pipe Diameter Inches 2.5

Pipe Length Ft 138.42

Friction on Suction Side 3.9896

Friction on Discharge Side 0.667

Suction Temp Fº 127.994

Suction Mass Flow Lb/hr 26489.7

Discharge Temp Fº 127.994

Discharge Mass Flow Lb/hr 26489.7

Table 6.12: STHIPMP Pump Parameters

68

Figure 6.17: Overhead View of STHIPMP Piping and Pump Network

69

7. Pinch and Exergy AnalysisPinch Analysis

Pinch analysis is a systematic process improvement approach for minimizing energy consumption

of chemical processes. It is achieved by setting thermodynamically feasible energy targets and reaching them by

optimizing heat recovery systems, energy supply methods and process operating conditions. Pinch analysis, when

applied to the design of heat exchangers in a plant, is an improvement on the traditional plants heat exchanging

network where heat exchange is not necessarily maximized nor utilities minimized. The pinch analysis removes

redundancies and inefficiencies associated with a traditional heat exchange network, making it more cost effective

with more desirable options even though it is more strenuous to design.

Ironically enough, before pinch analysis can be done on a heat exchange network, it is common practice to design an

“all utility plant”. An all utility plant is the worst possible scenario for a heat exchange network because it utilizes

utilities to reach its goals of all heating and cooling needs. This is expensive and inefficient, but it is easy to program

and gives you a good idea of what your heat exchange should not resemble when you complete a pinch analysis.

When conducting pinch analysis, the data for all the streams in the plant are combined to give two composite curves,

one for all hot streams and one for all cold streams. The composite curves are presented on a graph of temperature

(oF) versus heat load (BTU/hr) when doing the graphical method of pinch analysis. The point of closest approach

between the hot and cold composite curves is the pinch point with a hot stream pinch temperature and a cold stream

pinch temperature. The difference between the hot and cold stream can theoretically approach zero, but we limited

ourselves to an approach of about 10oF in the case that the heat exchange network become impractically large. Once

a pinch point has been established, the process can only be heated with heating utilities at temperatures above the

pinch and cooled with cooling utilities at temperatures below the pinch. Such a methodology will maximize the heat

recovery in the process with the establishment of a heat exchanger network based on pinch analysis principles. This

pinch point is typically where the design is most constrained, so by finding this point and starting the design at the

energy target, which is the minimum theoretical energy demand for the overall process, can be easily reached.

Figure 1 demonstrates the graphical pinch analysis of the styrene plant. Generating composite streams of

the hot (with and without methane) and cold streams, the amount of heating and cooling utilities can be calculated.

The difference between the methane and no methane streams shows how much heating duty will be required for the

styrene plant valued at 8.2901*107 BTU/hr. The remaining portions of the plant which require heating utilities will

utilize existing streams to accomplish this ultimately saving the styrene plant money. The cooling duty can be

observed by the amount of excess hot stream that requires cooling and does not have a cold stream to assist in this

cooling. The amount of cooling utility that the styrene plant will require is 1.2207*10 8 BTU/hr. Based on the

graphical pinch analysis, 252.68 lbmol/hr of methane is required for the styrene plant. The styrene plant will also

require 125,545.681 lb/hr of cooling water to account for the cooling duty needed for the facilty.

70

Figure 7.1: Graphical Pinch

An exact value of the pinch point temperatures can be observed below in figure 2. The closest these

streams can get to one another is 10°F. By shifting the cold composite stream by 1.2207*108 BTU/hr, the

temperatures were found to be 233.8°F and 223.8°F.

Figure 7.2: Pinch Point

71

The tabular pinch method provides similar results to the graphical pinch analysis and can be seen below in table 1. The results from the tabular method are as follows: the heating duty required for the styrene plant is 8.29*10 7 and the cooling duty required is 1.22*108. The pinch point is observed at 228.8°F which needs 5°F added and subtracted to observe the true pinch point temperatures of 233.8°F and 223.8°F.

Temperature

Delta T Add/Sub CP Delta H S/D Cascade Cascade, New

3366.403 0 S 0.00E+00 8.29E+07

1430 1936.403 -4 -2.00E+04 -3.88E+07 S 3.88E+07 1.22E+08

1068.03257 361.96743 7,-4 3.93E+04 1.42E+07 S 2.45E+07 1.07E+08

970.000018 98.032552 7,-4,-3 -5.49E+04 -5.38E+06 S 2.99E+07 1.13E+08

325.550951 644.449067 10,7,-4,-3 -1.89E+04 -1.22E+07 D 4.21E+07 1.25E+08

323.894055 1.656896 9,7,-4,-3 5.61E+06 9.30E+06 S 3.28E+07 1.16E+08

251.107661 72.786394 7,8,-4,-3 -2.25E+04 -1.63E+06 S 3.44E+07 1.17E+08

250.107661 1 8,6,-4,-3 1.16E+08 1.16E+08 D -8.16E+07 1.31E+06

245 5.107661 8,5,-4,-3 4.65E+04 2.37E+05 S -8.18E+07 1.08E+06

228.828105 16.171895 8,5,-3 6.65E+04 1.08E+06 S -8.29E+07 0.00E+00

187.863 40.965105 8,5,-2 -2.74E+06 -1.12E+08 S 2.95E+07 1.12E+08

177.07659 10.78641 8,5,-1 -6.62E+04 -7.14E+05 D 3.02E+07 1.13E+08

79.9956741 97.0809159 5,-1 -9.87E+04 -9.58E+06 D 3.98E+07 1.23E+08

75.0259561 4.969718 5 1.28E+05 6.38E+05 S 3.92E+07 1.22E+08

Table 7.1: Tabular Pinch

From table 2 below, it can be seen that the graphical and tabular methods both produce the same exact values for the heating/cooling utilities required. The amount of methane in both cases is 4055.5 lb/hr for the heating utility and the cooling water is at 125.5*106 lb/hr.

System CH4 lbmol/hr

CH4 lb/hr Heating Duty for

CH4 BTU/hr

Cooling Duty

BTU/hr

Cooling Water

Flow Rate 106 lb/hr

Heat Flowing Across Pinch

MMBTU/hr

Pinch Point

°F

Graphical 252.68 4055.5 8.29*107 1.2207*108 125.545 0 233.8/223.8Tabular 252.68 4055.5 8.29*107 1.2207*108 125.545 0 233.8/223.8

Table 7.2: Comparison of Graphical and Tabular Pinch

72

Heating Duty

Pinch Point

There are multiple benefits from utilizing the graphical pinch analysis method. The foremost being in the capital

costs of the facility. The heat exchangers used in the heat exchange network have a smaller area and thus cost less to

purchase. The graphical pinch analysis is also easier for an engineer to visually see the pinch point and analyze the

heating and cooling utilities required for the facility. The tabular method, compared to the graphical analysis, has

superior benefits. The heat exchanger network requires significantly less heat exchangers, mixers, and separators

which will reduce capital costs even though the heat exchangers will have to be of larger area. The tabular method is

more simplistic and faster to calculate producing highly accurate values. The tabular method is easier to

program/code than the graphical method which is why this method is an optimal choice when evaluating pinch

analysis.

When determining the heating and cooling utilities an all utilities plant was utilized and can be seen in

figure 3. The all utility plant uses all capital costs to reach the heating and cooling goals required for the styrene

plant. This is not a feasible approach because it wastes large amounts of energy increasing the cost to operate the

plant. However, it is useful when determining the optimum heat exchange network and the amount of utilities the

styrene facility must purchase to complete operations.

“Traditional” Plant and Tabular HX Network

In this section a traditional plant that incorporates process heat exchange is presented. A traditional plant

utilizes existing hot and cold streams to complete the required heating and cooling in the heat exchangers network

present within the facility while maintaining the first and second laws of thermodynamics. This simple heat

exchange network, found in figure 1, includes both process heat exchangers as well as utility heat exchangers. There

are in total three heat exchangers throughout the traditional plant. The fourth heat exchanger from the all utility

plant, used for cooling of stack gases, was removed so that the thermal energy of the stack gases can be utilized

elsewhere. The strategy for making the traditional styrene plant follows simple and easy guidelines. In general, by

rerouting hot streams to heat cold streams, and cold streams to cool hot streams, the amount of utilities required to

reach target temperatures is reduced. This, however, is not enough to make a thermodynamically feasible plant. The

hot streams that are paired with the streams that need cooling and vice versa must satisfy the second law of

thermodynamics. Based on this, the hot stack gases in the traditional plant are sent to the water feed heat exchanger.

The hot product stream leaving reactor two is sent to the ethyl benzene feed heat exchanger. The thermal heat from

the stack gases and reactor products are not sufficient in heating their respective feeds to their respective

temperature, therefore, a heating utility will be required for these two process heat exchangers. A utility heat

exchanger will be utilized to cool the hot products leaving the ethyl benzene heat exchanger. It is important to note

that the traditional plant is still not at optimum efficiency because it does not utilize these hot and cold streams to

their full potential. The reason this less-than-optimal plant is acceptable is due to the fact that it is simple, easy to

build and it represents how plants can be built without regard to overall heat transfer efficiency. The performance of

the traditional heat exchange network can be analyzed by using the information provided in progress report three.

Using the cooling and heating utilities required from the all utility plant and comparing them to the utilities required

in the traditional plant will produce a performance rating based off the all utilities plant.

73

Tables 1, 2 and 3 display the data within the traditional plant for the streams going into and out of the heat

exchangers WHX, EBHX and 3PHX respectively.

Substream:

MIXEDWATHIP WATHIPT HOTGAS STACK

Mass Fraction

H2O 1 1 0.112682 0.112682

EB 0 0 0 0

STYRENE 0 0 0 0

H2 0 0 0 0

BENZENE 0 0 0 0

TOLUENE 0 0 0 0

METHANE 0 0 0 0

C2H4 0 0 0 0

CO2 0 0 0.114666 0.114666

N2 0 0 0.72841 0.72841

O2 0 0 0.044242 0.044242

NAP 0 0 0 0

ENAP 0 0 0 0

Total Flow

(lbmol/hr)6540.010 6540.010 8007.926 8007.926

Total Flow (lb/hr) 117820 117820 220939 220939

Total Flow

(cuft/hr)2212.846 5085220 21526800 4142460

Temperature oF 70.026 1425 3222.133 249.445

Pressure (psia) 26 26 14.7 14.7

Vapor Fraction 0 1 1 1

Liquid Fraction 1 0 0 0

Solid Fraction 0 0 0 0

Enthalpy

(BTU/lbmol)-123840 -91954.70 -2799.32 -28841.20

Enthalpy (BTU/lb) -6874.260 -5104.260 -101.462 -1045.350

Enthalpy (BTU/hr) -8.1E+08 -6.0E+08 -2.2E+07 -2.3E+08

Entropy

(BTU/lbmol-oR)-40.375 -0.807 15.844 1.956

Entropy -2.241 -0.0448 0.574 0.0709

74

(BTU/lb-oR)

Substream:

MIXEDWATHIP WATHIPT HOTGAS STACK

Density

(lbmol/cuft)2.955 0.00129 0.000372 0.00193

Density (lb/cuft) 53.244 0.0231 0.0103 0.0533

Average MW 18.015 18.015 27.590 27.590

Liq Vol 60 oF

(cuft/hr)1890.934 1890.934 6084.097 6084.097

**VAPOR PHASE**

CPMX 10.00778 9.675824 7.418078

MUMX 0.039724 0.070457 0.02143

Average MW 18.01528 27.59002 27.59002

KMX 0.059462 0.077318 0.017625

**LIQUID PHASE**

CPMX

(BTU/lbmol-oR)19.47208

MUMX (cP) 0.62082

Average MW 18.01528

KMX

(BTU-ft/hr-sqft-R)0.347113

Table 7.3: Water Heat Exchanger (WHX) Stream Table

Substream:

MIXEDEBHIP EBHIPT R2OUT R2PROD

Mass Fraction

H2O 0 0 0.627501 0.627501

EB 0.951423 0.951423 0.148978 0.148978

STYRENE 0.048535 0.048535 0.17743 0.17743

H2 0 0 0.004364 0.004364

BENZENE 4.18E-13 4.18E-13 0.008045 0.008045

TOLUENE 4.16E-05 4.16E-05 0.015436 0.015436

METHANE 0 0 0.001652 0.001652

C2H4 0 0 0.002167 0.002167

CO2 0 0 0.010765 0.010765

75

N2 0 0 0 0

Substream:

MIXEDEBHIP EBHIPT R2OUT R2PROD

O2 0 0 0 0

NAP 1.13E-08 1.13E-08 0.00165 0.00165

ENAP 1.37E-08 1.37E-08 0.002011 0.002011

Total Flow (lb/hr) 67340.07 67340.07 185160 185160

Temperature (oF) 172.0766 965 1073.029 703.4774

Pressure (psia) 26 26 26 26

Vapor Fraction 0 1 1 1

Liquid Fraction 1 0 0 0

Solid Fraction 0 0 0 0

Enthalpy

(BTU/lbmol)1824.452 60706.49 -74459.4 -79404.9

Enthalpy (Btu/lb) 17.20093 572.3405 -3039.76 -3241.66

Enthalpy (Btu/hr) 1158310 38541500 -5.6E+08 -6E+08

Entropy

(Btu/lbmol-oR)-96.2935 -33.1109 -2.1089 -5.79298

Entropy

(Btu/lb-oR)-0.90786 -0.31217 -0.08609 -0.2365

Density

(lbmol/cuft)0.476743 0.001714 0.001583 0.00209

Density (lb/cuft) 50.5667 0.181828 0.038764 0.051188

Average MW 106.0671 106.0671 24.49513 24.49513

Liq Vol 60oF

(cuft/hr)1237.578 1237.578 3459.96 3459.96

**VAPOR PHASE**

CPMX

(BTU/lbmol-oR)66.67001 14.00602 12.7239

MUMX (cP) 0.015773 0.030093 0.022583

Average MW 106.0671 24.49513 24.49513

KMX

(BTU-ft/hr-sqft-R)0.037895 0.048971 0.033641

**LIQUID PHASE**

76

CPMX

(BTU/lbmol-oR)46.58015

MUMX (cP) 0.429738

Substream:

MIXEDEBHIP EBHIPT R2OUT R2PROD

Average MW 106.0671

KMX

(BTU-ft/hr-sqft-R)0.067408

Table 7.4: Ethyl Benzene Heat Exchanger (EBHX) Stream Table

Substream:

MIXEDR2PROD 3PH CLDUTIN CLDUTOUT

Mass Fraction

H2O 0.627501 0.627501 1 1

EB 0.148978 0.148978 0 0

STYRENE 0.17743 0.17743 0 0

H2 0.004364 0.004364 0 0

BENZENE 0.008045 0.008045 0 0

TOLUENE 0.015436 0.015436 0 0

METHANE 0.001652 0.001652 0 0

C2H4 0.002167 0.002167 0 0

CO2 0.010765 0.010765 0 0

N2 0 0 0 0

O2 0 0 0 0

NAP 0.00165 0.00165 0 0

ENAP 0.002011 0.002011 0 0

Total Flow (lb/hr) 185160 185160 3427730 3427730

Temperature (oF) 703.4774 85 70 119.9996

Pressure (psia) 26 26 14.7 14.7

Vapor Fraction 1 0.064208 0 0

Liquid Fraction 0 0.935792 1 1

Solid Fraction 0 0 0 0

Enthalpy

(BTU/lbmol)-79404.9 -103900 -123840 -122870

Enthalpy

(BTU/lb)-3241.66 -4241.77 -6874.32 -6820.3

77

Enthalpy

(BTU/hr)-6E+08 -7.9E+08 -2.4E+10 -2.3E+10

Entropy

(BTU/lbmol-R)-5.79298 -40.2363 -40.3755 -38.6196

Substream:

MIXEDR2PROD 3PH CLDUTIN CLDUTOUT

Entropy

(BTU/lb-R)-0.2365 -1.64262 -2.24118 -2.14372

Density

(lbmol/cuft)0.00209 0.06735 2.955475 2.900031

Density (lb/cuft) 0.051188 1.649751 53.24371 52.24487

Average MW 24.49513 24.49513 18.01528 18.01528

Liq Vol 60 oF

(cuft/hr)3459.96 3459.96 55012.81 55012.81

**VAPOR PHASE**

CPMX

(BTU/lbmol-R)12.7239 7.479035

MUMX (cP) 0.022583 0.012001

Average MW 24.49513 8.240705

KMX

(BTU-ft/hr-sqft-R)0.033641 0.07274

**LIQUID PHASE**

CPMX

(BTU/lbmol-R)19.89786 19.47247 19.47741

MUMX (cP) 0.620855 0.500208

Average MW 25.61041 18.01528 18.01528

KMX

(BTU-ft/hr-sqft-R)0.347101 0.367605

Table 7.5: Three Phase Heat Exchanger (3PHX) Stream Table

When evaluating the heat exchangers WHX, EBHX, and 3PHX in accord with the pinch temperatures found in

progress report 3, it is clear that the pinch points do not line up vertically. This means that there is heat flowing

across the pinch point of 233.8°F and 223.8°F. In table 4 below, the heat flowing across the pinch point can be

found. The total amount of heat flowing across the pinch point is 6.3119*107 BTU/hr. The following figures and

table show the heat flowing across the pinch.

78

Figure 7.3: 3PHX Heat Exchange Profile

Figure 7.4: WHX Heat Exchange Profile

79

Figure 7.5: EBHX Heat Exchange Profile

Heat Exchanger Heat Flow Across Pinch (BTU/hr)

WHX 1.9753*107

EBHX 1.5969*106

3PHX 4.1770*107

TOTALS 6.3119*107

Table 7.6: ASPEN-Traditional Plant Heat flow across the pinch (233.8°F/223.8°F)

Heat Exchangers Streams Matched Temperature (oF) Cooling Needed (BTU/hr)

1 Stream 2 Split

2A , 8 232.077 1.274*106

2 2B , 5 226.957 9.792*107

3 3 , 5 262.801 -

4 4 , 7 1804.445 3.311*105

5 4 , 6 245.462 -

6 3 , 6 698.362 -

7 4 , 8 320.528 -

8 3 , 10 826.655 -

9 3 , 9 732.992 -

10 3 , 8 700.180 -Table 7.7: Tabular Heat Exchanger Network Stream Matchups

80

A heat exchange network design was needed to ensure that zero amount of heat crossed the pinch point and that

approaches in each heat exchanger was no less than the pinch itself. Consulting the data obtained from the

traditional plant simulation, the ‘Dumbbell Method’ of designing a heat exchange network could be applied as seen

in figure 2. The dumbbell method show the cross-linking between streams in the heating and cooling of one another.

Figure 7.6: Tabular Heat Exchanger Network Stream Matchups

The creation of a traditional plant heat exchange network follow a basic thermodynamic CP based rule. Streams

that are to the right of the pinch point have a heat exchange network design to cool the hot streams, while streams to

the left of the pinch point were conversely designed to heat the cold streams to their temperature targets. The CP rule

states the following: Streams below the pinch point are only able to be paired if CP in < CPout or CPcold < CPhot; Streams

above the pinch point are only capable of pairing if CPout < CPin or CPhot < CPcold. These two guidelines ensure that the

heat loss from each hot stream is used by a corresponding cool stream which is critical in maintaining the plants

heating and cooling targets while utilizing a minimum amount of heat exchangers. This method is used to reveal the

amount of heat that crosses the pinch point and the number of heat exchangers needed for the network design

ensuring that the first and second laws of thermodynamics are not violated within the plant. Figure 3 describes each

stream and its corresponding stream that it is cross-linked with.

81

HX7HX9

HX4

HX2

HX1HX10

HX3HX6

HX5HX8

Figure 7.7: Individual Heat Exchanger Profiles

The traditional plant heat exchange network was also done utilizing a graphical technique. The graphical

method proved to be very similar to the tabular heat exchanger network providing similar cooling duties, however

this did not prove to be more efficient because it requires more heat exchangers. The following figures 4.8 and 4.9

depict the graphical method of heat exchanger networking.

82

0 50000000 100000000 150000000 200000000 250000000 300000000 350000000 4000000000

500

1000

1500

2000

2500

3000

3500Graphical Composition Diagram

Cold Composite

Hot Composite

Duty (BTU/hr)

Tem

pera

tre

°F

230.08°C

276.36°C

946.87°C

985.48°C1040.11°C

2571.26°C

226.46 °C

220.52°C 224.79 °C

379.85°C

Figure 7.8: Graphical Composition Diagram

83

Figure 7.9: Graphical “Dumbbell” Heat Exchanger Network

Table 4-6, shows the comparison between the graphical and tabular pinch methods in all utility plant, the

tabular heat exchange network, graphical heat exchange network and the traditional plant network. According to the

results, the traditional plant is not a suitable method to use when creating a heat exchange network. In fact, this

method proves worse than the all utility plant resulting in the heating and cooling duty required raised by a

significant amount. This means a traditional plant would require more expenses and thus is an insufficient way to

produce a heat exchanger network. The tabular heat exchanger and graphical heat exchanger networks resulted in

very similar results, however, the tabular method requires only 10 heat exchangers while the graphical requires 23

heat exchangers. This means that the graphical method will be much more expensive. The best method to use when

creating a heat exchanger network within the styrene plant is to utilize the tabular pinch analysis. It requires less

utility than the all the utility plant, less heat exchangers than graphical pinch analysis, and does not have heat

flowing across the pinch like the traditional heat exchanger network.

SystemCH4

lbmol/hrCH4 lb/hr

Heating Duty for CH4BTU/hr

Cooling Duty MM-

BTU/hr

Cooling Water Flow Rate lb/hr

Heat Flowing Across Pinch Point

MM-BTU/hr

Pinch Point°F

(High/Low)Graphical

Pinched Plant252.68 4055.5 8.29*107 1.2207*108 2.259*106 0 233.8 / 223.8

Tabular Pinched Plant

252.68 4055.5 8.29*107 1.2207*108 2.259*106 0 233.8 / 223.8

Tabular HX Network

252.68 4055.5 8.29*107 1.2241*108 2.290*106 0 233.8 / 223.8

Graphical HX Network

252.68 4055.5 8.29*107 1.2242*108 2.290*106 0 233.8/223.8

Trad. Plant Network

445 7120 1.46*108 1.852*108 3.428*106 6.319*107

Utility: 84.99/70Process:

1073.03/965Table 7.8: Comparison Table

Exergy Analysis

The first law of thermodynamics explores the concept of conservation of energy. This law states that

energy can be transferred between materials, namely thermal energy. The idea that these materials, before they come

into contact, hold a useful amount of energy is known as exergy. Exergy is the amount of energy available to be

used elsewhere as work. It is important to realize that the thermal energy being transferred from one source to

another cannot be recovered. This is because exergy is a more accurate representation of entropy. In our case, it is

the excess energy found in exiting streams of the styrene plant facility. The stack gases and the product streams from

reactor two have a large amount of thermal energy which could be utilized somewhere else in the styrene facility. In

the previous section, we developed heat exchange networks based off of this useful energy.

84

Simply put, exergy is the useful work available from a process stream. It is defined as the integral of the

ratio of temperature subtracted by initial temperature and temperature across the changes in heat or duty. Its formula

is represented by equation 1.

Exergy=∫ T−¿T

dQ Eqn 7.1

The Carnot cycle, which is the model for the most efficient means of transforming heat into work, is used

to derive the equation for exergy mentioned above. That being said, it is an idealized scenario. It also requires the

temperatures to be in the absolute scale of Rankine (oR). While rather straight forward, exergy analysis is usually

done in conjunction with data from a composite curve. When the Y-axis of the composite curve is transformed into

the exergetic temperature scale, with the X- axis remaining as the change in duty, a series of curves joined at

discontinuities is produced. The exergy of the hot stream is the area under its corresponding exergy curve, while the

exergy of the cold stream is the area under its exergy curve.

When the heat exchange network is vertical, the area between the hot exergy curve and the cold exergy

curve represents the work lost to the heat exchange network, or exergy loss. Exergy loss is the amount of useful

work that cannot be extracted from the process stream. Its analysis is particularly important to power plants, which

derive their profit from maximizing the extraction of available work. The analysis of exergy loss is also important to

our styrene plant because its minimization will increase efficiency in our plant, resulting in increased profits.

0 50000000 100000000 150000000 200000000 2500000000

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

WHX EXERGY

HotCold

DUTY (BTU/hr)

(T-T

o)/T

Figure 7.10. Water Heat Exchanger Exergy Analysis

85

Exergy Loss = 7.76E+07 BTU/hr

0 5000000 10000000 15000000 20000000 25000000 30000000 35000000 400000000

0.1

0.2

0.3

0.4

0.5

0.6

0.7

EBHX EXERGY

COLDHOT

DUTY (BTU/hr)

(T-T

o)/T

Exergy Loss: 6.50E+6 BTU/hr

Figure 7.11. Ethyl Benzene Heat Exchanger Exergy Analysis

0

20000000

40000000

60000000

80000000

100000000

120000000

140000000

160000000

180000000

2000000000

0.1

0.2

0.3

0.4

0.5

0.6

3PHX

COLDHOT

DUTY (BTU/hr)

(T-T

o)/T

Figure 7.12. Product Stream Heat Exchanger Exergy Analysis

86

Exergy Loss: 3.76E+07 BTU/hr

100000000 150000000 200000000 250000000 300000000 350000000 4000000000

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

Methane Target

Hot CompositeCold Composite

DUTY (BTU/hr)

(T-T

o)/T

Utility Loss

Figure 7.13: Methane Target Exergy Analysis

Comparison

Methane PlantExergy Loss due to Utilities 2.25E+07 BTU/hr

Exergy Loss due to Process 5.32E+07 BTU/hr

Methane Plant Total Exergy Loss 7.58E+07 BTU/hr

Traditional Plant

WHX Exergy Loss to Process 7.76E+07BTU/hr

EBHX Exergy Loss to Process 6.50E+6BTU/hr

3PHX Exergy Loss to Process 3.76E+07BTU/hr

Traditional Plant Total Exergy Loss 1.22E+08 BTU/hr

Figure 7.9. Comparison Table – Methane Plant vs. Traditional Plant Exergy Loss

From table 7.9, the total exergy loss of the traditional plant is significantly more than the exergy loss to the

methane plant. This clearly shows the inefficiencies of the traditional plant and that just simply matching hot

streams with cold streams and vice versa is not an optimal approach to creating a heat exchanger network.

87

8. PDF and Stream Table of the plant with Traditional HX Network

Figure 8.1: Process Flow Diagram of Traditional Plant

88

Styrene Plant Stream Tables

3PH 3POIL 3PVAP 3PWAT

Mole Frac

H2O 0.853204 0.014067 0.006684 1

EB 0.0343727

0.405118 4.07E-03 0

STYRENE 0.0417296

0.493048 3.34E-03 0

H2 0.0530301

0.000467 0.825288 0

BENZENE 0.0025228

0.027977 2.60E-03 0

TOLUENE 0.0041035

0.047749 1.29E-03 0

METHANE 0.0025228

0.000243 3.90E-02 0

C2H4 0.0018921

7.80E-04 0.028445 0

CO2 5.99E-03 0.003063 8.93E-02 0

N2 0.00E+00 0 0.00E+00

0

O2 0.00E+00 0.00E+00

0 0

NAP 3.15E-04 3.74E-03 1.36E-06 0

ENAP 3.15E-04 3.75E-03 1.39E-07 0

Mass Flow lb/hr

H2O 116188 161.2976 58.44041 115969

EB 2.76E+04 2.74E+04

2.10E+02

0

STYRENE 3.29E+04 3.27E+04

1.69E+02

0

H2 808.0835 0.599317 807.4837 0

BENZENE 1.49E+03 1390.954 98.66127 0.00E+00

TOLUENE 2858.101 2800.247 57.85516 0

89

METHANE 305.9329 2.482495 303.4502 0

C2H4 401.2356 13.93033 387.305 0

CO2 1993.242 85.78875 1907.453 0

N2 0 0 0 0

O2 0 0 0 0

NAP 305.5319 305.4477 0.084339 0

ENAP 372.4046 372.3942 0.010561 0

Mass Frac

H2O 0.6274994

0.002474 0.014611 1

EB 0.148979 0.419912 0.052484 0

STYRENE 0.1774312

0.50135 0.042264 0

H2 0.0043642

9.19E-06 0.201886 0

BENZENE 0.008045 2.13E-02 0.024667 0

TOLUENE 0.0154357

0.042953 0.014465 0

METHANE 0.0016523

3.81E-05 0.075868 0

C2H4 2.17E-03 2.14E-04 0.096833 0

CO2 1.08E-02 0.001316 0.476898 0

N2 0 0 0 0

O2 0.00E+00 0.00E+00

0 0

NAP 1.65E-03 4.69E-03 2.11E-05 0

ENAP 0.0020113

5.71E-03 2.64E-06 0

Total Flow lbmol/hr 7559.081 636.4804 485.3598 6437.241

Total Flow lb/hr 185161 65192.66 3999.706 115969

Total Flow cuft/hr 1.12E+05 1.22E+03

1.09E+05

1865.17

Temperature F 8.50E+01 8.50E+01

8.50E+01

85

Pressure psia 26 26 26 26

90

Vapor Frac 6.42E-02 0 1 0.00E+00

Liquid Frac 9.36E-01 1 0.00E+00

1

Solid Frac 0 0 0 0

Enthalpy Btu/lbmol -103900 19549.2 -15965.7 -122740

Enthalpy Btu/lb -4241.757 190.8601 -1937.42 -6813.06

Enthalpy kcal/hr -1.98E+08

3135500 -1952700

-2E+08

Entropy Btu/lbmol-R -40.23626 -85.1793 -1.43651 -38.718

Entropy Btu/lb-R -1.64E+00

-0.83161 -0.17432 -2.15E+00

Density lbmol/cuft 0.0673497

0.519712 0.004447 3.451289

Density lb/cuft 1.65E+00 5.32E+01

3.66E-02 6.22E+01

Average MW 24.49518 102.4268 8.240704 18.01528

Liq Vol 60F cuft/hr 3459.978 1172.326 426.4318 1861.22

*** VAPOR PHASE ***

CPMX Btu/lbmol-R 7.479035 7.479035

MUMX cP 0.0120011

0.012001

Average MW 8.240704 8.240704

KMX Btu-ft/hr-sqft-R 0.07274 0.07274

*** LIQUID PHASE ***

CPMX Btu/lbmol-R 19.89787 39.56396 17.95339

MUMX 0.555026 0.807185

Average MW 25.61047 102.4268 18.01528

KMX 0.076274 0.354864

91

AIRFEED

BENPROD

BENTOL

CH4FEED

92

Mole Frac

H2O 0 0 0 0

EB 0.00E+00

0 0.005322

0

STYRENE 0.00E+00

0 3.20E-06

0

H2 0 0 0 0

BENZENE 0 0.998297 0.367717

0

TOLUENE 0.00E+00

0.001703 0.626958

0

METHANE 0.00E+00

0 0 1

C2H4 0 0 0 0

CO2 0 0 0 0

N2 0.79 0 0 0

O2 0.21 0 0 0

NAP 0 0 7.31E-13

0

ENAP 0 0 7.31E-13

0

Mass Flow lb/hr

H2O 0 0 0 0

EB 0.00E+00

0 27.36152

0

STYRENE 0.00E+00

0 0.016116

0

H2 0 0 0 0

BENZENE 0 1390.258 1.39E+03

0.00E+00

TOLUENE 0 2.797447 2797.447

0

METHANE 0 0 0 7139.028

C2H4 0 0 0 0

CO2 0 0 0 0

N2 160934 0 0 0

93

O2 48866.01 0 0 0

NAP 0 0 4.54E-09

0

ENAP 0 0 5.53E-09

0

Mass Frac

H2O 0 0 0 0

EB 0.00E+00

0 0.00649 0

STYRENE 0.00E+00

0 3.82E-06

0

H2 0 0 0 0

BENZENE 0 0.997992 0.32994 0

TOLUENE 0.00E+00

0.002008 0.663566

0

METHANE 0.00E+00

0 0 1

C2H4 0 0 0 0

CO2 0 0 0 0

N2 0.767083 0 0 0

O2 0.232917 0 0 0

NAP 0 0 1.08E-12

0

ENAP 0 0 1.31E-12

0

Total Flow lbmol/hr 7272 17.82826 48.42533

445

Total Flow lb/hr 209800 1393.056 4215.778

7139.028

Total Flow cuft/hr 2.81E+06

26.4049 84.63593

171667

Temperature F 7.00E+01

176.3432 206.6813

70

Pressure psia 14.7 14.7 1.47E+01

1.47E+01

94

Vapor Frac 1 0 0.00E+00

1.00E+00

Liquid Frac 0 1.00E+00 1 0

Solid Frac 0 0 0 0

Enthalpy Btu/lbmol -52.3666 24580.72 16037.34

-32105.4

Enthalpy Btu/lb -1.81511 314.5827 184.2159

-2001.24

Enthalpy kcal/hr -95962.4 110432 195703 -3600200

Entropy Btu/lbmol-R 0.922993 -54.43 -64.1853

-19.373

Entropy Btu/lb-R 0.031992 -0.69659 -7.37E-01

-1.21E+00

Density lbmol/cuft 0.002588 0.675187 0.572161

0.002592

Density lb/cuft 7.47E-02 5.28E+01 4.98E+01

4.16E-02

Average MW 28.8504 78.13753 87.05729

16.04276

Liq Vol 60F cuft/hr 6238.748 25.28501 77.26153

381.7716

*** VAPOR PHASE ***

CPMX Btu/lbmol-R 6.980044 8.511541

MUMX cP 0.018143 0.01108

Average MW 28.8504 16.04276

KMX Btu-ft/hr-sqft-R 0.01468 0.019436

*** LIQUID PHASE ***

CPMX Btu/lbmol-R 33.08322 39.0477

MUMX 0.351342 0.292175

Average MW 78.13753 87.05729

KMX 0.073033 0.067503

95

CLDUTI

N

CLDUTO

UT

EBBENT

OL

EBFEE

D

Mole Frac

H2O 1 1 0 0

EB 0.00E+0

0

0.00E+00 7.64E-01 1.00E+

00

STYRENE 0.00E+0

0

0.00E+00 9.30E-02 0.00E+

00

H2 0 0 0 0

BENZENE 0 0 0.052792 0

TOLUENE 0 0 0.090101 0

METHANE 0 0 0 0

C2H4 0 0 0 0

CO2 0.00E+0

0

0.00E+00 0.00E+00 0

N2 0.00E+0

0

0.00E+00 0.00E+00 0

O2 0 0 0 0

NAP 0 0 1.76E-08 0

ENAP 0 0 1.76E-08 0

Mass Flow lb/hr

H2O 3427730 3427730 0 0

EB 0.00E+0

0

0.00E+00 2.74E+04 3.67E+

04

STYRENE 0.00E+0

0

0.00E+00 3.27E+03 0.00E+

00

96

H2 0 0 0 0

BENZENE 0 0 1390.954 0

TOLUENE 0 0 2800.247 0

METHANE 0 0 0 0

C2H4 0 0 0 0

CO2 0.00E+0

0

0.00E+00 0.00E+00 0

N2 0 0 0 0

O2 0 0 0 0

NAP 0 0 0.000759 0

ENAP 0 0 0.000925 0

Mass Frac

H2O 1 1 0 0

EB 0.00E+0

0

0.00E+00 7.86E-01 1.00E+

00

STYRENE 0.00E+0

0

0.00E+00 9.39E-02 0.00E+

00

H2 0 0 0 0

BENZENE 0 0 0.039946 0

TOLUENE 0 0 0.080418 0

METHANE 0 0 0 0

C2H4 0 0 0 0

CO2 0.00E+0

0

0.00E+00 0.00E+00 0

N2 0.00E+0

0

0.00E+00 0.00E+00 0

97

O2 0 0 0 0

NAP 0 0 2.18E-08 0

ENAP 0 0 2.66E-08 0

Total Flow

lbmol/hr

190268 190268 337.2999 346.01

Total Flow lb/hr 3427730 3427730 34821.15 36734.9

8

Total Flow

cuft/hr

6.44E+0

4

6.56E+04 6.63E+02 6.94E+

02

Temperature F 7.00E+0

1

1.20E+02 9.89E+01 7.00E+

01

Pressure psia 14.7 14.7 0.580103 14.7

Vapor Frac 0 0 0 0

Liquid Frac 1 1.00E+00 1 1

Solid Frac 0 0 0 0

Enthalpy

Btu/lbmol

-123840 -122870 3054.771 -5082.6

Enthalpy Btu/lb -6874.32 -6820.3 29.59046 -

47.8735

Enthalpy kcal/hr -5.9E+09 -5.9E+09 259650 -443170

Entropy

Btu/lbmol-R

-40.3755 -38.6196 -94.8284 -105.67

Entropy Btu/lb-R -2.24118 -2.14371 -0.91857 -

0.99531

Density

lbmol/cuft

2.955475 2.900031 0.508631 0.49888

6

Density lb/cuft 5.32E+0

1

5.22E+01 5.25E+01 5.30E+

01

98

Average MW 18.01528 18.01528 103.235 106.167

4

Liq Vol 60F

cuft/hr

55012.81 55012.81 638.3098 676.534

6

*** VAPOR PHASE ***

CPMX Btu/lbmol-R

MUMX cP

Average MW

KMX Btu-ft/hr-sqft-R

*** LIQUID PHASE ***

CPMX

Btu/lbmol-R

19.47247 19.47741 40.76167 40.4173

3

MUMX 0.620855 0.500207 0.562895 0.68132

7

Average MW 18.01528 18.01528 103.235 106.167

4

KMX 0.347101 0.367605 0.073613 0.07499

6

EBHIPT EBRECYCL EBTHIP EBTOTAL

Mole Frac

H2O 0 0 0 0

EB 0.950524 0.891262 0.764069 9.51E-01

STYRENE 0.049428 0.108633 0.093037 4.94E-02

H2 0 0 0 0

BENZENE 5.67E-13 1.25E-12 0.052792 5.67E-13

99

TOLUENE 4.79E-05 0.000105 0.090101 4.79E-05

METHANE 0 0 0 0.00E+00

C2H4 0 0 0 0

CO2 0 0 0 0.00E+00

N2 0 0 0 0.00E+00

O2 0 0 0 0

NAP 9.33E-09 2.05E-08 1.76E-08 9.33E-09

ENAP 9.33E-09 2.05E-08 1.76E-08 9.33E-09

Mass Flow lb/hr

H2O 0 0 0 0

EB 64069.14 27334.16 27361.52 6.41E+04

STYRENE 3268.415 3268.415 3268.432 3.27E+03

H2 0 0 0 0

BENZENE 2.81E-08 2.81E-08 1390.954 2.81E-08

TOLUENE 2.800247 2.800247 2800.247 2.800247

METHANE 0 0 0 0

C2H4 0 0 0 0

CO2 0 0 0 0

N2 0 0 0 0

O2 0 0 0 0

NAP 0.000759 0.000759 0.000759 0.000759

ENAP 0.000925 0.000925 0.000925 0.000925

Mass Frac

H2O 0 0 0 0

100

EB 0.951423 0.893116 0.785773 0.951423

STYRENE 0.048536 0.106792 0.093863 0.048536

H2 0 0 0 0

BENZENE 4.18E-13 9.19E-13 0.039946 4.18E-13

TOLUENE 4.16E-05 9.15E-05 0.080418 4.16E-05

METHANE 0 0 0 0

C2H4 0 0.00E+00 0 0

CO2 0 0 0 0

N2 0 0 0 0

O2 0 0 0 0

NAP 1.13E-08 2.48E-08 2.18E-08 1.13E-08

ENAP 1.37E-08 3.02E-08 2.66E-08 1.37E-08

Total Flow lbmol/hr 634.8846 288.8746 337.2999 634.8846

Total Flow lb/hr 67340.36 30605.37 34821.15 67340.36

Total Flow cuft/hr 370352 645.9151 663.1305 1.33E+03

Temperature F 965 278.8672 99.03346 1.72E+02

Pressure psia 26 14.7 14.7 14.7

Vapor Frac 1 0 0 0

Liquid Frac 0 1 1 1.00E+00

Solid Frac 0 0 0 0

Enthalpy Btu/lbmol 60706.5 10081.6 3064.837 1817.158

Enthalpy Btu/lb 572.3406 95.15705 29.68797 17.13216

Enthalpy kcal/hr 9712330 733892 260505 290724

Entropy Btu/lbmol- -33.1109 -86.6478 -94.8196 -96.2982

101

R

Entropy Btu/lb-R -0.31217 -0.81784 -9.18E-

01

-0.9079

Density lbmol/cuft 0.001714 0.447233 0.508648 0.476709

Density lb/cuft 0.181828 4.74E+01 5.25E+01 5.06E+01

Average MW 106.0671 105.9469 103.235 106.0671

Liq Vol 60F cuft/hr 1237.583 561.0483 638.3098 1237.583

*** VAPOR PHASE ***

CPMX Btu/lbmol-R 66.67

MUMX cP 0.015773

Average MW 106.0671

KMX Btu-ft/hr-sqft-

R

0.037895

*** LIQUID PHASE ***

CPMX Btu/lbmol-R 53.09234 40.76671 46.57914

MUMX 0.271708 0.56296 0.429554

Average MW 105.9469 103.235 106.0671

KMX 0.059553 0.073601 0.067416

FIREFEE

D

GIBBSE

B

GIBBSPR

O

GIBBSW

AT

Mole Frac

H2O 0.000395 0 0.820016 1

EB 0.000241 1 2.50E-02 0.00E+00

STYRENE 0.000198 0 7.75E-02 0.00E+00

102

H2 0.048835 0 0.077482 0

BENZENE 0.000154 0 0 0

TOLUENE 7.66E-05 0 0 0

METHANE 0.056559 0 0.00E+00 0

C2H4 0.001683 0 0 0

CO2 0.005284 0 0.00E+00 0.00E+00

N2 0.700394 0 0.00E+00 0.00E+00

O2 0.186181 0 0 0

NAP 8.02E-08 0 0 0

ENAP 8.24E-09 0 0 0

Mass Flow lb/hr

H2O 58.44041 0 144.1222 144.1222

EB 209.9186 1.06E+0

2

2.59E+01 0.00E+00

STYRENE 169.0446 0.00E+0

0

7.87E+01 0.00E+00

H2 807.4837 0 1.523821 0

BENZENE 98.66127 0 0 0

TOLUENE 57.85516 0 0 0

METHANE 7442.478 0 0 0

C2H4 387.305 0 0 0

CO2 1907.453 0 0.00E+00 0.00E+00

N2 160934 0 0 0

O2 48866.01 0 0 0

103

NAP 0.084339 0 0 0

ENAP 0.010561 0 0 0

Mass Frac

H2O 0.000265 0 0.575822 1

EB 0.00095 1 1.04E-01 0.00E+00

STYRENE 0.000765 0 3.15E-01 0.00E+00

H2 0.003655 0 0.006088 0

BENZENE 0.000447 0 0 0

TOLUENE 0.000262 0 0 0

METHANE 0.033686 0 0 0

C2H4 1.75E-03 0 0 0.00E+00

CO2 0.008633 0 0.00E+00 0.00E+00

N2 0.72841 0 0.00E+00 0.00E+00

O2 2.21E-01 0 0 0

NAP 3.82E-07 0 0 0

ENAP 4.78E-08 0 0 0

Total Flow

lbmol/hr

8202.36 1 9.755909 8

Total Flow lb/hr 220939 106.167

4

250.2896 144.1222

Total Flow

cuft/hr

3175230 2.00E+0

0

6.58E+03 2.71E+00

Temperature F 70.91102 7.00E+0

1

1.18E+03 7.00E+01

Pressure psia 14.7 14.7 26 14.7

Vapor Frac 1 0 1.00E+00 0.00E+00

104

Liquid Frac 0 1 0.00E+00 1.00E+00

Solid Frac 0 0 0 0

Enthalpy

Btu/lbmol

-2732.98 -5082.6 -65784 -123840

Enthalpy Btu/lb -101.462 -

47.8735

-2564.16 -6874.32

Enthalpy kcal/hr -5648900 -

1280.79

-161730 -249660

Entropy

Btu/lbmol-R

0.588367 -105.67 -0.77054 -40.3755

Entropy Btu/lb-R 2.18E-02 -

0.99531

-0.03003 -2.24118

Density

lbmol/cuft

0.002583 0.49888

6

0.001483 2.955475

Density lb/cuft 6.96E-02 5.30E+0

1

3.81E-02 5.32E+01

Average MW 26.93601 106.167

4

25.65518 18.01528

Liq Vol 60F

cuft/hr

7046.951 1.95524

6

4.826136 2.313066

*** VAPOR PHASE ***

CPMX

Btu/lbmol-R

7.090436 15.25854

MUMX cP 0.017643 0.031565

Average MW 26.93601 25.65518

KMX Btu-ft/hr-

sqft-R

0.017383 0.054703

105

*** LIQUID PHASE ***

CPMX Btu/lbmol-R 40.4173

3

19.47247

MUMX 0.68132

7

0.620855

Average MW 106.167

4

18.01528

KMX 0.07499

6

0.347101

H2LOW

P

HOTGA

S

HVYPRO

D

MAGJUN

K

Mole Frac

H2O 0.00668

4

0.17257 0 0.755477

EB 0.00407

4

0 1.65E-07 0

STYRENE 0.00334

4

0 0.005923 0

H2 0.82528

8

0 0 0.025086

BENZENE 0.00260

2

0 0 0

TOLUENE 0.00129

4

0.00E+0

0

0.00E+00 0

METHANE 0.03897

1

0 0 0.013057

C2H4 0.02844 0 0 0.041899

106

5

CO2 8.93E-

02

7.19E-

02

0 0.164481

N2 0.00E+0

0

7.17E-

01

0.00E+00 0

O2 0.00E+0

0

0.03814

6

0 0

NAP 1.36E-

06

0 0.494246 0

ENAP 1.39E-

07

0 0.499831 0

Mass Flow lb/hr

H2O 58.4404

1

24895.8

5

0 161.2976

EB 2.10E+0

2

0.00E+0

0

8.37E-05 0

STYRENE 1.69E+0

2

0.00E+0

0

2.941588 0

H2 807.483

7

0 0 0.599317

BENZENE 9.87E+0

1

0 0 0

TOLUENE 57.8551

6

0 0 0

METHANE 303.450

2

0 0 2.482495

C2H4 387.305 0 0 13.93033

CO2 1907.45 2.53E+0 0 85.78875

107

3 4

N2 0 1.61E+0

5

0.00E+00 0

O2 0 9774.58

1

0 0

NAP 0.08433

9

0 302.09 0

ENAP 0.01056

1

0 372.3704 0

Mass Frac

H2O 0.01461

1

0.11268

2

0 0.610748

EB 0.05248

4

0 1.24E-07 0

STYRENE 0.04226

4

0 0.004342 0

H2 0.20188

6

0 0 0.002269

BENZENE 0.02466

7

0 0 0

TOLUENE 0.01446

5

0.00E+0

0

0.00E+00 0

METHANE 0.07586

8

0 0 0.0094

C2H4 9.68E-

02

0 0 0.052747

CO2 4.77E-

01

1.15E-

01

0 0.324836

N2 0 7.28E- 0.00E+00 0

108

01

O2 0.00E+0

0

0.04424

1

0 0

NAP 2.11E-

05

0 0.445954 0

ENAP 2.64E-

06

0 0.549704 0

Total Flow

lbmol/hr

485.359

8

8007.92

8

4.768645 11.85129

Total Flow lb/hr 3999.70

6

220939 677.4021 264.0985

Total Flow cuft/hr 1.93E+0

5

2152680

0

11.85407 662.3681

Temperature F 8.49E+0

1

3222.13

9

249.4368 85

Pressure psia 14.7 14.7 0.580103 26

Vapor Frac 1.00E+0

0

1 0 2.49E-01

Liquid Frac 0.00E+0

0

0 1 7.51E-01

Solid Frac 0 0 0 0

Enthalpy

Btu/lbmol

-

15965.7

-

2799.33

38223.31 -120530

Enthalpy Btu/lb -

1937.42

-

101.462

269.0771 -5408.84

Enthalpy kcal/hr -

1952700

-

5648900

45932.12 -359970

Entropy

Btu/lbmol-R

-

0.30382

15.8437

3

-92.1113 -30.3684

109

Entropy Btu/lb-R -

0.03687

0.57425

6

-0.64843 -1.36277

Density

lbmol/cuft

0.00251

5

0.00037

2

0.402279 0.017892

Density lb/cuft 2.07E-

02

1.03E-

02

5.71E+01 3.99E-01

Average MW 8.24070

4

27.5900

1

142.0534 22.28437

Liq Vol 60F

cuft/hr

426.431

8

6084.09

7

10.10198 5.322676

*** VAPOR PHASE ***

CPMX

Btu/lbmol-R

7.47419

2

9.67582

9

9.015499

MUMX cP 0.01199

5

0.07045

7

0.01422

Average MW 8.24070

4

27.5900

1

35.13605

KMX Btu-ft/hr-

sqft-R

0.07272

7

0.07731

8

0.015013

*** LIQUID PHASE ***

CPMX Btu/lbmol-R 61.38189 19.46151

MUMX 0.727047 0.580796

Average MW 142.0534 18.01569

KMX 0.072985 0.353364

OILS2CO

L

R1FEE

D

R1PRO

D

R2FEE

D

110

Mole Frac

H2O 0 0.85577

2

0.83182

4

0.89569

3

EB 4.13E-01 0.13709

3

0.10527

2

0.06529

3

STYRENE 5.02E-01 0.00712

9

0.03491

3

0.02165

4

H2 0 0 0.02798

4

0.01735

6

BENZENE 0.028508 8.18E-

14

7.95E-

14

4.93E-

14

TOLUENE 4.87E-02 6.90E-

06

6.71E-

06

4.16E-

06

METHANE 0.00E+00 0 0 0

C2H4 0 0 0 0

CO2 0.00E+00 0 0 0

N2 0 0 0 0

O2 0 0 0 0

NAP 0.003815 1.35E-

09

1.31E-

09

8.11E-

10

ENAP 0.003816 1.35E-

09

1.31E-

09

8.11E-

10

Mass Flow lb/hr

H2O 0 67864.4

9

67864.4

9

117820

EB 2.74E+04 64069.1

4

50614.6

2

50614.6

2

STYRENE 3.27E+04 3268.41 16467.4 16467.4

111

5 6 6

H2 0 0 255.471 255.471

BENZENE 1390.954 2.81E-

08

2.81E-

08

2.81E-

08

TOLUENE 2800.247 2.80024

7

2.80024

7

2.80024

7

METHANE 0 0 0 0

C2H4 0 0 0 0

CO2 0 0 0 0

N2 0 0 0 0

O2 0 0 0 0

NAP 305.4477 0.00075

9

0.00075

9

0.00075

9

ENAP 372.3942 0.00092

5

0.00092

5

0.00092

5

Mass Frac

H2O 0 0.50193

8

0.50193

8

0.63631

4

EB 4.22E-01 0.47386

7

0.37435

5

0.27335

5

STYRENE 5.03E-01 0.02417

4

0.12179

6

0.08893

6

H2 0 0 0.00189 0.00138

BENZENE 0.021423 2.08E-

13

2.08E-

13

1.52E-

13

TOLUENE 4.31E-02 2.07E-

05

2.07E-

05

1.51E-

05

METHANE 0.00E+00 0 0 0

112

C2H4 0 0 0 0

CO2 0.00E+00 0 0 0

N2 0 0 0 0

O2 0 0 0 0

NAP 0.004704 5.61E-

09

5.61E-

09

4.10E-

09

ENAP 0.005735 6.84E-

09

6.84E-

09

5.00E-

09

Total Flow

lbmol/hr

624.6291 4401.93

6

4528.66

5

7301.63

4

Total Flow lb/hr 64928.56 135205 135205 185161

Total Flow cuft/hr 1.22E+03 2966520 2894510 4921740

Temperature F 8.50E+01 1174.9 1090.88

9

1175.00

1

Pressure psia 26 26 26 26

Vapor Frac 0 1.00E+0

0

1.00E+0

0

1.00E+0

0

Liquid Frac 1 0 0 0.00E+0

0

Solid Frac 0 0 0 0

Enthalpy

Btu/lbmol

22123.65 -69936.6 -67979.5 -77084.6

Enthalpy Btu/lb 212.8351 -2276.96 -2276.96 -3039.76

Enthalpy kcal/hr 3482350 -

7.8E+07

-

7.8E+07

-

1.4E+08

113

Entropy

Btu/lbmol-R

-86.4527 -4.46871 -4.09238 -2.68356

Entropy Btu/lb-R -0.8317 -1.45E-

01

-1.37E-

01

-0.10582

Density lbmol/cuft 0.512528 0.00148

4

0.00156

5

0.00148

4

Density lb/cuft 5.33E+01 4.56E-

02

4.67E-

02

3.76E-

02

Average MW 103.9474 30.7148

6

29.8553

4

25.3588

Liq Vol 60F

cuft/hr

1167.004 2326.76

3

2420.28

3

3222.04

*** VAPOR PHASE ***

CPMX Btu/lbmol-R 18.5179

1

17.6455

6

14.8128

5

MUMX cP 0.03058

4

0.02906

3

0.03194

4

Average MW 30.7148

6

29.8553

4

25.3588

KMX Btu-ft/hr-sqft-R 0.04907

8

0.04726

6

0.05028

3

*** LIQUID PHASE ***

CPMX Btu/lbmol-

R

39.99771

MUMX 0.586974

Average MW 103.9474

KMX 0.076323

114

R2OUT R2PRO

D

SH SH2

Mole Frac

H2O 0.85320

5

0.853205 0 0

EB 0.03437

3

0.034373 0.00044

9

0.00016

4

STYRENE 0.04172

9

0.041729 0.98296

2

0.85607

9

H2 0.05303 0.05303 0 0

BENZENE 0.00252

3

0.002523 1.50E-28 0

TOLUENE 0.00410

4

0.004104 2.51E-15 1.05E-16

METHANE 0.00252

3

0.002523 0 0

C2H4 0.00189

2

0.001892 0 0

CO2 0.00599

2

0.005992 0 0

N2 0 0 0 0

O2 0 0 0 0

NAP 0.00031

5

0.000315 0.00829

4

0.07151

1

ENAP 0.00031

5

0.000315 0.00829

6

0.07224

7

Mass Flow lb/hr

115

H2O 116188 116188 0 0

EB 27584.9

7

27584.97 13.6876 0.57509

9

STYRENE 32853.1

4

32853.14 29415.8

8

2941.58

8

H2 808.079

2

808.0792 0 0

BENZENE 1489.60

7

1489.607 3.36E-24 0

TOLUENE 2858.08

1

2858.081 6.66E-11 3.18E-13

METHANE 305.931

3

305.9313 0 0

C2H4 401.233

4

401.2334 0 0

CO2 1993.23

2

1993.232 0 0

N2 0 0 0 0

O2 0 0 0 0

NAP 305.530

2

305.5302 305.446

9

302.392

4

ENAP 372.402

6

372.4026 372.393

3

372.372

3

Mass Frac

H2O 0.62750

1

0.627501 0 0

EB 0.14897

9

0.148979 0.00045

5

0.00015

9

STYRENE 0.17743 0.177431 0.97703 0.81328

116

1 1 4

H2 0.00436

4

0.004364 0 0

BENZENE 0.00804

5

0.008045 1.12E-28 0

TOLUENE 0.01543

6

0.015436 2.21E-15 8.80E-17

METHANE 0.00165

2

0.001652 0 0

C2H4 0.00216

7

0.002167 0 0

CO2 0.01076

5

0.010765 0 0

N2 0 0 0 0

O2 0 0 0 0

NAP 0.00165 0.00165 0.01014

5

0.08360

5

ENAP 0.00201

1

0.002011 0.01236

9

0.10295

3

Total Flow lbmol/hr 7559.07

5

7559.075 287.329

2

32.9915

4

Total Flow lb/hr 185161 185161 30107.4

1

3616.92

8

Total Flow cuft/hr 4776630 3617250 570.133

6

67.2967

8

Temperature F 1073.02

9

703.4763 128.558

5

133.705

7

Pressure psia 26 26 0.58010 0.58010

117

3 3

Vapor Frac 1 1 0 0

Liquid Frac 0 0 1 1

Solid Frac 0 0 0 0

Enthalpy Btu/lbmol -74459.3 -79404.8 46938.9

7

45138.4

Enthalpy Btu/lb -3039.76 -3241.66 447.960

7

411.726

6

Enthalpy kcal/hr -

1.4E+08

-1.5E+08 3398650 375269

Entropy Btu/lbmol-

R

-2.1089 -5.79299 -74.519 -77.166

Entropy Btu/lb-R -0.08609 -0.2365 -0.71117 -0.70386

Density lbmol/cuft 0.00158

3

0.00209 0.50396

8

0.49024

Density lb/cuft 0.03876

4

0.051188 52.8076

5

53.7459

3

Average MW 24.4951

5

24.49515 104.783

7

109.632

Liq Vol 60F cuft/hr 3459.96

9

3459.969 528.693

9

61.8994

4

*** VAPOR PHASE ***

CPMX Btu/lbmol-R 14.0060

3

12.72391

MUMX cP 0.03009

3

0.022583

Average MW 24.4951

5

24.49515

118

KMX Btu-ft/hr-sqft-

R

0.04897

1

0.033641

*** LIQUID PHASE ***

CPMX Btu/lbmol-R 42.2441

2

43.8971

3

MUMX 0.47956

9

0.52503

9

Average MW 104.783

7

109.632

KMX 0.07531 0.07549

4

STACK STHIPRO

D

STLPRO

D

TOLPRO

D

Mole Frac

H2O 0.17257 0 0 0

EB 0 0.000486 0.000192 0.008423

STYRENE 0 0.99942 0.999724 5.06E-06

H2 0 0 0 0

BENZENE 0 0 0 0.000291

TOLUENE 0 2.83E-15 0 0.991281

METHANE 0 0 0 0

C2H4 0 0 0 0

CO2 0.07188

5

0 0 0

N2 0.71739 0 0 0

119

9

O2 0.03814

6

0 0 0

NAP 0 9.37E-05 8.36E-05 1.16E-12

ENAP 0 5.28E-07 4.34E-07 1.16E-12

Mass Flow lb/hr

H2O 24895.8

5

0 0 0

EB 0 13.1125 0.575015 27.36152

STYRENE 0 26474.3 2938.647 0.016116

H2 0 0 0 0

BENZENE 0 0 0 0.695477

TOLUENE 0 6.62E-11 0 2794.649

METHANE 0 0 0 0

C2H4 0 0 0 0

CO2 25334.3

2

0 0 0

N2 160934 0 0 0

O2 9774.58

1

0 0 0

NAP 0 3.054469 0.302392 4.54E-09

ENAP 0 0.020965 0.001915 5.53E-09

Mass Frac

H2O 0.11268

2

0 0 0

EB 0 0.000495 0.000196 0.009693

120

STYRENE 0 0.999389 0.999701 5.71E-06

H2 0 0 0 0

BENZENE 0 0 0 0.000246

TOLUENE 0 2.50E-15 0 0.990055

METHANE 0 0 0 0

C2H4 0 0 0 0

CO2 0.11466

7

0 0 0

N2 0.72841 0 0 0

O2 0.04424

1

0 0 0

NAP 0 0.000115 0.000103 1.61E-12

ENAP 0 7.91E-07 6.51E-07 1.96E-12

Total Flow

lbmol/hr

8007.92

8

254.3377 28.2229 30.59708

Total Flow lb/hr 220939 26490.48 2939.526 2822.722

Total Flow cuft/hr 414249

0

502.856 55.79947 58.64313

Temperature F 249.449

9

127.9395 127.944 234.1276

Pressure psia 14.7 0.580103 0.580103 14.7

Vapor Frac 1 0 0 0

Liquid Frac 0 1 1 1

Solid Frac 0 0 0 0

Enthalpy

Btu/lbmol

-

28841.3

47175.83 47190.7 11602.91

Enthalpy Btu/lb - 452.9396 453.0861 125.7705

121

1045.35

Enthalpy kcal/hr -

5.8E+0

7

3023590 335623 89462.32

Entropy

Btu/lbmol-R

1.95563

3

-74.2516 -74.2481 -71.1502

Entropy Btu/lb-R 0.07088

2

-0.7129 -0.71287 -0.77124

Density

lbmol/cuft

0.00193

3

0.505786 0.505792 0.52175

Density lb/cuft 0.05333

5

52.68006 52.68018 48.1339

Average MW 27.5900

1

104.1548 104.1539 92.25465

Liq Vol 60F

cuft/hr

6084.09

7

466.7944 51.79746 51.97652

*** VAPOR PHASE ***

CPMX Btu/lbmol-

R

7.41808

2

MUMX cP 0.02143

1

Average MW 27.5900

1

KMX Btu-ft/hr-

sqft-R

0.01762

5

*** LIQUID PHASE ***

CPMX Btu/lbmol-R 42.03476 42.0344 43.15118

MUMX 0.473925 0.4739 0.264231

Average MW 104.1548 104.1539 92.25465

122

KMX 0.075283 0.075284 0.064205

WATFEE

D

WATHI

P

WATHIP

T

WATR1 WATR2

Mole Frac

H2O 1 1 1 1 1

EB 0 0 0 0 0

STYRENE 0 0 0 0 0

H2 0 0 0 0 0

BENZENE 0 0 0 0 0

TOLUENE 0 0 0 0 0

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

NAP 0 0 0 0 0

ENAP 0 0 0 0 0

Mass Flow lb/hr

H2O 117820 117820 117820 67864.4

9

49955.8

EB 0 0 0 0 0

123

STYRENE 0 0 0 0 0

H2 0 0 0 0 0

BENZENE 0 0 0 0 0

TOLUENE 0 0 0 0 0

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

NAP 0 0 0 0 0

ENAP 0 0 0 0 0

Mass Frac

H2O 1 1 1 1 1

EB 0 0 0 0 0

STYRENE 0 0 0 0 0

H2 0 0 0 0 0

BENZENE 0 0 0 0 0

TOLUENE 0 0 0 0 0

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

NAP 0 0 0 0 0

ENAP 0 0 0 0 0

124

Total Flow

lbmol/hr

6540.02 6540.02 6540.02 3767.05

2

2772.96

8

Total Flow lb/hr 117820 117820 117820 67864.4

9

49955.8

Total Flow cuft/hr 2212.849 2212.85 5085230 292909

0

215614

0

Temperature F 70 70.0259

6

1425 1425 1425

Pressure psia 14.7 26 26 26 26

Vapor Frac 0 0 1 1 1

Liquid Frac 1 1 0 0 0

Solid Frac 0 0 0 0 0

Enthalpy

Btu/lbmol

-123840 -123840 -91954.7 -

91954.7

-

91954.7

Enthalpy Btu/lb -6874.32 -6874.26 -5104.26 -

5104.26

-

5104.26

Enthalpy kcal/hr -2E+08 -2E+08 -1.5E+08 -

8.7E+07

-

6.4E+07

Entropy

Btu/lbmol-R

-40.3755 -40.3748 -0.80701 -

0.80701

-

0.80701

Entropy Btu/lb-R -2.24118 -2.24114 -0.0448 -0.0448 -0.0448

Density lbmol/cuft 2.955475 2.95547

4

0.001286 0.00128

6

0.00128

6

Density lb/cuft 53.24371 53.2437 0.023169 0.02316

9

0.02316

9

Average MW 18.01528 18.0152

8

18.01528 18.0152

8

18.0152

8

125

Liq Vol 60F

cuft/hr

1890.937 1890.93

7

1890.937 1089.18 801.757

4

*** VAPOR PHASE ***

CPMX Btu/lbmol-R 10.00778 10.0077

8

10.0077

8

MUMX cP 0.039724 0.03972

4

0.03972

4

Average MW 18.01528 18.0152

8

18.0152

8

KMX Btu-ft/hr-sqft-R 0.059462 0.05946

2

0.05946

2

*** LIQUID PHASE ***

CPMX Btu/lbmol-

R

19.47247 19.4720

8

MUMX 0.620855 0.62082

Average MW 18.01528 18.0152

8

KMX 0.347101 0.34711

3

126

9. Process Equipment List

Figure 9.1 shows the equipment costs for the plant calculated by the ASPEN Process Economics Evaluator.

Equipment by Group Component NameTotal Direct Cost

Equipment Cost

(USD) (USD)Three Phase Separator 3PSEP - Block-flash vessel 299,900.00 98,600.00

3PHX - Block 324,400.00 150,000.00Heat Exchangers EBHX - Block 390,300.00 150,000.00

WHX - Block 1,327,300.00 720,300.00WATPUMP - Block 46,100.00 6,400.00

Pumps EBPUMP - Block 39,400.00 5,500.00EBTPMP - Block 37,300.00 4,700.00

Reactors R1 - Block 177,400.00 46,600.00R2 - Block 234,600.00 72,000.00RBTCOL - Block-cond 58,300.00 9,900.00RBTCOL - Block-cond acc 101,600.00 12,500.00RBTCOL - Block-reb 67,700.00 13,000.00RBTCOL - Block-reflux pump 31,700.00 4,300.00

Distillation Columns RBTCOL - Block-tower 248,500.00 89,500.00REBRECOL - Block-cond 72,500.00 14,500.00REBRECOL - Block-cond acc 104,600.00 13,600.00REBRECOL - Block-reb 139,500.00 59,900.00REBRECOL - Block-reflux pump 38,400.00 5,600.00REBRECOL - Block-tower 427,300.00 187,000.00RPRIMARY - Block-cond 179,600.00 62,600.00RPRIMARY - Block-cond acc 159,500.00 33,100.00

127

RPRIMARY - Block-reb 172,700.00 63,000.00RPRIMARY - Block-reflux pump 61,200.00 10,900.00RPRIMARY - Block-tower 2,376,200.00 1,355,100.00RSTHCOL - Block-cond 115,900.00 35,300.00

Distillation Columns RSTHCOL - Block-cond acc 110,200.00 19,500.00RSTHCOL - Block-reb 78,200.00 15,800.00RSTHCOL - Block-reflux pump 40,100.00 6,100.00RSTHCOL - Block-tower 398,100.00 164,800.00RSTLCOL - Block-cond 67,400.00 12,400.00RSTLCOL - Block-cond acc 98,800.00 15,100.00RSTLCOL - Block-reb 62,200.00 11,700.00RSTLCOL - Block-reflux pump 31,700.00 4,300.00RSTLCOL - Block-tower 1,122,000.00 533,400.00ebtank 660,200.00 455,900.00insulation 14,400.00 14,400.00benzene 146,100.00 62,400.00

Storage Tanks toluene 220,200.00 106,500.00hi-pur styrene 808,800.00 529,400.00Lo-pur styrene 215,800.00 103,800.00heavies 80,500.00 28,400.00

Table 9.1: Process Equipment Costs

128

10. Process Economics

Once the costs of the equipment has been created, the economic analysis can begin. In order to determine the

operating costs, the raw feed stream cost must be determined.  Once these costs are determined, total sales can be

found by inputting the final price of styrene, benzene, and toluene.  With these costs, the gross product sales can be

determined and subsequently the rate of return.  Rate of return shows investors whether they have a safe investment

or a terrible investment.   

Figure 9.1 shows the plot of rate of return vs price margin.  It should be noted that the market cost actually

provides a negative number so there is no margin available. This internal rate of return is the continuously

compounded discounted cash flow rate of return that accounts for the time value of money and also for the nature of

the cash flow.

0.15 0.2 0.25 0.3 0.35 0.4 0.45

-15

-10

-5

0

5

10

15

20

25

30

Rate of Return

Price Margin

Rat

e of

Ret

urn

(%)

Figure 10.1: Rate of Return vs Price Margin

129

A rate of return of 14.93% is achieved at price margin of 0.332 which corresponds to a styrene price of $1.002/lb.

The market price of styrene is $0.85/lb which would leave this plant with a negative rate of return of over 13%.

Figure 9.2 shows the plot of net present value (NPR) at a 15% rate of return and interest rates of 0, 10 and

15 %.

0 2 4 6 8 10 12 14 16 18

-150000000

-100000000

-50000000

0

50000000

100000000

150000000

200000000

Net Present Value

zero %ten %fifteen %

Years

Net P

rese

nt V

alue

($)

Figure 10.2: Net Present Value throughout Plant Life

Year zero designates the time the land is purchased. It can be seen from the graph that during the construction

phase, almost two years, there is no production and subsequently no profit.

Table 9.1 shows the critical investment parameters required for the plant.

FCI 27,813,666.70 dollarsWCI 1,390,683.34 dollarsTCI 29,204,350.04 dollars

Construction Time 24 weeksStartup Time 20.00 weeks

Economic Life of Plant 16.00 years

Total Operating Costs Per Year202,634,890.8

6 dollarsRate of Return 14.93 percentStyrene Price 1.002 $/lb

Annual Total Sales 227,381,283.2 dollars

130

7

Annual Operating Costs187,624,898.9

4 dollarsNet Cash Flow (after taxes and depreciation) 15,185,305.24 dollars

NPV at end of plant life (i=0%)158,574,355.5

8 dollarsNPV at end of plant life (i=10%) 25,207,588.10 dollarsNPV at end of plant life (i=15%) 276,927.91 dollars

Payout Time of Plant (i=0%) 7.71 yearsPayout Time of Plant (i=10%) 11.17 yearsPayout Time of Plant (i=15%) 0.00 years

Current Market Price of Styrene 0.85 dollars/lbCurrent Market Cost of Ethylbenzene 0.67 dollars/lb

Current Price of Benzene 0.57 dollars/lbCurrent Price of Toluene 0.46 dollars/lb

Table 10.1: Critical Investment Parameters

Based on these results, investing in this plant is not recommended. The plant cannot exist due to its

minimal size. Scaling up the plant would result in increased profit margin with the minimal increases in cost. For

example doubling the size of the reactors would only increase the price by a factor of 1.25 instead of 2. This is the

reason why most commercial styrene plants produce billions of pounds of styrene per year.

131

11. Conclusion: Design and Simulation of the Plant

Styrene is produced at 215 million pounds per year, with feeds of 350.8 lb-mol/hr of ethyl-benzene, 6540

lb-mol/hr of water and a recycle stream feed made of 257.45 lb-mol/hr of ethyl-benzene. This produces 254.18 lb-

mol/hr (26473.5 lb/hr) of 99.9 wt.% high purity styrene and 28.21 lb-mol/hr (2983.56 lb/hr) of 99.99 wt.% low

purity styrene. Simulations were completed to complete a pinch analysis on the optimum heat exchange network.

An all utility plant was simulated which gave heating and cooling duties for the traditional plant simulation. The

traditional plant simulation paired two hot and two cold streams together which proved to be an inefficient method

of heating and cooling streams. This designed pinched plant was not simulated but an analysis of the pinch was

completed and a new heat exchange network was designed based on the traditional plant and proven to be more

efficient. Pinch analysis ideally designs the perfect heat exchange network that minimizes the heating and cooling

utilities required.

Radfrac columns were also simulated which sized and modeled the distillation columns. The designed

distillation columns have strengths in simplifications where the newly designed Radfrac columns have a more

accurate output because of the ability to size the trays, tray types, and tray spacing potentially increasing the tray

efficiency.

Reactor Performance

This styrene planted uses two adiabatic packed catalyst bed reactors in series, where 1/3 of the catalyst is in

reactor 1 and the remaining 2/3 of catalyst is in reactor 2. The total conversion of ethyl benzene in reactor 1 is 21%

and 33% in reactor 2, while reactor 2 not only produces styrene but also a series of side reactions. Both reactors

utilize carbon steel with a wall thickness of 1/8 inch with an ellipsoidal cap design. Each of the reactors have 6

inches of insulation which is ideal for it to operate adiabatically at temperature of 1175 F and a pressure of 26 psia.

The heights and diameters of each reactor were set equal to each other. The height and diameter of reactor 1 and

reactor 2 is 10.5 ft and 12.76 ft respectively. The volume of catalysts in each reactor can be seen in table 14-1:

132

Conversion

Vol. of

Catalyst

(ft3)

Diameter

(ft)

Reactor Vol.

w/

Insulation

(ft3)

Catalyst

SplitSelectivity Yield

Reactor 1 21% 269.735 8.977 756.685 0.2929 0.82 47%

Reactor 2 33% 716.241 12.164 1413.461 0.7071 0.82 47%

Table 11.1: Specifications for reactors 1 and 2

Pinch Analysis

Pinch analysis was performed on the plant to determine the minimum utility of heating and cooling

requirements based on an approach of 10°F. Our heating and cooling utilities are 8.29E7 BTU/hr and 1.27E8

BTU/hr respectively. Pinch temperatures were determined to be 223.8°F and 233.8°F. The actual heat exchanger

network based on the simulation yields 10 heat exchangers and one stream splitter. Upon designing the tabular

method heat exchanger network the target heating and cooling values are approached but in order to maintain a

simple design our network yields an additional heating and cooling utility on top of the minimums of 1.59E6

BTU/hr and 1.0E6 BTU/hr.

Column sequencing

The column train is what separates the product stream into the marketable components and undesired by-

products of reaction. Many combinations of columns are possible to attain a specified goal. Heuristics are

guidelines that are required to be followed when choosing the most economically and robust distillation system.

Vacuum columns are necessary for the temperature sensitive separations involving styrene. The construction of

vacuum distillation columns have a high cost and should be as small as possible within the plant. The importance of

minimizing the number of vacuum columns should be taken into a higher consideration over performing difficult

separations last or removing ethyl benzene as a distillate.

Design columns simulated in Aspen determined the number of stages, stream compositions, and reflux ratio

which became the inputs to the RadFrac columns ultimately determining the column size and pressure drop. The

operating temperatures and pressures of the columns within the styrene plant, a 1/8 inch thick carbon steel wall is

acceptable for the material of construction. Columns operating at low pressure differentials were designed to

withstand at minimum of a 15 psi differential. The use of cold water as a cooling utility was found to satisfy the

condenser requirements for the entire distillation system. This ultimately reduced the amount of cooling utility

needed by utilizing one working fluid at a low cost.

133

Equipment Design

Two adiabatic plug flow catalyst packed bed reactors are used to perform the dehydrogenation of ethyl

benzene reaction. Due to the high temperatures reached during the reaction, each of the reactors are lined on the

inside of the carbon steel shell with 6 inches of kaolin fire brick. The kaolin fire brick allows carbon steel to be used

lowering the energy input to the reactor by reducing heat lost to the environment. The thermal resistance the kaolin

fire brick provides temperature drops across the brick to approximately 70°F interface at the steel. At the

temperature and pressure levels of each reactor, 1/8 inch thick carbon steel wall can be used as the shell of the

reactor. The volume was determined by Levenspiel plots. These plots provided the catalyst volume in reactor 1 and

2 at a chosen conversion in each reactor ultimately minimizing total volume requirements for the reaction.

Heat Exchangers were designed in the ASPEN Heat Exchanger Suite where a single pass counter-current

shell and tube configuration was used for the three phase (3PHX) and ethyl-benzene heat exchangers (EBHX).

Estimated from the layout of the traditional plant simulation, the total cost of each of these units is $154,492 and

$109,194.

Centrifugal Pumps were used throughout the plant due to the wide range of operating conditions they can

handle. Location of the pumps was chosen to provide optimum suction head as required by pump specifications.

Pressure drops due to gravity and friction of turbulent fluids flowing through the pipes was modeled in ASPEN.

Storage Tanks were designed out of carbon steel and had a fixed roof with a hatch and ventilation system.

The styrene high purity and low purity storage tanks require cooling for safety concerns associated with styrene’s

flash point. Cooling, along with nitrogen enriched oxygen decreases the polymerization of styrene. The other storage

tanks are blanketed with pure nitrogen to ensure safety of the tanks. The heavies tank requires a heating utility due to

the heavies properties become solid as it reaches ambient temperatures.

Economic Analysis

ASPEN Process Economic Analyzer software was used to determine the designing costs of the styrene

plant equipment. As far as equipment cost the most expensive was the distillation columns at 1.3857 million dollars

and the least expensive were the pumps at 16,600 dollars. An economic analysis was performed for a desired rate of

return (ROR) of 15% and net present value (NPV) at specified interest rates of 0%, 10% and 15%. The current

market value of styrene is $0.85/lb, which only can achieve a ROR value of 8.11%. To obtain a rate of return of

15% the price of styrene would need to be at $1.002/lb. At the final year of plant operation, year 17, at an interest

rate 0% the NPV of the plant is about $158.7 million, which is the largest compared to a 10% interest rate and 15%

rate at $25.2 million and $280 thousand respectively.

134

Feasibility of the Plant Becoming Successful

This plant requires selling styrene at $0.15 above market value to achieve an IRR of 15%. At this required

selling price, this styrene plant is not feasible because of the high selling price compared to the market value

producing a lower rate of return. Due to the unattractive nature of this return, this styrene plant needs to up-scale in

order to become a competitor within the styrene market.

Nomenclature

Ben Benzene mm Millimeters

BTU British Thermal Units NPSHA Net positive suction head actual

Conv Conversion NPSHR Net positive suction head required

dn Diameter of Nozzle NPV Net Present Value

DNC Di-nitro Cresol psia Pounds per square inch absolute

EB Ethyl-benzene psig Pounds per square inch gauge

FCI Fixed Capital Investment PV Vapor Pressure

ft Feet R Rankin

HP Horse power ROR Rate of Return

hr Hour RPM Rotations Per Minute

I Interest Rate s Second

in Inch SS Shell Side

IRR Internal Rate of Return Sty Styrene

KEB Ethyl-benzene overall rate constant TBC Tert-Butyl Catechol

Keq Equilibrium rate constant TCI Total Capital Investment

kg Kilogram Tol Toluene

lbs. Pounds TS Tube Side

LEL Lower Explosive Limit UEL Upper Explosive Limit

Lmin Minimum Length Vac Vacuum

WCI Working Capital Investment

135

References

1. "Catalytic Dehydrogenation of Ethylbenzene to Styrene." Catalytic Dehydrogenation of Ethylbenzene

to Syrene. N.p., n.d. Web. 20 Jan. 2015.

2. Behr, Prof. Dr. Arno. "Styrene Production from Ethylbenzene." Styrene Production from

Ethylbenzene (n.d.): n. pag. Styrene Production from Ethylbenzene. Dortmund University. Web. 19

Jan. 2015.

3. Lee, Won Jae. ETHYLBENZENE DEHYDROGENATION INTO STYRENE: KINETIC MODELING

AND REACTOR SIMULATION (n.d.): n. pag. ETHYLBENZENE DEHYDROGENATION INTO

STYRENE: KINETIC MODELING AND REACTOR SIMULATION. Texas A&M University. Web. 19

Jan. 2015.

4. "Styrene-butadiene Rubber (SBR) | Chemical Compound." Encyclopedia Britannica Online.

Encyclopedia Britannica, n.d. Web. 19 Jan. 2015.

5. "Thermoplastics." Plastipedia: The Plastics Encyclopedia. N.p., n.d. Web. 20 Jan. 2015.

6. "CIEC Promoting Science at the University of York, York, UK."Poly(phenylethene) (Polystyrene).

N.p., n.d. Web. 19 Jan. 2015.

7. "Styrene: 2014 World Market Outlook and Forecast up to 2018." Styrene: 2014 World Market Outlook

and Forecast up to 2018. N.p., n.d. Web. 20 Jan. 2015.

8. “Global Styrene Production Exceeded 26.4 Million Tonnes in 2012”, Merchant Research & Consulting

ltd., Web. 17 January, 2015. http://mcgroup.co.uk/news/20130830/global-styrene-production-

exceeded-264-million-tonnes.html

9. “Styrene: 2014 World Market Outlook and Forecast up to 2018”, Merchant Research & Consulting

ltd., Web 17 January, 2015. http://mcgroup.co.uk/researches/styrene

10. “Products for Processing.” Bricks and Blocks. N.p., n.d. Web. 26 Feb. 2015.

11. http://www.umich.edu/~elements/fogler&gurmen/html/web_mod/membrane/index.htm12. Itoh, N. "AIChE JournalVolume 33, Issue 9, Article First Published Online: 17 JUN 2004." A

Membrane Reactor Using Palladium. National Chemical Laboratory for Industry Tsukuba Science City, 17 June 2004. Web. 24 Apr. 2014.

13. Moustafa, T. M., and Sseh Elnahaie. "An Efficient Palladium Membrane Reactor to Increase the Yield

of Styrene in Ethylbenzene Dehydrogenation." An Efficient Palladium Membrane Reactor to Increase

the Yield of Styrene in Ethylbenzene Dehydrogenation. Journal of Membrane Science, 15 Sept. 2000.

Web. 24 Apr. 2014.

14. Turner, John Stewart. Buoyancy Effects in Fluids. Cambridge: U, 1973. Print

136

Appendix A: Membrane Reactors

A membrane reactor is a relatively common reactor type where the reactor acts as both a separator and a

reactor. The membrane reactor is a subset of the plug flow reactor; the differentiating factor is the presence of

membrane or porous shell within the reactor. The membrane is the barrier that permits specific particles to pass

through it. This selectively is determined by the size of the pores of the membrane. This is designed and it depends

on the substrate to be removed.

There are two main types of membrane reactors that are in use today. The first type is called the inert membrane

reactor with catalyst on feed side while the second is the simply called a catalytic membrane reactor. They are

abbreviated as IMRCF and CMR respectively11. The main feature of the IMRCF reactor type is that it allows the

pellets of the catalyst to flow on the feed side with the reactants11. The CMR on the other hand simply uses a

membrane that has been coated with the catalyst a variant of this is to instead utilize a membrane made of catalyst.

The main advantage of the membrane reactors is that by removing some of the reactants the equilibrium shifts to the

right. According to the Le Chateliers principle this allows for increased production of the target product. The

removal of the unwanted substrate during the reaction also reduces time as well as increasing output.

It has been determined that the styrene plant developed is not economic viability as the market price is below the sell

price needed to sustain a profit/entice investors/ recoup development costs. This option is being explored as a

method of intensely reducing costs and moving towards a viable process of styrene development. The styrene plant

would benefit greatly from the addition of membrane reactor.

Our Plant

The product of interest in the plant is styrene. The other reactants which include methane, benzene, toluene,

naphthalene and ethyl naphthalene that are removed using multiple distillations columns would diffuse through the

membrane as unwanted products during the reaction process. This would greatly reduce the cost of equipment in the

long run.

Currently there are two plug flow reactors in use that perform the adiabatic dehydrogenation of EB at high

temperature and low pressure over iron oxide catalyst. The new system would make use of platinum promoted

Aluminum oxide catalyst pellets instead of iron oxide12. The catalyst will be in mesh granules and the pellets should

be approximately 2m2/g.

The new proposed system would use a palladium membrane as only hydrogen gas can permeate it 12. The

concept to be used differs slightly from that described above. Instead of the membrane allowing unwanted reactants

to pass through it now instead only allows the favored reactant. This serves the same purpose. We have known for

many years that Palladium is only permeable by oxygen but only fairly recently has it been implemented in

137

membrane reactors. Research demonstrates that if implemented effectively the increase in yield is around 15% 13.

This is a significant increase over what was previously obtained.

One issue that must be considered when evaluating this reactor type is cost. They are significantly more

expensive than single purpose reactors. But they still are cheaper than the host of distillation columns that would

need to be put on place to simulate the function of the reactor.

Equilibrium Discussion:

The previous style of processing the styrene was limited thermodynamically, due to the limitations of the

design [2]. The use of the membrane would exceed and surpass those perceived limits.

The main chemical conversion that is supported in the shift to a palladium membrane reactor is the shown below:

The conversion for the reaction of ethyl benzene is originally 33% but with the change to the membrane

reactor this can improve greatly. The equilibrium shift sieves out unwanted products and increases the rate at which

the product is produced. There are many complex rate laws that further describe the phenomena.

Conclusion

The styrene plant is a good candidate for membrane reactor. Outside of the increase in reactor cost

(which is justified in column elimination) there are few negatives. The improved time and conversion would help

improve the effectiveness of the plant.

138

Appendix B: Primary Column Redesign For 99.9%

In the production of styrene, some of the most critical pieces of equipment are the distillation columns. The

separation of components and purification of the product (styrene) allows for the plant to provide a high quality

product. Heuristic analysis suggested the series of columns designed for the traditional plant is economical and

efficient to conduct separation and purify the final products of the styrene plant. The five main components entering

the Primary Column as seen in figure B.1 are ethyl benzene, styrene, benzene, toluene, naphthalene and ethyl

naphthalene. The separation and purification sequence can be seen in figure B.1.

Figure B.1 Best Column Sequence

One of the most difficult separations that occurs in the plant is within the primary column. Because of the

high volume of inlet flow rates and the similar boiling point ranges, this column’s operation can become difficult

without careful evaluation. The heavies and styrene exit at the bottom of the primary column as condensed liquid

while the ethyl benzene, benzene and toluene exit from the top of the primary column as vapor. When designing the

primary column a 90% recovery rate of the styrene from the feed stream was applied for economical reasons. Higher

recovery increases the height of the column and the number of stages. To simulate how the recovery percentage

affects the specifications of the primary column, a 99.9% recovery was applied to the ASPEN simulation. The

comparison and results of the primary column operating at 90% and 99.9% recovery are shown in table B.1.

139

RADFRAC PRIMARY90% Recovery 99.9% Recovery

Number of Trays 55 77Feed Tray Location 12 34Condenser Pressure (psia) 0.580103249 0.580103249Column Pressure Drop (psi) 6.61715038 11.6887954Column Diameter (ft) 24 25Tray Spacing (ft) 2 2Type of Tray Bubble Cap Bubble CapTray Thickness 10 Gauge 10 GaugeOpen Area Ratio 0.12 0.12Weir Height (in) 2 2Weir Length (ft) 15.50 18.07Condenser Utility Cooling Water Cooling WaterCondenser Temperature 98.93 97.89Condenser Duty (btu/hr) -31885241 -36057219Cooling Water Temperature 70 70Re-boiler utility Low- P Steam 100psia Low- P Steam 100psiaRe-boiler Utility Temperature (F) 328 328Re-Boiler Temperature (F) 249.33 128.53Re-boiler Duty (btu/hr) 34198480.90 36771364.70Material of Construction Carbon Steel A-135B Carbon Steel A-135BWall Thickness (Actual) 0.125 0.125Tray Efficiency (%) 100 100

Table B.1 – Comparison of 90% and 99.9% Recovery in Primary Column

The differences is the significant increase in the number of trays in the new 99.9% recovery column which

increased from 55 to 77, the changes in the location of the feed tray from 12 to 34, and the increase in pressure drop

across the trays from 6.62 psi to 11.69 psi. The difference in the pressure drop was expected as the pressure gradient

within the column is one of the driving forces that allow desired vapor to flow to the top and liquid flow to the

bottom. The weir length also changed from 15.50 to 18.07 ft. Within the new column the larger tray sizes allow for

gases to pass through more easily. The heating and cooling duties change from the 90% column to the 99.9%

column by 4 million btu/hr in cooling and 2.5 million in heating. Additional utilities are needed in order to

successfully operate the new column at the higher recovery rate. The temperature of the re-boiler was lowered to

128.53°F from 249.33°F. This indicates that it will require less passes to successfully complete the 99.9% recovery

rate desired.

Economically, the column designed for a 99.9% recovery rate of styrene would require significantly more

capital cost for the size and utility compared to the older 90% recovery design. To apply this recovery in terms of

feed and product streams, the respected column streams are shown in table B.2.

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90% Recovery 99.9% RecoveryTop (lbmol/hr) Bottom (lbmol/hr) Top (lbmol/hr) Bottom (lbmol/hr)

Water 0 0 0 0EB 261.80 0.26 261.80 0.01Styrene 31.86 286.75 16.42 302.21Benzene 18.08 ~ 0 18.08 ~ 0Toluene 30.86 ~ 0 30.86 ~ 0Naphthalene ~ 0 2.42 ~ 0 2.42Ethyl Naphthalene ~ 0 2.42 ~ 0 2.42

Table B.2 – Comparison of Top and Bottom Streams for 90% and 99.9% Recovery in Primary Column

The resulting streams end up with a recovery of over 15 lbmol/hr of styrene in the new primary column

design, which translates into 111,312 lbmol/year in additional production which is equivalent to 11,576,448 lbs of

additional production per year. The amount of extra costs for the 99.9% recovery column would require an

additional 1.8*1010 BTU/yr exceeding any potential benefits gained from the additional styrene produced.

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Appendix C: Separating heavies (naphthalene/ethyl naphthalene) from high purity styrene in a single column

The traditional styrene plant has a series of distillation columns responsible for separating styrene from

heavies (ethyl naphthalene and naphthalene) mainly from the primary column. This bottom stream enters a series of

distillation column separating the high purity styrene and heavies, while the second separates low purity styrene

from the heavies. An assumption that this one singular distillation column is capable of performing a full 99.9%

recovery of styrene under the condition of a 90% recovery from the primary column. A simulation was performed in

the Aspen simulator at the columns RADSTHIP and RSTLOCOL. The new design removed RSTLOCOL and

modified RADSTHIP, denoted as 99.9% Recovery, the results can be seen in table C.1:

RADSTHIP RSTLOCOL 99.9% RecoveryNumber of Trays 19 23 24Feed Tray Location 14 12 13Condenser Pressure (psia) 0.5801 0.5801 0.5801

Column Pressure Drop (psi) 0.6923 1.5923 1.6784

Column Diameter (ft) 19 3 10Tray Spacing (ft) 2 2 2Type of Tray Bubble Cap Bubble Cap Bubble CapCondenser Utility Cooling Water Cooling Water Cooling WaterCondenser Temperature 127.9427 127.9441 127.9427Condenser Duty (btu/hr) -5032950.9 -580106.2 -5712691.4

Cooling Water Temperature 70 70 70

Re-boiler utility Low- P Steam 100psia Low- P Steam 100psia Low- P Steam 100psiaRe-boiler Utility Temperature (F) 328 328 328

Re-Boiler Temperature 133.7340 249.5713 226.0335Re-boiler Duty (btu/hr) 5033807.92 605561.679 5733199.3Material of Construction Carbon Steel Carbon Steel Carbon Steel

Wall Thickness (Actual) 0.125 0.125 0.125Tray Efficiency (%) 100 100 100Tray Thickness 10 Gauge 10 10 GaugeOpen Area Ratio 0.12 0.12 0.12Weir Height (in) 2 2 2Weir Length (ft) 6.4424 3.4774 6.8597

Table C.1: Comparison between using two and one column

Though the simulation results, it is shown that the 99.9% Recovery Column is different compared to the

original designed columns. The largest differences were the number of trays and the pressure drop. Despite these

differences, the column stands to be reasonably feasible. The flow rates from the single column were compatible

with the results of the original design.

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Single Distillation Column First Column Second Column

Top (lbmol/hr)

Bottom (lbmol/hr) Top

(lbmol/hr)

Top (lbmol/hr)

Bottom (lbmol/hr)

Ethyl Benzene 0.0261 3.31E-07 0.02617 3.87E-06 5.40E-10

Styrene 286.4654 0.2845 286.4633 0.2864 0.000286Naphthalene 0.000364 2.4193 0.00241 0.00241 2.4148

Ethly Naphthalene 2.78E-14 2.4202 0.00011 2.21E-06 2.4201

Table C.2 Comparing Output of one and two columns

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Appendix D: Water Remediation

The three phase separator in the plant was designed to successfully separate the three phases of the

incoming stream from the three phase heat exchanger (liquids (oil), heavy liquids (water) and vapor). This separator

was designed to function ideally, where the three phases completely separate and their corresponding exit streams

are 100% of their respected phase. When turbulent flows occur, there are two ways in which the phases may mix.

Either through (1) organic components dissolving and (2) entrainment14. The three phase separators design has a

weir incorporated within where the exiting oil must exceed the weir height making the exiting stream have an

insufficient amount of water and vapor composition unless a surge occurs. The components of the vapor phase are

less dense than those of the liquid phases and exit through the top stream of the vessel. The primary concern of the

three phase separator design was to ensure that the effluent oil stream was high quality. Water may have a tendency

to mix with the other two phases because of its turbulent flow though the separator. Because this effluent stream

needs to be at the highest quality possible, extra treatment may be required to remove the impurities before it is

ready to be re-used. In terms to solve this, both methods described above of phase mixing must be addressed. First,

each dissolved organic component has a different degree of solubility and at the operating condition of the separator

these organic components have different degrees14.

In order to address this problem, both methods of phase mixing had to be addressed, dissolved organic

components and entrained components. In terms of dissolved organic components, each organic has a different

degree of solubility and at the operating conditions of this vessel, these soluble organics will likely not separate from

water14. This method requires several treatments before the water is removed. The first step of method 1 was to

determine the solubility of each component in water and the composition of the feed stream. The next step was

weighing the solubility of each organic component. The vapor phase components were excluded (except for carbon

dioxide) because the contributing mass of these components will be nominal. Each solubility was applied to the total

amount of water within the feed stream and the amount of each dissolved organic component was determined in

terms of mass flow rate composition14. The resulting composition of the water effluent stream is shown in table D.1.

Component Mass Flow Rate lbs/hr

Literaturesolubility (g/L)

Weighted solubility (g/L)

Dissolved Mass Flow Rate Composition (lb/hr)

Water 117820 --- --- 116188Ethyl Benzene 36733.9 0.15 0.1997 23.18

Styrene 29412.06 0.29 0.4598 53.38Benzene 1390.22 1.79 0.1287 14.94Toluene 2794.57 0.52 0.0717 8.33

Carbon Dioxide 1993.19 1.45 0.1395 16.19Naphthalene 305.439 0.03 0.0004 0.05

Ethyl Naphthalene 372.384 0.01 0.0002 0.02

Table D.1 – Three Phases Separator Effluent Dissolved Organic Composition

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The method discussed in which organics will mix and exit with the effluent water stream is through

entrainment. Entrainment occurs when particles of the oil phase become confined within the water phase. Due to the

turbulent flow the water-oil mixed particles are not able to move quickly enough to the oil phase. Again as in

method 1, the vapor phase excluding carbon dioxide is nominal. The next step was to estimate the total amount of

organic components of the oil phase that were entrained in the water phase. In order to complete this, the method

used was the modeling of a gravity three phase separator designed for oil production facilities14. The assumption

needed to analyze this separation was to assume that the oil particles form a ‘tail’ into the extended portion of the

separator. The ‘tail’ volume that exceeds the weir represents the oil volume that exits within the water (liquid) phase

represented in figure D-1.

Figure D.1 – Separation of Oil from Water Phase Representation

Once this was done, the next step was to calculate the volume fraction of oil in the water phase, β, which is

equal to V1/V2, where V1 is the theoretical calculated volume of oil, and V2 which is the actual calculated volume of

oil14. In figure D-1, the theoretical separator is represented by the top and bottom light gray shaded area, while the

actual separator is represented by the just the bottom light gray area. To calculate V1 and V2, parameters of

theoretical length, L1, actual length, L2, tank radius, R, and the angle formed from the bottom of the tank to the top

of the weir height, θ, as shown in figure D.214. Note Figure D.2 is not drawn to scale.

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Figure D.2 – Separation of Oil from Water Phase Representation

With L1 equal to 23.61, L2 equal to 26.46, R equal to 6.02 and θ equal to 98°, V1, V2 and β could be calculated using

equations below14:

V 1=R2 L1(θ−0.5 sin (2 θ )−3 sin (θ )−3θ sin (θ )−sin3(θ)3(1−cos (θ )) ) Eq. D.1

V 2=R2 L(θ−0.5 sin (2θ )−3 sin (θ )−3θ sin (θ )−sin3(θ)3(1−cos (θ )) ) Eq. D.2

β=1−V S1

V S2when L 2>L 1 Eq. D.3

Calculated Values

V1 203,464 ft3

V2 181,580 ft3

β 0.108

Table D.2: Calculated theoretical and actual separator values

With the volume fraction of oil in the water phase calculated, the volumetric flow rate of the total incoming

oil phase components was obtained and the composition of it was determined applying respected densities, then put

in terms of volume fractions14. The volume fraction of oil in the water phase effluent was then applied to the volume

fractions mentioned, resulting in the mass flow rate composition of oil remaining in the water effluent. The results

are shown in table D.3.

Component

Mass Flow Rate lb/hr

Conversion to m3/hr

Volume Fraction

(Oil Phase in Water)

Entrained Composition

(m3/hr)

Composition of Water Effluent

Stream(lb/hr)

Water 116188 --- --- --- 116188Ethyl

Benzene 27584.32 14.45 0.43 1.58 2979.11

Styrene 32852.4 16.41 0.49 1.80 3548.06Benzene 1489.572 0.77 0.023 0.08 160.874Toluene 2858.019 1.50 0.044 0.16 308.666Carbon Dioxide 1993.185 0 0.00 0

Naphthalene 305.523 0.149 0.0043 0.02 32.9965Ethyl

Naphthalene 372.3939 0.167 0.0051 0.02 40.2185

Table D.3 –Three Phases Separator Effluent Entrained Organic Composition

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As shown in tables D.1 and D.3, there are a significant amount of organic products both dissolved and

entrained in the effluent water phase coming from the styrene plant three phase separator, 3PSEP. Choosing the

correct approach to treat the components in the stream 3PWAT is an important step throughout the treatment

process. Treating the 3PWAT stream is important as these organic compounds pose as a risk to the purity of the final

product. Certain treatments will be required to filter out the organic compounds from the 3PWAT before it is

released. The dehydrogenation of ethyl benzene reaction occurring in the reactors will consists of organic

contaminants will consist of ethyl benzene that did not react, styrene, toluene, benzene, naphthalene, and ethyl-

naphthalene.

Granular Activated Carbon (GAC) is a widely utilized method of treating water with chemical substances.

Activated carbon essentially is carbon with increased porosity and in doing so improves its ability to absorb

chemicals. GAC has been utilized in effectively treating water contaminated with organic substances14.

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