CHE231- Production of Benzene (Full Report)

1

PRODUCTION OF BENZENE: GRACE PETROCHEMICALS

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[Document title]

UNIVERSITI TEKNOLOGI MARA KAMPUS PASIR GUDANG JOHOR
FACULTY OF CHEMICAL ENGINEERING
CHE231
INTRODUCTION TO CHEMICAL ENGINEERING PLANT DESIGN

PRODUCTION OF BENZENE

GRACE PETROCHEMICALS

NO.
NAME
POSITION
1.
NAZMI BIN MOHAMED
GENERAL MANAGER
2.
AHMAD SYAMIR BIN MOHAMAD ZAIN
ASSISTANT MANAGER
3.
NURDIANA SYAHIRA BINTI AZWAN
PROCESS ENGINEER
4.
MUHAMMAD AIZRUL BIN ROZALI
PROCESS ENGINEER
5.
KAMELIA BINTI KHAIRUDDIN
SAFETY ENGINEER

CLASS : EH1105C
LECTURER'S NAME : SIR ZAKI BIN SUKOR

Table of Content
Assignment 1- Identify and select process of plant 1-12
Introduction…………………………………………………………………………………………………….
Process Background………………………………………………………………………………………..
Process Selection…………………………………………………………………………………………….
Process Flow Diagram (PFD)……………………………………………………………………………
Conclusion………………………………………………………………………………………………………

Assignment 2- Site Selection 13-15
Details of the site…………………………………………………………………………………………….
Factors affecting……………………………………………………………………………………………..

Assignment 3- Plant Safety Analysis 16-33
Introduction…………………………………………………………………………………………………….
PFD Drawing……………………………………………………………………………………………………
Plant Hazop Analysis……………………………………………………………………………………….
Plant Safety Review…………………………………………………………………………………………

Assignment 4 – Separation Unit Sizing 34-53
Introduction…………………………………………………………………………………………………….
Process Flow Diagram (PFD)………………………………………………..………………………….
Mass Balance Calculations………………………………………………………………………………
Sizing Calculations…………………………………………………………………………………………..
Number of Tray
Feed Tray Location
Diameter of Tray
Height of Separation Column
Conclusion………………………………………………………………………………………………………

Assignment 5- Cost Estimation 54-71
Introduction…………………………………………………………………………………………………...
Process Flow Diagram…………………………………………………………………………………….
Capital Cost……………………………………………………………………………………………………..
Operation Cost………………………………………………………………………………………………..
Production Per Year………………………………………………………………………………………..
Conclusion………………………………………………………………………………………………………

6. Assignment 6- Energy Balance 72-90
Introduction…………………………………………………………………………………………………….
Equipment Drawing………………………………………………………………………………………..
Energy Calculation…………………………………………………………………………………………..
Summary…………………………………………………………………………………………………………

Assignment 1:
Identify and select process of the plant

Introduction

Benzene was isolated for the first time in 1825 by Michael Faraday as a product of whale oil pyrolysis. Benzene is a natural component of crude oil and is one of the most elementary petrochemicals. Benzene is a cyclic hydrocarbon with a molecular formula of C6H6 and its structure has alternating double bonds with hexagon shape.

The molecule of benzene is formed when 6 carbon atoms is joined in a ring with 1 hydrogen atom attached to each carbon. Benzene is a colourless gas, odourless, highly flammable, volatile and has a pleasant, sweet smell. It is widely used in producing chemical products such as detergent, gasoline, kerosene, pesticide, and plastic. Benzene is produced in large quantities from petroleum sources and is used for the chemical synthesis of ethyl benzene, phenol, cyclohexane and other substituted aromatic hydrocarbons. Benzene is also widely used as a solvent in chemical industry and production of numerous chemical.

There are few processes that can be used to produce benzene. One of the processes is catalytic reforming1. Catalytic reforming is the reforming of naphtha with the help of a catalyst which produces petrol with a high octane rating. This process involves dehydrogenation process which the benzene is used to separate from the other aromatic chemical substance by distillation. Also, this process uses platinum chloride as a catalyst and it requires the pressure ranging from 8-50 atm and temperature from 500-525 C.

Next, toluene hydrodealkylation2 is also one of the major processes used to produce benzene. Basically, the aim of this process is to turn toluene into benzene and it is a hydrogen intensive process where toluene is mixed with hydrogen by using a catalyst such as platinum or chromium oxide. Sometimes, the catalyst is being neglected when high temperature is used.

Steam cracking3 is among one of the most common processes used to produce benzene. This process produced ethylene and other alkenes from aliphatic hydrocarbons. Pyrolysis gasoline is the by-product of steam cracking of petroleum by products like paraffin gases, naphtha and gas oils.
Toluene disproportionation4 is also one of the processes to produce benzene. This process has a reaction in which a substance reacts simultaneously to be oxidized and reduced which will produce two different products. Benzene and xylene are described as a demanding product and it can be produce by toluene disproportionation as an alternative to toluene hydrodealkylation. The catalyst used in this process is zeolites.

Lastly, benzene can be produced by the production from coal tar5 .This process improved methods of recovery and purification that coke-oven benzene has been able to withstand the competition of petroleum-derived benzene as well as it has. Production of benzene from coal tar involves recovering benzene from coal tar.
All of these processes will then be discussed in detail in the process background section. This section will focus on how the benzene is produced and description of its process mechanism and how it reacts with the other chemical substances in a block flow diagram.

Physical and Chemical Properties of Benzene
Description
Colourless liquid
Molecular formula
C6H6
Molecular weigh t
78.1 g/mol
Density
o.879g/cm3 @ 25
Boiling point
80.1
Vapor pressure
100 torr @ 26.1
Solubillity
Soluble in ethanol, chloroform, ether, carbon desulfide, acetone, oils and glacial acetic acid, slightly soluble in water
Conversion factor
1 ppm – 3.2 mg/m3 @ 25

Process background
Toluene hydrodealkylation
Toluene hydrodealkylation is a process to convert toluene to benzene. The process is known as hydrogen intensive process and it uses chromium or platinum oxide as a catalyst under a certain condition. When the reaction temperature is too high, the catalyst is usually neglected. Toluene and hydrogen reacts in the reactor with the help of platinum oxide as a catalyst to produce benzene and methane.

C6H5CH3 + H2 C6H6 + CH4

Process condition:
Catalyst
Chromium or platinum oxide
Temperature
500-660 C
Pressure
20-60 atm

Process details:
Hydrogen is heated in a reactor in a presence of catalyst and it will produce benzene and methane.
Toluene is pumped to combine with the stream of mixed hydrogen and fresh hydrogen gas.
Mixture of toluene and hydrogen is pre-heated before it is introduced to the furnace and then it will be sent to the reactor.
The catalytic process occurs at low temperature and offer higher selectivity.
Then, the product will be cooled and next it will proceed to be separated from the unreacted hydrogen in a separator.
Unreacted hydrogen is compressed and it will be recycled back to the feed and reactor.
The product that left the separator are heated before it is being sent to the distillation column which the toluene is separated from the stream and recycled back to the feed.
Further fractionation will separate methane and toluene from benzene product.

Block flow diagram

Figure 1.1 Block flow diagram of the toluene dehydroalkylation of benzene

Catalytic Reforming
Catalytic reforming is a high temperature catalytic process to convert low-octane naphthas into high-octane gasoline blending components called reformates. Reforming involves
isomerization of alkanes
dehydrogenation of cyclohexanes to aromatic hydrocarbons
isomerization and dehydrogenation of alkylcyclopentanes
an dehydrocyclization of alkanes
These processes give high octane blending stock. Reforming process is also a source for feedstock for petrochemical plants. Reformates can be produced with very high concentrations of toluene, benzene, xylene, and other aromatics useful both for gasoline blending and petrochemical processing. Hydrogen, produced from dehydrogenation and dehydrocyclization reactions is separated from the reformate for recycling and use in other refinery processes like hydrodesulfurization.
Operating conditions
Temperature
500 – 530 C
Pressure
20 – 25 kg/ sq.cm @ 5- 45 atm
Catalyst
Platinum or rhenium chloride
Solvents
diethylene glycol or sulfolane

Block Flow Diagram
StabilizerStabilizerC, H GasC, H GasHigher aromaticsHigher aromaticsBenzeneBenzeneAromaticsAromaticsBenzene TowerBenzene TowerSolvent(Recycled)Solvent(Recycled)Aromatics+ SolventsAromatics+ SolventsNon-aromaticsNon-aromaticsExtractorExtractorSome C7, C8, C9 aromaticsSome C7, C8, C9 aromaticsStripperStripperSeparatorSeparatorReactorReactorPreheaterPreheaterNaphtaNaphta
Stabilizer
Stabilizer
C, H Gas
C, H Gas
Higher aromatics
Higher aromatics
Benzene
Benzene
Aromatics
Aromatics
Benzene Tower
Benzene Tower
Solvent
(Recycled)
Solvent
(Recycled)
Aromatics
+ Solvents

Aromatics
+ Solvents

Non-aromatics
Non-aromatics
Extractor
Extractor
Some C7, C8, C9 aromatics
Some C7, C8, C9 aromatics
Stripper
Stripper
Separator
Separator
Reactor
Reactor
Preheater
Preheater
Naphta
Naphta

Hydrogen gas, H2

Feed Preparation TowerFeed Preparation Tower
Feed Preparation Tower
Feed Preparation Tower

Figure 1.2 Block Flow Diagram of Catalytic reforming of Benzene

Process steps
First the naptha is hydrotreated to remove sulfur contaminant.
Recycled hydrogen is then added, mixed and heated.
Conversion of paraffins to aromatic compounds in catalytic reactors and in this reactors platinum or rhenium chloride is acts as catalyst.
In further step a stream is formed which is rich in aromatic compounds.
Then stream is sent to separation section to separate hydrogen and this hydrogen recycled as basic feedstock.
Liquid portion of stream fed to a stabilizer which separates hydrocarbons from liquids.
The liquid is then sent to a debutanizer
Benzene, toluene and xylenes are then extracted using glycol and sulfonate solvents

Process Description
The feed for this process is naptha, which may be straight-run, hydrocracked, thermally cracked, or catalytically cracked. First the naptha is hydrotreated to remove sulfur, which may be present. Recycled hydrogen is then added to it, mixed and heated. This stream is sent to the catalytic reactors where paraffins are converted to aromatic compounds. The catalyst involved in these reactors is usually platinum or rhenium chloride.
The exiting stream, is made up of excess hydrogen and a reformate which is rich in aromatics. This stream is then sent to the separation section of the process. Here, the hydrogen is separated from the liquid product, and recycled back to the initial feed. The liquid product, on the other hand, is fed to a stabilizer. A stabilizer separates the light, volatile hydrocarbons from the liquid product. The liquid is then sent to a debutanizer. Benzene, as well as toluene and xylenes (all called aromatics) are then extracted from this stable reformate.
The catalysts require regeneration after certain time period. Depending on catalyst type and severity of reaction, the cycle time and method of regeneration varies. Some catalytic reforming systems continuously regenerate the catalyst. In some systems, one reactor at a time is taken off-stream for catalyst regeneration, and some facilities regenerate all of the reactors during turnarounds.

Process parameters

The parameters depends on the catalyst, feedstock quality and product specifications. Aromatization reactions are the desirable reaction in the reactor.
The general rules are:

Increase in temperature increases aromatization and hydrocracking reactions.
Increasing pressure reduces aromatization, increases hydrocracking. Lower pressure increases coke formation.
Lower hydrogen recycle ratio increases coke formation.

Pyrolysis Gasoline
Pyrolysis gasoline is the by-product of steam cracking of petroleum by products like paraffin gases, naphthas, gas oils. Pyrolysis gasoline contains 5 percent diolefins. In addition, it also contains 60 percent aromatics compounds, 50 percent of benzene. Different techniques are applied on diolefins to produce benzene, these are:

Distillation of diolefins to olefins
Saturation of olefins to remove sulphur content
Execution of solvent extraction and distillation process to obtain benzene

Production from coal tar
This process improved methods of recovery and purification that coke-oven benzene has been able to withstand the competition of petroleum-derived benzene as well as it has. Production of benzene from coal tar involves recovering benzene from coal tar.
Extraction of lowest boiling point fractions
Applying of caustic soda for the removal of tar acids
Crude oil distillation
Crude oil purification through hydrodealkylation

Advantages and disadvantages

Process
Advantages
Disadvantages
Catalytic Reforming
Large amounts of hydrogen-rich gas will be produced as the by-product.
So, the reforming plant can act as a hydrogen supplier for the hydrodesulfurization plant and hydrocracking plant in the oil refinery, or even to be sold to the market.

There is not sufficient demand for the methylbenzene (toluene) produced
Requires many processes
High capital and operation cost

Hydrodealkylation of toluene
The use of catalyst can be substituted in its absence
Higher temperatures can also be applied with the similar reaction condition.
Process is irreversible
The uses of hydrogen requires explosion-proof plant construction and its use at high reaction conditions need use of chrome steel to avoid embrittlement

Pyrolysis gasoline
The reaction is very simple
No catalyst is required
A lot of heat energy is required for this process, which makes it very energy ineffiecient
The reactants are used up faster than usual

Production from coal tar
The process is very simple and reliable.
Low operating cost
Producing high purity sulphur-free nitration grade benzene at a high yield
Tar have been driven out or released from a carbonaceous material during the initial stage of combustion. Thus, carcinogenic materials are released, which can drastically harm people and environment.
A lot of by products will be produced

Conclusion

As a conclusion, benzene is one of the most highly produced chemicals around the globe. It is one of the most vital organic hydrocarbons (C6H6) in the petrochemical world. Naturally produced benzenes are commonly found in volcanoes and forest fires, whereas the synthetic benzenes are the centre of attraction for companies all over the world, ranging from the production of resins, synthetic fibers, rubber lubricants, dyes, to the more commercial products which are plastics, detergents, drugs and even pesticides. Hence, the most effective method for the production of benzene is highly required, mainly to increase the benzene production yield and also reducing the cost of production by a mile.

The best method for the production of benzene is catalytic reforming. This process utilizes the use of catalyst and high temperature to increase yield of production. This process gives a high octane blending stock. Reforming process is also a source for feedstock for petrochemical plants. Besides producing a reformate yield which consists of very high concentrations of aromatic compounds (benzene, toluene and xylene), this process also gives out high amounts of hydrogen gas. Due to the abundant amount of by-products (hydrogen) produced, the benzene production plant can act as a hydrogen supplier for the hydrodesulfurization plant and hydrocracking plant in the oil refinery, or even to be sold to the market. From this method of production, a proper site selection needs to be highly considerated.

An excellent site requires a high raw materials availability, easy market accessibility, an ample energy availability, a stable climate condition, good transportation facilities, sufficient water supply, proper waste disposal, a sufficient amount of labour supply, the right taxation and legal restrictions, a perfect site characteristics in terms of its physical conditions and its strategic location, a flood and fire protection facility and a good characteristic and facilities of a community that can have a huge impact in terms of supporting the site. Benzene is in an absolute demanding state nowadays, as companies all over the world are trying to squeeze every bits of profit from its production. Therefore, these types of characteristics in selecting a site need to be implemented to have a greater yield of production and a huge reduction in terms of cost, thus profiting a company in a very wide scale.

Assignment 2:
Site Selection

Site Selection

Location : Kerteh Industrial Area
Type of Industries : Petrochemicals
Price + Land Areas : RM4 500 000 /10 acres
Raw Materials Sources : Petronas Penapisan (Terengganu) Sdn. Bhd.
Transportation : 5.5km from Kerteh Port
Utilities : CEFS Respons Kerteh Interplant
Medical Clinic Sdn Bhd.
Onn Machinery Paka Sdn. Bhd.
Stesen Janakuasa TNB Sultan Ismail, Paka
Waste Treatment : Waste Management Paka

Factors Affecting

1. Location and Land Value
Kerteh now has changed into heavy industrial petrochemical centre. It houses the Petronas Penapisan (Terengganu) that supply raw material for our product. Kerteh Industrial area is located around 35 kilometres from Kemaman town. The price of land is RM 4.5 million per 10 acres that is cheaper than Pahang and Johor land. Kerteh is a suitable and strategic place because it quietly far from residential area. So, it can prevent harm from other residents.

2. The Accessibility of Utilities
In terms of utilities such as water supply and electricity, this area is provided with water from Terengganu Waterworks Department capacity of some 75.3 million gallons per day. Moreover, one of the main authorities that also provide water is Syarikat Air Terengganu Sdn Bhd (SATUWATER SDN BHD). For electrical and power supply, Kerteh, Terengganu petrochemical industry's main power source comes from Tenaga Nasional Berhad Distributor Kerteh Terengganu.

3. Availability of Raw Material
It is located near to a chemical plant that produces gasoline, kerosene, distillate fuel oils, residual fuel oils, and lubricants. It takes about 6 minutes or about 3.5 km from plant to transport the product. Therefore, the cost of transportation can be reduced.

Assignment 3:
Plant Safety Analysis
Introduction
Catalytic reforming is a high temperature catalytic process to convert low-octane naphtha into high-octane gasoline blending components called reformates. Reforming process include isomerization of alkanes, dehydrogenation of cyclohexanes to aromatic hydrocarbons, isomerization and dehydrogenation of alkylcyclopentanes and dehydrocyclization of alkanes. Hydrogen, which is the by-product of dehydrogenation process is produce in a significant amount will be fed into other refineries process such as hydrocracking. This process is normally used to produce high aromatic compound especially benzene and the other by-product such as toluene and xylene.
The process begins with the naphtha is being fed to the feed and then it will be pumped which the reaction pressure is 5-45 atm. The hydrogen-rich recycle gas will join the stream, resulting in the liquid-gas mixture being preheated by flowing through the heat exchanger. After that, the product from heat exchanger will be totally vaporized and it will be introduced into the fired heater before entering the fixed bed reactor which the reaction temperature is 495-520 C. Usually, the catalyst used to speed up the reaction is either platinum or aluminium oxide. When the vaporize reactants flow through the reactor, dehydrogenation of naphthene to aromatics occur and the reaction will absorb heat which is endothermic and resulting a large temperature decrease between the inlet and outlet of the reactor. To make sure the reaction temperature as well as the rate of reaction is stable, the vaporized streams need to enter the fired heater to proceed with preheating process again before entering the fixed-bed reactor. The condition inside the reactor will be the same as the first reactor which it will undergo a drastic decrease in temperature so preheating process must be occur again in the third fired heater before it flowing into the third reactor.
As the reaction temperature and the rate of reaction decrease in the third reactor, the amount of heat needed inside the reactor will become smaller. The hot reaction product from the third reactor will be cooled when flowing into the heat exchanger and then proceed to the water cooler for the further cooling process. Then, the stream may flow through the Pressure Controller (PC) before entering the separator column. At the separator column, the reaction temperature is 38 C and hydrogen-rich gas will be separated from the liquid mixture which the hydrogen-rich gas is recycled back to the initial stream. The hydrogen-rich gas will flow through the compressor at the pressure of 5 to 45atm so that the flow is being pressurized into the initial streams.
The liquid mixture from the separator column is directed flow into a fractionating column or known as stabilizer. At the stabilizer, Carbon and Hydrogen gas will exit the upper streams while the bottom product stream contains high-octane liquid reformates that will becomes a component of the refinery's product gasoline which then will be flow through the debutanizer. Some aromatic compound such as heptane, octane and nonane will exit to the upper stream which can be stored or sold to the other company.

After that, the product from debutanizer will continue to flow to the next equipment which is extractor column. Here, the non-aromatic compounds are extracted to the upper stream of the column and the remaining aromatic and the other solvent will be sent to the stripper for a further processing. Stripping is a process where one or more components are removed from a liquid stream by a vapour stream. Therefore, the solvent will be removed from the stripper column while the remaining aromatics were sent to the final equipment process which is benzene tower. Finally, the benzene is produced from the benzene tower and the higher aromatics such as toluene and xylene are also produced as a by-product of this catalytic reforming process.

PFD Drawing
Other aromaticsOther aromatics
Other aromatics
Other aromatics
Plant Hazop Analysis
Item
Process Parameters
Guide Word
Deviation
Possible Causes
Possible Effects
Actions Required

NO
No flow of reactants
1. Pipe blockage
1. No reaction occur
1. Install bypass line

into stabilizer
2. Tube leakage
2. No production
2. Regular scheduled

3. Pump failure
3. Column will dry out
maintenance

4. Valve failure
4. Possible harmful
3. Install flow indicator

concentrations
and flow alarm

4. Emergency shutdown

procedure

LESS
Less flow of reactants
1. Pipe blockage
1. Decreased rate of

into stabilizer
2. Tube leakage
reaction

Flow

3. Pump malfunction,
2. Production will be
maintenance

resulting in reduced
reduced greatly

pumping capacity
and flow alarm

4. Valve malfunction
4. Changes quality and

yield of desired products
procedure

MORE
More flow of reactants
1. Flooding will occur in
1. Install emergency

into stabilizer
resulting in increased
the column
drainage line and backup

pumping capacity
2. Temperature decrease,
valve to drain all excess

resulting in reduced rate
reactants to a certain

of reaction
desired concentrations


maintenance


and flow alarm


procedure

HIGH
High level of
1. Output pipe blockage
1. Overpressure of reflux

mixtures in
drum
maintenance

stabilizer

2 Condensed liquid will
2. Install level indicator

flow back to the stabilizer
and level alarm


Level





Stabilizer

LOW
Low level of
1. Pipe blockage
1. Decreased level in

mixtures in
2. Tube leakage
vessel, resulting in
maintenance

stabilizer
reduced rate of reaction

2. Back flow of material
and level alarm

3. Install backup valve

HIGH
High temperature
1. Failure of cooling system

of mixtures

maintenance

Temperature

in vessel

2. Off-specification of
2. Install temperature

products
indicator

LOW
Low temperature
1. Cooling system

of mixtures
malfunction, resulting in
maintenance

in vessel
the extreme cooling of

mixtures
products
indicator and temperature

3. Less recovery rate of
alarm

distillate products, thus

the conversion towards

desired products will be

less

NO
1. Pipe blockage

into debutanizer
2. Tube leakage
2. No production

3. Pump failure
maintenance

4. Valve failure
4. Possible harmful

concentrations
and flow alarm


Procedure

Flow
LESS
1. Pipe blockage

into debutanizer
2. Tube leakage
reaction

maintenance

reduced greatly
Debutanizer

pumping capacity
and flow alarm


Procedure

MORE

into debutanizer
the column

pumping capacity


of reaction


maintenance


and flow alarm


Procedure

HIGH
High level of mixtures

in debutanizer
drum
maintenance


and level alarm



Level




LOW
Low level of mixtures
1. Pipe blockage

in debutanizer
2. Tube leakage
maintenance


and level alarm


HIGH
High temperature of

mixtures in vessel

maintenance


products

Alarm

Temperature

LOW
Low temperature of
1. Cooling system

mixtures in vessel
maintenance


mixtures
products

alarm




Less

NO
1. Pipe blockage

into stripper
2. Tube leakage
2. No production

3. Pump failure
maintenance

4. Valve failure
4. Possible harmful

concentrations
and flow alarm


Procedure

LESS
1. Pipe blockage

into stripper
2. Tube leakage
reaction

maintenance

reduced greatly

pumping capacity
and flow alarm


Procedure

MORE

into stripper
the column

pumping capacity

Flow


of reaction


maintenance


and flow alarm

HIGH
High level of mixtures

in stripper
drum
maintenance


and level alarm






Level
LOW
Low level of mixtures
1. Pipe blockage
Stripper

in stripper
2. Tube leakage
maintenance



and level alarm


HIGH
High temperature of

mixtures in vessel

maintenance


products

alarm

LOW
Low temperature of
1. Cooling system

mixtures in vessel
maintenance

Temperature


mixtures
products

alarm




less

NO
No flow of cooling
1. Pipe blockage
1. Temperature of mixtures

liquid
2. Tube leakage
will be very high

3. Pump failure
2. Heat transfer rate will
maintenance

4. Valve failure
be very low

3. Desired temperature
and flow alarm

could not be achieved

procedure


5. In worst-case scenario,

explosion might occur

Flow

LESS
Less flow of cooling
1. Pipe blockage

liquid
2. Tube leakage
will be high

3. Pump failure
2. Heat transfer rate wil be
maintenance

very low

and flow alarm


procedure


MORE
More flow of cooling

will be low

pumping capacity

cooling water to a certain

desired amounts

Water

4. Production rate will be
maintenance
Cooler

reduced

and flow alarm


procedure

HIGH
High temperature
1. Water cooling valve
1. Performance of water

liquid
blocked
cooler will be tremendously

degraded
alarm

2. Water cooler can be
2. Install bypass stream

damaged

LOW
Low temperature
1. Fired heater does not

Temperature

work well

alarm


NO
No flow of feed inlet
Pipeline blockage causing

to the reactor
no flow of reactant
2. No production

inspection
Catalytic

HIGH
High flow of feed
1. Pump malfunction
Overflow of naphta
1. Install a backup valve
Reformer
Flow

inlet to the reactor
2. Valve failure

2. Level sensor alarm

is installed

LOW
Low flow of feed
1. Leakage of pipeline
Low level in the
1. Install low level

inlet to the reactor
2. Blockage of pipeline
reactor
indicator

2. Install alarm

3. Do inspection and

Maintenance

LOW
Low temperature
1. Chamber of heat
1. Low rate of reaction
1. Temperature sensor

of the reactor
sources is not
2. Product off spec
2. Temperature sensor

functioning
3. Slow production
alarm

2. Too much flow of

3. Regular inspection

Temperature

coolant

HIGH
High temperature of
Cooling system
1. Might cause fire
1. Install sensor

the reactor
malfunction


alarm

Flow
NO
No flow of the
1. Pipeline broken
No reaction occurs
1. Regular inspection
Fixed Bed

preheated naphta
2. Blockage or leakage of

and maintenance
Reactor

and hydrogen
Pipeline

HIGH
High flow of
1. Flowmeter malfunction
1. Undesired product

preheated naphta
2. Pump malfunction
achieved
indicator

and hydrogen

2. Reactant does not
2. Install alarm

mix well
3. Additional pump

3. Flooding of reactor

Pressure
LOW
Low pressure of
1. Pump malfunction
1. Level of reactant

reactor
2. Opening valve broken
decrease
2. Inspection and

3. Pipeline leakage
2. Product off spec
maintenance of pump

and pipelines

NO
No flow of
Pipe ruptured
No gas is recycled
1. Check maintenance

hydrogen and

procedure and schedule

recycled gas

2. Install alarm

HIGH
High flow of
Valve malfunction
1. Increased pressure in
1. Backup valve

Flow

hydrogen and

equipment
2. Install high flow alarm

recycled gas

2. Can cause failure of
3. Do inspection

Equipments

LESS
Less flow of
Leakage of pipe
Less gas are recycled
1. Install alarm

hydrogen and

2. Maintenance and

recycled gas

regular inspection

Pressure
HIGH
High pressure
Overflow of feed gas
May cause equipment
1. Install high pressure

stream
malfunctioning
alarm

Compressor


LOW
Low pressure
Leakage of pipe
Gas cannot be recycled
1. Install alarm



Flow
LESS
Less flow of cooled
Leakage of pipelines
1. Increase in temperature
Install temperature

gas

of the heat exchanger

2. Low heat transfer
alarm



MORE
More flow of cooled
Control valve broken
Decrease in the
Install automatic valve

gas

temperature of the heat

Heat Exchanger

exchanger

NO
No flow of cooled gas
Blockage of pipe or
Increase in the
Install flowmeter indicator

clogged pipelines
temperature of the heat
and flowmeter alarm

Exchanger

HIGH
Higher temperature
Cooling valve blocked
1. Damage of the heat

exchanger

2. Performance of heat
alarm

exchanger degraded
2. Install bypass stream

Temperature
LOW
Lower temperature
Fired heater does not
1. Reaction condition
Install automated

work well
cannot be achieved
temperature indicator


Deviates

Flow
NO
No flow of naphta and
Blockage or clogged
1. Desired reaction rate
1. Install flowmeter and

hydrogen to the fired
pipelines
is not obtained
flow indicator

heater

2. Waste of energy

LESS
Less flow of naphta and
Leakage of pipe
1. Low yield
1. Regular inspection and


2. Slow production rate
maintenance

heater

2. Work shifts of

maintenance

MORE
More flow of naphta and
Valve malfunction
1. Product off spec



heater

Unachieved

Fired
Temperature
LOW
Low temperature
1. Insufficient supply of
1. Reaction condition
1. Regular maintenance
Heater

reactants

2. Leakage in the heater
2. Low rate of reaction

inlet

HIGH
High temperature
Naphta and hydrogen
1. Product off spec
1. Install high temperature

gas are overheated
2. May affect other
alarm and temperature

equipments performance
indicator

Pressure
HIGH
High pressure of heater
Pressure meter damaged
1. Can cause damage to
1. Install pressure relief

equipment
valve

2. Minor explosion might
2. Install pressure alarm

occur
Sensor

Plant Safety Review

Based on the HAZOP analysis conducted in our plant, the flow of the feed inlet or naphtha inlet contain some problems. We need to install bypass line and regular scheduled inspection need to be performed due to the absence of flow in the inlet stream. Backup valve and level sensor alarm should be installed due to the high flowrate that can cause overflowing of naphtha. Low flowrate of inlet can be handled by installing a low level indicators and alarms. Regular inspection and maintenance should also be operated on our plant equipments.

Next, regular inspection also need to be done due to the low temperature of reactor. To prevent high temperature of reactor, we need to install a sensor and temperature alarm. Additional backup valves should be added to prevent high flow of preheated naphtha and hydrogen into the reactor. Low pressure of reactor can be handled by installing pumps and regular inspection and maintenance of pumps,valves and pipelines need to be done.

The flows in the compressor are either more flow, less flow or even absent. This can be prevented by installing alarms and backup valves. Moreover, regular inspection and maintenance should be performed and the schedule of maintenance and procedure must be regularly examined. To prevent the problem of having high or low pressures in the compressor, we should install alarm and the scheduled maintenance and procedure need to be checked once in a while. In the heat exchanger, the flows can either be more flow, less flow or absent. It can be prevents by installing indicator alarms and automatic valves. To avoid high or low temperature in the heat exchanger, we should install temperature indicator alarms and an emergency bypass stream.

Lastly, the flow in the fired heater can either be more flow, less flow or absent. It can be avoided by installing flowmeters, indicators, backup valves. Regular inspection and maintenance should also be regulated. High and low temperature problems in this equipment can be prevented by installing temperature alarms and indicators. Pressure relief valve and pressure sensor alarm should also be installed to prevent high pressure flow in the fired heater.

In a nutshell, the actions that are proposed above are executed to control the operation for the sake of safety in terms of human lives and environmental welfare. The operation is being controlled to prevent any mishaps from happening, and of course for the optimization of production process. The control equipments for the respective processes are of paramount importance for a chemical plant. That is why a HAZOP Analysis needs to be performed to analyze the possibilities of an accident to happen in terms of their causes, consequences and the course of actions that are required to handle this operation without any problems. Therefore, an operation can be operated smoothly if the plant safety review and possible control actions required are studied beforehand.

Assignment 4:
Separation Unit sizing
Introduction
Benzene is an organic chemical compound with the chemical formula of C6H6. The benzene molecule is composed of 6 carbon atoms joined in a ring with 1 hydrogen atom attached to each. As it contains only carbon and hydrogen atoms, benzene is classed as a hydrocarbon. Its structure also has alternating double bonds with hexagon shape. Benzene is a colourless gas, odourless, highly flammable, volatile and have a pleasant sweet smell. It is widely used in manufacturing chemical product for instance detergent, gasoline, kerosene, pesticide, and plastic. As benzene is used for the chemical synthesis of ethyl benzene, phenol, cyclohexane and other substituted aromatic hydrocarbons, it is produced in large quantities from petroleum sources. Benzene is also widely used as a solvent in chemical industry and production of numerous chemical.

There are few processes that can be used to produce benzene which are catalytic reforming, toluene hydrodealkylation and steam cracking. Our company are specialized in the manufacturing of benzene by using the catalytic reforming process. Catalytic reforming is the reforming of naphtha with the assist of a catalyst which produces petrol with a high octane rating. This process involves dehydrogenation process which the benzene is used to separate from the other aromatic chemical substance by distillation. This process also involves isomerization of alkanes, isomerization and dehydrogenation of alkylcyclopentanes and dehydrocyclization of alkanes. Besides, this process use platinum chloride as a catalyst and it necessitate the pressure ranging from 8-50 atm and temperature from 500-525 C
The process of catalytic reforming begin with naphta as feed which may be straight-run, hydrocracked, thermally cracked, or catalytically cracked is hydrotreated to remove sulphur which may be exist. Then, the liquid feed is pumped up to the reaction pressure 5–45 atm and is joined by a stream of hydrogen-rich recycle gas. The resulting liquid–gas mixture is preheated by flowing through a heat exchanger. As the feed mixture is completely vaporized, it is then heated to the reaction temperature of 495–520 °C before entering the first reactor.
The major reaction is the dehydrogenation of naphthenes to aromatics which is highly endothermic and results in a large temperature decrease between the inlet and outlet of the reactor in the fixed bed of catalyst. To maintain the required reaction temperature and the rate of reaction, the vaporized stream is reheated in the second fired heater before it flows through the second reactor. The temperature is again reduced across the second reactor and the vaporized stream must again be reheated in the third fired heater before it flows through the third reactor. As the vaporized stream proceeds through the three reactors, the reaction rates reduce and the reactors therefore become larger. At the same time, the amount of reheat required between the reactors becomes smaller.
Usually, three reactors are all that is required to provide the desired performance of the catalytic reforming unit. The hot reaction products from the third reactor are partially cooled by flowing through the heat exchanger where the feed to the first reactor is preheated and then flow through a water-cooled heat exchanger before flowing through the pressure controller into the gas separator.
Most of the hydrogen-rich gas from the gas separator vessel returns to the suction of the recycle hydrogen gas compressor and the net production of hydrogen-rich gas from the reforming reactions is exported for use in the other refinery processes that consume hydrogen such as hydrodesulfurization units or a hydrocracker unit.
The exiting stream is made up of excess hydrogen and reformates which is rich in aromatics. This stream is then sent to the separation section of the process. Here, the hydrogen is separated from the liquid product, and recycled back to the initial feed. The liquid product, on the other hand, is fed to a stabilizer. A stabilizer separates the light, volatile hydrocarbons from the liquid product.
The liquid is then sent to a debutanizer to separates some aromatics which are C7, C8, C9 aromatics and aromatic plus solvent mixture. Next, the aromatic and solvent mixture is being extracted by using glycol and sulfonate solvents to separate the non-aromatics from aromatics mixture. After that, the aromatics product from the extractor is then sent to the stripper to separate the solvent from aromatics product. Lastly, the aromatics which are still present in the stripper will enter the distillation column to separate the main product which is benzene as well as the higher aromatics which are toluene and xylene.

Process Flow Diagram
Aromatic + SolventAromatic + SolventHigher AromaticsHigher AromaticsAromaticsAromaticsAromaticsAromatics
Aromatic + Solvent
Aromatic + Solvent
Higher Aromatics
Higher Aromatics
Aromatics
Aromatics
Aromatics
Aromatics

Mass Balance Calculations

n2n2
n2
n2
Y1 - YY1 - Y
Y
1 - Y
Y
1 - Y

BENZENE TOWERBENZENE TOWER
BENZENE TOWER
BENZENE TOWER
n1n1
n1
n1

X1 - XX1 - X
X
1 - X
X
1 - X

n3n3
n3
n3

Z1 - ZZ1 - Z
Z
1 - Z
Z
1 - Z

Figure 4.1: The Flowchart of the Distillation Process of Benzene

Figure 1 above shows the last separator unit which involves the separation of benzene and higher aromatics by using a distillation column. We are going to substitute the higher aromatics to toluene. This is due to the fact that we are aiming at producing benzene as our main product. The higher aromatics are just the by-products of this production. These by-products are then sold to other companies, without further processing by us. That is why we can substitute higher aromatics to toluene due to the fact that higher and higher aromatics have higher and higher boiling points. Thus, for the sake of simplicity and understanding, toluene can be replaced as higher aromatics. Therefore, we want the benzene as our main product, separated from the toluene by separation using a distillation column.

The main aim of performing this mass balance is to identify the amount of feed flowrate required for this separation process to be carried out. Several assumptions are made:

We made an assumption that we are going to operate for only 340 days instead of 365 days. This is because we will be pausing the production operation to perform the shutdown operation for 25 days.

We assumed the annual production of benzene to be 120k metric tonnes per year. Therefore, the calculation of daily benzene production is as follows:

120000 metric tonnes 1 year = 352.95 metric tonnes/day
1 year 340 days

Thus, the molar flowrate, in kmol per hour of daily benzene production is calculated as follows: (Note that the molecular weight of benzene is 78.11 kg/kmol)

352.95 metric tonnes 1000 kg kmol 1 day = 188.28 kmol/hr
1 day 1 metric tonne 78.11 kg 24 hours

Degree of Freedom (D.O.F) Analysis

The Degree of Freedom (D.O.F) Analysis is the method of analyzing systems to see whether they are over or under specified, or if they are well-defined. The first step of performing the mass balance operation is always to make the Degree of Freedom (D.O.F) Analysis . The D.O.F Analysis for the distillation column is as follows:

D.O.F Analysis
D.O.F = Unknowns – Informations
Unknowns = n1,n2,n3,x,y,z 6
Informations =
Mass balances ( Benzene and Toluene) 2
The molar flowrate of benzene production is 188.28 kmol/hr 1

D.O.F = Unknowns - Informations
= 6 - 3
= 3 (UNSOLVABLE)

With the D.O.F of 3, it is very clear that we cannot solve this due to the abundance of unknown informations. Therefore, several assumptions are required to be performed:

We assume that 98 mol% of Benzene from the feed flowrate stream will be present at the overhead stream.

We assume that 96 mol% of Toluene from the feed flowrate stream will be present at the bottom stream.

We assume an equimolar feed molecular composition (molecular composition at feed is 0.5 mol Benzene/mol and 0.5 mol Toluene/mol).

New D.O.F Analysis

Unknowns = n1,n2,n3,x,y,z 6

Informations =

Mass balances (Benzene and Toluene) 2

Process specifications (equimolar feed molecular composition)
33 (98 mol% of Benzene from the feed flowrate stream
3
3
will be present at the overhead stream)
(96 mol% of Toluene from the feed flowrate stream
will be present at the bottom stream)

The molar flowrate of benzene production is 188.28 kmol/hr 1

= 6-6
= 0 (SOLVABLE)

Now, the current D.O.F of the distillation column is 0. Hence, the problem is solvable. The overall mass balance equation is as follows:

Overall balance
n1=n2+n3

Benzene balance
0.98 (0.5) (n1) = 188.28 kmol/h nBO
n1 = 384.24 kmol/h

nBB = (0.02) (0.5) (384.24)
= 3.84 kmol/h

Toluene balance
0.96 (0.5) (384.24) = 184.44 kmol/h nTB

nTO = (0.04) (0.5) (384.24)
= 7.68 kmol/h

n2 = nBO + nTO
= 188.28 kmol/h + 7.68 kmol/h
= 195.96 kmol/h

n3 = nTB + nBB
= 184.44 kmol/h + 3.84 kmol/h
= 188.28 kmol/h

Confirmation: n1= 384.24 kmol/h
n2 + n3 = 195.96 kmol/h + 188.28 kmol/h
= 384.24 kmol/h
Therefore, n1 = n2 +n3 confirmed.

Note:
nBO = Flowrate of Benzene at overhead stream (kmol/h)
nBB = Flowrate of Benzene at bottom stream (kmol/h)
nTO = Flowrate of Toluene at overhead stream (kmol/h)
nTB = Flowrate of Toluene at bottom stream (kmol/h)

Thus, the feed flowrate is 384.24 kmol/h.

Molar Compositions

Overhead Stream:

Y (n2) = nBO
Y (195.96 kmol/h) = 188.28 kmol/h
Y = 0.96 mol Benzene/mol

1 - Y = 1 - 0.96 mol Benzene/mol
= 0.04 mol Toluene/mol
Bottom Stream:

Z (n3) = nBB
Z (188.28 kmol/h) = 3.84 kmol/h
Z= 0.02 mol Benzene/mol

1 - Z = 0.98 mol Toluene/mol

Figure 4.2: The Complete Compositions of Components in Each StreamsFigure 4.2: The Complete Compositions of Components in Each Streams0.02 mol Benzene/mol0.98 mol Toluene/mol0.02 mol Benzene/mol0.98 mol Toluene/mol0.96 mol Benzene/mol0.04 mol Toluene/mol0.96 mol Benzene/mol0.04 mol Toluene/mol0.5 mol Benzene/mol0.5 mol Toluene/mol0.5 mol Benzene/mol0.5 mol Toluene/mol195.96 kmol/h195.96 kmol/h384.24 kmol/h384.24 kmol/hBENZENE TOWERBENZENE TOWER
Figure 4.2: The Complete Compositions of Components in Each Streams
Figure 4.2: The Complete Compositions of Components in Each Streams
0.02 mol Benzene/mol
0.98 mol Toluene/mol
0.5 mol Benzene/mol
0.5 mol Toluene/mol
0.5 mol Benzene/mol
0.5 mol Toluene/mol
195.96 kmol/h
195.96 kmol/h
384.24 kmol/h
384.24 kmol/h
BENZENE TOWER
BENZENE TOWER

188.28 kmol/h188.28 kmol/h
188.28 kmol/h
188.28 kmol/h

Sizing Calculations
Number of Trays
The sizing calculation for the distillation column was first calculated by identifying the number of stages needed for this operation. However, the vapour-liquid equilibrium data for the materials need to be identified first. Thus, the VLE data for benzene-toluene is listed as follow:

Mole Fraction of Benzene in Liquid, XA
Mole Fraction of Benzene in Vapour, YA
0.0200
0.0455
0.1000
0.2090
0.1800
0.3440
0.2600
0.4585
0.3400
0.5555
0.4600
0.6790
0.5400
0.7470
0.6200
0.8054
0.7000
0.8545
0.7800
0.9005
0.8600
0.9405
1.0000
1.0000

Figure 4.3: Vapour-Liquid Equilibrium Data for Benzene-Toluene System

Sources from: http://www.vaxasoftware.com/doc_eduen/qui/vle_en.pdf

The upper part of the column above the feed entrance is called enriching section. Since the entering feed of binary components A and B is enriched in this section, so that the distillate is richer in A than B.
Assumptions made:
Column operates at steady state condition
Constant molal overflow in the column

McCabe- Thiele Method
Calculation for Theoritical Stages/ Trays Equations for Enriching Section

REFLUX(L)XDREFLUX(L)XDDISTILLATE (D)XDDISTILLATE (D)XDCONDENSERCONDENSER
REFLUX
(L)
XD

REFLUX
(L)
XD

DISTILLATE (D)
XD

DISTILLATE (D)
XD

CONDENSER
CONDENSER

Figure 4.4: The Overview in the Distillate Stream of the Distillation Column

To find the equation for the enriching section operating line, the material balance calculation at the enriching section need to be carried out. Thus, the material balance calculation at the enriching section is as follows:

Overall Mass Balance: F = D + B (1)
Benzene Balance: FxF = DxD + BXb (2)

The material balance over the dashed line is as follow:

Overall Mass Balance: Vn+1 = Ln + D (3)
Benzene Balance: Vn+1yn+1 = Lnxn + DxD (4)

Based on the above assumptions made, the value of XF, XD and XB are 0.5, 0.96 and 0.02 respectively.
Since Vn+1 = Ln + D and R = L/D,
yn+1 = LnVn+1Xn+DxDVn+1
yn+1 = LnLn+DXn+DxDLn+D
yn+1 = LnDLnD+DDXn+DDxDLnD+DD
Enriching Section Operating LineEnriching Section Operating Lineyn+1 = RR+1Xn+xDR+1
Enriching Section Operating Line

Enriching Section Operating Line

Since we assumed the reflux ratio = 4,
Thus, the enriching section operating line for this distillation column is

yn+1 = 44+1Xn+0.964+1
yn+1 = 0.8 XN + 0.192

The operating condition for this distillation column is assumed to be at the boiling point of benzene. Hence, the q-line = 1

Figure 4.5: The Equilibrium Data Graph for Benzene – Toluene System

Feed Tray Location

The feed tray location for this distillation column is obtained by looking at the intersection between the Enriching Section Operating Line, Stripping Section Operating Line and q-line. The tray at the intersection between these three lines is the location of the feed tray. Thus, the feed tray location is at tray number 5.

EO = Number of ideal traysNumber of actual tray x 100%

80 = 9.5Number of actual tray x 100%

Number of actual tray = 11.875 ~ 12 trays plus reboiler

Where,
Efficiency: 80%
Ideal or Theoritical Number of Stages: 9.5

Diameter of Distillation Column

The diameter of the column is determined by first identify the allowable maximum velocity based on the cross sectional area of the column by using Sauder-Brown relation.

V = kρL- ρVρV
Where,
ρL = Vapour Density of Benzene
ρV = Liquid Density of Benzene
k = 0.107 ms-1 (at 1 atm)

V = 0.107 876 -2.77 2.77
V = 1.8998 m/s
So, we substituted the value into the equation, where
ρL = 876 kg/m3
ρV = 2.77 kg/m3

Next, the diameter of the distillation column is determined by applying the formula:

DC = 4VwπρVV
Where:
Vw = Mass Flowrate of Benzene at Distillate Stream (kg/s)
ρV = Liquid Density of Benzene
V = Allowable maximum velocity
The mass flowrate of benzene at the distillate stream is as follow:
Molecular weight of benzene = 78.11 kg/kmol
Mass flowrate of benzene at distillate = 195.96 kmol/hr X 0.96 = 188.1216 kmol/hr
Mass flowrate of benzene = 188.1216kmolhr×1 hr3600 s×78.11 kgkmol
= 4.0817 kg B/s
Thus, the diameter of the distillation column is:
DC = 4(4.0817)π(2.77)(1.8998)
DC = 0.99 m

Height of the distillation column
The height of column depends on the plate spacing. The plate spacing can be assumed to be 0.5 m because usually column above 1m diameter spacing between 0.3-0.6m. So the total number of space in distillation column is 11 spaces and the plate spacing between two trays is 0.5m, thus the height of distillation column is 5.5m.

Material

Ferritic Stainless Steel

Actual No of Stages

12

Diameter of Column (m)

0.99m

Height of Column (m)

5.5m

Conclusion

Overhead columnOverhead column1.0m1.0m
Overhead column
Overhead column

1.0m

1.0m

Tray areaTray area-The number of stage is 12 including reboiler-Tray spacing = 0.5m-Tray deck thickness = 2.0 mm-Height of distillation column = 5.5m-Diameter of distillation column = 0.99m-The number of stage is 12 including reboiler-Tray spacing = 0.5m-Tray deck thickness = 2.0 mm-Height of distillation column = 5.5m-Diameter of distillation column = 0.99m
Tray area
Tray area

-The number of stage is 12 including reboiler
-Tray spacing = 0.5m
-Tray deck thickness = 2.0 mm
-Height of distillation column = 5.5m
-Diameter of distillation column = 0.99m

-The number of stage is 12 including reboiler
-Tray spacing = 0.5m
-Tray deck thickness = 2.0 mm
-Height of distillation column = 5.5m
-Diameter of distillation column = 0.99m

Liquid holdup columnLiquid holdup column2.0 m2.0 m
Liquid holdup column
Liquid holdup column

2.0 m

2.0 m

Figure 4.6: The overall overview of the distillation column

In a conclusion, the VLE data for benzene-toluene need to be taken from the internet in order to sketch a graph to determine the theoretical number of stages of benzene. Firstly, the height of the distillation column will be determined from the plate spacing and diameter of the column. This is an important step because it will determine the overall height of the distillation column including overhead area and liquid holding area.

Also, plate spacing depends on the diameter of the distillation column. For example, if the diameter is above 1m diameter, the spacing can be estimate to be 0.3 to 0.6m. Since the diameter is almost 1m, 0.5m can be taken as an assumption for the spacing distance.

In figure 6, the equilibrium data of benzene-toluene was drawn to identify the number of theoretical stage and actual number of stages by using the McCabe-Thiele Method. The diameter and height of the distillation column can be calculated by referring to the number of actual stage.

Next, the assumption had to be made to the overhead column which is 1.5m so that it is easier for repair or maintenance purpose. Last but not least, the liquid holdup section is been assumed to 2.0m so that it can hold a low boiling point liquid that flow from upstream to downstream.

Thus, the total height of distillation column can be calculated which is 9m.The material used for the construction of distillation column is Ferritic Stainless Steel. The main reason to choose ferritic stainless steel is because it is less expensive, resistant to corrosion and contains a lot of chromium. Due to its structure, it contains better engineering properties compared to other stainless steel.

Parameter
Units
Function

Overhead area

1.0m

For repair or maintenance purpose
Tray area
Tray spacing

Tray deck thickness

Sieve trays

0.5m

2.0mm

Provide better economics trade-off between column height and diameter
the cost usually expensive when the tray is thicker
easy for worker to detach and assembles for maintenance

Low in cost
The efficiency is quite high ( 0.7-0.9 )
The cost for maintenance is low
Low fouling tendency

Liquid holdup section

2.0m

Static head to avoid pump cavitation
For piping and support structure
To provide driving force for thermosiphon reboiler

Assignment 6:
Cost Estimation

Introduction
Benzene is one of the important organic chemical compound with the chemical formula C6H6. The benzene molecule is composed of 6 carbon atoms joined in a ring with 1 hydrogen atom attached to each. As it contains only carbon and hydrogen atoms, benzene is classified as an aromatic hydrocarbon.

Benzene is a natural constituent of crude oil and is one of the elementary petrochemicals. Benzene is a colorless and highly flammable liquid with a sweet smell, and is the same as the aroma around petrol stations. It is used primarily as a precursor to the manufacture of chemicals with more complex structure, such as ethylbenzene and cumene. As benzene has a high octane number, it is an important component of gasoline.

Initially, there are few common processes involved in the manufacturing of benzene industrially but the main focus for the production of benzene would probably be from naphtha as the percentage yield of benzene from naphtha is much more higher compared to other processes.

The process of catalytic reforming begin with naphta as feed which may be straight-run, hydrocracked, thermally cracked, or catalytically cracked is hydrotreated to remove sulphur which may be exist. Then, the liquid feed is pumped up to the reaction pressure 5–45 atm and is joined by a stream of hydrogen-rich recycle gas. The resulting liquid–gas mixture is preheated by flowing through a heat exchanger. As the feed mixture is completely vaporized, it is then heated to the reaction temperature of 495–520 °C before entering the first reactor.

The major reaction is the dehydrogenation of naphthenes to aromatics which is highly endothermic and results in a large temperature decrease between the inlet and outlet of the reactor in the fixed bed of catalyst. To maintain the required reaction temperature and the rate of reaction, the vaporized stream is reheated in the second fired heater before it flows through the second reactor. The temperature again reduce across the second reactor and the vaporized stream must again be reheated in the third fired heater before it flows through the third reactor. As the vaporized stream proceeds through the three reactors, the reaction rates reduces and the reactors therefore become larger. At the same time, the amount of reheat required between the reactors becomes smaller.

Usually, three reactors are all that is required to provide the desired performance of the catalytic reforming unit. The hot reaction products from the third reactor are partially cooled by flowing through the heat exchanger where the feed to the first reactor is preheated and then flow through a water-cooled heat exchanger before flowing through the pressure controller into the gas separator.
Most of the hydrogen-rich gas from the gas separator vessel returns to the suction of the recycle hydrogen gas compressor and the net production of hydrogen-rich gas from the reforming reactions is exported for use in the other refinery processes that consume hydrogen such as hydrodesulfurization units or a hydrocracker unit.

The exiting stream is made up of excess hydrogen and reformates which is rich in aromatics. This stream is then sent to the separation section of the process. Here, the hydrogen is separated from the liquid product, and recycled back to the initial feed. The liquid product, on the other hand, is fed to a stabilizer. A stabilizer separates the light, volatile hydrocarbons from the liquid product.
The liquid is then sent to a debutanizer to separates some aromatics which are C7, C8, C9 aromatics and aromatic plus solvent mixture. Next, the aromatic and solvent mixture is being extracted by using glycol and sulfonate solvents to separate the non-aromatics from aromatics mixture. After that, the aromatics product from the extractor is then sent to the stripper to separate the solvent from aromatics product. Lastly, the aromatics which is still present in the stripper will enter the distillation column to separate the main product which is benzene as well as the higher aromatics which are toluene and xylenes.

Process Flow Diagram of Catalytic Reforming of Benzene

Capital Cost

Amount
Price (RM)
Major Equipments:

Industrial Distillation Column
5 units
1,058,374.85
Fixed-Bed Reactor
3 units
2,540,099.61
Continuous Industrial Furnace
3 units
205,000.00
Centrifugal Pump
1 unit
1,730.20
Dynamic Scraped Surface Heat Exchanger
1 unit
254,009.96

High Efficiency Gas Separator
1 unit
338,679.95
High Pressure Solenoid Valve
1 unit
328.00
Hydrogen Gas Compressor
1 unit
4,081.10
Installation Costs:

15% of total equipment cost
16 units
622,859.05
Layout of the Plant:

Warehouse
1 unit
138,211.40
Workshop
1 unit
285,680.00
Laboratory
1 unit
2,040.55
Cooling towers
3 units
122,400.00
Total 5, 573, 494.67

Notes:

Exchange rates : 1 USD = 4.08 MYR.

Price of Land : RM 10.33 per square feet.

Clearance of Land : RM 270 per acre.

Price Per Unit Of Major Equipment:

Industrial Distillation Column : RM 211,674.97
Fixed-Bed Reactor : RM 846,699.87
Continuous Industrial Furnace : RM 68,333.33
Centrifugal Pump :RM 1,730.20
Dynamic Scraped Surface Heat Exchanger :RM 254,009.96
High Efficiency Gas Separator :RM 338,679.95
High Pressure Solenoid Valve :RM 328.00
Hydrogen Gas Compressor :RM 4,081.10
Cooling tower :RM 40800

Price Based on the Layout of Plant:

Warehouse : RM 734.40 per square meter
Workshop : RM 285.68 per square meter
Laboratory : RM 2040.55 per square feet

References: www.alibaba.com
www.landserve.com.my
www.xe.com/currencyconverter

Operational Costs

Salary and Wages

No Of Personnel

Net Salary

Epf (11%)

Socso (2.5%)
Total Monthly Salary Based On Number Of Personnel

Monthly Gross Salary

Yearly Salary

RM
RM
RM

Administration

A
B
C
A+B+C*N
A+B+C

Manager

1

12,000.00

1,320.00

300.00

13,620.00

13,620.00

163,440.00

Executive

2

4,000.00

440.00

100.00

9,080.00

4,540.00

108,960.00

Assistant

2

2,800.00

308.00

70.00

6,356.00

3,178.00

76,272.00

Human Resources

Manager

1

8,000.00

880.00

200.00

9,080.00

9,080.00

108,960.00

Executive

1

4,000.00

440.00

100.00

4,540.00

4,540.00

54,480.00

Assistant

1

2,800.00

308.00

70.00

3,178.00

3,178.00

38,136.00

Finance

Accountant

2

4,000.00

440.00

100.00

11,350.00

5,675.00

136,200.00

Sales Marketing

Senior Executive

1

4,200.00

462.00

105.00

4,767.00

4,767.00

57,204.00
Junior Executive

1

1,500.00

165.00

37.50

1,702.50

1,702.50

20,430
Logistic

Manager

1

7,000.00

770.00

175.00

7,945.00

7,945.00

95,340.00

Executive

1

4,000.00

440.00

100.00

4,540.00

4,540.00

54,480.00

Officer

1

3,000.00

330.00

75.00

3,405.00

3,405.00

40,860.00

Assistant

2

1,500.00

165.00

37.50

3,405.00

1,702.50

40,860.00
Quality Control

Chemist

4

3,000.00

330.00

75.00

13,620.00

3,405.00

163,440.00

Lab Assistant

6

1,700.00

187.00

42.50

11,577.00

1,929.50

138,924.00
Maintenance

Senior Engineer

1

7,000.00

770.00

175.00

7,945.00

7,945.00

95,340.00
Senior Technician

1

2,200.00

242.00

55.00

2,497.00

2,497.00

29,964.00

Technician

3

1,600.00

176.00

40.00

5,448.00

1,816.00

65,376.00

Warehouse

Manager

1

6,000.00

660.00

150.00

6,810.00

6,810.00

81,720.00

Executive

1

4,000.00

440.00

100.00

4,540.00

4,540.00

54,480.00

Officer

4

3,000.00

330.00

75.00

13,620.00

3,405.00

163,440.00

Supervisor

4

2,500.00

275.00

62.50

11,350.00

2,837.50

136,200.00

Loading Team

10

1,400.00

154.00

35.00

15,890.00

1,589.00

190,680.00
Security

Manager

1

6,000.00

660.00

150.00

6,810.00

6,810.00

81,720.00

Security Guard

8

1,000.00

110.00

25.00

9,080.00

1,135.00

13,620.00
Safety

Safety Officer

1

4,000.00

440.00

100.00

4,544.35

4,544.35

54,532.20
Production (Process)

Manager

1

10,000.00

1,100.00

250.00

11,350.00

11,350.00

136,200.00

Supervisor

3

3,500.00

385.00

87.50

11,917.50

3,972.50

143,010.00

Board Man

6

2,000.00

220.00

50.00

13,620.00

2,270.00

163,440.00
Senior Technician

8

3,000

330.00

75.00

27,240.00

3,405.00

326,880.00
Senior Engineer

3

6,000.00

660.00

150.00

20,430.00

6,810.00

245,160.00
Junior Engineer

8

1,700.00

187.00

42.50

15,436.00

1,929.50

185,232.00

Technician

12

1,000.00

110.00

25.00

13,620.00

1,135.00

163,440.00

Total
Salary
RM1983314.20

2. Raw Materials
Substances
Price

Naphtha

= 100 000 000 kg x 0.51 USD x RM4.08
year 1 kg 1 USD
= RM 208 080 000
Total of raw material
RM 208 080 000

Note:
Total Number Of Personnel : 103 Personnel
Total Shift For Production Process, Quality Control, Warehouse: 3 Shift (4 Teams)

Production Process
Supervisor : 3 Personnel
Boardman : 6 Personnel
Senior Technician : 8 Personnel
Junior Engineer : 8 Personnel
Senior Engineer : 3 Personnel
Technician : 12 Personnel

Quality Control
Chemist : 4 Personnel
Lab Assistant : 6 Personnel

Warehouse
Manager : 1 Personnel
Executive : 1 Personnel
Officer : 4 Personnel
Supervisor : 4 Personnel
Loading Team : 10 Personnel
Notes:
Exchange rates : 1USD= RM 4.08

Particular
Price per year (RM)
Salaries and wages
1 983 314.20
Raw material
208 080 000
Total operational cost
210, 063,314.2

Annual Production Cost

Table 6.1 Typical factors for estimation of project fixed capital cost

Total physical plant cost (PPC) = PCE ( 1 + f1 + f2 + f3 +f4 )
= RM 1 766 337.38 + RM 1 766 337.38 (1 + 0.4 + 0.7 +
0.2 + 0.1)
= RM 6, 005, 547.092

Fixed Capital = PPC ( 1 + f10 + f12 )
= RM 6, 005, 547.092 ( 1+ 0.3 + 0.1 )
= RM 8, 407, 765.929

No.
Variable Costs
Typical Values
Total (RM)
1
Raw Materials
from flow sheet
208 080 000
2
Miscellaneous Materials
10% of item (5)
69 240. 426
3
Utilities
From flow sheet
416 160 000
4
Shipping and Packaging
Usually negligible
-

Sub Total A = ( 208 080 000 + 69 240.426 + 416 160 000 )
= RM 624 309 240.4

Fixed Costs

5
Maintenance
5- 10% of fixed capital
840, 776.6
6
Operating Labour
5% from fixed capital
420, 388.3
7
Laboratory costs
30% from item (6)
126, 116.49
8
Supervision
-
-
9
Plant overhead
50% from item (6)
210, 194.15
10
Capital charges
840, 776.6
11
Insurance
84, 077.66
12
Local taxes
Usually negligible
-
13
Royalties
-
-

Sub Total B = (840, 776.6 + 420, 388.3 + 126,116.49 + 210,194.15 +
840, 776.6 + 84, 077.66)
= RM 2, 522, 329.8

Direct Production Costs = A + B
= RM 624 309 240.4 + RM 2, 522, 329.8
= RM 626, 831, 570.2

14
Sales Expenses
20-30% of direct production costs
Negligible
15
General Overheads
Negligible
16
Miscellaneous Materials
Negligible

Sub Total C = 0
Annual Production Costs = A + B + C
= RM 624 309 240.4 + RM 2, 522, 329.8 + RM 0
= RM 626, 831, 570.2

Production Costs (RM/kg) = Annual Production Costs
Annual Production Rate
= RM 626, 831, 570.2
120 000 000 kg
= RM 5.22 /kg

Rate of Production Per Year
Based on the capital and total operational cost that was been calculated, the total amount of spent of the capital and total operational cost is RM 5, 573, 494.67 and RM 210, 063,314.2 per year respectively. The naphtha price for this month is RM 2350.0392 per metric tonnes as comparison to the previous month which is RM 2.194.86 metric tonnes. Also, our company require 120 000 metric tonnes of naphtha per year which is the raw material used in the production of benzene. After doing some calculation, our yearly cost of naphtha are estimate in Calculation 1 and probably would be around RM 282 004 800. For every 120 000 metric tonnes of naphtha, the amount of benzene, toluene and xylene that can be produced per metric tonnes is 60 000, 37 000 and 26 500 metric tonnes respectively. Based on the market price, a price per metric tonnes of benzene, toluene and xylene is RM 3 060.41, RM 2 049.55 and RM 3 051.60 respectively. Hence, the amount of money or profit that we can generate from the production of BTX compound is RM 340 325 350. So, our company has obtained the net profit after deducting the amount of money produced from our product with the amount of money spent for the raw material per year and the calculation can be shown in Calculation 2.

Calculation 1:
Cost of naphta per year = Raw material of naphta X naphta price
= 120 000 metric tonnes of naphta X RM 2 350.04 metric
= RM 282 004 800
Calculation 2:
Net Profit = RM 340 325 350 – RM 282 004 800
= RM 58 320 550 per year

The amount of investment of our company is obtained from the total capital cost and total operational cost. Hence, the total amount of our investment is RM 215, 636,808.9. Thus, our company have calculated the return of investment by following the formula below:

Return on Investment = Net cash flow
Total Investment
= RM 58 320 550
RM 215, 636,808.9
= 27%

Thus, after doing the calculation, we have identified that the Return on Investment (ROI) for our company is 33%. To conclude, we could eventually make half profit for our company for three years and the amount of money spent for the investment could also be recover as well.

Payback time = Total Investment
Average annual cash flow
= (RM 215, 636,808.9)
RM 58 320 550
= 3.69
= 4 years

Graph above depicts the payback time of the production which is on the fifth year. Based on the formula calculated earlier, the payback time calculated is to be about 4 to five years of the production. Hence, assuming the payback time it appears to be in the fifth year of the production.

Conclusion
As a conclusion, our plant basically composed of three departments involved in the production process which is the production team, quality control and warehouses with shift hours working. This is so because benzene typically operates for 24 hours. Thus, as to ensure the smooth sail of our plant processes, there would be three shift working with four teams from each department. By this scheduled working hours, each personnel working would have enough rest and not be stressed out by the extended working hours.

While the value of ROI that we obtained, our company would make profit within three years time. On the upcoming years after the profits are made, the amount of profit will appear to be inconsistent, because of certain factors like maintenance of the equipment, increment in salary or recruiting new ones, purchasing of newer and improved equipment, and the change in price of raw materials and products. Hence, though we make profit we still have to always be prepared for the worse case scenario and provide additional amount of money just in case our plant meets those factors above.

References
https://www.platts.com/news-feature/2014/petrochemicals/pgpi/paraxylene
http://markets.businessinsider.com/commodities/naphtha
https://www.platts.com/news-feature/2014/petrochemicals/pgpi/toluene
https://www.platts.com/news-feature/2014/petrochemicals/pgpi/benzene
www.alibaba.com
https://www.payscale.com

Assignment 7:
Energy Balance
Introduction

Benzene is one of the man-made aromatic hydrocarbon compounds which is a natural constituent of crude oil and is one of the elementary petrochemicals. It is used primarily as a precursor to the manufacture of chemicals with more complex structure, such as ethylbenzene and cumene, of which billions of kilograms are produced. Benzene is an important organic chemical compound with the chemical formula C6H6.Benzene is a colorless and highly flammable liquid with a sweet smell, and is responsible for the aroma around petrol stations. As benzene has a high octane number, it is an important component of gasoline. As benzene is a human carcinogen, most non-industrial applications have been limited.

Benzene is widely produced and commercially used in most of the industries present nowadays due to its numerous commercial and industrial applications. Benzene is used mainly as an intermediate to make other chemicals, above all ethylbenzene, cumene, cyclohexane, nitrobenzene, and alkylbenzene. More than half of the entire benzene production is processed into ethylbenzene, a precursor to styrene, which is used to make polymers and plastics like polystyrene and EPS. Some 20% of the benzene production is used to manufacture cumene, which is needed to produce phenol and acetone for resins and adhesives. Cyclohexane consumes 10% of the world's benzene production and it is primarily used in the manufacture of nylon fibers. These fibers are then processed into textiles and engineering plastics. Smaller amounts of benzene are used to make some types of rubbers, lubricants, dyes, detergents, drugs, explosives, and pesticides.The biggest consumer country of benzene was China, followed by the USA. Benzene production is currently expanding in the Middle East and in Africa, whereas production capacities in Western Europe and North America are stagnating. As a gasoline (petrol) additive, benzene increases the octane rating and reduces knocking. As a consequence, gasoline often contained several percents of benzene before tetraethyl lead replaced it as the most widely used antiknock additive. However, with the global phaseout of leaded gasoline, benzene has made a comeback as a gasoline additive in some nations.

Benzene can be produced from a few ways. These include toluene hydrodealkalation, toluene disproportionation, steam cracking and catalytic reforming. Other processes that produce benzene include the reduction of phenol and halobenzenes with metals, decarboxylation of benzoic acid and its salts to benzene, reaction of diazonium compound with hypophosphorus acid aniline and trimerization of acetylene. However, these production processes have no commercial significance. The most efficient method of producing benzene, with relatively higher cost is catalytic reforming.
Toluene hydrodealkylation process converts toluene to benzene. In this hydrogen-intensive process, toluene is mixed with hydrogen, then passed over a platinum oxide catalyst at 500–600 °C and 40–60 atm pressure. Sometimes, higher temperatures are used instead of a catalyst (at the similar reaction condition). Under these conditions, toluene undergoes dealkylation to benzene and methane. This irreversible reaction is accompanied by an equilibrium side reaction, that produces biphenyl (aka diphenyl) at higher temperatures. If the raw material stream contains much non-aromatic components (paraffins or naphthenes), those are likely decomposed to lower hydrocarbons such as methane, which increases the consumption of hydrogen. Sometimes, xylenes and heavier aromatics are used in place of toluene, with similar efficiency.

Toluene disproportionation (TDP) might be a more favoured alternative to the toluene hydrodealkylation if a chemical complex has similar demands for both benzene and xylene, . In the broad sense, 2 toluene molecules are reacted and the methyl groups rearranged from one toluene molecule to the other, yielding one benzene molecule and one xylene molecule. Given that demand for para-xylene (p-xylene) substantially exceeds demand for other xylene isomers, a refinement of the TDP process called Selective Toluene Disproportionation (STDP) may be used. In the steam cracking process, ethylene and other alkenes are produced from aliphatic hydrocarbons. Depending on the feedstock used to produce the olefins, steam cracking can produce a benzene-rich liquid by-product called pyrolysis gasoline. Pyrolysis gasoline can be blended with other hydrocarbons as a gasoline additive, or routed through an extraction process to recover the BTX aromatics (benzene, toluene and xylenes).

In catalytic reforming, a mixture of hydrocarbons with boiling points between 60–200 °C is blended with hydrogen gas and then exposed to a bifunctional platinum chloride or rhenium chloride catalyst at 500–525 °C and pressures ranging from 8–50 atm. The aromatic products of the reaction are then separated from the reformate by extraction with any one of a number of solvents, including diethylene glycol or sulfolane. Benzene is then separated from the other aromatics by distillation. The extraction step of aromatics from the reformate is designed to produce aromatics with lowest non-aromatic components. This catalytic reforming process are made up of several important process units which are reactors that consist of fixed bed catalyst, a furnace to provide a convenient reaction condition for the process, a cooling system, a gas separator and a hydrogen-gas recycle system. The process of this catalytic reforming begins when the stream of naphtha are merged together with a stream of recycled hydrogen. The mixture then will be preheated first before being further projected into the furnace where the mixture of naphtha and hydrogen stream is heated to a desired reaction condition. The exit stream of the mixture will eventually enter a series of reactor consist of fixed-bed reactor at a temperature of 495-525°C. The effluent product from the last reactor will next be cooled down by a cooler and transferred into a gas separator where the separation process of the hydrogen and the liquid reformate takes place. The top product from the distillation process which mainly consists of hydrogen are being compressed and recycled back to the process. The reformates are further sent to a stabilizer (to remove C and H compunds as overhead distillates), a debutanizer (to remove higher hydrocarbons), an extractor (to remove non-aromatics), a stripper (to remove the solvent) and a fractionator (which separate benzenes from the aromatics, mainly toluene).

One of the most vital processes involve is the recovery of benzene process. During the catalytic reforming process of the naphtha, high percentage of aromatic compounds such as benzene, toluene and xylene as well as some non-aromatic components are present. In this case, the best separation process for separating benzene and other aromatics is the distillation process. This is because there is only one aromatic compound (benzene) that need to be recovered at the end at its purest form. Thus, using distillation column to operate the distillation process might be the most optimum route towards the separation of benzene from the aromatics, as its operating cost is much cheaper than other separation processes. Therefore, it is most economical to use this method. In the benzene recovery process, the liquid reformate is first introduced to a distillation column in which liquefied petroleum gas (LPG) is separated from the mixture of liquid reformate which consists of the aromatics (benzene, toluene and xylene). Xylene leaves the bottom product while the overhead product which consists of benzene and toluene is projected to a second distillation column where benzene (main product) and toluene are finally separated.

From the calculated plant costing evaluation, the desired amount of benzene needed to be produced is 352.95 metric tonnes per day. The values obtained from this plant costing evaluation are of paramount importance when it comes to it being used in the separation unit sizing process. Based on the values obtained, all informations, (which include the number of trays required for the distillation column, feed tray location, the diameter of the tray and the height of the distillation column) are to be required for the most optimum benzene production.
In a nutshell, all the procedures involved finally led to the calculation of the energy required to be supplied at the separation unit. The main objective of the separation unit is basically to separate our main product which is benzene from toluene with respect towards their boiling points. The separation unit involved is assumed to be operated at a continuous condition and at atmospheric pressure. Based on the previous material balance calculation, the following aspects must be taken into considerations when designing the separation unit:

Overhead Area = 1.0 metre
Theoretical stages = 9.5 trays
Spacing between trays = 0.5 metres
Height of distillation column = 5.5 metres
Diameter of distillation column = 0.99 metres

The distillation column is also assumed to be operated at 80 % efficiency, with the total stages or trays involved in the distillation column being 12 trays.
Equipment drawing

Notes:
Area of the blue shading = Condenser
Area of the red shading = Reboiler

Energy balance calculation

Figure 1 above shows the flowchart of the distillation process between the binary mixture of benzene and toluene in the distillation column. First of all, it is crucial to determine the molar flow rate and composition of benzene and toluene in each system. The composition of the reboiler and condenser stream must also be identified in order for the energy balance calculation for both stream can be take place. Next, it is also important to perform the degree of analysis at both reboiler and condenser stream. This is done so as to identify whether there is enough information given in order to solve the material balance calculation. Hence, the degree of freedom analysis for reboiler and condenser stream is as follow:
DOF = 3 (unknown) – 2 (Independent Equation)
= 1-1 (Process Specification)
= 0 (Solvable for the stream of reboiler)
DOF = 3 (unknown) – 2 (Independent Equation)
= 1-1 (Process Specification)
= 0 (Solvable for the stream of condenser)
The composition of benzene and toluene at the first stream of the distillate is assumed to be at 0.70 kmol B/kmol and 0.30 kmol T/kmol respectively. It is because the distillation column is operated at only 80% efficiency and thus that is one of the reasons why reflux is needed in order to obtain the highest purity of the desired product out from the distillate stream. From the previous separation unit sizing, the value of reflux ratio is found to be at 4 and the D molar flowrate is 195.96kmol/h. Thus, the value of n2 at the condenser stream could be determined by applying the following formula:

R = LD
L = R × D
L = 4 × 195.96 kmol/h
L = 783.84 kmol/h

Since the value of L = n2. Hence, the molar flowrate of n2 is 783.84kmol/h. The molar flowrate of n1 is obtained from the following formula:

n1 = n2 + 195.96 kmol/h
n1 = 783.84 kmol/h + 195.96kmol/h
n1 = 979.80 kmol/h

After the molar flowrate at the reflux stream and the first stream of the distillate has been determined, the composition of toluene and benzene at the reflux stream is as follows:

Benzene Balance:
Input = Output
n1y1 = 0.96(195.96 kmol/h) + n2y2
(0.7)(979.80) = (0.96)(195.96) + 783.84y2
685.86= 188.1216 + 783.84y2
783.84y2 = 497.7384
y2 = 0.635kmol B/kmol
Thus, the molar composition of toluene is (1-y1) which is 0.365 kmol T/kmol. For the mass balance calculation at the reboiler stream, the same procedure is needed to be carried out. In this case, the molar composition of benzene and toluene is assumed to be at 0.7 kmol B/kmol and 0.30 kmol T/kmol respectively. For the reboiler stream, it is first important to determine the gradient of the stripping section operating line (SSOL). The gradient of the SSOL line is calculated by using the following formula:

M = y2 – y1
x2 – x1
= 0.50 - 0
0.42 - 0
= 1.19

The molar flow rate of n4 and n3 could be determined by using the calculation below:

Gradient of SSOL = L _
(L-B)
L = 1.19L – (Gradient of SSOL x B)
0.19L = 1.19 (188.28 kmol/h)
L = 1179.23 kmol/h
n4 = n3 + 188.28 kmol/h
n3 = 1179.23 kmol/h – 188.28 kmol/h
n3 = 990.95 kmol/h

Since L = n4. Hence, the molar flowrate of n4 is 1179.23 kmol/h. The determination of the composition for toluene and benzene at the reboiler stream could be determined since the molar flowrate has been identified by carried out the species balance calculation:

Benzene Balance:
Input = Output
0.65 (1179.23) = n3y2 + 0.02(188.28)
766.5= 990.95y2 + 3.7656
990.95y2 = 762.734
y2 = 0.77 kmol B/kmol
Thus, the molar composition of toluene is (1-y2) which is 0.23 kmol T/kmol.
Figure 7.2: The composition of component at each stream

Figure 7.3: The Overall Composition Of Component, Phase, Pressure And Temperature At Each Stream

The mass balance calculation for reboiler and condenser part have been determined. Then, energy balance equation must be calculated based on the following equation:
Q – W = ΔḢ + ḖK + ḖP
Q = Heat energy
W = Work done
ΔḢ = Enthalpy energy
ḖK = Kinetic Energy
ḖP = Potential Energy
Since there is no involvement from work done, potential energy and kinetic energy in the distillation column, they are neglected in the calculation. So, the assumption that can be made is:
Q = ΔḢ
The Figure 7.4 below shows the molar flowrate, composition of each stream, temperature and pressure around the condenser stream.
Figure 7.4: the composition of component, molar flowrate, temperature and pressure around the condenser

Figure 7.4 above is necessary in order to have the complete overview for the calculation of the energy balance. The next process involve for the energy balance calculation for distillation column is to identify the reference state for component of benzene and toluene. The reference state for component of benzene and toluene is set to be as follows:

Benzene (25 C, 1atm, liquid state)
Toluene (25 C, 1atm, liquid state)

After the reference state for the distillation column has been set and identified, it is important to construct the enthalpy table. Enthalpy table constructed must consist of the inlet and outlet stream component flow rates and specific enthalpies. Known values of the amount of flow rates and enthalpies and labels for the entries that are to be calculated must be inserted. The enthalpy table at the condenser stream is as follows:
Substance
m.inkmols
Ḣ.in(kJ/mol)
m.outkmols
Ḣ.out(kJ/mol)
Benzene (L)
-
-
0.0522
H3
Toluene(L)
-
-
0.00218
H4
Benzene (V)
0.19054
H1
-
-
Toluene (V)
0.08166
H2
-
-

Benzene
(m.in)(V) = 0.2722 x 0.7
= 0.19054 kmol B/kmol
(m.Out)(L) = 0.0544 x 0.96
= 0.0522 kmol B/kmol

Toluene
(m.in)(V) = 0.2722 x 0.3
= 0.08166 kmol T/mol
(m.out)(L) = 0.0544 x 0.04
= 0.00218 kmol T/mol

The next step involve in the calculation of the energy balance is by determining the hypothetical path for Ĥ1 and Ĥ2. The value of Ĥ for component of benzene and toluene at the distillate and reflux stream is equal to 0 because the pressure and temperature at that particular stream is similar to the reference state. The hypothetical path for Ĥ1 is as follows:
( Ḣ1):
Benzene Benzene Benzene Benzene
HCHCHBHBHaHa - 25 C - 80.1 C -80.1 C -120 C
HC
HC
HB
HB
Ha
Ha
-1 atm - 1 atm -1 atm -1 atm
-liquid - liquid - vapor -vapor
Ha = 2580.1Cp dT
= 2580.1 126.5x10-3+23.4x10-5T dT
= 7.648 kJ/mol
HB = Hv
= 30.7650 kJ/mol
HC= 80.1120Cp dT
= 80.1120 74.06x10-3+32.95x10-5T+-25.20x10-8T2+77.57x10-12T3dT
= 4.172 kJ/mol
Ḣ1 = Ha + Hb + Hc
= (7.648) + 30.7650 + (4.172)
= 42.585 kJ/mol

( Ḣ2):
Toluene Toluene Toluene Toluene
HaHaHBHBHCHC25 C - 110 C -110 C -120 C
Ha
Ha
HB
HB
HC
HC
1 atm - 1 atm -1 atm -1 atm
liquid - liquid - vapor -vapor
Ha = 25110Cp dT
= 25110 148.8x10-3+32.4x10-5T dT
= 14.507 kJ/mol
Hb = Hv
= 33.4700 kJ/mol
Hc = 110120Cp dT
= 110120 94.18x10-3+38.00x10-5T+-27.x10-8T2+80.33x10-12T3dT
= 1.344 kJ/mol
Ḣ2 = Ha + Hb + Hc
= 14.507 + 33.4700 + 1.344
= 49.321 kJ/mol
Substance
m.in kmols
Ḣ.in (kJ/mol)
m.outkmols
Ḣ.out (kJ/mol)
Benzene (L)
-
-
0.0522
H3
Toluene(L)
-
-
0.00218
H4
Benzene (V)
0.1904
42.585
-
-
Toluene (V)
0.8166
49.321
-
-

The Figure 7.5 below shows the molar flowrate, composition of each stream, temperature and pressure of toluene and benzene around the reboiler stream.

Figure 7.5: The composition of component, molar flowrate, temperature and pressure around the condenser
Figure 7.5 shows the composition of component, molar flowrate, temperature and pressure around the condenser. These informations are necessary in order to have the complete overview for the calculation of the energy balance. The next process involved for the energy balance calculation for distillation column is to identify the reference state for benzene and toluene. The reference state for component of benzene and toluene is set to be as follows:

Benzene (25 C, 1atm, liquid state)
Toluene (25 C, 1atm, liquid state)

The enthalpy table at the reboiler stream is as follows:
Substance
m.inkmols
Ḣ.in (kJ/mol)
m.outkmols
Ḣ.out (kJ/mol)
Benzene (L)
0.179
0
-
-
Toluene(L)
0.963
0
-
-
Benzene (V)
-
-
0.212
H3
Toluene (V)
-
-
0.0633
H4

The next step involved in the calculation of the energy balance is by determining the hypothetical path for Ĥ3 and Ĥ4. The value of Ĥ for component of benzene and toluene is equal to 0 because the pressure at temperature at that particular stream is similar to the reference state. The hypothetical path for Ĥ3 is as follows:
( Ḣ3):
Benzene Benzene Benzene Benzene
HCHCHBHBHaHa-25 C - 80.1 C -80.1 C -120 C
HC
HC
HB
HB
Ha
Ha
v-1 atm - 1 atm -1 atm -1 atm
Ha = 2580.1Cp dT
=2580.1 126.5x10-3+23.4x10-5T dT
= 7.648 kJ/mol
Hb = Hv
= 30.7650 kJ/mol
Hc = 80.1120Cp dT
= 80.1120 74.06x10-3+32.95x10-5T+-25.20x10-8T2+77.57x10-12T3dT
= 4.172 kJ/mol
Ḣ3 = Ha + Hb + Hc
= (7.648) + 30.7650 + (4.172)
= 42.585 kJ/mol
( Ḣ4):
Toluene Toluene Toluene Toluene
HBHBHCHCHaHa-25 C - 110 C -110 C -120 C
HB
HB
HC
HC
Ha
Ha
-1 atm - 1 atm -1 atm -1 atm
Ha = 25110Cp dT
= 25110 148.8x10-3+32.4x10-5T dT
= 14.507 kJ/mol
Hb = Hv
= 33.4700 kJ/mol
Hc = 110120Cp dT
= 110120 94.18x10-3+38.00x10-5T+-27.x10-8T2+80.33x10-12T3dT
= 1.344 kJ/mol
Ḣ4 = Ha + Hb + Hc
= 14.507 + 33.4700 + 1.344
= 49.321 kJ/mol

Substance
m.in
Ḣ.in (kJ/mol)
m.out
Ḣ.out (kJ/mol)
Benzene (L)
0.179
0
-
-
Toluene(L)
0.963
0
-
-
Benzene (V)
-
-
0.212
42.585
Toluene (V)
-
-
0.0633
49.321

Thus, the value of Q at the condenser and reboiler is as follows:
Q Condenser
Q = Ḣ
= Ḣoutput Ḣinput
= 0 - [ 0.1904kmols 42.585 kJkmol + (0.08166 kmols )( 49.321 kJkmol )]
= -12.1357 kJS
= -12.14 kW

Q Reboiler
Q = Ḣ
= Ḣoutput Ḣinput
= [ 0.212 kmols 42.585kJkmol +0.0633 kmols 49.321kJkmol
= 12.15 kJS
= 12.15 kW

Summary

Based on the energy balance calculation at the condenser and reboiler streams, it is identified that 12.14 kW of heat energy is required to be removed from the condenser. This is because we want the benzene and toluene at that particular stream to achieve their desired temperatures and pressures. Identically, it is required to supply 12.15 kW of energy to the reboiler to reheat back benzene and toluene so that they can be recycled back to the distillation column. Therefore, it is vital to identify the stream that releases excess heat and that excess heat energy could eventually be supplied to the reboiler stream in order to achieve such value at that respective stream. This is done to achieve the most optimum heat transfer. Subsequently, the additional supply of heat energy to the reboiler stream is unnecessary to be supplied at larger amounts in order to reheat the benzene and toluene at the bottom stream. The energy released by the condenser is also required to be optimized by channelling or connecting the condenser stream to any streams in the plant that require additional supply of heat energy. This will minimize the amount of heat supplied and will drastically help to reduce the operation cost for the production of benzene. From the summary above, we can conclude that the product where benzene came out of the distillation vapor is initially in vapor form. This benzene that came out as an overhead distillate is condensed into liquid form for several benefits. The advantages are easy leakage detection, easy to transport benzene in liquid form and more benzene can be filled into storage tank to be transported at a time (due to the volume of liquid benzene being much smaller compared to vapor benzene) .

CHE231- Production of Benzene (Full Report)

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