FINAL 03 Version

8/8/2019 FINAL 03 Version

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Table of Contents

Table of Contents ............................................................................................................................. 1

1.0 Project Brief ............................................................................................................................... 4 2.0 Project Plan and Objectives ....................................................................................................... 4

3.0 Chemical Engineering Design ................................................................................................... 7

3.1 Process Description ................................................................................................................ 8

3.2 Selected Process Stage ........................................................................................................... 9

3.2.1 Reactor ............................................................................................................................ 9

3.2.2 Reactor Selection .............................................................................................................. 10

3.3 Design Considerations of a fixed bed catalytic reactor (FBCR) ...................................... 11

3.3 Rate Kinetics ........................................................................................................................ 12

3.3.1 Langmuir and Hinshelwood (LH) Model .................................................................... 12

3.3.2 Rate of Reaction ............................................................................................................ 15

3.3 Catalyst Holdup ................................................................................................................... 16

3.4 Volume of Catalyst and Reactor .......................................................................................... 17

3.5 Reactor Sizing ...................................................................................................................... 19

3.5.1 Length and Diameter ..................................................................................................... 19

3.5.2 Bed Design .................................................................................................................... 21

3.5.3 Dimensional Design of Fixed Bed Catalytic Reactor .................................................. 23

3.5.4 Heat Exchanger ............................................................................................................. 25

3.5.5 Reactor Wall Thickness ................................................................................................ 27

3.6 Material of Construction ...................................................................................................... 28

3.7 Reactor Support ................................................................................................................... 28

3.8 Pressure Drop ....................................................................................................................... 30

3.9 Bursting Disc Design ........................................................................................................... 35

3.10 Pipe Diameter ..................................................................................................................... 39

3.11 Mean Resident Time .......................................................................................................... 41

3.12 Design Data Summary ....................................................................................................... 41

3.12 Vessel Data Sheet .............................................................................................................. 43

4.0 Control and Instrumentation .................................................................................................... 44

4.1 Introduction and Control Objectives .................................................................................... 44

8.2 Control and Instrumentation ................................................................................................ 86

4.2 Control Strategy ................................................................................................................... 45
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4.2.1 Reactor Instruments ...................................................................................................... 45

4.2.2Valves ............................................................................................................................ 45

4.2.3 Reactor Control Loops .................................................................................................. 46

4.2.4 Reactor Control Systems ............................................................................................... 48

4.2.5 Alarms and Safety Trips ............................................................................................... 51 5.0 Piping and Instrumentation Diagram (P&ID) .......................................................................... 53

5.1 Introduction ......................................................................................................................... 53

5.2 Piping and Instrumentation Diagram ................................................................................... 55

5.3 Piping and Instrumentation Diagram (P&ID) Item List .................................................. 56

6.0 Hazard and Operability Study (HAZOP) ................................................................................. 57

6.1 Introduction .......................................................................................................................... 57

6.2 Basics of a HAZOP .............................................................................................................. 58

6.3 Process Line for HAZOP ..................................................................................................... 60 6.4 Piping and Instrumentation Diagram of Chosen Process Line ........................................... 62

6.5 HAZOP Ethanol Production Process ................................................................................ 63

7.0 Economic Appraisal ................................................................................................................. 71

7.1 Estimating Capitol Cost ....................................................................................................... 71

7.1.1 Wilson Method .............................................................................................................. 72

7.1.2 Zevnik and Buchanan Method ...................................................................................... 74

7.2 Estimating Raw Material Cost ............................................................................................. 75

8.3 Economic Appraisal ............................................................................................................. 86 7.4 Estimating Operating Cost ................................................................................................... 77

7.5 Assumptions ......................................................................................................................... 79

7.6 Annual Income ..................................................................................................................... 80

7.7 Pay Back Time ..................................................................................................................... 80

7.8 Net Present Value and Rate of Return ............................................................................. 81

7.9 Raw Material and Product Cost Variation ....................................................................... 83

7.10 Conclusion ......................................................................................................................... 84

8.0 Reference ................................................................................................................................. 85 8.1 Chemical Engineering Design ............................................................................................. 85

9.0 Appendix
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1.0 Project Brief

Project Title Ethanol ProductionChosen Process Method Ethylene HydrationLocation USAPlant Capacity 100,000 tonnes per annumPlant Operating Time 8000 hours / 333 daysProduct EthanolFeedstock Ethylene

Steam

Brine (coolant)Catalyst Phosphoric AcidMass Flow of Feedstock Ethylene 7763.975 kg/hr

Steam 4991.127 kg/hr Mass Flow of Product Ethanol 12500 kg/hr

2.0 Project Plan and Objectives


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The Advanced Process Design part 2 is a more individual project where the investigated route in

part 1 continues to be investigated from a design perspective. There still remain parts which

remain group work but mostly the individuals ability to demonstrate their knowledge is tested.

The objective of this task is to:

Apply knowledge of reactor design

Gain a comprehensive understanding of the requirements of design requirements

To develop key skills - Initiative

Time management

Research skills

IT Skills

Report writing skills

The project objective is the design of an important piece of equipment from the process

researched and chosen. As can be seen on the project brief, an industrial ethanol plant was the

process. This report will contain the design of the reactor from that process.

Working from these objectives, a written report is to be compiled with the findings with the

following critical tasks:

Chemical Engineering Design

The design of the reactor: using relevant kinetic data, calculate the vessel dimensions,

volume and catalyst weight to ensure the desired reaction.

Control and Instrumentation

Devise and explain control strategies to ensure the safe operation of the reactor process.

Piping and Instrumentation Diagram (P&ID)

Based on the control and instrumentation strategy, devise a detailed diagram for use in

construction of the reactor process.

Hazard and Operability Study (HAZOP)


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A HAZOP is done on one line of the process to identify all possible hazards and come up

with the solutions.

Economic Appraisal

This is done evaluate the economic feasibility of the process.

Cash flow tables and the rate of investment return are calculated to justify it.

A Gantt chart (Appendix A) has been done to track the progress of the project.

To ensure the project is on track, meetings are to be held with the assigned tutor (Dr JT) weekly.

However, it is important to note that the group tasks are flexible dependant upon the availability

of other group members.

Review of the Project against the Plan

An overall review of the project is that though it was challenging, it was an enjoyable task that

was completed successfully. The initial schedule (Appendix A) was devised to help give

structure to time-management, however, this was a flexible and the resulting change can be seen

in Appendix B.

The obstacles encountered were mainly on two counts. The main section (Chemical Engineering

Design) over ran due to a number of numerical errors and extra design aspects initially missed.

This meant that these were corrected a number of times.

The HAZOP was planned in later than the schedule which led to the economic appraisal

investigated earlier than planned. This proved to be good initiative as this task proved to be a

tricky and time-consuming task.

The objectives I set out to achieve have been met to an acceptable degree, however, the mostimportant skill I used and enhanced are the following two:

Research this is the area I found the most difficult and frustrating, with no

immediate success in my search for relevant kinetic data I initially


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struggled. However, with some help and determination, my kinetic data

was found and calculated successful values.

Initiative I immediately realised that there was only so much assistance I could gain

from others and I was keen on testing myself to complete as much as I

could with as little help as possible. For this reason I spent a lot of time

trying out various methods to ensure I understood what I was doing.

I found that position of group leader was harder than I imagined, however,

I also feel that the role helped me to experience and utilise various skills.

This is one of the skills I feel will assist me considerably in my career.

3.0 Chemical Engineering Design


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3.1 Process Description

In Advanced Process Design Part 1, the objectives were to research and evaluate possible routes

for the manufacture of ethanol. A total of 5 routes were investigated of which two were synthetic

routes (direct and indirect hydration of ethylene) and three biological routes (fermentation of

corn and wood chips). The evaluation process looked at the raw material availability, process

complexity, conditions for example. The diagram below is a summarised form of the various

routes:

Figure 1: Process Routes Considered

The evaluation process decision was in support of the synthetic direct hydration of ethylene.

The process is a vapour phase hydration of ethylene at 6MPa and 573K (300o

C) with a solidPhosphoric Acid catalyst. There are two reactants (ethylene and steam) that are fed into a reactor

in vapour form which reacts on the catalyst bed to form ethanol, the by products formed are

diethyl ether and acetylene ( acetaldehyde impurity hydration reaction).

The products leave at a higher temperature because the reaction is exothermic (data suggests 20K

hotter). These are cooled in a heat exchanger into liquid and vapour streams, the purpose of this


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step is to reduce material costs of the pipes and equipment that follow. This also aids in the

separation of the products and reactant. This reaction has a low yield per pass of approximately

5% which results in high amounts of unreacted ethylene which is recycled, this remains in

vapour phase after cooling and is separated in an adsorber, compressed and fed back into the

reactor. The other products condense into liquid form.

The liquid products are then distilled to separate the desired product ethanol from the by

products.

Though the conversion per pass is low, the selectivity of ethylene is at acceptable 98%

conversion.

3.2 Selected Process Stage

3.2.1 Reactor

This report will focus on the reactor and its design. The reactions at this stage are:

C2H4 + H 2O CH3CH 2OH

Ethylene + Water Ethanol

C2H2 + H 2O CH 3CHO

Acetylene + Water Acetaldehyde

2CH 3CH2OH (CH 3CH 2)2O + H 2O

Ethanol Diethyl Ether + Water

Assumption : For this design the acetylene and diethyl ether reactions will be ignored

because the amounts produced are relatively small.

The design will also concentrate on the conversion as a whole due to the small

amounts produced.


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The ethanol production was specified at 100,000 tonnes of ethanol per annum. From this data,

taking into account of maintenance and shut down time (8000 hours operating time) the mass

flow rate of ethanol required is calculated to be 12,500 kg/hr.

3.2.2 Reactor Selection

There are various types of reactors available and the evaluation of them is described below:

Batch Reactor

This type of reactor has no flow in or out of reactants of products whilst the reaction occurs. It is

generally used for small scale production and can achieve high conversion if left in the tank for

long periods of time. However, this is unsuitable for this process because the product stream

needs to removed continuously. The other problem is the ability to apply this to a large scale

production successfully.

CSTR

This is a very common industrial processing unit where the reactants are fed in and products

removed continuously. It is used when intense agitation is required for the reaction to occur.

However, the reaction for this process occurs on the surface of the catalyst and hence, there is no

requirement for agitation.

Tubular

This type of reactor is where the reactor is length is large with respect to the diameter. With no

moving parts they are easier to maintain and usually have the highest conversion per reactor

volume. The disadvantage is the possible occurrence of hot spots and temperature control is

more difficult.

A packed bed reactor of this type is best suited to the reaction designed in this report because of

the need for a solid catalyst with no agitation required and for it to be a continuous process.


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The reactor type selected is known as a Fixed Bed Catalytic Reactor.

3.3 Design Considerations of a fixed bed catalytic reactor(FBCR) 1

In a FBCR for a fluid-solid reaction, the solid catalyst is present as a bed of relatively small

individual particles, randomly oriented and fixed in position. The gas moves by convective flow

through the spaces between the particles. Diffusive flow through the particle is also possible.

Adiabatic Operation

No attempt is made to adjust the temperature within the bed by means of heat transfer. The

exothermic or endothermic nature of the reaction determines the solution of a single stage reactor

whereas a multi stage reactor where the catalyst is divided into at least 2 beds can use heat

exchangers or a cold-shot cooling method (exothermic). The purpose of temperature adjustment

is to shift the equilibrium position to increase yield and maintain a high rate of reaction.

Non Adiabatic Operation

This is where temperature is adjusted using heat transfer, this means the reactor is essentially a

shell and tube heat exchanger.

Amount of Catalyst

This is important in the sizing of the reactor, the volume, bed depth and diameter can be

calculated from this.

Particle and Bed CharacteristicsParticles These include chemical composition which determines catalytic activity; Physical

properties (e.g. size, shape, density and porosity or voidage) determines its diffusion

characteristics.
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Bed These include density and voidage characteristic; volume of bed.

Pressure Drop

When a gas flows through a bed of particles, interactions between the gas and particles lead to africtional pressure drop (-P). Calculation of (-P) enables determination of both L and D for a

given W. Also Specifying D or L for a given W will allow one to calculate (-P) using the Ergun

equation.

Choice between allowable (-P) or D (or L) from a given W requires a trade-off between the

cost of the vessel and the cost of compressing the gas. The smaller D, the greater the L/D ratio

and the greater (-P); implying the cost of the vessel is less.

3.3 Rate Kinetics

These are generally calculated from experimental data where the rate has been calculated under

known conditions to which they can be directly related in the design. However, I was unable to

find such data where the exact conditions and process inputs matched. The closest data I found

was from a journal which had differences in that the catalyst was different and the temperaturewas lower. Nevertheless, its data is valuable in assisting in the design as the reaction is the same.

3.3.1 Langmuir and Hinshelwood (LH) Model

The LH model is based on a fluid solid catalyst reaction where adsorption and desorption

occurs on the catalyst surface. I am evaluating my rate equation based on this model.

The reaction that occurs here is a bi-molecular (i.e. two reactants)


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The reactants will have the following subscripts based on the above.

Ethylene = A

Steam = B

Ethanol = P

The general bi-molecular rate equation based on the LH model is:

mol/m 3 hr or

mol/kg hr

mol/ m 3 hr atm -1

atm -1

mol/m 3

mol/m 3

Based on assumptions on the adsorption or lack of adsorption this equation can be evaluated to

another form.

Assumption: Ethylene is strongly adsorbed on the catalyst surface

Water and Ethanol are weakly or not adsorbed on catalyst surface

The equation will now have this form:


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The adsorption constants and rate constant is based on experiments conducted on similar

conditions.

The two unknowns are the two concentration values

Calculating Concentration

By substituting equation (3) into the ideal gas equation (2), the resulting equation (4) can be used

to calculate the concentration.

The pressures required for this equation are the partial pressure as we are trying to find the

concentration of each reactant. However, as the mole fraction are the same the partial pressure

will also be the same, and hence half of the total pressure of 6 MPa.


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Substitute values into equation (4)

3.3.2 Rate of Reaction

The values calculated and those from the journal can be substituted into equation (1) to work out

the rate of reaction.

This rate value is expressed in terms of the volume, when designing a fixed bed catalytic reactor

it can also be expressed in terms of the mass. Therefore, by using the following equation this is

achieved with respect to the catalyst bed.


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3.3 Catalyst Holdup

When designing any reactor, there is a design equation for each type of reactor which

corresponds to calculating an important design aspect. For a fixed bed reactor, it is expressed in

terms of W (catalyst holdup) as the equation below shows.

W amount of catalyst kgFA molar flowrate kmol/hr XA conversion 0.98 (specified from TP1)(-rA) rate of reaction kmol/kg hr 0.07125 mol/kg hr

Calculating molar flowrate

The unknown from the above equation is the molar flowrate. The following expression to

calculate the molar flowrate is used because both the mass flowrate and relative molecular mass

are already known.


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Substitute values into the above equation.

Now substituting this into equation (7) gives the catalyst weight.

3.4 Volume of Catalyst and Reactor


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The reactor consists of a cylinder and a hemisphere which make the total volume.

When operating at high pressures, the pressure force acting against the roof is transmitted upto

the shell. For this reason curvature becomes necessary at the two ends of the reactor.

There are various types of circular ends that can be used, however, the hemispherical ends

provides the most strength. It is capable of resisting about twice the pressure of a torispherical

head of the same thickness. 2

Based on general assumptions commonly used, the total volume can can calculated.

Since the mass has been calculated, the volume it occupies can be calculated using the following

equation.

In this type of reactor, the catalyst volume takes up anywhere between 50 75% of the reactor

volume.

Assumption : the catalyst volume is two thirds of the cylindrical volume of the reactor

Therefore, the cylindrical volume is:


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The volume of the hemisphere can only be calculated once the diameter or radius is known.

Diameter of the reactor = 2.1 m (see section Reactor Sizing below for calculation)

Assumption: the height of will be set at half the diameter of the reactor

If D r = 2.1 m

3.5 Reactor Sizing

3.5.1 Length and Diameter

From the volume calculated above, the sizing of the reactor is calculated based on the following

volume equation.

It is calculated based on the volume of the cylinder.


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However, both the diameter and length are unknown. Hence, general correlations are to be tried

and compared.

Assumption: (1) Length is twice the size of the diameter

(2) Length is thrice the size of the diameter

(1) If L = 2D

Therefore

(2) If L = 3D

Therefore

After comparing the values of the length with the flowrate (554.262 kmol/hr), the first

assumption gives more realistic values of the reactor sizing.

Cylinder Diameter = 2.1 mCylinder Length = 4.2 m

The dimensions of the hemisphere need to be included.

The length has been specified in the section above.


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Hemisphere Diameter = 2.1 m

Hemisphere Length = 1.05 m

The dimensions of the total reactor are:

3.5.2 Bed Design

Assuming that the diameter of the catalyst is the same as the reactor diameter, the depth of the

catalyst can be calculated by using the volume it occupies in the reactor.

VC = 9.534 m 3

DR = 2.1 m

Catalyst Bed

The catalyst bed will be separated into two compartments of 1.375 m length and 2.1 m diameter

each. However, the reaction is exothermic (temperature increased by 20K) and this extra heat

needs to be removed. For this safety a heat exchanger will be added.


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Heat exchanger

The heat exchanger is required to remove heat generated by the process. As the temperature is

relatively small, the heat transfer area will also relatively small and is to be placed inside the

reactor between the catalyst beds as can be seen in the dimensional diagram below.


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3.5.3 Dimensional Design of Fixed Bed Catalytic Reactor

Raw Material Input(Ethylene and Steam)


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Bed 1

Bed 2

HE01 0.45m

4.2m

6.3m

2.1m

0.177m

0.5m

0.5m

1.375m

2.1m

1.05m

1.05m

Burst Disc

Product Output(Ethanol)


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Bed 1

Bed 2

HE01 0.45m

4.2m

6.3m

2.1m

0.177m

0.5m

0.5m

1.375m

2.1m

1.05m

1.05m

3.5.4 Heat Exchanger

The purpose of the design here is to calculate the area required for the specified duty (rate of heattransfer). The actual design here will be a simple calculation and will not explore all the specifics

regarding the number of tubes for example. It is also important to note that alternate options are

available.


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Q Heat Transferred W or J/sU Overall Heat Transfer Co-efficient W/m 2K A Heat Transfer Area m 2

Mean Temperature Difference K

The form of the equation required becomes:

Calculate the Heat Transferred (Duty)

This is calculated using the flowrate and heat capacity of ethanol using the following equation:

Mass flowrate of ethanol kg/sHeat capacity of ethanol at 593K 105.56 J/mol K Temperature into heat exchanger 593 oC

Temperature exiting heat exchanger 573 oC

The mass flowrate of ethanol into the heat exchanger will be half the total exiting the reactor, this

is because of the design to separate the catalyst bed into two compartments. The amount of

ethylene converted will be half the amount so its flowrate will also be half. This will be reflected

in the calculation below.

The heat capacity calculated in the energy balance is expressed in the form of J/mol K, however,

it is needed in the form of its mass as opposed to the number of moles.

By dividing by its relative molecular mass the form changes into what is required.


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Convert to J/s which is also 1 Watt.

Assumption : The Heat Transfer Coefficient (U) is 350 W/m 2 oC 3

Substitute values into equation (13) to calculate the area of heat transfer

3.5.5 Reactor Wall Thickness

Every reactor needs to have a designed thickness to ensure it has sufficient strength to cope with

the design conditions, it is ever more important in a design where the reaction requires high

pressure within the reactor.

The pressure value used will be the burst disc pressure (see bursting disc calculations below).

This is calculated using the following equation.

PBD Burst Disc Pressure 66 MPa or 66 bar DR Diameter 2.1 mf i design stress of material 101 N/mm 2 or 1010 bar (stainless steel (320)) 4

Corrosion Allowance
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Corrosion to the reactor is a consideration that needs to be taken into account when sizing thethickness of the reactor, though stainless steel is resistant to many forms of corrosion, anallowance still needs to be calculated into the thickness. Most design codes and standards specifya minimum of 1mm 5. For this reason, this value of 1mm is to be used.

3.6 Material of Construction

The pipes and vessel will all be made out of stainless steel as it is high strength and resistant tocorrosion except the steam pipe. The reasons are:

The high pressure and temperature required in the reactor means that stainless

steel is the safest option.The steam pipe does not require such high strength material and the cost savingcompared stainless steel is significant. However, due to the high temperature, anadditional alloy (chromium) will need to be added to improve the tensile strength.

3.7 Reactor Support

The reactor has a weight and will need to be supported because of the earths gravitational pull

on it. This force is calculated by working out the load of the contents (equation 16) and the load

of the vessel (equation 17).

Weight of Contents


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WM Weight of Contents Newton or kg m /s 2

VR Reactor Volume 19.149 m 3

Density of densest material

(ethanol)

789 kg/m 3

Gravitational force 9.81 m/s 2

Weight of Vessel

This equation is for a steel vessel.

WV Weight of the shell (excludes internal fittings) Newton or kg m/ s 2

CV factor that takes into account nozzles,

manholes, internal supports etcDM Mean diameter of vessel mHV Height between tangent lines mt wall thickness mm

CV For vessels with few fittings = 1.08

DM

HV Length of the cylindrical section = 4.2 m


This weight exerted by the vessel is to be supported by saddles, they are usually constructed of

bricks or concrete and have a contact angle between 120 o and 150 o.


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Wear plates are also to be welded to the shell wall to reinforce the wall over the area of contact

with the saddle.

A typical saddle design and dimensions is as follows for a vessel with a diameter of 1.2 m.

Figure 2: Saddle Design Dimensions 6

Vessel

Diameter

(m)

Maximum

Weight

(kN)

Dimensions (m) Dimensions (mm)V Y C E J G t 2 t1 Bolt

Diameter

Bolt

Holes

1.2 180 0.7

8

0.2 1.09 0.45 0.3

6

0.14 12 10 24 30

3.8 Pressure Drop

Pressure drop is a term used to describe the decrease in pressure from one point in a pipe or tube

to another downstream. This is usually the result of friction of the fluid against the tube. High

flow rates in small tubes give larger pressure drop. Low flow rates in large tubes give lower

pressure drop.

This can be ignored liquid phase reactions. However, not in gas phase reactions where the

concentration of the reacting species is proportional to the total pressure. Proper accounting of

the pressure drop on the reaction system in many instances can be a key factor in the success or

failure of the operation.


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To calculate the pressure drop the Ergun equation is used. The Ergun equation is a correlation for

the friction factor in a column as a function of the Reynolds number.

f friction factor ug superficial linear velocity m/s

gas density kg/m 3

LD bed depth mdp effective particle diameter m

Effective Particle Diameter

Assumption: Particle diameter = 2 x 10 -3 m

The effective particle diameter is twice the particle diameter

Superficial Linear Velocity

This is calculated using the following equation

u superficial linear velocity m/svo volumetric flowrate m 3/hr

SA surface area m 2

Calculate volumetric flowrate


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Mass flowrate 7763.975 kg/hr Ethylene Gas density 2.085 kg/m 3 7

Mass flowrate 4991.127 kg/hr Steam Gas density 15.009 kg/m 3 8

Calculate surface area

DR Diameter of the reactor 2.1 mr R Radius of the reactor 1.05 mLR length of the reactor 6.3 m



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Friction Factor

To calculate this value we use the following Ergun Correlation.

Bed voidageR E Reynolds Number

Bed Voidage

Bed density 800 kg/m 3

Particle density 1685 kg/m 3

Reynolds Number

There are many variations of the Reynolds number, for a packed bed reactor the following is

most suitable.

Average density 8.547 kg/m 3

ug superficial velocity 0.0232 m/sdp particle diameter 2 x 10 -3 m

viscosity 0.000017 kg/ms 12

bed voidage 0.474

The average density is calculated as follows.


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The viscosity is required in the form of Poiseuille (Pa s)


All the unknown values have been calculated and the pressure drop can be worked out by

substituting them into equation (18).

Pressure Drop

To convert this number into bar, the following conversions can be applied.


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The pressure drop and vessel dimensions from a given catalyst weight will have an impact on the

vessel cost. However, in a reactor where the pressure plays an important role in the conversion

process, it is important that the pressure drop is controlled.

The value calculated here is relatively small and acceptable for this reactor, when compared to

the pressure required (~60bar).

3.9 Bursting Disc Design

Over-pressure / pressure exceeding the system design pressure, is one of the most serious

hazards in chemical plant operation. Failure of a vessel or the associated piping can trigger a

sequence of events that culminate in a disaster. Pressure vessels are invariably fitted with some

form of pressure relief device, set at the design pressure so that potential over-pressure is

relieved in a controlled way. A bursting disc a thin disc of material that is manufactured and

designed to fail at a predetermined value, giving a full bore opening for flow. 9

The information required by a manufacture is the area and diameter of the disc. The area is

calculated through the following equation.


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QM mass flowrate discharged kg/hr Ao cross-sectional area of bursting disc m 2

PV relieving pressure bar F function of isentropic component k and ratio of back

absolute pressure to inlet absolute pressure Discharge coefficient of nozzle type

RMM molar mass kg/molZ compressibility factor

Mass flowrate

Q = mass flowrate out of reactor = 12755 kg/hr

Relieved Pressure

Assumption: Relieving pressure is 10% above the operating /design pressure.

If pressure = 6 MPa

F value

The isentropic exponent k for air we assume is 1.4(table 2 BS 2915 value for Air).

The back pressure pb we are assuming as 1 bar absolute, this is because when the burst disc

ruptures it will open to atmosphere.


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Coefficient of Discharge

This value depends on factors such as the manufacturer model and type, disk burstingcharacteristics and flow restrictions. However, for the purpose of this design a general value of

0.62 can be used if the following four conditions are met:

The disk must be installed within 8 pipe diameters of the vessel

The disk discharge pipe must not exceed 5 pipe diameters

The disk must discharge directly to the atmosphere

The inlet and outlet piping is at least the same nominal pipe size as the bursting disk. 10

Vent Temperature

This is calculated using the following equation.

PN Operating Pressure 6 MPaPV Relieving Pressure 6.6 MPaTN Operating Temperature 573 K TV Vent Pressure K

Compressibility Factor


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This value is a function of the reduced pressure and temperatures which are then used to

estimate the compressibility factor from the correlation in Appendix C.

Reduced Pressure (P r)

P Design Pressure 60 bar Pc Critical Pressure of Air. 37.69 bar

Reduced Temperature (T r)

T design temperature 573 K Tc critical temperature of air 132.45 K

Compressibility Factor = 1

Substitute values into equation (24) to first calculate the area of the bursting disc and from there

the diameter can be calculated.


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3.10 Pipe Diameter

The inlet and outlet pipe diameters are calculated using the following equation

d pipe diameter mvo volumetric flowrate m 3/su gas velocity m/s

Inlet Reactant PipesThe typical velocity of a vapour stream is 15-30 m/s. 11

Ethylene Pipe

The volumetric flowrate has been calculated in the pressure drop calculation but here is needed

in the form m 3/s


If u = 15 m/s

If u =30m/s


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Steam Pipe


If u = 15 m/s

If u =30m/s

Outlet Product (Ethanol) Pipe

The following values are known from previous calculations

mass flowrate 12755 kg/hr or 3.543 kg/sethanol density 789 kg/m 3 12


If u = 15 m/s

Assumption: The calculations give unrealistic values, hence the pipe diameter will be assumed

to be 1 inch.


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3.11 Mean Resident Time

Mean resident time (Resident Time Distribution) is the time spent by the reactant in the reactor.It is calculated using the following equation:

Mean resident time sReactor Volume 14.149 m 3

Volumetric Flowrate 1.1354 m3

/s

The volumetric flowrate is the rate of both reactant streams.

vA Ethylene = 1.043 m 3/s

vB Steam = 0.0924 m 3/s

3.12 Design Data Summary

Reactor Data

Volume 14.159 m 3

Diameter 2.1 mLength 6.3 mWall Thickness (inc. corrosion) 71.9 mmReactor Weight (support) 368 kNMaterial Stainless SteelPressure Drop 0.0928 bar Ethylene Pipe Inlet 1 inchSteam Pipe Inlet 1 inchEthanol Pipe Inlet 1 inch


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Catalyst Data

Catalyst Phosphoric AcidWeight 7623 kgVolume 9.534 m 3

Diameter 2.1 mBed Depth 2.75 mBed Density 800 kg/m 3

Particle Density 2 x 10 -3Effective Particle Diameter 0.004 mBed Voidage 0.474

Kinetic Data

Rate Constant 18 mol/m 3

Rate of Reaction 57 mol/m 3 hr Rate of Reaction 0.007125 mol/kg hr

Mean Resident Time 10.449 sSpace Velocity 2300 -hr

Activation Energy 80 cal/mol K


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3.12 Vessel Data Sheet

Vessel Data Sheet

EQUIPMENT NO. (TAG)

DESCRIPTION (FUNCTION)

SHEET NO.Operating Data

NO. REQUIRED 1 CAPACITY 19.159 m 3

SPECIFIC GRAVITY OF CONTENTS COMPUTATED (YES OR NO) X

SHELL JACKET FULL/HALF COIL INTERNAL COIL

CONTENTS X X X

DIAMETER 2.1 m X X

LENGTH 6.3 m X X

DESIGN CODE X X

MAX. WORKING PRESSURE 66 bar X X

DESIGN PRESSURE 60 bar X XDESIGN TEMPERATURE 573 K X X

TEST PRESSURE (HYDROSTATIC) X X X

TEST PRESSURE (AIR) X X X

MATERIALSSTAINLESSSTEEL X X

JOINT FACTOR X X

CORROSION ALLOWNACE 1 mm X X

THICKNESS 71.9 mm X X

END TYPE HEMISPHERE THICKNESS JOINT FACTOR

END TYPE HEMISPHERE THICKNESS JOINT FACTOR

TYPE OF SUPPORT SADDLES THICKNESS X MATERIAL X

WIND LOAD DESIGN XINTERNAL BOLTSMATERIAL X TYPE X NUTS XEXTERNAL BOLTSMATERIAL X TYPE X NUTS XINSULATION (SEP.ORDER) 2 INSULATION FITTING ATTACHMENT BY

GASKET MATERIAL X INSPECTION BY

PAINTING X

WEIGHT 388 kN EMPTY 240 kNINTERNALS AND

EXTERNALS XORDER NO. DRG NO.

ORDER NO. DATE OF ENQUIRY

MANUFACTURER

REMARKS AND NOTES - UNLESS OTHERWISE STATED ALL FLANGE BOLT HOLES TO BE

OFF CENTRE OF VESSEL CENTRE LINES N/S AND E/W (NOT RADIALLY)PREPARED 3 6


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CHECKED 2 5APPROVED 1 4

DATE

ENGINEERING

PROCESS REV BY

APPROVED

DATE

REV BY

APPROVED

DATE

4.0 Control and Instrumentation

4.1 Introduction and Control Objectives

Instruments are used in industrial plants to indicate, monitor, control and/or record the key

process variables during operation.

Process variables such as pressure can be incorporated into automatic control loops, manual

monitoring and/or automatic computer logging systems in order to ensure its safe level. 13

Critical processes are fitted with automatic alarms to ensure the safety of all.

The main objectives of using instrumentation and control when designing the packed bed reactor

(PBR) are the following:

Safe reactor operation:

This is achieved by ensuring the critical situations are prevented. For example, the pressurein the reactor is monitored and controlled using a pressure control instrument.

Production rate:

Instruments such as flow controller will ensure that the desired amount of flow enters the

reactor to produce the required amount of ethanol.

Product quality:

The product is achieved using specific temperature, pressure and ratio of reactants which are

maintained using the different instruments and control.

Cost:

It is always desirable to operate at the lowest production cost, however cost commensurate

with other objectives. For example: production rate will depend on cost.


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4.2 Control Strategy

4.2.1 Reactor Instruments

The following instruments are commonly used when designing a reactor and are all used in the

P&ID of this reactor.

Flow Indicator Control (FIC)

Pressure Indicator Control (PIC)

Pressure Indicator (PI)

Temperature Indicator Control (TIC)

Temperature Indicator (TI)

Burst Disc Indicator (BDI)

4.2.2Valves

There are 3 types of valves used in the reactor section.

Gate Valves

They are controlled manually and are designed to be either open or closed fully. They have no

role in flow regulation.

Control Valves

Control valves are vital components of modern manufacturing. Chemical plants consists of

hundreds or more control loops all networked together to ensure the process is run correctly and

producing the desired product.

In the reactor section, control valves (V02) and (V05) are used to control the amount of flow of

both ethylene and steam entering the reactor.


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Figure 3: Control Valve

V02

Closing/Stop Valves

Closing valves are used for safety trips. They prevent developments of hazardous situation.

Closing valves are used in the reactor section (figure 8) to stop the inlet feed (ethylene and

steam) when temperature and/or pressure go above or below the operating conditions.

Figure 4: Closing Valve

V03

4.2.3 Reactor Control Loops

The purpose of a control loop is to automate operators actions, by specifying a set point the

instrument will continuously monitor and adjust the process as required to achieve the desired

value.

The reactor operates at high pressure and temperature to achieve the required conversion, for this

reason control loops are required to ensure that both of these are at the set point or within an

acceptable tolerance.

Design Pressure = 60 bar

Due to the equilibrium nature of this reaction, lower pressure values will result in low conversion

whereas high pressure results in higher operating costs and poses a danger with respect to

material constraints.

The pressure indicator (PI-03) at the reactor is linked to the flow controller (FIC-101 and FIC102) which is able to control the flow of the reactant streams.

Design Temperature = 300 oC / 573 K

The temperature also has an effect on the reaction again due to the equilibrium nature of the

reaction similar to the pressure. Therefore, a series of temperature indicators (TI-01, TI-02, TI-


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03) are placed at various positions that provide feedback to the flow controllers (FIC-101 and

FIC 102) which increase or decrease the flowrate of the reactant streams.

Figure 5 below shows the piping and instrumentation diagram of the pressure control loop.

Figure 5: Pressure control loop at the reactor

V01

FIC

101

PI01

A-H

A-L

C01

S01

V04

FIC102

S02

The flowrate of the ethylene and water streams are controlled by the data interpretation from the

above pressure indicators by using a control valve (V02 and V05) as can be seen in Figure 6.

These valves will vary its position depending on the flow controllers.

Figure 6: Flowrate controlled Control Valves

V01

FIC

101

PI01

A-H

A-L

C01

S01

V04

FIC102

S02

V02

S03

V05

S04


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4.2.4 Reactor Control Systems

Ratio of Flow Control

Ratio control is necessary when a certain ratio is required to maintain a safe operation and/or

product quality, for this reaction it is important that the ethylene is kept ever so slightly in excess

(studies have found that when steam is in excess, the catalyst life decreases significantly). For

this reason a ration control system has been implemented with a set point.

If the ethylene stream flow was changed due to pressure indicators from the reactor, the ratio

controller will also implement the required change in the steam to ensure the ratio remains

constant. This is achieved by linking the ratio control to the flow controllers (FIC-01 and FIC102) which vary the flow through the control valves (V02 and V05) as can be seen in Figure 3.

Figure 7: Ratio Controller

V01

FIC

101

PI01

A-H

A-L

C01

S01

V04

FIC102

S02

V02

S03

V05

S04

RatioController

S06

S05

Pressure and Temperature Control

As stated before, this reaction occurs at high pressure and temperatures, for this reason an

emergency shutdown system also needs to be implemented. This needs to be an automatic and

immediate action so it is vital that this system is automated through computers.


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This can be accomplished by the stop valves (V03 and V06) as shown in Figure 8 that receive

input from the alarms from PI-01 and PI-02

Figure 8: Location of the Stop Valves

V01

FIC101

PI01

A-H

A-L

C01

S01

V04

FIC102

S02

V02

S03

V05

S04

RatioController

S06

S05

V03

V06

E02

E01

To ensure these are activated when needed it is linked to the high alarm highlighted in Figure 8,

this safety trip will notify the operators but will not wait for their response and will use the stop

valve immediately. This is shown by the label C01/C02 (cause) and E01/E02 (effect), the cause

is high pressure/temperature and the effect is the shutting of the valve.For the situation where the reduction of flowrate can resolve the problem, the control is achieved

through the usage of the control valves (V02 and/or V05).

This system is designed to be implemented into both the pressure and temperature control

systems.

As this reaction is exothermic the need to control the temperature is because of the effect it can

have on the efficiency and high temperatures that leads to sintering of the catalyst. 14

The importance of removing the heat generated is heightened because the catalyst bed has been

separated into two compartments, this increases the risk of catalyst damage in the second bed.

The solution for this is to include a heat exchanger in the reactor between the catalyst

compartments. To ensure that this process operates to spec, a temperature indicator controller


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(TIC-103) measures the temperature of the outlet coolant. This is linked to a control valve (V08)

placed at the inlet of the coolant to increase/decrease the flow as Figure 9 shows.

Figure 9: Control of Heat Exchanger Flowrate

R01

TIC01

V07

S07

V08

TI03

HE01

Cascade Control

Cascade control is where the output of the primary controller is used to manipulate the set point

of the secondary controller to ensure safe operation.

The example of its use here is the pressure control, this is partly maintained by varying the flow

of the reactants by use of the pressure indicator at the reactor.


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4.2.5 Alarms and Safety Trips

Alarm and safety trips in the reactor section are required not just to ensure a continued and

smooth run of operation but to comply with health and safety requirements.

Key instruments such as temperature and pressure are fitted with audible and visual alarms

where lack of response by operator or delays could cause a hazardous situation.

Alarms

Audio and visible alarms are fitted to the pressure and temperature indicators to notify operators

of any hazardous situation to ensure their safety, as can be seen on Figure 8 there are two types

of alarms used, a high alarm and a low alarm. They are used depending of the seriousness of the

situation.

Alarm high (A-H) if the pressure exceeded the design pressure for example

Alarm low (A-L) if the pressure dropped outside the tolerance set by design

Safety Trips

Safety trips are used in conjunction with the high alarms to automatically prevent any hazardous

situation. The main trips are those that are connected to the stop valve which would immediately

stop the flow of the reactants as can be seen in Figure 8.

Bursting Disc

This is used as a safety relief system in order to protect vessels, piping and equipment due tohigh pressure. It is designed to burst at a specific pressure releasing the contents of the vessel.


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Lagging

Lagging is to be used on the pipes and reactor. The reasons for this are:

Safety of the operators. The temperature in the pipes and reactor are significantly

high and the lagging will ensure that the risk will be as minimal as possible. In

addition to this, hazard signs will also be used to further minimize the risk.

To prevent heat loss from the pipes and the reactor. The reactants need to react at

high temperature so any heat loss will reduce the conversion.

Noise levels above 60 decibels are rendered areas where it is obligatory to provide

noise protection. By ensuring this limit is not exceeded, the operator risk is

reduced and cost is removed.


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5.0 Piping and Instrumentation Diagram (P&ID)

5.1 Introduction

A Piping and Instrumentation Diagram - P&ID, is a schematic illustration of functional

relationship of piping, instrumentation and system equipment components.

P&ID shows all of piping including the physical sequence of branches, reducers, valves,

equipment, instrumentation and control interlocks .

The P&ID are used to operate the process system.

A P&ID should include:

Instrumentation and designations Mechanical equipment with names and numbers All valves and their identifications Process piping, sizes and identification Miscellaneous - vents, drains, special fittings, sampling lines and reducers/increasers. Permanent start-up and flush lines Flow directions Interconnections references Control inputs and outputs, interlocks Interfaces for class changes Seismic category Quality level

Annunciation inputs Computer control system input Vendor and contractor interfaces Identification of components and subsystems delivered by others Intended physical sequence of the equipment


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A P&ID should not include:

Instrument root valves control relays manual switches equipment rating or capacity primary instrument tubing and valves pressure temperature and flow data elbow, tees and similar standard fittings extensive explanatory notes


55/87

5.2 Piping and Instrumentation Diagram

1-SS-001-2L

V01 V02

1-CS-002-2L

V04 V05

R01

101

102

01

A-H

A-L

V03

03

A-H

A-L

C01

C01

02

A-H

A-L

02

103

BD01

01A-H

V06

01

A-H

A-L

C01

RatioController

V08

03

C01

1-MS-003-XV07

1-SS-004-2L

FIC

FIC

PI

PI

PI

TI

TI

TIC

BDI

TI

STEAM

ETHYLENE

E-8

Piping & Instrumentation Diagram of a Fixed Bed Catalytic Reactor


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5.3 Piping and Instrumentation Diagram (P&ID) Item List

Equipment List

Equipment Description

R01 Reactor Vessel

HE01 Heat Exchanger

Located inside reactor BD01 Bursting Disc

Pipe Code Explanation

1-SS-001-2L pipe

diameter

1-SS -001-2L the material of construction i.e. stainless steel

1-SS- 001 -2L pipe identification number

1-SS-001- 2L lagging is required

1-SS-001- X no lagging

Valves ListPipe Description

V01 Gate Valve (Ethylene Pipe)

V02 Control Valve (Ethylene Pipe)V03 Stop Valve (Ethylene Pipe)V04 Gate Valve (Steam Pipe)V05 Control Valve (Steam Pipe)

Pipes ListPipe Description

1-SS-001-2L Ethylene Inlet

1-CS-002-2L Steam Inlet1-MS-003-X Heat Exchanger Inlet and Outlet1-SS-004-2L Product Outlet


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V06 Stop Valve (Steam Pipe)V07 Gate Valve (Heat Exchanger)V08 Control Valve (Heat Exchanger)V09 Gate Valve (Heat Exchanger)

Instrument ListInstrument Description

FIC-01 Flow Indicator Controller

Ethylene PipePI-01 Pressure Indicator (Ethylene Pipe)

FIC-02 Flow Indicator Controller

Steam PipePI-02 Pressure Indicator (Steam Pipe )PI-03 Pressure Indicator (Reactor Inlet)TI-01 Temperature Indicator (Bed 2)TI-02 Temperature Indicator (Bed 1)

TIC-01 Temperature Indicator Controller

Heat Exchanger OutletTI-03 Temperature Indicator Product

Pipe

6.0 Hazard and Operability Study (HAZOP)

6.1 Introduction

The technique of Hazard and Operability Study, commonly known as a HAZOP has been used

and developed over a few decades in order to identify potential hazards and operability

problems caused by deviations in new and existing plants. It is best described as a procedure for

the systematic, critical, examination of the operability of a process. 15

It has become an important tool in predicting hazards and operability issues in the planning stage

and generally finds solutions to problems that could be very costly to resolve once the plant has

been commissioned.


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6.2 Basics of a HAZOP

Firstly a HAZOP team is required which consists of a Chairperson (experienced in HAZOP but

not directly involved in design who ensures that the method is followed correctly), Scribe

(documents outcomes) and other team members (people from varied disciplines such as

maintenance and designer).

A HAZOP systematically questions the sections of the process to establish the deviations that

can occur, once these are identified an assessment can be made upon the causes and

consequences of these deviations. If considered important, the necessary action is sought to

rectify them.

As this tool has been developed, certain guidewords and deviations have been investigated to the

extent that they are considered necessary for all HAZOPs. They are as follows:

Guidewords

Flow Temperature

Pressure Level

Deviation

Word Meaning

No The design intent does not occur (e.g. Flow/No),

Less A quantitative decrease in the design intent occurs (e.g.

Pressure/Less)

More A quantitative increase in the design intent occurs (e.g.

Temperature/More)

Reverse The opposite of the design intent occurs (e.g. Flow/Reverse)

Early Usually used when studying sequential operations, this would

indicate that a step is started at the wrong time or done out of

sequence

Late As Early


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Other The activity occurs, but not in the way intended (e.g.

Flow/Other could indicate a leak or product flowing where it

should not, or Composition/Other might suggest unexpected

proportions in a feedstock)

The assessment is carried out for all the deviations relating to each guideword. Other guidewords

and deviations are also carried out that relate will relate to every plant, for example a plant shut

down is an important deviation. These can be seen in the full HAZOP below for the ethanol

production process. Figure 4 below explains the systematic approach applied to a HAZOP.

Figure 10: Method in Conducting a HAZOP 16


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6.3 Process Line for HAZOP


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The chosen process step for the HAZOP carried out by our group is not from the section

considered in this report, for this reason the Piping and Instrumentation Diagram (P&ID) of the

chosen process is included below.

The HAZOP is to be carried out at the process line between the heat exchanger and the separator,

during this stage partial condensation has occurred at the heat exchanger where the ethylene can

be separated (in order to be recycled and fed into reactor). The ethylene is the only material from

the reactor which will remain in its vapour form, the ethanol and other by products will be

condensed into liquid form by the heat exchanger.

The HAZOP team is as follows:

Chairperson: Dr J Titiloye

Core Team: Irfan Badat

Mohammed Abu Jafor Onyeka Muehlhoezer

Theophilus Domobiyo

Danara Basbayeva

Scribe: Irfan Badat

The HAZOP checklist signed by the Chairperson (Dr J Titiloye) can be seen in Appendix D.


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6.5 HAZOP Ethanol Production Process

Guide Word Deviation HazardNo.

Possible Causes Consequences ActionNo.

Action Required

Flow No 1 Blockage Over-pressure 1.1 Install an pressure indicator and

alarmBack-pressure 1.2 Relief valve

1.3 Use by-pass line1.4 Ensure line can withstand

pressure1.5 Use correct material

Rupture of Heat Exchanger pipes 1.6 see 1.4 & 1.5No material into separator 1.7 Level Indicator in Heat

Exchanger Power failure 1.8 Arrange back-up power

arrangements

2 No reaction No feed into separator 2.1 See 1.7Pump Malfunction (due to nousage)

2.2 Consider 1.3

3 Valves Malfunction No flow 3.1 Refer to hazard 1Pump Damage 3.2 Ensure Regular MaintenanceBlockage 3.3 Refer to hazard 1

4 Pump Failure No flow 4.1 Refer to 3.2

connect control pump to controlroom

4.2 Low level alarm in the separator

4.3 Pressure indicator after pumpPressure Build up 4.4 Refer to hazard 1


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5 Leakage/LineFracture

Loss/Less Feed into Separator 5.1 Consider welded pipe work

Pressure Drop on the Pump 5.2 See 3.25.3 Electrical zoning to avoid static

charge5.4 Emergency procedures

Spillage 5.5 Spillage handling/maintenanceFire 5.6 Ensure regular fire drills

LESS 6 Partial Blockage Pump Damage 6.1 Consider mass flow recorder andindicator

Pressure 6.2 Consider level indicator insufficient separation/loss of product

6.3 Ensure proper maintenance of valves

7 Leakage Refer to hazard 5 7.1 Refer to hazard 5

8 Pump not operating

at full capacityRefer to hazard 2 8.1 Refer to hazard 2

9 Insufficient material

from reactor Excessive cooling in heatexchanger

9.1 Consider 1.7

10 Control Valve not

operating properlySeparator level too low 10.1 See 6.3

MORE 11 Over pressure/

pump malfunctionPipe Burst 11.1 Refer to hazard 5

High Feed level in separator 11.2 Consider relief valve in separator Flooding in separator 11.3 See 11.2Inefficient cooling at heatexchanger (not cooled down)

11.4

12 Using both pumps

(back up line)High feed in separator 12.1 see 11.2


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12.2 Install level indicator alarm

13 Control ValveMalfunction

Separator level high 13.1 Instal l rel ief valve in separator

13.2 See 6.3

REVERSE 14 Faulty Valve Pump not fitted correctly 14.1 Commissioning: practice runNo feed into separator 14.2 Use back up line using alarm that

are linked to control room

Pressure build up on line (back pressure)

14.3 Use one way valve

Damage heat exchanger 14.4 See 14.3Flow sheet error 14.5 SOP of pump isolation

LATER 15 Not Applicable Not Applicable Not Applicable

SOONER 16 Not Applicable Not Applicable Not Applicable

WHERE ELSE? 17 flow sheet error Swapping gate for drain valve 17.1 see 14.1

18 later f low Inlet spli t between both l ines 18.1 Consider flow indicator on bothlines so operators know both arebeing used

Guide Word Devia tion

Pressure MORE 19 Valves fully open Over pressure 19.1 Consider pressure indicator

19.2 Refer to hazard 11

20 Pump malfunction Refer to hazard 3 20.1 Refer to hazard 3

21 High mass flow rate Line eruption 21.1 Check piping pressure rating(SOP)

21.12 Refer to hazard 5Pump damage 21.2 Consider flow indicator linked to


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control roomTank over flow 21.3 Consider drain valve in separator

22 valve malfunction Temporary blockage 22.1 Refer to hazard 1

High temperature 22.2 Consider vent/pressure relief

LESS 23 Pump Malfunction Refer to hazard 3 23.1 Refer to hazard 3

24 Leakage Refer to hazard 5 24.1 Refer to hazard 5

25 By pass Line Reduced pressure 25.1 See 18.1

Temperature MORE 26 Heat exchanger

faultyHigh temp in pipe and separator 26.1 Consider temp control alarm

26.2 see 14.1

27 Incorrect coolant /temp

High temp in pipe and separator 27.1 See 26.1

27.2 Consider material analysis uponentry

28 Coolant flowrate

lower than desiredHigh temp in pipe and separator 28.1 Consider flow control

29 Climate change Possible feed vapourisation

(product loss)29.1 Consider lagging and insulation

30 Fire Explosion 30.1 Linked to fire station

Damage to valves & other equipment

30.2 Ensure tank fire proof

30.3 SOP training on fire handling


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31 Inaccurate tempgauge

Incorrect actions 31.1 See 14.1

LESS 32 Too much coolant Frozen pipe 32.1 Check pipe freezing points 33 Climate change Less product 33.1 Install temp indicator control

33.2 see 29.133.3 tracing on lagging

33.4 consider heat tracer on linelinked to alarmGuide Word Devia tion

LEVEL HIGH 34 Refer to FLOW

(hazard 11 - 13)Refer to FLOW (hazard 11 - 13) 34.1 Refer to FLOW (hazard 11 - 13)

LOWER 35 Refer to FLOW

(hazard 6 -10)Refer to FLOW (hazard 6 -10) 35.1 Refer to FLOW (hazard 6 -10)

MORE 36 Not Applicable Not Applicable Not Applicable

LESS 37 Not Applicable Not Applicable Not Applicable

MISSING

COMPONENT38 Diethyl Ether

vapourisesEnters back into reactor 38.1 Soluble in ethanol so can be

separatedIncomplete separation

39 Missing coolantfluid

No temp reduction in heatexchanger

39.1 Refer to hazard 27

EXTRA 40 Impurities Product contamination 40.1 Consider quality analysis

DIFFERENT/WRONG STREAM

41 Not Applicable Not Applicable 41.1 Not Applicable

CONTAMINATION 42 Refer to hazard 40 Refer to hazard 40 42.1 Refer to hazard 40


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EXTRA PHASE 43 Not Applicable Not Applicable 43.1 Not Applicable


Sequence Not applicable, for batch Processes only

OTHER START UP 44 No reactants No product 44.1 Refer to SOP

44.2 Monitor process via control room

SHUT-DOWN 45 45.1 see 44.145.2 see 44.2

RELIEF SYSTEM 46 High pressure and

tempRefer to TEMPERATURE andPRESSURE

46.1 Refer to TEMPERATURE andPRESSURE

POWER/SERVICE

FAILURE47 Power Cut Product Loss 47.1 see 1.8

CORROSION/

ERROSION48 Roasting of

equipmentDeposit 48.1 see 3.2

Leakage

ContaminationEquipment damage

MATERIAL OF

CONSTRUCTION49 Corrosion Material Loss 49.1 see 3.2

Equipment Damage 49.2 Ensure material conforms toUSA standard

TOXIC/ASHYXIA 50 Inhalation- health problems 50.1 PPE utilisation

50.2 Refer to MSDS


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MAINTENANCE 51 Not completed Pipe worn out 51.1 Consider isolation valve

Equipment breakdown

DOUBLE VALVES 52 Not Applicable Not Applicable 52.1 Not Applicable

VALVE ACCESS 53 Inappropriatelocation

Time consuming to resolve 53.1 Ensure they are all accessible


INSTRUMENTS,TRIPS, TESTING

54 Faulty form supplier Not in operation 54.1 Refer to SOP

54.2 Ensure regular checks

FIRE 55 spark Fire 55.1 Refer to hazard 30

56 bad electricalzoning

Fire 56.1 Refer to hazard 30

STATIC

ELECTRICITY57 Spark Fire 57 Refer to hazard 30

NOISE 58 Noise pollution Communication and hearing

problems58.1 Consider noise meter and link to

control room58.2 Silencer

IONISING

RADIATION59 Not Applicable Not Applicable 59.1 Not Applicable

SAMPLING 60 Refer to hazard 40 Refer to hazard 40 60.1 Refer to hazard 40

THERMALRADIATION

61 refer to tempdeviation


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SPARES 62 Equipment Damage Process not running 62.1 Ensure adequate spares

WHAT ELSE 63 Operator errors Process Errors 63.1 Regular operator training63.2 Manual handling training


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7.0 Economic Appraisal

The main purpose of investing money into a chemical plant is to make a profit. Inorder to ensure that a profit would be made from a certain process, it is essential to

carry out an economic evaluation of the process.

By comparing the capital and operating costs of alternative processes, a decision

could be made in which process to use. This is usually done for simple projects and

can also be done for simple choices between alternative processes and which

equipment to use.

Design engineers need to be able to estimate a quick, rough cost of process equipment

and alternative process designs for project evaluation. Before the profitability of a

plant can be assessed, an estimate of the investment required and the cost of the

production are needed.

Various components make up the capital and operating costs of a plant. This section

calculates an estimate of the various costs and the net cash flow to prove or disprove

its economic viability.

7.1 Estimating Capitol Cost

Capital Costs are those required to bring a project to a commercially operable status

and are a one off payment. These include land, buildings, construction and equipment.There are various methods to estimate the capitol cost, for the purpose of this design

two have been chosen. The Wilson Method and another devised by Zevnik and

Buchanan.


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7.1.1 Wilson Method

This method is based on the number of plant items, its principle is that the average

cost of one item of equipment is a function of the process size and complexity.

The following equation is what Wilson devised to be the important factors in

obtaining an estimate of the capitol cost.

Capitol Cost

Installation factor

Plant items

Average cost of plant items

Construction material factor

Pressure factor

Temperature factor

However, this equation relates to the capitol cost in 1971. An adjusted UK plant cost

index has been calculated that gives a more accurate cost for 2004.

The updated version has an added factor of 27.3.

Plant Items (f)

There are 4 plant items which include the reactor, heat exchanger, adsorber and

distillation column.

Normally a heat exchanger is not included but due to the high usage and importance

of it, it has been included.


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Average Unit Cost (AUC)

V is the capacity in tonnes/year = 100,000

Installation Factor (f)

Using the AUC value the installation factor is found graphically using Appendix E.

f = 1.85

Construction Material Factor (FM)

The material chosen is stainless steel due to the high pressure and temperature in the

reactor.

FM = 1.3

Pressure Factor (FP) and Temperature Factor (FT)

These are also worked out graphically (see Appendix D)

If P = 60 bar FP = 1.1

If T = 300 oC FT = 1.1

Capitol Cost


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7.1.2 Zevnik and Buchanan Method

This is a method where the following conditions need to be met:

Capacity above 4500 tonnes / year

Temperature and Pressure above ambient

Petrochemical type process

The capitol cost is calculated using the following equation.

Where

Capitol Cost

Number of process steps

Capacity

Maximum Process Pressure

Maximum Process Temperature

Construction Material Factor

Number of Process Steps (N)

This is the same as above where there are 4 steps (reactor, heat exchanger, adsorber,

distillation column).

Capacity (Q)

Ethanol production specified at 100,000 tonnes per year

Factor of Construction Material, Pressure and Temperature

(X)

Stainless Steel is the chosen material which has a factor of 0.1.

The pressure used here is the design pressure of 59.215 atm.

The temperature used will be the reactor temperature after reaction of 593 K.


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Capitol Cost

7.2 Estimating Raw Material Cost

There are 2 reactant raw materials (ethylene and steam) and the coolant (brine) to be

calculated.

Ethylene

Cost per kg = 0.46 17

Mass flowrate = 7763.975 kg/hr

Total operating time = 8000 hours

Steam

Cost per kg = 7 per tonne 22

Mass flowrate = 4991.127 kg/hr



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Brine

This has been calculated as a function of its chemical composition of 25% sodium

chloride and 75% water.

NaCl Cost per kg = 0.025 /kg

Mass flowrate = 130,000,000 kg/yr

Water Cost per kg = 0.05 /tonne 22

Mass Flowrate = 390,000,000 kg/yr

Ethanol

The cost of ethanol has been taken from two sources, the first from the literature

which was an average value taken in 1998 and the second is a current price. The

updated price was used as the initial value used was significantly higher than

expected.

Price 1

Cost per kg = 4.2 22

Mass flowrate = 12,500 kg/hr



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Price 2

Cost per $ gallon = 3.7 18

To convert this into a mass form the following calculation was done.

19

Multiply the volume of 1 US gallon by the density of ethanol to work out the mass of

1 US gallon.

Density of ethanol = 789 kg/m 3

Therefore, the cost of 2.986 kg is $3.7

7.4 Estimating Operating Cost

Operating cost is the cost incurred to produce the product. It is important as it is

needed to judge the viability of the plant and make choices between possible processroutes if this is the stage of the project. They are two types of operating costs, Fixed

and Variable.

Fixed operating costs do not vary with production rate. These are bills that have to be

paid whatever the quantity produced.

Variable operating costs are dependant on the amount of product produced.

Two methods were used to calculate the operating costs.


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The first was using a typical plant cost breakdown based on the cost of the raw

materials.

The second is based on 3 factors, the raw material, fixed capitol cost and labour cost.

Method 1

Assumption: Use a typical plant operating cost breakdown

Raw material costs are approximately 31% of the operating costs.

Figure 11: Breakdown of a typical plant costs 20

Units % of Total Cost CostRaw Materials 31 32,120,432

Labour 11 11,397,573Supervision 2 2,072,286Maintenance 2 2,072,286

Plant Supplies 1 1,036,143Royalties and Patents 2 2,072,286

Utilities 8 8,289,144Payroll Overheads 2 2,072,286

Laboratory 2 2,072,286Plant Overhead 9 9,325,287Packaging 1 1,036,143Shipping 1 1,036,143

Depreciation 4 4,144,572Property Taxes 1 1,036,143

Insurance 1 1,036,143Administration 4 4,144,572

Sales 11 11,397,573Research 5 5,180,715

Finance 2 2,072,286Total 100 103,614,296Method 2

This method utilises the fixed capitol cost so both the Wilson and Zevnik & Buchanan

values will be used to calculate the operating costs.

Figure 12: Operating Cost Calculation 21


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Wilson Zevnik and

BuchananVariable Costs (A)

Raw Materials 32,120,432 32,120,432Miscellaneous

Materials

10% of M 158,250 634,951

Shipping Negligible Negligible

Fixed Costs (B)Maintenance(M) 10% of capital cost 158,249 6,349,506Operating Labour

(OL)

Estimate 600,000 600,000

Laboratory 20% of OL 120,000 120,000Supervision 20% of OL 120,000 120,000

Plant Overheads 50% of OL 300,000 300,000Capital Charges 15% of capital cost 2,373,743 9,524,258

Insurance 1% of capital cost 158,249.5 634,951Local Taxes 2% of capital cost 316,499 1,269,901

Royalties 1% of capital cost 158,249.5 634,951

Other(C)Sales Expense 20% of A&B 7,601,583 10,461,790

Total 45,609,501 62,770,738

Having compared all the values (Capital Cost, Raw Material Cost and Operating

Cost), the decision to exclude price 1 of ethanol and operating cost method 1 has been

taken because of irregularities in the cash flow table where there were significant

profits or significant losses.

The cash flow table which produces the most reasonable results is price 2 of ethanol

and the Zevnik and Buchanan cost methods.

7.5 Assumptions

Certain assumptions are important to carry out the analysis. They are as follows:

First 2 years are for construction

Capital Cost split 50:50 between first 2 years

Working Capital is included in Capital Cost


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Operating Plant Life is 10 years

Operating at full capacity

Supply = Demand

7.6 Annual Income

A major problem with every business is its ability to sell its products, the assumption

resolves this. However, as this is optimistic the ethanol will be sold at a more

competitive price during the plant life of 10 years.

From this the annual profit can be calculated (excluding capital cost).

7.7 Pay Back Time

From the information collected, a cash flow table can be constructed from which the

payback time of the project can be calculated.

Figure 13: Cash Flow Table

Year Capital Cost

()

Operating

Cost(/yr)

Income

(/yr)

Net Cash

Flow(/y)

Cumulative Cash

Flow(/yr)0 -31,747,528 -31,747,528 -31,747,528

1 -31,747,528 -63,495,056 -63,495,056

2 62,770,738 80,000,000 17,229,262 -46,265,795

3 62,770,738 80,000,000 17,229,262 -29,036,533

4 62,770,738 80,000,000 17,229,262 -11,807,271

5 62,770,738 80,000,000 17,229,262 5,421,990

6 62,770,738 80,000,000 17,229,262 22,651,252

7 62,770,738 80,000,000 17,229,262 39,880,514

8 62,770,738 80,000,000 17,229,262 57,109,7759 62,770,738 80,000,000 17,229,262 74,339,037


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10 62,770,738 80,000,000 17,229,262 91,568,299

11 62,770,738 80,000,000 17,229,262 108,797,560

12 62,770,738 80,000,000 17,229,262 126,026,822

Figure 14: Cash Flow Forecast

The table and graph show the pay back time of the investment into the project to be

accomplished during Year 6. As the plant life is 12 years, this means that there will be

7 years of 100% profit and approximately 33% profit in the year where the project

break evens.

The Return on Investment (ROI) accounts for the loss of value of investment and in

this project the cumulative net cash flow (NCF) is accounted for.

7.8 Net Present Value and Rate of Return

Net present value (NPV) analysis is used to calculate the present worth of future

earnings; it is sensitive to the interest rate assumed. By calculating the NPV at

different interest rates, is it possible to find a cumulative NPV at the end of the project

which is zero. This particular rate is called the Discounted-cash flow rate of return


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(DCFRR) and is the measure of the maximum rate that the project could pay and still

break even by the end of the project life.

NPV is calculated as follows:

NCF Net Cash Flow

i interest rate assumed

n year

See Appendix F for cash flow table of present values.

By plotting the cumulative NPV of various interest rates, the DCFRR can be

calculated of the project.

The table below is the Cumulative Present Values taken from Appendix F.

Figure 15: Present Values at Various Interest Rates

Interest Rate (%) Cumulative PV1 82,244,0795 44,079,451

10 12,261,607

15 -8,549,324

Figure 16: Discounted Rate of Return

The graph above shows that the DCFRR is approximately 12.5%.


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This means that this project can offer potential investors a 12.5% return and break

even.

7.9 Raw Material and Product Cost Variation

The process uses ethylene which is an oil derived substance and with the instability of

the market. The following cash flow table has been done to include a 2% increase in

its cost per annum from Year 3.

Figure 17: Cash Flow Table

Year Capital Cost

()

Operating

Cost(/yr)

Income

(/yr)

Net Cash

Flow(/y)

Cumulative Cash

Flow(/yr)0 -31,747,528 -31,747,528 -31,747,528

1 -31,747,528 -63,495,056 -63,495,056

2 62,770,738 80,000,000 17,229,262 -46,265,795

3 63,456,453 80,000,000 16,543,547 -29,722,247

4 64,155,881 80,000,000 15,844,119 -13,878,129

5 64,869,298 80,000,000 15,130,702 1,252,573

6 65,596,984 80,000,000 14,403,016 15,655,589

7 66,339,223 80,000,000 13,660,777 29,316,366

8 67,096,307 80,000,000 12,903,693 42,220,059

9 67,868,533 80,000,000 12,131,467 54,351,527

10 68,656,203 80,000,000 11,343,797 65,695,32411 69,459,626 80,000,000 10,540,374 76,235,697

12 70,279,118 80,000,000 9,720,882 85,956,579

Though the project still remains profitable, it can be clearly seen that the decrease in

annual income from the value in year 2 is significantly high. The annual income has

been halved from that original figure. However, even with these conditions the project

is profitable and the only concern would be in extending its life.

This concern can be looked at realistically by taking into account price increases of

ethanol. As stated earlier, a more competitive price will be used to gain market share.The following table shows the rate at which ethanol price would need to increase to

remain as profitable.

Figure 18: Cash flow table with material and product cost changes (see Appendix G)

Year Capital Cost

()

Operating

Cost(/yr)

Income

(/yr)

Net Cash

Flow(/y)

Cumulative Cash

Flow(/yr)

0 -31,747,528 -31,747,5281 -31,747,528 -63,495,056 -63,495,056


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2 62,770,738 80,000,000 17,229,262 -46,265,795

3 63,456,453 80,760,000 17,303,547 -28,962,247

4 64,155,881 81,527,220 17,371,339 -11,590,909

5 64,869,298 82,301,729 17,432,430 5,841,522

6 65,596,984 83,083,595 17,486,611 23,328,133

7 66,339,223 83,872,889 17,533,666 40,861,7998 67,096,307 84,669,682 17,573,375 58,435,174

9 67,868,533 85,474,044 17,605,511 76,040,685

10 68,656,203 86,286,047 17,629,844 93,670,529

11 69,459,626 87,105,764 17,646,138 111,316,667

12 70,279,118 87,933,269 17,654,151 128,970,817

The cost of ethanol needs to increase by 0.095% per year.

Break-even is achieved in the same year as initially projected (year 6).

The DCFRR increases by approximately by 1% which means the return for investorswould be about 13.5 % as can be seen below.

Figure 19: Discounted Rate of Return

7.10 Conclusion

This ec

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