Report Design Project

8/2/2019 Report Design Project

http://slidepdf.com/reader/full/report-design-project 1/32

Programmes in Advanced Process Design

MSc DesignProject

2011-2012Design of Heat Exchanger

Network -

14th March, 2012

Submitted To: Simmon Perry

Project Advisor

Submitted By: Farhan Hafeez

ID: 8197194



MSc Design Project 2011-2012

Heat Exchanger Network Design Page 1

Contents1. Introduction .............................................................................................................................................. 3

2. Site Description ......................................................................................................................................... 4

2.1 Plant A: ................................................................................................................................................ 4

2.2 Plant B: ................................................................................................................................................ 4

2.3 Plant C: ................................................................................................................................................ 4

3. Process Data Extraction ............................................................................................................................ 5

3.1 Stream Identification: ......................................................................................................................... 5

3.2 Temperature-enthalpy profile: ........................................................................................................... 5

3.3. Mixing: ............................................................................................................................................... 5

3.4 Effective Temperature: ....................................................................................................................... 6

.................................................................................................................................................................. 63.5 Soft Constraints: .................................................................................................................................. 6

4. Energy Target: ........................................................................................................................................... 7

4.1 Delta Tmin selection: ............................................................................................................................. 7

4.2 Composite Curves ......................................................................................................................... 8

4.3 Problem Table ............................................................................................................................... 9

4.4 GrandCompositeCurve: ............................................................................................................... 11

4.5 Utilities selection ............................................................................................................................... 11

4.5.1 Plant A ........................................................................................................................................ 11

4.5.2 Plant B ........................................................................................................................................ 12

4.5.3 Plant C ........................................................................................................................................ 12

4.6 Utilities Summary .............................................................................................................................. 13

5. Economics ............................................................................................................................................... 14

5.1 Capital Cost Estimation ..................................................................................................................... 14

5.2 Heat Exchanger Area Targeting ........................................................................................................ 15

5.3 HEN Area Cost Calculations............................................................................................................... 17

5.4 Energy Cost Calculations ................................................................................................................... 18

6. Capital – Energy cost Trade-off and ∆Tmin............................................................................................... 21

7. Heat Exchanger Network Design ............................................................................................................ 22

7.1 Plant Data .......................................................................................................................................... 22

7.2 HEN Design Methods ........................................................................................................................ 23





7.3 HEN Design Plant A ........................................................................................................................... 24

7.3 HEN Design Plant B ........................................................................................................................... 27

7.4 HEN Design Plant C ........................................................................................................................... 29

7.5 Sensitivity Analysis ............................................................................................................................ 30

8. Conclusion ............................................................................................................................................... 30

9. References: ............................................................................................................................................. 30





1. IntroductionThis report is written in reference to Research Techniques and Methods (skills) module for

MSc in Advanced Process Design. The work assigned in this course is to design a total site system and is

named as MSc Design Project.

Groups comprising of four students have been made to accomplish this task and the design

project is divided into two main sections based on individual member’s work and group work. The

individual member has to design heat exchanger network for each plant in the total site and group task

is to design site utility system.

This report covers the design of heat exchanger network for each plant on total site as an

individual task for the design project.





Figure 3: Flow sheet and Stream Data Plant C

2. Site DescriptionThe total site includes three plants named as Plant A, B & C respectively.

The data provided for each plant is limited to feed and product streams, process streams temperatures

and heat flows associated with each stream. Thus the thermodynamic analysis of each plant is based on

given data and each plant is discussed in detailed in the following sections.

A brief introduction of each plant is as follows:

2.1 Plant A:

Plant is consisted of two overall feed streams and two product streams with number of

intermediate streams from one unit to the other. Each stream has its own heating or cooling duties with

specified inlet and out temperatures from heat exchangers. The flow sheet for plant A is shown below:

2.2 Plant B:

Plant B is the most complex plant in the given case which involves four product streams

from two distillation columns in sequence and a feed stream to first distillation column after passing

through two more process units. The flow sheet of Plant B is shown below:

Figure 2: Plant B Flow Sheet

2.3 Plant C:

Plant C has two cold streams entering into a reactor with a single outlet hot stream

which is being cooled down. Moreover there is downstream process involves in plant C which requires

10MW of MP steam and 48MW of LP steam which are being supplied at 19.07and 3.62bara respectively.

Figure 1: Flow sheet Plant A





3. Process Data ExtractionFollowing basic rules have been employed in extracting data for each plant:

3.1 Stream Identification:

It is important to identify the streams which are to be included in

heat integration analysis from a process thus only streams with known temperatures and heat flows are

considered for heat integration analysis for each plant i.e. in fig 4 of plant A, the stream circled red is not

considered due to unknown material and energy flow.

Similarly in plant B the red circled streams are not considered in analysis due to unknown or not fully

defined stream properties.

3.2 Temperature-enthalpy profile:

As the physical properties data for the streams is

unknown so we have assumed that the heat capacity of the streams in each plant is constant over the

range of temperature change with gives straight line when plotted these streams on temperature-

enthalpy diagram for heat exchanger area targeting unless until specified in the given data as in case of

one stream in plant C which is shown as follows:

In this case we have assumed three streams based on heat capacity change with temperature

change from 100 to 264C, 264 to 264C with phase change and 264 to 650C of a single stream.

The purpose of selecting this data is to avoid significant errors in heat integration analysis and heat

exchanger area targeting.

3.3. Mixing:

Mixing of two streams causes hidden heat transfer and thus this fact is considered in data

extraction for our process. This problem can be seen in plant A feed streams heating as follows:

Figure 5: Plant A under defined stream Figure 4: Plant B Undefined and Unknown Streams

Figure 6: Plant C Stream 2 Temperature-Enthalpy Data





Here we have assumed heating of Stream 1 from 58

to 150C and stream 2 from 25 to 150C instead of 58 to

165C and 25 to 80C respectively to avoid heat of mixing

and thus error in heat integration analysis.

However the CP value for each stream is calculated with initial temperature change from 58 to165C and 25 to 80C respectively which is required to calculate heat load from heating these streams

from initial 58 and 25C temperature to final 150C temperature.

Stream No Type TS – ⁰C TT – ⁰C ∆H – kW CP – kW/⁰C

1 Cold 58 165 3380 31.59

2 Cold 25 80 185 3.36Table 1: Stream Data Plant A

These values of CP are used to calculate heat load of heat exchangers for each stream to be heated to

150C as follows:

Stream-1: Q = CP x (TT – TS) Similarly for; Stream-2: Q = CP x (TT – TS)

Q = 31.59 x (150-58) Q = 3.36 x (150-25)Q = 2906.28 kW. Q = 420 kW.

3.4 Effective Temperature:

Data is extracted from each plant based on availability of heat at its effective

temperature which can be used for heat recovery opportunities, e.g. the feed entering into the process

unit is at 380C in a section of plant B as follows:

However the unit outlet stream is at 400C which is being cooled

to 225C. Thus the availability of heat for heat recovery

opportunity is at 400C temperature and thus by extracting databased on this rule for every stream in each plant has provided

an opportunity to get error free heat integration analysis.

3.5 Soft Constraints:

Considering the flexibility of product storage temperature as a soft constraint we have

extracted the data for plant A and B following streams at different temperatures than given to reducethe complexity of heat exchanger network design which is shown in the later part of the report.

Figure 7: Plant A Mixing Streams

Figure 8: A Stream Data from Plant B

Figure 9: A Stream from Plant A and B -Soft Constraint





Following is the streams data from each plant for heat integration analysis based on above rules.

Plant A:

Stream No Type TS – C TT – C ∆H – kW CP – kW/C

1 Cold 58 150 2906.28 31.59

2 Cold 25 150 420 3.36

3 Cold 35 225 2150 11.324 Hot 258 257 4270 4270

5 Hot 280 110 5760 33.88

6 Hot 80 55 5200 208

7 Hot 45 35 2600 260

8 Hot 450 80 4193 11.33Table 2: Stream Data Plant A

Plant B:


1 Cold 25 380 30350 85.5

2 Cold 210 280 6475 92.53 Cold 350 351 8400 8400

4 Cold 278 279 4960 4960

5 Hot 400 225 28880 165

6 Hot 40 39 4080 4080

7 Hot 184 183 7430 7430

8 Hot 350 200 9700 64.66

9 Hot 200 199 1760 1760

10 Hot 278 67.5 6084 28.9Table 3: Stream Data Plant B

Plant C:


1 Cold 25 650 96000 153.6

2 Cold 100 264 15600

264 264.1 34800

264.1 650 21000

3 Hot 650 170 14800 30.833Table 4: Stream Data Plant C

4. Energy Target:Now we evaluate the energy target for each plant with the above data. We will use each

plant’s composite curve, grand composite curve and problem table analysis using SPRINT (CPI software)

to determine the energy targets for respective plant.

4.1 Delta Tmin selection:

To generate above mentioned tools for energy targeting using SPRINT

software we need an initial value of minimum approach temperature i.e. ∆Tmin. The value of ∆Tmin sets

the energy recovery target between hot and cold streams of the process which is quite evident from





composite curves of the process and is discussed in following section. We select an initial value of 10C

for each plant based on experience and high energy cost these days.

4.2 Composite Curves:

The composite curves are produced by plotting all the hot and cold

streams in a process on temperature-enthalpy diagram and the enthalpy change is the sum of individual

stream enthalpy change in a certain temperature range. Two different curves are plotted one is termed

as hot composite curve which contains all the hot streams in a process and second is termed as cold

composite curve which contains all the cold streams of the process. Composite curve is a tool which is

used to determine the process-to-process heat recovery, hot and cold utility targeting at certain

minimum temperature difference (∆Tmin) between hot and cold streams.

Following are the composite curves of our system for individual plant at ∆Tmin 10C.

The red and blue curve in above figure is hot and cold composite curves respectively for plant A.

It is noticeable from the curve above that cold stream does not extend beyond the start of hot stream

and ends at same enthalpy value indeed. This type of problem is called as threshold problem where

there is no need of either hot or cold utility in the process. Here in this case the threshold value for both

the curves is at start of hot composite curve of the process thus there is no need of hot utility in this

system as all the heat required for the process will be provided by process-to-process heat recovery,

whereas relatively large quantity of cold utility is required. The overlap area of the hot and cold

composite curve shows the process-to-process heat recovery, In our case the heat recovery comes out

to be 5475kW.

Figure 10: Composite Curve Plant A

Figure 11: Composite Curve Plant B





For plant B the composite curve shows that there is a huge process-to-process heat recovery

opportunity however in this case the hot composite curve extends beyond the start of cold composite

curve and vice versa. Thus we need both cold and hot utility for our process. The value for process-to-

process heat recovery in this case comes out to be 45821kW.

For plant C is also a threshold problem at ∆Tmin of 10⁰C with no cold utility requirement and

process heat recovery is 143800kW.

4.3 Problem Table:

The composite curve gives us the potential of process-to-process

heat recovery and hot and cold utility requirement but it does not clearly shows the temperatures at

which these utilities are required and exact values of utilities load as there could an error in reading

these values from graph. Moreover it’s difficult to construct composite curves without computer aided

programs. Thus to overcome this problem we haveused another technique which gives same results

with more accuracy and both heat loads with

required temperatures for utilities targeting.

In constructing problem table we follow the same

technique of dividing streams into temperature

intervals as in case of producing composite curves

but in this case we shift hot streams by -∆Tmin/2

and cold streams by ∆Tmin/2 to keep feasible

temperature difference for heat transfer. Theenthalpy change is then calculated in each interval

and finally the heat is cascaded down from high

temperature side to low temperature side of the

process. We get some negative heat flows in the

process of heat cascading down which is not

feasible for heat flow thus we add external heat

Plant A - ∆Tmin 10C

Interval Temperature* Enthalpy

[C] [kW]

445 0

275 1926.1

253 2920.72

252 7235.93

230 8230.55

155 10773.1

150 10767.8

105 10210.7

75 8822.97

63 10763.9

50 13277.2

40 13130.5

30 16546.9

Table 5: Problem Table Plant A

Figure 12: Composite Curve Plant C





equal to highest negative value in the system to make it at least zero and the temperature at which the

heat flow value becomes zero is termed as pinch temperature. This external heat is actually the hot

utility target and the heat left at the end of the table is cold utility target. Adding ∆Tmin/2 in highest

temperature in the table gives us the required temperature of the hot utility and subtracting ∆Tmin/2

from lowest temperature gives the required temperature for cold utility.

SPRINT software is again used to construct the problem table for each plant.

The threshold problem for plant A is also evident from problem table as the

Value of external heat requirement for process heating is zero at the shifted temperature of 445C.

It is the pinch temperature of the process which divided the process into heat source and heat sink

which will be discussed further in following part of the report. Furthermore the maximum cold utility

temperature required for the process is at 25C which is achieved by subtracting ∆Tmin/2 from lowest

temperature in the problem table corresponding to minimum cold utility requirement i.e. 16546.7 kW.

In case of plant B the pinch

temperature is at 355C and 4363.65kWof hot utility is required at the

temperature of 400C. (Adding ∆Tmin/2 in

highest temperature in the table) The

area below the pinch is heat source

area and it require cold utility which in

this case is 12124.4kW at 25C.

(Subtracting ∆Tmin/2 in lowest

temperature in the table).

Similarly for plant C the minimum hotutility temperature comes out to be

660C from problem table and there is

no need for cold utility requirement as

all the process streams lie above the pinch region. The heat duty

of hot utility is found to be 23600kW.

Plant B -

Interval Temperature* Enthalpy

[C] [kW]

395 4363.65

385 6013.93

356 8320.46

355 0

345 795.356

285 9447.49

284 9499.2

283 4590.9

273 5107.92

220 9380.03

215 8957.91

195 9119.46

194 10822.9

179 9974.03

178 17347.4

62.1 10788.7

35 8471.86

34 12466.4

30 12124.4

Plant C

Interval

Temperature* -

C Enthalpy – kW

655 23600

645 21518.5

270 55807.8

269 21153.7

165 26443.3

105 11520

30 0

Table 6: Problem Table Plant B

Table 7: Problem Table Plant C





4.4 GrandCompositeCurve:

After setting targets for process-to-process heat recovery

and hot and cold utility targets with their required temperature levels we need to select appropriate

utilities for the process. This objective is attained using grand composite curve which is obtained by

plotting problem table on temperature-enthalpy curve. The grand composite curve presents true

interface between the utilities and process conditions and clearly shows the process division into

heat sink and heat source regions in a graphical manner. The hot utility is supplied in heat sink area

and cold utility is supplied in heat source area. Moreover grand composite curve is used to target

multiple utilities in a process. Discussion on each plant grand composite curve is incorporated in the

following section of utilities selection.

4.5 Utilities selection:

SPRINT software is used again to produce grand composite curves for each plant as follows:

Figure 13: Balanced Grand Composite Curve Plant A

4.5.1 Plant A

Grand composite curve of plant A (only red curve in above figure) shows that sufficient

heat is available at higher temperatures so steam generation can be considered. Maximum saturated

steam temperature that can be achieved in plant A is around 250 which can be superheated to 400C at

∆Tmin of 10C, we need saturated MP steam at 19.07bara for downstream process heating of plant C so

we have considered to produce MP stream in plant A.MP steam produced in plant A is at 20 bar

Saturated conditions with 212.37C saturation temperature. MP steam pressure is considered to be a bar

above the required pressure in downstream process to accommodate pressure and temperature losses

and it is assumed that there is no more than one bar pressure drop in steam supply to downstream

process. The utilities targets obtained from balanced grand composite curve are as follows:

MP Steam = 9673.19kW and Cooling water = 6873kW





4.5.2 Plant B

The balanced grand composite curve for plant B is shown in the following figure:

Similar to plant A the plant B has also potential to use steam as cold utility and here in this case

we have considered to produced MP steam at same level as at in plant A production and for plant C

downstream process heating. The utility targets for plant B from balanced grand composite curve are:

MP Steam = 8471kW and Cooling water = 3652kW.

Hot utility is also required in plant B at minimum temperature of 360C with a duty of 4363kW.

This temperature is high enough to make steam heating unsuitable because maximum steam saturation

temperature (supercritical temperature) at supercritical pressure is 372.15C which is practically

infeasible to use for process heating, moreover superheated steam at high temperature is also not

feasible for process heating due to low heat transfer coefficients and local condensation at lowertemperatures. For these reasons hot oil can be used as hot utility in plant B. Hot oils can achieve higher

temperature at relatively lower pressures and can be used both in vapor and liquid phase for process

heating. Thus hot oil at 400C will be used for this process heating. The return temperature of hot oil is

also vital in setting energy targets and it is considered to be at hot stream pinch temperature of 360C for

our case.

4.5.3 Plant C

The balanced grand composite curve for plant C is shown as follows:

Figure 14: Balanced Grand Composite Curve Plant B

Figure 15: Balanced Grand Composite Curve Plant C





Two hot utilities have been used in this case one is flue gas at 750C and other is LP steam at

4 bar saturated conditions to produce balanced grand composite curve. It is clearly seen from the

balanced grand composite curve of plant C that there is a demand of hot utility at much higher

temperature but with relatively small heat load. The big heat pockets in the process have potential to

supply remaining high level heat. Moreover it is the most energy intensive process as total hot utility

requirement with maximum process-to-process heat exchange is 23600kW. Around 12905kW of heat

can be supplied by flue gas and remaining 10695kW can be provided by LP steam but we will not use LP

steam as a major portion of flue gas would be wasted at lower heat level so instead of using LP steam at

lower temperature level we can utilize lower temperature level flue gas heat left after exchanging at

higher temperature for heating. And if we look at the stream data of plant C then it is quite evident that

all the heat required for this plant is to be supplied to cold feed streams and the only heat source(hot

stream) is product stream which put a significant constraint on maximum process-to-process heat

recovery as more than 167400kW of heat would be required at least 660C for plant start-up which is to

be provided by some outside heat source. Thus for these practical constraints we have divided the plant

into two sections as follows:

Thus by dividing Plant C we can use furnace flue gas for cold stream heating and hot stream heat will

be used for high pressure steam generation which can be used for power production and plant C

downstream LP steam requirement, it can also supply heat to hot oil for plant B heating as heat is

available at high temperature and in sufficient quantity.

The MP steam requirement for plant C will be provided by MP steam generated in Plant A and B.

4.6 Utilities Summary

The following table contains the utilities selected for each plant:

Plant Hot Utility

Requirement -

kW

Cold Utility

Requirement - kW

Hot Utility

& Load - kW

Cold Utility

& Load – kW

A 0 16546 - MP Steam – 9673

CW – 6872

B 4363 12124 Hot Oil – 4363 MP Steam – 8472

CW – 3652

C 167400 143800 Flue Gas - 167400 HP Steam – 143800Table 8: Utilities Data - Plant A,B and C

Cold Streams

Hot Stream

Plant C +

Figure 16: Plant C Section





5. EconomicsOne of the most important factors to be considered in designing a process is the

process economics. There are number of parameters to be evaluated in the economic analysis of a

process but we will consider the capital cost of heat exchanger network and energy cost for our

processes in this part of report. In second part of the design project includes the capital cost estimation

of other major equipment in total sight utility system.

5.1 Capital Cost Estimation

Capital cost involves the cost of constructing a new plant or a piece of

equipment. There are five generally acceptable methods to estimate capital cost in process industry as

follows:

a. Order of Magnitude Estimate.

b. Study Estimate.

c. Preliminary design Estimate.

d. Definitive Estimate.e. Detailed Estimate.

A brief introduction of each method is carried out in the following table:

Method Main Features

Order of Magnitude Estimate This estimate is usually based on the previous cost

data of same process and it involves the whole

process estimate which is then corrected by using

scaling factor.

Study Estimate The approach to this method is based on order of

magnitude estimate method but in this case costestimation of major equipment in a process based

on their approximate sizing is carried out.

Preliminary Design Estimate This estimate requires more accurate equipment

sizing with some approximation of piping,

electrical and instrumentation requirements

Definitive Estimate This estimate is based on preliminary specification

of major equipment, utilities, piping, electrical and

instrumentation.

Detailed Estimate This Estimate is as close to actual cost as possible

and is based on detailed engineering of the

equipment, utilities etc. with vendors quotes.Table 9: Description of Capital Cost Estimation Techniques

We have selected to do the cost estimation of heat exchanger network based on Study Estimate due to

limited time and data resource.

Cost data available for equipment would be out of date, based on some years ago and using this data

could result in significant error in cost estimation. To take this thing into account we use “Escalating

Factor” which is defined as below:





A =

………………..eq-01

Where,

CEPCI is the chemical engineering plant cost index, it is published on monthly basis and includes major

equipment cost indices. CEPCIbase year is the year in which the equipment given cost is available.

There are other similar cost indices available in the market which can be used for updating the cost of equipment using same formula as given in equation 04,however we have data available for heat

exchangers based on CEPCI so our calculations are based on this index.

Thus by using escalating factor the cost of equipment in current year is calculated as;

Ccurrent year = Cbase year A

Where

Ccurrent year and Cbase year are the cost of equipment in current year and base year respectively and;

We might encounter a problem where the base year equipment cost would be based on certain capacity

which might be different from current equipment capacity. This problem can be resolve by using

following equation:

Ccurrent year = Cbase year ( ) ………………….eq-02

Where,

Ccurrent year and Cbase year are the cost of equipment in current year and base year respectively and;

Q current year and Q base year are the equipment size in current scenario and base year case

And “m” is the power factor which is mostly known for some major equipment in process industry and if

its value is unknown then it can be taken as 0.6 which turn the above equation into sixth-tenth rule as

most of the equipment cost due to change in size varies around sixth-tenth to the original equipment

size

Some other factors to be considered in cost estimation of an equipment are pressure factor(FP),

temperature factor(FT), design factor(FD) and material of construction factor(FCM). All these factors are tobe considered in cost estimation of equipment based on current case and base year case.

Thus the final equation for cost estimation of equipment can be written as:

Ccurrent year = Cbase year A FPFD FCM FT (

)

……………………eq - 03

5.2 Heat Exchanger Area Targeting

Heat exchanger area account for a significant cost of

manufacturing and it should be considered in designing a heat exchanger area network. The area is

inversely proportional to ∆Tmin so by increasing ∆Tmin the area can be reduced. Thus selection of ∆Tmin

for a given process heat exchanger network design is vital in terms of capital investment.

Composite curves not only give the energy targets for a process but these can also be used to

estimate the heat transfer area of a network. For area targeting we produce composite curve with

utilities included and these composite curves are termed as “balanced composite curves”. The balance

composite curves are divided into different enthalpy intervals based on kinks in the curve. Vertical heat

transfer in each interval is calculated and the cold and hot streams temperature considered are the





terminal temperature of the interval. We also consider that the overall heat transfer coefficient remains

constant in each interval and thus with this assumption we calculate area of heat exchanger in each

interval by using following equation:

Ainterval =

Where,

Ainterval = the area of an interval

Q interval = heat transfer between hot and cold stream in one particular interval

U = overall heat transfer coefficient

LMTD =

= log mean temperature difference between cold and hot stream in one

particular interval, h1& h2 and c1&c2 are the hot and cold stream inlet and outlet terminal temperature

respectively.

The network area is calculated by the summation of each interval area as follows:

ANetwork = ∑

…………………………….eq-04

Where ANetwork = heat exchanger area of vertical heat transfer for whole network

K = total number of enthalpy intervals

Now we consider the area targeting for plant A heat transfer by using following balanced composite

curve:

The balanced composite curve for plant A is divided into 16 intervals and by assuming 4 kW/m 2- C value

of heat transfer coefficient (U), we calculated the area of heat transfer in each interval using eq-04 and

the results are tabulated in the following table:

Figure 17: Blocked Balanced Composite Curve Plant A





3500

3550

3600

3650

3700

3750

0 10 20 30

A r e a - m 2

∆Tmin- ⁰C

HEN Area vs ∆Tmin

Similarly we have estimated the area of heat exchanger network for Plant A at different values of ∆Tmin.

The results are shown in the following table:

This data in above table and accompanying graph clearly shows that the area decrease with increase in

∆Tmin. This data is produced only for Plant A due to study the relationship of heat exchanger area vs

∆Tmin. Similarly same data can be produced from plant B and C but due to time constraint and complex

utilities involvement in plant B and C these calculations are restricted to Plant A only.

5.3 HEN Area Cost Calculations

Based on equipment cost estimation method as discussed

earlier we now calculate the cost of heat exchanger network (HEN) area (as calculated in above section )

for plant A.

The escalating factor is calculated for April, 2011 as the latest data available for CEPCI was in this month

and the base year is 2001.

Interval Q interval kW LMTD Ak – m2

1 2600 4.867923 133.5272

2 4290.34 9.259437 115.837

3 16.81 1.189795 3.532121

4 337.63 8.058595 10.47422

5 555.22 6.208295 22.35799

6 340 16.10384 5.278245

7 1279.39 27.97374 11.43385

8 2885.81 53.86313 13.39418

9 1819.9 50.50524 9.008472

10 661.6 2.00597 82.45387

11 4315.2 0.334114 3228.838

12 994.8 2.85725 87.04175

13 1784.9 17.55247 25.42234

14 141.7 12.52 2.829473

AreaNetwrok 3644.289Table 10: Enthalpy, LMTD and Area Data for Plant A - Decomposed Balanced Composite Curve at ∆Tmin 10C

∆Tmin - C Area – m2

2.5 3732.592

7.5 3663.14

10 3644.289

12.5 3626.003

15 3583.583

17.5 3565.436

20 3549.73

Table 11: HEN area vs ∆Tmin Data - Plant AFigure 18: HEN Area vs Tmin Graph





A =

= = 1.69

FCM and FDthe material of construction and design factors are considered to be the same as per original

equipment.

The pressure factor (FP) is considered based on stream pressure i.e. 20 bar and temperature factor (FT) is

based on maximum temperature of the system i.e. 450C. The values for these factors and power factor“m” in eq-06 is taken from “Chemical Process Design and Integration” by Robin Smith†.

FP = 1.5, FT = 2.1 and m = 0.68

Base Equipment capacity = 1000m2 and Base equipment cost = $ 105,000

Thus cost of HEN with area 3644.289m2 (∆Tmin 10C) for Plant A is as follows,

C2011 = C2001 A FPFD FCM FT

Thus by putting all the values in above equation we get the cost of HEN

C2011 = 105000x1.69x1.5x1x1x2.1x

= 1347000 $ or 1.347 MM$similarly the cost for each network for different Tmin values are calculated and tabulated in the following

table,

The cost for heat exchanger network follows the same downward trend as of area for increase in ∆Tmin

value which is quite evident from the above table.

5.4 Energy Cost Calculations

High energy price has put a strict condition in giving priority to design

minimum energy requirement processes. For this reason heat exchanger network has become more

significant than ever in order to meet process heating and cooling demands with minimum external

energy requirements. For energy analysis of a system the cost of energy is vital and there are certain

ways in which we can calculate the utilities cost. For this report steam pricing and hot oil cost estimation

is considered. Price of some common fuels is given in the following table:

Fuel Price

Coal 4.72 $/GJ

∆Tmin - C Area – m2 Cost – MM$

2.5 3732.592 1.369

7.5 3663.14 1.352

10 3644.289 1.347

12.5 3626.003 1.343

15 3583.583 1.332

17.5 3565.436 1.327

20 3549.73 1.323

Table 12: HEN Cost Data for Different ∆Tmin Values





Heavy Oil 909.01 $/t

Electricity 0.10608 $/kWh

Natural Gas 0.0319 $/kWhTable 13: Fuel Prices

Moreover the cost of cooling water is considered to be 1% of the price of electricity. The steam price

calculations are as follows:

Steam Pressure – bar a Temperature - C

HP 40 400 – super heated

MP 20 212.37 - Saturated

LP 4 143.7 - SaturatedTable 14: Steam Conditions for Total Site

The above table shows the different steam levels we have selected for our system. HP steam is at 40 bar

and 400C superheated temperature. Its cost calculation is as follows;

Amount of heat required to produce HP steam = Q HP = hsh-hBFW - kJ/kg

From steam table, hsh = 3214 kJ/kg

Boiler feed water is available at 105C, so hBFW = 4.2 (105-25)(relative to make-up water at 25C)

= 336 kJ/kg

Q HP = 3214 – 336

= 2878 kJ/kg

Thus the fuel required to produce HP steam = Q HP x Costfuel x

Assuming boiler efficiency of 0.90 and natural gas as fuel the cost of HP steam comes out to be

CostHP =

= 0.0283 $/kg or 28.3 $/t

Cost of MP steam = CostHP – Valuepower produced

HP steam is expanded isentropic ally to 20 bar MP steam conditions, thus;

W = hHP - hisp

Whereas hisp = isentropic enthalpy of steam after turbine

And hisp = hHP – ƞ(hHP – hg)

Whereas hg = enthalpy of steam at 20 bar and

Ƞ = turbine isentropic efficiency which is assumed to be 0.87 in our case

Again using steam table hHP = 3214 kJ/kg

Hg = 2799 kJ/kg

Thus, hisp = 3214 – 0.87(3214-2799)

= 2852.95 kJ/kg

From steam table this enthalpy value at 20 bar dictate the steam temperature of 231.5C which is higher

than the saturated temperature of 212.37C, thus the steam at turbine outlet is superheated and we





calculate our steam price on these basis as some degree of superheat would be preferable to

accommodate thermal losses in steam supply.

So W = 3214 – 2852.95

= 361.05 kJ/kg

Values of power =

= 0.0106 $/kg

Cost of MP steam = 0.0283 – 0.0106

= 0.0177 $/kg or 17.7 $/ton

Similarly the cost of LP steam is calculated as,

CostLP = CostMP – Valuepower generated

MP steam is expanded isentropically to 4 bar LP steam conditions, thus;

W = hMP - hisp

Whereas hisp = isentropic enthalpy of steam after turbine

And hisp = hMP – ƞ(hMP – hg)

Whereas hg = enthalpy of steam at 4 bar and

Ƞ = turbine isentropic efficiency which is assumed to be 0.87 in our case

Again using steam table hMP = 2852.95 kJ/kg

Hg = 2739 kJ/kg

Thus, hisp = 2852.95 – 0.87(2852.95-2739)

= 2753.81 kJ/kg

From steam table this enthalpy value at 4 bar dictate the steam temperature of 150.46C which is higher

than the saturated temperature of 143.7C, thus the steam at turbine outlet is superheated and we

calculate our steam price on these basis as some degree of superheat would be preferable to

accommodate thermal losses in steam supply.

So W = 2852.95 – 2753.81

= 99 kJ/kg

Values of power =

= 0.00292 $/kg

Cost of LP steam = 0.0177 – 0.00292

= 0.0147 $/kg or 14.7 $/ton

Fuel Oil Calculations:

Cost of hot oil = Cost of pumping + cost of fuel

Cost of pumping is not considered in this system as it requires detailed engineering calculations, thus

only fuel cost is considered as follows,

CHO = Q HO x Cfuel x





1.3

1.32

1.34

1.36

1.38

0 5 10 15 20 25

C o s t - M M $

∆Tmin - ⁰C

Capital-Energy Cost Trade-Off - ∆Tmin

Total Cost Capital Cost Energy Cost

Q HO = Cp (Ttarget – Tsupply)

Ttarget = 400C and Tsupply = 360C (temperatures selected using grand composite curve of plant B)

Cp = 2.527 kJ/kg K at 360C

Thus Q HO = 2.527 (400-360)

= 101.08 kJ/kg

CHO =

= 0.0001 $/kg or 0.1$/ton

Initial start-up cost of hot oil would be much higher than this running cost as more heat would be

required for hot oil heat up from storage temperature to required process heating temperature.

6. Capital – Energy cost Trade-off and ∆Tmin

We have seen so far that the capital cost

of heat exchanger network is associated with ∆Tmin, higher the value of Tmin less would be the overall

area but another significant parameter to be considered in selection of Tmin is energy cost. Energy cost

of a process increase with increase in Tmin due to missing opportunities of process to process heat

recovery in a heat exchanger network. Thus the optimum value of Tmin is determined by a trade-off

between capital and energy cost.

Now we consider the Tmin selection for Plant A. We have already calculated the heat exchanger

network area cost at different values of Tmin and we also know that Plant A has a threshold problem with

cold utility requirement only. The cold utility for plant A is MP steam generation and cooling water. Onlythe cost of cooling water will be used to evaluate the energy cost variation with change in Tmin as the

steam is being generated from process heating without external source.

Figure 19: Capital-Energy Trade Off Curve – Plant A





0

10

20

30

0 5 10 15 20 25 30 C o o l i n g w a t e r

t e m p e r a t u r e - ⁰ C

Tmin - ⁰C

CW Temperature vs ∆Tmin

Cooling Water Temperature

From the figure above it is quite evident that there is no change in utility cost for plant A with T min

as it is a threshold problem for certain value of T min so the total cost follows the capital cost downward

trend with Tmin. Thus in this case it is no possible to select optimum value of T min for the heat exchanger

network. However another interesting fact related to change in Tmin is the change in utility temperature

requirement for process heating or cooling. Quality of heat is equally important as the quantity of heat

and it is observed that the maximum cold utility temperature decrease with increase in Tmin as shown in

the following figure:

Thus from the above figure we have selected the Tmin value of 10C for Plant A heat exchanger network

design as below this value the required cooling water temperature decreases our set target of 25C.

similarly for plant B and C we have selected the Tmin of 10C based on experience of determining this

value for plant A.

7. Heat Exchanger Network DesignNow as we have produced enough data for each plant,

we can design the heat exchanger network for plant A, B and C.

7.1 Plant Data

Following table contains all the data that we have generated so far from the given site processes.

Plant ∆Tmin - ⁰C Heat Recovery -

kW

Q Hmin & Utilities

(kW)

Q Cmin & Utilities

(kW)

A 10 5476.62 Total - 16546

MP Steam – 9673CW – 6873

B 10 45821 Hot Oil - 4363 Total – 12124

MP Steam – 8471

CW – 3652

C 10 - Flue gas -167400 HP steam –

143800Table 15: Data Summary Plant A,B and C

Figure 20: Cooling water vs ∆Tmin Curve





As we have discussed already in utility selection section of this report that plant C has been divided

into two section containing cold and hot streams separately. Thus in above table all the hot utility

requirement will be fulfilled by furnace flue gas and HP steam will be generated from hot stream.

Another important information that is required during the design on heat exchanger network is

process and utility pinch temperatures(these are shifted temperature by ∆Tmin. Following table containsthis data for each plant:

Plant Process Pinch Temperature – ⁰C Utility Pinch Temperature - ⁰C

A 455 75

B 355 35

C 30 -Table 16: Pinch Data Plant A,B and C

7.2 HEN Design Methods

There are several methods available for design of heat exchanger network but we have discussed the

following methods briefly in this section:

a. Pinch Design Method.

b. Superstructure Optimization.

c. Stochastic Optimization (Simulated Annealing)

Pinch Design Method:

This method is based on pinch analysis in which a problem is divided across the pinch and design

starts from the pinch. The rules to design heat exchanger network based on pinch analysis are defined

near the pinch however there are no certain rules for design away from pinch. Energy targets are fully

met when a network is design on this method, moreover it gives a freedom to take decision by designerand incorporate all the process insight. However this method becomes complicated when a large

number of streams and multiple utilities are involved.

Superstructure Optimization:

In this method all the possible combinations for one stream’s heat exchange to other stream are

made using superstructure and the final result is optimized to acceptable number of heat exchangers

and streams. This method is based on automated design and it does not involve designer decisions thus

the process insight cannot be incorporated into this design method. Furthermore it become really

complex when large number of streams is involved due to difficult mathematical problem.

Stochastic Optimization:

In this method an initial feasible network is provided by the designer which is then exploited for

possible matches between each hot and cold streams. Perturbation moves like repiping, resequencing,

add/remove heat exchanger or stream split are set with their probabilities. Then the simulated

annealing parameters like initial temperature, final temperature, Markov chain length are set. This

method is analogous to physical annealing). Initial and final temperatures control the bad moves and





markov chain length governs time for each match. Finally the design is optimized to get a feasible

solution. This method produce different design in every solution thus a designer can select the best

solution for a given process. This method takes relatively large time to generate the design and thus an

optimum solution is not possible in practical time, to overcome this problem non-linear programming is

used to obtain optimum results.

For this design project we have selected Pinch Design Method with maximum energy recovery in

plant A and B because the number of streams and utilities involved in each process are manageable and

it will give us a non-reducible energy efficient heat exchanger network design.

7.3 HEN Design Plant A

Plant A has five hot streams and three cold streams and it is a threshold

problem with no hot utility requirement. Based on pinch design method we first divide the process

across process and utility pinch. As a threshold problem the process pinch is at the beginning of heat

source area of the problem(shown in problem table section 4.3) thus the problem is divided across the

utility pinch only. Table 2 has stream data for plant A and Table 14 & 15 section 7.1 contains the datarequired to design maximum energy recovery heat exchanger network for plant A. Using this data we

generate the grid diagram using SPRINT which is the best representation of all the process streams to be

included for design of heat exchanger network with their temperature, enthalpy and CP data. Moreover

it also shows the process division across the pinch as follows:

However it was a threshold problem but multiple utilities have changed into a pinch problem. The grid

representation of the process streams shows some interesting facts about the location of process

streams. It is quite clear from the above figure that almost half of streams lie away the pinch and it

increase the complexity in the design process as no certain rules are defined for matching streams away

the pinch. Moreover we start the design of network from zero utility end of the process as there is no

utility to achieve the required temperature and then we move towards the rest of the process.

Following general rules are applied for making matches at pinch:

Above the Pinch: Number of hot stream, NH < Number of cold Stream NC and CPH< CPC

Figure 21: Grid Diagram Plant A





1N: 1

258 257

25 7DH :4270 CP:4270

2N: 2

280 110

11 0DH :5760 CP:33.8 824

3N: 3

45 35

35DH :2600 CP:2 60

11

FF:0.152

13

FF:0.159

14

FF:0.629

12

FF:0.061

15

FF:1

4N: 4

80 55

55DH :5200 CP: 208

5N: 5

450 80.2 4

80D H:4193. 33 CP:1 1.3333

6N: 6

5815 0

15 0DH:379.065CP:31.5888DH:2527.1CP:31.5888

7N: 7

2515 0

15 0DH:151.364CP:3.36364DH:269.091CP:3.36364

8N: 8

3522 5

22 5DH:396.053CP:11.3158DH:1753.95CP:11.3158

9N: 9

2530

30DH:6873.52CP:1374.7

10N:10

10 5212.47

212.37DH:9673.19CP:90.0044

1

1

N:11

N:12

25 7

212.45

*Q:4270A:94.7223S: 0

2

2

N:13

N:14

184.58

212.39

*Q:3233.06A:67.3166S: 0

3

3

N:15

N:16

11 0

15 0

*Q:2527A:67.8843S: 0

8

8

N:23

N:24

35

26.89

*Q:2600A:190.4S: 0

7

7

N:25

N:26

68

70

*Q:379.065A:37.9018S: 0

9

9

N:27

N:28

68

70

*Q:396.053A:20.5583S: 0

10

10

N:29

N:30

67.99

70

*Q:151.364A:6.6912S: 0

11

11

N:31

N:32

68

30

*Q:1569.52A:35.3899S: 0

M1 N:33

68

12

12

N:34

N:35

55

28.86

*Q:2704A:81.1493S: 0

4

4

N:17

N:18

258.74

212.47

*Q:2167.57A:18.5349S: 0

5

5

N:19

N:20

23 5

15 0

*Q:269.091A:1.9944S: 0

6

6

N:21

N:22

80.24

22 5

*Q:1753.95A:173.318S: 0

80

70

H1

H2

H3

H4

H5

C1

C2

C3

Cooling Wate

MP Steam

Below the Pinch: Number of hot stream, NH > Number of cold Stream NC and CPH> CPC

CP rules are of significant importance as the temperature difference should increase on moving

away from pinch and this objective is achieved by following the CP rules. In matching streams away the

pinch it has been considered to supply heat from lowest minimum temperature e.g. in this case the

matches with hot streams and steam generation are made first at higher temperatures and then the low

temperature heat is supplied to cold process streams as they required heat at relatively lower

temperature than steam. Another problem associated with using steam generation as a cold utility is the

phase change of the water from liquid to gas. This change takes place at constant temperature and thus

a higher temperature than the saturation temperature of the steam is to be maintained in a relatively

bigger area of the exchanger to maintain minimum design approach temperature. This limit further

made the design more complex. Furthermore below the pinch only hot stream 4 (inlet temperature 80C)

can exchange heat with three cold streams(outlet temperature 70) due to temperature constraints and

one of cold stream i.e. stream 6 has inlet temperature of 58C which is higher than stream 6 outlet

temperature. This problem was addressed by setting stream 6 outlet temperature at 68C after

exchanging heat will all the streams through split stream rule. And to achieve exit temperature of 68C at

the end mixing point one of split stream is being cooled with cooling water. This constraint resulted in

an extra cooler for the process.

Following above mentioned rules and considering the process limitation we have design the

following maximum energy recovery heat exchanger network for plant A at Tmin of 10C:

This network has zero cross pinch heat transfer and following table contains the heat exchanger

network report generated by SPRINT software:

Total Hot

Utility =

0 [kW] MER Heat Exchanger Network Report Plant A

Total cold

Utility =

16544.1 [kW]

Figure 22: MER Heat Exchanger Network - Plant A





Original Optimized

Each Hot Cold Minimum Hot Hot Cold Cold Area Area

No. Stm Stm Approach In Out In Out [m^2] [m^2]

[C] [C] [C] [C] [C]

F 1[H1 ]cu 1 10 44.61 258 257 212.39 212.45 94.72 94.72

2[H2 ]cu 2 10 28.79 280 184.58 105 212.39 67.32 52.53

3[<Unna] 2 6 34.58 184.58 110 70 150 67.88 59.53

4[<Unna]cu 5 10 46.3 450 258.74 212.45 212.47 18.53 10.44

5[<Unna] 5 7 108.74 258.74 235 70 150 1.994 2.126

6[<Unna] 5 8 10 235 80.24 70 225 173.3 49.69

7[<Unna] 4:11 6 10 80 68 58 70 37.9 0

F

8[<Unna]cu

3 9 10 45 35 25 26.89 190.4 190.4

9[<Unna] 4:13 8 < 10.00> 80 68 35 70 20.56 0

10[<Unna] 4:12 7 < 10.00> 80 67.99 25 70 6.691 0

11[<Unna]cu4:14 9 39.14 80 68 28.86 30 35.39 0

12[<Unna]cu

4:15 9 28.11 68 55 26.89 28.86 81.15 136.3

Original Optimized

Process Minimum Approach Temperature 10 [C] 44.94 [C]

Utility Minimum Approach Temperature 10 [C] 10 [C]

Overall Minimum Approach Temperature 10 [C] 10 [C]

Table 17: Network Report MER Heater Network - Plant A and Area Comparison with Optimized Network

As it is clear from the table above that there is no violation of Tmin in any heat exchanger of the networkbut the highlighted areas of exchanger are too small and practically infeasible thus we optimized the

network for minimum number of heat exchangers, the results are compared in same table above and

following is the optimized heat exchanger network:

Figure 23:Optimized MER Heat Exchanger Network- Plant A

1 N: 1

2 5 8 2 5 7

25 7D H : 4 2 7 0 C P : 4 2 7 0

2 N: 2

2 8 0 1 1 0

11 0D H : 5 7 6 0 C P : 3 3 . 8 8 2 4

3N: 3

4 5 3 5

35D H : 2 6 0 0 C P : 2 6 0

13

FF:0

11

FF:*****

12

FF:0

14

FF:1

15

FF:1

4 N: 4

8 0 5 5

55D H : 5 2 0 0 C P : 2 0 8

5N: 5

4 5 0 8 0 . 2 3

80D H : 4 1 9 3 . 3 3 C P : 1 1 . 3 3 3 3

6N: 6

5815 0

15 0D H : 3 7 9 . 0 6 5C P : 3 1 . 5 8 8 8D H : 2 5 2 7 . 1C P : 3 1 . 5 8 8 8

7N: 7

2515 0

15 0D H : 1 5 1 . 3 6 4C P : 3 . 3 6 3 6 4D H : 2 6 9 . 0 9 1C P : 3 . 3 6 3 6 4

8N: 8

352 2 5 . 0 1

22 5D H : 3 9 6 . 0 5 3C P : 1 1 . 3 1 5 8D H : 1 7 5 3 . 9 5C P : 1 1 . 3 1 5 8

9N: 9

2530

30D H : 6 8 7 3 . 5 2C P : 1 3 7 4 . 7

10N : 1 0

10 52 1 2 . 4 7

2 1 2 . 3 7D H : 9 6 7 3 . 1 9C P : 9 0 . 0 0 4 4

1

1

N : 1 1

N : 1 2

25 7

2 1 2 . 4 5

* Q : 4 2 7 0A : 9 4 . 7 2 6 3S: 0

2

2

N : 1 3

N : 1 4

1 9 5 . 7 7

2 1 2 . 3 9

* Q : 2 8 5 3 . 9 9A : 5 2 . 5 3 2 4S: 0

3

3

N : 1 5

N : 1 6

11 0

15 0

* Q : 2 9 0 6 . 0 7A : 5 9 . 5 2 8 1S: 0

8

8

N : 2 3

N : 2 4

35

2 6 . 6 7

* Q : 2 6 0 0A : 1 8 9 . 1 1 4S: 0

7

7

N : 2 5

N : 2 6

80

58

* Q : 0A: 0S: 0

9

9

N : 2 7

N : 2 8

80

35

* Q : 0A: 0S: 0

10

10

N : 2 9

N : 3 0

80

25

* Q : 0A: 0S: 0

11

11

N : 3 1

N : 3 2

80

30

* Q : 0A: 0S: 0

M1 N : 3 3

80

12

12

N : 3 4

N : 3 5

55

30

* Q : 5 2 0 0A : 1 3 6 . 3 1 6S: 0

4

4

N : 1 7

N : 1 8

3 0 7 . 0 5

2 1 2 . 4 7

* Q : 1 6 2 0 . 1 5A : 1 0 . 4 3 6 1S: 0

5

5

N : 1 9

N : 2 0

2 6 9 . 9 5

15 0

* Q : 4 2 0 . 4 5 5A : 2 . 1 2 6 1 9S: 0

6

6

N : 2 1

N : 2 2

8 0 . 2 3

2 2 5 . 0 1

* Q : 2 1 5 0 . 1A : 4 7 . 6 9 0 4S: 0

80

70

H1

H2

H3

H4

H5

C1

C2

C3

Cooling Wate

MP Steam





1N:1

25380.1

380DH:427.465CP:85.493DH:27357.7CP:85.493DH:2564.79CP:85.493

2N:2

210280

280DH:6475CP:92.5

3N:3

350351

351DH:8400CP:8400

4N:4

278279

279DH:4960CP:4960

5 N:5

400 225

225DH:6601.14 CP:165.029DH:22278.9 CP:165.029

6 N:6

3 50 200

200D H: 97 00 C P: 64 .6 66 7

7 N:7

1 84 183

183DH:7430 CP:7430

8 N:8

40 39

39DH:4080 CP:4080

9 N:9

2 00 199

199DH:1760 CP:1760

10 N:10

2 78 67.1

67.1D H: 60 95 .7 5 C P: 28 .9 03 5

11N:11

2530

30DH:3652.54CP:730.507

12 N:12

400 360

360DH:4363.65 CP:109.091

13N:13

105212.47

212.37DH:8471.86CP:78.8266

360

350

40

30

C1

C2

C3

C4

H1

H2

H3

H4

H5

H6

CW

Hot Oil

MP

This network has eliminated all the process-to-process heat exchanger(dashed lines) below the

pinch and this load is shifted to process-to-process heat exchangers above the pinch but in doing so

there is 926kW of cross pinch exchanger heat flow which has resulted in load shift from MP steam utility

to cooling water thus total cooling requirement remained the same as shown in following cross pinch

heat transfer report of above network:

Minimum Hot Utility = 0 [kW]

Minimum Cold Utility = 16546.7 [kW]

Process heat recovery= 5476.62 [kW]

Total Hot utility = 0.00000 [kW]

Total Cold utility = 16544.1 [kW]

Total Cross Pinch heat transfer = 0.00000 [kW]

Exchanger cross pinch heat transfer

3 3 [<Unnamed> ] 379.065 [kW] Pinch Number 1

5 5 [<Unnamed> ] 151.364 [kW] Pinch Number 1

6 6 [<Unnamed> ] 396.053 [kW] Pinch Number 1Total exchanger cross pinch heat flow = 926.482 [kW]

Table 18: Cross Pinch Report - MER Heat Exchanger Network Plant A

As the cost of cooling water is not high and there is slight decrease in steam production load thus the

optimized network is feasible both in terms of exchanger cost and utility load. Additional benefit of

reduced area for some heat exchanger is obtained in optimized network due to high driving force as

increased process to process minimum approach temperature shown in table 16.

7.3 HEN Design Plant B

Following the same pinch design method as we used in plant A we designed

maximum energy recovery heat exchanger area for plant B. The grid diagram for plant B is as follows:

Figure 24: Grid Diagram Of Plant B





1N:1

25380.1

380DH:427.465CP:85.493DH:27357.7CP:85.493DH:2564.79CP:85.493

2N:2

210280

280DH:6475CP:92.5

3N:3

350351

351DH:8400CP:8400

17

FF:0.214

18

FF:0.786

19

FF:1

4N:4

278279

279DH:4960CP:4960

5 N:5

400 225

225DH:6601.14 CP:165.029DH:22278.9 CP:165.029

6 N:6

350 200

200DH:9700 CP:64.6667

7 N:7

184 183

183DH:7430 CP:7430

20

FF:0.105

21

FF:0.895

22

FF:1

8 N:8

40 39

39DH:4080 CP:4080

9 N:9

200 199

199DH:1760 CP:1760

10 N:10

278 67.1

67.1DH:6095.75 CP:28.9035

11N:11

2530

30DH:3652.54CP:730.507

15

FF:0.588

14

FF:0.412

16

FF:1

12 N:12

400 360

360DH:4363.65 CP:109.091

13N:13

105212.47

212.37DH:8471.86CP:78.8266

M2N:45

351

1

1

N:14

N:15

360

351

*Q:6601.14A:268.994S:0

6

6

N:24

N:25

318.05

350.08

*Q:6923.31A:283.138S:0

5

5

N:22

N:23

287.99

279

*Q:4960A:232.669S:0

7

7

N:26

N:27

247.02

269.1

*Q:6761.29A:195.901S:0

10

10

N:40

N:41

225

212.45

*Q:3633.97A:166.916S:0

4

4

N:20

N:21

249.87

280

*Q:6475A:120.961S:0

11

11

N:32

N:33

200

212.4

*Q:3225A:105.816S:0

8

8

N:28

N:29

183

169.43

*Q:7430A:166.983S:0

13

13

N:36

N:37

39

30

*Q:427.465A:35.9572S:0

15

15

N:38

N:39

39

30

*Q:3652.54A:307.245S:0

M3 N:46

39

9

9

N:30

N:31

199

190.01

*Q:1760A:97.543S:0

12

12

N:34

N:35

222.45

212.47

*Q:1605.72A:54.3733S:0

16

16

N:42

N:43

67.1

82.52

*Q:4489.9A:57.9618S:0

2

2

N:16

N:17

360

351

*Q:1798.46A:73.2969S:0

3

3

N:18

N:19

360

380.1

*Q:2566.3A:178.968S:0

M1 N:44

360

360

350

40

30

C1

C2

C3

C4

H1

H2

H3

H4

H5

H6

CW

Hot Oil

MP

The complexity of design in this case is increased due to two pinches a process pinch at 355C and

utility pinch at 35C. Moreover between these two pinch temperatures we can see from the grid diagram

that the process streams are quite far from pinch thus causing more design difficulty due to no certain

rules. In same manner as in plant A we will transfer heat at lowest possible temperature when matching

hot and cold streams. Steam generation here poses the same constant temperature phase change

problem in network. Another interesting feature of plant B is the availability of sufficient heat but not at

high temperatures which add more complexity to the process. Thus by considering the same phase rules

as in plant A with plant B process constraints the HEN design is as follows:

Since the hot oil returned temperature is selected by process pinch temperature of 360C so the hot oil

utility stream is spit to maintain Tmin constraint for heating cold streams whose inlet temperatures are at

360C. Hot streams match with MP steam generator heat exchanger is designed in a way that the outlet

temperature of hot stream remained above 222.27C i.e. 10C plus steam saturation temperature to

overcome constant temperature phase change problem. Streams at lower temperatures with high heat

content are used to heat cold streams to a certain temperature without violating Tmin constraint and

remaining heat to the cold streams is provided by hot streams at higher temperature. This approach

made the design more complex as one exchanger on each stream could not fulfill its heating or cooling

requirement.

The Network report for this heat exchanger network generated using sprint as follows:

Exch Hot Cold Minimum Hot Hot Cold Cold Area DutyNo. Stm Stm Approach In Out In Out [m^2] [kW]

[C] [C] [C] [C] [C]

1[H1 ] 5 3:18 10 400 360 350 351 269 6601

2[Hu1 ]hu 12:14 3:17 < 10.00> 400 360 350 351 73.3 1798

3[Hu2 ]hu 12:15 1 < 9.92> 400 360 350.08 380.1 179 2566

F 4[H2 ] 6 2 39.87 350 249.87 210 280 121 6475

Figure 25: MER Heat Exchanger Design Plant B





F 5[H3 ] 5 4 < 9.99> 318.05 287.99 278 279 232.7 4960

6[H4 ] 5 1 < 9.92> 360 318.05 269.1 350.08 283.1 6923

7[H5 ] 5 1 18.89 287.99 247.02 190.01 269.1 195.9 6761

F 8[H6 ] 7 1 14.57 184 183 82.52 169.43 167 7430

F 9[H7 ] 9 1 < 9.99> 200 199 169.43 190.01 97.54 1760

10[<Unna]cu 5 13 12.6 247.02 225 212.4 212.45 166.9 3634

F 11[G1 ]cu 6 13 13.58 249.87 200 105 212.4 105.8 3225

12[G2 ]cu 10 13 < 9.99> 278 222.45 212.45 212.47 54.37 1606

13[H10 ] 8:20 1 < 10.00> 40 39 25 30 35.96 427.5

15[Cu1 ]cu 8:21 11 10 40 39 25 30 307.2 3653

16[<Unna] 10 1 37.1 222.45 67.1 30 82.52 57.96 4490

Process Minimum Approach Temperature 9.92 [C]

Utility Minimum Approach Temperature 9.92 [C]

Overall Minimum Approach Temperature 9.92 [C]Table 19: Network Report MER Heat Exchanger Network Plant B

The minimum approach temperature of this network is found to be 9.92C and this values is a result of

two heat exchangers with 9.92C Tmin i.e. exchanger 3 and exchanger 6 and this could be result of some

calculation error as the heat load and temperature are considered to 3 decimal places and the

complexity of design involved large number of fractional values for the heat load and temperatures.

Similarly the heat transfer across the pinch is negligible and might be the result of same error because

only 6.7kW of cross pinch heat transfer is carried out from same heat exchanger 6 as shown below:

*DTmin = 10 [C]

Minimum Hot Utility = 4363.65 [kW]

Minimum Cold Utility = 12124.4 [kW]

Process heat recovery= 5476.62 [kW]

Total Hot utility = 4364.76 [kW]

Total Cold utility = 12117.2 [kW]

Total Cross Pinch heat transfer = 1.11646 [kW]

Exchanger cross pinch heat transfer

6 6 [H4 ] Pinch Number 1 -6.74845 [kW]

Total exchanger cross pinch heat flow = -6.74845 [kW]

There is no cross pinch mixingTable 20: Cross Pinch Report - Plant B

7.4 HEN Design Plant C

As we have discussed earlier in utility selection part of this report that

plant C hot stream will used to produce HP steam and its cold streams will be heated by furnace flue

gas. Thus the design of heat exchanger network for plant C is straight forward and does not involve any

complexity as follows:





All the cold stream heating is carried out by furnace flue gas and all the heat from hot stream is used to

produce steam.

7.5 Sensitivity Analysis

Impact of a change of process input on process out can be estimated by

defining equipment depended and equipment independent relationships. Equipment dependent

relations are usually called as design equation where equipment independent relationships are material

balance, energy balance etc. or those relationships which do not require equipment specification.

For heat exchangers equipment dependent relationships are area equations i.e. Q = U A ∆TLMTD.Thus change in heat load due to change in temperature or process demand in any plant can affect the

performance of heat exchanger area network.

8. ConclusionA number of tools and techniques used to process data in a way to achieve required

target of designing heat exchanger network for three plants in a total site. The design stages involved

complex data handling which seemed to be relatively simple and straight forward in the beginning.

Economic data gathering and analysis was the most challenging part of the design project followed by

maximum energy recovery heat exchanger design for each plant. However the results obtained fromheat exchanger network meet the design parameters and energy targets for the process.

9. References:1] Richard Turton/C.Bailie/Wallace B.Whiting/Joseph A Shaeiwitz “Analysis, synthesis, and Design of

Chemical Processes, Third Edition. P: 177-182, 230, 583-584, 923-941.

[2]Fuel Price Data –Departmen of Energy and Climate Change , Quarterly Energy Prices, December 2011-

Section3 –Industrial Prices Table 3.1.1

<http://www.decc.gov.uk/publications/basket.aspx?filepath=statistics%2fsource%2fprices%2fqep312.xl

s&filetype=4#basket

[3]Chemical Engineering (April 2011) “ Economic Indicators: Chemical engineering Plant Cost Index.”

Viewed 11/02/2012

< http://www.che.com/PCI>

[4]ROBIN SMITH. Chemical Process Design and Integration. P:17-20

Figure 26: Heat Exchanger Network Plant C

http://www.decc.gov.uk/publications/basket.aspx?filepath=statistics%2fsource%2fprices%2fqep312.xls&filetype=4#basket








[5]SIMON PERRY. Utility System Handouts 2011-2012. P:L08-6

[6]SIMON PERRY. Energy System Handouts 2011-2012. P:L03-L1

Report Design Project

Documents

Transcript of Report Design Project