Energy efficient styrene process: design and plantwide control

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Process Systems Engineering

Energy efficient styrene process: design and plantwide control Alexandre C. Dimian, and Costin Sorin Bildea Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.8b05560 • Publication Date (Web): 06 Mar 2019 Downloaded from http://pubs.acs.org on March 17, 2019

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Industrial & Engineering Chemistry Research

Figure 1 Catalyst activity and selectivity as function of temperature (Reprinted with permission from Lee & Froment, Ind. Eng. Chem. Res. 2008, 47, 9183-9194. Copyright 2018, American Chemical Society)

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Industrial & Engineering Chemistry Research 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46

Figure 2 - The influence of ratio steam/ethylbenzene on conversion and selectivity (Lee & Froment, 2009, with permission)


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Figure 3 Reactor-separator-recycle structure of styrene manufacturing process (base case)



Figure 4 Concentration and temperature profiles in the first dehydrogenation reactor


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Figure 5 Composite curves and grand composite curve in the base-case



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Figure 6 RSR structure by styrene manufacturing, MVR alternative EB-REC

2.0 bar 630 °C

H-1 EB-VAP

MIX-1

R1-IN

EB-FEED

R-1

MIX-2

BFW2 Q=8.2 MW

FURN-1B 0.5 bar 82.3 °C

2.5 bar 150 °C

STEAM

FURN-2

FURN-3 Q=2.93 MW

SEP

Q=4.06 MW

BFW1

W-EVAP COMP

LIGHTS

R-2

STYRENE

ORGANICS

FURN-1A

Q=22.15 MW W=3.67 MW Q=10.1 MW

1.2 bar 33 °C

R2-OUT

COOLER Q=47.48 MW


WATER

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Figure 7 Composite curves and grand composite curve in the MVR alternative



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Figure 8 Flowsheet for the separation section by styrene manufacturing BENZENE

C-4

H2

M-1

65°C 1.1 bar 115°C 1.15 bar

CH4+

TOLUENE

H-1

7 °C

Q=-770 kW

S-2

ETHYL BENZENE RECYCLE

P-1

102 °C 1.65 bar

CP-1

W=-70 kW 33°C 1.2 bar R-OUT

WATER

STYRENE

GASES

C-1

S-1

30.8°C 0.25 bar

C-2

76.4°C 0.15 bar

C-3

75.2°C 0.10 bar

ORGANIC 117.4 °C 0.50 bar

113.2°C 0.40 bar

114.2°C 0.22 bar HEAVIES

Qc=-470 kW Qr=1281 kW

Q=-9223 kW Qr=9071 kW


Q=-4196 kW Qr=3915 kW

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Figure 9 Energy saving in the base case



Figure 10 Flowsheet for energy saving with MVR and five FEHE units


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Figure 11 Plantwide control structure of the styrene plant

Ethylbenzene recycle LC

Ethylbenzene

FC

FEHE-2 FT

FY

FEHE-1

FC

BFW-2

REACTOR-1

REACTOR-2

FEHE-4 TC

FURN-1B FT

BFW-1

FC

FURN-2

TC

TC

PC

FURN-3

TC

FEHE-3

LC

FURN-1A

FEHE-5

PC

PC

TC

Hydrogen

Lights (V) LC

Lights (L)

TC

FC

TC LC

TC

COL-2

COL-1 COOLER

COMP-1

PC LC

SEP VLL

LC

LC

COMP-2

SEP VLL

Water

XC TC

TC

LC

Water

LC

XT

Styrene



(a)

(b)

240

1

1

Figure 12 - Dynamic

0.975

simulation results.

0.95

Production is changed by

0.925

increasing the reactor-inlet

220 0.9995

180

x / [-]

160

Styrene

0.999

Hydrogen

Hydrogen

140

0.9985 Styrene

120 100

0.998 0

5

10

15

20

0.9 0

5

t / [h]

(c) 160

100

1500

120

75

1000

80

Ethylbenzene

BFW1

BFW2 500

40

0 5

10

20

15

(d)

1

0.995 F 0.99

25

0.985

0

20

0.98 0

t / [h]

5

10

15

20

t / [h]

(e)

(f)

640

25

710

FURN-1A 20

635

700

FURN-2

FURN-3

630

690

Q / [GJ/h]

FURN-1A

15 FURN-3 10 FURN-1B

FURN-2

5

FURN-1B 625

680 0.5

1

1.5

t / [h]

2

2.5

0 0.5

1

1.5

2

t / [h]


ethylbenzene flow rate by 10%

50

0 0

15

x

F / [kmol/h]

F / [kmol/h]

2000

10

t / [h]

x / [-]

F / [kmol/h]

Ethylbenzene to reactor 200

T / [°C]

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46

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2.5

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(a)

(b)

200

1

1

Styrene

0.9995

Ethylbenzene to reactor 160

x / [-]

140

0.999

simulation results.

0.95

Production is changed by

0.925

decreasing the reactor-inlet

0.9

ethylbenzene flow rate by

Hydrogen Styrene

100

0

0.998 5

10

15

20

0

5

t / [h]

10

15

10%.

(d)

2000

160

100

120

75

1000

80

500

F / [kmol/h]

Ethylbenzene

1

x

BFW1 1500

20

t / [h]

(c)

F / [kmol/h]

0.975

Hydrogen 0.9985

120

Figure 13 - Dynamic

0.995 F

50

0.99

40

25

0.985

0

0

x / [-]

F / [kmol/h]

180

BFW2 0 0

5

10

15

20

0.98 0

5

10

t / [h]

15

20

t / [h]

(e)

(f)

640

720

25 FURN-1A

FURN-1B 635

710 FURN-1A

630

700 FURN-3 FURN-2 690 1

15 FURN-3 10 FURN-1B 5

625 0.5

Q / [GJ/h]

20

T / [°C]

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46


1.5

t / [h]

2

2.5

FURN-2

0 0.5

1

1.5

2

t / [h]


2.5


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Graphical TOC

Energy saving by MVR Ethylbenzene Separation

Reaction

H2 /Bz /Tol

Qin E-Bz MVR W

Styrene

Qout FEHE

Steam

Water FEHE


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Pinch Point Analysis Temperature (°C)

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Enthalpy (MW)




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1

1

Energy efficient styrene process: design and plantwide control

2

By Alexandre C. Dimian, Costin Sorin Bildea

3 4 5 6 7 8

Department of Chemical and Biochemical Engineering, University POLITEHNICA of Bucharest, 1, Gh. Polizu Street, RO-011061, Bucharest, Romania a corresponding author [email protected]

9 10 11 12

Abstract

13 14

The paper deals with conceptual design and simulation of an energy efficient process for manufacturing styrene

15

by ethyl benzene dehydrogenation in adiabatic reactors using superheated steam as inert. High performance

16

catalyst is employed with selectivity of 95% at conversion of 70%. An innovative solution leads to

17

a spectacular energy saving. The idea is running the steam generation under vacuum, followed by

18

mechanical vapour recompression (MVR). The resulting pressure allows matching the temperature-

19

enthalpy profiles of cold and hot streams in an evaporation/condensation zone that concentrates 40

20

% from the hot energy input. In addition, effective heat transfer is ensured by high transfer

21

coefficients. An efficient network of five feed-effluent-heat exchanger (FEHE) units diminishes the

22

utility consumption by 73% with respect to the base-case. The economic analysis demonstrates that

23

the cost of energy dominates over the cost of equipment. Despite the investment in compressor, the

24

MVR alternative brings a reduction in the Total Annual Cost by 36%. The feasibility of the proposed

25

solution is validated by dynamic simulation. The plantwide control adopts the strategy of keeping

26

constant the flow rate of ethylbenzene to the reaction section proportional to the required production

27

rate. The process can handle disturbances of +/- 10% in throughput while keeping the purity of styrene

28

over 99.5%.

29 30 31


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2 1

1

Introduction

2

Styrene is a major component of many polymer-based products so that a steady increasing in

3

worldwide production is expected in the next years. The manufacturing technology is dominated by

4

the dehydrogenation of ethylbenzene to styrene in adiabatic reactors in the presence of a large amount

5

of steam1. Several older studies have been devoted to the optimisation of the steam/hydrocarbon ratio,

6

considering the constraints set by the chemical equilibrium and catalyst activity and selectivity2-5.

7

However, the energy saving issue has received little attention, probably because at that time the cost

8

of energy was low.

9

In the last years the integration of process design and control became a central research topic in

10

Process System Engineering, stimulated by the contributions of Luyben and co-workers6. The styrene

11

process was the object of two important studies in this field. Vaseduvan7 developed a process design

12

and a simulation model in view of evaluating the performance of different plantwide control

13

structures. Then, Luyben8 reworked the same problem focusing on optimal process design and

14

dynamics. He demonstrated that higher economic advantage can be achieved by optimizing the

15

reaction yield that reduces the raw material costs. This effect was obtained working at lower

16

temperature and lower conversion, since the profit due to better raw materials use exceeds the higher

17

cost of recycles and separations. The above papers focused largely on the process dynamics but

18

disregarded the management of energy around the chemical reaction system, which in our view is a

19

key aspect. Besides, both studies made use of kinetic data for a catalyst that may be seen outdated.

20

The progress in catalyst performance recorded in the last decade allows tackling the styrene process

21

from a different perspective, namely of the involved energy.

22

Saving energy is a central issue of today’s chemical process design. Energy price is high and subject

23

to large market fluctuations. Styrene manufacturing may be counted among large energy consumers,

24

because of large endothermic reaction heat and difficult separations. The goal of this paper is to fill

25

the above gap, by presenting a method that can save significantly more energy than the usually

26

reported technology.1 The analysis considers a high performance catalyst9-10 that may reach a

27

conversion around 70% and a selectivity of 95%, as reported in modern processes11.

28

The paper deals with the conceptual design and simulation of an energy efficient plant with a capacity

29

of 100000 tonnes per year for producing styrene from ethyl benzene with purity of minimum 99.7

30

%wt. Employing vacuum steam generation followed by Mechanical Vapour Recompression (MVR)

31

leads to a substantial energy and equipment cost saving, compared with a base-case heat-recovery

32

solution. The feasibility of the proposed solution is validated by dynamic simulation by implementing

33

a robust plantwide control strategy capable of dealing with large disturbances in the production rate.


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3 1 2

2

Technology issues

3 4

2.1 Chemical equilibrium

5

The synthesis of styrene from ethylbenzene by dehydrogenation is an equilibrium-controlled reaction

6

favoured by low partial pressure of ethylbenzene and high reaction temperature. The preferred

7

industrial method for reducing the reactant partial pressure is using superheated low-pressure steam.

8

In addition, the steam protects the catalyst from coking and is applied to in-situ regeneration of

9

activity. The equilibrium conversion increases with the steam/hydrocarbon ratio, but this implies

10

higher energy costs. Similarly, higher temperature increases the cost of energy but leads to higher

11

conversion which in turn reduces the cost of separation. Thus, the operation parameters of reactors

12

are subject to optimization of an economic function, in which the pressure, temperature and

13

steam/ethylbenzene ratio are key variables.

14 15

2.2 Reactor technology

16

Adiabatic reactors are the most employed in industry. The following constraints should be considered

17

in design:

18

a. The molar ratio steam/ethyl benzene may take values from 5 to 18, or weight ratio from 0.85 to

19

31. The optimum may be determined from a cost function including the equipment and energy,

20

which in turn depend on catalyst performance in terms of temperature, conversion, and selectivity.

21

Sheppard et al.3 pointed out that, for optimisation, the selectivity is more important than activity.

22

Some studies indicate an optimum steam to ethylbenzene molar ratio around 12 (mass ratio of 2)5.

23

The disadvantage of the adiabatic mode is the rapid fall of the reaction rate with the temperature

24

so that several reaction stages with intermediate re-heating are necessary. Direct steam injection

25

at a certain height of the bed may be employed2. The inlet reactor temperature should be as high

26

as possible, in general above 600 °C, but below the maximum temperature that the catalyst can

27

tolerate, in this case 650 °C.

28 b.

Maximum achievable conversion is limited by chemical equilibrium. To cope with this, the

29

strategy is using several reactors in series, with intermediate re-heating. For economic reasons, the

30

number of reactors in industry is usually two and maximum three.

31 c.

Higher selectivity implies working at lower conversion, but this leads to higher costs of recycles.

32

However, since typically the material costs dominate in economics, achieving good selectivity

33

should be a target in process design, as demonstrated by Luyben8. Fulfilling this goal depends on



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4 1

catalyst. The performance catalyst considered in this study allows styrene selectivity better than

2

90% at conversion around 70%.

3 d.

The equilibrium conversion is favoured by low partial pressure of ethyl benzene, which can be

4

obtained by low pressure at the reactor inlet, and as discussed by higher steam ratio. In addition,

5

the reactor design should ensure small pressure drop across the catalyst bed. For this reason, radial

6

reactors have been developed for high throughput. More recent technologies make use of sub-

7

atmospheric pressure in reactor that lead to a significant reduction of steam ratio11.

8

From the above description it comes out that the design of an adiabatic reactor for styrene synthesis

9

is an optimisation problem that involves the inlet ratio steam/ethylbenzene, initial and intermediate

10

inlet temperatures, pressure and pressure drops across the bed.

11

However, the essential element in design is the catalyst. In the past, several optimisation studies have

12

been done but considering a catalyst with relatively poor selectivity that limited the optimal

13

conversion to 50 to 60%. In the present study the catalyst is equivalent with a performant industrial

14

catalyst12 reported to work at 70% conversion with selectivity around 95%. Note that the cost of the

15

catalyst, iron based, and of the reactor, simple adiabatic, is not relevant from the plant optimisation

16

viewpoint. As we demonstrate here the cost of energy used in the reaction section is a key component

17

in the operation expenses, which may be drastically reduced.

18 19

2.3 Catalyst and kinetics

20

The classical catalyst for ethyl benzene dehydrogenation is based on iron oxide. Operating conditions

21

are temperatures from 540 to 650 °C and pressure from sub-atmospheric to 3 bar. The iron catalyst

22

is known for an exceptional long life, from 2 to 3 years. Periodical regeneration takes place by steam

23

injection. The research for improving the performance continues today.

24

Testing alternative catalysts by simulation reveals significant differences in yield and productivity,

25

as shown Table 1. A modern industrial catalyst investigated by Lee and Froment9-10 is compared with

26

the catalysts employed by Vasudevan7 and Luyben8 in their studies. The reaction section is a series

27

of two-adiabatic units with intermediate re-heating. It may be observed that the first catalyst shows

28

much higher productivity close to 0.3 kg-styrene/h per kg-catalyst, as well as higher selectivity, over

29

95%. Another notable difference is the use of a lower steam ratio. This catalyst, with some physical

30

properties9-10 presented in Table 2, is selected for design.

31 32

Table 1 Performance of alternative catalysts for styrene dehydrogenation

33 Unit

Lee & Froment

Vaseduvan et al.


Luyben

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5 EB-in EB-out Catalyst weight Steam / EB Temperature# Pressure# Styrene production Conversion Styrene selectivity Productivity

kmol/h kmol/h kg °C

707.00 241.94 154970 11 613 / 625

225.51 83.85 71400 15 650/650

275.72 149.13 163400 15 560/560

bar kmol/h kg/kg-cat.h

1.25 / 1.06 442.28 0.6578 0.951 0.297

2.70/2.1 118.41 0.628 0.836 0.172

2.7/2.4 116.62 0.459 0.921 0.074

1 2 3

Table 2 Physical properties of the catalyst for ethylbenzene dehydrogenation12 Bulk density

kg-cat./m3

1 422

Pellet density

kg-cat./m

2 500

Void fraction of the bed

m3/m3

0.4312

Internal void fraction

m3/m3

0.4

Tortuosity factor Equivalent pellet diameter

3 mm

5.5

4 5

Figure 1 displays the influence of temperature on catalyst activity and selectivity10. On the range 600

6

to 640 °C the conversion rises from 60 to 80 %, close to equilibrium. The selectivity keeps remarkably

7

high over a large range, about 0.98 at 40% conversion and 0.94 at 70%, but starts to decline to 0.9 at

8

conversion above 80%. Note that for the second catalyst7,8 the selectivity is far below these values,

9

about 0.95 at 46% conversion, and 0.84 at 63% conversion.

10

11



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6 1 2

Figure 1 Catalyst activity and selectivity as function of temperature (Reprinted with permission from

3

Lee & Froment, Ind. Eng. Chem. Res. 2008, 47, 9183-9194. Copyright 2008, American Chemical

4

Society)

5 6

Figure 2 shows the influence of the steam ratio (7 and 11) on conversion and selectivity10.

7

Surprisingly, the activity is not affected, but the selectivity is somewhat better at higher steam ratio.

8

In the technological literature the steam to ethylbenzen ratio 7 is seen as a minimum, while the

9

maximum is about 15. Yee et al.5 points out that the optimum should be in the range 11 to 13, with a

10

rather flat shape depending on the cost of energy.

11

12 13

Figure 2 The influence of ratio steam/ethylbenzene on conversion and selectivity (Reprinted with

14

permission from Lee & Froment, Ind. Eng. Chem. Res. 2008, 47, 9183-9194. Copyright 2008,

15

American Chemical Society).

16 17

The chemical reactions included in the kinetic model used in the present study are:

18

𝐶𝐶6 𝐻𝐻5 𝐶𝐶𝐻𝐻2 𝐶𝐶𝐻𝐻3 ↔ 𝐶𝐶6 𝐻𝐻5 𝐶𝐶𝐶𝐶𝐶𝐶𝐻𝐻2 + 𝐻𝐻2

(R1)

𝐶𝐶6 𝐻𝐻5 𝐶𝐶𝐻𝐻2 𝐶𝐶𝐻𝐻3 + 𝐻𝐻2 → 𝐶𝐶6 𝐻𝐻5 𝐶𝐶𝐻𝐻3 + 𝐶𝐶𝐻𝐻4

(R3)

19 20 21 22 23 24 25

𝐶𝐶6 𝐻𝐻5 𝐶𝐶𝐻𝐻2 𝐶𝐶𝐻𝐻3 → 𝐶𝐶6 𝐻𝐻6 + 𝐶𝐶2 𝐻𝐻4

(R2)

𝐶𝐶6 𝐻𝐻5 𝐶𝐶𝐶𝐶𝐶𝐶𝐻𝐻2 + 2𝐻𝐻2 �⎯� 𝐶𝐶6 𝐻𝐻5 𝐶𝐶𝐻𝐻3 + 𝐶𝐶𝐻𝐻4

(R4)

The reaction rates are described by a Langmuir-Hinshelwood-Hougen-Watson (LHHW) model: 𝑟𝑟1 =

𝑟𝑟2 =

𝑃𝑃𝑆𝑆𝑆𝑆 𝑃𝑃𝐻𝐻 2 �� 𝐾𝐾𝑒𝑒𝑒𝑒

𝑘𝑘1 𝐾𝐾𝐸𝐸𝐸𝐸 �𝑃𝑃𝐸𝐸𝐸𝐸 −�

2

(1)

2

(2)

�1+𝐾𝐾𝐸𝐸𝐸𝐸 𝑃𝑃𝐸𝐸𝐸𝐸 +𝐾𝐾𝐻𝐻2 𝑃𝑃𝐻𝐻2 +𝐾𝐾𝑆𝑆𝑆𝑆𝑃𝑃𝑆𝑆𝑆𝑆� 𝑘𝑘2 𝐾𝐾𝐸𝐸𝐸𝐸 𝑃𝑃𝐸𝐸𝐸𝐸

�1+𝐾𝐾𝐸𝐸𝐸𝐸 𝑃𝑃𝐸𝐸𝐸𝐸 +𝐾𝐾𝐻𝐻2 𝑃𝑃𝐻𝐻2 +𝐾𝐾𝑆𝑆𝑆𝑆 𝑃𝑃𝑆𝑆𝑆𝑆�


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7

1 2

𝑟𝑟3 =

𝑘𝑘3 𝐾𝐾𝐸𝐸𝐸𝐸 𝑃𝑃𝐸𝐸𝐸𝐸 𝐾𝐾𝐻𝐻2 𝑃𝑃𝐻𝐻2

𝑟𝑟4 =

𝑘𝑘3 𝐾𝐾𝑆𝑆𝑆𝑆𝑃𝑃𝑆𝑆𝑆𝑆 𝐾𝐾𝐻𝐻2 𝑃𝑃𝐻𝐻2

(3)

2

�1+𝐾𝐾𝐸𝐸𝐸𝐸 𝑃𝑃𝐸𝐸𝐸𝐸 +𝐾𝐾𝐻𝐻2 𝑃𝑃𝐻𝐻2 +𝐾𝐾𝑆𝑆𝑆𝑆 𝑃𝑃𝑆𝑆𝑆𝑆�

(4)

2

�1+𝐾𝐾𝐸𝐸𝐸𝐸 𝑃𝑃𝐸𝐸𝐸𝐸 +𝐾𝐾𝐻𝐻2 𝑃𝑃𝐻𝐻2 +𝐾𝐾𝑆𝑆𝑆𝑆 𝑃𝑃𝑆𝑆𝑆𝑆 �

3

The rate constants for reaction and adsorption steps are given by Arrhenius-type equations:

4 5

𝑘𝑘𝑖𝑖 = 𝐴𝐴𝑖𝑖 ⋅ 𝑒𝑒𝑒𝑒𝑒𝑒 �− 𝑅𝑅𝑅𝑅𝑖𝑖 �

6

𝐾𝐾𝑗𝑗 = 𝐴𝐴𝑗𝑗 ⋅ 𝑒𝑒𝑒𝑒𝑒𝑒 �−

The parameters of the kinetic model are displayed in

7

Table 3 and Table 4. Note that the contributions of thermal reaction rates are neglected.

𝐸𝐸

Δ𝐻𝐻𝑎𝑎,𝑗𝑗 𝑅𝑅𝑅𝑅

(5) �

(6)

8 9

Table 3 The parameters of the kinetic model Parameter

Symbol A1 A2 A3 A4 AEB AST AH2 E1 E2 E3 E4 ΔHa,EB ΔHa,ST ΔHa,H2

Pre-exponential factor of reaction rate constants ki, [kmol/(kgcat·h)] Pre-exponential factor of adsorption constants Kj, [bar-1]

Activation energy [kJ/mol]

Adsorbtion enthalpy [kJ/mol]

10 11

The equilibrium constant is given by:

12

ln (Keq / [bar]) = 15.591 -14931 / T

Value 4.594 · 109 1.060 · 1015 1.246 · 1026 8.024 · 1010 1.014 · 10-5 2.678 · 10-5 4.519 · 10-7 175.38 296.29 474.76 213.78 -102.22 -104.56 -117.95

(7)

13 14

Table 4 The implementation of kinetic model in Aspen Plus Reaction

Rate constant

Pre-exp. [kmol/(kg·cat·h·bar)]

R1 R2 R3 R4

15

Activation energy [kJ/mol]

𝑘𝑘1∗ = 𝑘𝑘1 𝐾𝐾𝐸𝐸𝐸𝐸

0.4777·105

73.16


1.0748·1010

194.07


1.2634·1021

372.54

𝑘𝑘4∗ = 𝑘𝑘4 𝐾𝐾𝑆𝑆𝑆𝑆

2.1488·106

109.22

Adsorbtion constant [bar-1] 𝑙𝑙𝑙𝑙 𝐾𝐾𝐸𝐸𝐸𝐸 = −11.499 + 12300/𝑇𝑇[𝐾𝐾]

𝑙𝑙𝑙𝑙 𝐾𝐾𝐻𝐻2 = −14.6098 + 14193.7/𝑇𝑇[𝐾𝐾] 𝑙𝑙𝑙𝑙 𝐾𝐾𝑆𝑆𝑆𝑆 = −10.5279 + 12582.4/𝑇𝑇[𝐾𝐾]



Page 24 of 50

8 1

The above kinetic model considers the process rate constrained only by the chemical reaction. A more

2

rigorous model should include the effect of intra-particle diffusion. Lee and Froment10 calculated an

3

efficiency factor for the reaction R1 to R3 between 0.6 and 1. Comparing the reactor simulation with

4

homogeneous and heterogeneous models they found slight differences, of about 2% in both

5

conversion and selectivity, thus not of significance for the conceptual design. As noted by Rase12, the

6

difference between models might be initially true, when the reaction is very fast, and the temperature

7

is above 600 °C. In an adiabatic reactor the temperature drops rapidly, such that this correction is

8

limited only on a fraction of residence time, already very short. In practice the catalyst activity is

9

declining in time and the process gradually becomes reaction controlled.

10

The simulation showed that the catalyst selectivity was slightly overestimated when comparing with

11

the experimental data from the Figures 1 and 2. A convenient correction is obtained by setting the

12

pre-exponential factor k3* of the reaction R3 to 1.26341×1021.

13 14

2.4 Preliminary material balance

15

The plant capacity is 100 ktpy styrene for an annual operation time of 8000 hours. Table 5 presents

16

a preliminary material balance in molar and mass units. From the Figure 1 and Figure 2 one may

17

estimate styrene selectivity of 94% for conversion close to 70%. The selectivity sharply drops for

18

higher conversions. Supplementary data from Lee & Froment9 indicates benzene and toluene

19

selectivity of 2 and 4%, respectively. Other secondary reactions leading to waste and impurities are

20

neglected.

21 22

Table 5 Preliminary material balance of a styrene plant MW EB ST Bz Tol H2 CH4 C2H4 Total

106 104 78 92 2 16 28

Price Input USD/ton kmol/h kg/h 900 127.9 13553.6 1500 700 700 300 300 14000 127.9

13553.6

Output kmol/h kg/h 0 0 120.2 12500 2.6 199.5 5.1 470.5 115.1 230.2 5.1 81.8 2.6 71.6 250.6 13553.6

Reaction EB=ST+H2 EB=Bz+C2H4 EB+H2=TOL+CH4

Selectivity 0.94 0.02 0.04

23 24 25

3

Development of alternatives for the reaction section

26

A base-case design will be developed starting from known technology. In a modified hierarchical

27

methodology proposed by the authors the reactor design and energy integration around it should be ACS Paragon Plus Environment

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9 1

solved earlier in the Reactor-Separation-Recycle (RSR) structure13. The design of the reaction section

2

reveals a complex heat integration problem with impact on the energy requirements of the entire

3

process. Moreover, the energy cost dominates over the capital cost in the total annual cost. This

4

approach is different from the Douglas’ doctrine14, where the reactor design was constrained by the

5

cost of recycling the reactants, while the energy integration was left as a downstream activity.

6

Figure 3 presents the simulation of the reaction section, including black-box separation and recycles.

7

The flowsheet comprises a series of two adiabatic reactors PFR-1 and PFR-2 with intermediate re-

8

heating. There are two recycles of ethyl benzene and water. The first comes from the separation

9

section. The second starts from boiling-feed-water (BFW) recycled from the separation section,

10

which is evaporated and superheated at 700 °C in the furnace FURN-1. Fresh and recycled ethyl

11

benzene are mixed and evaporated in H-1, and then mixed with the superheated steam. Further, the

12

mixture, heated up in the furnace FURN-2 at a suitable temperature (630 °C), enters the first catalytic

13

bed (PFR-1), and then the second bed (PFR-2) after re-heating in the furnace FURN-3. The hot

14

mixture leaving the reactor is cooled down to 33 °C by passing it through a series of heat exchangers

15

lumped in the unit COOLER. This simple structure is the starting point for analysing the heat

16

integration problem for the reaction section. The inlet temperatures in PFR-1 and PFR-2 are both set

17

at 630 °C, while the pressures are 2.5 and 2.28 bar. Note that the pressure drop in reactors is calculated

18

rigorously by using the Ergun equation. The molar ratio steam/ethyl benzene at the reactor inlet is set

19

at 12. After some trials the size of both reactors is set at a diameter of 4 m and bed length of 2 m.

20

Note that the diameter is larger than the bed length for limiting the pressure drop, which simulates

21

also the behaviour of radial reactors.

22

Table 6 presents data regarding the reactor design that include the geometry, amount of catalyst, inlet

23

and outlet temperatures and pressures, as well as ethyl benzene conversion and styrene selectivity. A

24

pressure drop of 0.1 bar in the inter-stage furnace was assumed. Beside the base-case, Table 6 reports

25

also the results for the alternatives later explained. It may be seen that the operating pressure and the

26

bed pressure drop have a strong influence on conversion, and implicitly on selectivity. Thus, dropping

27

the inlet pressure from 2.5 bar to 1.2 bar and the bed depth from 2 to 1.3 m raises the conversion from

28

0.648 to 0.741, while the selectivity increases from 0.945 to 0.969, although the catalyst amount is

29

reduced by 35%.

30 31

Table 6 Reactor design and performance for supra- and sub-atmospheric pressures

Diameter / [m]

Base case, supraBase case, near MVR steam atmospheric pressure atmospheric pressure Reactor 1 Reactor 2 Reactor 1 Reactor 2 Reactor 1 Reactor 2 4 4 4 4 4.6 4.6 ACS Paragon Plus Environment


Page 26 of 50

10 Height /[m] Catalyst / [kg] Inlet temperature / [°C] Outlet temperature / [°C] Inlet pressure / [bar] Pressure drop / [bar] Residence time / [sec] Conversion Selectivity

2 35721 630 557.9 2.5 0.11 0.56 0.451 0.961

2 35721 630 593.7 2.29 0.14 0.46 0.648 0.945

1.3 23218 630 545.6 1.2 0.17 0.31 0.466 0.983

1.3 23218 630 581.2 0.93 0.66 0.18 0.741 0.969

1.4 33068 630 554.7 1.8 0.06 0.38 0.453 0.980

1.4 33068 630 593.2 1.64 0.08 0.32 0.702 0.950

1 2 EB-REC

P-1 2.1 bar 596 °C 3 bar 26 °C

MIX-1

2.5 bar 630 °C

H-1 EB-VAP R1-IN

Q=3.36 MW EB-FEED

630 °C

178 °C

R-1

LIGHTS

R-2

MIX-2 700 °C STEAM 3 bar 25 °C BFW

FURN-2 Q=4.42 MW

FURN-3 Q=2.93 MW

SEP

STYRENE

1.3 bar 33 °C ORGANICS

FURN-1 Q=44.14 MW

WATER R2-OUT

COOLER Q=50.36 MW

3 4

Figure 3 Reactor-separator-recycle structure of styrene manufacturing process (base case)

5 6

Figure 4 displays profiles of molar fractions for water-free ethyl benzene and styrene, as well as of

7

temperature in the base-case. The fall in temperature is a measure of the reaction rate. The profiles

8

indicate that the ethylbenzene dehydrogenation is an equilibrium-controlled reaction taking place

9

primarily on a half of the reactor length, the other half serving to a slight progress toward equilibrium.

10

This second reaction zone may provide a catalyst reserve for ensuring the flexibility in operation.

11


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11

1 2

Figure 4 Concentration and temperature profiles in the first dehydrogenation reactor

3 4

Figure 3 displays also data regarding temperature and pressure at key locations, as well as the duties

5

of units for determining the heating and cooling requirements. Thus, the ethylbenzene evaporation

6

needs 3.36 MW, while the steam generation 44.14 MW. Feed pre-heating requires 4.2 MW, while

7

the intermediate re-heating 2.93 MW. In total the hot utility requirements sum up to 54.85 MW. The

8

cooling of reactor effluent down to the first separation step at 33 °C gives a duty of 50.36 MW. Thus,

9

the process shows an energy deficit of 4.49 MW. In other words, only 8% from the energy injected

10

is used by the chemical reaction. The need of energy saving is obvious.

11

Pinch Point Analysis may identify energy saving opportunities and develops alternatives for the Heat

12

Exchangers Network (HEN). This work can be done directly in Aspen Plus 9.0 from a simulation run

13

via the Aspen Energy Analyser (AEA) tool. Several levels of analysis are available, starting with the

14

evaluation of actual consumptions and setting of targets for integration, namely for utilities and CO2

15

emission, as well as sizing exchangers and estimating costs. Note that HEN design is assisted by an

16

automatic tool for heat exchangers matching.

17

A physical insight can be obtained from the examination of composite curves, displayed in Figure 5

18

for ΔTmin of 10 °C. The plots indicate that the largest amount of energy is involved in BFW

19

evaporation. However, this cannot be compensated by condensation of hot reactor effluent because

20

of the lower temperature level.

21

As shown in Table 7, the base-case analysis indicates a possible saving of 43.4 % in total utilities,

22

namely 41.62 % for hot utilities and 45.34 % for cold utilities, as well as 32.6% reduction in carbon

23

emissions. There are possibilities for saving energy, as by using the hot outlet reactor stream for



Page 28 of 50

12 1

ethylbenzene feed conditioning and low-pressure steam generation, described in Ullmann’s

2

monograph1. This solution selected as base-case will be developed in the Process Integration section.

3

The examination of Figure 5 shows that most of the energy involved in the process is concentrated in

4

a pinch zone at about 130 °C due to the water evaporation for steam generation and the condensation

5

of reactor outlet. Moving the two curves would save a tremendous amount of energy, of about 22

6

MW. How could we do this? The answer will be given later in this section.

7 8

Figure 5 Composite curves and grand composite curve in the base-case

9 10

Table 7 Target for energy saving in the base-case Actual

Target

Available savings

% of Actual

Total utilities / [MW]

105.3

59.6

45.7

43.4

Heating utilities / [MW]

54.9

32.1

22.8

41.62

Cooling utilities / [MW]

50.4

27.5

22.8

45.34

11200

7580

3660

32.56

Carbon emissions / [kg/hr] 11 12

But before, let us examine another possibility for steam saving, namely operating at “the lowest

13

workable pressure”. Lower partial pressure of ethyl benzene increases significantly the equilibrium

14

conversion, and consequently less steam is necessary. In this case the inlet pressure in the first reactor

15

is reduced from 2.5 to 1.2 bar, near-atmospheric. Considering line pressure drops, the second reactor

16

operates at a lower pressure, in this case 0.9 bar. Short bed length is used in reactors, from 2 to 1.3

17

m. The first separation step is done under vacuum at 0.5 bar. The pressure drop on the hot effluent

18

line is about 0.5 bar. As shown in Table 6, the conversion increases from 0.648 to 0.74 for a catalyst

19

load reduced by 35%, while the selectivity remains very high at 0.969. If the amount of BFW is

20

adjusted in view of keeping the same conversion, the molar steam ratio may be reduced from 12 to 8.

21

The Aspen Economic Analyzer tool shows a drop in total raw utility consumption by 25%, from 105


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13 1

to 78.3 MW, but this may be further reduced by heat integration. Using low pressure in reactors is

2

applied in recent technologies. EB-REC

2.0 bar 630 °C

H-1 EB-VAP

MIX-1

R1-IN

EB-FEED

R-1

MIX-2

BFW2 Q=8.2 MW

FURN-1B 0.5 bar 82.3 °C

2.5 bar 150 °C

STEAM

FURN-2

FURN-3 Q=2.93 MW

SEP

Q=4.06 MW

BFW1

W-EVAP COMP

LIGHTS

R-2

STYRENE

ORGANICS

FURN-1A

Q=22.15 MW W=3.67 MW Q=10.1 MW

1.2 bar 33 °C

R2-OUT

WATER

COOLER

3 4

Q=47.48 MW

Figure 6 RSR structure by styrene manufacturing, MVR alternative

5 6

In a third alternative, presented in Figure 6, we propose water evaporation under vacuum conditions,

7

combined with mechanical vapour recompression (MVR) as innovative solution. Now the water

8

evaporation takes place at 0.5 bar and 81.3 °C. Then the water vapour is compressed in two stages

9

from 0.5 to 2.5 bar, with inter-stage cooling at 150 °C. The flowsheet remains as in the base-case,

10

with slight modifications. Thus, the inlet of PFR-1 is set at a lower pressure (2 bar). The reactor

11

diameter is slightly increased to 4.6 m, while the bed height is decreased at 1.4 m, which results in a

12

total pressure drop over the reaction section of 0.2 bar. The simulation gives an ethylbenzene

13

conversion of 70.2 % with styrene selectivity at 95 %. The pressure drops over the heat integration

14

line is limited at 0.5 bar, checked later by rigorous design of exchangers.

15

The above modification has a remarkable effect on energy saving.

16

Figure 7 shows the modified composite curves for ΔTmin of 10 °C, while Table 8 the results of the

17

targeting procedure. Note that the hot composite curve is now above the cold curve in the region of

18

water evaporation. The grand composite curve shows a large duty of about 26 MW for heat saving in

19

the pinch region. Indeed, the analysis gives an actual load of 50.1 MW for hot utility but possible

20

saving of 40.1 MW to 10 MW. The actual consumption is 47.3 MW for cold utility but with possible

21

saving of 40.1 MW to 7.2 MW. The reduction in utility requirements is of 80.2 MW from 97.4 MW

22

to 17.2 MW, thus of 82.3%! The explanation of this remarkable effect is that a large amount of energy



Page 30 of 50

14 1

available in the hot stream is recovered for steam generation with a duty of 22 MW. Indeed, the

2

difference in hot utility targets from the Tables 7 and 8 are 32.1-10=22.1 MW. This result is obtained

3

by spending 3.7 MW mechanical energy. The ratio of prices of thermal energy against electricity,

4

roughly three, indicates that this method could be feasible. This alternative will be developed in the

5

section devoted to Process Integration.

6

Figure 7 Composite curves and grand composite curve in the MVR alternative

7 8

Table 8 Utility requirements and targets for heat integration for MVR alternative Actual

Target

Available savings

% of Actual

Total utilities / [MW]

97.4

17.2

80.2

82.3

Heating utilities / [MW]

50.1

10.0

40.1

80.0

Cooling utilities / [MW]

47.3

7.2

40.1

84.6

11870

2377

9493

80.0

Carbon emissions / [kg/hr] 9 10 11

4

Separation system

12

For the assessment of the separations we consider the following mixture resulting from the reaction

13

section (flowrates in kmol/h): styrene 122, ethylbenzene 55, toluene 4, benzene 2.5, methane 4,

14

ethylene 2.5, hydrogen 121, water 2180 and α-methylstyrene 0.8. The last component was introduced

15

as a representative heavy-key impurity. Note that α-methylstyrene is typically produced by the

16

dehydrogenation of isopropyl benzene, originating from propene present as impurity in ethylene by

17

benzene alkylation. The ASTM norm D5135-9519 defines the styrene purity of minimum 99.74%, the

18

other impurities being ethylbenzene 0.043%, α-methylstyrene 0.028% and aromatics. Therefore, the

19

styrene must be obtained as the last top product in the separation sequence, after ethylbenzene/styrene

20

splitter. Note that the nbp of α-methylstyrene is 165 °C versus 145.5 °C for styrene. Neglecting this ACS Paragon Plus Environment

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15 1

operation, demanding as hardware, does not change the process recycle structure but underestimates

2

the capital cost.

3

Figure 8 presents the conceptual flowsheet. The first step consists in splitting the inlet mixture in gas,

4

water and organic streams. This operation may be done by a simple flash at atmospheric or sub-

5

atmospheric pressure. The temperature is set at 33 °C usually achieved with recycled cooling water.

6

Accurate simulation of a three-phase flash requires a suitable thermodynamic model. We select

7

NRTL model with Henry components, which can describe accurately a three-phase equilibrium

8

involving water, organics and supercritical components (hydrogen, ethylene, methane). The accuracy

9

is proven by the prediction of the low styrene solubility in water, which at 33 °C is 0.034 wt%

10

compared with 0.032 the experimental value (Ullmann, 2012).

11

The results show ca. 95 % recovery of styrene and ethyl benzene in the organic phase, with only very

12

low amounts passed in water. On the contrary, important amounts of styrene and ethylbenzene go in

13

the vapour phase together with benzene and toluene. This stream contains also hydrogen with smaller

14

amounts of methane, ethene and water.

15

The recovery of the aromatic components is required for economic reasons. Hence, the vapour phase

16

is submitted to compression in the unit CP-1. In this case, starting at a pressure of 1.1 bar and using

17

a compression ratio of 1.5 ends up in a pressure of 1.65 bar and temperature of 102 °C. Further the

18

vapour is cooled by air and refrigerated water to 7 °C, then condensed and split by a L/V flash. The

19

liquid phase containing recovered aromatics is sent to the distillation column C-1. The gas phase is

20

hydrogen-rich with 94 mol.% H2, 3 mol.% CH4 and 2 mol.% C2H4. Usually this stream is used as

21

fuel, but more profitable would be to treat it as saleable co-product for refinery use or purified.

22

Next, the distillation sequence is built up by considering five components, ordered by normal boiling

23

points (°C) as follows: benzene 80.1 / toluene 110.6 / ethylbenzene 132.6 / styrene 145.7 / α-methyl-

24

styrene 165.4. Two sequences may be proposed, direct (DS) and indirect (IS), each one of two splits.

25

The DS is ABC→A+BC, BC→B+C, while the IS is ABC→AB+C, AB→A+B. Accurate evaluation

26

has shown15 that the candidates are similar as total annual costs, with a slight advantage of indirect

27

sequence. However, we select the direct sequence following the heuristics “remove impurities first

28

and perform the most difficult task last”.



Page 32 of 50

16

BENZENE

C-4

H2

M-1

65°C 1.1 bar 115°C 1.15 bar

CH4+

TOLUENE

H-1

S-2

7 °C

ETHYL BENZENE RECYCLE

P-1

Q=-770 kW 102 °C 1.65 bar

CP-1

W=-70 kW 33°C 1.2 bar R-OUT

STYRENE

GASES

C-1

S-1

30.8°C 0.25 bar

C-2

76.4°C 0.15 bar

C-3

75.2°C 0.10 bar

ORGANIC

WATER

117.4 °C 0.50 bar

114.2°C 0.22 bar

113.2°C 0.40 bar

HEAVIES

1 2

Qc=-470 kW Qr=1281 kW

Q=-9223 kW Qr=9071 kW

Q=-4196 kW Qr=3915 kW

Figure 8 Flowsheet for the separation section of the styrene manufacturing process

3 4

Accordingly, the organic phase from S-1 enters the column C-1 removing in top benzene, toluene,

5

water and light impurities. The column operates under vacuum at 250 mbar. Then the ethyl

6

benzene/styrene mixture is split in high purity fractions in the tower C-2 with top pressure at 150

7

mbar. This difficult separation requires large number of stages and high reflux rate. Peng-Robinson

8

EOS or NRTL with Henry components give similar results. The top stream containing ethyl benzene

9

is recycled to the reaction section, while the bottom stream is sent to the vacuum distillation C-3 in

10

view of final styrene purification. Finally, the top distillate from C-1 is separated in components in

11

the distillation unit C-4. Note that here a small amount of water dissolved in the hydrocarbon phase

12

leaves in the top distillate together with benzene. A suitable simulation option for such operation is

13

SRK with Kabadi-Danner mixing rules and three-phase condenser. It is important to note that the

14

vapour compression after flash separation ensures an overall high recovery of styrene at 100% and

15

ethylbenzene at 99.5%.

16

Inhibitors employed to prevent fouling of internals by styrene polymerisation are typically aromatic

17

compounds with amino, nitro, or hydroxy groups. In the final product small amounts of tert-

18

butylcatechol (TBC) are added (10 – 50 mg/kg) for preservation during storage and transportation.

19

Key data regarding temperatures, pressures and energy loads are marked on the Figure 8. Details

20

regarding streams, sizing and economics of distillation units are given in Table 9. The sizing of

21

internals has been performed by employing the interactive tool available in the Aspen Plus. The ACS Paragon Plus Environment

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17 1

results are the geometric characteristics of column and internals, the operation point with respect to

2

flooding and the pressure drop. In addition, online economic evaluation can be performed by using

3

an interactive tool, in which both detailed sizing of components and cost estimation is performed.

4

The results include both capital and operation costs.

5

The ethyl benzene/styrene C-2 column dominates by far both the equipment and operation costs,

6

being characterised by large diameter, substantial number of stages, large duties of reboiler and

7

condenser. Employing vacuum makes necessary the use of performance internals, as structured

8

packing (Mellapack plus). Minimising these costs can be achieved at best by maximising the

9

conversion in the reaction section.

10 11

Table 9 Design and sizing of the separation section by styrene manufacturing Item Distillate rate Bottom rate NTP / feed location Pressure top Pressure drop Temp. top Temp. bot Reflux rate Q condenser Q reboiler Internals Diameter Height Equipment cost Installed cost Utility cost Total equip. cost Total install. cost Total utility cost

Unit kg/h kg/h

C-1 693 18651 35 / 15 bar 0.250 bar 0.04 °C 30.8 °C 97 kg/h 4000 MW 1.45 MW 1.96 Mellapack plus m 1.4 m 15 M$ 0.300 M$ 0.900 M$/year 0.613 2.9 million USD 4.80 million USD 4.25 million USD/yr

C-2 6001 12650 88/35 0.150 0.120 76.7 113.2 85000 9.26 9.10 Mellapak plus 3.3 28 1.60 2.44 2.482

C-3 12500 150 35/10 0.1 0.06 75.2 114.2 25000 4.7 4.4 Mellapak plus 2.5 12 0.491 1.07 1.053

C-4 438 255 15/7 1.1 0.01 73.0 115.3 2000 0.44 0.44 Pall 0.6 6 0.080 0.390 0.106

12 13

The inspection of Table 9 shows that the separation section requires substantial amounts of energy,

14

namely about16 MW hot utility as LP steam. However, this is only 1/3 from the energy to be injected

15

in the reaction section. One may observe that the utility cost (annual) is about the same with the total

16

installed cost of distillation units, which is charged over several years. This observation demonstrates

17

again that energy saving is a priority for this process.

18

Important energy saving may be further achieved by applying advanced methods, namely for the split

19

ethyl benzene/styrene. A recent study by Cui et al.15 showed that saving of ca. 30% may be obtained

20

by double-effect distillation, and up to 40% by heat pump assisted distillation.



Page 34 of 50

18 1 2

5

Process Integration

3

In this section we tackle the problem of saving energy in the reaction section. An opportunity for

4

considerable saving was identified: vacuum steam generation coupled with condensation of reactor

5

effluent stream. The interactive energy analysis tool in Aspen Plus 9.0 allows the systematic

6

development of a heat exchange network (HEN), then following the performance as saving and costs,

7

and defining targets for the next step.

8 9

5.1 Base-case design

10

As base-case we consider the scheme presented in Ullmans’ Encyclopaedia1 and depicted in Figure

11

9. This scheme includes two feed-effluent heat exchanger (FEHE) units that take advantage from the

12

reactor effluent enthalpy for ethylbenzene evaporation (FEHE-2) and pre-heating (FEHE-1), plus two

13

units for medium steam generation (MPS-GEN) at 10 bar and 160 °C and low-pressure steam

14

generation (LPS-GEN) at 3 bar and 135 °C. The total heat recovered is 14.88 MW, from which 7.03

15

by process/process heat exchange and 7.85 by steam generation. This rather simple scheme may be

16

compared with a maximum of 22.8 MW identified by the super-targeting method (see Table7).

17

Further improvement might be obtained by inserting supplementary equipment, but the economic

18

gain would remain limited. Thus, this scheme may be considered realistic as industrial reference.

19 EB-REC 185 °C

HPS-GEN

FEHE-2 MIX-1 EB-FEED

313 °C

1.54 bar 498 °C

Q=3.43 MW

1.76 bar 594 °C

FEHE-1

491 °C

Q=-6.0 MW

MIX-2

R2-OUT 2.0 bar 630 °C

Q=3.70 MW 493 °C 619 °C

R-2

R-1 700 °C

3 bar 25 °C

FURN-2 0.42 MW

FURN-1 Water

FURN-3 Q=2.93 MW

STYREN

Q=38.3 MW

LPS-GEN 203 °C

COOL-1 150 °C

Q=-29.3 MW

Q=-1.85 MW

20 21

LIGHTS

SEP

Figure 9 Energy saving scheme in the base-case

22


1.1 bar 33 °C

ORGANICS WATER

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19 1

Table 10 presents a summary of key sizing elements of the FEHE units in the base-case designed as

2

TEMA shell-and-tubes heat exchangers employing the interactive tool EDR. Rigorous calculation of

3

partial heat transfer coefficients takes place, including the hydrodynamic regime and pressure drop.

4

Fouling resistances were introduced as maximum 0.00035 W/m2hK. The pressure drop was limited

5

to 0.12 bar on both sides. The hot fluid was allocated in tubes in all cases. The evaporators are

6

designed as AKT kettle boilers, the other units as BEM type.

7 8

Table 10 Heat exchanger network for reaction section in the base-case Heat Exchanger FEHE-1 FEHE-2 MPS-GEN LPS-GEN COOL-1 FURNACE

TEMA type

BEM AKT AKT AKT BEM -

Duty

Hot inlet / outlet

Cold inlet / outlet

Area

Units parallel / series

MTD

kW

°C

°C

m2

-

°C

2900 3275 6000 2500 29600 42500

593.0 / 513.2 320.9 / 197.5 513.2 / 320.9 197.5 / 150 150 / 33 1800 / 700

180.3 / 425.9 32.8 / 180.1 160 / 161 134 / 135 20 / 25 33 / 700

107 180 294 540 930 -

1/1 1/1 1/1 1/1 2/1 -

231.3 100.4 203.9 46.4 33.9 -

HTC W/ m2 K 145 204 100 100 740 -

Δp hot Bar 0.07 0.07 0.07 0.15 0.3 -

9 10 11

5.2 MVR alternative

12

The pinch-point analysis presented in Figure 7 demonstrated that a radical improvement can be

13

obtained by upgrading the heat transfer conditions using vacuum steam generation combined with

14

mechanical vapour compression.

15 ACS Paragon Plus Environment


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20 1

Figure 10 Flowsheet for energy saving with MVR and five FEHE units

2 3 4

Figure 10 presents the energy saving flowsheet with five FEHE units, while Table 11 the results of

5

the heat exchange system using the EDR tool. Note that the hot reactor effluent in the FEHE units

6

circulates in tubes. The steam generation is split in two parts: BFW1 that supplies ca. 80 % from

7

steam sent directly the reaction section, and BFW2 injected with the ethylbenzene vapour. The

8

centrifugal compressor is a 3-stage unit with stage compression ratio of 1.61. The power for

9

compression is reduced by inter-stage cooling at 110 and 115 °C for the first two stages. The discharge

10

temperature of about 180 °C may be tolerated by modern compressors. Considering 80% efficiency

11

gives a mechanical power of 3.3 MW.

12

The unit FEHE-1 serves for preheating ethyl benzene stream after evaporation in FEHE-2, FEHE-3

13

for steam pre-heating after compression, and FEHE-4 for secondary steam generation. Finally, the

14

unit FEHE-5 is used for primary steam generation. These units save together 38.35 MW from the

15

40.1 MW initial target, which is 95.5 % or practically the whole energy available for process/process

16

exchange.

17

The units FEHE-2 and FEHE-1 handle the evaporation and preheating of ethyl benzene from 31 to

18

434 °C, while the hot fluid temperature drops from 593 °C to 429 ° C. They save 6.14 MW. The unit

19

FEHE-3 takes over the remaining enthalpy for heating up the vapour resulting from steam

20

compression from 170 to 415 °C, saving 4.3 MW. The hot fluid temperature drops further from 429

21

to 308 °C, so that the enthalpy can be used for secondary steam generation in FEHE-4 for a duty of 6

22

MW. The hot reactor effluent stream leaves FEHE-4 as vapour at a temperature of 133 °C and a

23

pressure of 1.3 bar.

24

The last unit FEHE-5 contributes to a very large energy saving of 22.2 MW, otherwise lost in water

25

or air cooling. Note that the heat transfer coefficients are very good on both sides, over 2000 W/m2K,

26

since phase transition takes place, condensation and vaporisation, respectively. The dew point of the

27

hot fluid at 1.3 bar is 111.5 °C, while the bubble point of water is at 86 °C at 0.6 bar. Thus, MTD of

28

about 25 °C can be ensured. The overall clean HTC is of 1864 W/m2K, but only 758 W/m2K if

29

considering fouling. The unit COOLER ends up the reactor effluent cooling removing a duty of 9.35

30

MW from 102 °C to 33 °C. Cooling water or a combination with air cooling may be used. A H-curve

31

plot indicates that the condensation process continues, about 5 MW being available down to 85 °C.

32

Regarding the size it can be observed that FEHE-1 and FEHE-2 are small units, while FEHE-3,

33

FEHE-3, and FEHE-5 are large units with similar areas, but different heat transfer conditions.


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21 1

The next question raised is if the steam compression is profitable. A quick economic assessment can

2

be done by comparing the cost of saved thermal energy versus purchasing the compressor.

3

Considering 8000 operation time, the energy cost of 10 USD/GJ and 22.4 MW duty, leads to a saved

4

energy of 6.3 Million USD/year. The purchasing cost of a centrifugal compressor may be estimated

5

with the relation 580000+20000×[P, kW]0.6 from Towler & Sinnott16. Considering a power of 3500

6

kW gives a purchasing cost of 3.26 Million USD, or about 1 Million USD per MW power. This value

7

is close to other evaluations, as proposed by Loh et al.16 from US DOE, or by Internet search. The

8

payback time of 3.26/6.3×12=6 months predicts good profitability.

9

Summing up, the five FEHE units can save an amount of 38.35 MW hot utility. Dividing this amount

10

by the requirement in the base case of 54.9 MW gives a saving of about 70%. Achieving low utility

11

consumption leads saving a considerable amount of CO2 emissions.

12 13

Table 11 Heat Exchangers Network of the reaction section in MVR alternative Heat Exchanger FEHE-1 FEHE-2 FEHE-3 FEHE-4 FEHE-5 COOLER FURN-1B FURN-1A FURN-2 FURN-3

TEMA type

BEM AKT BEM AKT AKT -

Duty

Hot inlet / outlet

Cold inlet / outlet

Area

Units parallel / series

MTD

kW

°C

°C

m2

-

°C

2900 3247 4300 5676 22200 9136 2490 5758 1013 2500

593.4 / 517.3 517.3 / 429.4 429.4 / 307.8 307.8 / 133 133 / 102.0 102 / 33 1500 / 800 1500 / 800 1500 / 800 1500 / 800

178.4 / 434.3 31.2 / 178.4 178.4 / 434.3 33 / 179.5 33.0 / 87.0 20 / 25 133.6 / 700 396.7 / 700 603.9 / 630 530 / 630

100 64 1380 1246 1235 364 79 243 99 162

1/1 1/1 2/1 2/1 2/1 -

231.3 304.9 53.1 48.6 24.4 33.9 630.8 474.3 204.9 308.3

HTC W/ m2 K 145 204 64 106 758 740 50 50 50 50

Δp hot Bar 0.07 0.07 0.03 0.15 0.08 -

14 15

The performance of the heat integration can be compared with the targeting analysis from Table 8.

16

The hot utility requirements fulfilled by furnaces have a total duty of 11.76 MW, while de cold utility

17

takes over 9.14 MW. These values are close to the minimum requirements of 10 and 7.2 MW from

18

Table 8 demonstrating a very good efficiency of the proposed HEN scheme. Further saving is

19

possible, about 1.8 MW for each hot and cold utility, but this will be marginal compared with 38.35

20

MW achieved by implementing the five FEHE units described before.

21 22

6

Plantwide control

23

In this section, the plantwide controllability of the design is evaluated by dynamic simulation. The

24

results show that, despite high integration, the styrene plant is easily controllable.



Page 38 of 50

22 1

Figure 11 presents the plantwide control structure. Note that in practice the water stream is sent to a

2

treatment section, make-up is added, and then recycled. This section effectively “cuts” the recycle,

3

therefore the modelling of BFW recycle is not necessary. Moreover, columns COL-3 and COL-4 are

4

not part of the recycle loop. They do influence the other units of the plant; therefore, their control is

5

a local problem and it is not relevant for plantwide control.

6

The implementation follows the ideas explained in Dimian et al13. The styrene plant is a one-reactant

7

process. The reactant conversion is close to the equilibrium value. In this case, the reactant flow at

8

reactor-inlet (fresh feed + recycle) is the dominant variable which determines the production rate.

9

The recommended control structure (Dimian et al., 2014, Luyben and Tyreus, 1999) makes use of

10 11

feedback control of the reactant inventory, through the following steps: -

12 13

manipulator; -

14 15 16

The flow-rate of reactor-inlet (recycle plus fresh feed) is fixed and is used as throughput

The reactant inventory is determined by measuring the level in the buffer vessel, where fresh and recycled ethylbenzene streams are mixed;

-

The manipulated variable is the fresh ethylbenzene feed, which is adjusted in order to keep the reactant inventory at a constant value.

17

The advantage of this strategy is that the reactor is decoupled from the rest of the plant. The stability

18

of the recycle system is guaranteed if the individual units are stable or stabilized by local control.

19

The steam/ethylbenzene ratio at reactor inlet is kept constant at the prescribed value in feedforward

20

fashion by changing the BFW-2 flowrate. Since ethylbenzene flow FEB1 and boiler feed water flow

21

FBFW-1 are measured, the required flowrate FBFW-2 is calculated as FBFW-2 = R×FEB1 - FBFW-1 and sent

22

as setpoint to a flow controller. Control of the other units (furnaces, heat exchangers, distillation

23

columns) is standard. Thus, the heating duties are manipulated to keep constant furnace-outlet

24

temperatures. Control of VLL separators (pressure, organic and aqueous levels) is achieved by

25

manipulating the vapour, organics and aqueous flow rates. Note that for the lights’ column (COL-1)

26

a two-point temperature control is implemented. This ensures good recovery of both light and heavy

27

components such that the loss of valuable ethylbenzene with the liquid distillate is avoided. Regarding

28

the recycle column (COL-2), a tight control of the distillate composition is not necessary, therefore

29

the reflux flow rate is set constant. The purity of the bottom product (styrene) is controlled by a

30

conventional composition / temperature cascade, where the manipulated variable is the reboiler duty.

31

Note that a linear mathematical model would be necessary in order to develop the control structure

32

based on controllability indices (such as Relative Gain Array, Condition Number, Disturbance

33

Condition Number, etc.). Although it is possible to obtain such model (for example using the


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23 1

linearization tools of Aspen Dynamics), the model is of extremely high order, therefore unsuitable

2

for control design purposes. Ethylbenzene recycle LC

Ethylbenzene

FC

FEHE-2 FT

FY

FEHE-1

FC

BFW-2

REACTOR-1

FEHE-4

TC

FURN-1B BFW-1

FT FC

FURN-2

TC

TC

PC

FURN-3

TC

FEHE-3 PC

TC TC

Hydrogen

LC

SEP VLL

LC

PC

LC

COMP-2

LC

Lights (L)

TC

FC

COL-2

COL-1 COMP-1

Lights (V)

LC

TC

COOLER

PC

LC

FURN-1A

FEHE-5

3 4

REACTOR-2

SEP VLL

Water TC

TC

XC

LC

Water

LC

XT

Styrene

Figure 11 - Plantwide control structure of the styrene plant

5 6

The performance of the above control structure was tested by flow-driven dynamic simulation

7

performed in Aspen Plus Dynamics. Note that the more rigorous pressure-driven dynamic simulation

8

requires sizing of all pumps and control valves and tuning of all flow control loops. In practice, this

9

can be quite easily done; therefore, it is reasonable to assume that good flow control can be achieved.

10

These loops have fast dynamics and are not significant when the plantwide control is considered. All

11

vessels were sized based on the assumption of a 10 minutes residence time. Regarding the heat-

12

exchangers, the following instantaneous models were used: HEATER, HEATX (process-process heat

13

exchange without phase change) and MHEATX (process-process heat exchange with phase change).

14

The dynamics of furnaces was treated by inserting a first-order element with 10 minutes time-constant

15

between the output of the temperature controllers and the actual duty. For all temperature

16

measurements, one-minute lag was assumed. The concentration measurement assumed 15 minutes

17

sampling time followed by 15 minutes delay. The temperature controllers were tuned by identifying

18

first-order plus dead time models and applying the Cohen-Coon settings. The concentration controller

19

was tuned by finding the stability limit using the ATV (auto-tune variation) method and Tyreus &

20

Luyben settings.

21

As key disturbance we consider the production rate. Figure 12 presents the dynamic behaviour of the

22

plant for production increase. After one-hour steady state, the reactor-inlet flow rate is ramped up by ACS Paragon Plus Environment


Page 40 of 50

24 1

10%, from 19473 kg/h (about 183 kmol/h) to 21420 kg/h (202 kmol/h). As result (Figure 12a), the

2

production rate increases proportionally by 10% (from 120 kmol/h to 133 kmol/h). However, the

3

transition is quite slow, taking more than 10 hours. As shown in Figure 12b, the purity of the styrene

4

product is well controlled at 99.95%. Figure 12c shows the evolution of ethylbenzene and boiler feed-

5

water flow rates. The ethylbenzene flow rate, initially at 127.9 kmol/h, grows to 140.4 kmol/h at

6

steady state. The first steam flowrate BFW1 remains constant at 1759 kmol/h, while the second one

7

BFW2 increases from 427.8 to 646.5 kmol/h. As depicted in Figure 12d, during the transition period

8

the recycle flow rate changes by less than 30%, but this variation should not affect the operation of

9

the separations, namely the distillation column COL-2 that runs at high reflux ratio of about 16. A

10

slight decrease of recycle purity is noticed. Figure 12e and Figure 12f illustrates the behaviour of

11

furnaces (note the change of time-axis). The maximum deviation from setpoint is about 10 °C (FURN-

12

1B). However, the temperatures quickly return to their setpoints. The change of FURN-1B duty is

13

rather large, for the following reason: it is preferable that the compressor works at the maximum

14

capacity, thus the flow of BFW-1 is kept constant. In these conditions, during production rate changes,

15

the constant steam / ethylbenzene ratio is achieved solely by changing the BFW-2 flow rate. Thus,

16

the duty of FURN-1B was allowed to increase more than twice the steady state value (steady state

17

value: 8.98 GJ/h; range 0 … 34.8 GJ/h), which provided the necessary overdesign.


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25

1 2

Figure 12 - Dynamic simulation results. Production is changed by increasing the reactor-inlet

3

ethylbenzene flow rate by 10%.

4 5

Figure 13 presents similar dynamic simulation results for the decrease of production rate. Again, the

6

production rate is proportional to the flow rate of ethyl-benzene at reactor inlet. The product purity

7

exceeds the required value (99.95 %). During the transition period moderate changes of recycle flow

8

rate are observed. The temperatures of furnaces are well controlled, only small changes of heating

9

duties being necessary.



Page 42 of 50

26

1 2

Figure 13 - Dynamic simulation results. Production is changed by decreasing the reactor-inlet

3

ethylbenzene flow rate by 10%.

4 5

7

Process performance

6

With the elements obtained so far, we can assess the process performance in terms of material and

7

heat balance. Table 12 presents the component material balance as resulting from the reaction section.

8

The selectivity of styrene formation is 0.9492, very close to the value 0.94 assumed in the preliminary

9

material balance (Table 5). The amounts of benzene and toluene were correctly estimated.


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27 1

On this basis an overall material balance may be formulated in terms of output products, as shown in

2

Table 13. The styrene product rate fulfils the plant capacity of 100000 tpy. One may note the

3

availability of a significant production of hydrogen at ca. 2000 tpy. The residue and different losses

4

are limited to 1% from the styrene production, but this amount could be reduced even more. The

5

material consumption including losses is 1.085 tonne ethyl benzene per tonne styrene. This value is

6

in good agreement with material consumption reported for modern processes. Thus, the Italian ENI-

7

VERSALIS company11 claims a consumption of 1.051 tonne ethyl benzene per tonne styrene with

8

99.95 % purity and per pass conversion at 71 %.

9 10

Table 12 Component material balance of the reaction section

Ethylbenzene ST Bz Tol H2 CH4 C2H4 Total

Input kmol/h kg/h 127.9 13557.4

127.9

13557.4

Output kmol/h kg/h 0.0 0.0 121.42 12627.2 2.47 192.7 3.94 362.5 121.42 242.8 3.94 63.0 2.47 69.2 255.65 13557.3

Reactions

Selectivity

EB=ST+H2 EB=Bz+C2H4 EB+H2=TOL+CH4

0.9492 0.0203 0.0325

1.000

11 12

Table 13 Input/output products’ material balance

Ethylbenzene Styrene Benzene Toluene Hydrogen Gas fuel Residue Total

Input kg/h 13557

13557

tpy 108459

108459

Output kg/h 12501 200 377 243 132 105 13557

Tpy 100007 1596 3012 1943 1058 843 108459

13 14

The cost of energy may be estimated from the energy balance. We consider an uncertainty margin of

15

25% with respect to the requirements of the reaction and separation sections. For the base-case and

16

MVR alternative the energy cost is 11.08 and 5.73 M$, respectively (Table 14), while for the

17

separations this is 4.25 M$ (Table 9). It results that the base-case the cost the energy cost per tonne

18

styrene is 15.33×106×1.25/105=191.25 USD/t, while in the MVR case this is reduced to

19

5.73×106×1.25/105=71.63 USD/t. One may conclude that saving energy is an important factor for

20

ensuring the process profitability. ACS Paragon Plus Environment


Page 44 of 50

28 1 2

8

Economic analysis

3

In this section we present an economic analysis that aims emphasising the advantage of this

4

innovative approach in economic terms. For comparing the conceptual alternatives, we follow the

5

approach of the Total Annual Cost (TAC) that combines both energy and capital costs8:

6 7

TAC = Energy cost + Capital cost/3

(8)

8 9

In the above relation the capital cost should be considered as the installed cost of equipment.

10

Table 14 presents the cost estimation of heat transfer equipment in the base-case and the MVR

11

alternative. The purchased cost of FEHE units is retrieved from Aspen Plus ver. 9.0 as calculated with

12

the EDR interactive tool. For other shell-and-tube heat exchangers, as COOL-1, the cost equation is

13

retrieved from Aspen as Chx=10000+800×A[m2]0.8. The heat recovery by steam generation was

14

assumed similar with kettle reboilers in which the vapour fluid circulates in tubes and water in shell.

15

The cost is 2.5 times the value given by the previous equation.

16

The steam for reaction is supplied by a furnace imbedding FURN-1, FURN-2 and FURN-3 duties.

17

The cost equation is17 Cf=43000+111000×Q[kW]0.8.

18

The purchasing cost of compressor is calculated with the relation17 Ccp=580000+2000×W[kW]0.6.

19

When not available from Aspen Plus, the installed cost of equipment is obtained by multiplying the

20

purchased cost by the following installation factors17: heat exchangers 3.5, furnace 2, compressor 1.5.

21

For calculating the cost of energy, the following prices are assumed:

22

-

Thermal energy as steam delivered by furnace or recovered: 10 USD/GJ.

23

-

Electrical energy: 0.08 cts/kWh or 22 USD/GJ.

24

-

Cooling water: 0.5 USD/GJ.

25

Table 14 compares the cost of heat transfer equipment in the two alternatives. The purchase cost of

26

process heat exchangers is similar, of 3.3 and 4 M$. The cost of furnace much higher in the base-case

27

gives finally an installed cost of 7.85 and 5.74 M$ respectively, thus 27% lower in the MVR

28

alternative.

29

Table 15 evaluates both the energy and the installed equipment costs as total annualized costs (TAC).

30

Since the MVR alternative includes the costs incurred by vapour compression, the total equipment

31

cost raises by about 15% with respect to the base-case. However, the cost of energy makes a clear

32

difference between the two alternatives. The annual energy cost in the base case is of about 11 M$,

33

despite the credit for steam generation. The cost of energy in the MVR alternative drops sharply by


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29 1

73%, mainly because the reduced hot utility load, from 12.24 to 3.38 M$. Thus, the saved thermal

2

energy can pay off largely the compression cost of 2.22 M$.

3

The comparison on Total Annual Costs highlights that the MVR method may drop significantly the

4

energy and heat transfer equipment costs involved in the reaction section, by about 36%.

5

From a plant design perspective, it is interesting to examine the economics of the separation section

6

too. Table 3 shows an installed equipment cost of 4.8 M$ (1.6 M$/y on annualized basis) and utility

7

cost of 4.25 M$/y. Again, the cost of energy surpasses the cost of equipment. Note that both costs are

8

substantially lower than in the reaction section.

9 10

Table 14 The installation cost of heat transfer equipment Base-case Q A MW m2 2.90 100 3.28 180 6.0 300 1.85 540

PCE M$ 0.06 0.15 0.21 0.27

COOL-1 29.6 930 Purchase cost of equipment Installed cost heat exchangers Furnaces 42.5 Installed cost of furnaces Installed cost of equipment

0.20 0.95 3.31 2.27 4.54 7.85

Item FEHE-1 FEHE-2 MPS-GEN LPS-GEN

MVR alternative Q A M m2 2.9 100 3.25 180 4.3 1500 5.7 1400 24.7 1150 9.1 300

Item FEHE-1 FEHE-2 FEHE-3 FEHE-4 FEHE-5 COOL-1

11.7

Furnaces

PCE M$ 0.06 0.15 0.21 0.27 0.37 0.09 1.15 4.02 0.86 1.72 5.74

11 12 13

Table 15 Comparison of economics of base case and MVR alternatives Energy MW Heat equipment Process/Process Steam generation Hot utility Cold utility Compressor Total cost An. cost Total Annual Cost

5.26 7.85 42.5 29.6 0

Base case Energy cost M$ 0 -2.26 12.24 1.10 0 11.08

Cap. Cost M$

Energy MW

0.73 1.89

38.35 0 11.75 9.1 3.5

4.54 0.70 0 7.85 2.62 13.70

14 15 ACS Paragon Plus Environment

MVR case Energy cost M$ 0 0 3.38 0.13 2.22 5.73

Cap. cost M$

3.72 0 1.72 0.30 3.26 9.00 3.00 8.73


Page 46 of 50

30 1

9

Conclusions

2

This paper deals with an innovative method for energy saving for manufacturing styrene by

3

ethylbenzene dehydrogenation in adiabatic type reactors. The reactor design considers high-

4

performance catalyst capable of ensuring selectivity of about 95% at conversion of 70%. Steam used

5

as inert brings the energy for the endothermic reaction, enhances the equilibrium conversion and

6

protects the catalyst against coke formation.

7

In the base-case the molar steam ratio is set at 12, reported optimal in several studies. The adiabatic

8

reactors are designed using detailed kinetics and rigorous pressure drop calculation. The reactors have

9

larger diameter than length such to reduce the pressure drop below 0.1 bar. This construction mimics

10

radial type reactors. Two reactors are employed with inter-stage re-heating at 630 °C.

11

The second alternative considers near-atmospheric pressure in reactors. Keeping the same conversion

12

is possible with lower catalyst load. The catalyst productivity increases by 53%, while a substantial

13

saving of 30% in utilities is obtained.

14

The third alternative develops an innovative solution that result in a spectacular energy saving. The

15

idea is based on generating the steam under vacuum, around 0.5 bar, followed by pressure rise by

16

mechanical vapour recompression. Water evaporation for steam generation is the largest energy-

17

consuming operation taking about 40% from the heating requirements for the reaction section.

18

Upgrading the steam by MVR allows matching the temperature-enthalpy profiles of the cold and hot

19

streams in an evaporation/condensation zone, within a driving force of 25 °C. A centrifugal

20

compressor may be used for rising the steam pressure. An efficient heat exchange network of five

21

FEHE units is developed by using Pinch Point Analysis. Employing five FEHE units drops utility

22

consumption by 73% with respect to the base case.

23

The design of the separation section is dominated by the presence of an expensive distillation column

24

for splitting styrene from ethyl benzene, further recycled. Rigorous hydraulic calculation considering

25

structured packing gives a pressure drop limited below 0.15 bar for 90 theoretical stages.

26

The feasibility of the proposed method is supported by dynamic simulation. Disturbances of +/- 10%

27

in production rate are analysed. The plantwide structure is based on keeping constant the ethylbenzene

28

flow rate entering the reactor and manipulating the fresh feed. On the steam side the disturbances are

29

handled by manipulating one water stream. The plant response is robust with a transient of about 10

30

hours. The styrene purity is kept over 99.8%. Only a slight variation of ethylbenzene purity may be

31

noticed, while the flowrate change is limited below 30% with negligible influence on separations.

32

An economic analysis demonstrates the profitability of the proposed method. The cost of energy

33

injected in the reaction section for steam generation dominates over the annualized equipment costs.

34

In the base case the ratio is 4.2, while in the MVR alternative this is reduced to 1.9. In the last case ACS Paragon Plus Environment

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31 1

the energy cost drops by 48 %, but the equipment cost raises by 15%. Both effects results in a total

2

annualized cost reduction by 36%.

3

As more generic value, the paper demonstrates that the cost of energy for running highly endothermic

4

reactions may be significantly larger than the cost of process equipment, in term of annualized costs.

5

If water is implied in reaction engineering, as inert or product, substantial energy saving may be

6

obtained by thermal upgrading of inlet or effluent streams using mechanical vapour recompression.

7

This method developed here has been applied recently by the authors to a highly exothermic reaction,

8

the MTO (methanol-to-olefin) process18.

9 10



32 1 2 3 4

TOC Graphic

5

6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24


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(1) James, D. H.; Castor, W.M. Styrene. In Ullmann’s Encyclopaedia of Industrial Chemistry; Wiley-VCH, 2012. (2) Clough, D. E.; Ramirez, W. F. Mathematical modelling and optimization of the dehydrogenation of ethyl benzene to form styrene. AIChE Journal, 1976, 22, 1097-1105. (3) Sheppard, C. M.; Maier, E. E.; Caram, H. S. Ethyl benzene dehydrogenation reactor model. Ind. & Eng. Chem. Proc. Des. Dev. 1986, 25, 207-/210. (4) Elnashaie, S. S. E. H.; Elshishini, S. S. Modelling, simulation and optimization of industrial fixed bed catalytic reactors; Gordon and Breach Science Publishers, 1993. (5) Yee, A.K.Y; Ray, A.K; Rangaiah, G.P. Multiobjective optimization of an industrial styrene reactor. Comp. Chem. Eng. 2003, 27, 111-130. (6) Luyben, W.L.; Tyreus, B.D. Plantwide Control; McGraw-Hill, 1999. (7) Vaseduvan, S.; Rangaiah, G. P.; Murthy Konda, N.V.S.N.; Tay, H.T. Application and evaluation of three methodologies for plantwide control of the styrene monomer plant. Ind. Eng. Chem. Res. 2009, 48, 10941–10961. (8) Luyben, W.L. Design and control of the styrene process. Ind. Eng. Chem. Res. 2011, 50, 1231–1246. (9) Lee, W.J. Ethylbenzene dehydrogenation into styrene: kinetic modelling and reactor simulation; Texas A&M University, 2005. (10) Lee, W.J.; Froment, G.F. Ethylbenzene dehydrogenation into styrene: kinetic modelling and reactor simulation. Ind. Eng. Chem. Res. 2008, 47, 9183-9194. (11) ENI-Versalis, Styrene, proprietary process technology. (https://www.versalis.eni.com, last accessed on June 2018) (12) Rase, H. Handbook of Commercial Catalysts; CRC Press, 2016. (13) Dimian A.C.; Bildea, C.S.; Kiss, A.A. Integrated design and simulation of chemical processes; Elsevier, 2014. (14) Douglas, J.M. Conceptual design of chemical processes; McGraw-Hill, 1988. (15) Cui, C.; Li, X.; Guo, D.; Sun, J. Towards energy efficient styrene distillation scheme: From grassroots design to retrofit. Energy 2017, 134, 193-205. (16) Loh,

H.P.;

Lyons,

J.;

White,

C.W.

Process

Equipment

(https://www.osti.gov/biblio/797810/, last accessed March 2019)


Cost

Evaluator;


(17) Towler, G.; Sinnott, R.K. Chemical engineering design: principles, practice and economics of plant and process design; Butterworth-Heinemann, 2013. (18) Dimian, A.C.; Bildea, C.S. Energy efficient methanol-to-olefins process. Chem. Eng. Res. & Des. 2016, 131, 41-54 (19) ASTM Standards, D5135-95. Standard Test Method for Analysis of Styrene by Capillary Gas Chromatography; ASTM International, West Conshohocken, 1995.


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Energy efficient styrene process: design and plantwide control

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