Selective hydrogenation of phenylacetylene from a C8-naphtha fraction in a monolith reactor and separation into styrene, ethylbenzene and xylenes

FVO Nr.

Fabrieksvoorontwerp

Vakgroep Chemische Procestechnologie

Onderwerp

Selective hydrogenation of phenylacetylene from

a C8-naphtha fraction in a monolith reactor and

separation into styrene, ethylbenzene and xylenes

Auteurs

R.V. Goede

M.M. Stork

M.l.

van der Weiden

B.l.

Zuurdeeg

Keywords

Telefoon

0181 - 640237

015 - 2615105

0184 - 684564

015 - 2614303

selective hydrogenation, catalysis, rnonolith, separation, Pd, phenylacetylene, styrene, ethylbenzene, xylenes, AEP

FVO Nr.

Fabrieksvoorontwerp

Vakgroep Chemische Procestechnologie

Onderwerp

Selective hydrogenation of phenylacetylene from

a C8-naphtha fraction in a monolith reactor and

separation into styrene, ethylbenzene and xylenes

Auteurs

R.V. Goede

M

.

M. Stork

M.l. van der Weiden

B.l. Zuurdeeg

Keywords

Telefoon

0181 - 640237

015 - 2615105

0184 - 684564

015 - 2614303

selective hydrogenation, catalysis, monolith, separation, Pd, phenylacetylene, styrene, ethylbenzene, xylenes, AEP

Datum opdracht

Datum

verslag

4 - 11 - 1996

31 - 1 - 1997

Abstract

This is a conceptual design, made as an assignment for the study of Chemical Technology at the Delft University of Technology. A design is presented for a phenylacetylene hydrogenation plant for the production and purification of 20 kton of styrene per year. In the plant phenylacetylene in a Cg-fraction from a naphtha cracker is hydrogenated selectively to styrene in a 0.1 % Pd/C monolith reactor to a maximum concentration of 10 ppmw. The overall selectivity is 30%. The reactor effluent is separated into ethylbenzene (7.6 kton; 99.7%), xylenes (14.3 kton; 99.1%) and styrene (19.9 kton; 99.9%). For this difficult separation 1-(2-amino-ethyl)-piperazine is used as an extractive agent. Byproducts are heavies (mainly oligomerised styrene) and lights (mainly methane). A hazard and operability study has been performed revealing a lot of possibilities for safe and steady processing. The economy se ems very favourable: the Return On Investment is 50% and the Internal Rate of Return is 42%.

Contents

1. Introduction . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1 1.1 General introduction . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .. 1 1.2 Assumptions . . .... . . .... .. . ... .. . . .. . . . .. .. . . 1 1.3 Location . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .. 3 1.4 Battery limits . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 3 1.5 Used chemicals . .. .. . . ... .. . .... . . .. .. . .... 5 2. Pracess structure ... .. .. . . . .. .... ... .. . .. .. . .. . . .. . .. . .... . . .. . 6

2.1 Motivation of the structure . .. .... ... . . . .... . . ... . . .. .. . . . 6

2.2 Summary of flowsheet .. . . ... . .. .. . ... . .. . .... ... 8

2.2.1 Reaction system . . . ... . . .. . . .. . . 8

2.2.2 Separation system .. . . . ... . . . . . . . . . . . . . . . . . . . . . . . 9

3. Calculations: flowsheet and units . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 10

3.1 Introduction . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .. 10

3.2 Simulation . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .. 11

3.2.1 Reactor model and kinetics . . . . . . . . . . . . . . . . . . . . . . . . .. 11

3.2.2 Thermodynamics: Margules parameters .. ... .. . . .. . 16

3.2.3 Distillation . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .. l7 3.3 Apparatus design . . .. . . .. . ... ... . . ... . . . .. . .. .. . . . 18

3.3.1 Reactor . . . .. . . . .. . . .. .... .. ... .... . .... .. . 18

3.3.2 Pumps ... . . . .. . . ... .. . .. . . 19

3.3.3 Heat exchangers and heat integration . . . . . . . . . . . . . . . . . . . . . 21

3.3.4 Knock out drum . . . .. ... . . . .... ... .... . . . ... . .. . 23

3.3.5 Distillation towers . . . . . . . . . . . . . . . . . . . . . . . . . . 24

3.3.6 Valves .. .... .... .. . . ... .. . . ... . . .. ... ... .. . . .... 25

4. Mass and heat balance . . . ... . .. . . . .... . . ... .. . . .. . ... . . 26

5. Process contral . . .. . ... .... . ... . .. . . .. . . . .. . .. .... . . 27

5.1 Introduction .. . . .. ... . .... . . .. .. . . . .... .. .. 27

5.2 Flow contral .. . .. .. ... ... .... . . .... . . . ... .. . . ... 27

5.3 Production level control . .. . ... . . . ... . . ... . ... . . ... 27

5.4 Level control . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 27

5.5 Pressure control . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 28

5.6 Heat transfer contral . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .. 28

5.7 Temperature control .... .. . . .. . . . .. .. .. .. .. .... . . .. ... ... 29

6. Process safety, health and environment . . . . . . . . . . . . . . . . . .. 30

6.1 Introduction . . . . . . . . . . . . . . . . . . . . . . . . . . . . 30

6.2 Substance related hazards .... . . . .... . . . .... . . .. . . .. .. .. . 30

6.2.1 Safety and health . . . . . . . . . . . . . . . . . .. 30

6.2.2 Environment . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .. 33

6.3 Hazard and operability study . . . . . . . . . . . . . . . . . . . . . . . . . 34

6.3.1 Hazop for reactor system . . . . . . . . . . . . . . . . . .. 34

6.3.2 Hazop for knock out drum . . . . . . . . . . . . . . .. 35

6.3.3 Hazop for distillation tower . . . .... . . .. . . .. 37

6.4 Main SHE problems .. .. . . ... . .. . ... . .. .. ... .. . .. 37

7. Economy .. ... .. . . .. ... .. . .. ... . .. .. . . . ... . . .. . . .. 38

7.1 Introduction . .. . . .. . .. . . . .. .. .. . . .. . . 38

7.2 Calculation of direct production cost .. .... . . .. ... . . 38

7.3 Calculation of wages .... . .. . . .. ... . ... .. ... . . .. .... 38

7.4 Calculation of total investment cost ... ... .. .... ... . .. . ... .. ... . 38

7.5 Calculation of total yearly cost .... . . ... . ... . . . .. 38

7.6 Yield . . . ... ... . .. . .... .. . ... . . .. ... . . .. . 39

7.7 Economical criteria . .. . .. . . .... . ... . . ... . ... 39

7.8 Discussion . . .. .. . . .. ... . . .... . . .. . . ... 40

8. Discussion and recommendations . . . .. .. . .... . . 41

9. Conclusions . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .. 44 10. Literature .. . .. . . .. ... .. . . ... .. . . 45 11. Acknowledgements . . . .. . . ... . ... . . .. .. . .... . . . ... . . ... 47 Appendix I II nI IV Justification of parameters for monolith . . . .. . . ... ... .. .. . Estimation of reactor cost . . . . . . . . . . . . . . . . . . . . . . . . . v

Optirnisation method .. .. ... . . . . . . . . . . .. . . viii

Specification forms . . . ... . . .... . ... . . .. . ... . ix REACTOR . .. ... . . .... .. .... . . .. . .. ix CENTRIFUGAL PUMPS .. .. .. . . .. . . . .. . . .. . .. . . .. . xi TOWERS . . . . .. . .. ... . . .. . . .... . . . xvi HEAT EXCHANGERS ... ... . .. ... . . .. . xx V AL VES . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . xxxvi

KNOCK OUT DRUM . . . . . . . . . . . . . . . xxxviiii

1. Introduction

1.1 General introduction

This conceptual design is made as an assignment for the study of Chemical Technology on the Delft University of Technology. A short description of the assignment is given below.

The Cg-fraction of a naphtha cracker contains a large amount of styrene. Apart from the styrene some phenylacetylene, ethylbenzene and a mixture of ortho-, meta- and para-xylene are present. Styrene represents a commercially attractive product but cannot be used easily because of the presence of phenylacetylene. Phenylacetylene is a polymerisation inhibitor and a crosslinker and it can hardly be separated from the styrene by means of simple distillation. Phenylacetylene can, however, be removed by selective catalytic hydrogenation to styrene. Simutaniously, styrene is hydrogenated to ethylbenzene. The reaction equations are:

Phenylacetylene + H2 Styrene + H₂ --> Styrene -->.Ethylbenzene (1.1 ) (1.2)

A high selectivity towards styrene is desired. Nowadays, in styrene polymerisation, longer linear chains are preferred because of their better properties, such as improved strength. The phenylacetylene acts as a brake on chain growth and as a crosslinker. Polymerising styrene in presence of phenylacetylene results in polystyrene containing large amounts of styrene monomer. This styrene monomer is toxic. Hence the amount of phenylacetylene in styrene is restricted. Normally, specifications are at a level of 10 ppmw.

Because the removal of phenylacetylene by hydrogenation becomes more difficult as specifications tighten (poor selectivity towards styrene), better reactors are needed.

Nowadays a trickle bed reactor is used. However, operation close to plug flow is desirabie when dealing with a consecutive reaction mechanism, as in this case. A monolith reactor is operated much closer to plug flow than a trickle bed reactor. One of the purposes of this preliminary design is the evalutation of a monolith reactor in order to make a comparison between a trickle bed reactor and a monolith reactor possible.

This paper presents the results of a preliminary design on the performance of a plant with a three-phase monolith reactor for the selective hydrogenation of phenylacetylene to styrene. The product stream is separated into styrene, xylenes and ethylbenzene.

The styrene produced by this plant can be used for mass polymerisation. If the separation of styrene from the ethylbenzene-xylene mixture is not included, the styrene can be used for solvent polymerisation.

1.2 Assumptions

The total production of styrene is 20 kton per year. The total feed flow is resulting. Because the feed streams are eontinuous and bigger than 45 ktona, the stated amount in [15], and the use of a monolithie reactor, a eontinuous production is ehosen. The feed streams are at a temperature of 50°C and a pressure of 3.0 bar. The plant is on stream for 8000 hours per year. The plant is depreciated in 10 years.

H-z.-Re

le

AEP-rnoke

U

L-V

split

Sepor<atJon

Figure 1.1: Schematie representation of the proeess

The monolith has to be regenerated onee a year. This can be done in situ by carefully burning of the deposited heavies with diluted oxygen; the heavies eon sist of oligomers of styrene. This will take about a week.

The required speeifieations for styrene, ethylbenzene and mixed xylenes ean be found in the table below. All other strearns are sold as fuel to other on-site faeilities.

Table 1.1: Required purity of product streams ProdUct stream Purity

Styrene 99.8% (max. 10 ppmw phenylaeetylene) Ethylbenzene 99.7%

The monolith reactor is used as a cocurrent downflow reactor, i.e. both liquid and gas flow down, caused by gravitational force. The velocity of the liquid and gas can be set by the pressure difference over a monolith bloek.

1.3 Location

The plant is designed for the DSM site in Geleen, the Netherlands. At this site, all relevant utitilities are present. There are, besides other units, a naphtha cracker, an ethylbenzene dehydrogenation plant and a styrene polymerisation plant. At. this moment, phenylacetylene is hydrogenated using a trickle bed reactor. Replacement of the trickle bed reactor by the monolithic reactor using the present separation system might be an option.

If the plant would be situated at another site, the following considerations should be taken into account. Since the feed stream comes from a naphtha cracker it seems logical to situate the plant near a naphtha cracker. An industrial area is favourable because of the availability of utilities, possible purchasers of products and the short transport lines for the raw materials and most products. Processing capacity for styrene, ethylbenzene and xylene mixture on site is advantageous. The low MAC value for some of the materiais, e.g. styrene and ethylbenzene, spe aks well for an area with a low population density. Generally, any big refinery, e.g. SHELL Pemis, the Netherlands, would be an appropiate location.

1.4 Battery limits

The Cs-fraction of the cracked naphtha stream is fed continuously to the phenylacetylene hydrogenation plant. The temperature of this stream is 50 °C, the pressure is 3 bar. The hydrogen needed for the reaction comes from a continuous hydrogen/methane (83 volume% hydrogen) gas stream, also from the naphtha cracker; the temperature is also 50°C; the pressure is 3 bar.

Energy can be supplied by means of three pressure levels of steam and by fired heaters. Cooling water is available at a maximum (design) temperature of 20°C from the Whilhelmina kanaal. According to governmental regulations, the cooIing water is not allowed to be drained at temperatures above 40

oe.

The styrene product stream is delivered continuously to the intermediate storage of the polymerisation plant at a temperature of 40°C and at ambient pressure. Some 1-(2-amino-ethyl)-piperazine (AEP, 10-50 ppmw) is left in the styrene stream as polymerisation inhibitor. The presence of this AEP in the styrene stream will not cause any problems because, in the polymerisation plant an excess amount of initiator is added to compensate for the presence of AEP. The ethylbenzene byproduct stream is sent directly to a dehydrogenation plant at ambient pressure and a temperature of 40

oe.

The xylenes are stored in a tank at ambient pressure at temperatures above 15°C because of crystallisation of p-xylene (tm=13.2 0c) until they are transported by ship. The storage tank is supplied with a heating-apparatus to prevent crystallisation in case of prolonged cold.

The separated gases from the knock out drum and the reactor purge are used as fuel on site as do the formed heavies in the distillation columns. This fuel is used on site.

The composition of the used feed is obtained from [30] and can be found in the table below.

Table 1.2: Typical composition of a Cg-naphtha fraction

Compound Massfraction [-] ethylbenzene 0.150 p-xylene 0.099 m-xylene 0.124 phenylacetylene 0.010 o-xylene 0.1l7 styrene 0.500 TOTAL: 1.000

The composition of the make-up AEP stream is unknown; therefore it will be taken as 100% pure AEP.

The separation of the three xylene isomers or the dehydrogenation of ethylbenzene is not

1.5 Used chemicals

A list with properties of the chernicals used in this design can be found below.

Table 1.3: List of used chemicals

N:lme Short Formul<l Mw t' , tm' n:une (kg/mol) (0C] (0C] hydrogen H2 H, 0.00202 ·252.9 -259.3 mcth:lIlc ME CH. 0.018 -162 -182 cthylbenzene EB c"H ,o 0.1062 136.2 -94.9 p-xylene PX c"H,O 0.1062 138.4 13.2 m-xylene MX c"H 10 0.1062 139.1 -47.8 phenyl<lcetyl~ne PH c"H. 0.1022 143.0 -44.8 o-xylene OX c"H,O 0.1062 144.5 -25.2 styrene ST c"H. 0.1042 145.1 -31.0 N-(:uninoethyl)- AEP C.H,~NJ' 0.1292 222 -18 piper<lZine distyrene DS C,.H ,• <l 0.2084 ±300 a: at atmosphenc pressure b: LD50_rat (mglkg) c: d:

M

I

I /

-J

'w-C-C-N

\-I

I

I

'\.

p (20·C) (kg/ml_] 0.088 0.7 867.0 861.1 864.2 930.0 880.2 906.0 1000.0 MAC _{P ..} LC.oc .... , ... (ppm] (mbar) (mg/I) 50 9.3 100 8.0 18 100 8.0 16 7.9 100 8.0 13 25 6.0 5000' 0.1

I

I

I

C=C-c.

I

-C-I

All prices are estimates of current marketprices and are provided by Shell Research and Technology Centre Amsterdam and DSM Geleen. All prices are calculated in US Dollars ($). The rate of exchange is based on the rate in January 1997: S 1

=

f. 1.80. Besides the prices

indicated in table 1.3, the prices indicated in table IA are used.

Table IA: Mixture prices.

Mixture Price [$ / ton]

Hydrogen / Methane 775 250 Xylenes 290 Fuel 190 pric.: (S/ton] 350 390 290 350 1100 25-!0 190

2. Process structure

2.1 Motivation of the structure

Before simulating the process the sequence of steps needed to reach the objectives in the introduction should be determined. The alternatives should be listed in order to obtain an overview of all possible ways to reach the objectives.

According to Douglas' heuristics [15] phenylacetylene should be separated from the styrene before entering the reactor. Unfortunately this separation is very difficult. In a U.S. patent from the DOW corporation [1] a method for this separation is described, but the hence obtained styrene stream only has a purity of 99.6% and a lot of phenylacetylene (far over 10 ppmw) in this stream is expected, due to the resemblance of both substances. So, af ter this separation the styrene stream would still have to go through the hydrogenation reactor to hydrogenate the remaining phenylacetylene.

No other literature nor patents describing high recovery separation of phenylacetylene from styrene are found. So, it is clear the phenylacetylene has to be hydrogenated in presence of styrene. When hydrogenating the phenylacetylene, the temperature rise should not be too high, to prevent catalyst fouling by deposits of oligomers of styrene which would increase the risk of a thermal runaway. So, in this conceptual design it is chosen that the whole Cs-fraction of the cracker is sent through the reactor and separated afterwards.

The aim of this conceptual design is to devellop a monolith reactor in order to be able to compare it with a trickle bed reactor. Therefore, the choice for a monolith reactor is tri vi al. The designed monolith reactor wiU be compared with data from DSM, concerning a trickle bed reactor operated using a 0.1 wt% Pd/C catalyst. Therefore, the choice for this catalyst is also trivial.

Often gas recycle is a profitable but expensive part of a process. It is profitable because it reduces raw material cost, but the required compressor consurnes a lot of energy and is thus expensive in use. Fortunately, when using a monolith reactor it is possible to recycle the gas by internal recirculation. Therefore, no compressor is needed for the gas recycle.

Industrial hydrogen of ten contains a large amount of methane. This is an inert in the hydrogenation reactor, which will accumulate unless a gas purge is present. Therefore, a gas purge is installed.

The liquids leaving the reactor contain a smaU amount of dissolved gasses. To prevent problems in the separation system a knock out drum is placed between the reaction system and the separation system.

Due to the very small difference in boiling point between o-xylene and styrene (0.6 °C) simple distillation of the Cs-stream to isolate styrene is almost impossible. Fortunately, literature suggests several alternatives for the isolation of styrene from a Cs-naphtha fraction.

One alternative is the use of zeolites. [27] indicates that the molecules with the smallest diameters (styrene, phenylacetylene, p-xylene and ethylbenzene) can be separated from the others (o-xylene and m-xylene) using zeolite ZSM-S. Subsequent simple distillation of the styrene is possible. However, at this time the recovery of styrene is restricted to a maximum level of 95% [21, p723]. Furthermore,. little information is found concerning the absorbtion kinetics and the selectivity. Finally, this method of separation is not proven technology yet.

Another alternative is crystallisation of p-xylene and o-xylene. Subsequently, again, simple distillation of the styrene is possible. This method results in streams of separated xylenes, which increases the yield somewhat (this would imply the installation of two crystallizers). However, the high energy consumption needed to cool the liquid below the melting point of o-xylene (T m=-25.2 0c) makes the process expensive.

AIso, [23] suggests the use of an Ag-membrane for this separation. The cost of tbis process increase linearly with capacity when sealing up, making the process very expensive.

Finally, many extractive distillation processes using different extractive agents (N-methylpyrrolidon, N-formylmorfoline, sulfolane [25], esters [26] and N-(aminoethyl)piperazine [1]) are found in literature and patents.

Because extractive distillation is proven technology, relatively cheap and relatively easy to operate and design, this way of separating is chosen in this conceptual design.

A good extractive agent should, according to the authors, have the following properties: - It should not be too expensive.

- It should have a low or na MAC value.

- It should have a boiling point that differs significantly (at least 20°C) from the boiling point of styrene to prevent an azeotropic mixture from being formed.

- It should oe polymerisation inhibiting.

N-(aminoethyl)piperazine (AEP) is, of the above mentioned extractive agents, the only one that fulfils all these criteria and is therefore chosen.

Af ter the extractive distillation the extractant has to be stripped from the styrene stream in a distillation column at a less reduced pressure.

Ta prevent accumulation of heavies formed in the distillation section the heavies are separated from the AEP recycle stream in another distillation column.

The separation of ethylbenzene can be performed using simple distillation. However, [21, p710] shows the required column has a lot of trays (300) and a high reflux ratio. Therefore, alternatives are investigated. [24] indicates the possibility of executing this separation by adsorption using a molecular sieve (sodium Y or potassium Y). This alternative introduces many extra process steps and has a low selectivity and is therefore not chosen. Another alternative, crystallisation, is very expensive because of the high energy consumption when cooIing the mixture (tm.m-xYlene = -47.8 0c) and is therefore also not chosen. So the best

alternative for the separation of ethylbenzene from the xylenes remains simple distillation, although this implies that the distillation column is very high. It is worth considering to avoid this separation at all by introducing the mixture directly to the PAREX-process [21, p723].

Normally a design temperature of 100°C is used for columns in which styrene is present but in presence of a polymerisation inhibitor, like ABP, the design temperature can be risen significantly (to 140°C). To keep the bottom temperatures below this design temperature, column pressure has to be low. Therefore, columns containing styrene are operated under reduced pressure.

Because cooIing in the condenser under 40°C is expensive, top temperatures should be above 40°C. So, the pressures in the top should not be too low. In order to keep the bottom pressures low enough, pressure drops in the columns should be smal!. Of all possible contacting devices (e.g. sieve plates, dumped packings (Pall rings and Ber! saddles) and

structured packings (Gempack, Max-pack and Mellapack)), the structured packing has the lowest pressure drop per theoretical tray. Although structured packings are more expensive in purchase than others, structured packings are chosen because of their low pressure drop. Mellapack 250Y is by far the most used structured packing and is therefore used in this conceptual design.

The boiling points of the Cs-chemicals are all in the same range so when the columns are operated under approximately the same pressure little intermediate heating or cooling is required. Therefore, to minimise the energy consumption in the columns where no styrene is present, these columns are also operated under reduced pressure. The pressures suggested in [l] are used in the design.

2_2 Summary of flowsheet

The flowsheet resulting from the considerations mentioned above can be found at the end of this chapter.

2.2.1 Reaction system

2.2.2 Separation system

The separation system consists of a knock out drum and four distillation columns in order to remove dissolved gasses from the reactor effluent and to obtain the separate components from the naphtha feed. As mentioned above, all columns are operated under reduced pressure. The required purities of product streams can be found in table 1.1.

In the knock out drum M 10, gasses dissolved in the liquid stream leaving the reactor are removed. A throttle valve is included in this apparatus.

As mentioned above, T 13 is an extractive distillation column using AEP as extractive agent. Without AEP, the separation of o-xylene and styrene is very difficult due to the close boiling points. Addition of AEP lowers the relative volality of styrene significantly. In other words, the boiling point of styrene is reduced by al most 20°C [1]. Because the bottom temperature of the column is kept below 140°C, little styrene is polymerised.

In the second column, T19, styrene is stripped from AEP. All other Cg-chemicals (mainly 0-xylene) leave this column over the top with styrene. Ergo, the purity requirement for styrene has to be achieved in the bottom flow of the first column. Some AEP (preferably 10-50 ppmw) is left in the styrene stream to prevent polymerisation during storage and transport between the hydrogenation and the polymerisation plant.

Removal of lights is not necessary, as lig hts that are not removed in the knock out drum will leave the separation system together with the ethylbenzene product stream. Removal of the formed heavies will be necessary to prevent their accumulation in the separation system. This

is done in the third column, T23. These heavies can be used as fllel.

In the fourth column, T31, the top stream of the first column the ethylbenzene is separated from the xylenes. This separation is very difficult due to small difference in boilingpoints of ethylbenzene and p-xylene.

AEP make-up ,tream ~~1.~O ________________________________ ~ ____________________________________________________________ - - ,

CB

FEED

r ' \ r ' \ r ' \

AEP recycle ,tream

,

r.1., ,R2,

,

,

,

LTJ

,

r.1., , R4 ,

,

,

,

,

,

,

,

,

, ,

LTJ

,

r.1., R6,

,

yJ

,

r.1., ,RB,

,

,

,

LyJ P12 Process flowsheet for production of

STYRENE from a CB fraction R.V.Goede B.J.Zuurdeeg FVO nr . .3192. January 1997 H20 -"HJJ , . ~ I -'""~7-' - -'""~7-' -,' HtS (-j-+,---_{, , I} - .. 4-'7-' - 4-'7-' Reactor purge

Knock out drum o"go,

HJ2 HJ4

Ethyibenzene

· 3. Calculations: flowsheet and units

3.1

Introduction

The reactor is modelled in Simulink (Matlab). All other apparatus are simulated with ChemCad lIL The properties of some streams are calculated with the ChernProp program. On the basis of the calculated flows and tower profiles, the dimensions of apparatus are calculated with prelimenary design methods. The heavies formed in the distillation columns are simulated as if they were formed in stream 17, i.e. the feed stream of the distillation section. The heavies are modelled as a dimer of styrene, C'6H'6' because, of all heavies this component is the most difficult to separate from AEP, for it has the closest boiling point. The reaction equation for this formation is:

2 Styrene ---> Distyrene (3.1)

There are no kinetic data available for this reaction; the formation of heavies is, according to [1], assumed to be 0.1 w% of the total styrene feed into the separation system.

In table 3.1 the feed and production streams are summarized including their ratio to the styrene production. The streams in and out can be found in the figure l.I.

Table 3.1: In and output process streams and relative consumption and production Flow number Description Massflow Massflow ratio

[kg/sJ [ tonlton,tyr<nJ lil 2 Cg-feed 1.4531 2.1 ., H₂-feed 0.0032 0.046 :) 1 AEP make-up 0.0001 0.000 Total 1.4564 out 13 offgas 0.0017 0.0025 16 reactor purge 0.0009 0.0013 42 styrene 0.6919 1.00 40 ethylbenzene 0.2641 0.38 41 xylene 0.4971 0.72 43 distyrene 0.0008 0.0012 Total 1.4565

3.2 Simulation

3.2.1 Reactor model and kinetics

Aim of modelling the monolith reactor

For the following reasons it is important to have a good model of the monolith reactor. First, this paper is written in order to be able to compare the benefits from a monolith reactor in this process to those of an, already existing, triclde bed reactor. Second, only with a good model it is possible to estimate reasonable conditions at which the reactor can be operated, allowing further design of the monolith reactor, the injectionpoints of hydrogen inc1uded. So rat her a lot of effort has been made building a good model and searching for the best operating conditions.

Assumptions of the model

In the monolith channels, a Taylorflow regime is assumed. This means gas bubbles and liquid plugs are moving alternatingly through the monolith channel. The liquid drop lets tridele down from a distribution section, entraining gas bubbles. The gas bubbles are assumed to be

bar-shaped with spherical ends. The frequency of downflowing droplets is assumed to be constant. The walls of the monolith are covered with a washcoat layer, in which 0.1 w% Pd is dispersed. At the surface of the washcoat a thin liquid-layer (the film) flows down very slowly.

At first, it is assumed th at concentrations in the film are equal to those in the washcoat. In reality, concentrations in the washcoat will differ from those in the film because of the reactions in the washcoat. This assumption only influences reaction rates and this effect is taken into account by introducing an effectiveness factor in the reaction rates.

Mass transfer from the gas to the film can be described with two mechanisms. First, the direct transfer between the gas and the film. Second, the transfer through the liquid plug to the film. The second mechanism is dominated by the resistance to mass transfer between the gas and the liquid phase, so the hydrogen concentration in the liquid phase equals this concentration in the film. To calculate the total mass transfer from the gas to the film the two streams of hydrogen resulting from these mechanisms are summed.

Other assumptions are:

-The values of the parameters are independent of temperature and composition changes. Due to the small changes of temperature and composition in the reactor, this assumption introduces only very small errors.

- The gas mixture behaves as an ideal gas and is ideally mixed in the reactor shell. The gas phase remains the same volume in spite of reaction of hydrogen and changes in temperature or pressure.

- The pressure drop due to friction in the gas phase is negligible compared to the pressure drop due to friction of the liquid phase.

- Gas and liquid streams are in a pseudo-steady state. The gas velocity is always equal to the liquid velocity.

- The reactor is adiabatic.

- Mass transfer of hydrocarbons from the liquid to the gas is neglected, because of their low saturation pressures.

- Between two monolith structures, there is a 0.5 m space, in which the hydrogen injection points (the term injection point should not be taken literally; meant is that fresh gas

(containing more hydrogen) is sucked into the next monolith structure), including the liquid distributers, are fitted.

Description of the model

The used model is based on equations given by Edvinsson [13]. Concentrations in the film and in the washcoat are assumed to be equal (where the effectiveness factor compensates for the concentration profile within the washcoat).

The concentrations of PH, ST, EB in the liquid and in the film are integrated with respect to the reactor length. The same was done with the concentration of Hl in the gas phase and in the film and with the temperature. The concentrations of PH, ST and EB in the liquid and the H₂ gas concentration changed because of the mass transfer to the washcoat. The concentrations of PH, ST, EB and Hl in the washcoat changed due to the hydrogenations and due to mass transfer from the liquid or the gasphase. The temperature changed due to convective heat transport and due to the hydrogenation reactions.

The most important equations of the model are given below. The justification of the parameters used in these equations is given in appendix 1.

The concentration profiles are calculated by integrating the following differential equations:

dCI,n_ k1ra*(CI,n-cr

,n)

dz -- ul*~ (3.2) where n

=

PH, ST or EB. k fa _g

*

(m

*

C _g, H -₂ C, _{I , 2} H ) (3.3) u _g *V _g

de

J, n _ Achannel

*k

_lJ

a

*

(c

I, n -c J, n) ~ashcoat '"

(v

n1

*r

1 + Vn }

"'r

2) dz - ____ -!-_ _ _{~m *U}-'----'-_..L.!..-'--'-_J + ---;---'---=----

_Arum

_*U J where n

=

PH, Vnt

=

-1, _Vn2

=

0, n = _{ST, Vn1}

=

1, _Vn2 = -1, n

=

EB, Vnt = _{0, Vn2} = l. dC f H Achannel *kgfa

*

_{(m*c g} H -Cf H) A

*

(r +r ) _ _ ' ~2 _ ' 2 ' 2 washcoat 1 2 dz Afilm*U_t Afilm*U f (3.4) (3.5)

The temperature profile

is

calculated by integrating the following differential equation with respect to the reactor length:

dt Awashcoat'" (rl'" -MI 1 + r 2 '" -MI 2 )

dz Achannel'" (u1 "'V1 "'P1 "'Cp,l+Ug",Vg*Pg"'Cp,g)

(3.6)

The activation energies for reactions (1) and (2) are found in Chaudhari et al. [12]. The other kinetic data for these reactions are taken from Mochizuki et al. [6]. The kinetics of reaction

(1) and (2) are assumed to follow a Langmuir Hinshelwood mechanism. The reaction rate equations are:

(3.7)

(3.8)

The pre-exponential factors of the Arrhenius equation for kt and k] are calculated using the kinetic data at 50°C from [6] and the activation energies from [12]. The reaction enthalpies are calculated with values found in [19].

Operating conditions

The inlet temperature of the monolith is the temperature of the feed, 50°C. In the monolith the temperature increases approximately 30°C, so the temperature always stays well below the polymerisation temperature of styrene (without AEP: 100°C). So, there is no need for intermediate cooling. Operating at higher temperatures is of no use, because this would only increase reaction rates, but these are not limiting. The convers ion of phenylacetylene must be 99.96%.

The reactor pressure is chosen to be 3 bar, comparable with the current trickle bed process. Because of the low hydrogen partial pressure in the reactor, hydrogen losses in the purge are very small at this pressure.

The monolith structure has 400 cpsi (cells per square inch) and the channels have a square-shape, both conform standard design. The washcoat has a thickness of 50 !lI11. Without catalyst deactivation, a thickness of 10 !lI11 would be sufficient. However, a washcoat thickness of 50 !lm is standard, is not much more expensive than a washcoat thickness of 10

!lI11 and remains operative, even af ter considerable catalyst deactivation, although this could reduce selectivity towards styrene a little, because of the larger transport limitation of phenylacetylene.

Simulation

The model is simulated using Simulink (Matlab ). Because of the simultaneous integration with respect to several variables, each simulation takes typically 10-90 minutes on a 486 DX-33 computer.

The simulations make clear that the selectivity towards styrene is best when the monolith is operated under low hydrogen pressure. The reason for this is that the mass transfer of phenylacetylene to the washcoat proceeds much slower than the mass transfer of styrene to the washcoat, due to its much lower concentration difference. To achieve good selectivity, hydrogen concentration must be low, to prevent too much hydrogenation of styrene. Due to the low hydrogen pressure, thus the low hydrogen concentration, hydrogen wil! be exhausted fast, so more hydrogen injection points are incorporated in the Simulink-model.

Due to the high required conversion of PH, the high velocity of the liquid and the low hydrogenation rate the monolith structure is quite long. Because the yearly throughput is smal! and the liquid velocity is rather high, the cross area of the monolith is also small. A very long and smal! monolith is therefore to be expected. So, the liquid is forced to pass the reactor several times (in different parts to prevent mixing). Af ter each pass the liquid is pumped up, giving rise to extra cost. Due to lack of design-time, temperature is not changed after each pass. When finally designing the monolith reactor intermediate cooling could prove to be profitable (because this would enable operation within a narrower temperature range, thus closer to the optimal temperature, which also has to be deterrnined then).

In short, the best operating conditions form a balance between high selectivity (low hydrogen pressure, many hydrogen injection points and high velocities) and short reactorlength (high hydrogen pressure, few hydrogen injection points and low velocities).

To find the conditions at which the reactor should be designed estimates of the costs of the monolith reactor plus its equipment are used. These estimates and their calculation can be found in appendix TI. The following parameters are optirnised: the injection concentration of hydrogen, the di stance between the hydrogen injection points, the liquid and gas velocity, the length of the liquid segments and the length of the gas bubbles.

- Injection concentration of hydrogen, c_H2.₀:

The advantage of increasing the hydrogen InJection concentration is the increase of the reaction rates, which leads to a shorter monolith. However, increasing the hydrogen concentration results in an increased hydrogenJliquid ratio. As explained earlier, this leads to a decrease of selectivity, due to the mass transfer lirnitation of phenylacetylene.

- The di stance between the hydrogen injectionpoints, ID:

A shorter distance between the injection points leads to a reduced length of the monolith structure. However, either more injection points are required (leading to a longer reactor) or the average hydrogen concentration should rise in order to achieve the same conversion of phenylacetylene (leading to a decreased selectivity). The optimal di stance between the hydrogen injection points is a balance between these effects. In practice, the length of a monolith structure is at most I m.

- InfIuence of the liquid and gas velocity , UI' Uc :

Increasing the velocity of the liquid (and of the gas) increases the Reynolds number. This leads to an improved mass transfer of the liquid phase, thus reducing the mass transfer lirnitation of phenylacety1ene. The se1ectivity towards the hydrogenation of pheny1acetylene is hereby improved. Disadvantage of an increased velocity is the resulting need for a longer and smaller monolith structure (which is, as exp1ained above, not desirabie as it complicates the design and results in more pumps).

- The length of the liquid segments and gasbubb1es, 11. 19:

An increase in the ratio of the 1ength of the liquid segments compared to the 1ength of the gas bubb1es leads to a better mass transfer of the liquids, leading to an improved selectivity. However, an increase in this ratio also means less hydrogen is injected for each unit of liquid, so more hydrogen injectionpoints are needed.

Optimisation

Integration is executed for each simulation until the ratio of phenylacetylene to styrene concentrations drops below

la

ppmw, giving the required length of the monolith structures. The results of the optimisation procedure can be found in tab Ie 3.2. The used optimisation parameters can be found in table 3.3. An explanation of the used optimisation technique can be found in appendix lIl.

Table 3.2: Results of optimisation procedure

Property Value

CPH.end 0.0379 [mole/m3 ]

CST.end 4059 [mole/m3

]

CEB.end 1542 [mole/m3]

Tend 80.3 [0C]

L 76 [m]

Selectivity (-~PH / L}EB) 30 [%]

Table 3.3: Optimal values of optimisation parameters

Optimisation parameter Optimal value

CH2.0 6.33 [mole/m 3 ] ID 1 [m] uI. ug 0.115

[mis]

11 6.47 '" 10. 3 _[m] 19 8.65 '" 10-3 [m]

3.2.2 Thermodynamics: Margules parameters

Calculation of the distillation profiles with use of UNIF AC parameters in the ChemCad program takes a lot of time, therefore Margules parameters were used.

These Margules parameters were obtained from Txy-data, generated by ChemCad with the UNIF AC parameters. These data files were converted by a simple selfwritten Pascal program to a file format which could be used in the Dortmund Data Bank (DDB) program. The DDB program calculates the Margules parameters which can be used in ChemCad as BIP's. The interactions between AEP and the ot her components, especially styrene, are not estimated correctly by this procedure. These values were set te typical values for nitrogenous components found in Dechema databank.

3.2.3 Distillation

ChemCad can estimate the location of the feed tray and the total number of trays needed for a separation on the basis of split factors of key components. Also from literature [1] some data is acquired, e.g. the height of a column, feed location and toplbottom temperatures. This information is needed to simulate the towers and generate tower profils.

The distillation towers are designed via a sizing procedure used in [20]. The tower is split in four pieces: top and bottom of rectification and top and bottom of stripping section. For each piece the mass flow rate for both liquid and vapour are extracted from the tower profile. The densities for each phase are calculated with ChemProp or taken from ChemCad. Then the flowparameter (F_1g) is calculated with:

F,

=111

~

g.

13.91

g

11

PI

Dependent on the type of contacting device, in this case Mellapack 250Y, the gas load coefficient (C_sp) IS determined with [20, figure 6.39]. Next the flooding gas velocity is

calculated with:

(3.10 )

The maximum gas velocity in the column is set to 80% of the flooding velocity. The

operating gas load (Fsp) with a gas velocity at 80% of the flooding velocity is calculated with:

Fsp=0.8ugmax{Pg (3.11)

The column diameter follows from:

(3.12 )

If the top and bottom diameter of a section differ the biggest one is chosen.

Ta estimate the height of the column, the height of a theoretical plate, the HETP, is needed, as is the number of trays (N). The HETP is estimated with [20, figure 6.40]. A typical value

for the HETP in vacuum distillation is 0.4 m. The height of the packed bed (h_sp) is calculated

with N times BETP. Each 15 theoretical trays the liquid flow has to be gathered and redistributed; the height of one redistributer is 1.5 m. This height is calculated with:

3.3 Apparatus design

3.3.1 Reactor

Operation of the reactor

The different numbers of the reactor (R2, R4, R6 and R8) are used to indicate the different series of monolith structures. The monolith structures have a totallength of 76 m. There are 76 injection points, where the liquid has to be redistributed over the monolith and where fresh gas, with a higher hydrogen concentration, enters the rnonolith channels. In contrast to the assumption used in appendix II that the reactor could be split into parts of 10 m, the monolith structure is split into 4 parts, each with a length of approximately 30 m. This number

is

estimated being the optimum between higher reactor cost (due to a higher reactor) and a more complicated process (more pumps). The liquid feed enters the reactor, is distributed and flows through the first series of monolith structures and hydrogen injection points. After the 19th monolith structure has been passed through, the liquid is pumped up and then passes through the second series, is pumped up again, passes through the third series, is pumped up again and finally it passes through the fourth series. All series are placed inside the same reactorsheil. The pumps are placed outside the reactorshell in order to be able to fix them more easily in case of problems. To avoid gas leaving the reactor at the liquid exits, a liquid column is maintained there. In this way, only some dissolved gas leaves the reactor with the liquid. gas inlet Figure 3.1: reactor

DO

DO

Top Vl ew gas outlet of the

The hydrogen / methane mixture supplies the hydrogen. A gas purge removes the methane and some hydrogen. To assure good mixing the hydrogenlmethane feed enters the reactor axially at the bottom while the purge is placed at the side of the top of the reactor. A top view of the reactor can be found in figure 3.1.

The monolith blocks have a square shape, because of its ease in use. The size of a block is 0.26 m

*

0.26 m

*

1 m. This means the minimum reactor diameter is 2

*

-/2

*

0.26

=

0.74 m. To provide some space for gas circulation and additional equipment, the reactor is designed to have a diameter of 1 m. The reactorsheIl has a thickness of 5* 10-3 _{m (calculations}

show that a much thinner wall would suffice, but this is the minimum advised wall thickness) [11, p 722]. . 11

~

_~

mm

I I I I I! j I f I I I I " • I I Figure Detail m 0 n 0 3 .2: of the 1 i t h structure 3.3.2 Pumps

In order to force the liquid (and the gas) to flow down in the monolith structure with the desired velocity the pressure at the end of each monolith structure has to be somewhat higher than 3 bar (in appendix I is calculated that this pressure should be: 3.02 bar). At the liquid distributers, the pressure should also be somewhat above 3 bar, to compensate for pressure losses at the orifices. So, the end of the monolith structure is connected with the liquid distributer, both operated at a pressure above 3 bar. The gas leaving the monolith structure flows through a pressure control valve, where the pressure is reduced to 3 bar. This is illustrated in figure 3.2

It is of great importance that the liquid distributers perforrn at optimal conditions, because small differences in flowrates in the monolith will set back the selectivity and may result in hot spots in the monolith channels. This can lead to polymerisation of the styrene resulting in blocked monolith channels. In this paper, they are not designed, because this would be to detailed for a preliminary design, but when designing them, this should be done with great care. Recent developments [29] indicate that even when a simple show er head ($ 10) is used a good liquid distribution can be achieved because of the formation of a layer of foam above the monolith structure. Further research is needed on this subject.

Centrifugal pumps are chosen in the entire design. Centrifugal pumps have the following advantages:

- They are relatively cheap and simple in construction. - No valves are used in centrifugal pumps.

- They can be coupled directly to an electric motor.

-No damage is done to them when the delivery line becomes blocked for a short time.

- They are smaller than other pumps of equal capacity and can therefore more easily be sealed (which is needed to prevent polymerisation of leaked styrene, blocking the pump).

PI raises the liquid feed to the entrance height of the monolith reactor (where the pressure should be 3.00 bar). P3,

PS

and P7 raise the reaction-intermediate ao-ain to the entrance heio-ht _{o 0} of the monolith reactor. The reaction intermediate contains both hydroo-en and methane in _{o ,} equilibrium with the liquid. To prevent cavitation, the streams entering P3 , PS and P7 flow down 3 m before entering the pumps, thus raising the suction pressure of the pumps satisfactory. P3, PS and P7 process essentially the same stream and therefore they have the same design.

P12 supplies the pressure needed to raise the hydrogenated and flashed stream to the extractÏve distillation column (TB; the pressure at the entrance of the column should be 9.1 kPa). The entrance of this column is at a height of 46 m. As the flash is placed 3 m above the ground the liquid flows down before entering the pump. Therefore, there will be no cavitation problems.

P 18 supplies the pressure needed to raise the AEP/Styrene-stream to the entrance of the stripper column (TI9). The height of the entrance in this distillation column is 8.5 m (3 m higher than needed; this leaves space, wruch wil! be needed for the next pump). The pressure at the entrance is 14 kPa. The stream entering trus pump is subcooled, so there will be no cavitation problems.

P22 pressurises the stream leaving T19, in order to flow to the entrance height of T23, 11.6 m (leaving 3 m free to avoid cavitation in P40). Before entering trus pump, to prevent cavitation, the stream flows down 3 m. The pressure at the entrance of T23 should be 15 kPa.

P26 pressurises the feed stream to 7 kPa at the height of the AEP entrance of the extractive distillation column (TB; 57.9 m [the entrance stream has a height of 13.6 m]). The stream is subcooled, so no cavitation problems are expected.

P30 supplies the pressure needed to raise the xylenes/ethylbenzene stream to the column which separates these (T31). Because, for easy maintenance, it is desirabie to place the pump

on the ground, the stream flows down first 63 m, before entering the pump. So, no cavitation problems are expected. The pressure at the entrance of the distillation column (which is located 73 m above the ground) is 26 kPa.

P38 brings the xylene product stream (stream 41) at 1 bar. The inlet stream is subcooled, so there are no cavitation problems expected.

P40 brings the heavies (stream 43) at 1 bar. Because of the danger of cavitation, the inlet stream flows down 3 m before entering the pump, as is already mentioned in the description of P22.

The specifications of the pumps are to be found in appendix IV and are summarised in table 3.4.

Table 3.4: Summary of the pumpspecifications

Pump: Capacity Psuction Pout Efficiency Wshaft

[m3/s] [bar] [bar] [-] [kW] PI 1.64

*

10-3 3.00 5.97 0.5 0.976 P3 PS 1.64

*

10-3 _{3.26 5_97 0_5 0_889} P7 P12 1.73

*

10-3 1.35 3.99 0.5 0.913 P18 6.76

*

10-3 0_17 1.01 0.5 1.136 P22 6.36

*

10-3 _{0.41 1.24 0.5 1.056} P26 5.85

*

10-3 1.00 4_29 0.5 3.849 P30 8.95

*

10-4 _{5.33 6.46 0.5 0.238} P38 5.84

*

10-4 0_39 1.10 0_5 0.083 P40 1.15

*

10-6 0.36 1.10 0.2 0_000

3_3_3 Heat exchangers and heat integration

During operation of a plant it is necessary that the streams have the right temperature before they enter a particLdar unit, e.g. a distillation column. A good method to achieve trus is the llse of heat exchangers_ For a good performance of the heat exchangers, it is required th at the temperature difference between both streams is larger lhan 10

0

c.

In this design use has been made of 16 heat exchangers, divided in coolers, condensers and partial condensers, using a cold water or hydrocarbon flow, and reboilers, fed with steam or hot hydrocarbons.

In general all heat exchangers can be fed with cooling water or steam. It is often interesting, from an econornical and environmental point of view, to combine the temperature requirements of two streams in order to reduce the costs of above mentioned energy carriers.

This can be done using a heat exchanger, wherein the heat is transferred between two process streams. Especially stearn is rather energy consurning, 50 it is profitable to put some effort

To make optimal use of the heat content of the streams the following streams were coupled:

Stream 19, containing AEP and styrene at t = 137.4 °C, must be cooled, while the heat is delivered to stream 33, containing ethylbenzene/xylene at t = 40

oe.

This is done in H16, using a floating-head heat exchanger, as this seems to be the best type for a temperature difference above 80°C [11, pS72].

The condenser inlet of T23, containing AEP at t

=

155.1 °C, is split into 3 streams. Two of them are used for the kettle reboilers HlS and H33, working at t

=

137.4 and 108.9 °C respectively. Although kettie reboilers are usually more expensive than thermosyphon reboilers, this type is chosen because of the low bottom pressures of the distillation columns [11, p655]. It is possible to build in a kettie reboiler within the distillation column. According to Coulson and Richardson [11, p671] it depends on the heat flux whether this is possible.

From a simple calculation, it follows th at the heat flux in the first and last column are 25 kW/m2 and 66 kW/m2 respectively. The diameters needed for the cooiing bundles are smaller than the diameters of the two columns, so these kettle reboilers can be built within the distillation columns.

The rest of the heat exchangers could not be integrated, so they have to be heated with steam or with a fired heater. The use of the reactor purge and the knock out drum offgas is not sufficient to cover the heat requirements of the kettle reboilers of T19 and T23. The heat content is only 0.13 MW, while the heat requirement exceeds 11 MW. The streams th at need cooling transfer their heat to cooIing water of 20°C, heating it up to 40°C:

The kettle reboiler of T19, H21, is fed with medium pressure steam (10 bar), superheated at 220

oe.

The kettle reboiler of T23, H2S, is fed with high pressure steam (40 bar), superheated at 410

oe.

The remainder of the top stream from T23 is condensed in a floating head condenser H24, because of the high temperature difference.

The other coolers are simple fixed-tube plate heat exchangers, for the temperature difference is relatively small.

For the design of the heat exchangers one formula is important:

(3.14 )

Q

is the heat duty calculated by Chemcad. The heat transfer coefficient, U, comes from [15, p486]. This coefficient is dependent on the streams through the apparatus and on the type of exchanger. Due to counter current operation the [.. T is the logarithmic temperature difference. But if one phase is evaporating and the other one is condensing the temperature difference is constant; this is the case with HlS and H33. The temperatures are determined by the way of process operation. The exchange area needed is resulting from equation (3.14).

All above mentioned heat exchangers are summarised in table 3.5. Table 3.5: Details of used heat exchangers

no. function

_Q

A U ~t type

[kW] [m2] [W/m20C] [aC]

H9 cooler 26.60 2.10 284 44.8 fixed-tube plate

H14 p. condenser 1773 355.1 170 29A fixed-tube plate

HLS reboiler 3114 123.9 1420 17.7 kettle'

H16 cooler/heater 76.60 2.59 284 104 floating head

H17 cooler 725.7 34.1 284 74.9 fixed-tube plate

H20 condenser 453.9 10A 8-7 J_ _51.1 fixed-tube plate

H21 reboiler 1446 25.8 1420 39.5 kettle ••

H24 condenser 2731 25.7 852 125 floating head

H25 reboiler 10860 65.1 1420 117 kettle"

H27 cooler 1206 59.7 284 71.1 fixed-tube plate

H29 condenser 322.2 18.7 852 20.2 fixed-tube plate

H"7 :J_ p. condenser 4942 671 170 43.3 fixed-tube plate

H ... _{:J:J reboiler 5044 76.9 1420 46.2} kettle'

H34 condenser 115.1 5.12 852 26A fixed-tube plate

H35 cooler 64.58 5.75 284 39.5 fixed-tube plate

H36 cooler 49.25 5.87 284 29.6 fixed-tube plate

...

bllllt-m mtegrated kettle reboller wlth AEP top stream from tower 2:J (also see text)

.. exceeding upper limit mentioned in [11]; cannot be built in

3.3A Knock out drum

The knock out drum is designed according to the design procedure in [11, p394-396]. Because

this procedure suggests that the drum would be extremely high, a larger diameter is chosen

(d

=

1 m) and the height is calculated accordingly (h

=

2.3 m). The drum is operated at l.1 bar (the inlet pressure is 3 bar) and at 70

oe.

3.3.5 Distillation towers

The dimensions of the distillation towers are given in table 3.6.

Table 3.6: Dimensions of distillation towers

Diameter Heigt of bed HETP Total height

[m] [m] [m] [m]

top bottom top bottom top bottom

Tl3 2.24 2.24 12.2 35.6 0.42 0.40 62.6

T19 1.11 1.64 2.3 2.5 0.45 0.36 10.8

T23 3.99 4.20 1.7 5.6 0.42 0.40 13.6

T31 2.89 2.60 59.6 56.6 0.40 0.38 149.0

The extraction 1Ower, T 13, may seem rather high, but this could be expected based on the used patent [1], where a height of 60 m was mentioned. Of course, in the final design, the height should be taken into account. The use of partial condensers was necessary because, with the thermodynamic model used, ChemCad could not cope with methane dissolved in the top stream. If, with a better thermodynamic model, a total condenser were 10 be used, the top

stream would be liquid, and would have a different temperature.

Towers T 19 and T23 are of reasonable proportions and are not expected to cause big problems when being built.

As was already known from literature and mentioned in chapter 1, tower 131 consists of a huge number of trays. This is inevitabie for the required separation. Due to the huge number of trays, the tower seems very high (149 m), but in reality the tower wil! be split into two, perhaps even three, parts, connected with a compressor and a pump to transport the two phases from the bottom of one part to the top of the other. However, for this preliminary design, the column is assumed to be 149 m high. It is emphasised though that the authors realise the column will probably not be built into one part. Therefore, although pump cost and compressor cost are not taken into account in the economics, an extra margin in the tower cost is inc1uded. For the same reasons mentioned for T13, this tower has a partial condenser. The same comment is applicable.

3.3.6 Valves

In the process there are several points, where a significant pressure drop is needed. This is accomplished using throttle-valves. The following table gives a short overview of the used throttle-valves:

Table 3.7: Overview valves

Valve number Pressure in Pressure out Mass flow Volume flow

[bar] [bar] [kg/sJ [10.3 _m3_/s] In out MlO 3.0 1.1 1.4556 1.73 1.73 MIl 3.0 1.1 0.0009 0.54 1.48 M28 1.3 1.0 5.5538 6.36 6.36 M37 1l.8 1.0 0.2641 0.31 0.31 M39 l.1 l.0 0.6919 0.78 0.78

MlO is an integated valve plus knock out drum. Liquid enters the valve first, where the pressure is reduced. Liquid and gasses are separated in the subseqllent knock out drum.

It is assllmed that the pressure drops over the valves have no significant effect on the temperature nor on the liquid density. This assumption introduces almost no error. The valves are controlled by simple PID-controllers.

4. Mass and heat balance

In appendix V the mass and heat balance can be found. The tot al mass and enthalpy of each stream and enthalpy changes in each unit are indicated. Thereby, the mass and enthalpy streams can be followed through the whole plant. In this way, the consistency of the calculated streams can be checked easily.

There are rather big height differences in the plant. The main purpose of most pumps is to raise liquids to the desired height. So, the power supplied by the pumps is converted into potential energy and in a very small increase of temperature. However, when flowing through the pipes, the liquids wil! transfer some heat to the surroundings. In order to keep the mass and heat balance sheet clear the enthalpy transfer to the surroundings is assumed to be equal to the increase of enthalpy caused by the pumps. In the mass and heat balance this is indicated by a continuous arrow leading to the pumps and a dashed arrow starting from the pumps, which both transport the same amount of enthalpy. So, the overall effect of the pumps is assumed to be zero. Though this is not entirely correct, the introduced error will be very smalI, because of the small energy requirements of the pumps.

For every unit, the mass balance is correct (differences only arise due to rounding off). The heat balance is not correct for the reactor, for all other units it is correct (differences only arise due to rounding off). The inconsistency of the heat balance for the reactor is caused by the use of different models and is only small (the enthalpy of the streams is evaluated using Chemcad, while the temperature difference of the streams in the reactor is calculated using the developed model for the monolith; the rise of the surrounding gas temperature is neglected in that model). .

When all mass streams entering and leaving the process are summed, the difference between the total mass stream in and out is only 0.1

gis,

while the total mass stream into the process is 196.3076 kg/s. This difference is caused by rounding off.

When all heat streams entering and leaving the process are summed, the difference between the total heat stream in and out is only 3.9 kW, while the total heat stream into the different process units is 38080.5 kW. A difference of 3.8 kW is introduced in the reactor (as explained above). The remaining difference, 0.1 kW, is caused by rounding off.

The consistency of mass and heat balances for all units and the very small differences in the total mass and heat streams indicate the mass and heat streams have been determined correctly. It should, however, be kept in mind that the heat balance is simplified by the assumption concerning the enthalpy of the pumps.

In appendix VI the composition, temperature, pressure, total mass and enthalpy of all streams can be found.

5. Process control

5.1 Introduction

A chemical plant cannot function properly without process contro!. All flows, levels, ternperatures and pressures within the plant can vary with process feed conditions, utility quality or even with the weather conditions. Without process control the conditions within the plant would drift from their design values causing economically unfavourable conditions and possibly dangerous situations. To avoid this, there are several feedback loops in a plant that keep the operating conditions close to the desired values. Other control mechanisms are also present in a large scale chemical plant, but this chapter emphasises feedback contro!. The control loops are divided into seven groups: flow, production level, level, pressure, heat transfer, temperature and quality contro!.

5.2 Flow control

All flows are controlled directly by valves. The setpoints of these flows are given by ot her control systems. One exception is the flow of stream 1 (from now on flow 1), that is determined by flows 28 and 31. In this case, flow 28 is set to meet the requirements of tower 13, and flow 31 is set to meet the requirements of tower 23. Installation of a buffer at the mixing point would increase the reliability of this control procedure. In the discussion of other con trol loops, direct control of all flows is assumed.

5.3 Production level control

The production level is controlled by flow 4. All other streams are adapted to the production

level by level, pressure and quality controls.

5.4 Level control

All levels are controlled by the first valve encountered by the bottom flow of the apparatus.