Production of benzene with the mobil selective toluene disproportionation process out of a naphtha cracker product stream

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,Ni

FVONr.

Preliminary Plant Design

Laboratory of Chemical Process Technology

Subject

Production of Benzene with the

Mobil Selective Toluene Disproportionation

Process out of a Naphtha Cracker Product Stream

Authors

A.J. Brinkerink (Alexander)

J.N. van Es (1os)

N.M. Hiemstra (Ninke)

M

.

Vos (Michiel)

Keywords

Toluene Disproportionation ZSM-5

Telephone

015-2122544

015-2132909

015-2613486

015-2142509

MSTDP

Date assignment

Date report

.

05-11-1996

11-02-1997

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FVO Nr.

Preliminary Plant Design

Laboratory of Chemical Process Technology

Subject

Production of Benzene with the

Mobi1 Se1ective To1uene Disproportionation

Process out of a Naphtha Cr:acker Product Stream

Authors

AJ. Brinkerink (Alexander)

J.N

.

van Es (Jos)

N.M. Hiemstra (Ninke)

M. Vos (Michiel)

Keywords

Toluene Disproportionation ZSM-5

Telephone

015-2122544

015-2132909

015-2613486

015-2142509

MSTDP

Date assignment

Date report

05-11-1996

11-02-1997

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FVO 3193 MSTDP

Summary

At Dow Chemical in Terneuzen athermal hydrodealkylation process (RDA) is used to transform a mixture of toluene, xylenes and ethyl benzene (TX) into more valuable benzene. The "Mobil Selective Toluene Disproportionation" process (MSTDP) is an alternative process to produce benzene and mainly para-xylene out ofthe TX feed.

A preliminary plant design is conducted to see whether the MSTD process is an economically feasible altemative for the present HDA process.The preliminary plant design is based on a feed stream of 450,000 tonnes TX per year.

The MSTD process has a 30 % toluene single pass conversion. The para-xylene selectivity of the catalyst, which is ZSM-5, equals 82 %. This selectivity is reached by a selectivation step preceeding the actual process. This step is performed during 112 hours at 839 K and 3 bar, 6.5 WHSV and 0.5 mol hydrogen per mol hydrocarbon.

The reactor is a downflow fixed bed reactor and operates at 728 K and 35 bar. The weight hourly space velocity equals 4 kg toluene per kg ZSM-5 per hour. In order to reduce coking the reactor is operated under hydrogen pressure in a ratio of 3 mol hydrogen per mol toluene.

Under reaction conditions present alkanes and also some toluene are cracked into methane, ethane, propane and n-pentane. Rydrogen is recyc1ed. A part of this gas stream is purged to prevent accumulation of impurities. In a stabilizer column the light ends are separated from the product stream. Benzene is obtained at a purity of 99.98 weight% and therefore satisfies the specifications provided by DOW Chemical. The toluene is separated from the TX feed before entering the reactor. In a separation section n-octane, ethyl benzene, meta-/para-xylene and heavies/ortho-xylene are obtained as product streams. Ethyl benzene has a purity of 99.58 weight% and therefore meets the given specifications.

After 500 operating days the catalyst is regenerated in situ in order to re move depositedcokes. During 68 hours hydrogen is passed over the catalyst at 811 K and 35 bar.

A hazard and operability study (RAZOP) is performed to investigate the process safety. Based on the results of this analysis the process control system of the plant was set up.

The economical feasibility of the process was investigated. The designed process appeared to be unprofitable and therefore not competitive with the existing HDA process. The reasons are found in the low margin between feed and product prices and the high investment costs.

Savings in the investment costs can be realized by using sieve tray columns instead of columns with structured packings and by taking over equipment of an existing RDA plant. Finally an increase of the toluene conversion per pass, by finding an improved catalyst or process altemative, will be very rewarding.

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FVO 3193 MSTDP

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6.3 Fired heater and reactor ... ... 31

6.4 Purge stream, hydrogen-feed stream ... ... ... .. ... .... .. 31

6.5 Flasher ... ... ... ... ... ... .... 32

6.6 Towers ... 32

7. Economy ... 33

7.1 Cost estimations ... 33

7.2 Production volume dependent operating costs ... ... .... 34

7.3 Direct operating labour ... ... 34

7.4 Investrnent costs ... 35

7.5 Purchase costs of installed equiprnent ... 35

7.6 Sales income .. .. ... 36

7.7 Net Profit ... ... ... 37

7.8 Evaluation of analysis of process economics . ... ... ... 37

7.9 Profitability analysis ... .. '.' ... 37

7.10 Possible profit contributors ... . ... ... ... ... 37

7.10.1 Sales incorne .... ... ... ... .. ... ... ... . 38

7.10.2 Direct operating labour ... 38

7.10.3 Investrnent costs ... .. .. ... 38

7.10.4 Production volume dependent costs ... ... 38

8. Conclusions and recommendations ... ... ... .. ... 39

8.1 Conclusions ... ... .. ... ... .... ... ... ... ... . 39

8.2 Recommendations .... ... ... .... ... ... ... 40

9. List of symbols ... .. ... ... ... ... ... .. ... ... 41

10. Literature ... 45

11. Acknowledgements ... ... ... ... .. .. ... ... .. 47

Appendix A: Original assignment B: Component list C: Flowsheet

D: Mass and heat balance

E: Stream and component balance F: Heat of reaction calculations G: Letters from Delaval and Stork H: I: J: Azeotrope Economy calculations Stabilizer calculations K: Equipment lists L: Specification forms

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1.Introduction

Since Faraday discovered benzene in 1825, more and more applications have been found for this special molecule and its derivatives. The discovery of the combustion engine increased the demand for fuel and increased, as aresult, the study for combustible fluids and gases. Benzene and its derivatives appeared to have high octane numbers and appeared to he very weil combustible. Unfortunately about one and a half a century after Faradays discovery another one was made. Benzene and benzene derivatives are carcenogenic and are very harmful for man and his environment. Simple combustion was not allowed anymore and other ways of processing of benzene derivatives were demanded. Since their use goes beyond combustion alone base chemicals like benzene and xylene are made out of the benzene derivatives right now.

At DOW Chemical in Terneuzen, the Netherland a mixture of toluene, xylenes and ethyl benzene is transformed into more valuable benzene with athermal hydrodealkylation process (HDA). This process is conducted at high temperature and pressure. With an excess of hydrogen, a high conversion and selectivity is reached. The "Mobil Selective Toluene Disproportionation" process (MSTDP) is an alternative process to produce benzene out of a mixture of toluene, xylenes and ethyl benzene (TX). Valuable para-xylene is produced as weIl and in much higher amounts than the less valuable meta- and ortho-xylene.

Both benzene and para-xylene are important base chemicals for the chemical industry and the market demand for these bulk chemicals is larger than the demand for toluene. It is expected that the demand for and the prices of benzene and para-xylene will increase even more in the future.

A Preliminary Plant Design (PPD) is conducted to see whether the MSTD process is an economically feasible alternative for the present HDA process. The original assignment cao be found in appendix A. During the PPD the hierarchical approach given by Douglas [Et. 8] is

followed. At first, mass balaoces of product and feed streams are made. Then a process design of the plant is made with the aid of a flowsheeting computer programme. When the total mass and energy balances are calculated, the equipment calculations can be done in order to size the equipment. An economical analysis is conducted to calculate the investment, the net profit, the Return On Investment (ROl) and the Internal Rate of Return (IRR). This analysis is also used to compare the MSTD process with the HDA process from Dow Chemical. Finally a health, safety and environment (SHE) study is performed.

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2.Base of Design

2.1 Block scheme

The process block scheme, given in figure 2.1, is made using the general design heuristics given by Douglas (lit. 8] and using the data on the MSTD process, described by Gorra [lit 9] and in Hydrocarbon Processing (lit. 32].

HYDROGEN FEED

i

c : -' '_EE_D-r----+,,_ I

TOLUEN~

I

i

l '

I

COLUMN I 1

i

~

T

I

I . I " IIfYO&OGON .~a.a

r

I 1 I REACTOR ~

I

v i i PUIlIFI·

,

0 i CATION i V' " RECYCLE STREAM

Figure 2.1 Block scheme of MSTD process

,.

1

1,/ PURGE

Cl

LIG HTS

I

DUCT BENZENE SEPA·

l

IlA.

TION i T I _I N·OCTANE i ! ETHYLBENZENE M·.P·XYLENES. HEAVIES

Because of the low conversion of toluene, only 30% per pass, we chose to use a recycle of the product stream. In the toluene column the feed stream is purified. Only toluene is to enter the reactor. Also benzene, n-heptane and methyl cyclohexane leave the column at the top. The feed stream consists of the same chemicals as the product stream from the reactor. Separating toluene from the product stream and sending it back to the reactor would have lead to a structure in which the separated toluene would be mixed up again with the heavily imp ure feed. The purifying step before toluene enters the reactor and the separation section af ter the reactor will have the same units. Integrating these units seemed a logical decision. Now feed and reactor effluent are purified together and a rather pure toluene feed enters the reactor.

In the reactor, the present linear alkanes are cracked into lighter products: metane, ethane, propane and n-pentane. The same wiU happen to the methyl cyclohexane. The main reaction is the disproportionation of toluene into xylenes and benzene. Hydrogen is added a ratio of 3 mol hydrogen per mol toluene to prevent coking in the reactor. It is also a cracking reactant. A small hydrogen consumption takes place.

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FVO 3193 MSTDP

This stream is partially purged to the hydrogen purification plant. The rest of this stream is mixed with the hydrogen feed stream and then sent back to the reactor. In the separation section lights are also removed and sent to the naphtha cracker. Benzene is distilled off and the rest is recycled to the toluene column. The advantage of putting the benzene separating column in the recycle is that a smaller stream has to be processed.

The bottom stream of the toluene column is sent to the purification section. In this section, three columns split the heavies into n-octane, ethyl benzene, para- and meta-xylene and a rest stream of heavies. In this last stream ortho-xylene and n-nonane are to be found. Because of an azeotrope this steam is hard to purify. Ethyl benzene can be sold to a styrene plant, because of its purity. In a crystallizer plant the and meta-xy lene stream can be purified and para-xylene can he obtained as a pure product.

2.2 Reaction

The main reaction in this process is the toluene disproportionation (figure 2.2). Due to the catalyst the reaction rate increases under less severe conditions and benzene and mainly para-xylene is produced.

+

Figure 2.2 Disproportionatioll reaction

Also ortho- and meta-xylene are produced, figure 2.2 only shows the para-xylene. The feed stream to the reactor also consists of some other components. These are methyl cyclohexane and n-heptane and n-octane. They are cracked in the reactor, due to the excess hydrogen. These cracking reactions are shown in figure 2.3-.5.

6

--0

Figure 2.3 Methyl cyclohexane cracking reaction

3 +

Figure 2.4 n-Heptane cracking reaction

6~ ----+. 28 C H₄

Figure 2.5 n-Octane cracking reaction

The main reaction is not performed 100%. There is a 30% toluene conversion. The other reactions take place completely. Methyl cyclohexane, n-heptane and n-octane are fully converted.

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MSTDP

The MSTD process is catalysed by ZSM-5. This catalyst is a fully regenerabie zeolite on porous spherical silica partic1es. According to Absil [lit 1], the process may be practiced over a range of conditions such as a temperature of 533 to 923 K, preferably 588 to 753 K, a pressure of about 15 to 70 bar, preferably 29 to 56 bar and a weight hourly space velocity of 0.5 to 20, preferably 4 to 10 kg/kg. The weight hourly space velocity is the weight of liquid flowing through the reactor every hour divided by the weight of zeolite in the catalyst. The silica to alumina ratio is less than 30 and the weight ratio of zeolite to the total composition is 0.5-0.8 kglkg.

2.3 General process data

The feed, 450 ktonne/year coming from the naphtha cracker, has the composition as given in table 2.1.

Table 2.1: composition of the feed strea m, weight%

component weight % benzene 1 n-heptane 4 methyl cyc10hexane 1 toluene 56 n-octane 3 ethy 1 benzene 16 para-xylene 2 rneta-xy lene 5 ortho-xy Ie ne 9 n-nonane 2 isopropy I benzene 1

A list of properties of the chemicals rnentioned in table 2.1 can be found in appendix B. The benzene must have a purity of at least 99.9 weight%, containing less than 0.05 weight% toluene and less than 0.10 weight% non-arornatics. The demanded purity of ethyl benzene is more than 99.3 weight%. It may contain up to 0.15 weight% rneta- and para-xylene and 170 ppm ortho-xylene.

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2.4 Utilities

The plant will be situated at the location of DOW Chemical,Temeuzen, the Netherlands. The available utilities are sea water and ethyl ene for cooling (respectively to 303 and 173 K), and steam at various pressures for heating up to 498 K.

Hydrogen is delivered by the hydrogen-net at 35 bar.

TabeL 2 2 UtiLities

Utility fuction min/max. T

(K)

Air cooling 318

sea water cooling 303

ethylene coaling 173

steam, 4.5 bar, 433 K heating 408

steam, 12 bar, 483 K heating 448

steam,32 bar, 533 K heating 498

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3. Process Structure

3.1. Process conditions and justifications

The reactor is a downflow fixed bed reactor. The reactor is filled with five different layers of particles. The upper layer consists of inert cerarnic balls, which serve to distribute the gas flow over the entire diameter of the reactor. The second layer consists of the actual catalyst, which is a fully regenerabie ZSM-5 zeolite on porous spherical silica particles. The third, fourth and fifth layers con sist of inert cerarnic balls of increasing diameter, which serve to support the catalyst layer. The advantage of this kind of support is the low pressure drop, compared to other supports. Because the reaction is exothermic, the reactor has to be cooled. Therefore the segment of the reactor which contains tbe catalyst, is surrounded by a cooling jacket. The cooling medium is sea water.

In order to increase the para-xylene selectivity of fresh catalyst, a selectivation step is

executed during 37 hours before the process is started. During the process the reactor operates at a temperature of 728 K and a pressure of 35 bar. A weight hourly space velocity (WHSV) of 4 kg liquid/kg zeolite is used. The above mentioned process conditions are based on the conditions in an actual MSTD process at an Italian refinery [lit. 9].

In the reactor hydrogen is added a ratio of 3 mol hydrogen per mol toluene to prevent excessive coke deposition in the reactor. It is also a cracking reactant and therefore a small hydrogen consumption takes place. Because some coke deposition still occurs, the catalyst has to be regenerated af ter 500 days [lit. 9] operating time.

The separation of hydrogen from the reactor exit stream takes place in a flash vessel. A horizontal flash drum is chosen because of the high liquid to gas mass flow ratio. In order to obtain an efficient separation, the pressure needs to be high and the temperature has to be low. The pressure is therefore chosen as high as possible and is set at 33 bar. Because

cryogenic cooling is much more expensive than cooling with sea water, the temperature in the flash drum is set at 303 K. This is the minimum temperature which can be achieved by

cooling with sea water.

The separation of tbe light ends from the product stream takes place in a sieve tray column, the stabilizer. Because the magnitudes of the strearns in the rectification and the stripping section differ a lot, the column has two different diameters. The column operates under a top pressure of 25 bar and the feed temperature is 480 K. The bottom temperature, which is 550 K, cannot be achieved by steam heating. Therefore the reboiler ofthe column is a fumace. All the other distillation columns are equipped with structured pakings, because of its higher separating efficiency than random packings or sieve trays. This gain in efficiency justifies the higher investment costs. In the end this is expected to be more profitable than siev trays. The benzene is separated as a product stream in a distillation column with random packing. The top pressure in this column is 3 bar and the feed temperature equals 410 K. The toluene

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is separated in a distillation column with structured packing, operating at a top pressure of 1 bar. The feed temperature is 387.7 K.

In the ethyl benzene/xylene-column the ethyl benzene and n-octane are separated from the xylenes and heavies. Because the separation between ethyl benzene and para-xylene is very difficuit, the column operates under vacuum in order to enlarge the relative volatility of these two components. The column is filled with structured packing and operates under a top pressure of 0.13 bar. The feed enters at 348 K. The difficult separation results in a very large number of trays (372). Because the column would be too high when it would be built as one tower, it consists oftwo separate columns, which are set in series [lit. 27].

N-octane and ethyl benzene are separated in a distillation column with structured packing, which operates under a top pressure of 1 bar. The temperature of the feed is 403 K. As it is a valuable byproduct, the ethyl benzene strearn from the column can be sold.

Para- and meta-xylene are separated from ortho-xylene in a distillation column with

structured packing, operating under a top pressure of 1 bar. The temperature of the feed is 413 K. The para- and meta-xylene mixture can be separated in a crystallization plant if desired. The ortho-xylene containing bottom strearn from the above mentioned column could be separated into an ortho-xylene stream and a heavies strearn. However azeotropic distillation would be necessary, because o-xylene and n-nonane have an azeotrope.

3.2. Process Flow Diagram

The process flow diagram can be found in appendix C and is described here.

The TX feed (stream 1), which is coming from a naphtha cracker, enters the process at a pressure of 12 bar. In an expansion valve the pressure is lowered to 1 bar and the TX feed is mixed with the bottom stream (stream 27) ofthe benzene column (T22). This stream also contains toluene and xylenes. The resulting TX strearn (stream 3) is sent to the toluene column (Tl). The bottom strearn from this column (stream 28) consists of n-octane, ethyl benzene, xylenes and heavies. It is expanded to 0.13 bar in an expansion valve and partially condensed and cooled to a temperature of 348 K in a cooler (H27). The top strearn consists mainly of toluene and it is sent to a pump (P5) where the pressure is elevated to 35 bar. The pump outlet stream (strearn 5) is sent to the reaction section.

Fresh hydrogen feed (stream 6), which comes from a Pressure Swing Absorption (PSA) unit, and recycle hydrogen (strearn 16) are added to stream 5 and the resulting stream is heated in a

feed-effluent heat exchanger (H7) and then heated to a temperature of 728 K in a fumace (F9). The furnace exit stream (stream 9) is sent to the reactor (RIO), where toluene is disproportionated to benzene and mainly para-xylene. The present n-heptane, n-octane and methyl cyclohexane and also some toluene are cracked and light ends are produced. The reactor outlet strearn is cooled in the feed-effluent heat exchanger (H7) and further cooled to a

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temperature of 303 K in a cooler (H11). In the flash vessel (V12) hydrogen is separated from the product stream and recycled to the reactor. Because the hydrogen recycle stream (stream 13) contains impurities, such as methane, ethane and propane, part of this stream is purged to prevent build up of these impurities.The purge stream (stream 14) is sent to the PSA unit. The rest of the hydrogen recycle stream (stream 15) is sent to a compressor (C8), where the

pressure is elevated to 35 bar, and then is mixed with stream 5 and 6.

The bottom stream of the flash vessel (stream 17) is expanded to a pressure of 25 bar in an expansion valve (E13) and heated to a temperature of 480 K in a heater (H14). Then the stream enters the stabilizer (TI5), where the light ends (stream 20), which were forrned in the reactor, are stripped off from the product stream. These light ends are sent to a naphtha cracker. The bottom stream (stream 21) is expanded to a pressure of 3 bar and cooled by stream 37 in a heat exchanger (H20). It is further cooled to a temperature of 410 K in a cooler (H21).

This stream (stream 24) enters the ben ze ne column (T22), where benzene is recovered as a product stream (stream 25). This stream is sold. The bottom stream (stream 26) is recycled to the reaction section by fust expanding it to a pressure of 1 bar and then mixing it with the TX feed (stream 2).

The outlet stream of cooler H27 (stream 30) enters the 2-stage column (T28a and T28b). The top stream of this column (stream 31) consists mainly of n-octane and ethyl benzene. The bottom stream (stream 36) contains the xylenes and the heavies.

The pressure of stream 31 is elevated to 1 bar in a pump (P37). Then the temperature is increased to 403 K in a heater (H38). The resulting stream (stream 33) is sent to the ethyl benzene column (T39), where ethyl benzene is obtained as a product stream from the bottom

(stream 35). This stream is sold. The top stream of the column consists mainly of impure n-octane (stream 34), which can be sent to a naphtha cracker, a catalytic refonner or a FCC plant.

The pressure of stream 36 is elevated to 1 bar in a pump (P44). Then stream 37 is heated by stream 22 in a heat exchanger (H20). Stream 38 enters the xylene column (T45). The top stream (stream 39) consists mainly of meta- and para-xylene and can be sent to a

crystallization plant in order to obtain pure para-xylene. The bottom stream (stream 40) contains ortho-xylene and heavies and is not further processed.

3.3. Thermodynamics

3.3.1 K-Value

The Soave-Redlich-Kwong (SRK) [lit. 17] equation of state is chosen for predicting the K-values of the hydrocarbon system. The SRK-equation is very effective for hydrocarbon systems at medium to high pressures and temperatures.

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3.3.2 Enthalpy model

For the same reasons mentioned in 3.3.1, SRK is chosen for predicting the enthalpy ofthe system.

3.3.3 Phase equilibria

The thermodynarnic phase equilibria of the components were appropriate to perform the separations under normal conditions (e.g. distillations, flash).

Only the phase equilibrium of ortho-xylene and n-nonane at 1.5 bar and 430 K shows a azeotrope. The Txy-plot is shown in apppendix H. The separation of ortho-xylene and n-nonane is beyond the scope of the prelirninary plant design. Thermodynarnic properties of the components are shown in appendix B.

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4. Calculations

4.1 Catalyst and reactor

4.1.1 Catalyst

The used catalyst is a shape-selective ZSM-5 catalyst, developed by Mobil. The catalyst is composed of 80 weight% ZSM-5 and 20 weight% binder material. These two materials are randomly mixed. The ZSM-5 has a silicalalumina ratio of 30. The specific surface area of the ZSM-5 catalyst is 500 m2/g. Silica is chosen as a binder material. The properties ofthe ZSM-5 catalyst are:

- relatively high stability of framework; - high stability in acids;

- low stability in bases;

- low concentration of acid groups; - high acid strength of acid groups; - hydrophobic.

In order to enhance the shape selectivity of the catalyst, a selectivating pretreatment step is necessary. During this step a thin layer of coke is deposited on the inside of the porous channels. This coke selectivation enhances para-selectivity by "molecular traffic control" . The smaller para-xylene isomer, as weIl as the benzene co-product, can exit the pores many times faster than the larger meta- and ortho-isomers. Depletion of the para-isomer in this manner causes the xylene mixture within the pore system to re-equilibrate, forming more para-xy lene.

Selectivation is accomplished using toluene feed, but at reactor conditions that differ

significantly from norm al operating conditions. Selectivation can be obtained by treating the catalyst with toluene for 112 hours at 839 K , 6.5 WHSV, 0.5 hydrogenJtoluene and 3 bar [lit. 12]. These conditions hold for a catalyst composition of 65 % ZSM-5 and 35 % alumina. At normal operating conditions coke formation will also occur, because of the high operating temperature and the acidity of the catalyst. Because this coke formation is only moderate, regeneration of the catalyst is necessary after 500 days [lit. 9].

4.1.2 Kinetics

The primary mechanism is a vapour phase benzylic chain reaction. This mechanism is shown in figure 4.1. There are two possible mechanisms, but according to Xiong [lit. 30] the one in figure 4.1 is favourable.

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Figure 4.1 reaction mechanism

The reaction rate can be described by the following equation [lit. 2]:

-r=k'.p (4.1)

Using the ideal gas law and assuming a density for the catalyst-bed of 1000 kg/m), k' can be easily converted to the pseudo-frrst-order rate constant k in hol units. The resulting equation is:

PElud r = k

-RTPbed101 ,325

(4.2)

The main reaction is exothermic. The heat of reaction equals -17.13 kJ/mol, see appendix F. In the reactor also side reactions take place. These side reactions are the cracking of

n-heptane, n-octane, n-nonane and methyl cyclohexane under consumption of hydrogen. The

conversion is assumed to be 100%. The products formed are methane (80 volume%) and ethane (20 volume%).

4.1.3 Reactor

A downflow fixed catalytic bed reactor is chosen, which is operated isothermally [lit. 1]. The reactor is filled with particles of several sizes (see figure 4.2). The layers A, C, D and E consist of spherical partic1es made of ceramic clay material. Layer B consists of the actual catalyst material. The sizes of the particles in the different layers are given in table 4.1. The relative volumes of the layers are estimated from the picture of the reactor in figure 4.2. The flowpattem in the reactor is the following: in the upper part the flow is in reverse V-form, in the middle part the flow approximates plug flow and in the lower part the flow is in V-form. Therefore, in order to use the catalyst efficiently, the catalyst has to be located in the middle part of the reactor. The function of layer A is to distribute the gas flow over the complete reactor width. The function of layers C, D and E is supporting the catalyst bed. The advantage of this form of support is the low pressure drop and the use of catalyst only on places where it is maximally active.

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Figure 4.2 Reactor design

The reactor is surrounded by a cooling jacket and cooled with sea water. The construction material of the reactor is Austenitic stainless steel, type AISI 304. This material is chosen because of the following properties:

- resistance against hydrogen embrittlement; - high temperature oxidation resistance; - high mechanical strength.

The cooling jacket is filled with sea water. Stainless steel, type 304, has no resistance against sea water corrosion and it is therefore inappropriate. The jacket is constructed of Mone!, because it is not susceptible to stress corrosion cracking in chloride solutions and is very resistant against salt solutions.

4.1.4 Reaction conditions

The reactor operates isotherrnally. The reaction takes place at a temperature of 728 K and a pressure of 35 bar. To preserve the catalyst from coking excess hydrogen is added at a 3 mol hydrogen/rnol toluene ratio. The process is continuo us and the weight hourly space velocity (WHSV) equals 4 kg liquid per kg zeolite per hour. The toluene conversion is 30 % and the para-xylene selectivity is 82 %.

4.1.5 Regeneration

In situ regeneration by hydrogen is suggested. This is accomplished by passing hydrogen at 811 K and 29 bar for a total of 68 hours [lit.12]. In the mentioned article the pressure of the regeneration equals the reaction pressure. In our case regeneration is suggested at 35 bar. Generally regeneration consists of two steps. In the flTst step low boiling hydrocarbons at moderate temperature are removed. The second step is performed at high temperature to remove coke by reaction with hydrogen. The time integration of pretreatment (selectivation), reaction and regeneration is schernatically represented in figure 4.3.

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--' S . . .

•

Figure 4.3 Cycle diagram

4.1.6 Calculations Reactor dimensions:

MSTDP

I I

_se _

TI ... (DoIp) _:>

The dimensions of the reactor are calculated as follows. The reactor vessel is assumed to be a cylinder. Starting with the eatalyst density, the given WHSV, the mass flow of the toluene feed and the catalyst composition, the necessary catalyst volume is calculated. With the bed porosity and the volume ratio of the different layers in the reactor, the volumes of these layers are caleulated. The bed porosity of the layers is assumed to be 1.25 times the porosity of a closest paeking of spheres. With the LID-ratio of the catalyst bed and the calculated catalyst bed volume the length and diameter of the catalyst bed are calculated. The lengths of layers A, C, D and E are obtained from their volumes and the reactor diameter (which equals the diameter of the catalyst bed). The calculated dimensions are given in table 4.1.

used data: WHSV= 4 kg toluene/(kg cat. h) Pc~t = 1780 kg/m3 <Ptoluene = 103151 kg toluenelh LID = 0.85 Pressure drop:

The pressure drop over the fixed bed is estimated with a modified form of the Kozeny-Carman equation [lit. 4], which is valid for streamline and turbulent flow and is given below. With these equations the total pressure drop over the different beds in the reactor is

caleulated. The necessary data on viscosity, density and volumetrie flow are taken from the flowsheet ealculations in ChemCAD.

u e p Re

=---=-.

.

I e s(l-e) 11 (4.3) L

_pu;

.,

. (4.4)

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R _ _ l, =5'(Re 1

r

l _+OA'(Relro_.₁ (4.5) pu 1

-The total pressure drop over the reactor, the reactor diameter and the total bed length are calculated at:

ilP = 0.737 bar Dreactor= 3A3 m

total reactor bed length = 7.28 m

Table 4 J Reactor dimensions

layer dpartic:le (m) volume ratio (-) bed length (m) ilP (bar)

A 0.0191 1 0.971 0.01146 B 0.00159 3 2.912 0.66480 C 0.00635 0.5 0.485 0.02016 D 0.0127 1 0.971 0.01814 E 0.0191 2 1.941 0.02292 total

-

- 7.28 0.737

4.1.7 Heat properties of reactor

By sirnulating an adiabatic reactor in ChemCAD the adiabatic temperature rise was

determined at 28 degrees. For isotherrnal operation this temperature rise is too high. In order to achieve isothermal operation the reactor needs cooling. This is obtained by surrounding the reactor with a cooling jacket around the heat producing catalyst section. The cooling medium is sea water. With this design it is important to ensure that the radial heat profile is

minimized. The Reynolds number of the reactant medium was calculated as Re=364. The resulting turbulence promotes radial heat transfer.

4.2 Tower design

The design is done for the stabilizer, equipment number T15. 4.2.1 Type of column

First the type of column is chosen. There are mainly three types of columns. A plate column, a structured packed column and a random packed column. For the stabilizer, the decision to use a plate column has been made upon the following aspects [lit. 5].

1. Plate columns are easier to clean than packed columns. Since coking is pos si bie with these components, there is always a chance that column internals become dirty.

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FVO 3193 MSTDP

easy to install, because the packing is placed into the column at the top.

3. Plate columns have a higher efficiency than packed columns when the flow rates of gas and liquid have a wider range.

In this design sieve trays are taken as plates in the stabilizer.

4.2.2 Tower diameter

When all the stream compositions are known, the design calculation can be done. At flTst the column diameter is calculated with the method described in Coulson vol 6[lit. 5]. This is done by calculating the column area.

1 d ?

-1t -=A

4 col col (4.6)

This column area is calculated from the net column area, the area needed for separation. The net column area is the area without the downcomers, which is ca1culated from the vapour velocity and flow in the column.

<I>

A (1- ) =A max.vap

col perc net Um;:u.,vap (4.7)

The maximum vapour velocity is based up on the mass flow of vapour and liquid in the column and upon their densities.

u _max=K (4.8)

This K-value depends on the mass flow of vapour and liquid and, empirically on the tray spacing and the flooding percentage of the column ..

(4.9)

With these relations the diameter of the column and the most important column areas can be calculated. This is do ne for the stripping and the rectification section. When these diameters are not in the range of 150% of each other, a design for a two-diameter column can be chosen for. The mentioned criterion applies to the stabilizer column.

4.2.3 Tower internals

After this is done, a plate design has to be made. The hole area has to be chosen, the weir height, the hole diameter, and the plate thickness. These are all parameters that determine the weeping rate, the pressure drop per plate, the downcomer backup and the entrainment. Weeping check

The minimal vapour velocity must be that high, that the vapour is capable of flowing through the holes in the plate or, to explain it otherwise, the liquid must flow over the plate to the downcomer and not through the holes.

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FVO 3193 MSTDP

The minimum vapour veloeity required is calculated as follows.

K? -O.9(25.4-d, _ 10 I) e

Ll =

vap,req (4.10)

The K₂-value depends on the weir height and the height of liquid on the plate. This height is calculated with equation (4.11)

<P ~

h =750(~)3

ow 1

PJiq w

(4.11)

When the minimum liquid veloeity in the column does not exceed the required veloeity,

weeping oecurs and the plate design has to be changed. Pressure drop

The pressure drop eonsists of several parts. The total pressure drop is deterrnined with equation (4.12) and is expressed in a liquid height.

h =h +h +h _tol _r _d _wrlr. +h _oW,m:u (4.12)

The values for ~ _{and hd are the contributions of the dry plate pressure (hd) and of the effe cts}

of froth on the plate (~). The dry plate pressure depends on the hole pitch. This hole pitch is best chosen between 2 to 5 times the hole diameter. The ratio of hole area and this pitch gives a certain Co-factor, whieh is used to ealculate hd'

Ll

P

h

d=51( m p)2 var Co P[jq

The hr-value is calculated from equation (4.14).

h = 12500

r

(4.13)

(4.14)

Then the pressure drop per plate is calculated by multiplying the h(O( with the liquid density and the gravitational accelleration.

Downcomer backup

The liquid on the plate may not exceed a certain height. When this does occur, the liquid can be transported up through the holes by the ascending vapour. This height depends on the height of liquid on the plate, but also on the height of liquid in the downcorner. The equation for the liquid in the down corner consists of contributions from the apron area and the weir length as is shown in equation (4.15)

I

h de

=

166( W )2

PJiqAapr

(4.15)

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FVO 3193 MSTDP

the height of liquid on the plate gives the downcorner backup height equation (4.16).

h =h +h _b +h +h

W oW,max t de (4.16)

This value is also needed to calculate the downcorner residence time. This is the estimated time the liquid needs to stay on the downcorner in order to get out all the vapour, so that no vapour will flow down the column.

AdChdcPliq

tres

Zw

(4.17)

The last check that has to be done is the entrainment check. No liquid is to flow up with the vapour. This is done by keeping the entrainrnent factor, which is calculated from the vapour veloeities, under a certain value. If this value is exceeded then the column design has to be adapted.

When these values are calculated for the column, there is nothing left but designing the column plates. After the liquid has come out of the downcorner and before it enters the new one, it can "rest" in an area without holes. This calming zone is a non-perforated area of nearly 5 cm. On the edges of a plate no holes are to be foud either. This gives a plate design, from which the total perforated area can be calculated and with which the number of holes can be found.

The efficiency

First the plate efficiency is calculated and then the column efficiency is calculated. The plate efficiency is dependent on the Reynolds number (Re), the Schmidt number (Sc) and the surface tension number (Dg) as is shown in equation (4.18).

(4.18)

This result ofthe so called Van Winkle's correlation is used in equation (4.19) to calculate the total column efficiency with the value of the equilibrium line and the molar vapour and liquid flows. mV log[l +E (--1)] mV'

L

Eo . -mV 100 ( - ) t:> L (4.19)

The initial number of plates is now divided by this efficiency to get the overall and real number of plates in the column. Then the fmal height and the final pressure drop can be calculated by multiplying the number of trays with the tray height and the pressure drop per plate.

The above mentioned calculation method results in the following column design. The total calculation can be found in appendix J.

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FVO 3193 MSTDP

T. a e bi 4 2 . sta I Izer eSlgn varia es b "L' d ' . bi

rectification stripping dcol(m) 1.768 3.733 ~ol(m2) 2.454 10.942 A.,;t (m2) 1.472 6.565 A_{dc (m}2) 0.491 2.188 l_w(m) 1.556 3.285 dh(mm) 4.5 4.5 hw(mm) 40.0 42.5 ltray(m) 0.60 1.00 ~Iate(mm) 5.0 5.0 Nho1es(-) 5.09*103 15.27*103 Eo(-) 0.858 0.885 Nplates(-) 30 26 ilP(bar) 0.520 hcol(m) 43.50

The column is made of stainless steel, type AIS! 304, and so are the plates and downcomers. This is based upon the fact that there are no corrosive materials, no excessive temperatures or pressures.

4.3 Flash design

For the flasher, equipmen number V12, a horizontal flash drum is chosen because of the high liquid to gas mass flow ratio (which equals 1.85). The dimensions of the flash drum are calculated with the help of various equations of Olujié [lit. 24] First the diameter of the horizontal flash drum is calculated. Therefore the gas velocity needs to be high enough to prevent excessive carryover of liquid droplets. The equation for vertical drums can be used (equation 4.20). The obtained gas velocity must be multiplied with 1.25.

1

1 ,...,- 1 2- C (PL-PC)?

u = .'!')·u = . ) "

-hsep vsep lIdrum

Pc

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FVO 3193

MSTDP

Using a standard demister, the capacity coefficient is set at 0.06.

With PL' PGat 814.7 and 12.27 kg/m3 , the gas velocity of a horizontal flash drum can be

calculated. The gas velocity is 0.61 mis.

The diameter of the horizontal flash drum is calculated with equation 4.21.

. I

M

-dhst'p= 1. 1284-( C ) 2

PCUhst'p<P C

(4.21)

When the diameter is known the length of the drum can be calculated with equation 4.22 and 4.23.The length ofthe horizontal flash drum depends on the residence time and the fractional gas cross section.

(4.22)

(4.23)

For a length to diameter ratio of about 5 the liquid height to diameter ratio (hL/dhsep) has to be between 0.5 and 0.6 .This is done by trying several different fractional gas cross sections. The liquid height follows from equation 4.24.

(4.24)

After several trials, the fractional gas cross section was set at 0.45. The residence time of the liquid in the drum is set at 300 seconds. This is a generally used value [lit. 24]. With a mass flow of 103227 kg/s and a liquid volumetrie flow of 126.7 kg/m3 the following dimensions were caIculated:

Dnsep: 3.3 m ~sep: 13.5 m

The construction material is stainless steel AISI 304, because the process fluids are non-corrosive and because hydrogen embrittlement could occur with carbon steel as a construction material.

4.4 Heat exchanger design

At two points, heat integration is applied. In H20 stream 37 is heated with stream 22. The heat exchanger calculated here is H7, heating the reactor feed with the effluent stream. The used design procedure is described in Coulson vol. 6 [lit. 5].

The chosen exchanger type is a shell and tube exchanger, because this is the most cornmonly used type. Heat exchanger H7 has an intemal floating head, because the temperature

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FVO 3193

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type is relatively easy to clean. A clamp ring is used to reduce the clearance between the outermost tubes in the bundie and the shell allowing fluid to bypass the tubes.

Equation (4.25) is used for calculating the required area, with known Q, U and L\ T m'

Q=U*A*L\T_m (4.25)

In this equation U is the overall heat transfer coefficient, taken from table 12.1 [lit. 5]. This gives, together with A and L\Tm,the mean temperature difference, calculated with (4.26), Q, which is the heat transferred per unit time (from ChemCAD)

(4.26) (4.27)

The stream with the higher pressure, in this case the feed stream, flows throught the tubes, because higher pressure tubes are cheaper than higher pressure shells. But, as the pressure difference between the feed and the effluent is small (0.75 bar), it will not be an

undiscussable choice.

The correction factor Ft is determined with R and S (dimensionless temperature ratios) and figure 12.20 [lit. 5]. The heat exchanger will be a two shell, four tube pass exchanger. This configuration will give a better approach to true counter current flow than a one shell pass heatexchanger.

(4.28) (4.29)

Now, the reguired area, ~eq can be calculated with equation (4.25)

The heat exchanging eguipment must be designed so that it has a value for U larger than used in equation (4.25). It depends on the tube and shell dimensions whether this can be acquired. This overall coefficient based on the outside area of the tube is calculated with equation (4.30). d ln( do) 1 1 1 0 dj do 1 do 1 - = - + - + +(-*-)+(-*-) Uo ho hl 2kw dj hjd dj hj (4.30)

The properties of the process streams determine the fluid film coefficients, the fouling factors are taken from table 12.2 [lit. 5] and the ~-factor is taken from table 12.6 [lit. 5].

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FVO 3193

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First, the tube side coefficient and the pressure drop are calculated. Equation (4.31) is used for ~. hi*di . R P 0.33 ( . .fi)0 14 - - =Jh

*

e

*

r_l

*

VlSC.corr. act. . k I I 'ft (4.31)

The index t stands for tube side. As the fluid is a gas, the viscosity correction factor (the actual viscosity divided by the viscosity of the fluid at the tube wall) may be considered unity flit. 5].

jh.t is a heat transfer factor taken from figure 12.23 [lito 5].

The dimensionless numbers of Reynolds and Prandtl are calculated with the stream properties. kr.t is the thermal conductivity, and the inner diameter dj is chosen, just like the outer tube diameter do and the tube length L.

With these tube dimensions, the area of the tube is known (equation 4.32). (4.32)

Then, with the number of tubes being Nt =

Arc/

_{Atube' the bundle diameter can be calculated}.

This Db is needed later in the shell side calculations.

N..!...

D =d *(_()n,

b n K, (4.33)

KI and n_{l ,}used in equation (4.33) are constants (table 12.4 [lil. 5]) for the tube configuration, in this case a triangular pitch is used.

The tube-side pressure drop is caused by contraction at tube inlets, expansion at the exits, flow reversals in the he aders and friction in the tubes.

The total pressure drop is calcu1ated with equation (4.34) 2 L p*u DoP =N *(8*Jj,' *(-)+2.5)*--1 I p ,I d. 2 I (4.34)

The linear velocity of the fluid through the tubes, Ut is calculated by dividing the volumetrie flow rate

<Pv

by the cut-through area of the tubes Act. jf.t Is the friction factor from figure 12.24 [lit. 5]. Np is the number of tube side passes, in this case 4. The loss in terms of velocity heads is estimated on 2.5 per pass flit. 5].

For the calculation of ho, the outer fluid film coefficient, Kern's method is used. Equation (4.35) is used to calculate this coefficient.

h *d o t ! · R P 0.33 - - = ]

*

e

*

r k h s s f (4.35)

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FVO 3193

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The index s stands for shell side.

jh

is

taken from figure 12.29 [lit. 5]. The baffle cut is 25%. The baffles used are single segmental baffles. For calculating the Reynolds number, the linear shell side velocity Us and

the equivalent diameter de are needed. The latter comes from equation (4.36), in which the triangular pitch arrangement is considered.

(4.36) Pt is the tube pitch, this is 1.25 times the tube outside diameter

do.

To obtain the shell side velocity, the crossflow area

Aa

is needed. Equation (4.37):

A _ (Pt-d)*Ds*IB

s

P

t

. (4.37) Ds is the shell inside diameter, determined from figure 12.10 [lit. 5]. Is is the baffle spacing, taken as O.4*Ds'

Us Is now calculated by dividing the shell side mass velocity by the density of the fluid. The

shell side mass velocity G_sis obtained by dividing the fluid flow rate on the shell side (Ws) by the crossflow area.

With equation (4.38)

(4.38),

the shell side pressure drop can be calculated. The friction factor jf comes from figure 12.30

[lit. 5].

With the properties of the fluids given in table 4.3, the chosen tube and shell dimensons, and factors taken from the tables and figures mentioned in the text, the results in tab Ie 4.5 were obtained.

Table 4.3 Properties of streams used in calculations

T[ 728 K T2 529 K t[ 370K

t₂ 528 K Cp,t 2690 Jlkg.K

_<Vv

_,

_!

1.97 m3/s Pt 23.57 kg/m3 ~t 3.2 104 Pa.s kr.t 0.1023 W/m.K Ws 43.6 kg/s _Ps 15.1 kg/m3 ~s 1.86 10.5 Pa.s

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FVO 3193

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Table 4.4 chosenfactors used in calculations

U 300 Ft 0.958 L 5m

W/m2.K

do 0.03 m di 0.027 m Kl 0.175

nl 2.285 Jh.t 0.0021 Jf.t 0.00215

Np 4 _Jh.s 0.0013 _Jf.s 0.028

kw

65W/m.K _hid 5OOOW/m2.K

_hoo

5000 W/m2.K Table 4.5 Resultsjrom calculations

ÀTm 171.2 K ~eq 516 m 2 Nt 1096 hi 1902 W/m2_.K _ÀP t 0.44 bar

Awbe

0.47m2 de 0.021 m ho 1651 W/m2_.K ÀPs 2.86 bar U_o 608 W/m2.K

Resuming, heat exchanger H7 is an exchanger with an internal floating head with clamp ring. The construction material is AISI 304 steel, because there are no extreme circumstances neither corrosive fluids, only the danger of hydrogen embrittlement. It has 1096 tubes with a length of 5 meter and an inner diameter of 27 mm and an outer diameter of 30 mmo The heat transfer area is 516 m2 and the overall coefficient based on the outside area of the tube is 608 W/m2.K. This is weU above the assumed transfer coefficient of 300 W/m2.K. The tube side pressure drop is 0.44 bar and the shell side pressure drop is 2.86 bar.

4.5 Pump design

The pump calculated here is the pump which brings the mixed feed and recycle streams from 1 to 35 bar, equipment number P5. The chosen pump is a two-stage centrifugal pump. With the help of the Stork Pompen B.V.-prograrnme "Handyselect" a pump with a horizontal shaft orientation and a speed of 1450 rpm is selected. The NPSH required will be approximately 3m and the total efficiency is 60 to 65%. Because the flow consists only of hydrocarbons, a carbon steel pumphouse (ASTM A 216 WCA) and a cast steel fan (ASTM A 48 - 40) meet the demands.

The isentropic power is calculated as follows [lit. 10].

(4.39) This leads to an isentropic power of 129 kW.

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FVO 3193

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p =p _s _isen

*"

_'lh (4.40)

The heat of dissipation is the difference between the isentropic power and the shaft power. This can also be written as:

Q=p -Po =tf.. c (T -T) s Isen '+' m p n I with:

<Pm

=

29.4 kg/s cp

=

182 J/mol K Ti

=

383 K (4.41)

Equation 4.41 can be used to calculate the pump outiet temperature, which becomes 388.6 K.

The total power delivered by the pump is:

P.

P =~ (4.42)

rotal 11 " 11 ~

Equation 4.42 can be simplified by knowing that

Th11_ =11(0'<11

and with 11lolal

=

0.64, the total power becomes 202 kW.

4.6 Compressor design

(4.43)

This design is made for the hydrogen recycle compressor, equipment number C8. The compressor selected is a single-section compressor with two impeller wheels and a verticalIy divided house. The material of the impellers is layered steel (ASTM A543 type B) and the compressor house is made of carbon steel (TSTE 355V)

The polytropic work is calculated with equation 4.44 [lit. 10]:

" -1 P --w _s.polytropl. _C=c _p*T*(-..!!..) _I n -1 cp

=

3634.7 Jlkg.K Ti

=

303 K Po

=

35 bar Pi

=

32.9 bar. Pi

The constant n is calculated with equation 4.45:

lC-1 n 11polytropic = - -*--1 lC n-with: 11polytropic = 0.85 lC

=

1.38 (4.44) (4.45)

11polytrepic is given by the manyfacturer of the compressor, see appendix G. With these constants, n becomes 1.48. Now, equation (4.44) can be calculated, _{W s}is 2.08 kj.

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FVO 3193 compressor becomes: TI ws.polytropic . Ipolytropic W S,r

MSTOP

(4.46)

The real work becomes 2.4 kJ. Now, the outgoing temperature can be calculated with T_o=

31OK.

-w _s.r=c _p*(T ₍₎-T.) _I (4.47)

The total efficiency TJtotal' which is 0.81, can be calculated with equation 4.48:

Tl tntnl =11 ,",rh *T} "0/''''''0";(' (4.48)

This value is taken from tabel 3.1 in Coulson vol. 6 [lit. 5], with the power of the motor in the compressor is given as 850 kW (appendix G)

The shaft power is calculated with equation 4.49:

CPm *ws,polytropic

TJlOtal p

.thaft

(4.49)

Here is:

cPm

=

12.8 kg/s. So, the shaft power becomes 356.4 kW.

4.7 Process

The process calculations have been perforrned with ChemCAD, a computer aided design prograrnrne for chemical plants. The results of these calculations are presented in appendices D & E. They are also used in sizing the equipment.

Overall, the feed, consisting of benzene, n-heptane, methyl cyclohexane, toluene, n-octane, ethyl benzene, para-, meta-, and ortho-xylene, n-nonane and isopropyl benzene is converted with a stream consisting of hydrogen and methane into seven product streams. These are a purge stream, a light ends stream, a n-octane stream a heavies stream and three more valuable streams. These last three are an ethyl benzene strearn, a meta- and para-xylene stream and a benzene stream.

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FVO 3193 MSTDP

Table 4.6 Results oftotal process

Feed (ktonne/yr) Products Purity Products

(ktonne/yr) (weight%) (tonne/tonne TX feed)

TX feed 450.0 Benzene 117.1 99.98 0.260

Hydrogen 42.7 para- and 138.1 99.76 0.307

feed meta-xy lene

ethyl 78.3 99.58 0.174 benzene purge 64.6 0.144 lights 14.9 0.033 . n-octane 15.9 79.06(oct) 0.035 heavies 63.7 68.15( ortho- 0.142 xyl)

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FV03193

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5. Safety, health and environment

5.1 Safety analysis

A restricted Hazard and Operability-analysis (HAZOP) flit. 3] is conducted to screen the design on process safety. The design is divided into three sections to enhance clarity ofthe analysis. The fITst section is the separation section including the feed, toluene distillation tower and the ethyl benzene/xylene separation train (unit Tl, T28, T39 and T45). The second section is the reaction section. This inc1udes the reactor (RIO), the hydrogen make-up feed, the gas compressor (unit C8). The flashdrum (V12) is excluded of this part. The rest of the design is called the product purification section, including stabilizer (TI5) and the benzene distillation tower (T22).

These sections are visualised in figure 2.1.

The results ofthe HAZOP-analysis for the sections are given in ttie following tables (5.1-.3). For a complete and acute shutdown of the plant in case of failing utilities or other disasters, the reactor is drained by using instalied fail-open valves. The drained contents will be stored temporarily. The hydrogen purge will be closed to prevent coking in the reactor. All the distillation columns will be equipped with fail-close valves to pre vent drainage. All pumps will be stopped and the fired heater is cooled with water.

In general pressure relief valves will be installed. These valves will open when the process pressure exceeds the design pressure by 10 percent.

One should also be aware of the fact that in the reactor, the temperature is above the auto ignition temperature of most of the reactants (see table B3 in appendix B). A runaway-warning and -prevention system is defmitely necessary and operating conditions should be weIl monitored.

Table 5.1 Results of HAZOP-analysis ofseparation section

Guide Deviation Possible causes Consequences Action required

Word

No No feed flow Naphtha no production, all Close valves of

cracker down equipment drained product streams

Piping/valve Environmental Close valves of

leakage pollution by loss of product streams and

content store feed

tempora-rily in vessel

More More feed Valve totally Bad operability by lack Flow control of

flow open ofcapacity product streams

More Bad cooling Bad separation Temperature

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FVO 3193

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Guide Deviation Possible causes Consequences Action required Word

More pressure High level of Bad separation Temperature condensate (flooding) control

Less Less feed Valve partially Overcapacity of Flow control of

flow blocked equipment product streams

Less product of Overcapacity of Flow con trol of naphtha cracker equipment product streams Pipinglvalve Overcapacity of Close valves of leakage equipment and product streams and

pollution store feed tempora-- rily in vessel Less Feed too cold Bad separation Temperature

temperature control

Bad Bad separation Temperature

heating/cooling control

Less pressure Pump Bad separation Pressure control deficiency

Pi ping/val ve Bad separation Close valves of

leakage product streams and

store feed tempora-rily in vessel

Table 5.2 Resu/ts of HAZOP analvsis ofreaction section

Guide Deviation Possible causes Consequences Action required Word

No No feed flow see table 5_1 Reactor runaway Temperature con trol on fired heater

No hydrogen hydrogen net is Coking in Reactor Close feed flow

make-up gas down valve or adapt

purge ratio No heating fired heater Too cold, no reaction Temperature

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FVO 3193

_MSTDP

Guide Deviation Possible causes Consequences Action required Word

More More runaway Reactor overheated frred heater off and

temperature . increase hydrogen

flow

Fired heater runaway in reactor Temperature

runaway control

More pressure Pump Equilibrium in reactor Pressure control malfunctioning shifts

Compressor Equilibrium in reactor Pressure control malfunctioning shifts

More feed Broken flow Low convers ion by· Install extra flow flow controller lack of capacity controller

Less Less Fired heater Low Conversion. Temperature temperature down Cracking capcity low. control

Less pressure Pump Coking in reactor Pressure con trol

malfunctioning and control flow

valve of reactor Compressor Coking in reactor Pressure con trol

malfunctioning and control flow

val ve of reactor Less feed see table 5.1 Runaway by Flow controller on

flow overheated feed or feed or temperature

local hot spots con trol on fired heater.

Less see table 5.1 Coking in reactor Close hydrogen

hydrogen purge

make-up gas

Part of Benzene Naphtha Conversion low by No content in cracker shift of equilibrium

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FVO 3193 MSTDP

Table 5.3 Results of HAZOP analysis of purification section

Guide Deviation Possible causes Consequences Action required

Word

No No feed flow see table 5.1 All equipment drained Close valves of product streams More More feed Valve totally Bad operation of flash Flow controller on

flow open drum and columns feed valve

More Runaway in Bad operation of flash Temperature

temperature reactor drum and columns control

More pressure Runaway in Bad operation of flash Pressure control

reactor drum and columns

Less Less flow Misoperation Bad separation in . Temperare and

flash columns pressure contral

Less Winter Bad operation of flash T emperature

temperature conditions drum and columns control

Steam net down Bad operation of flash Temperature drum and columns control

Less pressure see table 5.1 Bad operation of flash Pressure control drum and columns

5.2 Environment

The main purpose of this plant is, from an environmental point of view, to process the toluene and xylene stream from the naphtha-cracker, as these chemicals are less and less allowed as gasoline compounds. It is also necessary to make sure all the by-products can be used in other processes, to avoid wastage and pollution. In this MSTD-pracess, benzene, ethyl benzene and xylenes are further processed in other plants. The lights, corning off the stabilizer (T15), the heavies (bottom stream of T45) and n-octane are sent back to the naphtha-cracker. The purge stream in the hydrogen-recycle is purified and the hydrogen sent back to the PSA unit.

In this way, a full use of all the feed- and product streams is achieved.

The only emission th at takes place, is that of the furnace (F9). Here, emission control and off-gas treatment should be applied.

It should be noted that benzene is proven to be carcinogenic. Toluene and xylene should also be handled carefully. It is best to handle and store all chemicals according to the prescriptions given in e.g. Chemiekaarten flit. 22].

Energy losses must be prevented as much as possible. For this reason, heat integration is applied at two points: a feed/effluent heat exchanger (H7) and a heat exchanger in stream 22 and 37 (H20). Sea water is used as cooling medium, as it is available in excess. So it has no significant effect on the environment.

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FVO 3193 MSTDP

6. Process controll

6.1 Pumps, compressors and heat exchangers

Every pump and the compressor are bypassed. In this bypass stream a valve is installed. This valve normally is half open and is used to keep the pressure in the pipe at a certain value. Opening the valve means pressure decrease and vice versa. This is controlled with feedback pressure control in !he pipe after the pump or compressor.

All heat exchangers have a temperature feedback controller. Temperature fluctuations in the process streams to the heat exchangers or in the cooling or heating substances can be

counteracted by increasing or decreasing !he flow of this auxiliary substance.

6.2 Compressor

When the hydrogen-compressor (C8) breaks down, the hydrogen In the reactor wil not be sufficient to prevent coking. At this moment, the purge stream valve has to be closed and the feed to the reactor has to be stopped. Compressors are very expensive. Since breakdown of this compressor will cause the reactor, which is the heart of the plant, to stop, it is useful to have two compressors. One of them is running, the other is a spare one. In the economy calculations (chapter 7), the second compressor is not taken into account.

6.3 Fired heater and reactor

The fired heater (F9) is controlled with two temperature controllers. One is feedforward and is placed in the feed stream to the heater, !he other one is feedback and is placed in the reactor. They both influence the fired heater, but it is a master-slave configuration. Ihis is mainly done to prevent a runaway in the reactor. When the feed stream becomes bigger, more heating is needed. When this occurs at the same time as a temperature rise higher than 728 K in the reactor, the new feed will not be heated as extreme as wanted by the fust controller. This reactor temperature controller is an auctioneering system [lit. 29]. There will be several thermocouples in the reactor, but the controller will only be affected by the thermocouple with the worst signal.

There is also a third temperature controller in the cooling jacket. It operates so that when the temperature ofthe sea water changes, tbis will have no effect on the cooling capacity.

6.4 Purge stream, hydrogen-feed stream

Hydrogen is fed to the reactor in a certain ratio to the toluene. In the reactor feed stream is a composition controller. When the hydrogenltoluene ratio falls down because of higher toluene concentrations in the feed or lower hydrogen concentrations in the hydrogen stream, this stream is opened or shut.

Composition changes in the feed will also lead to other amounts of lower alkanes in !he reactor effluent due to Ie ss or more cracking reactions in the reactor. The flasher will give a larger top stream when there are more cracking reactions. The top stram then consists of less

(37)

FVO 3193 MSTDP

hydrogen and of more lower alkanes. The purge stream valve must be opened more in this case. This will also affect the hydrogenltoluene ratio and so the composition controller in the top stream of the flasher will have to be coup led to the composition controller in the hydrogen feed stream.

6.5 Flasher

In the flasher (V12) the liquid level is controlled with a valve in the bottom stream and the pressure is controlled with a valve in the top stream.

6.6 Towers

The top pressure in the columns is regulated with the flow of cooling water through the condenser. Increasing the cooling water flow reduces the amount of vapour and therefore also the pressure in the separating drum. The liquid level in this drum is controlled with a valve in the bottom stream. The reflux flow to the column is regulated with a valve in the recycle stream. The temperature in the bottom of the column is regulated with the amount of steam th at is sent through the bottom heat exchanger or with the amount of fuel bumt in the heaters. The level of fluid is controlled with a valve in the bottom stream.

(38)

FVO 3193

7. Economy

In this chapter an economical analysis of the designed MSTD process is performed. This analysis will show whether the new designed process is economically feasible.

MSTDP

The analysis is based on economie al analysis rnethods found in Montfoort [lil. 19] and Douglas [lito 8]. The calculations are expressed in US$. An exchange rate for US$ to Dutch Guilders of 1.7 DflIUS$ is used. For inflation corrections the cost index of 1993 (359.2) is used. The plant is assurned to operate 8000 hours a year. Taxes are neglected in all

calculations.

First the total costs and the sales incorne wil! be estirnated. With the criteria of return on investment (ROl) and intemal rate of return (IRR) the profitability of the process is

investigated. The pretreatment and regeneration processes of the ZSM-5 catalyst are neglected in this economical survey.

7.1 Cost estimations

For the meaning of alle the used symbols see the sybollist in chapter 9. The total costs of a process can be defined as follows:

(7.1)

In case of a preliminary design, all costs except the production volume dependent costs are based on the investments and the direct operating labour. Therefore a new definition of the costs is made where the ~ and the KM are put together with the terms KI and KL:

(7.2)

The production volume dependent costs are based on two elements: the quantity of raw materials and utility quantity and the costs per quantity.

Kp

=

a*~~ _I v.*q. _I _I (7.3) The investment costs are corrected with a factor f.

(7.4) The direct operating labour is calculated in the same way.

(7.5)

The values for the factors a, f and d used in equations 7.3, 7.4 and 7.5 are taken from the "Beste Model", Montfoort [lit. 19]

Production of benzene with the mobil selective toluene disproportionation process out of a naphtha cracker product stream

,Ni

FVONr.

Preliminary Plant Design

Laboratory of Chemical Process Technology

Subject

Production of Benzene with the

Mobil Selective Toluene Disproportionation

Process out of a Naphtha Cracker Product Stream

Authors

A.J. Brinkerink (Alexander)

J.N. van Es (1os)

N.M. Hiemstra (Ninke)

M

.

Vos (Michiel)

Keywords

Telephone

015-2122544

015-2132909

015-2613486

015-2142509

Date assignment

Date report

.

.

05-11-1996

11-02-1997

FVO Nr.

Preliminary Plant Design

Laboratory of Chemical Process Technology

Subject

Production of Benzene with the

Mobi1 Se1ective To1uene Disproportionation

Process out of a Naphtha Cr:acker Product Stream

Authors

AJ. Brinkerink (Alexander)

J.N

.

van Es (Jos)

N.M. Hiemstra (Ninke)

M. Vos (Michiel)

Keywords

Telephone

015-2122544

015-2132909

015-2613486

015-2142509

Date assignment

Date report

05-11-1996

11-02-1997

Summary

Contents

MSTDP

MSTDP

1.Introduction

MSTDP

2.Base of Design

2.1 Block scheme

i

TOLUEN~

I

i

I

i

T

I

r

I

,

1

Cl

I

l

IlA.

2.2 Reaction

--0

MSTDP

2.3 General process data

_pu;

_MSTDP

_<Vv

_,

_!

_MSTDP

_hoo