Table of appendices

(1)

Acknowledgements

We are grateful for the assistance, help and support received from the following people during the process design.

ir. drs. G. Bierman

em. prof. dr. ir. H. van Bekkum R.F. Caers

dr. ir. M.-O. Coppens prof. dr. E. Drent dr. ir. A.W. Gerritsen ir. J. de Glopper prof. ir. J. Grievink ir. A. Govaert, MBA prof. ir. G.J. Harmsen dr. ir. H. J. van der Kooi prof. ir. J.A. Moulijn ir. A.F.M. Paijens ir. A. Rooijmans

prof. dr. ir. W.P.M. van Swaaij ir. P.L.J. Swinkels

ing. M.C. Valentijn dr. ir. G.A. Verspui

(2)

Summary

This report describes the conceptual process design of the recovery of carbon monoxide and hydrogen from Low Joule Gas (LJG) followed by reaction with ethylene, from the

Unsaturated High Joule Gas stream (UHJG). LJG contains about 16% hydrogen and 8%

carbon monoxide. The remainder mostly consists of methane and nitrogen. Furthermore some sulphuric compounds are present in the ppm-level.

The hydroformylation of ethylene is opted as most favourable process, compared to, for example the copolymerisation of ethylene and carbon monoxide regarding the purification of the feed, the specialty chemistry involved and the higher pressure required. Moreover comparing the values of feedstock and product, the hydroformylation allows a larger investment than physical separation of individual compounds.

In the design the annually produced amount of propion aldehyde is 25 kt. Because the feedstock is impure, the required amount of feedstock is 9.37 t/t. Besides this impurity, this yield is due to 90% conversion and the product loss in the product purification. The weight percentage in this flow is 95.4 w% and the water content is 2.7 w%. Industrially, a water content between 1.0 and 2.5 w% is common.

Propion aldehyde has an internal market, as it will be used as an intermediate for further processing in ExxonMobil. Hydroformylation of ethylene is catalysed by rhodium with triphenyl phosphine as ligands, dissolved in tetraglyme ((tetra-ethylene glycol) dimethyl ether). The reaction takes place at 24 bar and 100ºC.

The off-gas of the stripper, which mainly separates aldehydes from gaseous components, is saleable at a pressure of at least 22 bars and at a specified Wobbe-index. The index of the off-gas is close to the index of UHJG. Since the off-gas is saleable and the reactant

concentration is relatively low, no recycle of the reactants is necessary.

The Total Capital Investment in this design is 18.4 M$. With the Total Production Costs of 37.9 M$/a the Discounted Cash-Flow Rate Of Return is 47% and 29% before and after taxes and depreciation, respectively.

A.J. Breugem L.J. Gerritsma R.A. Krul M.J.J. Over

(3)

Table of appendices

Appendix 1 : Project description, Conceptual Process Design Appendix 2 : Pure component properties

Appendix 3 : Additional information on propion aldehyde Appendix 4 : Market information on propion aldehyde Appendix 5 : Propionic acid usage

Appendix 6 : Product options and selection

Appendix 7.1 : Process Flow Scheme hydroformylation of ethylene Appendix 7.2 : ASPEN PLUS flow Sheet

Appendix 8.1 : Utilities summary

Appendix 8.2 : Calculation of heat exchange and utilities Appendix 9 : Process streams summary

Appendix 10 : Definition of feed and product streams Appendix 11 : BOD Margin

Appendix 12 : Azeotropes Appendix 13 : Process yields

Appendix 14 : Mass and heat balances Appendix 15 : Heat integration

Appendix 16 : Equipment calculation assumptions Appendix 17.1 : Equipment calculation of the compressor

Appendix 17.2 : Equipment calculation of the desulphurisation unit Appendix 17.3 : Equipment calculation of the hydroformylation reactor Appendix 17.4 : Equipment calculation of the impeller

Appendix 17.5 : Equipment calculation of the floating roof tank Appendix 17.6 : Equipment calculation of the heat exchangers Appendix 17.7 : Equipment calculation of the columns

Appendix 18 : Equipment summary specifications sheet

Appendix 19 : Guidewords and their explanations for a HAZOP study

Appendix 20 : Determination of a material factor for the Fire and Explosion Index Appendix 21 : Fire and Explosion Index of the Hydroformylation plant

Appendix 22 : Total equipment costs

Appendix 23 : Calculations of the equipment costs Appendix 24 : Calculations of the Wobbe index Appendix 25 : NPF, NFV and DCFROR

(6)

1 Introduction

Low Joule Gas (LJG) is an off gas stream originating from the Flexicoking process, which upgrades lower-quality crude oils into higher-valued gasolines and other fuels. Currently the principal uses this stream as fuel gas. However the LJG contains important chemical

building blocks, such as hydrogen and carbon monoxide. Due to considerations, which are elaborated on in chapter 2, the production of propion aldehyde is chosen. Propion aldehyde is produced by the hydroformylation reaction:

Cat

2 2 2 3 2

CH =CH + CO + H CH -CH -CHO [1.1]

The required ethylene is obtained from an Unsaturated High Joule Gas (UHJG) stream available onsite. Specifications of the LJG and UHJG streams are given in chapter 3 and appendix 1.

Propion aldehyde (propanal) is highly reactive, and is employed primarily as a chemical intermediate to prepare C3 and C6 compounds. It is generally produced by hydroformylation of ethylene. Some pure component properties of propion aldehyde are mentioned in

appendix 2. Additional physical properties of propion aldehyde are shown in table A3.1 in appendix 3.

Propion aldehyde is an intermediate and it serves as a source of a propyl group for chemical synthesis. Propion aldehyde is primarily converted to 1-propanol, propionic acid, and

trihydroxymethylethane. Minor applications include the production of pharmaceuticals, agricultural chemicals, rubber additives, and corrosion inhibitors. The most recent production levels of these compounds are those of 1988, which are stated in table A4.1 in appendix 4.

1.1 Main applications of the product

Since the discovery of the oxo reaction by Roelen (Ruhrchemie) in 1938 [1] companies have patented and implemented several processes to hydroformylate olefins. Generally, there are four types of hydroformylation processes. For the hydroformylation of terminal olefins, a division in two ways of processing can be made; reaction with a catalyst dissolved in an organic phase or in an aqueous phase solvent. Catalysts are rhodium triphenyl phosphine, with or without sulphonated phenyl groups, depending on the choice of solvent. Internal olefins are catalysed with (modified) cobalt. Choice depends on desired linearity, (un)desired co-production of paraffins, available pressure. These process types, taken from [2], [3]. [4]

and [5], are shown in table 1.1.

(7)

TABLE 1.1:COMPARISON OF HYDROFORMYLATION PROCESSES

Companies Union Carbide, Davy Powergas and Johnson Matthey

Ruhrchemie Rhône-Poulenc

Kuhlmann Shell

Catalyst [RhH(CO)(PR3)3] R=C6H5

[RhH(CO)(PR3)3]

R=3-C6H4SO3Na HCo(CO)4 [CoH(CO)3PR3] R=C6H5

Solvent Organic Water Organic Organic

P (bar) 20 50 200 70

T (K) 370 390 410 440

Olefin Terminal, C3 Terminal, C3 Internal, C3-C10+ Internal, C3-C10+

Product Aldehyde Aldehyde Aldehyde Alcohol

Linearity (%) 70-95 95 60-80 70-90

Paraffin

by-product (%) 0 0 2 10-15

Metal deposition No No Yes Yes

Heavy ends Little Little Yes Yes

Poison

sensitivity High High Low Low

Co/Rh costs High High Low Low

Ligand costs High High Low High

In this design the rhodium triphenyl phosphine in an organic phase solvent is chosen, due considerations elaborated on in chapter 2. The figure below depicts a typical flow scheme of the process layout.

FIGURE 1.1:FLOW SCHEME OF RHODIUM-CATALYSED HYDROFORMYLATION [2]

This conceptual design, however, differs from patents, since the process feed is impure. Due to the low concentrations no recycle is applied. A more thorough comparison with patents is made in chapter 2.

Oxo capacities by region in 1984 and 1994 are stated in table A4.2 in appendix 4. The consumption of some types of aldehydes by region in 1993 and 1998 are also shown in this appendix, in table A4.3.

1.2 Market situation

The amount of propion aldehyde produced globally is 196 kt/a (1998). This seems a small market, however, this does not include the internally used amount. The propionic acid production was in 1989 about 192 kt/a. The expected application of propion aldehyde is the oxidation to propionic acid. The distribution of producers and some commercial end usage areas of propionic acid are shown in appendix 5. Most propion aldehyde derivatives are produced in the United States and the demand is in Europe is larger than the supply and therefore the market opportunity is good.

370 K 20 bar

CO/H2

Olefin

Crude aldehyde to distillation

Off-gas Off-gas

Stripper

Gas-liquid separator Demister

Reactor

Regeneration

Rh catalyst Bleed Compressor

(8)

Because it is difficult to look at the direct market of propion aldehyde, some of the derivatives are also studied. Market outlook of propionic acid: Use of ammonium propionate or propionic acid as preservatives for feed and corn continues to drive overall market growth. Demand for older propionic-based herbicides is expected to remain flat or decline slightly, but newer phenoxypropionates are taking up the slack. Strong export demand for cellulose acetate propionate has compensated for a static US market. With the world market in 2000

estimated at about 180 kt/a and increasing at the rate of some 7 kt/a (~3.75%), supply is felt ample, [6]. At a production rate of 30 kt/a, the market will more ample, being increased 7%

additional to the annual increase.

1.3 Environmental issues

Because aldehydes have an unpleasant and intense odour, aldehyde-containing off-gases need to be drawn off centrally and burned. Wastewaters are usually treated chemically and biologically, but in this design the off-gases are burnt at the end-users of the Gasunie and the higher aldehydes are fed to the Gofiner. Some influences of the release of propion aldehyde are shown in appendix 3.

(9)

2 Product and process options and selection

As the objective of the project is to make a conceptual process design to make more economical use of two off-gasses originating from the Flexicoker (appendix 1) first a choice for the product has to be made. Considering the UHJG and LJG streams, there are several options to obtain economic gain. Important key components in these gas streams are hydrogen, ethylene and carbon monoxide. In this chapter some promising product and process options are considered.

2.1 Product options and selection

The following options are considered:

1. Recovery of pure hydrogen from the LJG streams

2. Recovery of pure carbon monoxide from the LJG streams 3. Recovery of pure ethylene from the UHJG streams 4. Production of methanol

5. Production of neoacids

6. Water/gas shift reaction to produce hydrogen

7. Copolymerisation of ethylene and carbon monoxide to polyketones 8. Hydroformylation of ethylene to propion aldehyde

Choosing the best option, several criteria are used. These criteria are economical potential, competition in the market, market demand, bulk versus specialty chemicals, process

conditions, the number of process operations and the kind of process operations. Hereof the economic potential is considered to be the most important criterion.

To assess the options economically in a more or less objective way, a base case scenario is used. In this base case an amount of 10% of the available LJG stream is used, because this is the maximum stoichiometric amount of LJG that can be used when reacting CO and H2 with the ethylene from the UHJG. For detailed evaluation of the processes see appendix 6.

T^ABLE2.1:PRO AND CONS OF PROCESS OPTIONS

Process Advantages Disadvantages

1 Internal market, Economic potential (830k$/kt) Low TCI (1M$), Expensive apparatuses 2 High mass percentage in feed stream,

Valuable chemical (140k$/kt), Internal market Low TCI (6M$), Expensive apparatuses 3 TCI (25M$), High margin (7.3M$/a) High pressure (50 bara), Low recovery 4 Interesting base chemical (157k$/kt) Competitive market, Low margins (0.5M$/a) 5 Valuable product Lack of proven technologies, Complex

processing

6 Internal market High investments, High pressure (45 bara), High temperature (700K)

7 Very high margin (87.1M$/a), High TCI (293

M$) Pure CO required, High operating costs, Lack of proven technologies

8 High margin (25.9 $/a), High TCI (87 M$), Proven technologies, Internal market Mild process conditions (25 bara, 100 ºC)

Expensive catalyst, Small external market

(10)

The recovery of one or more of the valuable compounds from LJG or UHJG appears to be economically unfeasible, due to the relatively low margin and the expensive process equipment (mostly compressors) or processing (cryogenic distillation). The purification is only feasible if the value of the manufactured pure components is below the price of the available pure streams.

The production of methanol is discharged because the market is competitive and the margins are low. Neo-acids are specialty chemicals. This implies high margins but also complex processing. A clear drawback is the lack of proven technology.

The water/gas shift reaction is rejected because the investments are high. Further the CO, which is used for the reaction, is more valuable and has a better internal market than H₂. The production of polyketone and the production of propion aldehyde are the most attractive options. In terms of net cash flow, the copolymerisation is the single best option. However, technological drawbacks are the minimum requirement of 97% purity of carbon monoxide, the purity of ethylene, the pressure of about 45 bara (doubled pressure if compared with the other product options) and the three phase separations and processing. This three phase processing is mostly a slurry type of batch, with expensive process costs. The most important drawback for the process is the lack of knowledge on the processing. The hydroformylation of the ethylene from the UHJG and the CO and H₂ from the LJG is more promising than copolymerisation. Advantages by choosing this process are more proven technology, milder process conditions and a better internal market.

2.1 Process options and selection

In the next paragraph the process for hydroformylation of ethylene from the UHJG and the carbon monoxide and hydrogen from the LJG is described.

2.1.1 Pre-treatment section

There are some undesired contaminations of sulphuric compounds and cokes fines present in the feed. The UHJG and LJG streams contain sulphuric, halogenic compounds and cokes fines. The sulphuric compounds and HCl are rhodium catalyst deactivating; the cokes fines are -alumina deactivating. The removal of these compounds is considered.

2.1.2 Removal of water

The removal of water could be considered because all water in the process ends up in the product stream. In the compressor a part of the water is removed by cooling the effluent of the first stage to 35^oC to condensate the water in the stream. The removal of water by distillation is not totally possible because of the azeotropic behaviour of water and propion aldehyde. The amount of water however is low, resulting in sufficiently high product specifications and recovery, even when performing normal distillation, where propion aldehyde and water are the top products.

2.1.3 Removal of sulphuric compounds

The feed of the reactor contains 60 t/a of sulphuric compounds. Because of the highly

deactivating behaviour of the catalyst pre-treatment is needed. Therefore the sulphur content has to be reduced to 50 wppb to ethylene [7].

One option to remove the sulphuric compounds is the washing of the feed by the product stream. Unfortunately this cannot be 100% efficient, despite the statements of [8], so catalyst deactivation still occurs and consequently pre-treatment is still necessary. In this design especially this is true, as the product stream is relatively small compared to the gas feed stream.

Furthermore the poisoning sulphuric and halogenic compounds remain present in the

product. For using the propion aldehyde as an intermediate, these poisoning compounds still have to be removed because of catalyst poisoning in a sequential process (e.g.

hydrogenation to n-propanol). When producing propionic acid from propion aldehyde the

(11)

specs of the product will most certainly not be satisfactory as propion acid is mainly used as an additive for cattle feed, [7].

In the Flexsorb unit the stream is desulphurised to the maximum allowed concentration of sulphur for burning. There is a possibility of further desulphurisation of the LJG stream but it is too expensive to desulphurise the whole stream (300,000 Nm³/h) while only a part of the stream is used (3%).

Therefore the desulphurisation is performed by conversion in a two stages. First the COS present is hydrolysed to H₂S in a zinc-promoted γ-alumina bed. In the second stage the remaining H₂S and CH₃SH are converted in a zinc oxide bed. As the reaction is irreversible, most COS, CH₃SH and H₂S is converted in these beds.

2.1.4 Removal of halogenic compounds

In the LJG some traces of HCl are present and this contamination should not exceed 50 wppb on ethylene basis too. In the washing step with water, before the LJG enters the battery limit, gas/liquid equilibrium is established and the amount of HCl in the water phase is known (3·10^-3 kg/m³ at 353 K and 2.3 bar, [9]). Using data on the solubility of HCl in water [10], the calculated weight based content of HCl in the LJG is 85 wppb on an ethylene basis.

The gas/liquid equilibrium data can be presented as:



HCl



^ ^ ^{H O}_HCl² ^

ln p 32.5 C 7.5 [2.1]

This leads to a vapour pressure of HCl of 0.08 Pa. At a pressure of 2.3 bar, this leads to a volume content of 0.066 vppm on an ethylene basis, the weight based is then 85 wppb. After mixing with UHJG this reduces to 23 wppb. This means the contamination does not exceed the maximum tolerated amount.

2.1.5 Removal of cokes fines

The low amounts of cokes fines (about 50 kg/a) could best be removed from the feed by an electrostatic precipitator [11]. The high investments however do not justify this pre-treatment, as most of the cokes will probably be kept away from the catalyst by the pre-treatment beds.

Moreover the fines that do come through will leave the process by the bottom (by-product) stream of the second column. However in conference with the principal it was decided to install a filter before the compressor. This filter is not designed, because this is too

specialised work, but the costs are taken into account in the economic calculations. The total erected costs are estimated at 1M$.

Impurities like H₂O, CH₄, CO₂ and N₂ are rather inert or harmless for the reaction.

The reactants are not purified from these inert components because this would need an intensive supplementary separation step and the reaction also takes place at low

concentrations of reactants according to [12] and [13]. Despite the need for larger units and an extra product purification process unit, the process will be more economical without purification of reactants.

2.1.6 Adding reactants

The addition of extra ethylene is an option to increase the production because more LJG is not limiting now. However in this design, there is chosen not to take this option into account, for it seems illogical to dilute pure ethylene with the very impure LJG to make use of the CO and H2 content.

(12)

2.1.7 Reaction section

The best catalyst to produce propion aldehyde is a rhodium complex. An alternative is a cobalt complex, but due to lower activity (1000 times less, [2]) and greater residence times this catalyst causes the formation of 1-propanol.

Using a rhodium catalyst for hydroformylation leaves some different ways of processing (see chapter 1). The aqueous biphasic catalyst system of Hoechst is based on the solubility of a hydrophilic sulphonated Rh-complex in water. The almost perfect separation of the water phase with the sulphonated rhodium complex from the organic product stream prevents the loss of costly rhodium. This is achieved by simply decanting the organic and water phase.

When hydroformylating in an organic phase, a high boiling oxygenated hydrocarbon is used as a solvent. The rhodium triphenyl phosphine complex is dissolved in this liquid and the propion aldehyde is separated from the catalyst solution by phase separation in a flash and demister.

The solubility of reactants (ethylene) in water is relatively low compared to the solubility in an organic phase. This is the reason why the reaction rate in aqueous phase hydroformylation is lower at the same vapour pressures. To reach the same reaction rate for aqueous phase hydroformylation, the pressure is increased. The pressure for industrial process aqueous phase reaction is 50 bar versus 15-20 bar for organic phase reaction.

Other techniques like Supported Aqueous Phase Catalyst (SAPC) and Supported Liquid Phase Catalyst (SLPC) are inferior because of the problems with stability of the carrier and particularly the leaching of catalyst [14]. Another disadvantage of such a catalyst system is the difficulty of heat removal.

Because of the significantly lower pressure, organic phase processing is preferred.

Operating with an aqueous phase on 50 bara would make it necessary to take one additional large compressor into account. The operating costs would be consequently higher. Another important consideration for choosing this option is the fact that there are only 2 operating aqueous phase plants compared to over 25 plants operating with organic phase [7]

worldwide.

As already stated in chapter 1 this design for the use of LJG and UHJG for the production of propion aldehyde differs on one major issue from patents: the feed is impure (only 6 mol%

ethylene, 6 mol% CO and 16 mol% H₂. According to some patents [12], [13] and [7] these low concentrations are no problem for processing, because in this design we satisfy the requirement of a partial pressure of ethylene of 0.15 bar, which prevent the catalyst from oligomerisation.

2.1.8 Reactor type

The production will be performed continuously for the following reasons:

1) The two feedstocks are also continuous streams

2) For large-scale processing continuous processing is more economical

Then, a continuously stirred tank reactor (CSTR) is the most applicable reactor type for the hydroformylation of olefins, because of the homogenous catalyst system. Another important advantage of using a CSTR is the ease of heat removal from the reactor, as the reaction is highly exothermic.

2.1.9 Recycle

Despite nearly all-industrial processes apply a recycle of non-converted reactants, in the design for the processing of LJG and UHJG a recycle is left out because of the low concentration of reactants in the first place and almost all reactants are reacted.

(13)

2.2 Recovery and purification section

Besides the ethylene the higher olefins present in the feedstock react with CO and H₂, resulting in higher aldehydes. Also small amounts of olefins are hydrogenated; the importance of this side reaction however is negligible, [15].

The first step to recover the aldehyde product stream is to separate the gaseous reactor outlet from the highly valuable catalyst and solvent. A demister, preventing droplets of solvent to leave the reactor, achieves recovery of nearly all catalyst. The losses can be considered negligible ([16] and [17]).

First the gas stream from the demister is fed to a stripping column to remove the gaseous components from the aldehydes. Secondly a distillation column separates the propion aldehyde from the higher aldehydes and most of the water.

A block scheme of the process is found in the chapter 3 in figure 3.2.

(14)

3 Basis of design

3.1

Description of the design

The objective of the project is to make a conceptual process design for the recovery and/or use of carbon monoxide and hydrogen from Low Joule Gas (LJG) originating from the Flexicoker to produce valuable products. Also an Unsaturated High Joule Gas stream (UHJG), containing 8.6% olefins (mostly ethylene), is available at the Esso refinery. See appendix 1 and table 3.4 for information on these streams. The reviewed processes are divided into two major groups, physical separation of pure products and chemical use of the products in the two streams. The different processes are briefly discussed in chapter 2 and in appendix 6 in more detail.

The hydroformylation of ethylene, hydrogen and carbon monoxide is the process chosen for the conceptual design. The process is divided into three main stages, which are discussed in more detail in the following paragraph.

1. Pre-treatment section 2. Reaction section 3. Recovery section

The pre-treatment section contains a compressor to raise the LJG pressure to 25 bar and desulphurisation units, which will be discussed below. The knockout drum in the compressor removes condensed water in the feed stream.

The reaction section contains a continuously stirred tank reactor, in which a gas/liquid reaction takes place in the liquid phase. The liquid phase is tetraglyme (tetra ethylene glycol dimethylether, a high-boiling, oxygenated organic solvent) in which a rhodium catalyst with phosphorous ligands is dissolved. In the liquid an excess of ligands is dissolved. The

reaction section also contains a phase separator to recycle the tetraglyme. This separator is a flash drum combined with a demister.

The recovery section is a stripper to separate gaseous products from the aldehydes and a distillation column to purify the propion aldehyde. This process results in a 95.4 w% product of propion aldehyde.

3.2 Process Definition

The utilities used for the processes described below can be found in the process flow sheet (appendix 7.1). The pure component properties of all components used are given in

appendix 2.

3.2.1 Process concepts chosen

In the following section all units will be discussed in more detail.

Cokes removal

Cokes fines, in this design, are mostly accumulated in the desulphurisation reactors. Some fines will accumulate in the hydroformylation reaction. The product stream will not contain cokes fines. Quantities however are very low (about 50 kg/a, assuming a cokes content of about 1 ppm).

LJG-compressor

The LJG gas stream of 2.3 bar has to be compressed to 25 bar, this is done by a two stage reciprocating compressor. The Low Joule Gas stream is cooled during the compression to prevent the lubrication oil from cracking. A knockout drum removes condensed water. Total

(15)

water removal is undesired, because this is needed for desulphurisation later. The outlet temperature is chosen to raise the feed temperature of the desulphurisation unit to 60 C.

Desulphurisation

The removal of sulphur is necessary to avoid poisoning of the catalyst. According to Esso the maximal tolerated amount of sulphur is 50 wppb to ethylene feed. The sulphur in the feed is present as COS, R-SH and H₂S, where R-SH is assumed to be CH₃SH. The

desulphurisation is a two-stage process. First the COS present is converted to H₂S in a Zn- promoted -alumina bed ([18] and [19]). In the second stage the H2S and CH₃SH are converted in a zinc oxide bed. The reactions and kinetics of the reactions are:

2 2 2

COS + H O CO + H S   H 30_mol^kJ [3.1]

 

 

1

2 2

[COS]

Rate k

1 k [H O] [3.2]

2 2

H S + ZnO H O + ZnS   H 63_mol^kJ [3.3]

3 3

CH SH + ZnO CH OH + ZnS   H 21_mol^kJ [3.4]

The equilibrium constant for the second reaction is assumed equal to the given constant for the first reaction:

 ²  ²

2 3

H O H O

p

H S CH SH

P P

K P P [3.5]

Because high pressure drop is undesired, three reactors are used in parallel and the two beds are used in the same reactor as demonstrated in figure 3.1:

FIGURE 3.1:SCHEMATIC OVERVIEW OF THE DESULPHURISATION UNITS

For optimal performance the desulphurisation will be performed at 60 C and an approximate pressure of 25 bar.

Reactor and catalyst

In the reaction section the following reactions take place:

^Rh

2 2 2 4 3

R-CH=CH + CO + H R-C H -CHO + R-CH(-CHO)-CH   H 120^mol^kJ [3.6]

With R = H, CH3, C2H5, etc. Herein R = H for the main reaction, the production of propion aldehyde from ethylene, H2 and CO. Selectivity towards linear products (first product in equation [3.6]) is no criterion in this design, contrary to most industrial hydroformylation processes.

ZnO bed for H2S and CH3SH conversion

-alumina bed for COS conversion

(16)

The reaction kinetics of the ethylene conversion in tetraglyme are given by [15] as a turn over frequency (TOF) in mol/(molRhs), where partial pressures are expressed in kPa and concentrations in mol/l:



  

    



   

C

T

C CO

CO

17000 1 1 P 0.0176 exp

1.987 373 T [TPP]

TOF P [TPP]

1 0.000496 31.8

[TPP] P

[3.6]

Therefore the reaction rate of ethylene is:

 

r_C^T  [Rh] TOF  _C^T [3.7]

The reactor is a continuously stirred tank reactor (CSTR) in which the gaseous reactants are fed into tetraglyme.

To design the reactor some requirements have to be met:

1 Kinetics [20]

- The partial pressure of the olefins in the reactor has to be at least 0.15 bar absolute, in order to prevent the catalyst from oligomerisation and consequent deactivation.

- The free triphenyl phosphine over rhodium ratio has to be at least 30.

- The rhodium concentration has to be between 0.01 and 10 mol/m³. - The concentration of free triphenyl phosphine has to be over 10 mol/m³. - The free triphenyl phosphine over partial pressure of CO has to be at least 0.1

mol/(m³kPa).

2 Mass transfer

The mass transfer of reactants from the gaseous phase to the liquid phase should not be limiting. Therefore the characteristic time of mass transfer is compared to the characteristic time of reaction. The following equation should hold:

masstransfer  reaction [3.8]

3 Heat transfer

To remove the heat of reaction besides a jacket also a coil is used. This is necessary when using one single reactor and still achieving a high extent of conversion. The content of the reactor again has to be mixed sufficiently.

4 Flow regime

The superficial velocity of the gas should be low, to avoid large amounts of liquid entrainment. The maximum superficial velocity is assumed to be 10 cm/s,

approximated from [21].

These requirements resulted in the design parameters listed in table 3.1:

TABLE 3.1:CSTR HYDROFORMYLATION SPECIFICATIONS

Catalyst Rhodium triphenyl phosphine (Rh-TTP)

Liquid solvent Tetra ethylene glycol dimethylether (tetraglyme)

Pressure 24 bar

Temperature 373 K

Conversion 90%

Partial pressure of ethylene 0.155 bar Superficial velocity 0.10 m/s

Reactor volume 40 m³

Gas hold up 0.1

Height over diameter ratio 3

(17)

Usage of one single CSTR is chosen instead of several smaller CSTR’s because

calculations made clear no advantages (e.g. less reactor volume, heat transfer without a coil) were achieved. An important reason is the need for a reactor of at least 40 m³, using a H/D ratio of 3 and preventing the solvent to entrain.

Flash-demister

The vapour/liquid separation isolates the solvent with the highly valuable catalyst from the gaseous products, reactants and inert compounds. A flash unit is performing this operation.

To prevent catalyst containing liquid droplets from leaving the reaction section along with the gasses, a demister unit is withholding even the smallest particles.

Off-gas stripper

In the off-gas stripper gaseous compounds are separated from the liquid aldehydes and water bottom stream. The top of the stripper column has a partial condenser and the top stream will be used as a High Joule Gas and will be sold to Gasunie at a price

corresponding to its heating value. The stripper has 8 stages and the feeding stage is on the first stage from the top. The bottom of the stripper is fed to a distillation column.

Distillation column

The distillation column has 28 stages and the feeding stage is the 16^thfrom the top. The top stream results in a 95.4 w% pure stream of propion aldehyde. The contents of the product stream are given in table 3.2.

TABLE 3.2:LIST OF IMPURITIES IN THE PRODUCT STREAM

Impurity w%

Water 2.7 Isobutane 1.1 n-Butane 0.4 2-Methyl-Butane 0.3

3.2.2 Block scheme

F^IGURE3.2:BLOCK SCHEME OF THE HYDROFORMYLATION PROCESS

3.2.3 Thermodynamic properties

The Redlich-Kwong-Soave with the Boston-Mathias (RKS-BM) alpha function was chosen as model for thermodynamic properties. This model is used for all units, except the hydrogen sulphide removal and the distillation columns. ASPEN recommends the RKS-BM model

LJG 9.53 t/h 3.02 t/t

UHJG 20.04 t/h

6.35 t/t 29.43 t/h

9.33 t/t Mixing

Section 25 bara 60 C

29.43 t/h 9.33 t/t Desulphuri

sation 25 bara

60 C

Reaction section 22.5 bara 100C

29.43 t/h 9.33 t/t

Stripper

22 bara 35-171C

Off-gas 25.92 t/h 8.22 t/t

Propion aldehyde 3.15 t/h 1.00 t/t

By-products 0.35 t/h 0.11 t/t Distillation

column 1 bara 44-78 C 3.51 t/h

1.11 t/t

Water 0.13 t/h 0.04 t/t

(18)

when hydrocarbons or synthesis gas are the main components. For the hydrogen sulphide removal the SOLIDS model is used and for the columns the Wilson model is used. Detailed information is given is chapter 4.

3.2.4 Pure component properties

The pure component properties are given in appendix 2.

3.3 Basic assumptions

3.3.1 Plant capacity

The feed streams are a Low Joule Gas stream (LJG) and an Unsaturated High Joule Gas stream (UHJG) originating from the Flexicoker at Esso Refinery in Rotterdam. The UHJG stream has a low concentration of ethylene (~8%) at 160.3 kt/a. To have an equimolar amount of carbon monoxide and ethylene in the reactor feed the LJG is kept at 76.2 kt/a.

There are four outgoing streams, the product stream, an off-gas stream, a rather small by- product stream (containing mostly higher aldehydes) and a water stream. The in- and outgoing are stated in table 3.3:

TABLE 3.3:IN- AND OUTGOING STREAMS

Stream ID Name Amount (kt/a)

IN <1> Low Joule Gas 76.2

<4> Unsaturated High Joule Gas 160.3

OUT <3> Water 1.0

<19> Propion aldehyde (Product) 25.0

<15> Off-gas 207.7

<20> Higher aldehydes 2.8

There will be 8000 operating hours per annum. The economical plant life is assumed to be 15 years. The capital investments and construction of the plant are done in the first 2 years.

Turnaround occurs every 4 years, simultaneously with the turnaround of the Flexicoking process.

3.3.2 Location

Because the feedstock of the process are the out going gas streams of the Flexicoker, the location of the plant is fixed at the Esso refinery in Rotterdam, the Netherlands. The process can also be used at other sites with a Flexicoker in operation.

3.3.3 Battery limit and definition of in- and outgoing streams

See appendix 10, for the exact composition of the in- and outgoing streams of the battery limit. Some assumptions were made about the two feed streams and these are stated in Table 3.4.A and 3.4.B.

(19)

TABLE 3.4.A:COMPOSITION OF THE LOW JOULE GAS FEEDSTREAM FOR MODELLING PURPOSES

Specification given by principal

Used in modelling the process

Explanation

CO 22 mol% CO 21.3 mol% The used percentage is lower because the LJG stream is normalised. The percentage is recalculated now including the water percentage.

H2 16 mol% H2 15.5 mol% See explanation CO.

CO₂ 8 mol% CO₂ 7.8 mol% See explanation CO.

N2 52.5 mol% N2 50.8 mol% See explanation CO.

CH4 1.5 mol% CH4 1.5 mol% See explanation CO.

H2S 300 vppm H2S 290 vppm See explanation CO.

COS 100-150 vppm COS 145 vppm See explanation CO. Upper value is used, so that the design is sufficient for the worst- case scenario.

Hydrocarbon ~10 wt ppm - - Neglected because the flow is small and because it is not catalyst inhibiting.

Flexsorb-SE Saturated at 1.3

barg and 313K - - Neglected because the flow is small and because it is not catalyst inhibiting.

H2O Saturated at 1.3

barg and 313K H2O 2.5 mol% Calculated amount of water if the gas stream is water saturated at 313 K and 1.3 barg.

Cokes ~ 1 ppm sub

micron particles Cokes - The cokes particles are neglected for modelling, because they are inert and accumulated in the stream.

TABLE 3.4.B:COMPOSITION OF THE UNSATURATED HIGH JOULE GAS FEEDSTREAM FOR MODELLING PURPOSES

Used in modelling the process

Explanation

H2 16.0-19.0 mol% H2 16.3 mol% Given value by Principal.

CH4 40.5-47.0 mol% CH4 46.80 mol% Given value by Principal.

C2H6 20.2-20.7 mol% C2H6 20.4 mol% Given value by Principal.

i-C4H8 0.06-0.012 mol% i-C4H8 0.092 mol% Given value by Principal.

n-C4H8 0.002-0.006 mol% n-C4H8 0.004 mol% Given value by Principal.

i-C4H10 0.1-0.5 mol% i-C4H10 0.4 mol% Given value by Principal.

n-C4H10 0.05-0.1 mol% n-C4H10 0.074 mol% Given value by Principal.

C5+ 0.01-0.03 mol% 2-C5H12 0.02 mol% Assumed is that all C5+ is present in the form of Methyl Butane, because C5+ is probably mostly C5and assumed is that these mostly consist of branched paraffins.

CO 0.5-3.5 mol% CO 0.8 mol% Given value by Principal.

N2 4.4-5.5 mol% N2 5.4 mol% Given value by Principal.

CO2 5-15 vppm CO2 10 vppm Given value by Principal.

H2O Dew point H2O 0.33 mol% Given value by Principal.

COS 0.5-1.5 vppm COS 0.7 vppm Given value by Principal.

RSH 4-8 vppm CH3SH 6.3 vppm Assumed is that all organic sulphur is in the form of methyl mercaptan, because this is the lightest form of organic sulphur.

H₂S 10-21 vppm H₂S 14.2 vppm Given value by Principal.

The incoming LJG stream is at 2.3 bara and 40 C and is first compressed to 25 bara, to be mixed with the UHJG stream. The mixing ratio of the two streams is UHJG : LJG = 2.73 : 1.

The UHJG is at 25 bara and 40 C and has a flow rate of 160 kt/a.

(20)

The product stream specification depends on the purpose. In case of an chemicals grade 99.5 w% is desired, although industrially produced propion aldehyde (99.5 wt % on dry basis) with a 1.0-2.5 w% water is considered to be a high-purity grade [22] and [23]. In case of Aldol condensation as next processing step (e.g. for the propionic acid production) the water content may be higher. Water is generally added to decrease peroxide formation on standing [24]. Examples of utilities are present on the Botlek (Rotterdam) refinery are stated in table 3.5.

TABLE 3.5:UTILITIES AT THE BOTLEK REFINERY

Utility Availability

Steam 1, 3, 9 and 40 barg Cooling water 300 m³/h (T_outlet^max = 54 ºC) Demineralised water Not known and used Boiler feed water Not known and used Clean condensate Not known and used Dirty condensate Not known and used Nitrogen Not known and used Pure oxygen Not known and used

Hydrogen 70%, 95% and 99.9% purity

3.4 Economic Margins

The off-gas will be sold as heating gas to Gasunie. The economic value of this stream is determined by comparing the heating value to natural gas. Further usage of the higher aldehydes waste stream exceeds the battery limits, but will be used as additional feed of the Gofiner.

  



comb fuelgas

fuelgas comb NG

NG

Pr ice H Pr ice

H [3.9]

To calculate the margin, these costs were subtracted from the value of propion aldehyde, the value of the higher aldehydes stream and the off-gas stream. The value of the off-gas

stream is determined analogously to the values of the LJG and UHJG streams.

Given the optimal mixing ratio of UHJG:LJG = 2.73 and the annual production of 25.0 kilo tonnes propion aldehyde, the margin is determined. The values are stated in table 3.6.

 propionaldehyde  higher aldehydes  off gas  LJG UHJG

MARGIN C C C C C [3.10]

TABLE 3.6:THE VALUE OF THE FEED AND THE PRODUCT, AND THE MARGIN OF THE PROCESS

Item Stream Price ($/t) Value

CLJG <1> 19 1M$/a

C_UHJG <4> 148 24M$/a

Coff-gas <15> 120 25M$/a

Chigher aldehydes <20> 120 0

Cpropion aldehyde <19> 942 24M$/a

MARGIN 24M$/a

At an internal rate of return = 60% and 13 years of production after 2 years of construction, the maximum allowed total capital investment is 19 M$. After taxes (35%) this value is 12 M$. Appendix 11 gives an overview.

(21)

4 Thermodynamic properties 4.1 Thermodynamic model

ASPEN recommends PENG-ROB, RK-SOAVE, LK-PLOCK, PR-BM and RKS-BM as appropriate thermodynamic models to use for non-polar, real compounds [25]. The PENG- ROB (Peng Robinson) and the RK-SOAVE (Redlich-Kwong-Soave) cubic equations of state are used in the models PR-BM and RKS-BM, respectively. The latter two models use the particular equation of state with the Boston-Mathias (BM) alpha function for all

thermodynamic properties and are therefore more extensive models. ASPEN also recommends using PR-BM or RKS-BM as thermodynamic model for hydrocarbon

separations and for synthesis gas reactions. In case of polar, non-electrolyte systems many models can be used. Dechema [26] shows the best modelling results with the NRTL and Wilson models.

There is little difference between these two models, the Redlich-Kwong-Soave with the Boston-Mathias alpha function is chosen. However for the absorption of hydrogen-sulphide in the zinc oxide bed an other thermodynamic models has to be used, because the zinc oxide is a solid component, therefore for this section the SOLIDS model is used. For the columns the Wilson model is used, because this model corresponds best with the found literature, which is shown in figure 4.1.

The components 2-methylbutanal, 3-methylbutanal and dimethyl propion aldehyde were not available in the ASPEN databanks and their thermodynamic properties were therefore estimated with ASPEN using the Joback method.

4.2 Data validation

The chosen model was tested by comparing calculated values of enthalpy, entropy and specific heat for the five most important components present in the system, H₂, CO, C₂H₄, C₅H₁₀O and H₂S. The equations used are:

        

ig 2 2

cp R (A B T C T D T ) [4.1]

   

T 0 pT 0

H H c (T T ) [4.2]

   

       

   

T 0 pT

0 0

T P

S S c ln R ln

T P [4.3]

A, B, C, D, H₀, and S₀ were taken from the Handbook [27]. The following tables show the values of the heat capacity, enthalpy and entropy of the above-mentioned compounds at 4 pressure and temperature combinations of standard temperature and pressure and elevated temperature and pressure. The temperature of 120C and pressure of 40 bara is shown to check a large validity range.

(22)

TABLE 4.1:VALIDATION OF THE ASPEN MODEL AT 298K AND 40 BARA

c_pT [kJ/(moleK)] HT [kJ/mole] S_T [kJ/(moleK)]

H₂ Calculated 0.0248 0 0.100

RKS BM 0.0290 0.02 -0.031

Difference -0.0042 -0.02 +0.131

CO Calculated 0.0292 -110.5 0.167

RKS BM 0.0315 -110.8 0.058

Difference -0.0023 +0.3 +0.109

C₂H₄ Calculated 0.0442 52.5 0.189

RKS BM 0.0664 50.3 -0.090

Difference -0.0222 +2.2 +0.279

C3H6O Calculated 0.0807 -185.6 0.274

RKS BM 0.0261 -191.3 -0.238

Difference +0.0546 +5.7 +0.512

H2S Calculated 0.0366 -20.6 0.175

RKS BM 0.0342 -24.8 0.001

Difference +0.0024 +4.2 +0.174

TÂBLE4.2:VÂLIDATIONO^FT^HEA^SPENM^{ODEL AT}398KÂND40^BARA

cpT [kJ/(moleK)] H_T [kJ/mole] ST [kJ/(moleK)]

H₂ Calculated 0.0266 2.66 0.108

RKS BM 0.0293 2.94 -0.022

Difference -0.0027 -0.28 +0.130

CO Calculated 0.0298 -107.5 0.176

RKS BM 0.0305 -107.7 0.067

Difference -0.0007 +0.2 +0.109

C₂H₄ Calculated 0.0536 57.9 0.204

RKS BM 0.0590 56.2 -0.072

Difference -0.0054 +1.7 +0.276

C₃H₆O Calculated 0.0975 -175.8 0.302

RKS BM 0.0004 -184.8 -0.221

Difference +0.0971 +9.0 +0.523

H₂S Calculated 0.0377 -16.8 0.186

RKS BM 0.0433 -18.6 0.020

Difference -0.0056 +1.8 +0.166

(23)

TABLE 4.3:VALIDATION OF THE ASPEN MODEL AT 298K AND 1 BARA

c_pT [kJ/(moleK)] HT [kJ/mole] S_T [kJ/(moleK)]

H₂ Calculated 0.0248 0 0.131

RKS BM 0.0299 -0.23 0.000

Difference -0.0051 +0.23 +0.131

CO Calculated 0.0292 -110.5 0.198

RKS BM 0.0380 -110.5 0.089

Difference -0.0088 0 +0.109

C₂H₄ Calculated 0.0442 52.5 0.220

RKS BM 0.1759 52.5 -0.053

Difference -0.1317 0 +0.273

C3H6O Calculated 0.0807 -185.6 0.305

RKS BM 0.1265 -186.5 -0.207

Difference -0.0458 -0.1 +0.512

H2S Calculated 0.0366 -20.6 0.206

RKS BM 0.0846 -20.7 0.043

Difference -0.0480 +0.1 +0.163

TÂBLE4.4:VÂLIDATIONO^FT^HEA^SPENM^{ODEL AT}398KÂND1^BARA

cpT [kJ/(moleK)] H_T [kJ/mole] ST [kJ/(moleK)]

H₂ Calculated 0.0266 2.66 0.138

RKS BM 0.0300 2.90 0.009

Difference -0.0034 -0.24 +0.129

CO Calculated 0.0298 -107.5 0.206

RKS BM 0.0348 -107.6 0.098

Difference -0.0050 +0.1 +0.108

C2H4 Calculated 0.0536 57.9 0.235

RKS BM 0.0784 57.3 -0.040

Difference -0.0248 +0.6 +0.275

C3H6O Calculated 0.0975 -175.8 0.333

RKS BM 0.1582 -177.7 -0.182

Difference -0.0507 +1.9 +0.515

H2S Calculated 0.0377 -16.8 0.217

RKS BM 0.1543 -17.2 0.053

Difference -0.1166 +0.4 +0.164

The values calculated and found by ASPEN for the heat capacity and enthalpy are

practically identical. The results for the calculations of the entropy do differ significantly. The difference however remains almost constant, which might indicate that ASPEN has used a different value for S0. As this will make no difference in the calculations that ASPEN

performs, the RKS-BM is a suitable thermodynamic model, in the parts of the design where the aldehyde and water interactions are not important.

However the values for propion aldehyde differ the most. This can be explained by the polar characteristics of aldehydes. Therefore it is sensible to use another model for the final separation in the process. As mentioned above the Wilson model corresponds well with literature.

The vapour-liquid diagram for a binary system with propion aldehyde and water is shown in the figure below. In the design, both in the stripper and the aldehyde distillation column, the Wilson model is used. Validation of this model is done by comparing the x and y values of the binary system, which are predicted by the Wilson model, and the experimental data, which stem from [28]. As shown in figure 4.1, the Wilson model properly describes the

(24)

vapour liquid behaviour of an aldehyde water mixture. The azeotrope in a mixture of propion aldehyde and water is visible at x = y (no driving force for phase separation). The x,y-

diagram at 1 atm and 22 bar is shown in appendix 12. This appendix also contains an overview of azeotropic behaviour of some aldehydes and water.

F^IGURE4.1:COMPARISON BETWEEN THE WILSON MODEL AND EXPERIMENTAL DATA

Also the activity coefficients are compared. In figure 4.2 these coefficients are shown for propion aldehyde and water at 22 bar and atmospheric pressure. Also the literature values of the activity coefficients of both compounds at infinite dilution(^_WATER,lit and ^N PRO 01,lit_ _ ) are indicated. This figure properly shows interaction between aldehydes and water.

0 10 20 30 40

0 0.2 0.4 0.6 0.8 1

x_N-PRO-01

liquid _WATER,lit

 

N PRO 01,lit

N-PRO-01,1 atm WATER,1 atm

FIGURE 4.2:ACTIVITY COEFFICIENTS FOR THE PROPION ALDEHYDE/WATER SYSTEM

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1

0 0.2 0.4 0.6 0.8 1

x_N-PRO-01 yN-PRO-01

1 atm

(25)

5 Process structure and description

5.1 Process Flow Scheme

The Process Flow Scheme can be found in appendix 7.1.

5.2 Process stream summary

The process stream summary can be found in . In this summary mass flows are tabulated for all components separately and in total for each stream. Furthermore,

temperature, pressure, phase (V, L, V/L) and enthalpy of the streams are specified.

5.3 Utilities

The utility summary can be found in appendix 8. In this summary information is given on consumption of utilities fro each apparatus. For the most important units a short description is given below.

TABLE 5.1:SHORT DESCRIPTION OF SOME PROCESS UNITS

Number Unit Description

K101 LJG two-stage

compressor

A two stage cooling reciprocating compressor elevates the pressure to 25 bara

R101/R102/

R103

COS and H₂S converters

In the first stage COS is converted to H₂S in a Zn- promoted -alumina bed. In the second stage the H2S and CH₃SH are converted in a zinc oxide bed. Three beds are operated parallel to reduce pressure drop and possibility to stay in operation if one of the beds is clogged.

R201 Hydroformylation reactor

A continuous stirred tank reactor in which the gaseous reactants CO, H₂ and ethylene are converted into propion aldehyde. The catalyst rhodium-triphenyl phosphine is dissolved in tetraglyme. A conversion of 90% is reached. The pressure and temperature are 24 bara and 100 ^oC respectively.

V201 Flash/demister Flash and demister to recover and recycle the liquid solvent tetraglyme with the catalyst and ligands to the reactor.

C301 Stripper The off-gas leaves the stripper on the top. The bottom is further purified in distillation column C302.

C302 Distillation column The column purifies the propion aldehyde from the higher aldehydes.

5.4 Process yields

Process yields are important parameters for monitoring a process and comparing it with other processes. The yields can be found in appendix 13.

(26)

6 Process control

Process control is needed to suppress external disturbances. The control system has to ensure the stability of the process and optimise the performance of a chemical process. This is done by a rational arrangement of equipment and human intervention. In this chapter all processing units with their possible controllers are discussed. The control equipment can be found in the Process Flow Scheme shown in appendix 7.1.

6.1 Process controllers

Only basic controllers are used in the design. In the following table an overview of all

process units and their process controllers is shown. FC, PC, TC and LC are flow, pressure, temperature and level controllers, respectively.

TABLE 6.1:OVERVIEW PROCESS CONTROLLERS

Unit/stream Description Controller(s)

K101 LJG 2-stage compressor FC, intern and extern: PC, TC Stream 1 and 4 UHJG/LJG stream FC

Stream 5 Feed desulphurisation unit PC, TC, FC R101/R102/R103 COS and H₂S converter PC

Stream 7 Output desulphurisation unit FC

E201 Pre-heater TC

E301 Effluent cooler TC

E302/E305 Condenser TC

E303/E306 Reboiler FC

E304 Stripper bottom cooler TC

R201/R202 Hydroformylation reactor TC, LC, PC

V201A Flash vessel LC

V201B Demister vessel LC

PV301/302 Pressure valve PC

Stream 11 Solvent recycle FC

V301/V302 Reflux accumulator LC

C301/C302 Stripper and distillation column TC, LC 6.1.1 Compressor

The compressor temperature and pressure are internally controlled with a TC and PC controller. The temperature is compensated by a two stage cooling system.

The external controllers are in charge of controlling the pressure and temperature of the compressor-outgoing stream. A flow controller is used to control the power of the

compressor. The internal control system can be represented as:

F^IGURE6.1:SCHEMATIC CONTROL SYSTEM OF THE RECIPROCATING COMPRESSOR

6.1.2 LJG and UHJG stream

The UHJG stream is the limiting stream and is preferably totally used. To realise a

stoichiometric relation of UHJG:LJG of 2.7:1 the LJG flow has to be controlled by a valve.

G

TC PC TC

G

(27)

6.1.3 COS, CH3SH and H2S converter

The feed of the desulphurisation unit has to be controlled to sense pipe blockage, process disorder etc. The pressure in the COS, CH3SH and H2S converters has to be controlled with a valve in case the flow through the bed is clogged. In case of a sudden pressure rise in one of the beds the stream will be rerouted to the other two beds. Placing a pressure controller coupled with a valve can be used to monitor and control abnormalities.

6.1.4 Reactor

The gas hold-up in the reactor has to be controlled to keep the actual reactor volume constant. The level controller and the feedback loop to the feeds realises this.

The temperature and pressure is controlled by the throughput of the cooling water in the jacket. The FC in the solvent recycle has to make sure no catalyst will be lost.

6.1.5 Flash and demister

The level in the flash drum and demister is controlled to avoid drying. This is also valid for the two reflux drums in the condensers of both columns.

6.1.6 Pressure valve

The pressure has to be controlled to make sure the stream is sufficient decompressed.

6.1.7 Heat exchangers/condensers

In all heat exchangers the output temperature is controlled by the throughput of the cooling/heating medium.

6.1.8 Columns

A valve in the bottom stream controls the level in the reboiler so it will not flood nor dry up.

In the stripper the top is partially condensed, therefore the condensation rate determines the reflux ratio. In the distillation column, condensation is total. The reflux ratio is set by a switch, which splits the condensate.

Additional controls are discussed in chapter 10.

Table of appendices

Acknowledgements

Summary

Table of contents

Table of appendices

1 Introduction

1.1 Main applications of the product

1.2 Market situation

1.3 Environmental issues

2 Product and process options and selection

2.1 Product options and selection

2.1 Process options and selection





2.2 Recovery and purification section

3 Basis of design

Description of the design

3.2 Process Definition

3.3 Basic assumptions

3.4 Economic Margins

4 Thermodynamic properties 4.1 Thermodynamic model

4.2 Data validation

5 Process structure and description

5.1 Process Flow Scheme

5.2 Process stream summary

5.3 Utilities

5.4 Process yields

6 Process control

6.1 Process controllers