TECHNICAL FIELD OF THE INVENTION

The present invention relates to a plant and a catalytic process for the hydrogenation of phenol on an industrial scale whereby mainly cyclohexanone and cyclohexanol are formed.

BACKGROUND OF THE INVENTION

The vast majority of cyclohexanone is consumed in the production of £-caprolactam, which is an intermediate in the manufacture of Nylon 6. Mixtures of cyclohexanone and cyclohexanol are used for the production of adipic acid, which is mainly converted into Nylon 6,6. In addition, cyclohexanone can be employed as an industrial solvent or as an activator in oxidation reactions. It can also be used as an intermediate for the production of cyclohexanone resins.

In the 1930's, the production of cyclohexanone started on an industrial scale in parallel with the commercial production of e-caprolactam, adipic acid, Nylon 6 and Nylon 6,6. Ever since, the production volume of cyclohexanone has been growing and nowadays the annual production of cyclohexanone is over 7 million tons.

The three main commercial routes for production of cyclohexanone and cyclohexanol are based on the oxidation of cyclohexane, the hydration of cyclohexene and the hydrogenation of phenol, respectively. The first two routes generate large quantities of undesirable by-products and require large energy inputs. The third route, i.e. , the hydrogenation of phenol to cyclohexanone is either done in a “two-step” or in a “one-step” process. In the “two-step” process, phenol is first reacted with hydrogen by, e.g., using a Ni-comprising catalyst, to form cyclohexanol, which is consequently dehydrogenated to give cyclohexanone. In the “one-step” process, phenol is directly and with high selectivity hydrogenated into cyclohexanone. This “one-step” phenol hydrogenation process is known to combine a high product yield with a low energy demand. In the “one-step” process, cyclohexanone is prepared from phenol by catalytic hydrogenation in a phenol hydrogenation reactor, e.g., using a Pt- or a Pd-comprising catalyst. During phenol hydrogenation cyclohexanol is formed as main by-product next to only minor amounts of other compounds. Hydrogenation of phenol can be carried out in the liquid phase or the vapor phase (also often called gas phase). See, e.g., “Cyclohexanol and Cyclohexanone”, Kirk- Othmer Encyclopedia of Chemical Technology, e.g., 3rd Edition, Vol. 7 (1979) p. 410- 416; I. Dodgson et al. “A low Cost Phenol to Cyclohexanone Process”, Chemistry & Industry, 18, December 1989, p. 830-833; or M T. Musser “Cyclohexanol and Cyclohexanone”, Ullmann's Encyclopedia of Industrial Chemistry (7th Edition, 2007).

The (exothermal) hydrogenation of phenol to the two major products, cyclohexanone and cyclohexanol, can be described by the following stoichiometric equations: phenol hydrogen cyclohexanone phenol hydrogen cyclohexanol

The (endothermal) dehydrogenation of cyclohexanol to cyclohexanone can be described by the following stoichiometric equation: cyclohexanol cyclohexanone hydrogen

GB890095 describes a process for preparing cyclohexanone from phenol by catalytic hydrogenation, which comprises passing gaseous phenol together with hydrogen over a catalyst containing a metal belonging to the palladium group at a temperature below 250 °C. It mentions that the gaseous reactants may be diluted with an inert gas such as nitrogen, argon or propane.

GB1316820 describes a method for producing cyclohexanone by the hydrogenation of phenol in the vapor-phase in the presence of a catalyst composed of a carrier having deposited thereon a layer of palladium and an alkali metal carbonate as a promoter. It teaches that selectivity with respect to cyclohexanone is influenced by the presence of a promotor.

GB1332211A describes a process for the production of cyclohexanone by catalytic hydrogenation of phenol in the vapor phase at a temperature between 75 and 250 °C in the presence of a catalyst containing between 0.1 and 5% by weight of the metal of the platinum group on active alumina and between 5 and 50% by weight alkaline earth metal hydroxide and/or alkaline earth metal oxide based on the total weight of the catalyst.

W02016075047A1 describes an industrial scale continuous process for the production and recovery of cyclohexanone from phenol and hydrogen. It teaches that the net steam consumption of such a process can be less than 1.5 kg steam per kg produced cyclohexanone.

WO2011073233A1 describes a method for preparing cyclohexanone, cyclohexanol or a mixture thereof in a continuous way by catalytically hydrogenating phenol fed into a reactor comprising a supported hydrogenation catalyst, comprising a dopant. It teaches that an increased conversion of phenol, and/or an increased selectivity towards cyclohexanone and/or cyclohexanol can be obtained by feeding water into the reactor during the hydrogenation of phenol.

The combination of a (very) high selectivity towards cyclohexanone and thus a (very) low selectivity towards cyclohexanol with a (very) high per pass conversion of phenol in the reactor section is highly desirable due to its economic advantages: smaller sized equipment for cyclohexanol and phenol recovery from the product stream and for cyclohexanol dehydrogenation, and reduced energy consumption for cyclohexanol and phenol recovery from the product stream and for cyclohexanol dehydrogenation. As a consequence, the inventors realized that industrial scale phenol hydrogenation plants for the production of cyclohexanone should be operated in an optimal process window with high per pass selectivity towards cyclohexanone in combination with high per pass conversion of phenol. The research that led up to this invention showed that optimum combination of per pass selectivity towards cyclohexanone and per pass conversion of phenol for a phenol hydrogenation plant depends on many variables, including the origin, sources and quality of the feedstocks phenol and hydrogen, type of catalyst, process type, plant lay-out, and desired production rate.

In practice, an increase in per pass selectivity towards cyclohexanol, and thus a decrease in per pass selectivity towards cyclohexanone, is usually observed when the per pass conversions of phenol is increased. Especially at (very) high per pass conversions of phenol, a steep increase in per pass selectivity towards cyclohexanol is observed, which does not allow one to obtain a 100 % per pass conversion of phenol and 100 % per pass selectivity towards cyclohexanone in combination.

The catalyst systems used in phenol hydrogenation plants can be significantly affected by degradation due to aging of the catalyst. “Aging” is typically described as catalyst deactivation or loss of catalyst activity and/or selectivity as a function of time on stream (TOS). In commercial phenol hydrogenation plants, the time on stream of a given catalyst batch can range from a period of time of several months up to many years. Without wishing to be bound by theory, it has been hypothesized that aging is caused by deposition of organic compounds on the catalyst (also often called coking) and/or by change of the crystal structure of the catalytic metal.

In general, the costs of cyclohexanone production by hydrogenation of phenol are increasing due to aging of the catalysts (due to, e.g., lower selectivity to cyclohexanone, lower per pass conversion of phenol, reduced productivity and higher purge losses of hydrogen).

Generally, the end of run (EOR) of a given hydrogenation catalyst batch is determined by technical (e.g., arriving at maximum allowable temperature, or arriving at maximum allowable pressure) or economic (see before) considerations.

In order to maximize the catalyst life, regenerations can be performed under carefully controlled (combustion) conditions to remove organic deposits on the catalyst. Such a catalyst regeneration might optionally include a) a thermal treatment in a substantially inert (free of oxygen) atmosphere to remove carbon-containing volatile material present on the catalyst, b) a thermal treatment in an oxygencontaining atmosphere to oxidize carbonaceous material present on the catalyst, c) cooling down of the catalyst after step a) and/or step b).

Maximizing the catalyst life is very desirable, because catalyst replacements and catalyst regenerations are time-consuming and expensive activities (e.g., due to reduced product output).

Up to now, it is still a challenge to hydrogenate phenol to cyclohexanone with high per pass selectivity towards cyclohexanone in combination with high per pass conversion of phenol during the entire life-time of a catalyst batch, while maintaining the production rate at a high level and the production costs at a low level.

Although the preparation of cyclohexanone from phenol has been known for many decades and ways to improve known preparation methods have been investigated thoroughly over the years, presently known industrial processes, which are generally of a continuous nature, still suffer from drawbacks.

In particular, in known continuous phenol hydrogenation processes, a decrease in per pass selectivity towards cyclohexanone in time is a problem. In addition, the heterogeneous hydrogenation catalyst slowly loses activity, resulting in a reduced per pass conversion of phenol. This results in a reduced cyclohexanone production rate in time. This is not only disadvantageous because stable production rates cannot be maintained, but also results in higher conversion costs for the dehydrogenation of produced cyclohexanol to cyclohexanone. Therefore, there exists a need to compensate the decrease in selectivity and activity of the heterogeneous hydrogenation catalyst in a phenol hydrogenation reaction.

It is a further object of the present invention to provide a method for preparing cyclohexanone in a continuous process that may serve as an improvement to known methods, in particular a method that overcomes one or more drawbacks of a known method, such as that referred to above.

It is a further object to provide a method for preparing cyclohexanone that produces less by-product(s), especially less cyclohexanol, that has to be converted to cyclohexanone by high temperature catalyzed dehydrogenation, whereby undesired by-product(s), like, e.g., benzene, are formed.

In addition, it is an object of the invention to provide a method of controlling a phenol hydrogenation reaction which requires no manual measurements and no manual adjustment of process parameters.

More in particular, it is an object of the invention to prepare cyclohexanone with improved cyclohexanone selectivity while maintaining a target cyclohexanone production rate compared to a conventional process that was operated in the same production facility. This results in benefits, such as a more environmentally friendly process, and a cheaper process compared to a conventional process operated in the same production facility. One or more further objects may become apparent from the remainder of the description.

The present inventors have discovered a process that solves the above-mentioned objects or at least mitigates them to a large extent.

SUMMARY OF THE INVENTION

The present invention provides a process for controlling an exothermic vapor phase hydrogenation of phenol that is catalyzed by a palladium comprising catalyst in an industrial scale hydrogenation reactor, wherein the reactor is charged with an ingoing mixture comprising phenol and hydrogen and from which is discharged an outgoing mixture comprising cyclohexanone, cyclohexanol, phenol and hydrogen, and wherein at least part of the hydrogen in the outgoing mixture is recycled to the ingoing mixture, wherein cyclohexanone is recovered from the outgoing mixture and at least part of the hydrogen from the outgoing mixture is discharged as a purge flow, and wherein at least part of the heat of reaction of the phenol hydrogenation reaction is transferred to a coolant via indirect heat exchange, wherein the process comprises the steps of: A) measuring the pressure in the reactor in an inline measurement;

B) measuring the concentration of hydrogen in the outgoing mixture in an inline measurement;

C) measuring the concentration of phenol in the outgoing mixture in an inline measurement;

D) measuring the concentration of cyclohexanol in the outgoing mixture in an inline measurement;

D2) optionally, measuring the concentration of cyclohexanone in the outgoing mixture in an inline measurement;

E) comparing the pressure in the reactor in step A) with a first set point;

F) comparing the concentration of hydrogen obtained in step B) with a second set point;

G) comparing the concentration of phenol obtained in step C) with a third set point;

H) comparing the concentration of cyclohexanol obtained in step D) with a fourth set point;

H2) optionally, comparing the concentration of cyclohexanone obtained in step D2) with a fifth set point;

I) adapting by an automatic mode of operation at least one process parameter if the pressure in the reactor obtained in step A) deviates from the first set point;

J) adapting by an automatic mode of operation at least one process parameter if the concentration of hydrogen in the outgoing mixture obtained in step B) deviates from the second set point.

The inventors found that, due to aging of the catalyst system and other process disturbances, if optimal process conditions in terms of selectivity and conversion are desired, the pressure and the output of phenol hydrogenation reactors should be carefully monitored and if required process conditions (like, e.g., flow rate of the ingoing mixture, in particular phenol pressure and hydrogen pressure, flow rate of the purge flow, reactor temperature, concentration of gas for dilution) adapted immediately in order to maintain production under optimal conditions. The invention achieves this by installing inline measurements of process parameters that were found to be critical and coupling these to an automatic comparison and adaptation based on pre-defined set points and adaptation rules. As shown in the Examples, the invention allows the phenol hydrogenation reactor to be operated under optimal conditions over a period of months to years at a constant high production rate without any human intervention. This was surprising and could not be predicted based on the prior art.

In the prior art, phenol hydrogenation processes were controlled, if at all, by offline measurement (manual measurement) to determine the concentrations of individual components such as hydrogen, phenol, cyclohexanone or cyclohexanol in the gas mixture that is discharged from the phenol hydrogenation reactor. Offline measurement means a measurement performed through the manual manipulation of a measurement system in order to obtain measurements for any given point, like, e.g., pulling a sample from the process by an operator at a certain moment in time and injecting the collected material in a certain type of analyzer, which determines the concentration of a component in the sample. The outcome of these individual measurements is used by the operators of the plant to determine whether process conditions can remain unchanged or whether changes are required.

Performing offline measurements relies on the skill of an operator and therefore the results of such measurements can also vary depending on the skill of the operator. In addition, offline measurements are typically performed at a low frequency (e.g., a few times per week or day). The inline measurements (automatic measurements) employed according to the invention are much more precise due to the lacking human element and the ability to measure continuously or at least with a high frequency. Inline measurement allows constant or even continuous monitoring of a parameter without interaction of an operator, which enables high frequency of repetition of a measurement and a stable measurement that is not affected by the skill of the operator.

Using prior art processes, which did not employ inline measurement, it was very difficult to impossible to track precisely when unacceptable results began to appear due to a lack of sufficiently high measurement frequency and statistical data. The research that led up to the invention showed that it has a surprisingly strong impact when and how fast after becoming necessary, decisions regarding the modification of process conditions are taken. Incorrect and not timely decisions regarding modification of process conditions were found to negatively impact the costs of and quality of cyclohexanone production (including reduced life time of the hydrogenation catalyst). The invention overcomes these disadvantages by allowing for an immediate, automated reaction to changes in process conditions that are measured inline and compared to pre-defined set points. This not only has the advantage that no human intervention is required and that a much more consistent and optimized process control is achieved, but most importantly also to reduced costs (including increased life time of the hydrogenation catalyst) and increased product quality as optimal process conditions are maintained more or less continuously throughout the process.

It was surprising that the process of the invention yields high production rates of the phenol hydrogenation reactor over a long period of time and at low production costs with very limited human interaction. This is achieved by implementing a combination of several inline measurements of special process parameters and coupling these measurements to an automated decision making process that results in an advantageous automatic process control, in which process parameters are adapted automatically according to the teaching of the invention. Up to now it was assumed that this was not possible for a phenol hydrogenation reactor because automation results in reduced versatility compared to the flexibility and variety of tasks that a well-trained operator could do.

The features and components of the plant according to the invention and the steps of the process according to the invention are described in more detail in the following.

In the process of the invention, an industrial scale hydrogenation reactor for the production of cyclohexanone from phenol is continuously fed with both phenol and hydrogen gas. The total incoming feed comprising at least phenol and hydrogen gas is referred to herein as “incoming mixture”. The incoming mixture can enter the reactor as a pre-mixed mixture via one line or mix inside the reactor. In the hydrogenation reactor, phenol is catalytically converted in the vapor phase, under the influence of a palladium-comprising catalyst, mainly into cyclohexanone and to a minor part into cyclohexanol. At least part of the heat of the exothermic phenol hydrogenation reaction is transferred to a coolant via indirect heat exchange. The product stream leaving the hydrogenation reactor comprises next to the desired product cyclohexanone and the by-product cyclohexanol also unconverted phenol and hydrogen. The product stream leaving the hydrogenation reactor, as well as any (further processed) streams derived therefrom is referred to herein collectively as “outgoing mixture”. At least part of the unconverted hydrogen in the outgoing mixture is recycled to the ingoing mixture and at least part of it is purged. Cyclohexanone is recovered as main product from the outgoing mixture.

According to the invention, the phenol hydrogenation reactor includes an inline measurement device that allows for the measurement of the pressure in the reactor in step A). The pressure can be measured by various (electrical) devices that are well-known to the skilled person. Especially, strain gauges are often applied for pressure measurements. Piezoresistive strain gauges (also called piezoresistors), which use the change in electrical resistance of a material when stretched to measure the pressure, are among the most common types of pressure sensors. The location of the inline pressure measurement can vary in the hydrogenation reactor. A very convenient location, which is also preferred according to the invention, is at the inlet of the reactor before the feed stream contacts the catalyst. For example, in a shell and tube type hydrogenation reactor, the inline pressure measurement can be located in the hood of the reactor before the feed stream entered the catalyst-filled tubes. In this way, the measured pressure corresponds to the total pressure of the feed stream charged to the shell and tube type hydrogenation reactor. But another possibility would be to have the inline pressure measurement at the outlet of the hydrogenation reactor or at another place within the reactor. The optimal set point for the pressure will vary slightly depending on the location of the pressure measurement as there usually is a drop in pressure when the reaction mixture passes through the catalyst. If the pressure is measured downstream of the catalyst, it may also change as the catalyst ages.

In the scope of the present invention the term “inline measurement” refers to an automatic measurement via a permanently installed device, so in general without the action of an operator being necessary, which is in contrast to offline measurement (also often called “manual measurement”).

Analysis can be performed by continuous or by discontinuous methods. Continuous analysis methods are used by placing a sensor directly in, e.g., a reactor, in a line with process fluid, or any another piece of equipment, or in a small side-stream that is withdrawn from, e.g., a reactor, a line with process fluid, or any another piece of equipment. Discontinuous implies that samples are withdrawn from, e.g., a reactor, in a line with process fluid, or any another piece of equipment, and injected into an analytical instrument. Continuous analyses give a continuous response signal, while discontinuous methods give a response at distinct, separate "points in time".

In the scope of the present invention the term “set point” refers to a value for an essential variable, or process value of a system. Deviation of such a variable from its set point is one basis for process control. Action might only take place if the measured value exceeds or falls below a given set point. In other cases an action is desired if the measured value deviates from the set point by a minimum amount (deadband).

In step E) of the process of the invention, the pressure in the reactor that was measured in step A) is compared with a first set point. Persons skilled in the art will know the optimal pressure of the industrial scale hydrogenation reactor that they are operating. It is related to the chemistry, catalyst type and reactor setup (e.g., length of tubes). If the optimal pressure set point is not known, this can be easily determined by performing test runs with different reactor pressures and determining the reactor pressure that provides optimal phenol conversion and selectivity towards cyclohexanone. In case the pressure is too high, the selectivity goes down. In case the pressure is too low, the throughput and therefore yield of cyclohexanone is too low. The targeted pressure, i.e. , the first set point, should always be super-atmospheric and preferably is in the range of from 0.1 to 1 MPa, more preferably from 0.15 to 0.6 MPa and most preferably from 0.2 to 0.5 MPa.

In step I) of the process of the invention, at least one process parameter is adapted by automatic operation if the pressure in the reactor deviates from the first set point. As explained above, the pressure can be too high or too low and should be monitored carefully and frequently. In practice the first set point can also be a pressure range around an optimal set point. For example, the range can be +/- 0.1 MPa, +/- 0.05 MPa or +/- 0.01 MPa from a pre-determined set point lying in the ranges indicated in the previous paragraph.

Steps A), E) and I) are preferably carried out at least once per day, more preferably at least once per hour, and most preferably at least once per minute or even every 5-30 seconds to achieve an optimal control. Because they need to be carried out at this frequency to achieve optimal results, the invention foresees for the adaptation in step I) to be carried out by automatic operation.

Process parameters that can be adapted by automatic operation in step I) to return the pressure again near or to the set point in a particular effective manner include a) adjusting the flow rate of the ingoing mixture and/or b) adjusting the flow rate of the purge flow.

Adjusting the flow rate of the ingoing mixture can be achieved in particular by adjusting the flow of the hydrogen-containing stream that enters the reactor and/or the flow of the phenol-containing stream that enters the reactor if these components enter the reactor via separate streams. If the pressure in the reactor is determined in step E) to be higher than the first set point, the flow rate of the ingoing mixture can be automatically reduced in step I). Conversely, if the pressure in the reactor is determined in step E) to be lower than the first set point, the flow rate of the ingoing mixture can be automatically increased in step I).

As explained above, in the plant or process according to the invention, at least part of the hydrogen from the outgoing mixture is discharged as a purge flow. This purge flow is an important way to regulate not only the concentration of hydrogen (relevant for step J) described below), but also the pressure in the reactor in step I. The flow rate of the purge stream can be adjusted by way of a valve that regulates the flow rate of the purge stream. If the pressure in the reactor is determined in step E) to be higher than the first set point, the flow rate of the purge stream can be automatically increased in step I). Conversely, if the pressure in the reactor is determined in step E) to be lower than the first set point, the flow rate of the purge stream can be automatically decreased in step I). In this way, the pressure can be returned again near or to the set point in a particular effective manner by simply opening or closing a valve.

The term “phenol hydrogenation reactor” as used herein takes on the usual meaning in the field. A phenol hydrogenation reactor is an enclosed volume in which hydrogenation of phenol takes place. The conversion of phenol to reaction products in a phenol hydrogenation reactor may be complete or just partial. The phenol hydrogenation reactor can be operated as a single reactor. Alternatively, multiple phenol hydrogenation reactors can be operated in parallel (e.g., to increase the total phenol hydrogenation capacity) and/or in series (e.g., to increase the overall conversion of phenol).

The hydrogenation reactor may in particular be any type of reactor suitable for hydrogenation of phenol. In particular, the reactor may be selected from fixed bed reactors, slurry reactors, shell and tube heat exchange reactors with catalyst in tubes and with generation of steam, and any other suitable type of reactor.

Preferably, the vapor phase hydrogenation of phenol according to the invention is carried out in a shell and tube heat exchange reactor and most preferably with water as coolant, resulting in the generation of steam. More preferably, the vapor phase hydrogenation of phenol according to the invention is carried out in a vertical shell and tube heat exchange reactor.

In one embodiment, the phenol hydrogenation reactor is a shell and tube heat exchange reactor with catalytic material arranged in the tubes. Preferably, it is a vertical shell and tube heat exchange reactor. The tubes are fed with a gaseous mixture comprising hydrogen and phenol and a coolant circulates externally around the tubes.

The term “heat exchanger” as used herein is a device for transferring heat from one stream to another. A heat exchanger may be direct (wherein the streams are mixed) or indirect (wherein the streams remain separated by a dividing wall). All heat exchangers referred to herein are indirect heat exchangers. An indirect heat exchanger comprises at least two chambers with a dividing wall. Heat is transferred from a stream in a first chamber, through the dividing wall, to a stream in a second chamber. Each chamber independently may have a long pathway, and a large surface area to volume ratio to facilitate heat transfer. Indirect heat exchangers are well-known to the person of skill in the art. Examples of indirect heat exchangers suitable for the present invention are shell and tube, plate, and tubular type heat exchangers.

In the scope of the present invention the expression “at least part of the heat of reaction of the phenol hydrogenation reaction is transferred to” refers to the heat that is released by the reaction of phenol with hydrogen and is transferred via a wall of a heat exchanger. The reaction of phenol with hydrogen is an exothermal (heat producing) reaction.

In the scope of the present invention the term “at least part” in the aforementioned expression means that at least 10%, even more preferably at least 50%, and most preferably at least 85% of the heat that is released by the reaction of phenol with hydrogen is transferred via a wall of a heat exchanger to the coolant.

The number of reactor tubes in the (vertical) shell and tube heat exchange reactor that is preferred according to the invention is typically more than 5. Preferably, it is more than 10. More preferably, it is more than 25. The number of reactor tubes is typically less than 100,000. Preferably, it is less than 50,000. More preferably, it is less than 20,000. Typically, in the process of the present invention, the number of reactor tubes in said (vertical) vertical shell and tube heat exchange reactor is from 25 to 20,000.

The length of a reactor tube in the (vertical) shell and tube heat exchange reactor that is preferred according to the invention is typically more than 0.25 m. Preferably, it is more than 0.5 m. More preferably, it is more than 1.0 m. The length of a reactor tube is typically less than 24.0 m. Preferably, it is less than 12.0 m. More preferably, it is less than 9.0 m. Typically, in the process of the present invention, the length of a reactor tube in said (vertical) vertical shell and tube heat exchange reactor is from 1.0 to 9.0 m, which allows good conversion of phenol while not requiring extensive amounts of catalyst. Preferably, all reactor tubes in the shell and tube heat exchange reactor have (almost) the same length.

The internal diameter of the shell of the shell and tube heat exchange reactor is typically more than 50 mm. Preferably, it is more than 100 mm. More preferably, it is more than 200 mm. The internal diameter of the shell of the (vertical) shell and tube heat exchange reactor is typically less than 10 m. Preferably, it is less than 8 m. More preferably, it is less than 6 m. Typically, in the process of the present invention the internal diameter of the shell and tube heat exchange reactor is from 0.2 to 6 m, which allows good reactor capacities while still being able to be transported. The internal diameter of the reactor tubes of the shell and tube heat exchange reactor is typically more than 2 mm. Preferably, it is more than 5 mm. More preferably, it is more than 10 mm. The internal diameter of the reactor tubes is typically less than 500 mm. Preferably, it is less than 250 mm. More preferably, it is less than 120 mm. Typically, in the process of the present invention the internal diameter of the reactor tubes in said shell and tube heat exchange reactor is from 10 to 120 mm. Preferably, all reactor tubes in the shell and tube heat exchange reactor have (almost) the same internal diameter.

In one embodiment, the temperature of the ingoing mixture comprising phenol and hydrogen that is charged to the hydrogenation reactor is more than 85 °C. Preferably, it is more than 100 °C. More preferably, it is more than 125 °C. This allows optimal conversion rates. To avoid damage of the catalyst, decomposition and the formation of by-products, the temperature of the ingoing mixture should be below 300 °C, preferably below 250 °C, and more preferably, it is below 220 °C. Most preferred it is from 125 °C to 220 °C.

Temperatures, as stated herein, can be measured by various (electrical) devices. Most commonly used sensors to measure temperature in the chemical industry are resistance thermometers (also called resistance temperature detectors; RTDs) and thermocouples in applications below 600 °C. Pt100 type sensors, which fall into the group of RTD’s, are the most commonly used.

In the tubes in the shell and tube heat exchange reactor heat is generated due to the hydrogenation of phenol, thereby heating up the mixture of components in the tubes. Heat of the mixture of components is removed by indirect cooling with a coolant.

The temperature of the mixture comprising cyclohexanone and cyclohexanol that is discharged from the hydrogenation reactor should be more than 60 °C. Preferably, it is more than 80 °C. More preferably, it is more than 100 °C. To avoid damage of the catalyst, decomposition and the formation of by-products, the temperature of the mixture comprising cyclohexanone and cyclohexanol that is discharged from the hydrogenation reactor should be less than 260 °C. Preferably, it is less than 240 °C. More preferably, it is less than 220 °C. Best results are achieved if the temperature of the mixture comprising cyclohexanone and cyclohexanol that is discharged from the shell and tube heat exchange reactor is from 100 to 220 °C. The lower temperature boundaries are determined by the reaction rate in the phenol hydrogenation reaction. At low temperatures the reaction rates will be low and as a consequence large amounts of catalyst and large sized reactors will be required. In addition the temperature of the heated coolant will be low and as a consequence many potential applications of the absorbed heat are excluded. The upper temperature boundaries are determined inter alia by the stability of the catalyst. In addition, at high temperatures, more undesired by-products will be formed (e.g., tar, and benzene) due to decomposition of phenol, thereby lowering the yield of the phenol hydrogenation.

According to the invention, the temperature in the hydrogenation reactor is adjusted by the temperature of the coolant used in the indirect heat exchange and the temperature of the incoming mixture. The higher the reactor temperature, the higher are the reactions rates, but above a certain temperature threshold also the selectivity goes down and more undesired by-products are formed.

In this regard it should be noted that the actual temperature in the reactor cannot be measured or reported with accuracy. The temperature in the catalytic bed is not constant: it varies both in location and in time. For example, the local temperature inside the catalytic bed in a shell and tube heat exchange reactor that is used for the hydrogenation of phenol is a result of the heat that is generated by the hydrogenation of phenol and the heat that is transferred to the coolant. The temperature varies both in longitudinal direction and in radial direction of the catalytic bed. In general, the temperature in longitudinal direction of the catalytic bed goes through a maximum (‘peak’). The shape of the temperature profile and the position of the maximum temperature in the bed is depending on many process parameters, like e.g., flow rates of the feed, concentration of both phenol and hydrogen in the feed, the type of catalyst and the condition of the catalyst and the temperature of the coolant outside the catalytic bed (outside the tubes in case of a shell and tube heat exchange reactor). As time proceeds, the location of the maximum temperature in the bed shifts gradually from a location near the entrance of the gaseous feed in the bed to a location near the exit of the gaseous products from the bed. In general, the exact temperature is not known at every location inside a catalytic bed. However, a global indication can be obtained by measuring the temperature at various heights in a catalytic bed by installing a series of temperatures sensors (e.g., Pt100 type sensors) in the longitudinal direction of a catalytic bed.

In a particularly preferred embodiment of the invention, the temperature of the coolant is therefore used to control the hydrogenation reaction of phenol. In case the temperature of the coolant is too low then too much heat will be absorbed by the coolant and the exothermic hydrogenation reaction will only proceed to a smaller extent in the reactor. In case the temperature of the coolant is too high, then not enough heat will be absorbed by the coolant and the exothermic hydrogenation reaction will locally overheat the catalytic bed and thereby deteriorating the catalyst and forming undesired by-products. The inventors found that good results are achieved if the temperature of the coolant is maintained in the temperature ranges stated above for the mixture comprising cyclohexanone and cyclohexanol that is discharged from the reactor.

The temperature of the coolant is one of the process parameters that may be adapted (preferably automatically) in steps I) to L) of the invention. It is a particularly useful process parameter to adapt in steps K), L) and/or L2) in response to a finding in steps G), H) or H2) that the concentration of phenol or cyclohexanol in the outgoing mixture exceeds the respective third and fourth set points or if the concentration of cyclohexanone obtained in step D2) is lower than the fifth set point. In case the concentration of phenol in the outgoing mixture exceeds the third set point, the reactor temperature should be increased in order to increase conversion to cyclohexanone. However, care must be taken as a too high increase in temperature and conversion can negatively affect selectivity towards cyclohexanone. In case the concentration of cyclohexanol in the outgoing mixture exceeds the fourth set point, the reactor temperature should be reduced in order to reduce conversion to cyclohexanol.

Typically, in a shell and tube heat exchange reactor wherein phenol is hydrogenated, heat is removed by a coolant. Aqueous or organic solvents or mixtures thereof can be applied as coolant. Preferably, in the process of the present invention, water is applied as coolant in the shell and tube heat exchange reactor. By absorbing heat, the coolant that is charged to the shell and tube heat exchange reactor is heated up, and preferably evaporated (partly or completely) or a combination thereof. Preferably, at least 10 wt.% of the coolant that is charged to the shell and tube heat exchange reactor is evaporated. More preferably, at least 50 wt.%. Even more preferably, at least 90 wt.%. Typically, in the process of the present invention, water is applied as coolant in the shell and tube heat exchange reactor and more than 90 wt.% of the water evaporated.

Preferably, water is used as the coolant in the indirect heat exchange employed in the phenol hydrogenation reactor. This has the advantage not only of good obtainability and costs of the coolant, but also that steam can be produced in the heat exchanger. The co-generated steam has the advantage that it can be applied for heating purposes. Preferably, the generated steam is applied in the process for the production of cyclohexanone by hydrogenation of phenol. Preferably, the generated steam is applied for the evaporation of phenol that is charged to the phenol hydrogenation reactor (e.g., as heating source in a phenol evaporation section). More preferably, the generated steam is applied for the recovery of cyclohexanone from the outgoing mixture of the phenol hydrogenation reactor (e.g., by driving a reboiler of a distillation column).

When water is used as a coolant, then according to a further preferred embodiment, boiling water is used as the coolant. Not only is steam produced in this way, but also the hydrogenation reactor temperature can be easily set by the temperature of boiling water that is used to absorb heat of the exothermic gas-phase hydrogenation. The boiling temperature of the water can be conveniently set by adjusting the pressure of the boiling water. Adjusting the pressure of the boiling water is a particularly convenient and accurate way of adjusting the temperature of the coolant. Typically, the pressure of the steam, in case water is used as coolant, is measured and is compared with a pre-determined set point. The control system can be programmed to automatically open one or more valves that increase the flow of the outgoing steam stream in response to a determination that the pressure of the steam is too high, or to automatically close one or more valves that decrease the flow of the outgoing steam stream in response to a determination that the pressure of the steam is too low.

The inventors have found that the pressure of the boiling water can be adjusted within the range of from 0.1 MPa to 5 MPa, depending on the amount of cooling and reactor temperature that is desired. Preferably, in the process of the invention, the pressure of the boiling water is adjusted to a value of at least 0.1 MPa, more preferably at least 0.15 MPa. Typically, in the process of the invention, the pressure of the boiling water is adjusted to a value less than 5 MPa, more preferably less than 1.5 MPa. Preferably, in the process of the invention, the pressure of the boiling water is adjusted within the range of from 0.15 MPa to 1.5 MPa.

The capacity of the phenol hydrogenation reactor is typically selected based on the volume of phenol charged to the phenol hydrogenation reactor.

The process of the invention uses an industrial scale hydrogenation reactor for the production of cyclohexanone. This means a phenol hydrogenation reactor that can be charged with several thousands of tons of phenol per year (kilotons per annum; kta). In an embodiment of the invention, the industrial scale phenol hydrogenation reactor can be or is charged with more than 10,000 tons of phenol per annum (10 kta). The charging of phenol to the industrial scale phenol hydrogenation reactor is not bound to maximum value. Preferably, the charging of phenol to the industrial scale phenol hydrogenation reactor is from 20 kta to 350 kta phenol. More preferably, the charging of phenol to the industrial scale phenol hydrogenation reactor is from 25 kta to 250 kta phenol. The mode of operation of the phenol hydrogenation reactor can be selected from fed-batch operation and continuous operation. In a fed-batch mode of operation, hydrogen is charged over time to the phenol hydrogenation reactor that was charged in advance with phenol and the reaction products remain in the reactor until the end of the run. In a continuous mode of operation both hydrogen and phenol are charged continuously to the phenol hydrogenation reactor and the reaction products and unconverted feed-stocks are discharged continuously from the phenol hydrogenation reactor. Preferably, the industrial scale phenol hydrogenation reactor is operated in a continuous mode.

In the scope of the present invention the expression “at least part of the hydrogen in the outgoing mixture is recycled to the ingoing mixture” refers to hydrogen that is present in the outgoing mixture of the industrial scale phenol hydrogenation reactor. This hydrogen was not consumed by the vapor phase hydrogenation of phenol that is catalyzed by a palladium comprising catalyst in the industrial scale hydrogenation reactor.

In the scope of the present invention the term “at least part” in the aforementioned expression preferably means that at least 10%, even more preferably at least 50%, and most preferably at least 85% of the hydrogen in the outgoing mixture is recycled to the ingoing mixture.

In step B) of the process according to the invention, the concentration of hydrogen in the outgoing mixture is measured via an inline measurement. “Concentration of hydrogen in the outgoing mixture” can mean the hydrogen concentration in any hydrogen-containing stream downstream of the phenol hydrogenation reactor as this will be proportional to the hydrogen concentration in the mixture directly exiting from the reactor. In practice, a particularly convenient location for measuring the concentration of hydrogen in the outgoing mixture is in the line that transports the gaseous purge stream (see Example 1 below, line [4] in Fig. 2), in the line that recycles part of the hydrogen in the outgoing mixture to the ingoing mixture (line [22] in Fig. 2), in the line that transports the compressed gas mixture directly existing the compression section (line [32] in Fig. 2), in the line directly exiting the phenol hydrogenation reactor (line [29] in Fig. 2), or in the lines exiting the optional heat exchanger or gas/liquid separator (lines [30] or [31] in Fig. 2) . The streams in lines [32], [22] and [4] in Fig. 2 are compressed and contain only minor amounts of phenol, cyclohexanol and cyclohexanone. They mainly contain hydrogen and inert. Thus, these locations are of particularly advantage for performing an inline measurement of the concentration of hydrogen in the outgoing mixture. If not defined otherwise, the term “concentration” as used herein refers to molar ratios (mol%). Hydrogen concentrations can be measured by various technologies. Chromatographic methods (especially gas chromatography (GC) are commonly used to separate the individual components in samples of mixtures of these components, which are then quantified by a detector (e.g., thermal conductivity detectors (TCDs), also known as a katharometers). The detector signal is processed by a controller and calculates the hydrogen concentration. For continuous measurements of hydrogen concentrations in gaseous mixtures with a limited amount of components (calibrated), thermal conductivity detectors are commonly directly applied (without chromatographic pre-treatment).

In step F) of the process of the invention, the concentration of hydrogen obtained in step B) is compared with a second set point. Persons skilled in the art will know the hydrogen concentration that should be maintained in the industrial scale hydrogenation reactor that they are operating to obtain optimal conversion of phenol and selectivity towards cyclohexanone. If the hydrogen concentration in the reactor is too high, selectivity towards cyclohexanone decreases, as more cyclohexanol is formed (see reaction schemes and stoichiometries depicted above). If the hydrogen concentration in the reactor is too low, the conversion of phenol will be too low. The optimal set point for the hydrogen concentration in the outgoing mixture may vary depending on the age of the catalyst and other process parameters. For example, with a very fresh catalyst, it is often recommended to temper the reaction rate by operating with very little hydrogen being present in the outgoing mixture. When the catalyst is aging, then more hydrogen needs to be added to obtain sufficient conversion of the phenol.

In the mixture directly exiting the hydrogenation reactor (line [29] in Fig. 2, same for line [30] in Fig. 2), the second set point for the optimal hydrogen concentration can be in the range of from 0.1 to 40 mol%, in particular from 1 to 30 mol%, based on the total moles of components present in the mixture. If the gas composition in the line that recycles part of the hydrogen in the outgoing mixture to the ingoing mixture is measured to determine the H2 concentration of the outgoing mixture, the second set point can be in the range of 1 to 100 mol% (if no gas for dilution, e.g., inerts are added, the recycle can be almost 100 mol% H2), preferably from 2 to 85 mol%, more preferably from 3 to 70 mol%, and most preferably from 5 to 50 mol%. The lower values apply when fresh (or regenerated) catalyst is present, the upper ones take into account aging of the catalyst. The gas composition in terms of molar ratios is the same in lines [31], [32], [22] and [4] depicted in Fig. 2 (temperature and pressure are not) and thus the concentration of hydrogen can be measured in and the aforementioned ranges apply to all of these locations. In step J) of the process of the invention, at least one process parameter is adapted by automatic operation if the concentration of hydrogen in the outgoing mixture deviates from the second set point. As explained above, the concentration of hydrogen can be too high or too low, and changes depending on the state of the catalyst and other process parameters. It should therefore be monitored carefully and frequently. In practice the second set point can also be a concentration range around an optimal set point. For example, the range can be +/- 25%, +/- 10%, +/- 1% or +/- 0.5 % from a pre-determined set point lying in the ranges indicated in the previous paragraph.

Steps B), F) and J) are preferably carried out at least once per day, more preferably at least once per hour, and most preferably at least once per minute or even every 5-30 seconds to achieve an optimal control. Because they need to be carried out at this frequency to achieve optimal results, the invention foresees for the adaptation in step J) to be carried out by automatic operation.

Process parameters that can be adapted by automatic operation in step J) to return the hydrogen concentration again near or to the set point in a particular effective manner include a) adjusting the flow rate of the ingoing mixture and/or b) adjusting the flow rate of the purge flow.

Adjusting the flow rate of the ingoing mixture in this context means in particular adjusting the flow of the hydrogen-containing stream that enters the reactor if the components hydrogen and phenol enter the reactor via separate streams. If the hydrogen concentration in the outgoing mixture is determined in step F) to be higher than the first set point, the flow rate of the ingoing mixture can be automatically reduced in step J). Conversely, if the hydrogen concentration in the outgoing mixture is determined in step F) to be lower than the first set point, the flow rate of the ingoing mixture can be automatically increased in step J).

As explained above, in the plant or process according to the invention, at least part of the hydrogen from the outgoing mixture is discharged as a purge flow. This purge flow is an important way to regulate the concentration of hydrogen. The main purpose of the purge stream is to purge inerts (gas for dilution). Increasing the flow of the purge stream results in lower inerts concentration and thus higher hydrogen concentration (and vice versa). The flow rate of the purge stream can be adjusted by way of a valve that regulates it. If the hydrogen concentration in the outgoing mixture is determined in step F) to be higher than the first set point, the flow rate of the purge stream can be automatically decreased in step J). Conversely, if the hydrogen concentration in the outgoing mixture is determined in step F) to be lower than the first set point, the flow rate of the purge flow can be automatically increased in step J). In this way, the hydrogen concentration in the outgoing mixture can be returned again near or to the set point in a particular effective manner by simply opening or closing a valve.

According to a preferred embodiment of the invention, the ingoing mixture further comprises a gas for dilution. The gas for dilution should be an inert gas such as methane and/or nitrogen. As it is inert, the gas for dilution is present both in the ingoing and in the outgoing mixture. Herein, the gas for dilution will sometimes simply be referred to as “inerts”. The skilled person often likes to operate the hydrogenation process at a fixed pressure, independent of the changes in the amounts (and associated pressures) of the phenol and hydrogen in the feed streams, which may also need to be adapted according to the process control implemented by the invention. For example, if the combined partial pressure of phenol and hydrogen that are fed in the reactor equals to 0.21 MPa, then the partial pressure of the inerts that are fed in the reactor can be 0.09 MPa to reach the desired pressure set point of 0.30 bar (in case no other components are fed to the reactor). As hydrogen is consumed in the hydrogenation reactor, the ratio of hydrogen to inerts is changed by the reaction in the recycled gas stream that contains unreacted hydrogen and inerts as compared to the ingoing mixture. To maintain a sufficient reaction rate, the inventors have found that it is possible to play with the amount of inerts vs. excess of hydrogen. Accordingly, a decrease in the inerts concentration leads to an increase in hydrogen concentration and therefore in an increased reaction rate.

In a further preferred embodiment, next to the hydrogen in the outgoing mixture that is recycled to the ingoing mixture also at least part of the gas for dilution that is present in the outgoing mixture is recycled to the ingoing mixture. This offers a particular convenient way to minimize the amount of fresh gas for dilution that must be fed to the ingoing mixture. By employing a gas for dilution in the ingoing mixture, the molar ratio of this gas for dilution to hydrogen gas is higher in the purge than in the incoming mixture, This offers particular convenient possibilities for regulating the hydrogen concentration. In steps C) and D) of the process of the invention, the concentrations of phenol and cyclohexanol, respectively, are measured in an inline measurement. Optionally, the process of the invention can comprise an additional step D2) in which the concentration of cyclohexanone is measured in an inline measurement. Phenol, cyclohexanol and cyclohexanone can all be measured with the same inline measurement device. Thus, steps C), D) and optionally D2 can be carried out in one step. Phenol, cyclohexanol and cyclohexanone concentrations can be measured by various technologies. Chromatographic methods (especially gas chromatography (GC), are commonly used to separate the individual components in samples of mixtures of these components, which are then quantified by a detector (e.g., thermal conductivity detectors (TCDs), also known as a katharometers). The detector signal is processed by a controller and calculates the various concentrations.

The location of the inline measurement of the concentration of phenol, cyclohexanol and cyclohexanone in the outgoing mixture can be at various positions downstream of the hydrogenation reactor. The measurement can be performed directly in the mixture directly exiting the phenol hydrogenation reactor (line [29] in Fig. 2), or in the lines exiting the optional heat exchanger or gas/liquid separator (lines [30] or [5] in Fig. 2). A particular convenient location for the inline measurement of the phenol, cyclohexanol and/or cyclohexanone concentration in the outgoing mixture is in the liquid comprising phenol, cyclohexanol and cyclohexanone that is present in the gas-liquid separation section [H] or exits therefrom (line [5] in Fig. 2). In both positions a calibrated inline gas chromatograph (GC) can be installed. The concentrations of phenol, cyclohexanol and cyclohexanone, as used herein, are expressed as mol%, wherein the sum of phenol, cyclohexanol and cyclohexanone in the analyzed mixture is taken as 100 mol%.

The set point for the concentration of phenol that is used for comparison in step G) of the process of the invention and also that for the concentration of cyclohexanol in step H) of the process of the invention should each independently be in the range of 0 to 20 mol%, preferably from 0 to 12 mol%, more preferably from 0 to 10 mol% or below 10 mol%, always based on the sum of phenol, cyclohexanol and cyclohexanone in the analyzed mixture as 100 mol%. The converse applies to the fifth set point that relates to the concentration of cyclohexanone, which can be calculated as [anone] = 100 mol% - [anol] - [phenol], as the sum of the concentrations of these three components is taken as 100 mol% for indicating their respective concentration set points herein. Thus, the concentration of cyclohexanone, i.e. , the fifth set point, preferably is in the range of 80 to 100 mol%, preferably from 88 to 100 mol%, more preferably from 90 to 100 mol% or above 90 mol%, always based on the sum of phenol, cyclohexanol and cyclohexanone in the analyzed mixture as 100 mol%.

As it is generally desired to have as little cyclohexanol (high selectivity to cyclohexanone) and as little unconverted phenol (high conversion) as possible in the outgoing reaction mixture, the third and fourth set points are such that only require an adaptation of process parameter(s) in steps K) and L) of the invention if the concentrations measured in steps C) and D) are higher than the respectively applicable set point. Similarly, as it is generally desired to have as much cyclohexanone (high selectivity to cyclohexanone and high conversion of phenol) as possible in the outgoing reaction mixture, the fifth set point is such that it only requires an adaptation of process parameter(s) in step L2) of the invention if the concentrations measured in step D2) are lower than the applicable set point. This is in contrast to the set points referred to in steps E), F), I) and J) of the process of the invention, which require adaptation if the measured value is below or above the set point.

A further difference between the set points referred to in steps E), F), I) and J) and those referred to in steps G), H), K), L) and L2) of the process of the invention is that the former require very frequent attention, with inline measurements and comparisons to the set points being recommended in the ranges of at least once per day, once per hour, once per minute or every 5 to 30 seconds. By contrast, the phenol and cyclohexanol concentrations were found out to require adaptations only much less frequently. The inventors found that in many instances, it is sufficient to only measure, compare and possibly adapt process parameters based on the results in the range of once per day or once per week. This is also the reason, why an automated decision making process, while still desirable, is not required for steps K) and/or L) of the process of the invention. Only according to a preferred embodiment, the at least one process parameter that is adapted in steps K), L), and/or L2) is adapted by an automatic mode of operation.

In step K) of the process of the invention, at least one process parameter is adapted if the concentration of phenol in the outgoing mixture deviates from the third set point. In step L) of the process of the invention, at least one process parameter is adapted if the concentration of cyclohexanol in the outgoing mixture deviates from the fourth set point. In step L2) of the process of the invention, at least one process parameter is adapted if the concentration of cyclohexanone obtained in step D2) is lower than the fifth set point. Steps K), L) and/or L2) can be optional, for example, when the catalyst is still fresh or regenerated. The inventors found steps K), L) and/or L2) to be particularly important for scenarios where the catalyst starts to age, which occurs after weeks to months. Catalyst aging results in a (slight) decrease in conversion (resulting in an increased concentration of phenol in the outgoing mixture) and a decrease in selectivity (resulting in an increased concentration of cyclohexanol in the outgoing mixture). Normally, catalyst aging is a slow process and as a consequence the period of time between start-up with a fresh or regenerated catalyst bed and replacement or regeneration of the aged catalyst bed ranges from several weeks up to several years. The actual figures are depending on many parameters, including quality of the fresh phenol, quality of the fresh hydrogen, water dosing, desired specifications of product(s), type of catalyst and plant lay-out. However, due to process disturbances (like unusual high concentrations of CO in the fresh hydrogen) aging of the catalyst sometimes also is much faster than under normal conditions and as a consequence the period of time between the start of the process disturbance and replacement or regeneration of the aged (poisoned) catalyst bed might be reduced to just several days or even less. Also for these instances, steps K), L) and/or L2) can be beneficial in counteracting the negative effects of catalyst aging.

If the concentration of cyclohexanone is also determined (see step D2 above), the process of the invention can include a further step H2, which comprises comparing the concentration of cyclohexanone obtained in step D2 with a fifth set point and, optionally, a further step L2) which comprises adapting at least one process parameter if the concentration of cyclohexanone obtained in step D2) is lower than the fifth set point. The fifth set point, i.e. , the desired concentration of cyclohexanone in the outgoing mixture is from 75 to 100 mol%, preferably from 85 to 100 mol% or 90 to100 mol%, and most preferably higher than 95 mol%, 98 mol% or 99 mol%, each time based on the sum of phenol, cyclohexanol and cyclohexanone in the analyzed mixture as 100 mol%.

An amount of deviation from the set point that is still deemed allowable without triggering an adaptation in steps K), L) or L2) can be defined by the skilled person according to the invention (e.g., the range can be +/- 25%, +/- 10%, +/- 1% or +/- 0.5 % from a pre-determined set point lying in the ranges indicated above).

Steps C), G) and K) as well as steps D), H) and L) are preferably carried out at least once per month, more preferably at least once per week, and most preferably at least once per day, hour or minute.

Process parameters that can be adapted by automatic operation in steps K) and L) to return the phenol and cyclohexanol concentrations, respectively, again below the set point in a particular effective manner include a) adjusting the temperature of the coolant, b) adjusting the hydrogen concentration in the outgoing mixture, c) adjusting the concentration of the gas for dilution in the outgoing mixture, and d) combinations thereof.

It is well-known that the reactor temperature influences conversion of phenol and selectivity towards cyclohexanone in a phenol hydrogenation reactor. As explained above in connection with the heat exchanger, adjusting the temperature of the coolant (in particular via the pressure of the boiling water if the latter is used as the coolant) is a particular convenient way to adjust the reactor temperature. Typically, the temperature of the coolant (and thereby the reactor temperature) should be increased in order to increase conversion to cyclohexanone. However care must be taken as too high increase in temperature and conversion can negatively affect selectivity towards cyclohexanone.

The term “adjusting the hydrogen concentration in the outgoing mixture” as used herein in most instances will mean “adjusting the hydrogen concentration in the ingoing mixture”, in particular by increasing the amount of hydrogen charged to the process. Increasing the amount of hydrogen that is fed into the reactor, increases the hydrogen concentration and hydrogen partial pressure in the reactor, which increases the reaction rate, i.e. , conversion. Care should be taken, however, as a too high increase in the hydrogen concentration and associated conversion into cyclohexanol decreases the desired selectivity towards cyclohexanone. A particularly convenient way of adjusting the hydrogen concentration in the outgoing mixture is by adjusting the second set point that is used for comparison in step F) and for adaptation of process parameters in step J). This is exemplified in Example 2 below. In this way, the hydrogen concentration in the outgoing mixture can also be indirectly adjusted by way of process parameters.

Adapting the concentration of gas for dilution in the outgoing mixture as used herein in most instances will mean adapting the concentration of gas for dilution in the ingoing mixture. For example, a decrease of the concentration of the gas for dilution in the outgoing mixture is realized by decreasing the concentration of the gas for dilution in the ingoing mixture. A particular convenient way of achieving this is by adding less fresh hydrogen or adding more recycle gas mixture. As described above, adjusting (in particular decreasing) the concentration of gas for dilution is a convenient way of adjusting (in particular increasing) the concentration of hydrogen and thereby the reaction rate and associated conversion.

In a preferred embodiment, the process parameter(s) that can be adapted by automatic operation in step K) to return the phenol concentration again below the set point is selected from a) increasing the temperature of the coolant, b) increasing the hydrogen concentration in the outgoing mixture, c) decreasing the concentration of the gas for dilution in the outgoing mixture, and d) combinations thereof.

In a preferred embodiment, the process parameter(s) that can be adapted by automatic operation in step L) to return the cyclohexanol concentration again below the set point is selected from a) decreasing the temperature of the coolant, b) decreasing the hydrogen concentration in the outgoing mixture, c) increasing the concentration of the gas for dilution in the outgoing mixture, and d) combinations thereof.

If the process includes a step L2) of adapting at least one process parameter if the concentration of cyclohexanone obtained in step D2) is lower than the fifth set point, the process parameters to be adapted are the same as described above for the third set point relating to the concentration of phenol. The reason is that the means described for decreasing the concentration of phenol in the outgoing mixture are all means of increasing conversion to cyclohexanone and the selectivity towards cyclohexanone, respectively. Thus, these means of adapting process parameters can also be used in step L2) in order to increase the concentration of cyclohexanone in the outgoing mixture above the fifth set point. For example, an increase in the temperature of the coolant can be used in step L2) as shown in Example 4 below. A further way to overcome too high phenol and too low cyclohexanone concentrations is to reducing the phenol concentration in the ingoing mixture, i.e. , by feeding less phenol to the hydrogenation reactor. This, however, will reduce the output of the plant.

Next to the process described above, the invention also provides an industrial-scale phenol hydrogenation plant configured to carry out the process according to the invention. This includes the components that have been described in connection with the process steps above, but also a control system (e.g., a software) that can carry out steps E) to J), and preferably also steps K) and L) and/or D2) and L2) of the invention. The control system carries out not only the decision making process in steps E) to J), and preferably also steps K) and L) and/or D2) and L2) of the invention, but preferably also exerts as output of the decision making process the adaptation of the process parameters explained above. For example, the control system can be programmed to automatically open one or more valves that decrease the flow of the purge stream and/or increase the flow of the ingoing mixture in response to a determination in steps E) and/or F) that the pressure in the reactor or the concentration of hydrogen is too low. In this way, the plant is configured to carry out steps I) and J) of the process. This is just one example of possible inputs and outputs for the control system. The detailed description of the process of the invention above applies mutatis mutandis also to the corresponding parts and features of the plant of the invention.

The plant of the invention at least comprises a hydrogenation reactor comprising a palladium comprising catalyst, at least one inlet through which the reactor can be charged with an ingoing mixture comprising phenol and hydrogen, at least one outlet from which an outgoing mixture comprising cyclohexanone, cyclohexanol, phenol and hydrogen can be discharged, a line for recycling at least part of the hydrogen in the outgoing mixture to the ingoing mixture, a recovery section in which cyclohexanone can be recovered from the outgoing mixture, a purge through which at least part of the hydrogen from the outgoing mixture can be discharged as a purge flow, an indirect heat exchanger, which can transfer at least part of the heat of reaction of the phenol hydrogenation reaction to a coolant.

The plant of the invention is characterized in particular by the presence of inline measurement devices that can measure the following parameters: the pressure in the hydrogenation reactor; the concentration of hydrogen in the outgoing mixture; the concentration of phenol in the outgoing mixture; the concentration of cyclohexanol in the outgoing mixture; and optionally, the concentration of cyclohexanone in the outgoing mixture.

The plant also comprises an automated control system that can carry out steps E) to J), and preferably also steps K) and L) and/or D2), H2) and L2) of the invention as described above.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 illustrates schematically a plant and process for the preparation and recovery of cyclohexanone from phenol according to the invention.

FIG. 2 illustrates schematically an embodiment of a phenol hydrogenation reaction section [A] according to the invention. The phenol hydrogenation reaction section [A] comprises a hydrogen purification section [D], a phenol evaporation section [E], a phenol hydrogenation section [F], a heat exchange section [G], a gas-liquid separation section [H] and a compression section [J],

In the description of the drawings, the following reference numerals will be used:

[A] phenol hydrogenation reaction section

[B] separation and purification section

[C] cyclohexanol dehydrogenation reaction section

[D] hydrogen purification section [E] phenol evaporation section

[F] phenol hydrogenation section

[G] heat exchange section

[H] gas-liquid separation section

[J] compression section

[I] hydrogen comprising stream

[2] fresh phenol stream

[3] stream comprising recovered phenol

[4] gaseous purge stream

[5] mixture comprising phenol, cyclohexanone, cyclohexanol and hydrogen

[6] stream comprising cyclohexanol and cyclohexanone

[7] stream rich in cyclohexanone

[8] first stream rich in cyclohexanol

[9] second stream rich in cyclohexanol

[10] light components stream

[11] heavy components stream

[12] produced hydrogen

[21] hydrogen gas

[22] recycled hydrogen gas

[23] hydrogen-containing stream

[24] combined phenol stream

[25] stream of gaseous components

[26] not evaporated components

[27] water

[28] feed stream

[29] gas mixture comprising cyclohexanone, cyclohexanol, phenol and hydrogen

[30] mixture comprising hydrogen gas and liquid phenol, cyclohexanone and cyclohexanol

[31] gas mixture comprising hydrogen

[32] compressed gas mixture

DETAILED DESCRIPTION OF THE DRAWINGS

A plant and corresponding process for the preparation and recovery of cyclohexanone from phenol is schematically shown in FIG. 1. Such a process usually consists of two sections with an optional third section. All three sections are described.

Cyclohexanone is catalytically prepared by reacting phenol with hydrogen in phenol hydrogenation reaction section [A], Cyclohexanone is recovered in separation and purification section [B], In optional cyclohexanol dehydrogenation reaction section [C], the undesired by-product cyclohexanol is catalytically converted into cyclohexanone and hydrogen.

Phenol hydrogenation reaction section [A] comprises at least one hydrogenation reactor that is charged with a hydrogen comprising stream via line [1], a fresh phenol stream via line [2] and optionally a stream comprising recovered phenol via line [3] and may comprise additional equipment. The hydrogen-comprising stream can optionally comprise inerts (e.g., nitrogen and/or methane) as gas for dilution and optionally carbon monoxide. Optionally, inerts (e.g., nitrogen and/or methane) and optionally carbon monoxide are charged separately to phenol hydrogenation reaction section [A] (not shown in FIG. 1). The catalytic phenol hydrogenation according to the invention takes place in a vapor phase process. From the phenol hydrogenation reaction section [A], a gaseous purge stream comprising hydrogen and optionally inerts (e.g., nitrogen and/or methane) and optionally carbon monoxide is discharged via line [4], and a liquid mixture comprising phenol, cyclohexanone, cyclohexanol and hydrogen is discharged via line [5], Hydrogenation reaction section [A] is described in more detail in connection with Fig. 2 below as are ways to obtain the purge stream [4] and product stream [5],

The mixture comprising phenol, cyclohexanone, cyclohexanol and hydrogen is supplied via line [5] to separation and purification section [B], Optionally, a stream comprising cyclohexanol and cyclohexanone is supplied via line [6] from cyclohexanol dehydrogenation reaction section [C] to separation and purification section [B], Usually, separation and purification section [B] comprises one or more distillation columns. In separation and purification section [B] a stream rich in cyclohexanone and optionally a stream rich in cyclohexanol and optionally a stream comprising phenol are recovered. From this separation and purification section [B] a stream rich in cyclohexanone is discharged as primary product via line [7], This stream rich in cyclohexanone might be used as a solvent or as a raw material, e.g., for the production of e-caprolactam, caprolactone or adipic acid (not shown in FIG. 1). Optionally, a first stream rich in cyclohexanol is discharged via line [8] to outside the process. This first stream rich in cyclohexanol might be used as a solvent or as a raw material, e.g., for the production of adipic acid (not shown in FIG. 1). Optionally, a second stream rich in cyclohexanol is discharged via line [9] to cyclohexanol dehydrogenation reaction section [C]. Optionally, a stream comprising recovered phenol is discharged via line [3] to phenol hydrogenation reaction section [A], Optionally, a light components stream, optionally comprising benzene, cyclohexane and water, is discharged via line [10], Optionally, a heavy components stream comprising phenol and components with boiling points higher than phenol is discharged via line [11],

The optional cyclohexanol dehydrogenation reaction section [C] comprises at least one dehydrogenation reactor and one or more heat exchangers. In the cyclohexanol dehydrogenation reaction section [C] cyclohexanol is catalytically converted into cyclohexanone and hydrogen. In general, the dehydrogenation of cyclohexanol is a gas phase reaction that is performed at elevated temperatures. Optionally, a product stream comprising cyclohexanol and cyclohexanone is charged via line [6] to separation and purification section [B], The hydrogen that is produced in cyclohexanol dehydrogenation reaction section [C] is discharged as produced hydrogen via line [12], Optionally, the produced hydrogen in cyclohexanol dehydrogenation reaction section [C] is supplied to phenol hydrogenation reaction section [A] (not shown in FIG. 1). Optionally, the produced hydrogen in cyclohexanol dehydrogenation reaction section [C] is supplied to another hydrogen consuming process (not shown in FIG. 1). Optionally, the produced hydrogen in cyclohexanol dehydrogenation reaction section [C] is supplied to a heat generation unit (not shown in FIG. 1).

In FIG. 2 a scheme of an embodiment according to the invention of a phenol hydrogenation reaction section [A] (area enclosed by dashed line) is given.

A hydrogen comprising stream is charged via line [1] to the phenol hydrogenation reaction section [A], The hydrogen comprising stream can, e.g., originate from a naphtha cracker, a methane reformer, gasification of coal process, an electrolysis process or from a dehydrogenation process. Typically, the hydrogen comprising stream contains inert components like nitrogen and/or methane. Optionally, inert components like nitrogen and/or methane can be added as gas for dilution to the hydrogen comprising stream (not shown in FIG. 2). In case the hydrogen comprising stream contains components that are harmful to the hydrogenation, like CO and/or H2S, then a hydrogen gas purification step is required. The presence of these harmful components in the charged hydrogen can be temporary, e.g., due to upset conditions in the hydrogen gas production unit, or can be permanent. In such a hydrogen purification step the harmful impurities can be converted into inert components or removed from the hydrogen comprising stream.

For this purpose, the hydrogen comprising stream is charged via line [1] to the optional hydrogen purification section [D], Prior to charging to hydrogen purification section [D], the temperature of the hydrogen comprising stream can be modified in a heat exchange section to a temperature required in hydrogen purification unit [D] (not shown in FIG. 2]). In general, in such a heat exchange section, the temperature of the hydrogen comprising stream is increased.

Hydrogen purification section [D] may comprise one or more catalysts for the conversion of harmful components into inert components and/or one or more adsorbents for the removal of harmful components. Preferably, hydrogen purification section [D] comprises a catalyst for conversion of CO and/or an adsorbent for the removal of H2S. Hydrogen gas is discharged from hydrogen purification section [D] via line [21], The hydrogen gas in line [21] and the recycled hydrogen gas in line [22] are combined, thereby forming a hydrogen-containing stream that is charged via line [23] to phenol evaporation section [E], Hydrogen purification section [D] can comprises one or more reaction and/or adsorption units that are operated in parallel and/or in series. Optionally, water, e.g., in the form of vapor, can be charged to the hydrogen gas before, in or after the hydrogen purification section [D] (not shown in FIG. 2). The hydrogen purification section [D] can be absent or by-passed (not shown in FIG. 2), for example, when quality of the hydrogen comprising stream in line [1] is sufficient.

A fresh phenol stream is charged via line [2] optionally, but preferably, a stream comprising recovered phenol is charged via line [3], thereby forming a combined phenol stream that flows via line [24], The stream comprising recovered phenol that is charged via line [3] is discharged from the separation and purification section [B] that is shown in Fig. 1. Optionally, the stream comprising recovered phenol is not combined with the fresh phenol stream that is charged via line [2] and is charged separately to phenol evaporation section [E], or is absent (not shown in FIG. 2). Optionally, prior to charging to phenol evaporation section [E], the temperature of the combined phenol stream that flows via line [24] is increased in a heat exchange section that comprises one or more heat exchangers that can be operated in parallel and/or in series (not shown in FIG. 2).

In phenol evaporation section [E] virtually all components that entered via line [24] are evaporated. A stream of gaseous components is discharged from phenol evaporation section [E] via line [25], Components that entered via line [24] and are not evaporated are discharged (either continuously or batch wise) from phenol evaporation section [E] as not evaporated components via line [26], These are usually only small amounts. In general phenol evaporation section [E] involves heating of the incoming feed, in particular heating by steam. Preferably, phenol evaporation section [E] comprises a device, e.g., a wire-mesh demister, for removal of entrained droplets from the stream of gaseous components that is discharged. Phenol evaporation section [E] comprises one or more evaporators that are operated in parallel and/or in series. Optionally, the stream of gaseous components in line [25] is temperature-adjusted in a heat exchange section that comprises one or more heat exchangers that are operated in parallel and/or in series (not shown in FIG. 2). In general, in this heat exchange section, the stream of gaseous components is raised in temperature. Optionally, some water, e.g., in the form of vapor, is charged via line [27] to the stream in line [25], thereby forming a feed stream that flows via line [28] and is charged to the phenol hydrogenation section [F],

Phenol hydrogenation section [F] consists of one or more hydrogenation reactors that are operated in series and/or in parallel. In phenol hydrogenation section [F], cyclohexanone and cyclohexanol are obtained in a continuous process by heterogeneously catalyzed hydrogenation of phenol. A gas mixture comprising cyclohexanone, cyclohexanol, phenol and hydrogen is discharged from phenol hydrogenation section [F] via line [29],

The gas mixture comprising cyclohexanone, cyclohexanol, phenol and hydrogen that is discharged from phenol hydrogenation section [F] via line [29] is cooled down in heat exchange section [G], whereby at least a fraction of the phenol, cyclohexanone and cyclohexanol is condensed. Heat exchange section [G] comprises one or more heat exchangers that are operated in parallel and/or in series. As coolant, one or more coolants might be used in order to cool down the gas mixture comprising cyclohexanone, cyclohexanol, phenol and hydrogen that is discharged from phenol hydrogenation section [F] via line [29], At least one of the applied coolants should have a temperature sufficiently low to condense at least a fraction of the phenol, cyclohexanone and cyclohexanol present in the gas mixture comprising cyclohexanone, cyclohexanol, phenol and hydrogen that is discharged from phenol hydrogenation section [F] via line [29], Optionally, the cooling down of the gas is performed in several stages in series, whereby, optionally, in each stage another coolant is used. Optionally, at least one of the coolants is a process flow from phenol hydrogenation reaction section [A], from separation and purification section [B] and/or from cyclohexanol dehydrogenation reaction section [C] (not shown in FIG. 2). Optionally, one of the coolants is the recycled hydrogen gas that is charged via line [22] (not shown in FIG. 2). A mixture comprising hydrogen gas and liquid phenol, cyclohexanone and cyclohexanol is discharged from heat exchange section [G] via line [30] and is charged to gas-liquid separation section [H], Gas-liquid separation section [H] comprises one or more gas-liquid separators that are operated in parallel and/or in series. A liquid mixture comprising phenol, cyclohexanone, cyclohexanol and hydrogen is discharged from gas-liquid separation section [H] via line [5] and is charged to separation and purification section [B] shown in FIG. 1. A gas mixture comprising hydrogen is discharged from gas-liquid separation section [H] via line [31] and is charged to compression section [J], Any type of gas-liquid separator may be applied in gas-liquid separation section [H], however particularly good results are achieved when vertical vessels are used in which the liquid settles to the bottom of the vessel. In general, gas-liquid separation section [H] comprises a device, e.g., a wire-mesh demister, for removal of entrained droplets from the stream of gaseous components that is discharged.

Compression section [J] comprises one or more devices to compress a gas mixture. These can be operated in parallel and/or in series. The compressed gas mixture is discharged from compression section [J] via line [32] and is split in a recycled hydrogen gas that is transported via line [22] and in a gaseous purge stream that is transported via line [4], Optionally, the recycled hydrogen gas that is transported via line [22] is charged to a heat exchange section, where it is heated (not shown in FIG. 2). The recycled hydrogen gas, that is optionally heated, is discharged via line [22] and is then combined with the hydrogen gas in line [21], When carrying out the process of the invention, the gaseous purge stream that is transported via line [4] usually comprises hydrogen and one or more inert components, like nitrogen and/or methane and optionally carbon monoxide. This gaseous purge stream discharged via line [4] can, e.g., be used as fuel (not shown in FIG. 2). Optionally, prior to discharging from phenol hydrogenation reaction section [A], the gaseous purge stream that is transported via line [4] is charged to a heat exchange section, wherein the gaseous purge stream is cooled down (not shown in FIG. 2). Optionally, liquid formed in this heat exchange section is charged to gasliquid separation section [H] (not shown in FIG. 2). EXAMPLES

The following examples serve to explain the invention in more detail, in particular with regard to certain forms of the invention. The examples, however, are not intended to limit the present disclosure.

The Examples were carried out in the phenol hydrogenation reaction section [A] of a chemical plant for the preparation and recovery of cyclohexanone from phenol which was very similar to the embodiment of the invention depicted in FIG. 1. The chemical plant for the preparation and recovery of cyclohexanone from phenol further comprised a separation and purification section [B], in which high grade cyclohexanone is obtained by distillative separation from components with higher and lower boiling points, and a cyclohexanol dehydrogenation reaction section [C], in which cyclohexanol is converted into cyclohexanone. The phenol hydrogenation reaction section [A] comprised of a hydrogen purification section [D], a phenol evaporation section [E], a phenol hydrogenation section [F], a heat exchange section [G], a gas-liquid separation section [H], and a compression section [J], which was very similar to the embodiment of the invention depicted in FIG. 2.

A hydrogen comprising stream comprising about 94 vol.% hydrogen, about 6 vol.% nitrogen, traces of CO and traces of H2S was charged to hydrogen purification section [D] via line [1], In hydrogen purification section [D] CO was catalytically converted into CH4 and H2S was removed by an adsorbent. Under normal operation conditions, both the CO content and the H2S content of the hydrogen gas that was discharged from hydrogen purification section [D] via line [21] were each well below 1 ppm. The hydrogen containing stream that was obtained by combining recycled hydrogen gas transported via line [22] and the freshly fed hydrogen gas transported via line [21] was charged to phenol evaporation section [E] via line 23.

Fresh phenol stream charged via line [2] and a stream comprising recovered phenol charged via line [3] were combined and charged to a steam heated phenol evaporation section [E], In phenol evaporation section [E] virtually all phenol is evaporated. A small flow of not evaporated components was discharged from phenol evaporation section [E] via line [26] and the remainder was discharged as a stream of gaseous components via line [25], Steam was added via line [27] to the stream of gaseous components in line [25] so that the amount of water to the amount of phenol in line [28] was about 1 wt.%, based on the weight of phenol. In this way, the feed stream (transported via line [28]) of phenol hydrogenation section [F] was obtained.

The phenol hydrogenation section [F] was equipped with one industrial scale, shell and tube type phenol hydrogenation reactor with an annual capacity of about 120 kta phenol that was operated in a continuous mode for phenol hydrogenation in the vapor phase. The tubes were filled with a supported hydrogenation catalyst: Pd/AhCh (0.9 wt.%) with 1 wt.% Na (as NaHCOs) added as promoter. The number of tubes was about 5600. The height of the catalyst bed in each reactor tube was about 2 m. The hourly mass feed flow rate (WHSV) was 5 (kg phenol/hr) I (kg catalyst). As coolant boiler feed water was charged to the volume outside the tubes for removal of reaction heat, whereby steam was produced. More than 85% of the heat of reaction was transferred to a coolant.

EXAMPLE 1

The shell and tube type hydrogenation reactor was filled with fresh catalyst. Starting 72 hours after start-up of the process, the following process conditions were recorded during a continuous production period of 7 months:

The pressure of the boiling water that was used as coolant was set at 0.48 MPa (the associated temperature was ca. 150 °C). The produced vapor was used as energy source of a reboiler of a distillation column.

The pressure in the reactor was measured by measuring the total pressure of the feed stream charged to the shell and tube type hydrogenation reactor in an inline measurement in the hood of the reactor before the feed stream entered the catalyst-filled tubes. The pressure sensor continuously gave an electrical signal, which was a measure of the magnitude of the pressure in the reactor. The pressure set point (first set point) was 0.38 MPa. In case the pressure deviated from this set point, the flow of the gaseous purge stream in line [4] was automatically adjusted so that the first set point was maintained (too high pressure initiated opening of a valve in line [4]; too low pressure initiated closure of a valve in line [4]).

The concentration of hydrogen in the outgoing mixture was continuously measured by an inline measurement of the gaseous purge stream in line [4] via a calibrated inline thermal conductivity detector (katharometer). The thermal conductivity detector continuously gave an electrical signal, which was a measure of the hydrogen concentration. The set point of the hydrogen concentration (second set point) was 4 mol%, wherein the sum of the concentrations of all gaseous components (mainly hydrogen and inert gas, such as methane and/or nitrogen) is always taken as 100 mol%. In case the hydrogen concentration deviated from this set point, the flow of the hydrogen gas in line [1] was automatically adjusted so that the second set point was maintained (too low hydrogen concentration initiated an increased flow of the hydrogen gas in line [1]; too high hydrogen concentration initiated a reduced flow of the hydrogen gas in line [1]).

The concentrations of phenol, cyclohexanol and cyclohexanone in the outgoing mixture were determined by an inline measurement of the liquid in the liquid filled bottom section of the vertical vessel-type gas-liquid separator in gas-liquid separation section [H] by a calibrated inline gas chromatograph (GC). The set points for the concentrations of phenol, cyclohexanol and cyclohexanone (third to fifth set point) of the liquid in the liquid filled bottom section of the vertical vessel-type gasliquid separator in gas-liquid separation section [H] were 3 mol%, 3 mol% and 95 mol%, respectively (the sum of the concentrations of phenol, cyclohexanol and cyclohexanone was always taken as 100 mol%). During the first 7 months after startup, the concentrations of phenol, cyclohexanol and cyclohexanone remained < 3 mol%, < 3 mol% and > 95 mol%, respectively (the sum of the concentrations of phenol, cyclohexanol and cyclohexanone was always taken as 100 mol%).

A gas mixture comprising cyclohexanone, cyclohexanol, phenol and hydrogen [29] was discharged from the phenol hydrogenation reactor and was cooled down in heat exchange section [G] to 40 °C. The resulting mixture comprising hydrogen gas and liquid phenol, cyclohexanone and cyclohexanol was discharged via line [30] and charged to a vertical vessel in gas-liquid separation section [H], From the bottom section of that vertical vessel, a liquid mixture was discharged via line [5] and charged to separation and purification section [B] (see FIG. 1). In separation and purification section [B] a stream rich in cyclohexanone (> 99 mol% cyclohexanone), a stream rich in cyclohexanol, a light components stream and a heavy components stream were obtained by distillative separation and purification. The stream rich in cyclohexanol was charged to cyclohexanol dehydrogenation reaction section [C] via line [9], In cyclohexanol dehydrogenation reaction section [C] cyclohexanol was catalytically dehydrogenated, whereby hydrogen and cyclohexanone were produced. The produced hydrogen was discharged via line [12] and a stream comprising cyclohexanol and cyclohexanone was discharged via line [6] and charged to separation and purification section [B], A gas mixture comprising hydrogen was discharged from the top section of the vertical vessel in gas-liquid separation section [H] via line [31] and was charged to a compressor in compression section [J], The compressed gas mixture comprising hydrogen that was discharged from the compressor in compression section [J] was split up in a gaseous purge stream that was transported via line [4] and a recycled hydrogen gas that was transported via line [22],

The results of Example 1 show that due to the combination of inline measurements of pressure in the reactor and hydrogen concentration in the outgoing mixture and automatic control according to the invention, the phenol hydrogenation reactor could be operated during a period of about 7 months at a constant high production rate without any human intervention.

EXAMPLE 2

The preparation and recovery of cyclohexanone from phenol that was carried out in the same chemical plant as described in Example 1 was continued after the initial period of 7 months.

A few days later, the concentrations of phenol in the liquid in the liquid filled bottom section of the vertical vessel-type gas-liquid separator in gas-liquid separation section [H] as determined by the calibrated inline gas chromatograph exceeded the set point that was set at 3 mol% (the sum of the concentrations of phenol, cyclohexanol and cyclohexanone was always taken as 100 mol%). As a reaction to this, the set point of the hydrogen concentration was increased by 1 mol% (from 4 mol% to 5 mol%). As a result of this action the concentration of phenol in the liquid in the liquid filled bottom section of the vertical vessel-type gas-liquid separator in gas-liquid separation section [H] as determined by the calibrated inline gas chromatograph (GC) decreased to below 3 mol%, while the concentrations of cyclohexanol and cyclohexanone remained < 3 mol% and > 95 mol%, respectively (the sum of the concentrations of phenol, cyclohexanol and cyclohexanone was always taken as 100 mol%). This new situation could be remained for more than about 1 month without any further adjustments of the set point of the hydrogen concentration. In the period up to 12 months after start of the run, the set point of the hydrogen concentration had to be increased several times by 1 mol% (until the value of the set point of the hydrogen concentration became 10 mol%), in order to adjust the concentration of phenol in the liquid in the liquid filled bottom section of the vertical vessel-type gas-liquid separator in gas-liquid separation section [H] to a value below 3 mol%.

A consequence of increasing the set point of the hydrogen concentration from 4 mol% to 10 mol% was that the flow of the gaseous purge stream in line [4] was increased over time.

The results of Example 2 show that adaptation of the concentration of hydrogen in the outgoing mixture is used to control the concentration of phenol in the outgoing mixture. Every time that the concentration of phenol in the outgoing mixture exceeded its set point of 3 mol%, the set point of the hydrogen concentration in the outgoing mixture was step-wise increased by 1 mol% (from 4 mol% to 10 mol%) and as a consequence, the concentration of hydrogen in the outgoing mixture was increased step wise from 4 mol% to 10 mol%. After every raise of the concentration of hydrogen in the outgoing mixture, the concentration of phenol in the outgoing mixture decreased again to below 3 mol%. Due to this control strategy, the constant high production rate of the phenol hydrogenation reaction section [A] of the chemical plant could be extended up to 12 months after start of run.

EXAMPLE 3

The preparation and recovery of cyclohexanone from phenol that was carried out in the same chemical plant as described in Example 2 was continued after the initial period of 12 months.

A few days later, the concentration of cyclohexanone in the liquid in the liquid filled bottom section of the vertical vessel-type gas-liquid separator in gasliquid separation section [H] as determined by the calibrated inline gas chromatograph (GC) fell below the set point that was set at 95 mol% (the sum of the concentrations of phenol, cyclohexanol and cyclohexanone was always taken as 100 mol%). As a reaction to this, the pressure of the boiling water that was used as coolant was raised by 0.02 MPa (to 0.50 MPa). This results in an associated temperature increase of about 1.5 °C to about 152 °C. As a result of this action, the concentration of cyclohexanone in the liquid in the liquid filled bottom section of the vertical vessel-type gas-liquid separator in gas-liquid separation section [H] as determined by the calibrated inline gas chromatograph increased again to above 95 mol%, while the concentrations of phenol and cyclohexanol remained each < 3 mol% (the sum of the concentrations of phenol, cyclohexanol and cyclohexanone was always taken as 100 mol%). This situation could be maintained for more than 2 months without any further adjustments of the pressure of the boiling water that was used as coolant.

The results of Example 3 show that by once adapting the pressure of the boiling water that was used as coolant if the concentration of cyclohexanone in the outgoing mixture was less than the set point for the concentration of cyclohexanone in the outgoing mixture, the constant high production rate of the phenol hydrogenation reaction section [A] of the chemical plant could be maintained by another 2 months.

EXAMPLE 4

The preparation and recovery of cyclohexanone from phenol that was carried out in the same chemical plant as described in Example 3 was continued after the initial period of 14 months. Due to market conditions the production rate of the phenol hydrogenation reaction section [A] had to be reduced by 30%.

Directly after reducing the phenol feed rate to the phenol hydrogenation reaction section [A], without adapting the set points, the concentration of cyclohexanol in the liquid in the liquid filled bottom section of the vertical vesseltype gas-liquid separator in gas-liquid separation section [H] as determined by the calibrated inline gas chromatograph (GC) exceeded the set point that was set at 3 mol% (the sum of the concentrations of phenol, cyclohexanol and cyclohexanone was always taken as 100 mol%). As a reaction to this, the set point of the hydrogen concentration (second set point) was decreased by 2 mol% (from 10 mol% to 8 mol%), to which the system reacted by decreasing the flow of the hydrogen gas in line [1] as explained in connection with Example 1 above. As a result of this action, the concentration of hydrogen in the ingoing and the outgoing mixture was decreased, which led to the concentrations of cyclohexanol in the liquid in the liquid filled bottom section of the vertical vessel-type gas-liquid separator in gas-liquid separation section [H] as determined by the calibrated inline gas chromatograph to decrease again to below 3 mol%, while the concentrations of phenol and cyclohexanone remained both < 3 mol% and > 95 mol%, respectively (the sum of the concentrations of phenol, cyclohexanol and cyclohexanone was always taken as 100 mol%). This situation could be maintained for more than one month without any further adjustments of any of the set points. The results of Example 4 show that by lowering the set point of the concentration of hydrogen in the outgoing mixture and as a result of that the concentration of hydrogen in the ingoing and the outgoing mixture, the concentration of cyclohexanol in the outgoing mixture could be reduced to a value below its set point of 3 mol%.

EXAMPLE 5

The preparation and recovery of cyclohexanone from phenol was carried out in the same chemical plant as described in Example 1 , except that now in case the pressure deviated from the set point, the flow of the gaseous purge stream in line [4] was not automatically adjusted, and that in case the hydrogen concentration deviated from the set point, the flow of the hydrogen gas in line [1] was not automatically adjusted.

Instead, in case the pressure deviated from the set point, the flow of the hydrogen gas in line [1] was automatically adjusted (too low pressure initiated an increased flow of the hydrogen gas in line [1]; too high pressure initiated a reduced flow of the hydrogen gas in line [1]). And, in case the hydrogen concentration deviated from the set point, the flow of the gaseous purge stream in line [4] was automatically adjusted (too high hydrogen concentration initiated a reduced flow of the gaseous purge stream in line [4]; too low hydrogen concentration initiated an increased flow of the gaseous purge stream in line [4]).

And again, during the first 7 months after start-up, the concentrations of phenol, cyclohexanol and cyclohexanone of the liquid in the liquid filled bottom section of a vertical vessel-type gas-liquid separator in gas-liquid separation section [H] remained < 3 mol%, < 3 mol% and > 95 mol%, respectively (the sum of the concentrations of phenol, cyclohexanol and cyclohexanone was always taken as 100 mol%).

The results of Example 5 show that due to the combination of inline measurements of pressure in the reactor and hydrogen concentration in the outgoing mixture and automatic control according to the invention, the phenol hydrogenation reactor could be operated during a period of about 7 months at a constant high production rate without any human intervention.