Process control for phenol hydrogenation.

Abstract

The present invention provides an improved method for controlling the exothermic vapor phase hydrogenation of phenol catalyzed by a palladium-containing catalyst in an industrial scale hydrogenation reactor, the reactor being charged with an inlet mixture and discharged from the reactor with an outlet mixture, and at least a portion of the hydrogen in the outlet mixture being recycled to the inlet mixture and at least a portion of the heat of reaction being transferred to a cooling medium, the method comprising the steps of in-line measuring and controlling certain process parameters. The present invention also provides a plant configured to carry out the method of the present invention.

Description

[Technical field]

[0001] The present invention relates to a plant and catalytic process for industrial scale phenol hydrogenation in which primarily cyclohexanone and cyclohexanol are formed. [Background technology]

[0002] Most cyclohexanone is consumed in the production of ε-caprolactam, an intermediate in the production of nylon 6. A mixture of cyclohexanone and cyclohexanol is primarily used in the production of adipic acid, which is converted to nylon 6,6. In addition, cyclohexanone can be used as an industrial solvent or as an activator in oxidation reactions. Cyclohexanone can also be used as an intermediate for the production of cyclohexanone resins.

[0003] In the 1930s, cyclohexanone production began on an industrial scale in parallel with the commercial production of ε-caprolactam, adipic acid, nylon 6 and nylon 6,6. Since then, cyclohexanone production has continued to increase, and today, annual production of cyclohexanone exceeds 7 million tons.

[0004] The three main commercial production routes for cyclohexanone and cyclohexanol are based on the oxidation of cyclohexane, the hydration of cyclohexene, and the hydrogenation of phenol, respectively. The first two routes produce large amounts of undesirable by-products and require large energy inputs. The third route, i.e., the hydrogenation of phenol to cyclohexanone, is carried out in either a "two-step" or "one-step" process. In the "two-step" process, phenol is first reacted with hydrogen to form cyclohexanol, e.g., by using a catalyst containing Ni, which is then dehydrogenated to obtain cyclohexanone. In the "one-step" process, phenol is directly and highly selectively hydrogenated to cyclohexanone. This "one-step" phenol hydrogenation process is known to combine high production yields with low energy demands. In the "one-step" process, cyclohexanone is prepared by catalytic hydrogenation from phenol in a phenol hydrogenation reactor, e.g., by using a catalyst containing Pt or Pd. During phenol hydrogenation, cyclohexanol is formed as the main by-product apart from trace amounts of other compounds. Phenol hydrogenation can be carried out in the liquid phase or in the vapor phase (often called gas phase). See, for example, "Cyclohexanol and Cyclohexanone", Kirk-Othmer Encyclopedia of Chemical Technology, e.g. 3rd Edition, 1979, Vol. 7, p. 410-416; I. Dodgson et al., "A low Cost Phenol to Cyclohexanone Process", Chemistry and Industry, December 1989, Vol.18, p. 830-833; or M T. Musser, "Cyclohexanol and Cyclohexanone", Ullmann's Encyclopedia of Industrial Chemistry, 7th Edition, 2007.

[0005] The (exothermic) hydrogenation of phenol to the two main products cyclohexanone and cyclohexanol proceeds according to the following stoichiometric equation: C6H5OH + 2H2 → C6H 10 O + heat Phenol Hydrogen Cyclohexanone C6H5OH + 3H2 → C6H 11 OH + heat Phenol Hydrogen Cyclohexanol It can be expressed as:

[0006] The (endothermic) dehydrogenation of cyclohexanol to cyclohexanone has the following stoichiometry: C6H 11 OH + heat → C6H 10 O + H2 Cyclohexanol Cyclohexanone Hydrogen It can be expressed as:

[0007] GB 890095 describes a process for preparing cyclohexanone from phenol by catalytic hydrogenation which involves passing gaseous phenol with hydrogen over a catalyst containing a metal belonging to the palladium group at a temperature below 250° C. The document states that the gaseous reactants may be diluted with an inert gas such as nitrogen, argon or propane.

[0008] GB 1316820 A describes a process for producing cyclohexanone by hydrogenation of phenol in the vapour phase in the presence of a catalyst consisting of a support on which is deposited a layer of palladium and an alkali metal carbonate as a promoter. The document teaches that the selectivity for cyclohexanone is influenced by the presence of a promoter.

[0009] GB 1332211A describes a process for producing cyclohexanone by catalytic hydrogenation of phenol in the vapour phase at temperatures between 75 and 250° C. in the presence of a catalyst containing 0.1-5% by weight of a platinum group metal supported on activated alumina and 5-50% by weight of an alkaline earth metal hydroxide and / or an alkaline earth metal oxide, based on the total weight of the catalyst.

[0010] WO2016075047A1 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 of steam per kg of cyclohexanone produced.

[0011] WO2011073233A1 describes a process for continuously preparing cyclohexanone, cyclohexanol or a mixture thereof by catalytically hydrogenating phenol fed to a reactor equipped with a supported hydrogenation catalyst containing a dopant. The document teaches that by feeding water to the reactor during the hydrogenation of phenol, an increased conversion of phenol and / or an increased selectivity to cyclohexanone and / or cyclohexanol can be achieved.

[0012] The combination of (very) high selectivity to cyclohexanone, and therefore (very) low selectivity to cyclohexanol, in the reactor section with (very) high one-pass conversion of phenol is highly desirable due to its economic advantages: smaller equipment for recovery of cyclohexanol and phenol from the product stream and cyclohexanol dehydrogenation, and reduced energy consumption for recovery of cyclohexanol and phenol from the product stream and cyclohexanol dehydrogenation. As a result, the inventors have realized that an industrial-scale phenol hydrogenation plant for cyclohexanone production should be operated in an optimal process window with high one-pass selectivity to cyclohexanone combined with high one-pass conversion of phenol. The research that led to the present invention has shown that the optimum combination of single pass selectivity to cyclohexanone and single pass conversion of phenol for a phenol hydrogenation plant depends on many variables, such as the origin, source and quality of the phenol and hydrogen feedstocks, catalyst type, process type, plant layout and desired production rate.

[0013] Indeed, an increase in the single-pass selectivity to cyclohexanol, and therefore a decrease in the single-pass selectivity to cyclohexanone, is usually observed with increasing single-pass conversion of phenol. In particular, at (very) high single-pass conversions of phenol, a sharp increase in the single-pass selectivity to cyclohexanol is observed, which makes it impossible to simultaneously achieve 100% single-pass conversion of phenol and 100% single-pass selectivity to cyclohexanone.

[0014] Catalyst systems used in phenol hydrogenation plants can be significantly affected by degradation due to catalyst aging. "Aging" is typically defined as catalyst deactivation, i.e., 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 few months to many years. Without wishing to be bound by theory, it is hypothesized that aging is caused by the deposition of organic compounds on the catalyst (often referred to as coking) and / or by modification of the crystal structure of the catalyst metal.

[0015] Generally, the cost of producing cyclohexanone by hydrogenation of phenol continues to rise as the catalyst ages (e.g., due to decreasing selectivity to cyclohexanone, decreasing single-pass conversion of phenol, decreasing productivity, and increasing hydrogen purge losses).

[0016] Generally, the end of reaction (EOR) of a given hydrogenation catalyst batch is determined by technical (eg reaching the maximum allowable temperature or reaching the maximum allowable pressure) or economic (see above) considerations.

[0017] To maximize catalyst life, regeneration can be carried out under carefully controlled (combustion) conditions to remove organic deposits on the catalyst. Such catalyst regeneration may optionally include a) heat treatment in a nearly inert (oxygen-free) atmosphere to remove carbon-containing volatile materials present on the catalyst, b) heat treatment in an oxygen-containing atmosphere to oxidize carbonaceous materials present on the catalyst, and c) cooling the catalyst after steps a) and / or b).

[0018] Maximizing catalyst life is highly desirable since catalyst replacement and regeneration are time consuming and expensive (eg, due to lost production) operations.

[0019] So far, it remains a challenge to hydrogenate phenol to cyclohexanone with high single-pass selectivity to cyclohexanone and high single-pass conversion of phenol over the life of the catalyst batch, while maintaining high levels of production rate and low levels of production costs.

[0020] The preparation of cyclohexanone from phenol has been known for decades and although avenues for improving the known preparation have been thoroughly investigated over the years, the currently known industrial processes, which are generally continuous, still suffer from shortcomings.

[0021] In particular, in known continuous phenol hydrogenation processes, it is problematic that the first-pass selectivity to cyclohexanone decreases over time. In addition, the heterogeneous hydrogenation catalyst gradually loses activity, resulting in a decrease in the first-pass conversion of phenol. This leads to a decrease in the cyclohexanone production rate over time. This is not only disadvantageous because a stable production rate cannot be maintained, but also leads to an increase in the conversion cost of dehydrogenating the produced cyclohexanol to cyclohexanone. Therefore, it is necessary to compensate for the decrease in selectivity and activity of the heterogeneous hydrogenation catalyst in the phenol hydrogenation reaction.

[0022] It is a further object of the present invention to provide a method for preparing cyclohexanone in a continuous process which may serve as an improvement over known methods, in particular a method which overcomes one or more of the disadvantages of known methods as mentioned above.

[0023] A further object is to provide a process for preparing cyclohexanone which produces fewer by-products, in particular less cyclohexanol which must be converted to cyclohexanone by high temperature catalytic dehydrogenation resulting in the formation of undesirable by-products such as benzene.

[0024] Additionally, it is an object of the present invention to provide a method for controlling a phenol hydrogenation reaction that does not require manual measurement and adjustment of process parameters.

[0025] More specifically, the object of the present invention is to prepare cyclohexanone with improved cyclohexanone selectivity while maintaining a target cyclohexanone production rate compared to conventional processes operated at the same production facility. This provides advantages such as a more environmentally friendly process and a cheaper process compared to conventional processes operated at the same production facility. One or more further objects may become apparent from the remaining description.

[0026] The present inventors have discovered a way to solve or at least significantly alleviate the above mentioned objectives. Summary of the Invention

[0027] The present invention relates to a method for controlling an exothermic vapor phase hydrogenation reaction of phenol catalyzed by a catalyst comprising palladium in an industrial scale hydrogenation reactor, the reactor being charged with an inlet mixture comprising phenol and hydrogen and discharging an outlet mixture comprising cyclohexanone, cyclohexanol, phenol and hydrogen, at least a portion of the hydrogen in the outlet mixture being recycled to the inlet mixture, cyclohexanone being recovered from the outlet mixture, at least a portion of the hydrogen from the outlet mixture being discharged as a purge stream, and at least a portion of the heat of reaction of the phenol hydrogenation reaction being transferred to a cooling medium by indirect heat exchange, the method comprising: A) measuring the pressure in the reactor by in-line measurement; B) measuring the hydrogen concentration in the exit mixture with an in-line measurement; C) measuring the phenol concentration in the exit mixture with an in-line measurement; D) measuring the cyclohexanol concentration in the exit mixture with an in-line measurement; D2) optionally measuring the cyclohexanone concentration in the exit mixture by in-line measurement; E) comparing the pressure in the reactor of step A) with a first set value; F) comparing the hydrogen concentration obtained in step B) with a second set value; G) comparing the phenol concentration obtained in step C) with a third set value; H) comparing the cyclohexanol concentration obtained in step D) with a fourth set value; H2) optionally comparing the cyclohexanone concentration obtained in step D2) with a fifth set value; I) performing an adaptation of at least one process parameter in an automatic operating mode if the pressure in the reactor obtained in step A) deviates from a first set value; J) adapting at least one process parameter in an automatic operating mode if the hydrogen concentration in the exit mixture obtained in step B) deviates from a second set value. The present invention provides a method comprising:

[0028] The inventors have 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 output of the phenol hydrogenation reactor should be carefully monitored and, if necessary, immediate adaptations should be made to the process conditions (e.g., flow rates of the inlet mixtures, especially the phenol pressure and hydrogen pressure, flow rates of the purge streams, reactor temperatures, dilution gas concentrations, etc.) to maintain production under optimal conditions. The present invention achieves this by introducing in-line measurements of process parameters found to be crucial and by combining this measurement with automatic comparison and adaptation based on predefined set points and adaptation rules. As shown in the examples, the present invention allows the phenol hydrogenation reactor to be operated under optimal conditions for months to years at a constant high production rate without human intervention. This is surprising and would not have been predicted from the prior art.

[0029] In the prior art, the phenol hydrogenation process was controlled, if at all, by off-line measurements (manual measurements) to measure the concentration of individual components such as hydrogen, phenol, cyclohexanone or cyclohexanol in the gas mixture discharged from the phenol hydrogenation reactor. By off-line measurements, we mean measurements made by manual manipulation of a measurement system to obtain a measurement for any given time, for example, an operator at a given moment withdrawing a sample from the process and injecting the withdrawn material into some sort of analyzer that measures the concentration of the component in the sample. The results of these individual measurements are used by the plant operator to determine whether the process conditions can remain unchanged or whether changes are required.

[0030] Performing offline measurements depends on the skill of the operator, and therefore the results of such measurements may vary depending on the skill of the operator. In addition, offline measurements are typically performed infrequently (e.g., several times a week or a day). In-line measurements (automated measurements) used by the present invention are much more accurate due to the absence of the human element and the ability to measure continuously or at least frequently. In-line measurements allow periodic or even continuous monitoring of parameters without operator intervention, which allows for frequent measurement repetition and stable measurements that are not affected by the operator's skill.

[0031] Using prior art processes without in-line measurements, due to the lack of sufficiently high measurement frequency and statistical data, it was extremely difficult to impossible to track exactly when unacceptable results began to appear. The research that led to the present invention showed that when and how quickly after the need for process condition correction decisions are made has a surprisingly strong impact. It was found that inaccurate and untimely decisions on process condition corrections have a negative impact on the cost and quality of cyclohexanone production, including shortening the life of the hydrogenation catalyst. The present invention overcomes these shortcomings by enabling immediate automatic reactions to change process conditions that are measured in-line and compared to predefined set points. This not only has the advantage that no human intervention is required and a much more consistent and optimized process control is achieved, but most importantly, it results in reduced costs (including extended hydrogenation catalyst life) and improved product quality, since optimal process conditions are almost continuously maintained throughout the process.

[0032] It was surprising that the process of the present invention produces high production rates in phenol hydrogenation reactors at low production costs over long periods with very limited human intervention. This is accomplished by implementing a combination of several in-line measurements of specific process parameters and linking these measurements to an automated decision process that provides advantageous automated process control where adaptation of process parameters is performed automatically in accordance with the teachings of the present invention. Up until now, this was thought to be impossible for phenol hydrogenation reactors because automation reduces versatility compared to the versatility and variety of tasks that can be performed by a well-trained operator.

[0033] The features and components of the plant according to the invention and the steps of the method according to the invention are described in more detail below.

[0034] In the process of the present invention, both phenol and hydrogen gas are continuously fed to an industrial-scale hydrogenation reactor for the production of cyclohexanone from phenol. The entire inlet feed, including at least phenol and hydrogen gas, is referred to herein as the "inlet mixture". The inlet mixture may enter the reactor in one line as a premixed mixture or may be mixed inside the reactor. In the hydrogenation reactor, phenol is catalytically converted in the vapor phase under the influence of a catalyst comprising palladium to mainly cyclohexanone and partly to cyclohexanol. At least a part of the heat of the exothermic phenol hydrogenation reaction is transferred to a cooling medium by indirect heat exchange. The product stream leaving the hydrogenation reactor contains unconverted phenol and hydrogen apart from the desired product cyclohexanone and by-product cyclohexanol. The product stream leaving the hydrogenation reactor and any (further processed) product streams derived therefrom are collectively referred to herein as the "outlet mixture". At least a portion of the unconverted hydrogen in the exit mixture is recycled to the entry mixture and at least a portion is purged. Cyclohexanone is recovered from the exit mixture as the primary product.

[0035] According to the invention, the phenol hydrogenation reactor is equipped with an in-line measuring device that allows the measurement of the pressure in the reactor in step A). ​​The pressure can be measured by various (electrical) devices well known to the skilled person. In particular, strain gauges are often used for pressure measurement. Piezoresistive strain gauges (also called piezo-resistive elements), which use the change in the electrical resistance of a material when stretched to measure pressure, are the most common type of pressure sensor. The location of the in-line pressure measurement can be various in the hydrogenation reactor. A very convenient location, which is also preferred for the present invention, is the inlet of the reactor before the feed stream comes into contact with the catalyst. For example, in a shell-and-tube hydrogenation reactor, the in-line pressure measurement can be located in the hood of the reactor before the feed stream enters the tube filled with catalyst. In this way, the measured pressure corresponds to the total pressure of the feed stream charged to the multi-tube hydrogenation reactor. However, another possibility would be an in-line pressure measurement at the outlet of the hydrogenation reactor or at another location in the reactor. The optimum setting for pressure will vary slightly depending on the location of pressure measurement since pressure usually drops as the reaction mixture passes through the catalyst, and if the pressure is measured downstream of the catalyst, may also change as the catalyst ages.

[0036] Within the scope of the present invention, the term "in-line gauging" refers to automated gauging by permanently installed devices and therefore generally does not require operator action, as opposed to offline gauging (often referred to as "manual gauging").

[0037] The analysis can be performed by continuous or discontinuous methods. Continuous analysis methods are used by placing the sensor directly in, for example, the reactor, in a line with the process fluid, or in any other equipment, or in a small side stream drawn from, for example, the reactor, the line with the process fluid, or any other equipment. Discontinuous means that the sample is drawn from, for example, the reactor, the line with the process fluid, or any other equipment, and injected into an analytical instrument. A continuous analysis gives a continuous response signal, while a discontinuous method gives responses at different, discrete "times".

[0038] Within the scope of the present invention, the term "setpoint" refers to a value relating to an essential variable or process value of a system. If such a variable deviates from its setpoint, it is one basis for process control. Action can be taken only if the measured value exceeds or falls below a given setpoint. In other cases, action is called for if the measured value deviates minimally (by a deadband) from the setpoint.

[0039] In step E) of the process of the present invention, the pressure in the reactor measured in step A) is compared with a first setpoint. A person skilled in the art will know the optimal pressure for the industrial scale hydrogenation reactor he is operating. The optimal pressure is related to the chemistry, the type of catalyst and the reactor configuration (e.g. tube length). If the optimal pressure setpoint is unknown, it can be easily determined by carrying out trial runs at various reactor pressures to determine the reactor pressure that provides the optimal phenol conversion and cyclohexanone selectivity. If the pressure is too high, the selectivity will decrease. If the pressure is too low, the throughput and therefore the cyclohexanone yield will be too low. The targeted pressure, i.e. the first setpoint, should always be above atmospheric pressure and is preferably in the range of 0.1-1 MPa, more preferably 0.15-0.6 MPa and most preferably 0.2-0.5 MPa.

[0040] In step I) of the method of the invention, an automatic adaptation of at least one process parameter is performed if the pressure in the reactor deviates from a first setpoint. As explained above, the pressure can be too high or too low and should be monitored carefully and frequently. In practice, the first setpoint may be a pressure range around the optimal setpoint. For example, the range may be ±0.1 MPa, ±0.05 MPa or ±0.01 MPa from a predefined setpoint lying within the range indicated in the previous paragraph.

[0041] Steps A), E) and I) are preferably carried out at least once a day, more preferably at least once an hour, and most preferably at least once a minute or even every 5-30 seconds to achieve optimal control. As the above steps need to be carried out at such frequencies to achieve optimal results, the present invention foresees that the adaptation in step I) is carried out automatically.

[0042] Process parameters that can be automatically adapted in step I) to bring the pressure back near or to the set point in a specific and effective manner include a) adjusting the flow rate of the inlet mixture and / or b) adjusting the flow rate of the purge stream.

[0043] The regulation of the flow rate of the inlet mixture can be achieved in particular by adjusting the inflow of the hydrogen-containing stream entering the reactor and / or the inflow of the phenol-containing stream entering the reactor (if these components enter the reactor in separate flow paths). If the pressure in the reactor is measured in step E) to be higher than a first set value, the flow rate of the inlet mixture can be automatically reduced in step I). Conversely, if the pressure in the reactor is measured in step E) to be lower than the first set value, the flow rate of the inlet mixture can be automatically increased in step I).

[0044] As explained above, in the plant or method according to the invention, at least a part of the hydrogen from the exit mixture is released as a purge stream. This purge stream is an important way of regulating the pressure in the reactor in step I) as well as the hydrogen concentration (related to step J) below). The flow rate of the purge stream can be adjusted by a valve regulating the flow rate of the purge stream. If the pressure in the reactor is measured in step E) to be higher than a first set value, the flow rate of the purge stream can be automatically increased in step I). Conversely, if the pressure in the reactor is measured in step E) to be lower than the first set value, the flow rate of the purge stream can be automatically decreased in step I). In this way, the pressure can be brought back near or to the set value again in a specific and effective manner by simply opening and closing a valve.

[0045] The term "phenol hydrogenation reactor" as used herein has its usual meaning in the art. A phenol hydrogenation reactor is an enclosed space in which the hydrogenation of phenol takes place. The conversion of phenol to reaction products in the phenol hydrogenation reactor may be complete or partial. The phenol hydrogenation reactor may be operated as a single reactor. Alternatively, multiple phenol hydrogenation reactors may be operated in parallel (e.g., to increase the total phenol hydrogenation capacity) and / or in series (e.g., to increase the overall phenol conversion).

[0046] The hydrogenation reactor may in particular be any type of reactor suitable for the hydrogenation of phenol, in particular it may be chosen from fixed bed reactors, slurry reactors, multi-tube heat exchange reactors with catalyst in the tubes and steam generation, and any other suitable type of reactor.

[0047] Preferably, the vapor phase hydrogenation of phenol according to the present invention is carried out in a multi-tubular heat exchange reactor, and most preferably using water as the cooling medium, resulting in the production of water vapor. More preferably, the vapor phase hydrogenation of phenol according to the present invention is carried out in a vertical multi-tubular heat exchange reactor.

[0048] In one embodiment, the phenol hydrogenation reactor is a multi-tube heat exchange reactor with the catalytic material arranged in the tubes. Preferably, it is a vertical multi-tube heat exchange reactor. The tubes are fed with a gaseous mixture containing hydrogen and phenol, and a cooling medium is circulated around the tubes on the outside.

[0049] The term "heat exchanger" as used herein is a device for transferring heat from one stream to another. Heat exchangers can be direct (streams are mixed) or indirect (streams remain separated by a partition). All heat exchangers referred to herein are indirect heat exchangers. An indirect heat exchanger comprises at least two chambers with a partition. Heat is transferred from a stream in the first chamber through the partition to a stream in the second chamber. Each chamber can independently have long passages and a large surface area to volume ratio to facilitate heat transfer. Indirect heat exchangers are well known to those skilled in the art. Examples of indirect heat exchangers suitable for the present invention are shell and tube, plate, and tube heat exchangers.

[0050] Within the scope of the present invention, the expression "at least a part of the heat of reaction of the phenol hydrogenation reaction is transferred to" refers to the heat released by the reaction of phenol with hydrogen and transferred through the walls of the heat exchanger. The reaction of phenol with hydrogen is an exothermic (heat-producing) reaction.

[0051] Within the scope of the present invention, the term "at least a portion" in the above expression means that at least 10%, even more preferably at least 50%, and most preferably at least 85% of the heat released by the reaction of phenol with hydrogen is transferred through the walls of the heat exchanger to the cooling medium.

[0052] The number of reactor tubes in a preferred (vertical) multi-tubular heat exchange reactor according to the present invention is typically more than 5. Preferably, the number is more than 10. More preferably, the number is more than 25. The number of reactor tubes is typically less than 100,000. Preferably, the number is less than 50,000. More preferably, the number is less than 20,000. Typically, in the process of the present invention, the number of reactor tubes in said (vertical) multi-tubular heat exchange reactor of the vertical type is between 25 and 20,000.

[0053] The length of the reactor tube of the preferred (vertical) multi-tubular heat exchange reactor according to the present invention is typically greater than 0.25 m. Preferably, the length is greater than 0.5 m. More preferably, the length is greater than 1.0 m. The length of the reactor tube is typically less than 24.0 m. Preferably, the length is less than 12.0 m. More preferably, the length is less than 9.0 m. Typically, in the process of the present invention, the length of the reactor tube of the (vertical) multi-tubular heat exchange reactor is between 1.0 and 9.0 m, which allows good phenol conversion while not requiring a large amount of catalyst. Preferably, all reactor tubes of the multi-tubular heat exchange reactor are (almost) the same length.

[0054] The inner diameter of the shell of the multi-tubular heat exchange reactor is typically greater than 50 mm. Preferably, the inner diameter of the shell is greater than 100 mm. More preferably, the inner diameter of the shell is greater than 200 mm. The inner diameter of the shell of the (vertical) multi-tubular heat exchange reactor is typically less than 10 m. Preferably, the inner diameter of the shell is less than 8 m. More preferably, the inner diameter of the shell is less than 6 m. Typically, in the process of the present invention, the inner diameter of the multi-tubular heat exchange reactor is 0.2 to 6 m, which allows the reactor to have good operating capacity and at the same time be transportable.

[0055] The inner diameter of the reactor tube of the multi-tubular heat exchange reactor is typically larger than 2 mm. Preferably, the inner diameter is larger than 5 mm. More preferably, the inner diameter is larger than 10 mm. The inner diameter of the reactor tube is typically smaller than 500 mm. Preferably, the inner diameter is smaller than 250 mm. More preferably, the inner diameter is smaller than 120 mm. Typically, in the reaction of the present invention, the inner diameter of the reactor tube of the multi-tubular heat exchange reactor is 10-120 mm. Preferably, all the reactor tubes of the multi-tubular heat exchange reactor have (almost) the same inner diameter.

[0056] In one embodiment, the temperature of the inlet mixture comprising phenol and hydrogen charged to the hydrogenation reactor is greater than 85° C. Preferably, the temperature is greater than 100° C. More preferably, the temperature is greater than 125° C. This allows for optimal conversion rates. To avoid catalyst attrition, decomposition and by-product formation, the temperature of the inlet mixture should be less than 300° C., preferably less than 250° C., more preferably less than 220° C. Most preferably, the temperature is between 125° C. and 220° C.

[0057] Temperature can be measured by various (electrical) devices, as described herein. The most commonly used sensors for measuring temperature in the chemical industry are resistance thermometers (also called resistance temperature detectors, RTDs) and thermocouples for temperatures below 600° C. Within the RTD group, Pt100 type sensors are the most commonly used.

[0058] In the tubes of the multi-tube heat exchanger reactor, heat is generated by hydrogenation of phenol, thereby heating the mixture of components within the tubes, which is then removed by indirect cooling using a cooling medium.

[0059] The temperature of the mixture containing cyclohexanone and cyclohexanol discharged from the hydrogenation reactor should be above 60°C. Preferably, the temperature is higher than 80°C. More preferably, the temperature is higher than 100°C. In order to avoid catalyst wear, decomposition and the formation of by-products, the temperature of the mixture containing cyclohexanone and cyclohexanol discharged from the hydrogenation reactor should be lower than 260°C. Preferably, the temperature is lower than 240°C. More preferably, the temperature is lower than 220°C. Best results are achieved when the temperature of the mixture containing cyclohexanone and cyclohexanol discharged from the multi-tubular heat exchange reactor is between 100 and 220°C. The lower limit temperature is determined by the reaction rate in the phenol hydrogenation reaction. At lower temperatures, the reaction rate is lower and, as a result, a large amount of catalyst and a large reactor would be required. In addition, the temperature of the cooling medium to be heated is lower and, as a result, many possibilities for use of the absorbed heat are lost. The upper limit temperature is determined, among other things, by the stability of the catalyst. In addition, at high temperatures, the yield of phenol hydrogenation will decrease due to the formation of more undesirable by-products (e.g., tar and benzene) due to the decomposition of phenol.

[0060] According to the invention, the temperature in the hydrogenation reactor is regulated by the temperature of the cooling medium used for indirect heat exchange and by the temperature of the inlet mixture. The higher the reactor temperature, the higher the reaction rate, but above a certain temperature threshold, there is also a decrease in selectivity and more undesirable by-products are formed.

[0061] In this regard, it should be noted that the actual temperature in the reactor cannot be precisely measured or reported. The temperature in the catalyst bed is not constant: it varies both with location and with time. For example, the local temperature inside the catalyst bed of a multi-tubular heat exchange reactor used for the hydrogenation of phenol is the result of the heat generated by the hydrogenation of phenol and the heat transferred to the cooling medium. The temperature varies both along the length and along the radius of the catalyst bed. Generally, the temperature along the length of the catalyst bed goes through a maximum value ("peak"). The shape of the temperature profile and the location of the highest temperature in the bed vary depending on a number of process parameters, such as the feed flow rate, the concentration of both phenol and hydrogen in the feed, the type and condition of the catalyst, and the temperature of the cooling medium outside the catalyst bed (outside the tubes in the case of a multi-tubular heat exchange reactor). As time progresses, the location of the highest temperature in the bed gradually shifts from a location near the entrance of the gaseous feed into the bed to a location near the exit of the gaseous product from within the bed. Generally, the exact temperature is not known anywhere inside the catalyst bed. However, a global indication can be obtained by measuring the temperature at various heights within the catalyst bed by placing a series of temperature sensors (eg Pt100 type sensors) along the length of the catalyst bed.

[0062] In a particularly preferred embodiment of the present invention, the temperature of the cooling medium is therefore used to control the hydrogenation reaction of phenol. If the temperature of the cooling medium is too low, too much heat will be absorbed by the cooling medium and the exothermic hydrogenation reaction will proceed to a smaller extent in the reactor. If the temperature of the cooling medium is too high, insufficient heat will be absorbed by the cooling medium and the exothermic hydrogenation reaction will locally overheat the catalyst bed, causing catalyst degradation and the formation of undesirable by-products. The inventors have found that good results are achieved when the temperature of the cooling medium is maintained in the temperature ranges mentioned above for the mixture comprising cyclohexanone and cyclohexanol discharged from the reactor.

[0063] The temperature of the cooling medium is one of the process parameters that can be adapted (preferably automatically) in steps I) to L) of the present invention. It is a particularly useful process parameter for adapting in steps K), L) and / or L2) in response to the finding in steps G), H) or H2) that the concentration of phenol or cyclohexanol in the exit mixture exceeds the third and fourth setpoints, respectively, or if the concentration of cyclohexanone obtained in step D2) is lower than the fifth setpoint. If the concentration of phenol in the exit mixture exceeds the third setpoint, the temperature of the reactor should be increased to increase the conversion to cyclohexanone. However, care must be taken, since too high a temperature and conversion can adversely affect the selectivity to cyclohexanone. If the concentration of cyclohexanol in the exit mixture exceeds the fourth setpoint, the temperature of the reactor should be decreased to reduce the conversion to cyclohexanol.

[0064] Typically, in the multi-tubular heat exchange reactor where phenol is hydrogenated, heat is removed by a cooling medium. An aqueous or organic solvent, or a mixture thereof, can be applied as the cooling medium. Preferably, in the process of the present invention, water is applied as the cooling medium in the multi-tubular heat exchange reactor. By absorbing heat, the cooling medium charged in the multi-tubular heat exchange reactor is warmed and preferably evaporated (partially or completely), or a combination thereof. Preferably, at least 10% by weight of the cooling medium charged in the multi-tubular heat exchange reactor is evaporated. More preferably, at least 50% by weight. Even more preferably, at least 90% by weight. Typically, in the process of the present invention, water is applied as the cooling medium in the multi-tubular heat exchange reactor and more than 90% by weight of the water is evaporated.

[0065] Preferably, water is used as the cooling medium for the indirect heat exchange used in the phenol hydrogenation reactor. This is advantageous not only in terms of availability and cost of the cooling medium, but also in that steam can be produced in the heat exchanger. This co-produced steam has the advantage that it can be used for heating purposes. Preferably, the produced steam is used in a process for the production of cyclohexanone by hydrogenation of phenol. Preferably, the produced steam is used for the evaporation of phenol fed to the phenol hydrogenation reactor (e.g., as a heating source in the phenol evaporation section). More preferably, the produced steam is used for the recovery of cyclohexanone from the outlet mixture of the phenol hydrogenation reactor (e.g., by driving a reboiler of a distillation column).

[0066] When water is used as the cooling medium, according to a further preferred embodiment, boiling water is used as the cooling medium. Not only is steam produced in this way, but the temperature of the boiling water used to absorb the heat of the exothermic gas-phase hydrogenation can also be easily set by the temperature of the hydrogenation reactor. The boiling point 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 cooling medium. Typically, the pressure of the steam (when water is used as the cooling medium) is measured and compared with a predefined set value. The control system can be programmed to automatically open one or more valves that increase the outflow of the exiting steam stream in response to a determination that the steam pressure is too high, or to automatically close one or more valves that reduce the outflow of the exiting steam stream in response to a determination that the steam pressure is too low.

[0067] The inventors have found that it is possible to adjust the boiling water pressure in the range of 0.1 MPa to 5 MPa depending on the desired amount of cooling and reactor temperature. Preferably, in the process of the present invention, the boiling water pressure is adjusted to a value of at least 0.1 MPa, more preferably at least 0.15 MPa. Typically, in the process of the present invention, the boiling water pressure is adjusted to a value of less than 5 MPa, more preferably less than 1.5 MPa. Preferably, in the process of the present invention, the boiling water pressure is adjusted to a range of 0.15 MPa to 1.5 MPa.

[0068] The operating capacity of the phenol hydrogenation reactor is typically selected based on the amount of phenol to be charged to the phenol hydrogenation reactor.

[0069] The process of the present invention uses an industrial-scale hydrogenation reactor for the production of cyclohexanone. This means a phenol hydrogenation reactor that can be charged with several thousand tons (kilotons per year; kta) of phenol per year. In an embodiment of the present invention, the industrial-scale phenol hydrogenation reactor can be or is charged with more than 10,000 tons (10 kta) of phenol per year. The charge of phenol to the industrial-scale phenol hydrogenation reactor does not have to be a maximum value. Preferably, the charge of phenol to the industrial-scale phenol hydrogenation reactor is between 20 kta and 350 kta of phenol. More preferably, the charge of phenol to the industrial-scale phenol hydrogenation reactor is between 25 kta and 250 kta of phenol.

[0070] The operation mode of the phenol hydrogenation reactor can be selected from fed batch operation and continuous operation. In the fed batch operation mode, hydrogen is charged to the phenol hydrogenation reactor, which is previously charged with phenol, over time, and the reaction product remains in the reactor until the end of the reaction. In the continuous operation mode, both hydrogen and phenol are continuously charged to the phenol hydrogenation reactor, and the reaction product and unconverted feedstock are continuously discharged from the phenol hydrogenation reactor. Preferably, the industrial-scale phenol hydrogenation reactor is operated in a continuous mode.

[0071] Within the scope of the present invention, the expression "at least a portion of the hydrogen in the exit mixture is recycled to the entry mixture" refers to the hydrogen present in the exit mixture of the industrial-scale phenol hydrogenation reactor, which hydrogen was not consumed in the vapor-phase hydrogenation of phenol catalyzed by a catalyst comprising palladium in the industrial-scale hydrogenation reactor.

[0072] Within the scope of the present invention, the term "at least a portion" in the preceding expressions preferably means that at least 10%, even more preferably at least 50%, and most preferably at least 85% of the hydrogen in the exiting mixture is recycled to the entering mixture.

[0073] In step B) of the method according to the invention, the hydrogen concentration in the exit mixture is measured by in-line measurement. "Hydrogen concentration in the exit mixture" can mean the hydrogen concentration of any hydrogen-containing stream downstream of the phenol hydrogenation reactor, since this concentration will be proportional to the hydrogen concentration in the mixture leaving directly from the reactor. In practice, particularly convenient places for measuring the hydrogen concentration in the exit mixture are in the line transporting the gaseous purge stream (see Example 1 below, line [4] in FIG. 2), in the line recycling part of the hydrogen in the exit mixture to the incoming mixture (line

[22] in FIG. 2), in the line transporting the compressed gas mixture directly from the compression section (line

[32] in FIG. 2), in the line leaving directly from the phenol hydrogenation reactor (line

[29] in FIG. 2) or in the line leaving a heat exchanger or gas-liquid separator, if any, present (line

[30] or

[31] in FIG. 2). The streams in lines

[32] ,

[22] and [4] in Figure 2 are compressed and contain only traces of phenol, cyclohexanol and cyclohexanone. They contain mainly hydrogen and inerts. Thus, the above locations are particularly convenient for performing in-line measurements of hydrogen concentration in the exit mixture. Unless otherwise defined, the term "concentration" as used herein refers to molar ratio (mol%).

[0074] Hydrogen concentration can be measured by various techniques. Chromatography, especially gas chromatography (GC), is commonly used to separate individual components in a sample that is a mixture of components, which are then quantified by a detector (e.g., a thermal conductivity detector (TCD), also known as a catharometer). The detector signal is processed by a controller to calculate the hydrogen concentration. For continuous (calibrated) measurement of hydrogen concentration in a gaseous mixture containing limited amounts of components, a thermal conductivity detector is generally applied directly (without chromatographic pretreatment).

[0075] In step F) of the method of the present invention, the hydrogen concentration obtained in step B) is compared with a second set value. Those skilled in the art will know the hydrogen concentration that should be maintained in the industrial scale hydrogenation reactor they operate to obtain optimal phenol conversion and cyclohexanone selectivity. If the hydrogen concentration in the reactor is too high, the selectivity to cyclohexanone will decrease as more cyclohexanol is formed (see reaction scheme and stoichiometry above). If the hydrogen concentration in the reactor is too low, the conversion of phenol will be too low. The optimal set value for the hydrogen concentration in the exit mixture may vary depending on the age of the catalyst and other process parameters. For example, when using a very fresh catalyst, it is often recommended to moderate the reaction rate by operating with only a small amount of hydrogen present in the exit mixture. If the catalyst is aged, more hydrogen needs to be added to obtain sufficient phenol conversion.

[0076] In the mixture directly exiting the hydrogenation reactor (line

[29] in FIG. 2, line

[30] in FIG. 2), the second set point for the optimum hydrogen concentration may be in the range of 0.1-40 mol%, in particular 1-30 mol%, based on the total number of moles of components present in the mixture. If the gas composition in the line recycling part of the hydrogen in the exit mixture to the incoming mixture is measured to determine the H2 concentration in the exit mixture, the second set point may be in the range of 1-100 mol% (if no dilution gas, e.g. inert components, is added, the recycling can be almost 100 mol% H2), preferably 2-85 mol%, more preferably 3-70 mol%, most preferably 5-50 mol%. Lower values ​​are possible in the presence of unused (or regenerated) catalyst, the upper limit taking into account the aging of the catalyst. The gas composition in molar ratios is identical at lines

[31] ,

[32] ,

[22] and [4] shown in Figure 2 (although the temperature and pressure are not identical), and therefore the hydrogen concentration can be measured and the aforementioned ranges apply to all of these locations.

[0077] In step J) of the method of the invention, an adaptation of at least one process parameter is performed automatically if the hydrogen concentration in the exit mixture deviates from a second setpoint. As explained above, the hydrogen concentration can be too high or too low and varies depending on the state of the catalyst and other process parameters. Careful and frequent monitoring should therefore be performed. In practice, the second setpoint can be a concentration range around the optimal setpoint. For example, the range can be ±25%, ±10%, ±1% or ±0.5% from a predefined setpoint in the range indicated in the previous paragraph.

[0078] Steps B), F) and J) are preferably carried out at least once a day, more preferably at least once an hour, and most preferably at least once every minute or even every 5-30 seconds to achieve optimal control. Since the steps need to be carried out at this frequency to achieve optimal results, the present invention foresees that the adaptation in step J) is carried out automatically.

[0079] Process parameters that can be automatically adapted in step J) to bring the hydrogen concentration back near or to the set point in a specific and effective manner include a) adjusting the flow rate of the inlet mixture and / or b) adjusting the flow rate of the purge stream.

[0080] Adjusting the flow rate of the inlet mixture in this context means in particular adjusting the inflow of the hydrogen-containing stream entering the reactor when the components hydrogen and phenol enter the reactor in separate streams. If the hydrogen concentration in the exit mixture is measured in step F) to be higher than a first set value, the flow rate of the inlet mixture can be automatically reduced in step J). Conversely, if the hydrogen concentration in the exit mixture is measured in step F) to be lower than a first set value, the flow rate of the inlet mixture can be automatically increased in step J).

[0081] As explained above, in the plant or method according to the invention, at least a part of the hydrogen from the exit mixture is discharged as a purge stream. This purge stream is an important way to regulate the concentration of hydrogen. The main purpose of the purge stream is to remove inert components (dilution gases). Increasing the outflow of the purge stream leads to a lower concentration of inert components and therefore a higher hydrogen concentration (and vice versa). The flow rate of the purge stream can be adjusted by a valve that regulates it. If the hydrogen concentration in the exit mixture is measured in step F) to be higher than a first set value, the flow rate of the purge stream can be automatically decreased in step J). Conversely, if the hydrogen concentration in the exit mixture is measured in step F) to be lower than the first set value, the flow rate of the purge stream can be automatically increased in step J). In this way, the hydrogen concentration in the exit mixture can be brought back again near or to a set value in a specific and effective manner by simply opening and closing a valve.

[0082] According to a preferred embodiment of the present invention, the inlet mixture further comprises a diluent gas. The diluent gas should be an inert gas, such as methane and / or nitrogen. Being inert, the diluent gas is present in both the inlet and the outlet mixture. Herein, the diluent gas is sometimes simply referred to as "inerts". Those skilled in the art often prefer to operate the hydrogenation process at a constant pressure, independent of changes in the amount (and associated pressure) of phenol and hydrogen in the feed stream, which may require further adaptations according to the process control implemented by the present invention. For example, if the total partial pressure of phenol and hydrogen fed to the reactor is 0.21 MPa, the partial pressure of the inerts fed to the reactor may be 0.09 MPa (if no other components are fed to the reactor) to reach the desired pressure set point of 0.30 bar. As hydrogen is consumed in the hydrogenation reactor, the ratio of hydrogen to inerts is changing due to reactions in the recycle gas stream containing unreacted hydrogen and inerts compared to the inlet mixture. To maintain a sufficient reaction rate, the inventors have found that the amount of inerts and hydrogen excess can be varied: consequently, a decrease in the concentration of inerts leads to an increase in the hydrogen concentration and therefore an increase in the reaction rate.

[0083] In a further preferred embodiment, apart from the hydrogen in the exit mixture being recycled to the inlet mixture, at least a part of the dilution gas present in the exit mixture is also recycled to the inlet mixture. This offers a particularly convenient way of minimizing the amount of unused dilution gas to be fed to the inlet mixture. By using dilution gas in the inlet mixture, the molar ratio of this dilution gas to hydrogen gas is higher in the purge than in the inlet mixture. This offers a particularly convenient possibility for adjusting the hydrogen concentration. In steps C) and D) of the method of the invention, the concentrations of phenol and cyclohexanol, respectively, are measured by in-line measurement. Optionally, the method of the invention can comprise an additional step D2) in which the concentration of cyclohexanone is measured by in-line measurement. Phenol, cyclohexanol and cyclohexanone can all be measured using the same in-line measurement device. Steps C), D) and optionally D2 can thus be carried out in one step. The concentrations of phenol, cyclohexanol and cyclohexanone can be measured by various techniques. Chromatographic techniques, particularly gas chromatography (GC), are commonly used to separate individual components in a sample that is a mixture of components, which are then quantified by a detector (e.g., a thermal conductivity detector (TCD), also known as a catharometer). The detector signals are processed by a controller to calculate the various concentrations.

[0084] The location of the in-line measurement of the phenol, cyclohexanol and cyclohexanone concentrations in the exit mixture can be at various locations downstream of the hydrogenation reactor. The measurements can be performed directly in the mixture leaving the phenol hydrogenation reactor (line

[29] in FIG. 2) or in the line leaving the heat exchanger or gas-liquid separator that may be present (lines

[30] or [5] in FIG. 2). A particularly convenient location for the in-line measurement of the phenol, cyclohexanol and / or cyclohexanone concentrations in the exit mixture is in the liquid containing phenol, cyclohexanol and cyclohexanone that is present in or leaving the gas-liquid separation section [H] (line [5] in FIG. 2). A calibrated in-line gas chromatograph (GC) can be installed at any location. Concentrations of phenol, cyclohexanol and cyclohexanone, as used herein, are expressed as mol %, with the sum of phenol, cyclohexanol and cyclohexanone in the analyzed mixture equaling 100 mol %.

[0085] The setpoint for the phenol concentration used for comparison in step G) of the method of the invention, and also the setpoint for the cyclohexanol concentration in step H) of the method of the invention, must each independently be in the range of 0-20 mol%, preferably 0-12 mol%, more preferably 0-10 mol%, i.e. less than 10 mol%, always taking the sum of phenol, cyclohexanol and cyclohexanone in the analyzed mixture as 100 mol%. The opposite applies to the fifth setpoint for the concentration of cyclohexanone, which can be calculated as [anone]=100 mol%-[anol]-[phenol], since the sum of the concentrations of the three components is taken as 100 mol% to indicate the respective concentration setpoints herein. Therefore, the concentration of cyclohexanone, i.e., the fifth set value, is preferably in the range of 80 to 100 mol%, preferably 88 to 100 mol%, more preferably 90 to 100 mol%, or greater than 90 mol%, with the sum of phenol, cyclohexanol and cyclohexanone in the analyzed mixture always being 100 mol%.

[0086] Since it is generally desirable to have as little cyclohexanol as possible (high selectivity to cyclohexanone) and as little unconverted phenol as possible (high conversion) in the exiting reaction mixture, the third and fourth setpoints are such that they only require an adaptation of the process parameters in steps K) and L) of the invention if the concentrations measured in steps C) and D) are higher than the corresponding setpoints, respectively. Similarly, since it is generally desirable to have as much cyclohexanone as possible in the exiting reaction mixture (high selectivity to cyclohexanone and high conversion of phenol), the fifth setpoint is such that it only requires an adaptation of the process parameters in step L2) of the invention if the concentration measured in step D2) is lower than the corresponding setpoint. This is significantly different from the setpoints referred to in steps E), F), I) and J) of the process of the invention, which require an adaptation if the measured values ​​are lower or higher than the setpoints.

[0087] A further difference between the setpoints referred to in steps E), F), I) and J) of the method of the invention and those referred to in steps G), H), K), L) and L2) is that the former require very frequent attention with in-line measurement and comparison with the setpoint, which is recommended to be in the range of at least once a day, once an hour, once a minute or every 5 to 30 seconds. In contrast, it has been found that the concentrations of phenol and cyclohexanol require adaptation much less frequently. The inventors have found that it is often sufficient to measure, compare and possibly adapt the process parameters based on the results in the range of once a day or once a week. This is also why an automatic decision process is not required, although it is still desirable, for steps K) and / or L) of the method of the invention. According to only a preferred embodiment, at least one process parameter adapted in steps K), L) and / or L2) is adapted in an automatic operating mode.

[0088] In step K) of the method of the invention, an adaptation of at least one process parameter is performed if the concentration of phenol in the exit mixture deviates from a third set value. In step L) of the method of the invention, an adaptation of at least one process parameter is performed if the concentration of cyclohexanol in the exit mixture deviates from a fourth set value. In step L2) of the method of the invention, an adaptation of at least one process parameter is performed if the concentration of cyclohexanone obtained in step D2) is lower than a fifth set value. Steps K), L) and / or L2) may be optionally performed, for example, when the catalyst is still fresh or has been regenerated. The inventors have found that steps K), L) and / or L2) are particularly important for the situation occurring after several weeks to months, when the catalyst begins to age. Catalyst ageing leads to a (slight) decrease in conversion (with the result that the concentration of phenol in the exit mixture increases) and a decrease in selectivity (with the result that the concentration of cyclohexanol in the exit mixture increases). Usually, catalyst aging is a slow process, and as a result, the period between start-up with a fresh or regenerated catalyst bed and replacement or regeneration of the old catalyst bed can range from a few weeks to a few years. The actual figures vary depending on many parameters, such as the quality of the fresh phenol, the quality of the fresh hydrogen, the feed water, the desired product specifications, the catalyst type and the layout of the plant. However, due to process disturbances (such as abnormally high concentrations of CO in the fresh hydrogen), catalyst aging can be much faster than under normal conditions, and as a result, the period between the onset of the process disturbance and the replacement or regeneration of the old (poisoned) catalyst bed can be reduced to just a few days or even less. In this case too, steps K), L) and / or L2) can be beneficial in counteracting the adverse effects of catalyst aging.

[0089] If the concentration of cyclohexanone is also measured (see step D2 above), the method of the invention can comprise a further step H2 consisting of comparing the concentration of cyclohexanone obtained in step D2 with a fifth set value and, optionally, a further step L2) consisting of carrying out an adaptation of at least one process parameter if the concentration of cyclohexanone obtained in step D2) is lower than the fifth set value. The fifth set value, i.e. the desired cyclohexanone concentration in the exit mixture, is 75-100 mol%, preferably 85-100 mol% or 90-100 mol%, most preferably more than 95 mol%, 98 mol% or 99 mol%, in all cases taking the sum of phenol, cyclohexanol and cyclohexanone in the analyzed mixture as 100 mol%.

[0090] The amount of deviation from the set value that is considered acceptable without still triggering the adaptation in steps K), L) or L2) can be defined by the person skilled in the art according to the invention (for example, the range may be ±25%, ±10%, ±1% or ±0.5% from a predefined set value lying within the ranges indicated above).

[0091] Steps C), G) and K) as well as steps D), H) and L) are preferably carried out at least once a month, more preferably at least once a week, and most preferably at least once a day, hour or minute.

[0092] Process parameters that can be automatically adapted in steps K) and L) to bring the phenol and cyclohexanol concentrations back below the set values, respectively, in a specific and effective manner include a) adjusting the temperature of the cooling medium, b) adjusting the hydrogen concentration in the exit mixture, c) adjusting the concentration of the dilution gas in the exit mixture, and d) combinations thereof.

[0093] It is well known that the reactor temperature affects the conversion of phenol and the selectivity to cyclohexanone in a phenol hydrogenation reactor. As explained above with respect to the heat exchanger, adjusting the temperature of the cooling medium (particularly by the pressure of boiling water when boiling water is used as the cooling medium) is a particularly convenient way to adjust the temperature of the reactor. Typically, the temperature of the cooling medium (and therefore the temperature of the reactor) should be increased to increase the conversion to cyclohexanone. However, care must be taken because too high a temperature and conversion can adversely affect the selectivity to cyclohexanone.

[0094] The term "adjusting the hydrogen concentration in the exit mixture" as used herein will in most cases mean "adjusting the hydrogen concentration in the inlet mixture", in particular by increasing the amount of hydrogen charged to the process. By increasing the amount of hydrogen fed to the reactor, the hydrogen concentration and hydrogen partial pressure in the reactor will increase, which will increase the reaction rate, i.e. the conversion. However, care should be taken not to increase the hydrogen concentration and the associated conversion to cyclohexanol too much, as this will decrease the desired selectivity to cyclohexanone. A particularly convenient way of adjusting the hydrogen concentration in the exit mixture is to adjust the second set value used for the comparison in step F) and for the adaptation of the process parameters in step J). This is illustrated in Example 2 below. In this way, it is also possible to adjust the hydrogen concentration in the exit mixture indirectly using the process parameters.

[0095] Adapting the concentration of the diluent gas in the exiting mixture as used herein will in most cases mean adapting the concentration of the diluent gas in the inlet mixture. For example, reducing the concentration of the diluent gas in the exiting mixture is achieved by reducing the concentration of the diluent gas in the inlet mixture. A particularly convenient way to achieve this is by adding less fresh hydrogen or by adding more recycled gas mixture. As mentioned above, adjusting (particularly reducing) the concentration of the diluent gas is a convenient way to adjust (particularly increase) the concentration of hydrogen and thus the reaction rate and the associated conversion.

[0096] In a preferred embodiment, the process parameters that can be automatically adapted in step K) to bring the phenol concentration back below the set value are selected from a) increasing the temperature of the cooling medium, b) increasing the hydrogen concentration in the exit mixture, c) decreasing the concentration of the dilution gas in the exit mixture, and d) combinations thereof.

[0097] In a preferred embodiment, the process parameters that can be automatically adapted in step L) to bring the cyclohexanol concentration back below the setpoint again are selected from a) reducing the temperature of the cooling medium, b) reducing the hydrogen concentration in the exit mixture, c) increasing the concentration of the dilution gas in the exit mixture, and d) combinations thereof.

[0098] If the method comprises a step L2) of adapting at least one process parameter if the concentration of cyclohexanone obtained in step D2) is lower than the fifth setpoint, the process parameter adapted is the same as described above for the third setpoint related to the concentration of phenol. The reason is that all the measures described for reducing the phenol concentration in the exit mixture are measures for increasing the conversion to cyclohexanone and the selectivity towards cyclohexanone, respectively. Thus, the above measures for adapting the process parameters can also be used in step L2) to increase the concentration of cyclohexanone in the exit mixture above the fifth setpoint. For example, an increase in the temperature of the cooling medium 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 by reducing the phenol concentration in the inlet mixture, i.e. by feeding less phenol to the hydrogenation reactor. However, this would reduce the output of the plant.

[0099] Next to the above-mentioned method, the present invention further provides an industrial-scale phenol hydrogenation plant adapted to carry out the method according to the invention. It comprises the components described with respect to the method steps above, but also comprises a control system (e.g. software) capable of carrying out steps E) to J) of the invention, and preferably further steps K) and L) and / or D2) and L2). The control system not only carries out the decision-making process in steps E) to J) of the invention, and preferably further steps K) and L) and / or D2) and L2), but also preferably further adapts the process parameters as output of the decision-making process. For example, the control system can be programmed to automatically open one or more valves, which reduce the outflow of the purge stream and / or increase the inflow of the inlet mixture, in response to a measurement in steps E) and / or F) that the pressure or hydrogen concentration in the reactor is too low. In this way, the plant is adapted to carry out steps I) and J). This is only one example of possible inputs and outputs for the control system. The detailed description of the method of the invention above applies mutatis mutandis to the corresponding parts and features of the plant of the invention.

[0100] The plant of the present invention comprises at least a hydrogenation reactor comprising a catalyst containing palladium, at least one inlet by means of which an inlet mixture comprising phenol and hydrogen can be charged to the reactor, at least one outlet capable of releasing an exit mixture comprising cyclohexanone, cyclohexanol, phenol and hydrogen, a line for recycling at least a portion of the hydrogen in the exiting mixture to the entering mixture; a recovery section capable of recovering cyclohexanone from the exit mixture, a purge section capable of releasing at least a portion of the hydrogen from the exit mixture as a purge stream; - an indirect heat exchanger capable of transferring at least a portion of the reaction heat of the phenol hydrogenation reaction to a cooling medium; Equipped with:

[0101] The plant of the invention is characterized in that it has the following parameters: - pressure in the hydrogenation reactor; - hydrogen concentration in the exit mixture; - phenol concentration in the exit mixture; - the concentration of cyclohexanol in the exit mixture; and - possibly the concentration of cyclohexanone in the exit mixture; The present invention is characterized in that there is an in-line measurement device capable of measuring

[0102] The plant further comprises an automatic control system capable of carrying out steps E) to J), and preferably steps K) and L) and / or D2), H2) and L2) of the invention as described above. [Brief description of the drawings]

[0103] [Figure 1] FIG. 1 shows a schematic diagram of a plant and process for the preparation and recovery of cyclohexanone from phenol according to the present invention. [Diagram 2] 1 is a schematic diagram showing an embodiment of a phenol hydrogenation reaction section [A] according to the present invention, which 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]. DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS

[0104] In the description of the drawings the following reference numerals will be used: [A] Phenol hydrogenation reaction section [B] Separation and purification department [C] Cyclohexanol dehydrogenation reaction section [D] Hydrogen Refining Department [E] Phenol evaporation section [F] Phenol hydrogenation unit [G] Heat exchange section [H] Gas-liquid separation section [J] Compression section [1] Hydrogen-containing flow [2] Unused phenol stream [3] Recovered phenol-containing stream [4] Gaseous purge flow [5] Mixtures containing phenol, cyclohexanone, cyclohexanol and hydrogen [6] Stream containing cyclohexanol and cyclohexanone [7] Cyclohexanone-rich stream [8] The first stream is rich in cyclohexanol. [9] The second stream is rich in cyclohexanol.

[10] Flow of light components

[11] Heavy component flow

[12] Hydrogen produced

[21] Hydrogen gas

[22] Recycled hydrogen gas

[23] Hydrogen-containing stream

[24] Combined phenol stream

[25] Flow of gaseous components

[26] Components that did not evaporate

[27] Water

[28] Supply flow

[29] A gas mixture containing cyclohexanone, cyclohexanol, phenol and hydrogen.

[30] Mixtures containing hydrogen gas and the liquids phenol, cyclohexanone and cyclohexanol.

[31] Gas mixtures containing hydrogen

[32] Compressed gas mixture

[0105] [Detailed description of the drawings] A plant and a corresponding process for the preparation and recovery of cyclohexanone from phenol are shown diagrammatically in Figure 1. Such a process usually consists of two parts, possibly with a third part. All three parts are described.

[0106] Cyclohexanone is prepared catalytically by reacting phenol with hydrogen in the phenol hydrogenation section [A]. Cyclohexanone is recovered in the separation and purification section [B]. In the optional cyclohexanol dehydrogenation section [C], the undesired by-product cyclohexanol is catalytically converted to cyclohexanone and hydrogen.

[0107] The phenol hydrogenation reaction section [A] comprises at least one hydrogenation reactor, into which a hydrogen-containing stream via line [1], a fresh phenol stream via line [2] and optionally a recovered phenol-containing stream via line [3] are fed, and may also comprise additional equipment. The hydrogen-containing stream may optionally comprise inert components (e.g. nitrogen and / or methane) and optionally carbon monoxide as diluting gases. Optionally, the inert components (e.g. nitrogen and / or methane) and optionally carbon monoxide are fed separately to the phenol hydrogenation reaction section [A] (not shown in FIG. 1). The catalytic phenol hydrogenation according to the invention is carried out in a vapor phase process. From the phenol hydrogenation reaction section [A], a gaseous purge stream comprising hydrogen and optionally inert components (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]. The hydrogenation reactor section [A] is described in more detail below in connection with FIG. 2 as it is the process by which the purge stream [4] and the product stream [5] are obtained.

[0108] A mixture containing phenol, cyclohexanone, cyclohexanol and hydrogen is fed via line [5] to the separation and purification section [B]. Optionally, a stream containing cyclohexanol and cyclohexanone is fed via line [6] from the cyclohexanol dehydrogenation section [C] to the separation and purification section [B]. Typically, the separation and purification section [B] comprises one or more distillation columns. In the separation and purification section [B], a cyclohexanone-rich stream and optionally a cyclohexanol-rich stream and optionally a phenol-containing stream are recovered. From this separation and purification section [B], a cyclohexanone-rich stream is discharged as a primary product via line [7]. This cyclohexanone-rich stream can be used as a solvent or as a raw material for, for example, the production of ε-caprolactam, caprolactone or adipic acid (not shown in FIG. 1). Optionally, the first cyclohexanol-rich stream is discharged outside the process via line [8]. It is conceivable that this first cyclohexanol-rich stream can be used as a solvent or as a raw material for the production of, for example, adipic acid (not shown in FIG. 1). Optionally, the second cyclohexanol-rich stream is discharged via line [9] to the cyclohexanol dehydrogenation reactor [C]. Optionally, a stream containing recovered phenol is discharged via line [3] to the phenol hydrogenation reactor [A]. Optionally, a light stream, optionally containing benzene, cyclohexane and water, is discharged via line

[10] . Optionally, a heavy stream, containing phenol and components with a boiling point higher than phenol, is discharged via line

[11] .

[0109] 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 to cyclohexanone and hydrogen. Generally, the dehydrogenation of cyclohexanol is a gas-phase reaction carried out at high temperature. Optionally, the product stream containing cyclohexanol and cyclohexanone is fed via line [6] to the separation and purification section [B]. The hydrogen produced in the cyclohexanol dehydrogenation reaction section [C] is discharged as product hydrogen via line

[12] . Optionally, the product hydrogen from the cyclohexanol dehydrogenation reaction section [C] is fed to the phenol hydrogenation reaction section [A] (not shown in FIG. 1). Optionally, the product hydrogen from the cyclohexanol dehydrogenation reaction section [C] is fed to another hydrogen-consuming process (not shown in FIG. 1). Optionally, the product hydrogen of the cyclohexanol dehydrogenation reaction section [C] is fed to a heat generating unit (not shown in FIG. 1).

[0110] FIG. 2 shows a scheme of an embodiment of the phenol hydrogenation reaction section [A] (area enclosed by a dashed line) according to the invention.

[0111] A hydrogen-containing stream is fed to the phenol hydrogenation reactor [A] via line [1]. The hydrogen-containing stream may originate, for example, from a naphtha cracker, a methane reformer, a coal gasification process, an electrolysis process, or a dehydrogenation process. Typically, the hydrogen-containing stream contains inert components such as nitrogen and / or methane. Optionally, inert components such as nitrogen and / or methane can be added to the hydrogen-containing stream as dilution gas (not shown in FIG. 2). If the hydrogen-containing stream contains components harmful to hydrogenation, such as CO and / or H2S, a hydrogen gas purification step is necessary. The presence of these harmful components in the hydrogen fed may be temporary, for example due to upset conditions in the hydrogen gas production unit, or may be permanent. In such a hydrogen purification step, harmful impurities may be converted to inert components or removed from the hydrogen-containing stream.

[0112] For this purpose, the hydrogen-containing stream is fed via line [1] to an optionally present hydrogen purification section [D]. Before feeding into the hydrogen purification section [D], the temperature of the hydrogen-containing stream can be changed in a heat exchange section (not shown in FIG. 2) to the temperature required in the hydrogen purification unit [D]. Typically, in such a heat exchange section, the temperature of the hydrogen-containing stream is increased.

[0113] The hydrogen purification section [D] may comprise one or more catalysts for converting harmful components to inert components and / or one or more adsorbents for removing harmful components. Preferably, the hydrogen purification section [D] comprises a catalyst for the conversion of CO and / or an adsorbent for the removal of H2S. Hydrogen gas is discharged from the hydrogen purification section [D] via line

[21] . The hydrogen gas in line

[21] and the recycled hydrogen gas in line

[22] are combined to form a hydrogen-containing stream that is charged to the phenol evaporation section [E] via line

[23] . The hydrogen purification section [D] may comprise one or more reaction and / or adsorption units operated in parallel and / or in series. Optionally, water, e.g. in the form of steam, may be added to the hydrogen gas before, in or after the hydrogen purification section [D] (not shown in FIG. 2). The hydrogen purification section [D] may be absent or may be bypassed (not shown in FIG. 2), for example if the quality of the hydrogen-containing stream in line [1] is sufficient.

[0114] A fresh phenol stream is optionally charged via line [2], but preferably a recovered phenol-containing stream is charged via line [3] to form a combined phenol stream flowing through line

[24] . The recovered phenol-containing stream charged via line [3] is discharged from the separation and purification section [B] shown in FIG. 1. Optionally, the recovered phenol-containing stream is not combined with the fresh phenol stream charged via line [2] and is charged separately to the phenol evaporation section [E] or is not present (not shown in FIG. 2). Optionally, prior to charging to the phenol evaporation section [E], the temperature of the combined phenol stream flowing through line

[24] is increased in a heat exchange section (not shown in FIG. 2) comprising one or more heat exchangers which can be operated in parallel and / or in series.

[0115] In the phenol evaporator [E], practically all the components entering via line

[24] are evaporated. A stream of gaseous components is discharged from the phenol evaporator [E] via line

[25] . The components entering via line

[24] and not evaporated are discharged (continuously or batchwise) from the phenol evaporator [E] via line

[26] as non-evaporated components. These are usually in very small quantities. Generally, the phenol evaporator [E] requires heating of the incoming feed, in particular heating with steam. Preferably, the phenol evaporator [E] comprises a device for removing entrained liquid droplets from the stream of gaseous components that is discharged, for example a wire mesh demister. The phenol evaporator [E] comprises one or more evaporators operated in parallel and / or in series. Optionally, the stream of gaseous components in line

[25] is temperature-conditioned in a heat exchanger (not shown in FIG. 2) comprising one or more heat exchangers operated in parallel and / or in series. Typically, in this heat exchange section, the stream of gaseous components is increased in temperature. Optionally, some water, e.g., in the form of steam, is charged via line

[27] to the stream in line

[25] to form a feed stream which passes through line

[28] to the phenol hydrogenation section [F].

[0116] The phenol hydrogenation section [F] consists of one or more hydrogenation reactors operated in series and / or in parallel. In the 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 the phenol hydrogenation section [F] via line

[29] .

[0117] The gas mixture comprising cyclohexanone, cyclohexanol, phenol and hydrogen discharged from the phenol hydrogenation section [F] via line

[29] is cooled in a heat exchange section [G], whereby at least a fraction of the phenol, cyclohexanone and cyclohexanol is condensed. The heat exchange section [G] comprises one or more heat exchangers operated in parallel and / or in series. As cooling medium, it is conceivable that one or more cooling media are used to cool the gas mixture comprising cyclohexanone, cyclohexanol, phenol and hydrogen discharged from the phenol hydrogenation section [F] via line

[29] . At least one of the applied cooling media should be at a temperature low enough to condense at least a fraction of the phenol, cyclohexanone and cyclohexanol present in the gas mixture comprising cyclohexanone, cyclohexanol, phenol and hydrogen discharged from the phenol hydrogenation section [F] via line

[29] . Optionally, the cooling of the gas is carried out in several stages in succession, so that optionally a different cooling medium is used for each stage. Optionally, at least one of the cooling media is a process stream from the phenol hydrogenation reactor [A], from the separation and purification reactor [B] and / or from the cyclohexanol dehydrogenation reactor [C] (not shown in Figure 2). Optionally, one of the cooling media is recycled hydrogen gas, which is charged via line

[22] (not shown in Figure 2).

[0118] The mixture containing hydrogen gas and liquid phenol, cyclohexanone and cyclohexanol is discharged from the heat exchange section [G] via line

[30] and charged to the gas-liquid separation section [H]. The gas-liquid separation section [H] comprises one or more gas-liquid separators operated in parallel and / or series. The liquid mixture containing phenol, cyclohexanone, cyclohexanol and hydrogen is discharged from the gas-liquid separation section [H] via line [5] and charged to the separation and purification section [B] shown in Figure 1. The gas mixture containing hydrogen is discharged from the gas-liquid separation section [H] via line

[31] and charged to the compression section [J]. Any type of gas-liquid separator can be utilized in the gas-liquid separation section [H], but particularly good results are obtained when a vertical vessel is used in which the liquid settles to the bottom of the vessel. Typically, the gas-liquid separation section [H] comprises a device for removing entrained liquid droplets from the flow of gaseous components being discharged, for example a wire mesh demister.

[0119] The compression section [J] comprises one or more devices for compressing the gas mixture. The devices can be operated in parallel and / or in series. The compressed gas mixture is discharged from the compression section [J] via line

[32] and is divided into recycled hydrogen gas, which is transferred via line

[22] , and a gaseous purge stream, which is transferred via line [4]. Optionally, the recycled hydrogen gas transferred via line

[22] is fed to a heat exchange section, where it is heated (not shown in FIG. 2). The recycled hydrogen gas (optionally heated) is discharged via line

[22] and then combined with the hydrogen gas in line

[21] .

[0120] When carrying out the process of the present invention, the gaseous purge stream transferred via line [4] usually contains hydrogen and one or more inert components, such as nitrogen and / or methane, and optionally carbon monoxide. This gaseous purge stream discharged via line [4] can be used, for example, as a fuel (not shown in FIG. 2). Optionally, before being discharged from the phenol hydrogenation reaction section [A], the gaseous purge stream transferred via line [4] is fed to a heat exchange section, in which the gaseous purge stream is cooled (not shown in FIG. 2). Optionally, the liquid formed in this heat exchange section is fed to a gas-liquid separation section [H] (not shown in FIG. 2). EXAMPLES

[0121] The following examples serve to illustrate the invention in more detail, particularly with respect to certain specific embodiments of the invention, but are not intended to limit the disclosure.

[0122] The example was carried out in a phenol hydrogenation reaction section [A] of a chemical plant for preparing and recovering cyclohexanone from phenol, which is very similar to the embodiment of the present invention shown in Figure 1. The chemical plant for preparing and recovering cyclohexanone from phenol further included a separation and purification section [B] in which high-quality cyclohexanone is obtained by distillation separation from higher-boiling components and lower-boiling components, and a cyclohexanol dehydrogenation reaction section [C] in which cyclohexanol is converted to cyclohexanone. The phenol hydrogenation reaction section [A] included 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 are very similar to the embodiment of the present invention shown in Figure 2.

[0123] A hydrogen-containing stream containing about 94% by volume of hydrogen, about 6% by volume of nitrogen, traces of CO and traces of H2S was fed via line [1] to the hydrogen purification section [D]. In the hydrogen purification section [D], CO was catalytically converted to CH4 and H2S was removed by an adsorbent. Under normal operating conditions, the CO and H2S contents in the hydrogen gas discharged from the hydrogen purification section [D] via line

[21] were both well below 1 ppm. A hydrogen-containing stream obtained by combining the recycled hydrogen gas transferred via line

[22] with the freshly supplied hydrogen gas transferred via line

[21] was fed via line 23 to the phenol evaporation section [E].

[0124] The unused phenol stream fed via line [2] and the stream containing recovered phenol fed via line [3] were combined and fed to the steam-heated phenol evaporation section [E], where virtually all of the phenol was evaporated. A small amount of non-evaporated components was discharged from the phenol evaporation section [E] via line

[26] , and the remainder was discharged as a gaseous component stream via line

[25] . Steam was added to the gaseous component stream in line

[25] via line

[27] so that the amount of water relative to the amount of phenol in line

[28] was about 1% by weight based on the amount of phenol. In this way, the feed stream for the phenol hydrogenation section [F] (transported via line

[28] ) was obtained.

[0125] The phenol hydrogenation section [F] had one industrial-scale multi-tubular phenol hydrogenation reactor with an annual operating capacity of about 120 kta of phenol, operated in a continuous mode for phenol hydrogenation in the vapor phase. The tubes were packed with supported hydrogenation catalyst Pd / Al2O3 (0.9 wt%) and 1 wt% Na (as NaHCO3) added as a 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) / (kg catalyst). Boiler water was charged as a cooling medium into the space outside the tubes to remove the reaction heat, thereby producing steam. More than 85% of the reaction heat was transferred to the cooling medium.

[0126] [Example 1] The multi-tubular hydrogenation reactor was loaded with fresh catalyst. Beginning 72 hours after the start-up of the process, the following process conditions were recorded for a 7 month continuous production period:

[0127] The pressure of the boiling water used as the cooling medium was set at 0.48 MPa (the associated temperature was approximately 150° C.). The steam produced was used as the energy source for the reboiler of the distillation column.

[0128] The pressure in the reactor was measured by in-line measurement in the reactor hood of the total pressure of the feed stream entering the multi-tubular hydrogenation reactor before it entered the catalyst-filled tubes. A pressure sensor continuously emitted an electric signal that was a measure of the magnitude of the pressure in the reactor. The pressure setpoint (first setpoint) was 0.38 MPa. If the pressure deviated from this setpoint, the outflow of the gaseous purge stream in line [4] was automatically adjusted to maintain the first setpoint (too high a pressure initiated the opening of the valve in line [4]; too low a pressure initiated the closing of the valve in line [4]).

[0129] The hydrogen concentration in the exit mixture was continuously measured by in-line measurement of the gaseous purge stream in line [4] with a calibrated in-line thermal conductivity detector (Catharometer). The thermal conductivity detector continuously emitted an electrical signal that was a measure of the hydrogen concentration. The setpoint of the hydrogen concentration (second setpoint) was 4 mol%, while the sum of the concentrations of all gaseous components (mainly hydrogen and inert gases, e.g. methane and / or nitrogen) was always 100 mol%. If the hydrogen concentration deviated from this setpoint, the flow of hydrogen gas in line [1] was automatically adjusted to maintain the second setpoint (too low a hydrogen concentration initiated an increase in the flow of hydrogen gas in line [1]; too high a hydrogen concentration initiated a decrease in the flow of hydrogen gas in line [1]).

[0130] The concentrations of phenol, cyclohexanol and cyclohexanone in the outlet mixture were measured by in-line measurement of the liquid in the bottom of the vertical vessel-type gas-liquid separator in the gas-liquid separation section [H] with a calibrated in-line gas chromatograph (GC). The set points (third to fifth set points) for the concentrations of phenol, cyclohexanol and cyclohexanone in the liquid in the bottom of the vertical vessel-type gas-liquid separator in the 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 seven months after start-up, the concentrations of phenol, cyclohexanol and cyclohexanone 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%).

[0131] A gas mixture containing cyclohexanone, cyclohexanol, phenol and hydrogen

[29] was discharged from the phenol hydrogenation reactor and cooled to 40 °C in the heat exchange section [G]. The resulting mixture containing hydrogen gas and liquid phenol, cyclohexanone and cyclohexanol was discharged via line

[30] and charged into the vertical vessel of the gas-liquid separation section [H]. From the bottom of the vertical vessel, a liquid mixture was discharged via line [5] and charged into the separation and purification section [B] (see Figure 1). In the separation and purification section [B], a cyclohexanone-rich stream (>99 mol% cyclohexanone), a cyclohexanol-rich stream, a light component stream and a heavy component stream were obtained by distillation separation and purification. The cyclohexanol-rich stream was charged into the cyclohexanol dehydrogenation section [C] via line [9]. In the cyclohexanol dehydrogenation reactor [C], cyclohexanol was catalytically dehydrogenated to produce hydrogen and cyclohexanone. The hydrogen was discharged through line

[12] , and the stream containing cyclohexanol and cyclohexanone was discharged through line [6] and fed to the separation and purification reactor [B].

[0132] The hydrogen-containing gas mixture was discharged from the top of the vertical vessel of the gas-liquid separation section [H] through line

[31] and fed to the compressor of the compression section [J]. The hydrogen-containing compressed gas mixture discharged from the compressor of the compression section [J] was separated into a gaseous purge stream transferred through line [4] and recycled hydrogen gas transferred through line

[22] .

[0133] The results of Example 1 show that by combining in-line measurement and automatic control of the pressure in the reactor and the hydrogen concentration in the exit mixture in accordance with the present invention, it was possible to operate a phenol hydrogenation reactor at a constant high production rate for about seven months without human intervention.

[0134] [Example 2] The preparation and recovery of cyclohexanone from phenol carried out in the same chemical plant as described in Example 1 continued after the initial seven month period.

[0135] After a few days, the phenol concentration in the liquid at the bottom of the vertical vessel-type gas-liquid separator in the gas-liquid separation section [H], measured by a calibrated in-line gas chromatograph, exceeded the setpoint set at 3 mol% (the sum of the concentrations of phenol, cyclohexanol and cyclohexanone was always taken as 100 mol%). In response to this, the setpoint of the hydrogen concentration was increased by 1 mol% (from 4 mol% to 5 mol%). As a result of this action, the phenol concentration in the liquid at the bottom of the vertical vessel-type gas-liquid separator in the gas-liquid separation section [H], measured by a calibrated in-line gas chromatograph (GC), dropped to less than 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 maintained for more than about one month without further adjustment of the setpoint of the hydrogen concentration. During the first 12 months after the start of operation, the set value of the hydrogen concentration had to be increased by 1 mol% several times (until the set value of the hydrogen concentration reached 10 mol%) in order to adjust the phenol concentration in the liquid at the bottom of the liquid reservoir of the vertical vessel-type gas-liquid separator in the gas-liquid separation section [H] to a value less than 3 mol%.

[0136] The consequence of increasing the hydrogen concentration set point from 4 mol% to 10 mol% was that the outflow of the gaseous purge stream in line [4] increased over time.

[0137] The results of Example 2 show that the adaptation of the hydrogen concentration in the outlet mixture is used to control the phenol concentration in the outlet mixture. Each time the concentration of phenol in the outlet mixture exceeds its set value of 3 mol%, the set value of the hydrogen concentration in the outlet mixture is increased stepwise by 1 mol% (from 4 mol% to 10 mol%), and as a result, the hydrogen concentration in the outlet mixture is increased stepwise from 4 mol% to 10 mol%. After the increase in the hydrogen concentration in the outlet mixture, the phenol concentration in the outlet mixture always falls again to less than 3 mol%. This control strategy allows the constant high production rate of the phenol hydrogenation reaction section [A] of the chemical plant to be extended for up to 12 months after start-up.

[0138] [Example 3] The preparation and recovery of cyclohexanone from phenol carried out in the same chemical plant as described in Example 2 continued after the initial 12 month period.

[0139] After a few days, the cyclohexanone concentration in the liquid at the bottom of the vertical vessel gas-liquid separator in the gas-liquid separation section [H], measured by a calibrated in-line gas chromatograph (GC), fell below the set point of 95 mol% (the sum of the concentrations of phenol, cyclohexanol and cyclohexanone was always taken as 100 mol%). In response to this, the pressure of the boiling water used as the cooling medium was increased by 0.02 MPa (to 0.50 MPa). As a result, the associated temperature increased by about 1.5°C to about 152°C.

[0140] As a result of this action, the cyclohexanone concentration in the liquid at the bottom of the vertical vessel gas-liquid separator in the gas-liquid separation section [H], as measured by a calibrated in-line gas chromatograph, again exceeded 95 mol%, while the concentrations of phenol and cyclohexanol remained <3 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 two months without further adjustment of the pressure of the boiling water used as the cooling medium.

[0141] The results of Example 3 show that by adapting once the pressure of the boiling water used as the cooling medium when the cyclohexanone concentration in the outlet mixture is lower than the set value for the cyclohexanone concentration in the outlet mixture, it was possible to maintain a constant high production rate of the phenol hydrogenation reaction section [A] of the chemical plant for another two months.

[0142] [Example 4] The preparation and recovery of cyclohexanone from phenol, carried out in the same chemical plant as described in Example 3, continued after an initial 14 month period. Due to market conditions, the production rate of the phenol hydrogenation reactor [A] had to be scaled down by 30%.

[0143] Immediately after reducing the phenol feed rate to the phenol hydrogenation reaction section [A] without adapting the setpoint, the cyclohexanol concentration in the liquid at the bottom of the liquid sump of the vertical vessel-type gas-liquid separator in the gas-liquid separation section [H], as measured by a calibrated in-line gas chromatograph (GC), exceeded the setpoint set at 3 mol% (the sum of the concentrations of phenol, cyclohexanol and cyclohexanone was always taken as 100 mol%). In response to this, the setpoint of the hydrogen concentration (second setpoint) was reduced by 2 mol% (from 10 mol% to 8 mol%), to which the system responded by reducing the inflow of hydrogen gas in line [1] as described for Example 1 above. As a result of this action, the hydrogen concentrations in the inlet and outlet mixtures decreased, and thus the cyclohexanol concentration in the liquid at the bottom of the vertical vessel gas-liquid separator in the gas-liquid separation section [H], as measured by a calibrated in-line gas chromatograph, decreased again to less than 3 mol%, while the concentrations of phenol and cyclohexanone both remained <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 further adjustment of any of the set points.

[0144] The results of Example 4 show that by lowering the set point of hydrogen concentration in the exiting mixture, and the resulting reduction in hydrogen concentration in the inlet and exiting mixtures, it was possible to reduce the concentration of cyclohexanol in the exiting mixture below its set point of 3 mol%.

[0145] [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 in this example, the outflow of the gaseous purge stream in line [4] was not automatically adjusted if the pressure deviated from the set value, and the inflow of hydrogen gas in line [1] was not automatically adjusted if the hydrogen concentration deviated from the set value.

[0146] Instead, the inflow of hydrogen gas in line [1] was automatically adjusted if the pressure deviated from the set value (if the pressure was too low, an increase in the inflow of hydrogen gas in line [1] was initiated; if the pressure was too high, a decrease in the inflow of hydrogen gas in line [1] was initiated), and the outflow of the gaseous purge stream in line [4] was automatically adjusted if the hydrogen concentration deviated from the set value (if the hydrogen concentration was too high, a decrease in the outflow of the gaseous purge stream in line [4] was initiated; if the hydrogen concentration was too low, an increase in the outflow of the gaseous purge stream in line [4] was initiated).

[0147] Also in this case, during the first seven months after the start of operation, the concentrations of phenol, cyclohexanol and cyclohexanone in the liquid at the bottom of the liquid reservoir of the vertical vessel-type gas-liquid separator in the 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 set to 100 mol%).

[0148] The results of Example 5 show that by combining in-line measurement and automatic control of the pressure in the reactor and the hydrogen concentration in the exit mixture in accordance with the present invention, it was possible to operate a phenol hydrogenation reactor at a constant high production rate for about seven months without human intervention.

Claims

1. 1. A process for controlling the exothermic vapor phase hydrogenation of phenol catalyzed by a catalyst comprising palladium in an industrial scale hydrogenation reactor, comprising: the reactor is charged with an inlet mixture comprising phenol and hydrogen and the reactor discharges an outlet mixture comprising cyclohexanone, cyclohexanol, phenol and hydrogen; and at least a portion of the hydrogen in the exiting mixture is recycled to the entering mixture; Cyclohexanone is recovered from the exit mixture, and at least a portion of the hydrogen from the exit mixture is released as a purge stream; and At least a portion of the heat of reaction of the phenol hydrogenation reaction is transferred to the cooling medium by indirect heat exchange; The method comprises: A) measuring the pressure in the reactor by in-line measurement; B) measuring the hydrogen concentration in the exit mixture with an in-line measurement; C) measuring the phenol concentration in the exit mixture with an in-line measurement; D) measuring the cyclohexanol concentration in the exit mixture with an in-line measurement; D2) Optionally, measuring the cyclohexanone concentration in the exit mixture by in-line measurement; E) comparing the pressure in the reactor of step A) with a first set value; F) comparing the hydrogen concentration obtained in step B) with a second set value; G) comparing the phenol concentration obtained in step C) with a third set value; H) comparing the cyclohexanol concentration obtained in step D) with a fourth set value; H2) optionally comparing the cyclohexanone concentration obtained in step D2) with a fifth set value; I) performing an adaptation of at least one process parameter in an automatic operating mode when the pressure in the reactor obtained in step A) deviates from a first set value; J) performing an adaptation of at least one process parameter in an automatic operating mode if the hydrogen concentration in the exit mixture obtained in step B) deviates from a second set value.

2. The at least one process parameter to be adapted in steps I) and / or J) is a) the flow rate of the incoming mixture, and / or b) Purge flow rate The method of claim 1 , comprising:

3. 3. The method of claim 1, wherein the incoming mixture further comprises a diluting gas.

4. 4. The method according to claim 3, wherein apart from the hydrogen in the exiting mixture which is recycled to the entering mixture, at least a portion of the diluting gas present in the exiting mixture is also recycled to the entering mixture.

5. K) adapting at least one process parameter if the phenol concentration obtained in step C) is higher than a third set value. The method of any one of claims 1 to 4, further comprising:

6. L) performing an adaptation of at least one process parameter if the cyclohexanol concentration obtained in step D) is higher than a fourth set value; and / or L2) adapting at least one process parameter if the cyclohexanone concentration obtained in step D2) is lower than a fifth set value. The method of any one of claims 1 to 5, further comprising:

7. 7. The method according to claim 5, wherein at least one process parameter adapted in steps K), L) and / or L2) is adapted in an automatic operating mode.

8. At least one process parameter for which adaptation is performed in steps K, L) and / or L2) is a) the temperature of the cooling medium; b) the hydrogen concentration in the exit mixture; c) the concentration of the dilution gas in the exit mixture, and d) A combination of these The method according to any one of claims 5 to 7, wherein the compound is selected from the group consisting of:

9. The at least one process parameter for which adaptation is performed in step K is a) an increase in the temperature of the cooling medium; b) an increase in hydrogen concentration in the exit mixture; c) A reduction in the concentration of the dilution gas in the exit mixture; and d) A combination of these The method of claim 8, wherein the compound is selected from the group consisting of:

10. At least one process parameter for which adaptation is performed in steps L and / or L2 is a) A decrease in the temperature of the cooling medium, b) A decrease in the hydrogen concentration in the exit mixture; c) an increase in the concentration of the dilution gas in the exit mixture; and d) A combination of these The method of claim 8, wherein the compound is selected from the group consisting of:

11. 11. The method according to claim 1, wherein the cooling medium in the indirect heat exchange according to claim 1 is water, and the temperature of the hydrogenation reactor is set by the temperature of the boiling water used to absorb the heat of the exothermic gas phase hydrogenation, and the boiling point of the water is set by adjusting the pressure of the boiling water.

12. The method according to claim 11, wherein the pressure of the boiling water is adjusted within the range of 0.15 MPa to 1.5 MPa.

13. A process according to any one of claims 8 to 12, wherein the hydrogen concentration in the exiting mixture is increased by increasing the amount of hydrogen charged to the process.

14. 14. The method according to any one of claims 1 to 13, wherein the industrial scale hydrogenation reactor is a shell-and-tube heat exchange reactor using water as a cooling medium and wherein water vapor is produced by said reactor as a result of indirect heat exchange.

15. 15. An industrial scale phenol hydrogenation plant adapted to carry out the method according to any one of claims 1 to 14, the plant comprising: a hydrogenation reactor comprising a catalyst containing palladium, at least one inlet allowing an inlet mixture comprising phenol and hydrogen to be charged to the reactor, at least one outlet capable of releasing an exit mixture comprising cyclohexanone, cyclohexanol, phenol and hydrogen; a line for recycling at least a portion of the hydrogen in the exiting mixture to the entering mixture; a recovery section capable of recovering cyclohexanone from the exit mixture, a purge section capable of releasing at least a portion of the hydrogen from the exit mixture as a purge stream; - an indirect heat exchanger capable of transferring at least a portion of the reaction heat of the phenol hydrogenation reaction to a cooling medium; Equipped with The plant has the following parameters: - the pressure in the hydrogenation reactor; - hydrogen concentration in the exit mixture; - phenol concentration in the exit mixture; - Cyclohexanol concentration in the exit mixture The plant is characterized by being equipped with an in-line measuring device capable of measuring all of the above.

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