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 mainly relates to a plant and a catalytic process for industrial-scale phenol hydrogenation in which cyclohexanone and cyclohexanol are formed.

Background Art

[0002] Most cyclohexanone is consumed in the production of ε-caprolactam, an intermediate in nylon 6 production. The mixture of cyclohexanone and cyclohexanol is mainly 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, the production of cyclohexanone was started on an industrial scale in parallel with the commercial production of ε-caprolactam, adipic acid, nylon 6 and nylon 6,6. Since then, the production volume of cyclohexanone has been continuously increasing, and today the annual production volume 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 byproducts and require significant energy input. The third route, hydrogenation from 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, for example, using a catalyst containing Ni, to form cyclohexanol, which is then dehydrogenated to obtain cyclohexanone. In the "one-step" process, phenol is hydrogenated directly and with high selectivity to produce cyclohexanone. This "one-step" phenol hydrogenation process is known to combine high production yield with low energy demand. In the "one-step" process, cyclohexanone is prepared by catalytic hydrogenation from phenol in a phenol hydrogenation reactor using a catalyst, for example, containing Pt or Pd. During phenol hydrogenation, cyclohexanol is formed as the main byproduct, separate from trace amounts of other compounds. Phenol hydrogenation can be carried out in the liquid phase or the vapor phase (often called the gas phase). For example, see “Cyclohexanol and Cyclohexanone”, Kirk-Othmer Encyclopedia of Chemical Technology, e.g., 3rd Edition, 1979, Vol. 7, pp. 410-416; I. Dodgson et al., “A low Cost Phenol to Cyclohexanone Process”, Chemistry and Industry, December 1989, Vol. 18, pp. 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 its two main products, cyclohexanone and cyclohexanol, is expressed by the following stoichiometric formula: C6H5OH + 2H2 → C6H 10 O + heat Phenol Hydrogen Cyclohexanone C6H5OH + 3H2 → C6H 11 OH + heat Phenol Hydrogen Cyclohexanol It can be expressed as follows.

[0006] The (endothermic) dehydrogenation of cyclohexanol to cyclohexanone is expressed by the following stoichiometric formula: C6H 11 OH + heat → C6H 10 O + H2 Cyclohexanol, cyclohexanone, hydrogen It can be expressed as follows.

[0007] British Patent Application Publication No. 890095 describes a method for preparing cyclohexanone from phenol by catalytic hydrogenation, comprising passing gaseous phenol with hydrogen through 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] British Patent Application Publication No. 1316820 describes a method for producing cyclohexanone by hydrogenation of phenol in the vapor phase in the presence of a catalyst consisting of a palladium layer and a support on which an alkali metal carbonate as a co-catalyst is deposited. The document teaches that the selectivity for cyclohexanone is affected by the presence of the co-catalyst.

[0009] The British Patent Application Publication No. 1332211A describes a method for producing cyclohexanone by catalytic hydrogenation of phenol in the vapor phase at a temperature of 75 to 250°C in the presence of a catalyst containing 0.1 to 5% by weight of activated alumina-supported platinum group metal and 5 to 50% by weight of alkaline earth metal hydroxide and / or alkaline earth metal oxide, based on the total weight of the catalyst.

[0010] International Publication No. 2016075047A1 describes an industrial-scale continuous process for the production and recovery of cyclohexanone from phenol and hydrogen. The document teaches that the net steam consumption of such a method can be less than 1.5 kg of steam per 1 kg of cyclohexanone produced.

[0011] International Publication No. 2011073233A1 describes a method for continuously preparing cyclohexanone, cyclohexanol, or mixtures thereof by catalytic hydrogenation of phenol supplied to a reactor equipped with a dopant-containing supported hydrogenation catalyst. The document teaches that supplying water to the reactor during the hydrogenation of phenol can increase the conversion rate of phenol and / or increase the selectivity for cyclohexanone and / or cyclohexanol.

[0012] The combination of (very) high selectivity for cyclohexanone and therefore (very) low selectivity for cyclohexanol in the reactor section, along with a (very) high single-pass conversion rate of phenol, is highly desirable due to its economic advantages: smaller equipment for recovering cyclohexanol and phenol from production logistics and for cyclohexanol dehydrogenation, and reduced energy consumption for recovering cyclohexanol and phenol from production logistics and for cyclohexanol dehydrogenation. As a result, the inventors have found that industrial-scale phenol hydrogenation plants for cyclohexanone production should be operated within an optimal process window that combines high single-pass selectivity for cyclohexanone with a high single-pass conversion rate of phenol. Research leading to the present invention has shown that the optimal combination of single-pass selectivity for cyclohexanone and single-pass conversion rate of phenol for a phenol hydrogenation plant is determined by numerous variable factors, such as the origin, source, and quality of the feedstock phenol and hydrogen, the type of catalyst, the type of process, the plant layout, and the desired production rate.

[0013] In fact, an increase in one-pass selectivity for cyclohexanol, and consequently a decrease in one-pass selectivity for cyclohexanone, is usually observed as an increase in the one-pass conversion rate of phenol. In particular, when the one-pass conversion rate of phenol is (very) high, a rapid increase in one-pass selectivity for cyclohexanol is observed, making it impossible to achieve both 100% one-pass conversion rate for phenol and 100% one-pass selectivity for cyclohexanone simultaneously.

[0014] Catalyst systems used in phenol hydrogenation plants are susceptible to significant 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 flow time (TOS). In commercial phenol hydrogenation plants, the flow time of a given catalyst batch can range from a few months to many years. While not theoretically bound, it is hypothesized that aging is caused by the deposition of organic compounds on the catalyst (often called coking) and / or by changes in the crystal structure of the catalyst metal.

[0015] Generally, the cost of cyclohexanone production by hydrogenation of phenol continues to rise due to catalyst aging (e.g., decreased selectivity for cyclohexanone, decreased single-pass conversion rate of phenol, decreased productivity, and increased hydrogen purging loss).

[0016] Generally, the end of reaction (EOR) of a given hydrogenation catalyst batch is determined by technical (e.g., reaching the maximum allowable temperature or 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 substances present on the catalyst; b) heat treatment in an oxygen-containing atmosphere to oxidize carbonaceous substances present on the catalyst; and c) cooling of the catalyst after step a) and / or step b).

[0018] Catalyst replacement and regeneration are time-consuming and costly operations (for example, due to reduced production), so maximizing catalyst life is highly desirable.

[0019] To date, hydrogenating phenol to cyclohexanone while maintaining high production rates and low production costs, and while simultaneously achieving high single-pass selectivity for cyclohexanone and a high single-pass conversion rate for phenol throughout the entire lifespan of the catalyst batch, remains a significant challenge.

[0020] The preparation of cyclohexanone from phenol has been known for decades, and pathways to improve known preparation methods have been thoroughly investigated over the years, but currently known industrial processes (which are generally continuous) still have shortcomings.

[0021] In particular, a problem in known continuous phenol hydrogenation processes is the decrease in single-pass selectivity for cyclohexanone over time. In addition, heterogeneous hydrogenation catalysts gradually lose activity, resulting in a decrease in the single-pass conversion rate of phenol. This leads to a decrease in the cyclohexanone production rate over time. This is not only disadvantageous because it makes it impossible to maintain a stable production rate, but it also increases the conversion cost of dehydrogenating the produced cyclohexanol back into cyclohexanone. Therefore, in phenol hydrogenation reactions, it is necessary to compensate for the decrease in selectivity and activity of heterogeneous hydrogenation catalysts.

[0022] A further object of the present invention is to provide a method for preparing cyclohexanone in a continuous process that can function as an improvement over known methods, in particular a method that overcomes one or more drawbacks of known methods such as those mentioned above.

[0023] A further objective is to provide a method for preparing cyclohexanone in which fewer byproducts are produced, in particular, fewer cyclohexanols that must be converted to cyclohexanone by high-temperature catalytic dehydrogenation, which would result in undesirable byproducts such as benzene.

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

[0025] More specifically, an object of the present invention is to prepare cyclohexanone while maintaining a target cyclohexanone production rate and having improved cyclohexanone selectivity compared to conventional methods operated in the same production equipment. This provides advantages such as a more environmentally friendly method and a less expensive method compared to conventional methods operated in the same production equipment. One or more further objects may become apparent from the remaining description.

[0026] The inventors have discovered a method for solving or at least significantly reducing the above-mentioned objects.

Summary of the Invention

[0027] The present invention is a method for controlling a vapor-phase hydrogenation reaction with heat generation of phenol catalyzed by a catalyst containing palladium in an industrial-scale hydrogenation reactor, wherein an inlet mixture containing phenol and hydrogen is charged into the reactor, and an outlet mixture containing cyclohexanone, cyclohexanol, phenol and hydrogen is discharged from the reactor, and at least a part of the hydrogen in the outlet mixture is recycled to the inlet mixture, cyclohexanone is recovered from the outlet mixture, and at least a part of the hydrogen derived from the outlet mixture is discharged as a purge stream, and at least a part of the reaction heat of the phenol hydrogenation reaction is transferred to a cooling medium by indirect heat exchange, and the method comprises: A) measuring the pressure in the reactor by in-line measurement; B) measuring the hydrogen concentration in the outlet mixture by in-line measurement; C) measuring the phenol concentration in the outlet mixture by in-line measurement; D) measuring the cyclohexanol concentration in the outlet mixture by in-line measurement; D2) optionally, measuring the cyclohexanone concentration in the outlet mixture by in-line measurement; E) A step of comparing the pressure inside the reactor in step A) with a first set value; F) A step of comparing the hydrogen concentration obtained in step B) with a second set value; G) A step of comparing the phenol concentration obtained in step C) with a third set value; H) A step of comparing the cyclohexanol concentration obtained in step D) with a fourth set value; H2) Depending on the case, compare the cyclohexanone concentration obtained in step D2) with a fifth set value; I) If the pressure inside the reactor obtained in step A) deviates from the first setpoint, the automatic operation mode adjusts at least one process parameter; J) If the hydrogen concentration in the discarded mixture obtained in step B) deviates from the second set value, the automatic operation mode adjusts at least one process parameter. This provides a method that includes [something].

[0028] The inventors have found that, due to aging of the catalyst system and other process disturbances, when optimal process conditions for selectivity and conversion rate are desired, the pressure and yield of the phenol hydrogenation reactor should be carefully monitored, and, if necessary, process conditions (e.g., flow rate of the input mixture, particularly the phenol and hydrogen pressures, flow rate of the purge flow, reactor temperature, concentration of diluent gas, etc.) should be immediately adjusted to maintain production under optimal conditions. The present invention achieves this by introducing in-line measurement of process parameters that have proven to be critical, and by linking this measurement with automatic comparison and adjustment based on predefined setpoints and adjustment rules. As shown in the examples, the present invention enables the phenol hydrogenation reactor to operate under optimal conditions for several months to several years at a constant high production rate without human intervention. This is remarkable and not something that could have been predicted from the prior art.

[0029] In the prior art, the phenol hydrogenation process, even when controlled, relied on offline measurements (manual measurements) to determine the concentrations of individual components such as hydrogen, phenol, cyclohexanone, or cyclohexanol in the gas mixture released from the phenol hydrogenation reactor. Offline measurement refers to measurements performed by manually operating a measurement system to obtain a measurement value at any given point in time, for example, by an operator withdrawing a sample from the process at a given moment and injecting the collected material into a type of analyzer that measures the concentration of components in the sample. The results of these individual measurements are used by plant operators to determine whether process conditions can remain unchanged or need to be changed.

[0030] The implementation of offline measurements depends on the operator's skill, and therefore the results of such measurements may vary depending on the operator's skill. In addition, offline measurements are typically performed infrequently (e.g., a few times a week or a few times a day). Inline measurements (automated measurements) used in this invention are far more accurate because they are devoid of human intervention and can be performed continuously or at least frequently. Inline measurements enable periodic or even continuous monitoring of parameters without operator interference, thereby allowing for frequent measurement repetition and stable measurements unaffected by operator skill.

[0031] Using conventional processes that do not employ in-line measurement, the lack of sufficiently high measurement frequency and statistical data made it extremely difficult—or impossible—to accurately track when unacceptable results began to appear. Research leading to the present invention showed that when and how quickly decisions regarding modifications to process conditions are made have a surprisingly strong impact. It was found that inaccurate and timely decisions regarding modifications to process conditions negatively impact the cost and quality of cyclohexanone production (including shortening the lifespan of the hydrogenation catalyst). The present invention overcomes these shortcomings by enabling immediate, automated reactions to modify process conditions measured in-line and compared to predefined setpoints. This not only has the advantage of achieving much more consistent and optimized process control without requiring human intervention, but most importantly, it leads to reduced costs (including extended hydrogenation catalyst lifespan) and improved product quality because optimal process conditions are maintained almost continuously throughout the entire process.

[0032] It was remarkable that the method of the present invention produced high production rates for phenol hydrogenation reactors over long periods at low production costs while severely limiting human interference. This is achieved by performing a combination of several inline measurements of specific process parameters and linking these measurements to an automated decision process that results in advantageous automated process control where the process parameters are automatically adapted according to the teachings of the present invention. Until now, the above was considered impossible for phenol hydrogenation reactors because automation would reduce versatility compared to the flexibility and diversity of tasks that can be performed by well-trained operators.

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

[0034] In the method of the present invention, both phenol and hydrogen gas are continuously supplied to an industrial-scale hydrogenation reactor for the production of cyclohexanone from phenol. The entire influent feed, including at least phenol and hydrogen gas, is referred to herein as the “influent mixture.” The influent mixture may enter the reactor in a single line as a pre-mixed mixture, or it may be mixed within the reactor. In the hydrogenation reactor, phenol is converted in the vapor phase by a catalytic reaction under the influence of a palladium-containing catalyst, mainly to cyclohexanone, and partly to cyclohexanol. At least some of the heat from the exothermic phenol hydrogenation reaction is transferred to a cooling medium by indirect heat exchange. The production feed exiting the hydrogenation reactor includes unconverted phenol and hydrogen, in addition to the desired product, cyclohexanone, and the by-product, cyclohexanol. The production feed exiting the hydrogenation reactor, and any (further processed) production feed derived therefrom, are collectively referred herein as the “exit mixture.” At least a portion of the unconverted hydrogen in the exit mixture is reused in the entry mixture, and at least a portion is purged. Cyclohexanone is recovered from the exit mixture as the main product.

[0035] According to the present invention, the phenol hydrogenation reactor is equipped with an in-line measuring device that enables the measurement of the pressure inside the reactor in step A). ​​The pressure can be measured by various (electrical) devices well known to those skilled in the art. In particular, strain gauges are often used for pressure measurement. Piezoresistive strain gauges (also called piezoresistive 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 may vary in the hydrogenation reactor. A very convenient location, also preferred for the present invention, is the reactor inlet, before the feed material comes into contact with the catalyst. For example, in a multi-tube (shell-tube) hydrogenation reactor, the in-line pressure measurement may be located inside the reactor hood before the feed material enters the tubes filled with catalyst. In this method, the pressure measured corresponds to the total pressure of the feed material charged into the multi-tube hydrogenation reactor. However, other possibilities include the in-line pressure measurement at the discharge of the hydrogenation reactor or at another location within the reactor. The optimal pressure setting will vary slightly depending on where the pressure is measured, as the pressure typically drops as the reaction mixture passes through the catalyst. If the pressure is measured downstream of the catalyst, it may also change as the catalyst ages.

[0036] Within the scope of the present invention, the term "inline measurement" refers to automatic measurement using a permanently installed device, and therefore generally does not require the action of an operator, in contrast to offline measurement (often called "manual measurement").

[0037] Analysis can be performed by continuous or discontinuous methods. Continuous analysis methods involve placing the sensor directly in, for example, a reactor, a line containing process fluid, or any other instrument, or in, for example, a small side flow drawn from a reactor, a line containing process fluid, or any other instrument. Discontinuous analysis involves drawing the sample from, for example, a reactor, a line containing process fluid, or any other instrument and injecting it into the analytical instrument. Continuous analysis yields a continuous response signal, while discontinuous methods yield responses at different, distinct "points in time."

[0038] Within the scope of this invention, the term "setpoint" refers to a value relating to an essential system variable or process value. If such a variable deviates from its setpoint, it constitutes a basis for process control. Action may be taken only if the measured value exceeds or falls below a given setpoint. In other cases, action is required if the measured value deviates from the setpoint by a minimum amount (including the dead zone).

[0039] In step E) of the method of the present invention, the pressure in the reactor measured in step A) is compared to a first setpoint. Those skilled in the art will know the optimal pressure for the industrial-scale hydrogenation reactor they are operating. The optimal pressure is related to the chemical action, the type of catalyst, and the reactor configuration (e.g., the length of the tubing). If the optimal pressure setpoint is unknown, it can be easily determined by conducting trial runs at various reactor pressures to determine the reactor pressure that provides the optimal phenol conversion rate and cyclohexanone selectivity. If the pressure is too high, the selectivity will decrease. If the pressure is too low, the processing capacity and therefore the yield of cyclohexanone will be too low. The target pressure, i.e., the first setpoint, should always be above atmospheric pressure and preferably in the range of 0.1 to 1 MPa, more preferably 0.15 to 0.6 MPa, and most preferably 0.2 to 0.5 MPa.

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

[0041] Steps A), E), and I) are preferably performed at least once a day, more preferably at least once an hour, most preferably at least once a minute, or even every 5 to 30 seconds, in order to achieve optimal control. Since the above steps need to be performed at the above frequencies to achieve optimal results, the present invention anticipates that the adaptation in step I) will be performed in autonomous driving.

[0042] Process parameters that can be adapted in automatic operation in step I) to return the pressure to near or at the setpoint in a specific effective manner include a) adjusting the flow rate of the inlet mixture and / or b) adjusting the flow rate of the purge flow.

[0043] The flow rate of the incoming mixture can be controlled, in particular, by controlling the inflow of hydrogen-containing flow and / or phenol-containing flow into the reactor (if these components enter the reactor through separate channels). If the pressure inside the reactor is measured to be higher than the first setpoint in step E), the flow rate of the incoming mixture can be automatically reduced in step I). Conversely, if the pressure inside the reactor is measured to be lower than the first setpoint in step E), the flow rate of the incoming mixture can be automatically increased in step I).

[0044] As described above, in the plant or method according to the present invention, at least a portion of the hydrogen derived from the discharge mixture is released as a purge flow. This purge flow is an important method for regulating not only the hydrogen concentration (related to step J below) but also the pressure in the reactor in step I. The flow rate of the purge flow can be adjusted by a valve that controls the flow rate of the purge flow. If the pressure in the reactor is measured to be higher than a first setpoint in step E), the flow rate of the purge flow can be automatically increased in step I). Conversely, if the pressure in the reactor is measured to be lower than a first setpoint in step E), the flow rate of the purge flow can be automatically decreased in step I). In this way, the pressure can be returned to near or to the setpoint in a specific and effective manner simply by opening and closing a valve.

[0045] The term "phenol hydrogenation reactor" as used herein has the common meaning in the art. A phenol hydrogenation reactor is a closed space in which the hydrogenation of phenol takes place. The conversion from phenol to the reaction product in a phenol hydrogenation reactor may be complete or partial. A phenol hydrogenation reactor can 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 sequentially (e.g., to increase the overall phenol conversion rate).

[0046] The hydrogenation reactor may be any type of reactor suitable for the hydrogenation of phenol. In particular, the reactor can be selected from a fixed-bed reactor, a slurry reactor, a multi-tube heat exchange reactor having a catalyst in the tubes and generating steam, 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-tube heat exchange reactor, most preferably using water as a cooling medium, and as a result, water vapor is generated. More preferably, the vapor phase hydrogenation of phenol according to the present invention is carried out in a vertical multi-tube heat exchange reactor.

[0048] In one embodiment, the phenol hydrogenation reactor is a multi-tube heat exchange reactor configured with catalyst material arranged inside tubes. Preferably, it is a vertical multi-tube heat exchange reactor. A gaseous mixture containing hydrogen and phenol is supplied to the tubes, and a cooling medium circulates around the tubes on the outside.

[0049] As used herein, the term “heat exchanger” refers to a device for transferring heat from one flow to another. Heat exchangers may be direct (flows are mixed) or indirect (flows 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 the flow in the first chamber through the partition to the flow in the second chamber. Each chamber may 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 multi-tube, plate, and tubular heat exchangers.

[0050] Within the scope of this invention, the expression "at least a portion of the reaction heat of the phenol hydrogenation reaction is transferred to ~" refers to the heat released by the reaction between phenol and hydrogen and transferred through the walls of the heat exchanger. The reaction between phenol and hydrogen is an exothermic reaction (it produces heat).

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

[0052] The number of reactor tubes in the preferred (vertical) multi-tube 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 method of the present invention, the number of reactor tubes in the (vertical) vertical multi-tube heat exchange reactor is 25 to 20,000.

[0053] The length of the reactor tubes in the preferred (vertical) multi-tube heat exchange reactor according to the present invention is typically longer than 0.25 m. Preferably, the length is longer than 0.5 m. More preferably, the length is longer than 1.0 m. The length of the reactor tubes 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 method of the present invention, the length of the reactor tubes in the (vertical) vertical multi-tube heat exchange reactor is 1.0 to 9.0 m, which allows for a good phenol conversion rate without requiring a large amount of catalyst. Preferably, all reactor tubes in the multi-tube heat exchange reactor are (almost) the same length.

[0054] The inner diameter of the shell of a multitube 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 a (vertical) multitube heat exchange reactor is typically less than 10 mm. Preferably, the inner diameter of the shell is less than 8 mm. More preferably, the inner diameter of the shell is less than 6 mm. Typically, in the method of the present invention, the multitube heat exchange reactor shell The inner diameter is 0.2 to 6 m, which allows for good reactor operation capacity while also enabling transportation.

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

[0056] In one embodiment, the temperature of the entry mixture containing phenol and hydrogen charged into the hydrogenation reactor is higher than 85°C. Preferably, the temperature is higher than 100°C. More preferably, the temperature is higher than 125°C. This allows for an optimal conversion rate. To avoid catalyst depletion, decomposition, and byproduct formation, the temperature of the entry mixture should be less than 300°C, preferably less than 250°C, and 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 temperature measurement in the chemical industry are resistance thermometers (also called RTDs) and thermocouples for temperatures below 600°C. The Pt100 type sensor, which belongs to the RTD group, is the most commonly used.

[0058] In a multi-tube heat exchange reactor, heat is generated by the hydrogenation of phenol, which heats the mixture of components inside the tubes. The heat from the mixture of components is removed by indirect cooling using a cooling medium.

[0059] The temperature of the mixture containing cyclohexanone and cyclohexanol released from the hydrogenation reactor should be above 60°C. Preferably, the temperature should be above 80°C. More preferably, the temperature should be above 100°C. To avoid catalyst depletion, decomposition, and byproduct formation, the temperature of the mixture containing cyclohexanone and cyclohexanol released from the hydrogenation reactor should be below 260°C. Preferably, the temperature should be below 240°C. More preferably, the temperature should be below 220°C. The best results are achieved when the temperature of the mixture containing cyclohexanone and cyclohexanol released from the multitubular heat exchange reactor is between 100 and 220°C. The lower temperature limit is determined by the reaction rate in the phenol hydrogenation reaction. At lower temperatures, the reaction rate will be lower, and as a result, a large amount of catalyst and a large reactor will be required. In addition, the temperature of the heated cooling medium will be lower, and as a result, many potential uses for the absorbed heat will be lost. The upper temperature limit is determined, in particular, 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 byproducts (e.g., tar and benzene) caused by the decomposition of phenol.

[0060] According to the present invention, the temperature inside the hydrogenation reactor is controlled by the temperature of the cooling medium used for indirect heat exchange and the temperature of the influent mixture. A higher reactor temperature results in a higher reaction rate, but above a certain temperature threshold, a decrease in selectivity occurs, leading to the formation of more undesirable byproducts.

[0061] In this regard, it should be noted that it is not possible to accurately measure or report the actual temperature inside the reactor. The temperature in the catalyst bed is not constant; it varies both in location and over time. For example, the local temperature inside the catalyst bed of a multitube heat exchange reactor used for phenol hydrogenation is a result of the heat generated by the hydrogenation of phenol and the heat transferred to the cooling medium. The temperature varies both longitudinally and radially within the catalyst bed. Generally, the temperature in the longitudinal direction of the catalyst bed passes through a maximum value ("peak"). The shape of the temperature profile and the location of the highest temperature within the bed vary depending on numerous process parameters, such as the feed flow rate, the concentrations of both phenol and hydrogen in the feed, the type and state of the catalyst, and the temperature of the cooling medium outside the catalyst bed (outside the tubes in the case of a multitube heat exchange reactor). As time progresses, the location of the highest temperature within the bed gradually shifts from near the inlet of the gaseous feed into the bed to near the outlet of the gaseous products from the bed. In general, the exact temperature at any point within the catalyst bed is unknown. However, by installing a series of temperature sensors (e.g., Pt100 type sensors) along the longitudinal direction of the catalyst bed, it is possible to obtain an overall indication by measuring the temperature at various heights within 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 on a smaller scale within 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, degrading the catalyst and forming undesirable byproducts. The inventors have found that good results are achieved when the temperature of the cooling medium is maintained within the above-described temperature range for the mixture containing cyclohexanone and cyclohexanol released 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. This 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 concentrations of phenol or cyclohexanol in the exit mixture exceed the third and fourth setpoints, respectively, or when 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 reactor temperature should be increased to increase the conversion rate to cyclohexanone. However, care must be taken not to increase the temperature and conversion rate too much, as this may adversely affect the selectivity for cyclohexanone. If the concentration of cyclohexanol in the exit mixture exceeds the fourth setpoint, the reactor temperature should be decreased to reduce the conversion rate to cyclohexanol.

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

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

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

[0067] The inventors have found that the boiling water pressure can be adjusted within a range of 0.1 MPa to 5 MPa depending on the desired cooling amount and reactor temperature. Preferably, in the method 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 method 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 method of the present invention, the boiling water pressure is adjusted within a range of 0.15 MPa to 1.5 MPa.

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

[0069] The method of the present invention uses an industrial-scale hydrogenation reactor for the production of cyclohexanone. This means a phenol hydrogenation reactor capable of charging several thousand tons (kilotons per year; kta) of phenol per year. In embodiments of the present invention, the industrial-scale phenol hydrogenation reactor is capable of, or has been charged, more than 10,000 tons (10 kta) of phenol per year. The amount of phenol charged into the industrial-scale phenol hydrogenation reactor does not have to be the maximum amount. Preferably, the amount of phenol charged into the industrial-scale phenol hydrogenation reactor is 20 kta to 350 kta. More preferably, the amount of phenol charged into the industrial-scale phenol hydrogenation reactor is 25 kta to 250 kta.

[0070] The operating mode of a phenol hydrogenation reactor can be selected from feed batch operation and continuous operation. In feed batch operation mode, hydrogen is added over time to a phenol hydrogenation reactor that has been pre-charged with phenol, and the reaction product remains in the reactor until the reaction is complete. In continuous operation mode, both hydrogen and phenol are continuously charged into the phenol hydrogenation reactor, and the reaction product and unconverted feedstock are continuously discharged from the phenol hydrogenation reactor. Preferably, industrial-scale phenol hydrogenation reactors are operated in continuous mode.

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

[0072] Within the scope of the present invention, the term "at least a portion" in the foregoing expression preferably means that at least 10%, more preferably at least 50%, and most preferably at least 85% of the hydrogen in the exit mixture is reused in the entry mixture.

[0073] In step B) of the method according to the present invention, the hydrogen concentration in the exit mixture is measured by in-line measurement. The "hydrogen concentration in the exit mixture" can mean the hydrogen concentration in any hydrogen-containing flow downstream of the phenol hydrogenation reactor, since this concentration will be proportional to the hydrogen concentration in the mixture directly exiting the reactor. In practice, particularly convenient locations for measuring the hydrogen concentration in the exit mixture are in the line transporting the gaseous purge flow (see Example 1 below, line in Figure 2 [4]), in the line reusing some of the hydrogen in the exit mixture into the inlet mixture (line in Figure 2

[22] ), and directly from the compression section. It's out These locations include lines transferring compressed gas mixtures (line

[32] in Figure 2), lines directly exiting the phenol hydrogenation reactor (line

[29] in Figure 2), or lines exiting heat exchangers or gas-liquid separators where present (line

[30] or

[31] in Figure 2). The flows in lines

[32] ,

[22] , and [4] in Figure 2 are compressed and contain only trace amounts of phenol, cyclohexanol, and cyclohexanone. These flows primarily contain hydrogen and inert components. Therefore, these 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 using various techniques, including chromatography (especially gas chromatography (GC)). ) However, it is commonly used to separate individual components in a sample that is a mixture of components, and these components are then quantified by a detector (e.g., a thermal conductivity detector (TCD), also known as a casarometer). The detector signal is processed by a controller to calculate the hydrogen concentration. For continuous measurement (calibrated) of hydrogen concentration in gaseous mixtures containing a certain limit of components, thermal conductivity detectors are 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 to a second set value. Those skilled in the art will know the hydrogen concentration that should be maintained in their industrial-scale hydrogenation reactor to obtain the optimal phenol conversion rate and cyclohexanone selectivity. If the hydrogen concentration in the reactor is too high, the selectivity for cyclohexanone will decrease as more cyclohexanol is formed (see the reaction scheme and stoichiometry above). If the hydrogen concentration in the reactor is too low, the phenol conversion rate will be too low. The optimal set value for the hydrogen concentration in the effluent mixture may vary depending on the age of the catalyst and other process parameters. For example, when using a brand new catalyst, it is often recommended to adjust the reaction rate by operating with only a small amount of hydrogen present in the effluent mixture. If the catalyst is aged, more hydrogen needs to be added to obtain a sufficient phenol conversion rate.

[0076] For the mixture coming directly out of the hydrogenation reactor (lines

[29] and

[30] in Figure 2), a second setpoint for the optimal hydrogen concentration is preferably in the range of 0.1 to 40 mol%, particularly 1 to 30 mol%, based on the total number of moles of components present in the mixture. If the gas composition in the line that reuses some of the hydrogen from the exit mixture into the entry mixture is measured to determine the H2 concentration of the exit mixture, the second setpoint may be in the range of 1 to 100 mol% (if no diluting gas, e.g., inert components, is added, the reuse can be approximately 100 mol%, preferably 2 to 85 mol%, more preferably 3 to 70 mol%, and most preferably 5 to 50 mol%). A lower value is used if unused (or regenerated) catalyst is present, and the upper limit takes catalyst aging into account. The gas composition in molar ratio is identical at lines

[31] ,

[32] ,

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

[0077] In step J) of the method of the present invention, if the hydrogen concentration in the exit mixture deviates from a second setpoint, automatic operation is performed to adjust at least one process parameter. As described above, the hydrogen concentration can be too high or too low and varies depending on the catalyst state and other process parameters. Therefore, careful and frequent monitoring should be performed. In practice, the second setpoint may be a concentration range around the optimal setpoint. For example, the range may be ±25%, ±10%, ±1%, or ±0.5% from a predefined setpoint within the range shown in the previous paragraph.

[0078] Steps B), F), and J) are preferably performed at least once a day, more preferably at least once an hour, most preferably at least once every minute or even every 5 to 30 seconds, in order to achieve optimal control. Since the above steps need to be performed at this frequency to achieve optimal results, the present invention anticipates that the adaptation in step J) will be performed in autonomous driving.

[0079] Process parameters that can be automatically adapted in step J) to return the hydrogen concentration to near or at a setpoint in a specific effective manner include a) adjusting the flow rate of the incoming mixture and / or b) adjusting the flow rate of the purge flow.

[0080] In this context, adjusting the flow rate of the incoming mixture specifically means adjusting the inflow of the hydrogen-containing flow into the reactor when the components hydrogen and phenol enter the reactor in separate flows. 2 If the measured value is higher than the set value, the flow rate of the incoming mixture can be automatically reduced in step J). Conversely, if the hydrogen concentration in the outgoing mixture is higher in step F), 2 If the measured value is lower than the set value, the flow rate of the incoming mixture can be automatically increased in step J).

[0081] As described above, in the plant or method according to the present invention, at least a portion of the hydrogen derived from the exit mixture is released as a purge flow. This purge flow is an important method for adjusting the hydrogen concentration. The main purpose of the purge flow is to remove inert components (dilution gases). Increasing the flow rate of the purge flow lowers the concentration of inert components and therefore increases the hydrogen concentration (and vice versa). The flow rate of the purge flow can be adjusted by a valve that controls it. When the hydrogen concentration in the exit mixture reaches the first stage in step F) 2 If the measured value is higher than the set value, the purge flow rate can be automatically reduced in step J). Conversely, if the hydrogen concentration in the discharge mixture is higher in step F), 2 If the measured value is lower than the set value, the purge flow rate can be automatically increased in step J). In this way, the hydrogen concentration in the exit mixture can be returned to near or to the set value in a specific and effective manner simply by opening and closing the valve.

[0082] According to a preferred embodiment of the present invention, the incoming mixture further includes a dilution gas. The dilution gas should be an inert gas, such as methane and / or nitrogen. Being inert, the dilution gas is present in both the incoming and outgoing mixtures. In this specification, the dilution gas may also be simply referred to as the “inert component.” Those skilled in the art often prefer to operate the hydrogenation process at a constant pressure, regardless of changes in the amounts (and associated pressures) of phenol and hydrogen in the supply flow, although this may require further adaptation according to the process control implemented by the present invention. For example, if the combined partial pressure of phenol and hydrogen supplied to the reactor is 0.21 MPa, the partial pressure of the inert component supplied to the reactor may be 0.09 MPa to reach a desired pressure setpoint of 0.30 bar (if no other components are supplied to the reactor). As hydrogen is consumed in the hydrogenation reactor, the ratio of hydrogen to the inert component changes by reaction in the reuse gas stream containing unreacted hydrogen and inert component compared to the incoming mixture. To maintain a sufficient reaction rate, the inventors found that various adjustments to the amount of inert component and hydrogen excess are possible. Consequently, a decrease in the concentration of the inert component leads to an increase in the hydrogen concentration, and therefore an increase in the reaction rate.

[0083] In a more preferred embodiment, in addition to the hydrogen in the exit mixture that is reused in the incoming mixture, at least a portion of the diluent gas present in the exit mixture is also reused in the incoming mixture. This presents a particular convenient method for minimizing the amount of unused diluent gas that should be supplied to the incoming mixture. By using the diluent gas in the incoming mixture, the molar ratio of this diluent gas to hydrogen gas is higher during purging than in the incoming mixture. This presents a particular convenient possibility for adjusting the hydrogen concentration. In steps C) and D) of the method of the present invention, the concentrations of phenol and cyclohexanol are measured by in-line measurement, respectively. Optionally, the method of the present invention may include 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. Therefore, steps C), D) and optionally D2 can be performed in one step. The concentrations of phenol, cyclohexanol, and cyclohexanone can be measured by various techniques. Chromatographic methods (especially gas chromatography (GC)) ) However, it is commonly used to separate individual components in a sample that is a mixture of components, and these components are then quantified by a detector (for example, a thermal conductivity detector (TCD), also known as a casarometer). The detector signal is processed by a controller to calculate various concentrations.

[0084] The locations for in-line measurement of the concentrations of phenol, cyclohexanol, and cyclohexanone in the exit mixture may be various locations downstream of the hydrogenation reactor. Measurements can be performed directly in the mixture exiting directly from the phenol hydrogenation reactor (line

[29] in Figure 2) or in lines exiting a heat exchanger or gas-liquid separator, if applicable (line

[30] or [5] in Figure 2). A particular convenient location for in-line measurement of the concentrations of phenol, cyclohexanol, and / or cyclohexanone in the exit mixture is in the liquid containing phenol, cyclohexanol, and cyclohexanone, either in or exiting the gas-liquid separator [H] (line [5] in Figure 2). A calibrated in-line gas chromatograph (GC) can be installed at any of these locations. The concentrations of phenol, cyclohexanol, and cyclohexanone are expressed as mol%, as used herein, and the sum of phenol, cyclohexanol, and cyclohexanone in the analyzed mixture is 100 mol%.

[0085] The set value for the phenol concentration used for comparison in step G) of the method of the present invention, and further the set value for the cyclohexanol concentration in step H) of the method of the present invention, must each be independently in the range of 0 to 20 mol%, preferably 0 to 12 mol%, more preferably 0 to 10 mol%, i.e., less than 10 mol%, with the sum of phenol, cyclohexanol, and cyclohexanone in the analyzed mixture always being 100 mol%, respectively. The reverse is true for the fifth set value for the cyclohexanone concentration, which can be calculated as [anone] = 100 mol% - [anol] - [phenol], since the sum of the concentrations of the three components is set to 100 mol%, representing the respective concentration set values ​​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 higher than 90 mol%, with the sum of phenol, cyclohexanol, and cyclohexanone in the analyzed mixture always being 100 mol%.

[0086] Generally, it is desirable to have as little cyclohexanol as possible in the exiting reaction mixture (high selectivity for cyclohexanone) and as little unconverted phenol as possible (high conversion rate). Therefore, the third and fourth settings are designed to require the process parameters in steps K) and L) of the present invention to be adapted only if the concentrations measured in steps C) and D) are higher than the corresponding settings, respectively. Similarly, since it is generally desirable to have as much cyclohexanone as possible in the exiting reaction mixture (high selectivity for cyclohexanone and high conversion rate of phenol), the fifth setting is designed to require the process parameters in step L2) of the present invention to be adapted only if the concentration measured in step D2) is lower than the corresponding setting. This is significantly different from the setting values ​​referenced in steps E), F), I), and J) of the method of the present invention, which require adaptation if the measured value is lower or higher than the setting value.

[0087] A further difference between the setpoints referenced in steps E), F), I), and J) of the method of the present invention and those referenced in steps G), H), K), L), and L2) is that the former require much more frequent attention using inline measurements and comparisons with setpoints, which are recommended to be at least once a day, once an hour, once a minute, or every 5 to 30 seconds. In contrast, the concentrations of phenol and cyclohexanol were found to require adjustment at a much lower frequency. The inventors found that in many cases it is sufficient to perform measurements, comparisons, and possibly adjustments of process parameters based on results in the range of once a day or once a week. This is also why an automated decision process is still desirable but not required with respect to steps K) and / or L) of the method of the present invention. According to preferred embodiments only, at least one process parameter that is adjusted in steps K), L), and / or L2) is adjusted by an automated driving mode.

[0088] In step K) of the method of the present invention, at least one process parameter is adjusted if the concentration of phenol in the exit mixture deviates from a third setpoint. In step L) of the method of the present invention, at least one process parameter is adjusted if the concentration of cyclohexanol in the exit mixture deviates from a fourth setpoint. In step L2) of the method of the present invention, at least one process parameter is adjusted if the concentration of cyclohexanone obtained in step D2) is lower than a fifth setpoint. Steps K), L), and / or L2) may be carried out as appropriate, 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 situations that occur several weeks to several months later, when the catalyst begins to age. Catalyst aging results in a (slight) decrease in conversion rate (resulting in an increase in the concentration of phenol in the exit mixture) and a decrease in selectivity (resulting in an increase in the concentration of cyclohexanol in the exit mixture). Typically, catalyst aging is a slow process, resulting in a period ranging from several weeks to several years between the start of operation with a new or regenerated catalyst bed and the replacement or regeneration of an old catalyst bed. Actual figures vary depending on numerous parameters, such as the quality of the new phenol, the quality of the new hydrogen, feedwater, the specifications of the desired product, the type of catalyst, and the plant layout. However, process disturbances (such as abnormally high concentrations of CO in the new hydrogen) can cause catalyst aging to be much faster than under normal conditions, potentially shortening the period between the onset of the disturbance and the replacement or regeneration of the old (poisoned) catalyst bed to just a few days or even less. In this case as well, steps K), L), and / or L2) can be beneficial in mitigating the adverse effects of catalyst aging.

[0089] If the concentration of cyclohexanone is also measured (see step D2 above), the method of the present invention may include a further step H2, which includes comparing the concentration of cyclohexanone obtained in step D2) with a fifth setpoint, and optionally a further step L2), which includes adjusting at least one process parameter if the concentration of cyclohexanone obtained in step D2) is lower than the fifth setpoint. The fifth setpoint, i.e., the desired cyclohexanone concentration in the exit mixture, is in any case 75 to 100 mol%, preferably 85 to 100 mol%, or 90 to 100 mol%, most preferably greater than 95 mol%, 98 mol%, or 99 mol%, with the sum of phenol, cyclohexanol, and cyclohexanone in the mixture to be analyzed being 100 mol%.

[0090] A deviation from a set value that is still considered acceptable and does not trigger compliance in step K), L), or L2) can be defined by a person skilled in the art in accordance with the present invention (for example, the range may be ±25%, ±10%, ±1%, or ±0.5% from a predefined set value within the range shown above).

[0091] Steps C), G), and K), and steps D), H), and L) are preferably performed 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 adapted by automated operation in steps K) and L) to return the concentrations of phenol and cyclohexanol to below set values ​​in a specific and effective manner include a) adjusting the temperature of the cooling medium, b) adjusting the hydrogen concentration in the elimination mixture, c) adjusting the concentration of the dilution gas in the elimination mixture, and d) a combination thereof.

[0093] It is well known that the reactor temperature affects the conversion rate of phenol and the selectivity for cyclohexanone in a phenol hydrogenation reactor. As described above with respect to heat exchangers, adjusting the temperature of the cooling medium (especially by the pressure of the boiling water if boiling water is used as the cooling medium) is a particular convenient method for adjusting the reactor temperature. Typically, the temperature of the cooling medium (and therefore the reactor temperature) should be increased to increase the conversion rate to cyclohexanone. However, care must be taken, as excessively high temperatures and conversion rates can adversely affect the selectivity for cyclohexanone.

[0094] As used herein, the term “adjusting the hydrogen concentration in the exit mixture” will most often mean “adjusting the hydrogen concentration in the input mixture,” particularly by increasing the amount of hydrogen charged into the process. Increasing the amount of hydrogen supplied to the reactor increases the hydrogen concentration and partial pressure within the reactor, thereby increasing the reaction rate, i.e., the conversion rate. However, care should be taken because if the hydrogen concentration and the associated conversion rate to cyclohexanol increase too much, the desirable selectivity for cyclohexanone will decrease. A particularly convenient method for adjusting the hydrogen concentration in the exit mixture is to adjust the second setpoint used for comparison in step F) and for fitting 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 process parameters.

[0095] The concentration of the diluent gas in the exit mixture used herein is Adjust In most cases, this means that the concentration of the diluent gas in the incoming mixture AdjustThis would mean that, for example, a decrease in the concentration of diluent gas in the exit mixture is achieved by decreasing the concentration of diluent gas in the entry mixture. A particular convenient way to achieve this is by adding less fresh hydrogen or more recycled gas mixture. As mentioned above, adjusting (especially decreasing) the concentration of diluent gas is a convenient way to adjust (especially increase) the hydrogen concentration, and thereby the reaction rate and associated conversion rate.

[0096] In a preferred embodiment, process parameters that can be adjusted by automatic operation in step K) to return the phenol concentration back to below a set value are selected from a) increasing the temperature of the cooling medium, b) increasing the hydrogen concentration in the effluent mixture, c) decreasing the concentration of the diluent gas in the effluent mixture, and d) a combination thereof.

[0097] In a preferred embodiment, process parameters that can be adapted by automatic operation in step L) to return the cyclohexanol concentration to below a set value are selected from a) lowering the temperature of the cooling medium, b) lowering the hydrogen concentration in the effluent mixture, c) increasing the concentration of the diluent gas in the effluent mixture, and d) a combination thereof.

[0098] If the method includes step L2) of adjusting at least one process parameter when the concentration of cyclohexanone obtained in step D2) is lower than the fifth setpoint, the process parameter to be adjusted is the same as that described above for the third setpoint relating to the concentration of phenol. This is because all the means described for reducing the phenol concentration in the exit mixture are means for increasing the conversion rate to cyclohexanone and the selectivity for cyclohexanone, respectively. Therefore, the means for adjusting the process parameter can also be used in step L2) to raise the concentration of cyclohexanone in the exit mixture above the fifth setpoint. For example, increasing the temperature of the cooling medium can be used in step L2) as shown in Example 4 below. Further methods for overcoming excessively high phenol concentrations and excessively low cyclohexanone concentrations are by reducing the phenol concentration in the entry mixture, i.e., supplying less phenol to the hydrogenation reactor. However, this would reduce the plant's output.

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

[0100] The plant of the present invention is at least — A hydrogenation reactor equipped with a palladium-containing catalyst, — At least one inlet into which an inlet mixture containing phenol and hydrogen can be charged into the reactor, - At least one discharge port capable of releasing an exit mixture containing cyclohexanone, cyclohexanol, phenol, and hydrogen, — A line for reusing at least a portion of the hydrogen in the exit mixture into the entry mixture. — A recovery unit capable of recovering cyclohexanone from the discarded mixture. — A purging section capable of releasing at least a portion of the hydrogen derived from the exit mixture as a purging 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. It is equipped with.

[0101] The plant of the present invention has the following parameters in particular: — Pressure inside the hydrogenation reactor; — Hydrogen concentration in the effluent mixture; — Phenol concentration in the effluent mixture; —Concentration of cyclohexanol in the elimination mixture; and — Depending on the circumstances, the concentration of cyclohexanone in the discarded mixture. A key feature is the existence of an inline measurement device capable of measuring [the specified value].

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

[0103] [Figure 1] A schematic diagram illustrating a plant and method for preparing and recovering cyclohexanone from phenol according to the present invention. [Figure 2] A schematic diagram showing an embodiment of the phenol hydrogenation reaction unit [A] according to the present invention. The phenol hydrogenation reaction unit [A] comprises a hydrogen purification unit [D], a phenol evaporation unit [E], a phenol hydrogenation unit [F], a heat exchange unit [G], a gas-liquid separation unit [H], and a compression unit [J]. [Modes for carrying out the invention]

[0104] The following reference numbers will be used in the description of the drawings: [A] Phenol hydrogenation reaction section [B] Separation and purification department [C] Cyclohexanol dehydrogenation reaction site [D] Hydrogen Purification Section [E] Phenolic evaporation section [F] Phenolic Hydrogenate [G] Heat exchange section [H] Gas-liquid separation section [J] Compression section [1] Hydrogen-containing flow [2] Unused phenol stream [3] Flow containing recovered phenol [4] Gaseous purge flow [5] A mixture containing phenol, cyclohexanone, cyclohexanol, and hydrogen. [6] Flow containing cyclohexanol and cyclohexanone [7] Flows containing a large amount of cyclohexanone [8] The first stream, which is rich in cyclohexanol. [9] The second stream, which is rich in cyclohexanol.

[10] Flow of light components

[11] Flow of heavy components

[12] hydrogen produced

[21] Hydrogen gas

[22] Reused hydrogen gas

[23] Hydrogen-containing stream

[24] Merged phenol flow

[25] Flow of gaseous components

[26] Components that did not evaporate

[27] Water

[28] Flow of supplies

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

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

[31] Gas mixture containing hydrogen

[32] Compressed gas mixture

[0105] [Detailed description of the drawing] A plant and corresponding method for preparing and recovering cyclohexanone from phenol are schematically shown in Figure 1. Such a method typically consists of two parts, and sometimes a third part. All three parts are described.

[0106] Cyclohexanone is prepared by a catalytic reaction in the phenol hydrogenation section [A] by reacting phenol with hydrogen. The cyclohexanone is recovered in the separation and purification section [B]. In the cyclohexanol dehydrogenation section [C], which may be present, the undesirable byproduct cyclohexanol is converted to cyclohexanone and hydrogen by a catalytic reaction.

[0107] The phenol hydrogenation reactor [A] comprises at least one hydrogenation reactor into which a hydrogen-containing stream is charged via line [1], a stream of unused phenol via line [2], and optionally a stream of recovered phenol via line [3], and may also be equipped with additional equipment. The hydrogen-containing stream may optionally contain an inert component (e.g., nitrogen and / or methane) and optionally carbon monoxide as dilution gases. Optionally, the inert component (e.g., nitrogen and / or methane) and optionally carbon monoxide are charged separately into the phenol hydrogenation reactor [A] (not shown in Figure 1). The phenol hydrogenation of the catalytic reaction according to the present invention is carried out in a vapor phase process. From the phenol hydrogenation reactor [A], a gaseous purge stream containing hydrogen and optionally an inert component (e.g., nitrogen and / or methane) and optionally carbon monoxide is released via line [4], and a liquid mixture containing phenol, cyclohexanone, cyclohexanol and hydrogen is released via line [5]. The hydrogenation reaction section [A] is a method for obtaining the purge flow [4] and the product flow [5], and will be explained in more detail below in relation to Figure 2.

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

[10] . In some cases, a stream of heavy components containing phenol and components with a higher boiling point than phenol is released via line

[11] .

[0109] The cyclohexanol dehydrogenation reaction section [C], which may be present, comprises at least one dehydrogenation reactor and one or more heat exchangers. In the cyclohexanol dehydrogenation reaction section [C], cyclohexanol is converted to cyclohexanone and hydrogen by a catalytic reaction. Generally, the dehydrogenation of cyclohexanol is a gas-phase reaction carried out at high temperature. Depending on the circumstances, the product stream containing cyclohexanol and cyclohexanone is charged to the separation and purification section [B] via line [6]. The hydrogen produced in the cyclohexanol dehydrogenation reaction section [C] is released as generated hydrogen via line

[12] . Depending on the circumstances, the generated hydrogen from the cyclohexanol dehydrogenation reaction section [C] is supplied to the phenol hydrogenation reaction section [A] (not shown in Figure 1). Depending on the circumstances, the generated hydrogen from the cyclohexanol dehydrogenation reaction section [C] is supplied to another hydrogen-consuming process (not shown in Figure 1). In some cases, the hydrogen produced in the cyclohexanol dehydrogenation reaction section [C] is supplied to the exothermic unit (not shown in Figure 1).

[0110] Figure 2 shows a scheme of the phenol hydrogenation reaction section [A] (area enclosed by dashed lines) according to the present invention.

[0111] A hydrogen-containing stream is charged into the phenol hydrogenation reaction unit [A] via line [1]. The hydrogen-containing stream may originate from, for example, 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. In some cases, inert components such as nitrogen and / or methane may be added to the hydrogen-containing stream as a diluent gas (not shown in Figure 2). If the hydrogen-containing stream contains components harmful to hydrogenation, such as CO and / or H2S, a hydrogen gas purification step is required. The presence of these harmful components in the charged hydrogen may be temporary or permanent, for example, due to a disruption in the hydrogen gas generation unit. In such a hydrogen purification step, the harmful impurities may be converted into inert components or removed from the hydrogen-containing stream.

[0112] For this purpose, the hydrogen-containing flow is charged via line [1] into an optionally present hydrogen purification unit [D]. Before being charged into the hydrogen purification unit [D], the temperature of the hydrogen-containing flow can be changed in a heat exchanger to the temperature required by the hydrogen purification unit [D] (not shown in Figure 2). Generally, the temperature of the hydrogen-containing flow is increased in such a heat exchanger.

[0113] The hydrogen purification section [D] may comprise one or more catalysts for converting harmful components into 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 released from the hydrogen purification section [D] via line

[21] . The hydrogen gas in line

[21] and the recycled hydrogen gas in line

[22] merge to form a hydrogen-containing flow that is charged into the phenol evaporation section [E] via line

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

[0114] A stream of unused phenol may be charged via line [2], but preferably a stream containing recovered phenol is charged via line [3] to form a combined phenol stream that flows through line

[24] . The stream containing recovered phenol charged via line [3] is discharged from the separation and purification section [B] shown in Figure 1. Optionally, the stream containing recovered phenol may be charged separately into the phenol evaporator [E] without being combined with the stream of unused phenol charged via line [2], or it may not exist (not shown in Figure 2). Optionally, before being charged into the phenol evaporator [E], the temperature of the combined phenol stream flowing through line

[24] is increased in a heat exchange section (not shown in Figure 2) comprising one or more heat exchangers capable of parallel and / or continuous operation.

[0115] In the phenol evaporator [E], virtually all components entering through line

[24] are evaporated. The gaseous component flow is discharged from the phenol evaporator [E] through line

[25] . Components entering through line

[24] and not evaporating are discharged from the phenol evaporator [E] through line

[26] as non-evaporable components (continuously or batch by batch). These are usually in very small amounts. Generally, the phenol evaporator [E] requires heating of the incoming feed, particularly with steam. Preferably, the phenol evaporator [E] is equipped with a device for removing droplets mixed in with the discharged gaseous component flow, such as a wire mesh demister. The phenol evaporator [E] comprises one or more evaporators operating in parallel and / or continuously. Optionally, the gaseous component flow in line

[25] is temperature-controlled in a heat exchanger (not shown in Figure 2) comprising one or more heat exchangers operating in parallel and / or continuously. Generally, in this heat exchange section, the temperature of the gaseous component flow is increased. If necessary, a small amount of water, for example in the form of steam, is introduced into the flow of line

[25] via line

[27] to form a feed flow that flows through line

[28] and is introduced into the phenol hydrogenation section [F].

[0116] The phenol hydrogenation section [F] consists of one or more hydrogenation reactors operating continuously and / or in parallel. In the phenol hydrogenation section [F], cyclohexanone and cyclohexanol are obtained in a continuous process by hydrogenation of phenol via a heterogeneous catalytic reaction. A gas mixture containing cyclohexanone, cyclohexanol, phenol, and hydrogen is released from the phenol hydrogenation section [F] via line

[29] .

[0117] A gas mixture containing cyclohexanone, cyclohexanol, phenol, and hydrogen, released from the phenol hydrogenation section [F] via line

[29] , is cooled in a heat exchange section [G], thereby condensing at least one fraction of the phenol, cyclohexanone, and cyclohexanol. The heat exchange section [G] comprises one or more heat exchangers operated in parallel and / or continuously. One or more cooling media may be used to cool the gas mixture containing cyclohexanone, cyclohexanol, phenol, and hydrogen, released from the phenol hydrogenation section [F] via line

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

[29] . In some cases, the gas cooling is carried out in several stages in succession, and therefore, in some cases, a different cooling medium is used at each stage. In some cases, at least one of the cooling mediums is a process flow from the phenol hydrogenation reaction section [A], from the separation and purification section [B], and / or from the cyclohexanol dehydrogenation reaction section [C] (not shown in Figure 2). In some cases, one of the cooling mediums is recycled hydrogen gas charged via line

[22] (not shown in Figure 2).

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

[30] and charged into the gas-liquid separation section [H]. The gas-liquid separation section [H] comprises one or more gas-liquid separators operated in parallel and / or continuously. The liquid mixture containing phenol, cyclohexanone, cyclohexanol, and hydrogen is discharged from the gas-liquid separation section [H] via line [5] and charged into 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 into the compression section [J]. Any type of gas-liquid separator can be used in the gas-liquid separation section [H], but particularly good results are obtained when a vertical container is used in which the liquid settles at the bottom of the container. Generally, the gas-liquid separation section [H] comprises a device for removing droplets mixed in with the discharged gaseous component, such as a wire mesh demister.

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

[32] and separated into recycled hydrogen gas, which is transported through line

[22] , and a gaseous purge flow, which is transported through line [4]. Optionally, the recycled hydrogen gas transported through line

[22] is charged into a heat exchanger, where it is heated (not shown in Figure 2). The recycled hydrogen gas (which is optionally heated) is released through line

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

[21] .

[0120] When the method of the present invention is carried out, the gaseous purge flow transported through line [4] typically contains hydrogen and one or more inert components, such as nitrogen and / or methane, and optionally carbon monoxide. This gaseous purge flow released through line [4] can be used, for example, as fuel (not shown in Figure 2). Optionally, before release from the phenol hydrogenation reaction section [A], the gaseous purge flow transported through line [4] is charged into a heat exchange section where it is cooled (not shown in Figure 2). Optionally, the liquid formed in this heat exchange section is charged into a gas-liquid separation section [H] (not shown in Figure 2). [Examples]

[0121] The following examples serve to illustrate the present invention in more detail, particularly with respect to certain specific embodiments of the invention. However, these examples are not intended to limit the present 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 comprises a separation and purification section [B] from which high-quality cyclohexanone is obtained by distillation separation from higher boiling point components and lower boiling point components, and a cyclohexanol dehydrogenation reaction section [C] from which cyclohexanol is converted to cyclohexanone. The phenol hydrogenation reaction section [A] comprises a hydrogen purification section [D], a phenol evaporation section [E], a phenol hydrogenation section [F], a heat exchange section [G], a gas-liquid separation section [H], and a compression section [J], which are very similar to the embodiment of the present invention shown in Figure 2.

[0123] A hydrogen-containing stream containing approximately 94% by volume of hydrogen, approximately 6% by volume of nitrogen, trace amounts of CO, and trace amounts of H2S was charged to the hydrogen purification section [D] via line [1]. In the hydrogen purification section [D], CO was converted to CH4 by a catalytic reaction, and H2S was removed by an adsorbent. Under normal operating conditions, the CO and H2S content in the hydrogen gas released from the hydrogen purification section [D] via line

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

[22] with the newly supplied hydrogen gas transported via line

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

[0124] A stream of unused phenol, charged via line [2], and a stream containing recovered phenol, charged via line [3], were combined and charged into the steam-heated phenol evaporator [E]. In the phenol evaporator [E], virtually all of the phenol was evaporated. A small flow of non-evaporating components was released from the phenol evaporator [E] via line

[26] , and the remainder was released as a stream of gaseous components via line

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

[25] via line

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

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

[28] ) was obtained.

[0125] The phenol hydrogenation section [F] was equipped with one industrial-scale multi-tube phenol hydrogenation reactor with an annual operating capacity of approximately 120 kta of phenol, operated in continuous mode for phenol hydrogenation in the vapor phase. The tubes were filled with a supported hydrogenation catalyst Pd / Al2O3 (0.9 wt%) and 1 wt% Na (as NaHCO3) added as a co-catalyst. There were approximately 5600 tubes. The height of the catalyst bed in each reactor tube was approximately 2 m. The mass feed rate per hour (WHSV) was 5 (kg phenol / hr) / (kg catalyst). Steam was produced by charging boiler water into the space outside the tubes to remove the heat of reaction as a cooling medium. More than 85% of the heat of reaction was transferred to the cooling medium.

[0126] [Example 1] The multi-tube hydrogenation reactor was filled with unused catalyst. Starting 72 hours after the start of process operation, the following process conditions were recorded for a 7-month continuous production period.

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

[0128] The pressure inside the reactor was measured by in-line measurement within the reactor hood of the feed flow being charged into the multi-tube hydrogenation reactor, before the feed flow entered the catalyst-filled tubes. A pressure sensor continuously emitted an electrical signal, which was a measure of the magnitude of the pressure inside the reactor. The pressure setpoint (first setpoint) was 0.38 MPa. If the pressure deviated from this setpoint, the outflow of the gaseous purge flow in line [4] was automatically adjusted to maintain the first setpoint (if the pressure was too high, the valve on line [4] was opened; if the pressure was too low, the valve on line [4] was closed).

[0129] The hydrogen concentration in the exit mixture was continuously measured by in-line measurement of the gaseous purge flow in line [4] using a calibrated in-line thermal conductivity detector (casarometer). The thermal conductivity detector continuously emitted an electrical signal, which was a measure of the hydrogen concentration. The hydrogen concentration setpoint (second settingpoint) was 4 mol%, but the sum of the concentrations of all gaseous components (primarily hydrogen and inert gases, such as methane and / or nitrogen) was always 100 mol%. If the hydrogen concentration deviated from this settingpoint, the hydrogen gas inflow into line [1] was automatically adjusted to maintain the second settingpoint (if the hydrogen concentration was too low, the hydrogen gas inflow into line [1] was increased; if the hydrogen concentration was too high, the hydrogen gas inflow into line [1] was reduced).

[0130] The concentrations of phenol, cyclohexanol, and cyclohexanone in the discharge mixture were measured by in-line measurement using a calibrated in-line gas chromatograph (GC) of the liquid at the bottom of the liquid reservoir of the vertical container type gas-liquid separator in gas-liquid separation section [H]. The set values ​​(3rd to 5th set values) for the concentrations of phenol, cyclohexanol, and cyclohexanone in the liquid at the bottom of the liquid reservoir of the vertical container type gas-liquid separator in gas-liquid separation section [H] were 3 mol%, 3 mol%, and 95 mol%, respectively (the sum of the concentrations of phenol, cyclohexanol, and cyclohexanone was always set to 100 mol%). During the first 7 months after the start of operation, 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 set to 100 mol%).

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

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

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

[12] , and the stream containing cyclohexanol and cyclohexanone was released through line [6] and charged into the separation and purification section [B].

[0132] The hydrogen-containing gas mixture was released from the top of the vertical container of the gas-liquid separation section [H] via line

[31] and charged into the compressor of the compression section [J]. The hydrogen-containing compressed gas mixture released from the compressor of the compression section [J] was separated into a gaseous purge flow transported via line [4] and a recycled hydrogen gas transported via line

[22] .

[0133] The results of Example 1 demonstrate that, by combining in-line measurement and automatic control of the pressure inside the reactor and the hydrogen concentration in the discharged mixture according to the present invention, it was possible to operate the phenol hydrogenation reactor at a constant high production rate for approximately 7 months without human intervention.

[0134] [Example 2] The preparation and recovery of cyclohexanone from phenol, carried out at the same chemical plant as described in Example 1, was continued after the initial 7-month period.

[0135] After several days, the phenol concentration in the liquid at the bottom of the reservoir of the vertical container gas-liquid separator in section [H], as measured by a calibrated inline gas chromatograph, exceeded the set value of 3 mol% (the sum of the concentrations of phenol, cyclohexanol, and cyclohexanone was always set to 100 mol%). In response to this, the hydrogen concentration set value 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 reservoir of the vertical container gas-liquid separator in section [H], as measured by a calibrated inline gas chromatograph (GC), decreased to less than 3 mol%, while the concentrations of cyclohexanol and cyclohexanone remained at <3 mol% and >95 mol%, respectively (the sum of the concentrations of phenol, cyclohexanol, and cyclohexanone was always set to 100 mol%). This new situation could be maintained for more than one month without further adjustment of the hydrogen concentration set value. Within the first 12 months of operation, the hydrogen concentration setting had to be increased by 1 mol% several times (until the hydrogen concentration setting reached 10 mol%) in order to adjust the phenol concentration in the liquid at the bottom of the liquid reservoir of the vertical container-type gas-liquid separator in the gas-liquid separation section [H] to less than 3 mol%.

[0136] When the hydrogen concentration setting was increased from 4 mol% to 10 mol%, the outflow of gaseous purge flow from line [4] increased over time.

[0137] The results of Example 2 demonstrate that adjusting the hydrogen concentration in the elimination mixture is used to control the phenol concentration in the elimination mixture. Each time the phenol concentration in the elimination mixture exceeded its set value of 3 mol%, the set value of the hydrogen concentration in the elimination mixture was increased stepwise by 1 mol% (from 4 mol% to 10 mol%), resulting in a stepwise increase in the hydrogen concentration in the elimination mixture from 4 mol% to 10 mol%. After each increase in the hydrogen concentration in the elimination mixture, the phenol concentration in the elimination mixture was again reduced to less than 3 mol%. This control strategy allowed for a consistently high production rate in the phenol hydrogenation reaction section [A] of the chemical plant to be extended up to 12 months after commissioning.

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

[0139] Several days later, the cyclohexanone concentration in the liquid at the bottom of the reservoir of the vertical container-type gas-liquid separator in the gas-liquid separation section [H], as measured by a calibrated inline gas chromatograph (GC), fell below the set value of 95 mol% (the sum of the concentrations of phenol, cyclohexanol, and cyclohexanone was always assumed to be 100 mol%). In response, 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 rose by approximately 1.5°C to approximately 152°C.

[0140] As a result of this effect, the cyclohexanone concentration in the liquid at the bottom of the reservoir of the vertical container-type gas-liquid separator in the gas-liquid separation section [H], as measured by a calibrated inline gas chromatograph, again exceeded 95 mol%, while the concentrations of phenol and cyclohexanol remained <3 mol% (the sum of the concentrations of phenol, cyclohexanol, and cyclohexanone was always assumed to be 100 mol%). This condition 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 demonstrate that by once adjusting the pressure of the boiling water used as a cooling medium when the cyclohexanone concentration in the discarded mixture was lower than the set value for the cyclohexanone concentration in the discarded mixture, it was possible to maintain a consistently high production rate in the phenol hydrogenation reaction section [A] of the chemical plant for an additional 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 the initial 14 months. Due to market conditions, the production rate of the phenol hydrogenation reaction section [A] had to be reduced by 30%.

[0143] Immediately after reducing the phenol supply rate to the phenol hydrogenation reaction section [A] without adjusting the setpoint, the cyclohexanol concentration in the liquid at the bottom of the liquid reservoir of the vertical container type gas-liquid separator in the gas-liquid separation section [H], as measured by a calibrated in-line gas chromatograph (GC), exceeded the setpoint of 3 mol% (the sum of the concentrations of phenol, cyclohexanol, and cyclohexanone was always assumed to be 100 mol%). In response to this, the hydrogen concentration setpoint (second setpoint) was reduced by 2 mol% (from 10 mol% to 8 mol%), and the system reacted to this setpoint by reducing the hydrogen gas inflow into line [1] as described in Example 1 above. As a result of this action, the hydrogen concentrations in the incoming and outgoing mixtures decreased, causing the cyclohexanol concentration in the liquid at the bottom of the reservoir of the vertical container-type gas-liquid separator in the gas-liquid separation section [H] to again fall below 3 mol%, as measured by a calibrated inline gas chromatograph, while the concentrations of phenol and cyclohexanone remained at <3 mol% and >95 mol%, respectively (the sum of the concentrations of phenol, cyclohexanol, and cyclohexanone was always set to 100 mol%). This situation could be maintained for more than one month without further adjustment of any of the setpoints.

[0144] The results of Example 4 show that by lowering the set value of the hydrogen concentration in the exit mixture, and as a result the hydrogen concentrations in both the input and exit mixtures decreased, it was possible to reduce the concentration of cyclohexanol in the exit mixture to a value lower than the set value 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 flow in line [4] was not automatically adjusted when the pressure deviated from the set value, and the inflow of hydrogen gas in line [1] was not automatically adjusted when the hydrogen concentration deviated from the set value.

[0146] Instead, if the pressure deviated from the set value, the hydrogen gas inflow into line [1] was automatically adjusted (if the pressure was too low, the hydrogen gas inflow into line [1] was increased; if the pressure was too high, the hydrogen gas inflow into line [1] was reduced). Also, if the hydrogen concentration deviated from the set value, the gaseous purge flow outflow into line [4] was automatically adjusted (if the hydrogen concentration was too high, the gaseous purge flow outflow into line [4] was reduced; if the hydrogen concentration was too low, the gaseous purge flow outflow into line [4] was increased).

[0147] In this case as well, for 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 container type gas-liquid separator in the gas-liquid separation section [H] remained at <3 mol%, <3 mol%, and >95 mol%, respectively (the sum of the concentrations of phenol, cyclohexanol, and cyclohexanone was always assumed to be 100 mol%).

[0148] The results of Example 5 demonstrate that by combining in-line measurement and automatic control of the pressure inside the reactor and the hydrogen concentration in the discharged mixture according to the present invention, it was possible to operate the phenol hydrogenation reactor at a constant high production rate for approximately 7 months without human intervention.

Claims

1. A method for controlling the exothermic vapor phase hydrogenation of phenol catalyzed by a palladium-containing catalyst in an industrial-scale hydrogenation reactor, An input mixture containing phenol and hydrogen is charged into the reactor, and an exit mixture containing cyclohexanone, cyclohexanol, phenol and hydrogen is discharged from the reactor, At least a portion of the hydrogen in the exit mixture is reused in the entry 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, At least a portion of the heat of reaction in the phenol hydrogenation reaction is transferred to the cooling medium by indirect heat exchange. The method is, A) A step of measuring the pressure inside the reactor using in-line measurement; B) A step of measuring the hydrogen concentration in the discharged mixture using in-line measurement; C) A step of measuring the phenol concentration in the discarded mixture by in-line measurement; D) A step of measuring the cyclohexanol concentration in the elimination mixture by in-line measurement; D2) Depending on the circumstances, measure the cyclohexanone concentration in the discarded mixture using in-line measurement; E) A step of comparing the pressure inside the reactor in step A) with a first set value; F) A step of comparing the hydrogen concentration obtained in step B) with a second set value; G) A step of comparing the phenol concentration obtained in step C) with a third set value; H) A step of comparing the cyclohexanol concentration obtained in step D) with a fourth set value; H2) Depending on the case, compare the cyclohexanone concentration obtained in step D2) with a fifth set value; I) If the pressure inside the reactor obtained in step A) deviates from the first set value, the automatic operation mode adjusts at least one process parameter; J) A method comprising the step of adjusting at least one process parameter by an automatic operation mode if the hydrogen concentration in the discarded mixture obtained in step B) deviates from a second set value.

2. At least one process parameter that is adapted in step I) and / or J) is a) Flow rate of the incoming mixture, and / or b) Flow rate of the purge flow The method according to claim 1, including the method described in claim 1.

3. The method according to any one of claims 1 or 2, wherein the incoming mixture further comprises a dilution gas.

4. The method according to claim 3, wherein, in addition to the hydrogen in the exit mixture that is reused in the incoming mixture, at least a portion of the diluent gas present in the exit mixture is also reused in the incoming mixture.

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

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

7. The method according to any one of claims 5 and 6, wherein at least one process parameter that is adapted in step K), L), and / or L2) is adapted by the automatic driving mode.

8. At least one process parameter that is adapted in steps K, L) and / or L2) is, a) Temperature of the cooling medium, b) Hydrogen concentration in the leached mixture, c) The concentration of the diluent gas in the elimination mixture, and d) These combinations The method according to any one of claims 5 to 7, selected from the group consisting of the following.

9. At least one process parameter that is adapted in step K is: a) Rise in the temperature of the cooling medium, b) Increase in hydrogen concentration in the leached mixture, c) A decrease in the concentration of dilution gas in the discarded mixture, and d) These combinations The method according to claim 8, selected from the group consisting of the following.

10. At least one process parameter that is adapted in step L and / or L2 is: a) Decrease in the temperature of the cooling medium, b) Decrease in hydrogen concentration in the discharged mixture, c) Increase in the concentration of diluent gas in the detached mixture, and d) These combinations The method according to claim 8, selected from the group consisting of the following.

11. The method according to any one of claims 1 to 10, 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 boiling water used to absorb the heat of gas-phase hydrogenation which is exothermic, and the boiling point of 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 to within the range of 0.15 MPa to 1.5 MPa.

13. The method according to any one of claims 8 to 12, wherein the hydrogen concentration in the discharged mixture is increased by increasing the amount of hydrogen charged into the method.

14. The method according to any one of claims 1 to 13, wherein the industrial-scale hydrogenation reactor is a multi-tube heat exchange reactor that uses water as a cooling medium, and steam is generated by the reactor as a result of indirect heat exchange.

15. An industrial-scale phenol hydrogenation plant configured to carry out the method described in any one of claims 1 to 14, wherein the plant is — A hydrogenation reactor equipped with a palladium-containing catalyst, — At least one inlet into which an input mixture containing phenol and hydrogen can be charged into the reactor, - At least one discharge port capable of releasing an exit mixture containing cyclohexanone, cyclohexanol, phenol, and hydrogen, — A line for reusing at least a portion of the hydrogen in the exit mixture into the entry mixture. — A recovery unit capable of recovering cyclohexanone from the discarded mixture. — A purging section capable of releasing at least a portion of the hydrogen derived from the exit mixture as a purging 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. It is equipped with, The plant has the following parameters, namely — Pressure inside the hydrogenation reactor; — Hydrogen concentration in the effluent mixture; — Phenol concentration in the effluent mixture; — Cyclohexanol concentration in the elimination mixture A plant characterized by being equipped with an inline measuring device capable of measuring all of the following.

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