Distillation Handbook BP Oil.pdf

BP Oil Distillation Handbook Rev: 0 Volume 1 General Design & Column Internals Date: Feb-97 Section A Typical Distillation Process Flowschemes Page: 1A.1

Contents page

Section yet to be written

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This information is of a confidential or proprietary nature. For internal BP use only, not to be transmitted to others.

BP Oil Distillation Handbook Rev: 2 Volume 1 General Design and Column Internals Date: Jul-01 Section B Simulation & Distillation Environment Page: 1B.1

Contents page

Physical properties and thermodynamics 1B.2 Thermodynamics Recommended K-values Properties Crude oil properties Water handling Azeotropes Steam stripper followed by a vacuum drier Computational Fluid Dynamics 1B.9 Description Capabilities and application of CFD modelling Benefits of CFD simulation and caveats relating to its use Theory of CFD modelling Examples of the application of CFD within BP Oil Flow distribution through mist eliminator pad Vacuum distillation column flash zone modelling Liquid temperature and residence time calculation for a chimney tray Heater flue design

Distillation Environment 1B.17 Scope Availablility Confidentially Programs available Description of programs Simulation 1B.23 Background Simulation of vacuum distillation units Optimisation opportunities for VDU's Simulation of gas plants with absorber strippers Gas plant optimisation opportunities Setting up a gas plant simulation Generation of a pseudo McCabe-Thiele plot

Revision History Rev 1 Original Issue Rev 2 Section amalgamated from following RH 07/01 Section ba (05/98), bb (06/95), bc (06/98), bd (05/98)

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Physical properties and thermodynamics Thermodynamics Simulation models of distillation processes commonly employ the equilibrium stage method for which accurate volatility predictions are vital in order to determine separation efficiency. Typically, a wide menu of options for computation of K-values is available to suit the diverse applications to which simulators are put, but only a sub-set is of relevance in most refinery applications which, in the main, deal with hydrocarbon systems at modest pressure. Recommendations on which options to use for a particular circumstance are given in this section of the handbook. There is rarely a single best option which explains in part why there are so many alternatives available. If further guidance is needed, please seek assistance from a specialist at the Oil Technology Network. Most separations of interest will be hydrocarbon, as already stated, but also included in the tabulations are recommendations for other separations including HF alkylation unit columns, sour water strippers and amine units. Even some hydrocarbon separations require special thermodynamics e.g. propane/propene splitters and reformate splitters to properly account for relative volatilities. Frequently, simulation vendors have tuned the standard equation of state K-value methods not least by providing large data banks of interaction parameters which improve the predictions for fluid mixtures such as hydrogen/hydrocarbons or light gases/hydrocarbons. Hence, an apparently identical choice of method between packages need not necessarily give the same answer.

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Recommended K-values

(Alternatives are identified by italics) Application GENESIS HYSIM PRO/II CDU atms. column ZJ GS

PR ZJ Esso_tab GS PR SRK

VDU vac. column ZJ Esso_tab BK10 ZJ

BK10 GS

Steam stripper & Vacuum drier (see text below)

- Steam stripper: PR Vacuum drier: K-D

-

FCCU main frac. and gas plant

ZJ PR GS

Visbreaking main frac

ZJ PR GS

Hydrocracker main frac

ZJ PR GS

CRU columns ZJ PR SRK LPG columns ZJ

PR SRK PR GS

DHT reactor (hydrogen systems)

Soave PR

PR SRK

SRK GS

Alkylation (HF systems)

Uniquac PRSV+NRTL+Tabular Proprietary

Sour water stripping Not available PR_Sour Proprietary Amine columns (vapour/liquid)

Not available AMSIM

CO2 absorption (Benfield-type unit)*

Not available Not available

Etherification (MTBE/TAME)

NRTL NRTL NRTL

Light hydrocarbons (near critical)

Soave PR

PR SRK

BWRS SRK PR

Reformate splitters** NRTL PR PRSV NRTL UNIQ

SRKM

Reformate stabilisers ** PR Glycol systems NRTL PR Glycol Furfural systems NRTL NRTL

UNIQUAC PRSV SRKM

* For aqeuous electrolyte systems the OLI/FraChem program is the best available technology (contact GRE) ** Where distribution of aromatics is important, appropriate interaction parameters are required for either an 'activity' K-method or an EOS method. Recent (1996) experience in OTN has shown the HYSIM PR method with fitted parameters is satisfactory.

Key: PR Peng Robinson SRK Soave Redlich Kwong ZJ Zudkevitch Joffe GS Grayson Stread BK10 Braun K10 Esso_tabular Esso adaptation of API NRTL Non Randon Two Liquid (Renon) BWRS Benedict Webb Rubin Starling Glycol Proprietary SimSci interaction coefficients SRKM Soave Redlich Kwong Modified (Panagiotopoulos and Reid) PRSV Peng Robinson Stryjek Vera K-D Kabadi-Danner

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Properties Equally important to the process engineer is a knowledge of physical and transport properties and of petroleum qualities of hydrocarbon products. This requirement is satisfied by the process simulation packages in one of two ways: either by correlation or via a data bank. Where a link exists between crude oil assay information e.g. the BP Crude Oil Assays and the particular simulator there can be extra confidence in the calculated petroleum properties. Links do exist with HYSIM, GENESIS and PRO/II. The link to HYSIM is made using BP’s CLIP (Crude Library Interface Program) program. CLIP caters for individual crudes or blends or fractions and will produce a comprehensive file of basic and petroleum properties split up into the type of hypothetical components expected by simulation software. Otherwise, simulation packages employ general correlations or, in the case of many petroleum properties, none. General correlations for certain basic properties are perfectly adequate e.g. density, but for transport properties esp. viscosity, and for most petroleum properties, these are unreliable. API methods exist for a selection of petroleum properties such as pour point, flash point and refractive index, and it is these that are programmed into the commercial simulation packages. One reason why properties such as viscosity and cloud point are difficult to predict for oils is that the constituent component contributions to the overall property do not blend linearly. Standard features nowadays in most process simulators include options to calculate tray or packing hydraulics. In such cases the required physical and transport are automatically available. The same properties, of course, may be displayed and used in external software for column internals design, if required.

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Crude oil properties BP assays are available for 400+ internationally traded crudes and are updated regularly to reflect the changing characteristics of produced oil with time. These assays, in electronic format, comprise sets of 'property' vs. temperature profiles for upwards of 40 properties, listed below. A full description of these packs, if required, may be found in the PPEFA4 User Guide.

Petroleum Distillate Properties from BP Crude Oil Assays

Specific gravity T36:1 temperature * Cetane index * API Gravity * RON Wax content Molecular weight MON Total nitrogen Total sulphur Lead appreciation Vanadium content Mercaptan sulphur Flash point Nickel content Paraffin content Pour point Acidity Naphthene content Cloud point Asphaltenes Aromatic content Freezing point Soft asphaltenes Carbon/hydrogen ratio Aniline point Conradson carbon Watson K * Diesel index * Smoke point * Reid vapour pressure Kinematic viscosity

* Derived from other properties in the list

The following properties, in particular, are routinely used as quality indicators:

Property Range PONA Naphtha fractions to 205C (400F) Octanes Light naphtha fractions RVP Light naphtha fractions Naphthalenes Naphthas 205 - 315C (400F - 600F) Pour point Distillates and Gas oils Cloud point Middle distillates Freeze point Distillates Smoke point Middle distillates Nitrogen Heavy distillates through Gas oils Viscosity Heavy distillates through Bottoms Conradson carbon Gas oils through to Bottoms Metals Gas oils through to Bottoms Aniline point Middle distillates Naphthenic acid All fractions and Residue Mercaptans Naphthas and Gas oils Compositional analysis Naphthas to 150C (300F)

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Water handling A complicating factor in simulation hydrocarbon systems arises if water is present. All simulators provide for water decant, at least from the overhead condenser, if not also directly from trays. The traditional approach in simulation has been uniform: the favoured procedure is to assume that the decant water is pure water and the solubility of water in hydrocarbon is calculated based on the API Data Book Figure 9A1.4 for kerosene solubility. Exceptions to this rule, which provide rigorous treatment of water are standard PR in HYSIM and a variant on ZJ in GENESIS. Additionally, the HYSIM PR_Sour method has the ability to treat acid gas systems rigorously: a temperature dependent interaction parameter is used to match the solubility of the acid component in the water phase. On the topic of hydrocarbon-in-water solubility (more difficult to predict than water-in-hydrocarbon) there have been advances in recent years and the commercial simulators now provide a new method (Kabadi-Danner) which may be specified in conjunction with the regular equation of state methods such as SRK.

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Azeotropes Although the possibility of azeotropes should be checked with OTN, azeotropes formed in common systems are discussed below. Benzene / nohexane. May occur when fractionating cat reformate to remove benzene. Figure BA-1 gives a guide to when the azeotrope might form. Consult OTN for confirmation in all significant design work involving this azeotrope.

0 0.02 0.04 0.06 0.08 0.1 0.12 0.140

10

20

30

40

50

60

70

80

Mol fract of benzene in benzene/n-hexane binary

Azeo

trope

tem

pera

ture

d

egC

0 0.02 0.04 0.06 0.08 0.1 0.12 0.140

10

20

30

40

50

60

70

80

Mol fract of benzene in benzene/n-hexane binary

Azeo

trope

tem

pera

ture

d

egC

Figure BA-1: Azeotropic xomposition for benzene / n-hexane The azeotropic temperature may be estimated by the following equation. Temp = 76.651 - 484.58 * mol fract.

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Steam stripper followed by a vacuum dryer Work on a study of diesel steam strippers and their associated vacuum driers by Hydroprocessing Group, using HYSIM, considered the Peng-Robinson and the Kabadi-Danner packages. The study concluded that the steam stripper was best simulated using P-R because of its good handling of hydrocarbon systems, and the K-D package for the subsequent vacuum drier. The use of these packages not based on actual plant comparison, but rather a theoretical HYSIM based study. The K-D method's claim to fame is that it is one of the very few methods available anywhere to deal with the dilute region for water/hydrocarbons, particularly the solubility of hydrocarbon in water - something of interest to environmental engineers. K-D is a modification of SRK. Our more usual concern, from the perspective of mainstream processing, is the solubility of water in hydrocarbon i.e at the reflux drum, water-on-trays etc. HYSIM PR is fine for that situation.

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Computational Fluid Dynamics Description Computational Fluid Dynamics (CFD) is an important analysis tool with application to many areas within chemical, aeronautical and mechanical engineering as well as many other scientific fields. BP has much experience of, and has gained much benefit from, the application of CFD simulation techniques having used them for the analysis of the internal and external flow patterns associated with process plant for more than 10 years. CFD may be defined as the numerical simulation of the motion or behaviour of a fluid that is subject to influences and constraints. As such its scope extends from the everyday application of simple hydraulic correlations to the computer simulation of the detailed fluid dynamics of a complex system. The information in this section of the design manual will be restricted to the application of CFD as applied to distillation processes. A considerable number of BP Oil distillation columns and other units have been simulated during the past 5 years and many successful modifications made on the basis of CFD models. This has given enhanced understanding of the conditions within such units and their operation as well as an opportunity to assess design modifications. In this way CFD has proved to be a valuable analysis tool assisting in both unit design work and troubleshooting. Although in general it is straight forward to set up CFD models of simple systems, most refinery units are complex and require careful detailed simulation by experienced personnel. It is considered unlikely that an individual refinery will find it cost effective to perform CFD analyses in-house for the next few years. CFD modelling expertise exists within the TDU Separations, Energy and Treatments team (SET) generally using the general purpose software PHOENICS. BP also has access to FLUENT and monitors the development and capabilities of other CFD packages. This expertise may be called upon as required for process design, analysis and troubleshooting work. Capabilities and application of CFD modelling General purpose CFD simulation packages allow the modelling of many of the phenomena involved in distillation units to acceptable accuracy. Model capabilities include the simulation of turbulent flows of liquids and gases including thermal effects and heat radiation with variable fluid properties (density, viscosity, etc.). Particles (solid or liquid droplets) may be tracked through a system as well as phase changes and chemical reactions. Combustion simulation is now becoming possible with the recent development of more widely applicable and accurate physical models of the combustion process itself and the ability to use very detailed geometries.

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Within the field of distillation examples of common applications of CFD are: • Simulating vapour flow patterns in flash zones to enable the most appropriate

modifications to be made to reduce liquid entrainment. • Computing pressure drops through sections of VDUs. • Examining temperature profiles and residence times for liquid on chimney trays. • Predicting concentration profiles and mixing rates in fractionators. • Evaluating flow distribution and pressure drops in pipe and channel networks. • Designing heater flues to ensure uniform flow across tube banks. Benefits of CFD simulation and caveats relating to its use Increasingly numerical simulations are being used alongside more traditional methods in the design and analysis of process plant. CFD and other simulation tools can give unique insights into the plant operation and what effect modifications may have. Simulations can be beneficial in many situations for a number of reasons: • It is possible to evaluate parameters unmeasurable in experiments. • Scale-up uncertainties are largely eliminated as systems may be modelled full-scale. • Test conditions are repeatable and modifications are easy to assess and compare. • Where simulation is feasible a solution is often available quicker and cheaper than using

pilot plant or experiments and there are no safety or environmental hazards. It should be emphasised that the solution from any simulation, not only CFD, will only be as good as the representation of the detail of the physical system and the accuracy of any sub-models included. Each CFD model must be carefully developed in order to be fit-for-purpose including sufficient detail for a reasonable prediction of the behaviour of the fluid to be obtained but without being overly complex. In addition, after a solution is obtained it is essential that the results are interpreted carefully and appropriately for the modelling to be of value. CFD techniques are of course under continual development to improve the accuracy and generality of existing models and to devise new methods of modelling complex fluid dynamic phenomena. There are some areas where CFD techniques are not yet developed sufficiently or where there are too many uncertainties for reliable modelling to be possible. For instance full simulation of 2-phase flow on a tray or in a packed bed is still at the research level and will not be easily or widely applicable for a few more years. Theory of CFD modelling CFD works by dividing the flow domain (which may be two or three dimensional) into a mesh of grid cells. The appropriate flow equations (momentum, transport and energy) are then solved for each cell to predict the values of all flow parameters (pressure, velocity, density, concentration, temperature etc.) throughout the domain. Additional terms and equations are introduced to allow for phenomena such as turbulence effects, fluid property variations, buoyancy, two-phase conditions and chemical reactions. Internal components are

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modelled via blockages or porosities in the grid. Through the use of body-fitted co-ordinates the grid may be distorted in order to follow any non-linear or complex shaped flow region. In almost every case the equations must be solved iteratively, starting from an initial guess, until the solution converges to one that satisfies all the equations to sufficient accuracy. Modelling sections of distillation columns usually requires 6 or more variables to be computed at each of 40,000 or more grid cells in the domain requiring quite a few hours for the simulation even on a high-power workstation. Hence a balance has to be made between the requirement for a detailed and exact solution and the computational, staff and real time cost of obtaining it. Examples of the application of CFD modelling within BP Oil (a) Flow distribution through mist eliminator pad in vapour-liquid separator Oil carry over from vessels in the DiMe and Furfural solvent extraction units at Llandarcy Refinery were reducing plant capacity. GENESIS simulations for the units showed that the expected hydrocarbon carry over was considerably lower than that being obtained in practice. The difference indicated entrainment of liquid droplets with the vapour passing through the vessels. The simplest solution was to install a mist eliminator near the top of the vessels. CFD simulation was performed to ensure the optimal placement of the mist eliminators. The simulations showed that the tangential inlet causes rapid, large-scale circulation with most of the up-flow near the vessel walls and down-flow in the centre. Although the resistance of the knit-mesh pad tends to confine this non-uniform velocity profile to beneath the mist eliminator, high velocities remain through the mist eliminator near the vessel walls and there is little flow, sometimes even recirculation, through the centre. The range of velocities predicted in the pad (-1.0 to 5.0 m/s) was considerably outside the normal efficient operating range for such a mist eliminator. CFD was used to predict the flow distribution through the mist eliminator with a series of vortex breakers installed between the inlet and the knit mesh pad until an optimal design was identified. This sheet metal X-shaped device limits the circulation of the vapour whilst restricting the tendency for the vapour to be deflected upwards at high speed towards the pad. Vapour velocities of 1.5 to 2.5 m/s were now predicted in the mist eliminator with the high velocities confined to a relatively insignificant area. Figure B-1 shows the vapour flow at the inlet elevation and the vapour velocity distribution mid-way through the mist eliminator before and after the installation of the vortex breaker. The CFD simulation gave considerable insight into the operation of these vessels and allowed the comparison of various internal modifications much faster and more cheaply than would have been possible in physical experimentation. Upon subsequent start up the oil carry over was eliminated. Had the mist eliminator been installed without the vortex breaker oil carry over would have remained.

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(b) Vacuum distillation column flash zone modelling During shutdown and inspection considerable coking of the packed bed was discovered in a VDU at Castellon. The cause was attributed to poor liquid distribution onto the top of the bed and liquid entrainment from the flash zone through the chimney tray. A model of the flash zone including the inlet ducting, chimney tray and packed bed was established to demonstrate how the entrainment might be occurring and thus how it may be reduced. The simulations showed very poor flow distribution between the risers in the chimney tray - indeed there was negative (down-flow) through the central risers because the tangential inlet arrangement induces considerable swirl throughout the flash zone. Additional simulations demonstrated improvement by installing a slightly higher pressure drop tray and using baffles within the inlet duct or flash zone to remove much of the circulation. Figure B-2 shows the flow distribution between risers for the original design and the improvement made by installing a new tray and baffles in the flash zone. CFD investigation has allowed features of the vapour flow within distillation columns to be identified that would have been very difficult to obtain from analysis of existing units or rig testing. The ease of defining a different tray, inlet arrangement or baffle/deflector configuration in the computer model of the full scale unit allows rapid and cheap assessment of alternate designs and modifications. (c) Liquid temperature and residence time calculation for a chimney tray It is important that liquid on a chimney tray above the flash zone is removed as rapidly as possible particularly in VDUs used for deep cutting. Liquid that remains too long at a high temperature can thermally crack and cause coking. A CFD model of the liquid flow across a surface with blockages has been developed as part of the BP Deep Cut Technology which has the capability to predict flow patterns, temperature profiles and residence times for liquid falling from the packed bed above and from any side feeds. The model has been applied to the chimney tray above the flash zone of Grangemouth Refinery VDU where a cool quench feed is provided to reduce oil temperatures. Figure B-3 shows the flow patterns across the tray, oil temperatures and residence times for this case. It is then possible to model the tray with modified quench feed arrangements (including the use of spargers, weirs etc.) and flow rates and even with different chimney layouts to reduce the high temperatures and large residence times to acceptable levels.

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(d) Heater flue design A recent problem at Nerefco has been hot spots on the walls of the cross-over ducts between the convection and radiant sections of the CDU3 heater causing damage to the refractory linings. Three-dimensional cfd simulation showed that how the cross-over ducts merge and the design of bends have a significant effect on flue gas velocities and the balance between the flows from the radiant section outlets. The design of the cross-over ducts at Nerefco is causing a significant imbalance between the flow from the East and West outlets from the radiant section as well as high velocities beside the refractory lining at the locations where damage was noticed. Various changes were considered and 'tested' using CFD in order to improve flue gas flow patterns though the heater. The preferred modification to obtain a balance between the flows from the radiant section's east and west outlets would be to have separate ducts from each outlet that merge at the convection section so that the pressure drop through each duct is equal. Also, with careful design of expansions and bends, the possibility of refractory damage can be minimised.

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Figure B-1 Mist Eliminator Performance Improvement

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Figure B-2 Chimney Tray Vapour Flow Improvement

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Figure B-3 Chimney Tray Temperature Profile Analysis

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Distillation Environment Scope The Distillation Environment (DE) is a structure in which resides a number of PC programs for performing calculations on column internals. The DE allows access to a wide variety of programs from many different sources. Availability The DE is issued to all registered holders of the Distillation Handbook, and may be obtained by others on request. Where a recipient has more than one Manual, one set of disks only have been provided, the assumption being that those disks will be mounted on their local network. The current version is Version 6 (May 1998). Confidentiality Distillation Environment has no copy protection. This has been done to make the programs more easily accessible to BP users, or others authorised to access the disk. It is the responsibility of the designated disk holder to ensure that no unauthorised copying is made.

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Programs available The following programs are available on the current (1998) version of the DE (Version 6) Program function

Program Application Restriction on use

Source (1) Version Tray FRI 3.0 (Jan 1996) Dualflow trays Min of 50%

shareholders must be FRI members

Tray FRI 4.1 (1997) 4.2a (1998)

Sieve trays Min of 50% shareholders must be FRI members

Tray BP/Glitsch 2.1 (Feb 1998) Glitsch valve trays None Tray BP/FRI 1.3 (Feb 1998) Baffle trays None Tray Nutter 3.03 (Apr 19

96) Nutter valve trays None

Tray Norton not dated Norton standard valve tray None Tray FRI Aug 1993 FRI bubble caps Min of 50%

shareholders must be FRI members

Packing FRI Aug 1993 Various random packings Min of 50% shareholders must be FRI members

Packing Sulzer 6.0 (May 1998)

Sulzer structured packings None

Packing Nutter 3.03 (Apr 19 96)

Nutter random packings None

Packing Norton not dated Norton random and structured packings None Packing BP 1.2 (Oct 1997) Structured and random packing

hydraulics from tests at Sunbury plus the Generalised DP correlation

None

Packing ETA Nov 1993 ETA random packings, using the Generalised DP correlation

None

Direct Contact Heat Transfer

BP 1.2 (Mar 1998)

Heat transfer in PA spray region None

Distributor BP 1.7 (Apr 1998) For hydraulic design of gravity liquid distributors

None

Distributor BP 1.1 (Mar 1998)

For design of spray liquid distributors None

Liquid-liquid extraction

SPS Lexcol

1.11 (1994) Calculate mass transfer data and column diameter

100% BP companies only.

Collector trays BP 1.3 (Mar 1998)

Calculate accumulator tray hydraulics None

Notes: (1) Although disks of programs from Koch and Glitsch are available, they cannot be incorporated in the DE due to their anti-copy protection which the vendors will not remove. The tray routines fot the Glitsch and BP valve tray programs are the same.

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Description of programs FRI Dualflow tray. Although dualflow trays are not normally used in BP refineries, one Dualflow tray is in service at an OUS refinery and others are used by BP Chemicals This program gives the FRI rating procedure for calculating tray capacity and efficiency. FRI sieve tray. A rating program for 1 to 4 pass sieve trays based on FRI correlations. This is the standard program for sieve trays within BP Oil. Full tray hydraulics, capacity and efficiency are calculated. The interface for the FRISIV tray programs is old-fashioned compared to many of the other tools in DE. FRISIV version 4.1 has been in use since 1996 years and has proved acceptable for rating sieve trays. Version 4.2a was released in January 1998. This requires more memory to run than the previous version and so may not be useable on many COE computers within BP. When users access the FRISIV programs via DE6, version 4.1 will be loaded if the computer does not have sufficient memory to load version 4.2a. The changes recently made by FRI include the following: 1. The Topical Report 128 effect of entrainment and weeping on mass transfer efficiency has

been added to the mass transfer efficiency correlation of TR 126. 2. The revised Jet-Flood correlation, TR has replaced TR 112 correlations. 3. The standard deviation of design and the probability of flood has been replaced by target

confidence limits for jet-flood and downcomer flood. The numbers are used to calculate the maximum load at flooding conditions.

4. Maximum load as percent jet flood is reported in the summary page. 5. A new section titled "FLOODING CONFIDENCE LIMITS" has been added. In this

section Target Confidence Limit, Confidence limit at Design Loads and Max load at Target Confidence limit are displayed. This is done for both Jet-Flood and Downcomer- Flood at constant L/V and constant liquid rates.

6. If the downcomer clearance are is not specified, the program uses the downcomer length at bottom and the downcomer clearance from tray and the seal pan depth to calculate the side escape area.

7. A new output page has been added which displays the vapor load, capacity factors and percent downcomer flood for constant L/V and constant liquid rate for downcomer flood.

8. Three new help screen has been added to the Center and Off- Center dimensions screens. The screens help users to identify the required dimensions for the downcomers seal pans.

9. Few formats and comment in the screens and output have been changed to make the hard copy output consistent with the screen output.

FRI also provide their Sieve Tray Hydraulics correlations in the form of spreadsheet. This format is not as robust or user-friendly as the stand-alone programs but has advantages for in-depth analysis. Contact the OTN Separation team at Sunbury for a copy if you would like to use this version.

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BP/Glitsch valve tray program. This program is now (DE06) in Windows format. The program is based on the Glitsch valve tray rating and design procedure and is the standard program in BP Oil for evaluating the hydraulics of circular valve trays and the Glitsch V-0 fixed valve tray. The facility to use VG-0 valves has now been added. The original program (believed to have been acquired from Glitsch in the 1970's) has been extensively modified by BP although the basic procedure remains unchanged. The procedure can also be used for Koch trays. Glitsch (now Koch-Glitsch) have made available a version of their own program for Nye trays and Superfrac trays. However, yhey will not permit its use in DE-6 and the program is licensed (at no cost) only to named Engineers in the OTN Distillation Group. Thus, these Engineers will be able to design/rate the trays on behalf of sites. BP Baffle tray program. Written by BP incorporating the FRI algorithms. Also includes BP correlations for cracking on the tray. Nutter tray. Covers the Nutter rectangular valve tray. Note that the calculated jet flood predicted by this program is higher than would be expected - refer to Volume 1, Section D for further information. FRI bubble caps. The FRI bubble cap procedure applies only to circular caps, not to rectangular caps. The program predicts capacity and efficiency. FRI packing. This program calculates the hydraulics, capacity and HETP of a range of packings tested by FRI including Raschig and Pall rings and ceramic saddles. Sulzer packing. New program. Covers design and sizing of columns using Sulzer structured packings. The procedure is valid for a pressure range of vacuum to 2 bars. Norton structured and random packings. Covers the hydraulics of Norton packings but is not their latest program. Their latest program has been held pending some revisions and is expected to be issued on general release in 4Q98 and will be Windows based. Norton trays.. This program is expected to be issued on general release in 4Q98 and will be Windows based. It will include their Triton high capacity trays. BP Packing. Covers all packings (structured and random) tested by Sunbury. The program is applicable from vacuum to 50 bars and is the preferred program for structured packings. The program incorporates the ETA Generalised pressure drop correlation. ETA packing. Covers ETA random packings, including polypropylene packings, and uses the Generalised pressure drop correlation. Direct Contact Heat Transfer. Procedure for calculating the heat transfer duty available in the spray region of a pump around zone.

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Liquid distributors. Facilitates the hydraulic design of gravity liquid distributors, based on Sunbury correlations. Hydraulic conditions, spray geometry and level tolerance are calculated. Spray Distributors. This program permits the rating, optimisation and full design of spray distributors. It can calculate the location of spray nozzles, the quality of the spray distribution obtained, the amount of liquid being sprayed onto the wall and the bed area not covered (if any). Liquid-liquid extraction. Program for assessing packed bed depth of liquid-liquid extraction columns plus some other forms of contactor. The program has a column diameter calculation routine, but based on limited investigation this is not favoured by BP Oil; Dell Pratt is preferred. Accumulator Trays. Procedure for calculating accumulator tray hydraulics. It incorporates an improved pressure drop correlation (CFD showed that loss past the caps were already in the BP correlation), better emergence area calculation, better data formatting and additional warning messages for circular risers.

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Simulation Background The data contained in this Section is based almost entirely on the presentation given by Stuart Fraser at the Simulation PIT in 1996. The aim is to illustrate improved profitability by using simulation for process monitoring. The data given should be an invaluable aid to all engineers simulating distillation columns. The section covers examples of how relatively simple HYSIM/PRO2 unit models can be used to identify opportunities for improving unit profitability. These models should be used by unit engineers to check out the performance of their units. Simulation notes are presented and examples of optimisation opportunities are given for vacuum units and absorber/stripper type gas plants. Notes on generating McCabe-Thiele plots are also given. Simulation of vacuum distillation units Vacuum units are more difficult to simulate accurately than crude units - the results can be significantly influenced by the selection of the thermodynamic system, feed characterisation, and also by the way the column internals are represented. Main points of note are as follows: Set FBP of pseudo-components to 1000°C as per the Crude unit simulation. Thermodynamic systems recommended are summarised in Volume 1, Section BA. For HYSIM, the Essotab method is recommended. For feed characterisation, it is usually not practical to backblend the Vacuum unit products since: • mass balance closure on VDU’s is often poor (less than 98%) • analysis of the vacuum residue is poor covering less than 30 % of the residue boiling

range. Although higher recoveries (70%) are possible with High Temperature Simdis equipment. If possible, High Temperature Simdis analyses should be used for prediction of VDU performance. Therefore, it usually best to characterise the VDU feed from a Simdis sample of the Atm Residue feed (assume that the SimDis is a TBP weight analysis). Alternatively, generate the feed TBP input from a crude unit simulation. The efficiency of the crude unit stripping zone should be varied to match the front end analysis of the feed TBP analysis. The FBP for the pseudo-components generated should be set to 1000 C (to generate the correct special properties).

A fixed amount of cracked gas should be added to the feed equivalent to 0.1-0.2 %wt on feed. In the absence of any compositional data, a typical gas composition is:

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Component mol% CO2 1.6 H2S 15.3 C1 32.0 C2 12.9 C3 14.0 iC4 1.1 nC4 8.6 iC5 1.4 nC5 4.1 nC6 9.0

The sour gas rate may be a critical parameter if the ejector set is a bottleneck. The cracked gas rate is difficult to estimate with no direct metering facilities. Methods used for estimating the sour gas rate include radio-active tracer and estimation from measuring dP across the foul gas burner. The sour gas rate will increase the harder the furnace is fired. The heater and transfer line for the VDU should be simulated in the same manner as the CDU. Coil steam is often added to VDU heaters and this should be included in the feed to the VDU heater. Heater outlet temperatures for VDU’s are typically 410 - 425 C; the heater outlet pressure is normally in the range 0.15 - 0.4 Bara. The heater outlet pressure will have a significant impact on the process heat duty, and therefore, it is important that an accurate value is used. A check should be carried out to confirm that the transfer line is operating below sonic velocity. The inlet velocity at the end of the transfer line (flash zone inlet) should be calculated and compared with the sonic velocity from : Sonic Velocity (m/s) = 91.2 (Cp/Cv *T/Mw)^0.5

where T - temperature (K) Cp/Cv ratio of specific heats at constant pressure/volume (can be approximated to

1.03 for VDU feeds) Mw - feed gas molecular weight The transfer line pressure will increase to ensure that the operating velocity is always below sonic, but this will adversely affect the heater duty and also the VDU cut point attainable. An estimate of the tray count and pressure profile should be generated. It is important that realistic pressures are used for the trays. Most vacuum units are packed and consequently, the column pressure drop between the column top and the flash zone will normally be in the range 10 - 20 mbar. A higher pressure drop profile should be used for the trayed section in the residue zone. The tray efficiency in the stripping zone should not be over-estimated. Typically, no higher than 50% efficiency should be used. An example of a ‘tray’ pressure profile and column feed arrangement is given below. The pressure allocated to the wash tray is fairly critical and will affect the required wash oil

CONFIDENTIAL This information is of a confidential or proprietary nature. For internal BP use only, not to be transmitted to others.


flowrate. The wash zone tray pressure should be lower than the flash zone pressure to reflect the pressure drop through the wash zone bed. A dummy tray can be added to estimate the overflash quality if there is any intent on drawing overflash. In this case, some estimate for flash zone entrainment will be required. The entrainment rate can only be confirmed by sampling the overflash and matching the simulation overflash quality to the sample.

SET AT HEATER COT AND OUTLET PRESSURE

FLASH CONSTANT ENTHALPYAT FLASH ZONE PRESSURE

VDU FEED ARRANGEMENT

FEED VAPOUR

OVERFLASH

WASH OIL

HVGO

DUMMY TRAY

STEAM

VDU FEED +CRACKED GAS

ENTRAINMENT

VERY LOW EFFICIENCY

FEED LIQUID

tray 5

tray 6

tray 7

tray 8

typical column pressure profile :tray 7 - flash zone presstray 8 - f/z press + 10 mBar tray 6 - f/z press - 4 mBar

Virtually all VDU’s in BP are now packed columns, and therefore, an initial estimate of the tray count should be based on the depths and grades of packing in the various zones. Pumparounds should be modelled with one or two stages (can be confirmed from test run match). The wash zone is a very important part of the VDU model. The simulation of the wash zone will have a significant impact on the required wash oil rate and the residue cut point of the unit. As an initial assumption, it is recommended that the wash zone efficiency is set at 1 theoretical trays. Accurate specification of the wash zone pressure is also important. The wash zone efficiency will affect the calculated overflash rate. For a fixed wash oil rate, the higher the wash zone efficiency the lower the calculated overflash rate. To avoid wash zone coking, it is essential that we operate the VDU above a minimum overflash rate, and therefore, it is important that we do not under-estimate the wash zone efficiency.

CONFIDENTIAL



Initial specifications for VDU simulations are as follows : • liquid reflux rate below draw trays (set at nominal zero flow if tray is a total draw tray) • overflash rate at 0.1 - 0.2 USGPM/ft2 (0.24 - 0.48 m3/m2 hr). The required wash oil rate

to meet this overflash rate will be calculated from the simulation. The overflash should be drawn out of the column if there is an actual overflash draw in operation.

• overhead temperature specification (70 - 80 C) or 10 - 20 C above the calculated water dew point

• bottom pumparound rate and return temperature • These specifications should allow the simulation to converge. Thereafter, the simulation

specifications should be tightened up to match : • vacuum distillate yields. The qualities (TBP’s) should be matched by varying the tray

efficiencies. Avoid using individual tray efficiencies if possible, and do not use efficiencies on trays with coolers or heaters. A good match for vacuum distillate concarbon should also be sought.

• vacuum residue yield and TBP’s. The TBP data can be matched by varying the stripping efficiency. The residue yield and TBP quality can also be varied by changing the wash zone efficiency. Wash zone efficiencies below 0.6 theoretical trays should not be used.

• pumparound heat duties. The column temperature profile is predominantly affected by the distillate draw rates. The addition of trays in the pumparound trays will increase the pumparound draw temperature. Small changes to the heater inlet temperature and pressure may be required to improve the simulation match of pumparound duties.

Optimization opportunities for VDU’s For most VDU’s there is typically a $3 - 5/bbl benefit in maximising the vacuum residue cut point (assuming the incremental vacuum distillate can be processed in the FCC). Good VDU operations would normally achieve residue cut points of 1080+ F. Key parameters in optimising the VDU operation are typically as follows : • heater outlet temperature • wash oil rates • develop operating constraint plots showing column flood limits (this will allow optimum • operating flash zone pressure to be identified) • review slop oil yields (this is usually rerun through CDU or FCC) • review vacuum system performance Vacuum heater outlet temperature. The effect of raising or lowering the heater outlet temperature is shown below. The effect of varying the heater outlet pressure is also shown as a parameter. As would be expected, increasing the heater temperature has a big impact on residue yield. However, - beware of increasing heater coil temperatures (COT) and the potential to coke the heater. Check the heater design for operation at higher temperatures. Prolonged operating experience at higher COT’s from other VDU’s should be obtained before raising the heater temperature.

CONFIDENTIAL



The simulation results below indicate that lowering the heater outlet pressure from 5 to 3 psia reduces vacuum residue yield by about 0.9% vol on feed. This illustrates the importance of the transfer line design. Increasing the transfer line size may be an attractive revamp option if the unit is operating with a high heater outlet pressure. If required, the transfer line pressure drop can be modelled in the simulation.

ALLIANCE VAC RESIDUE V'SHEATER TEMP AND PRESS

16

18

20

22

24

26

760 770 780 790 800

HEATER TEMP (F)

RES

IDU

E YI

ELD

(M

BPD

)

5 psia

3psia

4 psia

Review wash oil rates. Depending on the severity of the heater outlet temperature, consider reducing overflash rates to a minimum of 0.1- 0.15 USGPM/ft2. In most units, the overflash is neither drawn or measured, therefore, the overflash is a calculated value. Be careful how the overflash is calculated within the simulation - the wash zone efficiency, wash zone pressure drop and feed rate, will have a significant impact on the calculated overflash. If the wash zone is run dry, it will certainly coke-up ! Very careful consideration is required before lowering wash oil rates. Operating experience at lower wash oil rates should be sought from other VDU’s at similar operating conditions, to confirm that coking of the wash zone was not a problem. Estimated overflash rates for Alliance VDU are shown below :

ALLIANCE VDU - ESTIMATED OVERFLASH RATES

0.00

0.05

0.10

0.15

0.20

0.25

400 500 600 700WASH OIL RATE (BPH)

OVE

RFL

ASH

RA

TE

(USG

PM/F

T2)

17000

18000

19000

20000

VAC

RES

ID R

ATE

(B

PD)

CONFIDENTIAL



Operate the VDU with the lightest quality feed. in general, available capacity in the VDU should be utilised by lowering the CDU cut point. For a maximum VDU heater outlet temperature, the lighter feed will result in an increased vacuum residue cut point (higher % vaporisation). This option may only be attractive where Vacuum Gas Oil is drawn to the diesel pool. As a general guideline, the VDU should be operated to ensure that the Vac Gas Oil draw is fully maximised up to some equipment constraint. This mode of operation will ensure that the VDU residue cut point is maximised. Minimise slop oil yield. in most refineries, slop oil is re-processed to the crude or FCC unit and this consumes capacity. There is an incentive to maximise top gas oil recovery and minimise VDU slop oil. Target VDU top temperatures around 20 C above the calculated water dew point. This will normally be equivalent to a top temperature of around 70 C. Confirm whether there is scope to increase the top pumparound capacity, and also consider option of increasing lower pumparound duties (feasibility of increasing bottom pumparounds will depend on product quality specifications). Typical slop oil rates should be no higher than 0.5 - 1 % on feed. Develop VDU operating plots showing column flood limits versus column flash zone pressure with heater outlet temperature as a parameter. An example of these plots is shown in the attached Figures 2 - 5. The column flood limit can be plotted as a vapour phase Cs factor : Cs (ft/s) = vap vel *(rhov/(rhol-rhov)^0.5 vap vel = ft/s The maximum operating Cs limit for the VDU internals can be shown as a horizontal line. A good general design limit for modern structured and grid packings is a Cs limit of 0.4 ft/s. These plots can be used to determine the optimum heater temperature and column pressure operation within the hydraulic constraints of the column internals.

CONFIDENTIAL



VDU COLUMN LOAD FACTORCs (ft/s) v's FLASH ZONE PRESSURE

97.3 MBPD FEED

0.35

0.37

0.39

0.41

0.43

0.45

0.47

25 30 35 40 45

FLASH ZONE PRESS (mmHga)

CO

LUM

N C

s FA

CTO

R (f

t/s)

HEATER COT = 780 F

COT = 765 F

FIG 2

VTB YIELD V'S FLASH ZONE PRESS97.3 MBPD VDU FEED

22000

23000

24000

25000

26000

27000

28000

25 30 35 40 45


VTB

YIE

LD (B

PD)

HEATER COT = 780 F

COT = 765 F

FIG 3

VDU COLUMN LOAD FACTORCs (ft/s) v's FLASH ZONE PRESSURE

89 MBPD VDU FEED

0.35

0.36

0.37

0.38

0.39

0.4

0.41

0.42

0.43

25 30 35 40 45


VAC

CO

LUM

N C

s (ft

/s) HEATER COT = 780 F

COT = 765 F

FIG 4

VTB YIELD V'S FLASH ZONE PRESSVDU FEED RATE 89 MBPD

20000

21000

22000

23000

24000

25000

26000

25 30 35 40 45


VTB

YIE

LD (B

PD)

HEATER COT = 780 F

COT = 765 F

FIG 5

CONFIDENTIAL



Simulation of Gas Plants with Absorber/Strippers Conventional gas plants with absorber strippers are normally quite difficult to match up to plant data. Plant data is often incomplete with little information on recycle rates and qualities. These units usually have 3 or 4 LPG recycle streams and few independent variables. Main points of note in modelling gas plants (FCC or Sat Gas plants) are as follows : • Thermodynamic systems to be used as summarised in Distillation Handbook. For

HYSIM, PR is favoured. The gas plant feed characterisation is normally generated by back-blending the products. A good mass balance in required (better than 98% closure).

• A typical gas plant configuration (Alliance) is shown below.

LEAN OIL FROM MAIN FRACT DRUM

COMP DISC

HP DRUMV - 15

SPONGE OIL

RICH OIL TO M

SEC ABSORBER V-19

TAIL GASREC1

REC2

REC3

PRIM ABS V-16

STRIPPER V-16

COKER LIQS

TO DEPENT V-20

LEAN OIL FROMDEPENT BTMS REC4

X-18

V-18

When modelling gas plants, the simulation can be simplified by considering the Absorber/Stripper and HP feed drum as a single column. This will eliminate 2 recycles and speed up convergence. The pressure profile in the ‘single column’ should be set up to match the pressure drop between the HP drum and primary absorber and stripper.

CONFIDENTIAL



HP DRUMV - 15

model as tray with specified temp

V-18 model as tray with specified temp

Q

Q

PRIM ABS V-16

STRIPPER V-16

The C3 recovery is a key parameter which the simulation should match with a deviation of less than 3 %. However, in practice it can be difficult to match the actual gas plant performance due to the interaction between many process parameters. Factors influencing the C3 recovery include : • lean oil rate, quality and temperature • sponge oil rate, quality and temperature • absorber and stripper tray efficiency • stripper base temperature (or C2 spec in base) • high pressure feed drum temperature and pressure • heat removal in absorber • heat removal from HP drum feed inlet streams (HP compressor discharge, stripper

overheads) • extraneous feeds (particularly those containing hydrogen) The lean and sponge oil rate, quality and temperature should be set to match plant data. Extraneous feeds should also be set to match the unit data. In general, the tray efficiencies for the primary absorber tend to be low. The number of theoretical stages for the primary and secondary absorber should be set initially to 25 % of the actual tray count. There is usually no significant benefit in having more than 5 - 6 theoretical trays in the primary absorber. The stripper tray count should initially be set to 50 % of actual trays. The process conditions for the HP drum should be set as per the plant operating data. The ethane content in the stripper (deethaniser) bottoms is a key parameter, and the simulation should be set up to match the measured lab analysis (ethane in propane product). The stripper operation will have a major impact on the gas plant simulation and the recycle

CONFIDENTIAL



rates. Most units tend to operate in an over-stripped condition (base temperature is too high), often to preferentially reject H2S to tail gas (rather than LPG). This generates LPG recycles which are difficult to match by simulation. If there is plant data to confirm the stripper overhead rate, then this should be matched. Also the reboiler duty for the stripper should also be matched closely. The stripout rate can also be varied by changing the number of stripping trays. The stripper base temperature should be matched within a 5 C deviation. The C3 recovery in the secondary absorber tail gas should be matched as best as possible. This is a difficult process since there is a significant interaction between the deethaniser, primary and secondary absorber. Lean oil rates, and process stream conditions should be set up to match the plant data. Recommended simulation fitting method is as follows : • match deethaniser stripout (in particular match deethaniser reboiler) • fix HP drum temperature and try to match HP drum liquid rate (by varying stripper and

absorber tray efficiency). • use temperature profile in primary and secondary absorber to obtain an indication of tray

efficiency. • vary tray efficiency in both primary absorber and secondary absorber to match propane,

propylene, and C4 recovery. Secondary absorber efficiency will have a more pronounced effect on C4+ recovery. Primary absorber will mainly influence C3 recovery. Where possible, reduce the number of theoretical stages rather that vary the tray efficiency. Never use fractional trays for feed trays and for trays with heaters or pumparounds.

Typical C3 concentrations (%mole) in absorber tail gas are : propylene,8/propane,2. Good FCC gas plant C3 recoveries are around 85 - 90 %. C4 recoveries should be around 98 %. The H2S distribution between the absorber tail gas and LPG’s will vary from 50 - 100 % depending on the deethaniser stripping severity. Gas Plant optimisation opportunities LPG recovery is the key performance indicator for gas plants. The effect of key operating parameters on LPG recovery are quantified below by way of a number of simulation cases. Note - the LPG recovery improvements shown below are relative to a current base case operation - the LPG recovery improvements predicted are not cumulative. The incremental improvements in LPG recovery by optimising gas plant operation, will diminish the better the base case operation. Stripping severity is one of the most important variables affecting LPG recovery. For Alliance FMC gas plant, the incremental LPG recovery by varying the stripper base temperature is shown below :

CONFIDENTIAL



LPG RECOVERY V'S STRIPPING SEVERITY

85.0

86.0

87.0

88.0

89.0

0 0.5 1

C2 IN PROPANE (% WT)

C3

REC

OVE

RY

(%)

1500

2000

2500

LPG

IN T

AIL

GA

S (B

PD)

In the case of Alliance, unsaturated LPG’s are routed to the Alkylation unit, so the maximum acceptable C2 content in the Alky feed will be limited by the condensing capacity of the Alky depropaniser. However, as the above graph shows, if the ethane content in propane is zero (as is the case at many BP Refineries), there is a significant LPG recovery benefit in increasing the ethane content even by a small amount (less than 0.25 % ethane in propane). The stripper base temperature difference over the range of 0 - 1 % ethane in propane is low at approximately 6 F. Therefore, GC analysis of depentaniser (or debutaniser) overheads is required for close control of the deethaniser operation. HP drum temperature has a significant affect on LPG recovery as shown below. The X-13’s exchangers are water cooled, and should have an unrestricted supply. Fouling monitoring would be important issue for this service.

C3 RECOVERY V'S HP DRUM TEMP

84.0

85.0

86.0

87.0

88.0

89.0

105 110 115 120

HP DRUM TEMP (F)

C3

REC

OVE

RY

(%)

2000

2500

3000

LPG

IN T

AIL

GA

S (B

PD)

Maximising lean oil rate has a very significant affect on C3 recovery as shown below. Usually, all of the main fractionator overhead liquid is used as a primary lean oil in the Primary Absorber. Additional lean oil can be obtained by recycling depentaniser bottoms to

CONFIDENTIAL



the Primary Absorber (as shown below). Although this improves LPG recovery, it may adversely affect the depentaniser operation.

C3 RECOVERY V'S SEC LEAN OIL

85.0

86.0

87.0

88.0

89.0

90.0

0 5000 10000

SEC LEAN OIL (BPD)

C3

REC

OVE

RY

(%)

1500

2000

2500

LPG

IN T

AIL

GA

S (B

PD)

Minimise recontactor temperature. The LPG recovery by recontacting primary absorber overheads with secondary lean oil is shown below. This shows that the LPG recovery benefits of the Alliance recontactor are fairly marginal.

LPG RECOVERY V'S RECONTACTOR TEMP

87.5

87.6

87.787.8

87.9

88.0

105 110 115 120 125

RECONTACTOR TEMP (F)

C3

REC

OVE

RY

(%)

2000

2100

2200

LPG

IN T

AIL

GA

S (B

PD)

Sponge oil rate and temperature have a significant affect on Alliance FCC gas plant LPG recovery as shown below. Most of the incremental LPG recovery is due to improvements in C4+ recovery. C3 recovery is not significantly affected by the sponge oil absorber.

CONFIDENTIAL



C3+ RECOVERY V'S SPONGE OIL RATESPONGE OIL TEMP AS A PARAMETER

2000

2100

2200

2300

2400

9000 11000 13000 15000SPONGE OIL RATE (BPD)

LPG

IN T

AIL

GA

S (B

PD)

108 F

88 F

Avoid extraneous feeds with a high H2 content. FCC gas plants are often used to recover small quantities of LPG from several extraneous gas streams such as stabiliser offgas and coker wet gas. These streams are often lean and contain high H2 contents. This will have a negative impact on FCC gas plant C3 recovery. Where possible, alternative disposal options should be considered. The net effect on LPG recovery from adding extraneous streams into the gas plant should be reviewed by the simulation.

CONFIDENTIAL



Setting up a gas plant simulation Conventional gas plants with absorber strippers are normally quite difficult to match up to plant data. Plant data is often incomplete with little information on recycle rates and qualities. These units usually have 3 or 4 LPG recycle streams and few independent variables. Main points of note in modelling gas plants (FCC or Sat Gas plants) are as follows : • Thermodynamic systems to be used as summarised in Volume 1, Section BA. For

HYSIM, PR is favoured. The gas plant feed characterisation is normally generated by back-blending the products. A good mass balance in required (better than 98% closure).

• A typical gas plant configuration (Alliance) is shown below.

LEAN OIL FROM MAIN FRACT DRUM

COMP DISC

HP DRUMV - 15

SPONGE OIL

RICH OIL TO M

SEC ABSORBER V-19

TAIL GASREC1

REC2

REC3

PRIM ABS V-16

STRIPPER V-16

COKER LIQS

TO DEPENT V-20

LEAN OIL FROMDEPENT BTMS REC4

X-18

V-18

When modelling gas plants, the simulation can be simplified by considering the Absorber/Stripper and HP feed drum as a single column. This will eliminate 2 recycles and speed up convergence. The pressure profile in the ‘single column’ should be set up to match the pressure drop between the HP drum and primary absorber and stripper.

CONFIDENTIAL



HP DRUMV - 15

model as tray with specified temp

V-18 model as tray with specified temp

Q

Q

PRIM ABS V-16

STRIPPER V-16

The C3 recovery is a key parameter which the simulation should match with a deviation of less than 3 %. However, in practice it can be difficult to match the actual gas plant performance due to the interaction between many process parameters. Factors influencing the C3 recovery include : • lean oil rate, quality and temperature • sponge oil rate, quality and temperature • absorber and stripper tray efficiency • stripper base temperature (or C2 spec in base) • high pressure feed drum temperature and pressure • heat removal in absorber • heat removal from HP drum feed inlet streams (HP compressor discharge, stripper

overheads) • extraneous feeds (particularly those containing hydrogen) The lean and sponge oil rate, quality and temperature should be set to match plant data. Extraneous feeds should also be set to match the unit data. In general, the tray efficiencies for the primary absorber tend to be low. The number of theoretical stages for the primary and secondary absorber should be set initially to 25 % of the actual tray count. There is usually no significant benefit in having more than 5 - 6 theoretical trays in the primary absorber. The stripper tray count should initially be set to 50 % of actual trays. The process conditions for the HP drum should be set as per the plant operating data. The ethane content in the stripper (deethaniser) bottoms is a key parameter, and the simulation should be set up to match the measured lab analysis (ethane in propane product). The stripper operation will have a major impact on the gas plant simulation and the recycle

CONFIDENTIAL



rates. Most units tend to operate in an over-stripped condition (base temperature is too high), often to preferentially reject H2S to tail gas (rather than LPG). This generates LPG recycles which are difficult to match by simulation. If there is plant data to confirm the stripper overhead rate, then this should be matched. Also the reboiler duty for the stripper should also be matched closely. The stripout rate can also be varied by changing the number of stripping trays. The stripper base temperature should be matched within a 5 C deviation. The C3 recovery in the secondary absorber tail gas should be matched as best as possible. This is a difficult process since there is a significant interaction between the deethaniser, primary and secondary absorber. Lean oil rates, and process stream conditions should be set up to match the plant data. Recommended simulation fitting method is as follows : • match deethaniser stripout (in particular match deethaniser reboiler) • fix HP drum temperature and try to match HP drum liquid rate (by varying stripper and

absorber tray efficiency). • use temperature profile in primary and secondary absorber to obtain an indication of tray

efficiency. • vary tray efficiency in both primary absorber and secondary absorber to match propane,

propylene, and C4 recovery. Secondary absorber efficiency will have a more pronounced effect on C4+ recovery. Primary absorber will mainly influence C3 recovery. Where possible, reduce the number of theoretical stages rather that vary the tray efficiency. Never use fractional trays for feed trays and for trays with heaters or pumparounds.

Typical C3 concentrations (%mole) in absorber tail gas are : propylene,8/propane,2. Good FCC gas plant C3 recoveries are around 85 - 90 %. C4 recoveries should be around 98 %. The H2S distribution between the absorber tail gas and LPG’s will vary from 50 - 100 % depending on the deethaniser stripping severity.

CONFIDENTIAL



Generation of a pseudo McCabe Thiele plot The following gives some simple guidelines for generating these plots from the HYSIM environment. With HYSIM, there are two methods for generating a McCabe Thiele plot : (a) An easy way (b) A more difficult (but still easy) way Easy Way BP has developed a calculator routine in HYSIM called !MTO. Type this when in the worksheet environment. A prompt will appear asking you in which column in the worksheet you want to generate a pseudo McCabe Thiele plot. Highlight the appropriate column. A new prompt will ask you the number of trays for the McCabe Thiele plot. Type in the number of trays in the column (e.g. 24). Another prompt will ask you to highlight the light keys. The light and heavy keys in a column are the components which will normally be specified. For example, in the case of a debutaniser, the light key will be nC4 and the heavy key will be iC5. For the light keys, you should highlight (using the Ins key) the light key and every component lighter than this. Once you have done this, a prompt will ask for the heavy keys. You should highlight the heavy key and all of the components heavier than this (e.g. iC5, NC5, NBP100……). That's all - and a pseudo McCabe Plot should appear on the screen. You can print this out and then add the feed location and Q line. The feed location should be plotted from the X-Y diagonal. The feed point is the liquid phase mole fraction of the light key and lighter in the feed. The slope of the feed line (off the 45 degree diagonal) is defined as : Q/(Q-1) where Q is the energy required to fully vaporise the feed, divided by the latent heat of vaporisation of the feed. So, for a saturated liquid feed, the feed line will be vertical. For a vapor feed, it will be horizontal. The slope of the feed (q-line) from the X-Y diagonal should intercept the equilibrium line. If the feed location is optimal, then the q-line should also intercept the stripping and rectification operating lines. More Difficult Way In the HYSIM column environment, select the print options to ensure that you print out - Mole and Fractions (Print, Format, Mole; Print Format Fractions). Also, set up the Print options, so that you are printing to a Spreadsheet and also to a File (Print, Format, Spreadsheet, and then Print, File - then supply a file name). All subsequent screen outputs will be routed to the named file. In the column unit operation environment, print out the component mole fractions for the light and lighter keys by : Print, Composition, and then INS all of the light key and lighter



components. Once this is carried out, type - Print, File and Esc. This will close off the output file. Now go into Excel and open up the above output file. This is an Excel Text PRN file, and when you open this, using the import wizard, click on the Delimited Option, and then click on the Comma delimiter option. You should now see a series of columns of mole fraction data for each tray in the column, for both the liquid and the vapor phase. There is also a column showing the equilibrium constant for each tray. The columns of K values should be deleted - these are not required. Once all of the columns of K Values are deleted, you should cut and paste the data into a single array - one row for each tray, and one column for each component in the liquid phase and one column for each component in the vapor phase. So if your fractionation column contains 10 light and lighter key components and 30 trays, you should have a block of data of 30 rows deep and 20 columns wide. Then simply sum up all of the liquid phase mole fractions for each tray. Also sum up all of the vapor phase mole fractions for each tray, and then using an X-Y plot, plot the liquid mole fractions versus vapor phase mole fractions. This will generate the Pseudo McCabe Thiele plot. Add the q-line as described above.

CONFIDENTIAL


BP Oil Distillation Handbook Rev: 1 Volume 1 General Design & Column Internals Date: Jul-95 Section C Design Pressure & Temperature Page: 1C.1

Contents page Vessel Design 1C.2 Design pressure Design temperature Selection of operating pressure Revision History Rev 1 Original Issue DR 07/95

CONFIDENTIAL



Vessel Design Design pressure In determining the design pressure of a column a margin shall be allowed above the normal operating pressure to accommodate normal pressure fluctuations and to allow for any pressure alarms and pressure limiting instrumentation required. As a general guide the following shall apply:

Operating pressure Margin above operating pressure Minimum margin barg psig percent barg psig

up to 70 up to 1000 Not less than 10 1.0 14.7 above 70 above 1000 See BP Oil - -

Vessels subject to a vacuum shall be designed for full vacuum unless a vacuum break valve is fitted. Vessels designed to less than full vacuum, i.e., half vacuum, may be economical in certain applications. The possibility of vacuum conditions occurring due to low ambient temperatures shall be taken into consideration in the design when necessary, e.g., in butane service. Alternative operating conditions away from the normal shall also be considered. As an example, most vessels are subject to steam out prior to entry. It is suggested that a minimum steam out pressure of 1 barg be quoted. Vessels which normally run full of liquid shall be designed to withstand the shut-in head conditions of pumps so that full flow liquid relief is avoided. The shut-in pressure shall, for example, be based on the maximum diameter impeller that can be fitted to the pump or the trip speed of any turbine driver. When the relief valves of low pressure vessels vent into a common closed system, such as a flare or blowdown system, into which relief valves from higher pressure vessels can also vent, consideration shall be given to raising the design pressure of the lower pressure vessels if this would result in savings on the overall cost of the pressure relieving system and the pressure vessels. There may be considerable savings on the flare/blowdown system should the design pressure of a column be raised to a level that there is no heat input from a reboil heat source at the column design pressure. That is, the column bottoms temperature is the same or higher than that of the reboil medium. The designer should be aware of the effect of design pressure/temperature on flange ratings. Savings can occur on the vessel and upstream/downstream piping, valves, equipment, etc., if the design pressure/temperature combination be such that the excursion into a higher flange rating is avoided. This is not always possible.

CONFIDENTIAL



Design temperature The maximum design temperature which is used to determine the appropriate nominal design strength for the selected material shall not be less than the actual maximum metal temperature expected in service.

Operating temperature (metal temperature)

Margin above operating temperature

°C °F °C °F Up to 425 Up to 797 15 27 Above 425 Above 797 See TDU

A minimum design temperature based on the local climatic conditions shall be stated if the vessels are intended to be pressurised at ambient temperatures, and/or are likely to be left unprotected on site. In the latter case, the vessel designer should take into account the minimum design temperature with ambient pressure only. The minimum design temperature which is used to determine the suitability of the material to resist brittle fracture shall be the lowest metal temperature expected in service. In the case of components thermally insulated externally, the lowest metal temperature shall be taken as the minimum temperature of the contents of the vessel at the appropriate loading condition. In the case of components not thermally insulated, the minimum temperature of the components under operating conditions and the method used for assessing the lowest metal temperature shall be reviewed with TDU. In cases where the calculated membrane stress can vary with the minimum design temperature, e.g., auto-refrigeration during depressurisation, the various combinations of stress and temperature shall be evaluated to determine which is the most onerous for material selection. Where different metal temperatures can be confidently predicted for different parts of the vessel, the design temperature for any section of the vessel may be based on the predicted metal temperature. This is often the case in crude columns where there may be three or more design temperatures. However, care should be taken in assessing the effect of stripping steam temperature on the column metal temperatures. Stripping steam is usually in the range 350-375°C (662-707°F).

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The possibility of auto-refrigeration due to plant depressurising shall be considered in the design. Two distinct operating conditions shall be taken into account when calculating the minimum design temperature: • No requirement for repressurisation • Partial or full repressurisation within 20 minutes so that there is no significant ambient

heat input For vessels containing fluids with a boiling point less than 0°C at atmospheric pressure, the minimum temperature of the boiling liquid shall be established from a single stage flash calculation. If the temperature is less than -50°C (-58°F) at atmospheric pressure, stepwise flash calculations should be performed if it is expected that a temperature higher than -50°C (-58°C) will be obtained. If the minimum temperature falls below -50°C (-58°C), heat transfer calculations shall be performed to establish the minimum vessel wall temperature, taking into consideration the thermal capacity of the vessel and the atmospheric heat inflow. The design minimum temperature for the vessel shall be 10°C lower than the minimum calculated temperature. When statutory requirements state a minimum temperature, or the boiling point of a liquid is specified as the minimum temperature, a further 10°C (18°F) reduction in temperature shall not be applied. When the minimum temperature is lower than -50°C (-58°C) but the vessel will not be repressurised within 20 minutes, the design minimum temperature may be higher than the minimum temperature by a temperature margin allowed in Appendix D of BS 5500. When the minimum temperature is lower than -50°C(-58°C) but the vessel is to be fully repressurised or partially repressurised, the design minimum temperature may be higher than the minimum temperature by a temperature margin allowed in Appendix D of BS 5500, depending on the stress level to be achieved. Once the design minimum temperature has been established, the material selection in relation to the impact test requirements and the associated test temperature shall be established using Figure 4 or 5 in BP Recommended Practice 4-6-1 (formerly CP8), Unfired Pressure Vessels, depending on the use of post-weld heat treatment. For design minimum temperatures down to -50°C(-58°C), appropriate grades of post-weld heat treated low temperature carbon steel may be used for any metal thickness. When a vessel is not to be post-weld heat treated, the temperature/thickness restrictions of Figure 4 in RP 4-6-8 shall be observed. For design minimum temperatures lower than -50°C, advice on material selection should be sought from TDU. Vessels subject to steam-out should have a note on the data sheet indicating the steam-out conditions (typically 1 barg and 120°C).

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Selection of operating pressure The operating pressure of any column is generally set as close to atmospheric pressure as reasonably possible, as determined by the available cooling medium and possibly by the availability of the heating medium, coupled with any thermal degradation properties of the bottoms product. Two important factors which can affect the column operating pressure are the heat sources and sinks to be utilised. The operating pressure should be set high enough to produce economic condensing approach temperatures. Typical approach temperatures to various heat sinks are as follows:

Condensing medium Approach temperature °C °F

Refrigeration 3-10 5-18 Cooling water 6-20 11-36

Pressurised fluid 10-20 18-36 Boiling water (steam generation) 20-40 36-72

Air 20-50 36-90 However, the pressure should not be high enough to give rise to reboiler temperatures which could degrade the bottom product. Typical approach temperatures to heat sources are as follows:

Reboiling medium Approach temperature °C °F

Process fluid 10-20 18-36 Steam 10-60 18-108 Hot oil 20-60 36-108

The designer should discuss the approach temperatures to be used with a heat exchanger specialist to establish the practicality of the approach temperature. One should also be aware of the following before finalising the column operating pressure: • Higher pressure operation reduces relative volatilities and increases the required reflux

rate and/or the number of stages required

• Higher pressure, after allowing for temperature effects, produces a net increase in vapour density and hence column capacity, or reduced column diameter for a given throughput

• Pressure may need to be set to allow operation outside an azeotropic or critical region

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BP Oil Distillation Handbook Rev: 3 Volume 1 General Design & Column Internals Date: Jun-01Section F Accumulator (Draw) Trays Page: 1F.1

Contents page Scope 1F.3 Definitions

1F.3

Accumulator trays

1F.4

Tray types The need for total liquid collection Liquid Availability Internal Reflux Control Minimising Flash Zone Contamination Process design criteria - chimney trays with separate caps 1F.6

Open area Pressure Drop Cap Design (separate caps) Vapour Distribution Vapour velocity between caps Chimney Height Residence time Riser orientation Riser cap hydraulics Liquid holdup on a shaped cap Liquid pouring distance from a vee shaped cap Liquid pouring from other pipes and ducts Liquid drawoff - external Static head for column outlet nozzles Self venting flow Liquid drawoff – overflash circuits Liquid drawoff - internal Downpipe sizing Vortex fomation Liquid seal Tray leakage Static leakage Dynamic leakage Use of gaskets Peripheral Sealing Sealants Seal welding Heat transfer on tray Condensation under chimney tray Interaction with adjacent equipment

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Chimney tray with attached cap 1F.31 Riser disposition Pressure drop Vane type tray

1F.32

Pressure drop Residence time Tray leakage Multi-channel accumulator tray

1F.34

Pressure drop Mechanical data

1F.35

Design Forces Manways Nomenclature for Section F

1F.37

References

1F.39

Revision History Rev 2 Original Issue DR 06/98 Rev 3 Liquid draw for overflash circuits added RH 06/01

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Scope This section includes all column internals which do not directly contribute to the separation of components, but whose use is nonetheless crucial to the proper functioning of a distillation or absorption column. Definitions Total drawoff tray A tray which positively collects all liquid falling onto it from the

whole column cross sectional area, thus preventing any of this liquid progressing to the tray below.

Chimney tray A specific type of total draw-off tray having vapour risers resembling

chimneys and separate caps over the risers. Vane type tray A liquid collection device comprising numerous Z-shaped or similar

vanes. Known also as a lamella tray. Several variations are used. Sloped cap tray A variation of the vane type tray which uses conventional channels,

but which uses a sloped vane as a cap, one end of which is attached (or integral with) the top of the riser.

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Accumulator Trays Tray types By design, the risers on a bubble cap tray allow it to come into the category of a total trap out tray. However, these trays are rarely used in modern refinery units, although they are still encountered in some older units. The design of bubble cap trays is not included in this section. There are numerous forms of liquid collection devices generally referred to as accumulator or chimney trays. Whereas many design principles are common to all types, for the purpose of this design manual the tray types are discussed separately. A chimney tray is perhaps the most commonly used form of total draw tray. Early designs tended to use circular chimneys. Due to the relative ease of fabrication (and therefore lower cost), and to the growing requirement for larger vapour flow areas it has become common practice to use square, rectangular or trough type chimneys with appropriately sized caps to prevent liquid falling into the risers. Whilst such chimneys are now widely used, there is appreciable variation in cap design, the latter depending on the nature of the column (vacuum or pressure) and to some extent on the user or tray manufacturer's philosophy towards minimising pressure drop. It should be noted that unless otherwise instructed a manufacturer will seldom provide the lowest pressure drop device possible, since to do so inevitably increases the cost. The detailed design of the chimney caps is discussed later. It is common practice in large columns (above approximately 5m diameter) to combine the chimney tray and the packed bed support into one structure. The resulting composite trestle beam thus needs to be located between chimneys thereby restricting access. The metal work above the caps will interfere to some extent with the vapour flow to the bed or trays above. This is discussed further under the paragraph dealing with vapour distribution. The need for total liquid collection The function of a total draw off tray is to collect all the liquid falling from the tray or packing above and to control the liquid discharge to the column below or to an external stream. At the same time the tray must allow the passage of rising vapour, and, at least for vacuum columns, must be of such a design as to minimise the pressure drop over the tray. Total liquid collection is generally required for any or all of the following reasons: • to ensure an adequate availability of liquid for the pumparound or product offtake • to control the quantity of internal reflux subsequently returned to the column • to minimise flash zone contamination • to thoroughly mix all the liquid at that point in the column before discharging to the next



zone These points are discussed in detail below. Liquid Availability When liquid needs to be withdrawn from the column - for example, as a product or for a pumparound circuit - it is necessary to ensure that an adequate flow can be collected and maintained, particularly under start-up conditions. If such a flow cannot be initiated, or maintained, the column will be inoperable, or at best, very inefficient. It is often assumed that if a part only of the calculated liquid load leaving a valve tray (for example) needs to be withdrawn, then there is no need for a total trap out tray. Whilst that may well be so in theory, experience has shown that this is not always seen in practice, and that unexpected weeping can cause high liquid losses. Total trap out trays should always be installed where a product or pumparound stream is required to be withdrawn from the vacuum column. Internal Reflux Control With a total draw system any internal reflux necessarily has to be returned externally to the column below the trap-out tray. It can then be metered and so some control over the quantity returned can be exercised. The need to know the flowrate is often desirable to ensure adequate irrigation of the packed section below. Minimising Flash Zone Contamination This particular aspect is specifically relevant to the overflash draw off tray located immediately above the flash zone. Overflash liquid will have already washed the lower part of the packed bed above the flash zone to remove asphaltic particles of such a size to have otherwise risen up the column from the flash zone. The liquid is thus heavily laden with metals and high carbon residue material and steps need to be taken to prevent that liquid from being recycled back up to the bed. If such recycle does occur there is a danger of fouling the packed bed above the flash zone. Without a chimney tray in this position, there is a danger that some of the falling overflash will be re-entrained up into the packed bed.

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Process design criteria - Conventional chimney trays with separate caps Open area The open area is the area available for vapour flow and will be a function of the desired maximum pressure drop. The following guide indicates typical values for conventional accumulator trays.

Typical application Percent open area (= riser area/column

area)

Normal pressure drop mb

Restriction to increased open area

Vacuum unit (conventional tray)

30-45 <1 access for seal welding

Vacuum unit (conventional tray)

45-80 <1

CDU 15-35 1-4 FCCU main fractionator 20-35 1-2 Pressure columns 5-15 1-2 Pressure Drop The incentive for considering chimney tray design in detail lies partly from the requirement to reduce the overall column pressure drop to a minimum when used in vacuum columns (not so important in pressure columns), and partly to ensure efficient liquid collection. Whereas such a philosophy may adversely affect vapour distribution, the cost of unnecessarily high pressure drop in a vacuum unit (approximately $70000/mmHg with crude at 18$/bbl) usually makes it advisable to minimise pressure drop. Pressure drop over a chimney tray refers to the tray plus the associated cap. Pressure drop can be considered to result from two features of the tray - vapour passage through the chimneys, and vapour passing through the open areas around the caps. Therefore, as the riser open area increases, and the associated pressure drop decreases, so the pressure drop through the caps which cover the chimneys increases.

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The vapour path shown will result in pressure losses as follows:

1

2

3 4 5

6

Figure F.1 Vapour path through risers and caps Point 1 - vapour converging into the chimney (riser) Point 2 - vapour emerging to space below the cap Point 3 - direction change to exit under cap Point 4 - vapour emerging to space between caps Point 5 - direction change to space between caps Point 6 - vapour passage through space between caps Experimental work at Sunbury (F1) covered various open areas, cap shapes and vapour emergence areas which account for points 1 to 4 above. Although a relationship for pressure drop between caps (points 5 and 6 above) was separately derived using FRI data for flow between bubble caps, the magnitude of this when compared with more recently performed CFD work suggested that the basic Sunbury correlation should be entirely adequate for industrial trays. Thus, although the pressure drop between caps is given below for the record, it is not included in the program in DE-6. The following relationships are used Ptray = 0.348*(Dv/Dlw)*(V/Cd)^2 Brit Ptray = 3.736*(Dv/Dlw)*(V/Cd)^2 Metric Cd = S (1.157*Aris*ln C + .488) S = .00273*ANG + 1 DE-6 incorporates the above correlations which have also been industrially verified (within 10%) at pressure drops up to approximately 10mmHg. For the record, but not to be used, is the estimated pressure drop over adjacent caps. For points 5 and 6 above: Pcaps = .0056*K*V2*Dv Where K = ((Ac-Ar) / Abc)^2 The effects of riser area and of vapour emergence area between the cap and the top of the riser are illustrated below. The distance the cap overlaps the riser should be minimised. A cap

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overlap of 25mm (1 inch) is recommended - there is no real need for any higher overlap. A lower (or zero) overlap is sometimes used, but the advantages are negligible. See later notes on cap design.

0 20 40 600

2

4

6

8

10

12

80Riser area (% of column area)

Tray

pre

ssur

e dr

op

(mm

Hg)

Parameter:

Vapour emergence area / riser arearatio of

0.8

1.0

1.2

1.5

2.0

Figure F-2 Pressure drop versus riser open area Figure F-2 shows the pressure drop for a range of vapour emergence areas each with a constant total cap area of 1.1 times the riser area - a typical ratio approximately equivalent to using a cap overlap of 25mm. Whilst absolute minimum pressure drop could be achieved with a vapour emergence area equal to twice the riser area, the extra column height required will usually be uneconomic. The recommended "optimum" value for vapour emergence area is 1.2 times the riser area. Figure F.3 shows the general effect of vapour emergence area ratio on pressure. Note that the vapour emergence area is reduced if one or more ends are blanked, and to maintain the same pressure drop the cap to riser distance will need to be increased.

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0.6 0.8 1 1.2 1.4 1.6 1.8 2 2.20.6

0.8

1

1.2

1.4

1.6

1.8

2

2.2

Vapour emergence area/riser area (ratio)

Pres

sure

dro

p

(mm

Hg)

Figure F.3 Effect of pressure drop on vapour emergence area (40% riser area)

As riser area increases, so cap area increases. With large cap areas (large cap overlaps) a situation could be reached where the reduction in pressure drop resulting from increased riser area is more than offset by the increasing pressure drop of vapour flowing between caps. This situation would not normally occur however until large open areas of approximately 60-70% or more were used. For mechanical reasons associated with access, the open area of a chimney tray with risers is unlikely to exceed 45%. Generally a pressure drop of about 0.3-0.5 mm Hg should be sought for vacuum columns; higher pressure drops are permissible for non vacuum units. Although BP does not specify an upper pressure drop limit as such, excessive pressure drop in a vacuum unit does directly affect the flash zone pressure achievable and thus the unit economics. A pressure drop exceeding 1.5mmHg would be regarded as excessive in a VDU.

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Cap Design (separate cap) Cap shape has an effect on pressure drop and on the way liquid drains from the cap. There is incentive to ensure that the flow path of the vapour leaving the risers is such as to minimise 90° turns. Such considerations lead to the use of an upturned V-shaped, or chevron, cap, as depicted in figure F-4.

slope

Vapour emergencearea blanked at thepouring end

Riser

Figure F.4 Recommended cap design for standard chimney riser in vacuum service A vee shaped cap is used partly to minimise the pressure drop and partly to ensure controlled liquid pouring from the cap. Whilst flat caps would not now be expected on a design for a new vacuum column, equally a contractor will seldom volunteer a change to a vee shaped cap in a revamp. There are however a number of units in BP with flat caps. The edges of these may be turned either up or down (or both) depending on the philosophy of the designer. Where they are turned up it is to prevent liquid falling over the side; if turned down it is to prevent liquid creeping to the underside of the cap and entering the riser that way. Some designers use a roof shaped cap which allows liquid to be directed to either side of the cap. This design suffers from having a 90 degree (at least) vapour direction change and hence an associated higher pressure drop than need be the case. It is not recommended. There is incentive to maximise the open area around the caps which effectively means reducing the extent of cap overlap to a minimum. The recommended overlap is 25mm, but occasionally some designers use a zero overlap. Whereas the idealised system assumes vertically falling liquid droplets and no vapour emerging from the risers, CFD studies have demonstrated that this does not always occur in practice. The use of a cap to riser open area ratio of 1:1 will result in an overlap of approximately 25mm. It is unlikely that this overlap will significantly affect pressure drop. The upturned vee-shaped cap (chevron shaped) as described should discharge liquid from one end. One of the major advantages of that shape is that it positively controls the liquid flow path and, unlike a flat cap, minimises unintentional flow over the side. It is extremely important that steps are taken, if necessary by blanking the area at one end of the riser, to ensure that liquid does not discharge into the vapour space of the adjacent or next cap. A typical arrangement is shown in figure F-5.

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discharge end

of cap

DowncomerDowncomer

L

V

Figure F.5 Chimney cap orientation A phenomenon of liquid pouring from a jug or cup is the so called "tea pot effect", where instead of liquid pouring in the normal way, some liquid clings to the container and falls elsewhere. The phenomenon is caused by the Bernoulli effect. In the case of a chimney tray cap some of the liquid pouring from the cap will run on the underside of the cap and down the riser in an analogous manner unless a suitable preventative device is installed. To avoid this phenomenon completely it is recommended that the vapour emergence area under the pouring end or ends of the cap is completely blanked. If the end is not fully blanked the cap should have a lip. This will significantly reduce the leakage, but may not eliminate it altogether. A research topic carried out by Sunbury quantified this effect. The work (F-2) illustrated that where liquid was free to drain from either end of a vee shaped cap, about 6% of the liquid falling onto the cap migrated to the underside and poured down the riser (that is, approximately 3% from each end of the riser). This total migratory liquid reduced to a total of 1.5% when a suitable lip was installed. The amount was also a function of the liquid rate. A cap is commonly inclined at an angle of between 1-2 degrees to the horizontal to assist drainage. With normal length caps (up to approximately 2000mm long) this does not significantly affect the uniformity of vapour emergence. However care is needed to ensure that, where the cap is unusually long (for example, the whole diameter of a large column), the difference in height between the two ends of the cap does not restrict the vapour emergence area. It is therefore recommended that the slope is fixed such that the total difference in level is between 20-50mm. In fact, providing the vee shaped cap rim is no less than 25mm, it would normally be acceptable to have no slope. A procedure for calculating liquid hold-up on a shaped cap, and the liquid pouring distance, is given in later in this section. Vapour Distribution



Good vapour distribution to the packed bed above is regarded as being essential to obtain the full efficiency from the packing, particularly in low pressure drop situations. Vapour leaving the caps of a multi-riser chimney tray will be maldistributed to some extent as there is no mechanism for vapour leaving a riser on one side of the tray to mix with vapour leaving the other side of the tray. A minimum distance of 12 inches (300 mm) should be used between the top edge of the cap and the underside of the packing support - subject to any other considerations, such as to allow for a manway. Allowing the packing to rest directly on the chimney trays is not acceptable. An item of particular concern is the presence of bed support beams intruding into the space above the riser caps. If this space contains a support beam or beams, as it will in columns of diameter greater than approximately 3m, care must be taken to ensure that these do not block the vapour space between adjacent caps. Such beams should run parallel to and above the caps. In some cases it will be necessary to perforate the beam to allow vapour cross-mixing. A case is known (F-3) where a poorly located beam was a contributory cause of premature bed flooding above. Vapour velocity between caps When the vapour velocity passing between the caps reaches a certain value falling droplets of liquid are re-entrained thus overloading the bed above, the amount of entrainment being a function of droplet size. This effective hold up of liquid reduces the liquid availability on the tray and can result in product shortage. The average velocity between caps should always be checked in highly loaded columns. The following criteria should be used:

Significance of vapour velocity between caps

Average velocity between caps Entrainment situation C-factor (British) C-factor (metric)

<=0.37 <=0.11 Acceptable 0.37-0.4 0.11-0.12 Possible entrainment 0.4 - 0.50 0.12-0.15 Entrainment highly likely 0.50 - 0.8 0.15-0.24 Entrainment will occur but should still be operable in

non-VDU situations providing there are no other restrictions

>0.8 >0.24 Expected to cause entrainment. Avoid. >=1.2 >=0.36 Will cause entrainment. Avoid.

Whilst the above are recommended values, it is recognised that chimney trays do operate apparently satisfactorily with high C-factors between caps and that in many cases it is nor practicable to avoid this. The following data from Kwinana CDU2 are given by way of example for a conventional riser type chimney tray. This unit performed well. Position Superficial velocity

in column

Velocity between caps

Velocity emerging between cap and

riser

Velocity in riser

Cs (metric)

Cs (Brit)

Cs (metric)

Cs (Brit)

Cs (metric)

Cs (Brit)

Cs (metric)

Cs (Brit)

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Preflash chim above feed 0.056 0.184 0.095 0.312 0.105 0.345 0.172 0.565 UIR 0.131 0.430 0.176 0.578 1.054 3.46 0.698 2.292 LIR 0.132 0.433 0.187 0.613 1.052 3.45 0.444 1.458 HGO 0.110 0.361 0.148 0.486 0.706 2.32 0.586 1.924 Overflash 0.094 0.309 0.127 0.417 0.603 1.98 0.501 1.645 Debut side 0.035 0.115 0.043 0.141 0.412 1.35 0.305 1.001 Debut reboiler feed 0.047 0.154 0.057 0.187 0.547 1.80 0.404 1.326 Chimney Height The height of the chimney should be such that the expected liquid level (or HLL if controlled) is a minimum of 100mm from the riser top. The operating liquid level on the tray will be set by one or more of the following criteria: • To provide sufficient hold up, or residence time, to protect the pump in cases of flow

fluctuation • To provide an adequate static head to ensure required liquid offtake rate is achieved • To provide an adequate NPSH for the particular pump in question. Note that in some

cases, for example, when liquid flows to a sidestream stripper, no pump is used, flow being by gravity only.

• To provide a reasonable control range for any level control located on the tray. • To limit, or control, residence time to meet thermal degradation criteria (high temperature

operation only). Flowrate fluctuations can result from step changes (increases or decreases) made during start up or normal operation, or from column upsets. Also, pump failure could result in build up of liquid on the tray and adequate time needs to be allowed to start the spare pump so avoiding liquid overflow down the chimneys. The definition of "adequate" is open to interpretation, and depends on a number of factors such as availability of high level alarms, size of valves on the pump, etc. On the other hand excessive chimney height results in the need for stronger and more expensive trays, and greater column height. Also, penalties for (temporary) liquid overflow need to be considered. The required net positive suction head depends upon the pump design, further discussion on which is given (for example) in the BP Handbook of Chemical Engineering Calculations (F-4). The available NPSH depends on the characteristics of the liquid being pumped and the physical location of the pump. In some specific cases - for example on trays from which liquid passes to a sidestream stripper - liquid is withdrawn from the tray by gravity only. In such cases the height of liquid on the chimney tray needs to be such as to provide sufficient head to overcome the frictional and acceleration losses as liquid leaves the column through the nozzle. Where the tray operates above approximately 315 degC (599deg F), thermal degradation can occur, the extent of which is a function of the residence time on that tray. This is normally of significance on a VDU overflash tray only and as such residence time on that tray (plus the recessed seal pan and lines) should be kept to a minimum. A cooled residue quench may be used, but is not recommended as the practice can reduce gas oil yield - . See Volume 5.



Residence time Adequate residence time is needed to provide time to take action in the event of pump failure. Sufficient residence time is also required to allow proper vapour disengagement so that non-aerated liquid can discharge from the offtake nozzle, distributor or other liquid destination. A minimum static head is required to allow liquid to flow from the nozzle; if this liquid is aerated the effective head will be lower and the flowrate will be reduced. The only known published residence times on a chimney tray are given in reference F-5, and reproduced below. A minimum residence time of 1 minute based on the volume of oil on the tray, neglecting that in the draw pan, is considered to be acceptable in most cases except for high temperature service (for example, the overflash tray when deep cutting - see TDU for advice). With debottlenecked units there is pressure to minimise residence time to save space. In such cases, where times may be as low as 15-30 seconds or less, the designer must ensure that unit safety and control are adequate. Cases are known (F-6,F-7) where the residence time on the tray has reduced to 5 and 4.5 seconds respectively. These times represent an absolute limit, but any time less than approximately 15 seconds is likely to be associated with a restricted drawoff capability unless the nozzles are sized for self-venting flow.

Recommended residence times for design of new drawoff trays with risers

Situation Minimum residence time (mins)

Liquid withdrawn by LC and fed to another unit 2 Liquid withdrawn by LC and pumped away using a manual start pump

3

Liquid withdrawn by LC and pumped away using an autostart pump

1

Liquid withdrawn by LC and fed to another unit that is some distance away, or has its instruments on a different board

5-7

Liquid withdrawal by flow control 3-5 Liquid flow to a thermosyphon reboiler without a level controller to maintain a level on the draw-off tray

1

External vessels can be utilised in cases where there is inadequate residence time on the tray. Indeed, use of an external vessel can sometimes replace a chimney tray with large risers, so creating space for additional packed bed depth or other internals. Riser orientation Several possible arrangements are shown below:

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Not recommended

Figure F.6 Possible arrangements for riser layouts Note that the riser orientation should be such as to minimise the hydraulic resistance to flow across the tray.

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Riser cap hydraulics Liquid hold up on a shaped cap The following procedure may be used for calculating the depth of liquid on a shaped cap. The cap should not overflow - normally a 25mm cap rim is adequate for most services.

hc

dh

x

y'yh

liquid level

angle ariser cap

riser

tray deck

liquid throw

cqp width Wc

Figure F.7 Liquid hold up and pouring distance - Vee shaped cap Assumptions: (1)Total liquid load is evenly distributed over column cross-sectional area. (2) Liquid level on caps is horizontal, that is, caps are installed without a slope towards the liquid exit. Qc = (C1*Vis^.02 *(1/tan a)^.996*(y)^2.47)+(C2*Wc*(hc)^1.5) From which h can be calculated y =0.5*Wc*tan a Metric units: C1=1.335 C2=1.838 British units C1=2.503 C2=3.333 If calculated level is less than or equal to y, repeat with h = 0. Value of y will then become the unknown level. Liquid pouring distance from a vee shaped cap For a chimney tray cap the horizontal velocity is calculated from the cross-sectional area of the liquid on the cap and the flowrate. The following applies for shaped or flat caps - refer to figure F-7 dimensions. Level in Vee part of cap only: 0.5 * (y’)^2 * tan(90-a) Level at top of Vee part of cap: 0.25 * W * y Level above Vee part of cap: (0.25 * W * y) + (h * w) Level in a flat cal: h * W



Liquid pouring from other pipes or ducts Liquid issuing from a pipe or duct will fall downwards due to the effects of gravity, calculation of which is made using the general theory of projectiles. Where this liquid passes through a rising vapour stream there will be an opposing force, although that is ignored in the procedures given below. The calculations are particurlarly applicable when estimating the position where liquid from a bottom tray seal pan will fall, or where reflux will impinge in the unlikely situation that a false downcomer has been omitted. Calculate the time taken to fall a vertical distance, and the horizontal distance travelled:- From the equation of motion: hv=uT + ½gt^2, if the initial vertical velocity is taken as zero, the time taken to fall a distance hv is as follows: T = SQR(hv / (0.5 * g) The horizontal distance from fluid source is then calculated as follows:- hh = Vel * T

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Liquid drawoff - external As a guide, different criteria are used for sizing offtake nozzles.

1/3 TS

LC

1/3 TS

LC

1/3 TS1/3 TS

dmc52

With level control Without level control Figure 7a: Sizing criteria for offtake nozzles Where level control is used the nozzle is sized for the head available based on the level being at the tray deck - note that the lower level tapping is usually 50mm above the deck in a chimney tray. For uncontrolled gravity flow the sizing criteria should be based on self-venting flow. Static head for column outlet nozzles It is necessary to ensure a minimum static head above the nozzle to overcome the frictional resistance associated with the nozzle and nearby pipe bends etc. The relationship for flow through a short tube nozzle is used, which is based on the following equations (reference F-8, for example). F = .76*A*(2*g*h)^0.5 Application of this equation to operating plant requires a safety factor to allow for the possible aeration of liquid due to poor vapour/liquid disengagement. A safety factor of 2.5 has been found to be acceptable in practice and is recommended. The following equations, which includes a safety factor of 2.5, can therefore be derived, and are applicable to each nozzle. h = 2327.2*F^2/d^4 or h = (.00017476*W^2)/(DL^2*d^4) The minimum liquid head above the centreline of the nozzle should be equal to one pipe diameter. Thus the following equation can be derived: d = 0.2913 * (W^0.4/DL^0.4)

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Self venting flow Where ingress of vapour is a possibility it can severely restrict the flow through an otherwise adequately sized nozzle. The following examples are given where the more conservative criteria of designing for self venting flow is recommended (F-16). • Gravity fed lines to sidestream strippers • Draw off pan from the base of a downcomer • Direct draw from chimney trays where residence time is less than 30 seconds d = Q^0.37 This will give a pipe diameter giving an average velocity of 70% of that from a conventionally sized nozzle. In many cases only the nozzle and the upper part of the pipe need be sized for self-venting flow. Figure F-8 illustrates recommended minimum pipe lengths required.

Reducer

2D

5-7D

Liquid headH

Diameter DNozzle flush with wall

If nozzle protrudesuse 4D

4D

Figure F-8 Offtake nozzle pipework Note that the design liquid head above the nozzle must exceed the pressure drop through the nozzle, the tee, and the reducer by at least 25%. This is adequately accounted for by the safety factor in the static head equation given above. Note that the minimum static head should be one diameter.

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Liquid drawoff – overflash circuits Overflash draw trays are not always total draw arrangements. Since it is important that these trays do not leak, where these are active trays the hole area must be checked to ensure minimal weeping. Seal welded or bubble cap trays are preferred for this reason. In pumped overflash systems (see Volume 4, Section E), some systems include a liquid return to the tray to boost the liquid rate over the tray thereby assisting liquid de-entrainment of the flash zone vapour. This is usually not recommended due to the associated increase in coking potential of the liquid (due to its nature) remaining on the tray for longer at high temperature. For vacuum columns, where the temperatures are high, usually above 400°C, the draw pan should be designed to avoid any stagnant zones. In these cases, draw pans with overflows are not recommended as liquid in the normal overflow section will be stagnant and prone to coking, as shown in Figure F-9. If internal overflows are required alternative designs as shown in Figures F-9b or c should be used.

Chimney Tray

SideDowncomer

Stagnant LiquidArea - possiblecoking potential

LC

DrawNozzle

Sidedowncom

er

'Sacrificial Downcomer'- will overflow before

other risers to preventlarge pressure drop

increase

Figure 9(a) Figure 9(b) Figure 9(c) Overflash draw prone

to coking Overflash draw

trap out pan Sacrificial chimney

In the case of F-9b, with correct design of the tray, the ‘sacrificial’ chimney will overflow without causing excessive pressure drop over the tray in the event of an overflash circuit failure. However, care is required in the placement of this small chimney to avoid placing the potential overflow in a direct flashzone feed path. The designs shown are based on side tray draws, however BP has several units fitted with central draws that work equally well based on the same principles.

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Liquid drawoff – internal Downpipe sizing With a partial drawoff, some liquid flows to the tray or packed bed below. This may be via segmental, rectangular or circular downpipes or ducts. The size of pipe needed is based on the hydraulics of flow down circular pipes which are sealed, the three hydraulic regimes of which can be described by reference to figure F-9 below.

Condition A Condition B Condition C

Froude No =<0.3 Froude No >0.3 to 2.0 Froude No > 2.0

Figure F-9. Hydraulic regimes for sealed downpipes Condition A occurs at low flowrates or when the downpipe size is large in relation to the flow. There is adequate facility for self-venting flow. However, since there is also facility for vapour to flow up the downpipe the seal at the base of the pipe must be adequate and dependable. Conditions B and C occur when the liquid flow is higher, or when the downpipe size is small in relation to the flow. Condition B could be subject to some vapour ingress at the top of the pipe. Condition C has an adequate depth of liquid such that vapour disengagement should occur before any liquid enters the downpipe. The maintenance of a seal at the base of the downpipe is less critical with Conditions C and D providing the depth of liquid in the pipe exceeds the pressure drop over the tray. This may not always be the case, particularly if the tray has an abnormally high pressure drop. It is recommended that designs are based on Condition A as this allows vapour disengagement by self venting. With sealed pipes bubbles tend to be trapped when the Froude number exceeds approximately 0.3 (Condition B) at which condition bubbles will be swept downwards (F-9). When the Froude number increases the tendency reduces until at Froude No = 2.0 (Condition C) the

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pipe will be full and any increase will cause backup. The minimum pipe diameter should thus be set to satisfy a maximum Froude number of 0.3 (Condition A). Nf = (Vl/(SQR(g*D))*SQR(Dl/(Dl-Dv)) For no trapped vapour the vapour density can be ignored. Thus: Nf = Vl/(SQR(g*Df)) but Vl = cfs / (0.7855 * Df^2) therefore D = 6.6*cfs^0.4/Nf^0.4 Max diameter is when Nf = 0.31 therefore Dmax (ins) = 10.544 * CFSL^0.4 Min diameter is when Nf = 2.0 therefore Dmin (ins) = 5.0 * CFSL^0.4 Vortex formation Vortex formation can occur regardless of the outlet nozzle orientation. One equation relates the formation of a vortex to the head, thus: If Hv < (Vv-1), then a vortex is possible Liquid seal As discussed above it important, particularly with Condition A, to ensure the seal is sufficient to balance the pressure drop over the chimney tray, plus a margin of safety. Figure F-10 defines the seal required and the pressure drop over the tray above:

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liquid seal

Traypressure

drop

vapour


Figure F.10 Downpipe liquid seal The pressure drop over an accumulator tray does not normally exceed 1mmHg (approximately 15mm liquid for middle distillates), although poor design or a highly overloaded column can increase that pressure 10 times or more. The liquid seal should be set to be a minimum of 50mm liquid in excess of that required to balance the pressure drop of the accumulator tray plus riser caps. Thus, typically a seal would not be less than approximately 75mm liquid. Note that if the seal is inadequate and vapour is allowed to rise up the downpipe (vapour will always find the path of least resistance) the liquid on the accumulator tray will become aerated and may eventually overflow the risers. This will increase the accumulator tray pressure drop even further (it restricts the riser to cap distance), and may also reduce the static head available to discharge the product and lead to a severe restriction in the quantity of product offtake rate. Tray Leakage This section primarily addresses the type of chimney tray with separate caps. However, many of the principles equally apply to other cap forms In theory, an accumulator tray, with caps over the risers, should not leak. The fact that it can and does is attributable to occurrences in two categories - mechanical defects or unexpected fluid dynamics. Leakage from these sources are additive. For non seal welded trays, leaks are likely to occur at points where three panels join - whether or not gaskets are used. This is made worse by expansion during high temperature operation or if gaskets become damaged. Tray buckling, inadequate bolt spacing or lack of bolt washers can also contribute to the problem. With a seal welded tray where both panels and the periphery may be seal welded, leakage should not occur. However, occasionally welding is inexpertly done, or sections are left unwelded by mistake. It is recommended that critical chimney trays are fully seal welded. This not only ensures that they are substantially leak tight after installation, but ensures that they remain so during operation. The latter is frequently not the case with gasketted trays. Metal thickness of fully seal welded trays should not be less than 3mm (10ga is 3.4mm). Static Leakage Where the tray allows, static leakage can occur through joints that are supposed to be leak tight but nonetheless leak as a result of poor design or workmanship. Static leakage should be checked after construction in which case a number of possible leak test procedures may be used. Such tests are based on the assumption that a certain degree of

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leakage is deemed to be acceptable. Although most of the time leakage is far lower than the maximum specified, it is instructive to have knowledge of what the leak rate would have been had the test only just satisfied the particular criteria adopted.

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Recommended tray leak test criteria as given in GS-146-8

BP leak test category Maximum level loss at prevailing liquid head (mm/hour)

Typical applications

1 10 Lowest distillate draw in residue vacuum columns

2 50 Vacuum columns - except category 1 3 200 All other columns

Note: Category 1 should be used for all columns where there is incentive to minimise loss of the lowest side-draw material to the residue. The categories given are designed to cover all units and recognise that different situations require different criteria. Approximately, the categories cover the following typical situations. Dynamic Leakage Dynamic leakage trays from with separate caps can occur where liquid is deviated from the expected flowpath by unanticipated hydraulic behaviour, and which then passes down through the vapour riser. This form of leakage occurs from two probable sources. Firstly leakage is known to result from liquid migrating from the top of the cap to its underside due to the "teapot" effect. Secondly, some leakage is thought to result from deviation of liquid flowing down between adjacent caps due to vapour turbulence, perhaps where negative velocities are encountered due, for example, to vortices under the tray. Sunbury have carried out tests (F-10) to determine leakage from the first source; no test data covering the second source are yet available. Liquid migration to the underside of the cap can be substantially reduced by providing a lip on the pouring edge. It can be practically eliminated by blocking the vapour emergence area at the pouring end(s). Complete blocking of the appropriate vapour emergence end is recommended for category 1 and 2 trays. Use of gaskets The use of a gasket at the tray periphery is still found and probably does not give rise to large leaks because of rigidity of the tray support ring. However, this will also depend on the clamp spacing used, washer size and the tray metal thickness, etc. Clamp spacing is not normally specified by BP. It is suggested however that spacing does not exceed 6 inches. Washers should be 1.5 ins (38mm) diameter. Where a gasket is used for tray panels there is always the danger that during operation it will become damaged. It is considered essential therefore that alternative means of ensuring a leak tight peripheral seal at critical locations (e.g. overflash tray) are used.

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Peripheral sealing One such method is shown below, the principle of which will be self explanatory.

5mm space

Figure F-10 A peripheral sealing arrangement Sealants Where an existing (gasketted system) is found to leak at the periphery the use of sealants have been used (F-11) although we have no direct knowledge of its use or success. However, we do have experience of using a proprietary sealant (Belzona) to rectify leaks in a poorly seal welded tray which does not contain gaskets.. Note that wherever possible, seal welding is definitely preferred. Belzona 1111 Super Metal sealant was used to seal leaks at the periphery of a chimney tray at Coryton in April 1998. The sealant extended approximately 50mm up the wall and into the deck. Consideration was given to covering the Belzona 1111 with Belzona 1391 (a ceramic coating) for additional chemical resistance, but, since the No 1111 already had excellent resistance to benzene and a wide range of hydrocarbons present, it was decided not to use the additional coating. The data sheet indicates Belzona 1111 is thermally stable to 200degC dry and 93 degC wet (immersed in the process fluids), whereas the Belzona ceramic coatings 1391 and 1391X have maximum temperatures of 185 and 130 degC respectively. The tray in question operated at 142 degC (design), although the 125 psi steam used for steam-out has a temperature of approximately 180 degC. In the event, the sealant proved to be totally intact after sevceral days of operation during which time the sealant was immersed in the liquid at ca 155 degC and having benzene and toluene contents of 6.1%wt and 9.3%wt respectively. The treatment area was cleaned using air operated wire brushes - grinding to clean the surfaces would have been preferable but a hot working permit could not be obtained. The exposed area was dried using paper sheets. The ambient temperature during application was 0-5 degC. A portable heater was hired to increase the local tray temperature to ensure



adequate curing before start-up. The heat was applied for a total of 10.5 hrs at a temperature of approximately 30 degC. For Belzona 1111, following curing times apply: Curing times for 6mm thickness (hours) Temperature degC 5 10 15 20 25 30 Movement or use involving no loading or immersion

4.0 3.0 2.25 1.75 1.0 0.75

Machining and/or light loading 6.0 4.0 3.0 2.0 1.5 1.0 Curing times for 6mm thickness (days) Full mechanical or thermal loading 4.0 2.0 1.5 1.0 0.83 0.67 Immersion in chemicals 5.0 4.0 3.0 2.0 1.5 1.0 Thus, curing time was approximately half that recommended for full immersion chemical resistance. Inspection after a week of operation confirmed the sealant was fully intact. Seal Welding Seal welding is the preferred method of ensuring a leak tight tray. Note however, that even seal welded trays can leak if the welding has not been performed properly. Therefore, where absolute leak tightness is required (for example, the tray from which HVGO is drawn) this should be specifically noted. A tray intended for seal welding is normally made from 3mm stainless steel and numerous examples of such trays operating successfully are known. A thinner gauge is not recommended unless very specific design methods are used to verify the choice. Note that BPC seal weld trays made from 2mm sheet metal when using expensive alloys or metals. At the temperatures normally encountered in VDU's (up to approximately 400 deg or 750 degF,) fully seal welded trays are acceptable although the design and welding procedures must be checked and approved by BP. Welding procedures may be obtained from OTN. Chimney trays are normally fabricated from AISI 410 which has a different coefficient of thermal expansion to CS and under operation it is necessary to ensure that these resultant expansion forces can be contained. BP has operated seal welded trays for at least 20 years without any known problem. It is necessary however to use material 3mm thick (or 10 gauge (3.4mm)) and to ensure the welds at key points are sound. Seal welded trays may buckle slightly due to thermal expansion, but this is not in itself seen as a problem. In some cases, particularly where naphthenic acid corrosion is possible, AISI 316 or 317 is used for these trays. AISI 300 series has a greater thermal expansion coefficient than that of AISI 410 and thus greater forces will apply. Koch report (F-12) that they have had two all welded trays in AISI 316 operating for 5 years without problem and that these trays have been inspected during that 5 year period. Tray thickness in this case is 4mm .

Mean coefficients of thermal expansion



mm x 10-6 per m from 0 degC to temperature 100 degC 300 degC 315 degC 500 degC 538 degC AISI 316 15.9 15.2 17.5 AISI 317 15.9 16.2 17.5 AISI 405 (similar to 410S) 10.8 11.6 12.1 AISI 410 9.9 11.4 11.6 mm x 10-6 per m from 25 degC to temperature Carbon steel (1022) 12.2 13.1 13.9 Heat transfer on tray Rising vapour impinging onto the underside of a cooler chimney tray will suffer partial condensation, thus in theory creating more internal reflux than expected. To reduce this possibility some designers insulate the chimney tray. The extent of condensation depends on the effective overall heat transfer coefficient, but is likely to be small. Also there is visual evidence from film of operating test columns that some condensation formed on the underside of a tray migrates up the riser due to high local velocities. A procedure for estimating the overall heat transfer coefficient is given in Appendix 2, although data from one unit where insulation was used suggests that the actual effective condensation (liquid reaching the bed below) is significantly (5 times plus) less. When a tray is insulated, insulation is invariably installed within, rather than under, the tray. The procedure is to tack weld steel mesh onto the deck and to lay the refractory material to a depth of 50mm. The steel mesh provides a means of ensuring a firm attachment. In most cases the additional weight of the refractory material will require the installation of a new suitably strong tray. Condensation under chimney trays The chimney tray is considered as a heat exchanger, with the oil on the shell side and vapour on the tube side. Note: units to be rationalised in a later Manual update. 1) Calculate tube side heat transfer coefficient (vapour to tray) Use relationship: Nu = .26*Re^.6*Pr^.3 Thus, Ht*Dc/k = .26*(353.7*W/(vis*D))^.6*(Cp*vis/k)^.3 2) Calculate shell side heat transfer coefficient (liquid to tray) Use relationship for segmental baffle-exchanger. Thus, hs*Do/k = .22*(Do*Ge/vis)^.6*(C3*vis/k)^.33*(vis/visw)^.14 (W/m2.C = 5.6783*Btu/ft3.h.F) 3) Overall tray heat transfer coefficient



1 = 1 + 1 U ht hs Where: U = overall tray heat transfer coefficient, W/m2 deg C 4) Heat extracted from vapour to form condensate Hc = U*Acn*DT/1000 5) Calculation of condensed liquid - calculate dew point temperature of vapour to tray (GENESIS) - perform a series of flash calculations down to approximately 10 deg C below the dew point - establish by interpolation the condensation temperature at which Hc kw has been extracted - note the corresponding condensed liquid flowrate. This liquid rate will be the amount condensed on the under side of the chimney tray. 6) Approximate values From previous work the following numerical can be regarded as typical: Overall HTU (U) W/m2.C 350-480 Vapour dew point reduction degC 10 Vapour temperature reduction per kw heat extraction degC 0.09 Condensed liquid per 1 deg C temperature reduction kg/h 23

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Interaction with adjacent equipment Whilst chimney tray design procedure does not and cannot take account of other adjacent equipment, the presence of beams etc above the tray must be carefully considered as they can interfere with the correct tray operation . Normally a chimney tray drawing does not show such adjacent items, even though the data are shown on other drawings. Unless all items are on the same (chimney tray) drawing, and adverse interaction may well be missed. Examples of beams restricting the vapour emergence from risers are known. See photo Library: Trouble.doc No 2

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Chimney tray with attached cap In concept these trays differ from conventional chimney trays only in that the caps are attached to the risers on their longer side. This feature affects the pressure drop and the vapour disposition after leaving the risers. It also increases the "between cap" area and thus reduces its velocity which, in some circumstances, might otherwise be sufficiently high to cause entrainment. Riser disposition The slope of the cap imparts a directional force on the emerging vapour. In order to achieve as uniform a vapour distribution as possible to the bed above the riser dimensional data given in figure F-12 should be used, which have been derived from extensive CFD analysis of actual trays (F-13)

2P P RR

C

2P

P/2

Yes No

Cap lip position

Cx

H

Acceptableforms of cap

2R

Figure F-12 Chimney tray with attached caps - recommended dimensional criteria Pressure drop The following approximate relationships may be used. This is not based on direct Sunbury data, but assumes the same equation as a conventional tray with a cap slope of 45 degrees. Ptray = 0.348*(Dv/Dlw)*(V/Cd)^2 Cd = S (1.157*Aris*ln C + .488) S = 1.123

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Vane type tray This type of tray is commonly used by Sulzer. It is variously known as a chevron or lammella type collector tray. Open area can be very high (about 97%) and the device has a low pressure drop. Residence time on the tray is relatively low but is increased by using an annular collector ring. Pressure drop Sunbury have performed pressure drop measurements on various forms of non conventional tray (F-14) and have derived the following relationships. DP = Kd*((Fv*Dvm^0.5)/(Acm*Aris))^2 Where Kd = 0.03 for a vane type tray

basis for open areaKd = 0.03

Figure F-13 Value for constant K for a vane type accumulator tray

Residence time Residence time on the tray itself is very low, relying on a peripheral channel to provide the main residence time. The peripheral channel width is commonly 10% of the column diameter and has a similar depth. Overall residence times of approximately 30 seconds are normal for these trays, although in general the guidelines given earlier for conventional trays would still apply. Tray leakage As with trays with risers, other trays can suffer static and dynamic leakage. Static leakage occurs through poor seal welding or bolted joints. In one case (F-15) this leakage was caused by too large a bolt spacing and, possibly, inadequate bolt tightening. Leakage from this source amounted to approximately 10% of liquid inventory. Dynamic leakage can occur at the ends of the vanes where they rest by gravity on the edge of the ring channel. As liquid pours into the channel, a small proportion can run back along the underside of the vane because of the "teapot" effect. The tray in reference (F-15) also suffered from dynamic leakage in this way. Sunbury have produced the following correlations (F-14) to estimate the extent of dynamic leakage. Modifying the vane ends to form a lip greatly reduces dynamic leakage. Equations are also given to estimate the leakage after modification. Note that the correlations were



developed on a 600mm diameter simulator and are not necessarily valid for any significantly larger diameter. However on one occasion when the relationship was tested on a larger column, the equation predicting leakage from the unmodified tray was plausible, but results from the modified tray equation were invalid. Non modified tray Overall: LK = 24.643*Fs*1.98*LN(LQ*.00019)+266.405 Fs<1.751 LK = LQ*Fs*.037-350.312*Fs*ln(LQ*.0005)+340.489 Fs>1.751 LK = LQ*Fs*-.0025+95.034*Fs*ln(LQ*.0004)+108.477 Modified tray Overall: LK = LQ*Fs*-1.014+LQ*(Fs+.0234)*1.019-88.867 Fs<1.751 LK = LQ*(Fs+8.151)-2.811)*1.335+6.5 Fs>1.751 LK = LQ*Fs*-1.07+LQ*(Fs.0897)*1.05-266.621

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Multi-channel accumulator tray These trays are occasionally encountered but are not recommended. They are likely to occupy more height (than a vane tray) and their inherent leak proof characteristics are unknown. Data are given however as there is sometimes a need to evaluate such designs. Pressure drop Pressure drop is calculated using the equations for vane trays, but using the values for Kd given in figure F-14 below.

Kd = 0.065 Kd = 0.02

Figure F-14 Typical arrangements for multi-channel accumulator trays

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Mechanical data Design Forces Chimney trays are mechanically designed taking account of both downward and upward forces. The following criteria should be used: Downward force Weight of the tray when full of water plus the point weight of two

men. For design purposes the weight of two men is taken as 150 kg.

Upward force A minimum value of 60 lb/ft2 (300 kg/m2) should be used. However, an upward force of 150 lb/ft2 (750 kg/m2) is recommended for those trays subject to high forces - such as immediately above the flash zone.

Some companies use 200 lb/ft2 (1000 kg/m2) for those trays but this is not normally considered necessary, particularly with high open area trays. Trays other than the tray immediately above the flash zone are normally designed on the basis of downward forces only.

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Manways Access through chimney trays is necessary both for inspection of the tray itself and for inspection of internals above and below the tray. Normal access is through one riser which has a removable cap. Minimum width of that riser should be 18 ins (450 mm). More than one manway may be required in large columns. Columns with a diameter exceeding 30 ft (9m) should have two such manways. With tall chimneys it may be preferable to have the manway in the side of the riser although the design will need to ensure that the manway is leak tight. Gasketted joints will normally be provided. In any event all access risers equal to or exceeding 18 ins (450mm) height should be equipped with internal rung ladders.

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Nomenclature for Section F Symbol Description Metric units British units a Angle of cap Vee to the horizontal deg deg A Cross-sectional area of pipe - ft2 Ar Riser cross-sectional area - ft2 Aris Riser area as a fraction of column area - - Abc Open area between riser caps - ft2 Ac Column cross-sectional area - ft2 Acm Non riser cross sectional area of tray m2 - ANG Cap angle (slope) to the horizontal - degrees Ar Riser cross-sectional area - ft2 C Vapour emergence as a fraction of riser area - - Cx Distance between riser and cap for vapour emergence - - C1,C2 Constants for calculating depth of liquid on a Vee cap - - C3 Constant for heat transfer under tray - Use C=1.0 cfs Volumetric flow of liquid - ft3/s cp Specific heat of vapour kj/kg degC - d Nozzle diameter - ins dh Distance between tray deck and riser cap base m ft D Pipe diameter - ins Dc Column diameter m - Df Pipe diameter - ft Dl Liquid density - lb/ft3 Dlw Density of water 998 kg/m3 62.4lb/ft3 Dmax Maximum pipe diameter - ins Dmin Minimum pipe diameter - ins Do Riser diameter - ft DP Tray pressure drop mmHg Drec Recommended maximum pipe diameter - ins Dt Temperature difference between vapour and liquid C - Dv Vapour density - lb/ft3 Dvm Vapour density kg/m3 - F Volumetric flowrate - ft3/s Fs Vapour velocity function m3/sec*(Dv)^0.5 Fsb Vapour velocity function - ft3/sec*(Dv)^0.5 Fv Flow of vapour through tray m3/s - g Gravitational constant 9.81 m/s/s 32.2ft/s/s Ge Liquid mass velocity - lb/h.ft2 h Head of liquid from nozzle centre line - ft hc Liquid depth in rectangular part of cap (vee or otherwise) m ft hh Horizontal distance traversed by liquid m - hs Film coefficient - Btu/ft3.h.F ht Film coefficient W/m2.degC - hv Vertical distance of falling liquid m - H Height of riser - - Hc Heat extracted from vapour kw - Hv Depth of liquid avove the orifice (vortex formation) ft k Thermal conductivity W/m2.degC - Kd Constant for non conventional tray pressure drop - - K1-K4 Constants for leaking tray from lip (std tray with risers) - - LK Vane tray leakage rate (per column area) kg/h/m2

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LTO Liquid flowrate - lb/h LQ Liquid rate to vane tray kg/h/m2 Nf Froude Number - - P Width of pan - - Pcaps Pressure drop between adjacent riser caps - mmHg Ptray Pressure drop over riser and cap mmHg mmHg Q Volumetric liquid flow - USGPM Qc Liquid flowrate leaving one cap m3/s ft3/s R Width of riser - - T Time for liquid to fall a vertical distance secs - U Overall tray heat transfer coefficient W/m2.deg C - V Vapour velocity through riser m/s ft/s Vel Initial horizontal velocity of fluid m/s - Vis Viscosity cp cp Visw Viscosity of water cp cp Vl Superficial liquid velocity in pipe - ft/s Vv Orifice velocity (vortex formaton) ft/s W Liquid flowing through nozzle - lb/h Wc Width of riser cap m ft Wv Vapour flowrate through tray kg/h - X Horizontal liquid flow from a cap m ft y Height of sloped part of cap m ft y’ Level of liquid in sloped part of vee shaped cap m ft

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References F-1 RCS TM 134471 dated 28th January 1985, "Distillation column internals - an

experimental assessment of the operating characteristics of chimney trays". F-2 Unpublished RCS Report (May 1985). F-3 BPE Report "SRC CDU 1 - Troubleshooting investigations",BPE.92.ER.100, D W

Reay F-4 BP Handbook and Chemical Engineering Calculations, Vol. 1. F-5 Hydrocarbon Processing. July 1968, Vol. 47, No. 7. Page1190120 F-6 BPE Report "Bulwer Island Primary absorber 203-E.Investigation into pumparound

draw-off problems", BPE.92.ER.216, D W Reay F-7 BPE Report "CPRL CDU Column V-101. Investigation of an AGO draw-off

problem", BPOI.93.ER.001, D W Reay. F-8 Piping Handbook. Crocker and King, Edition 5. 1976. F-9 Chemical Engineering/Deskbook Issue 14 April 1969 F-10 RCS Technical Memorandum No. 135279, dated June 5th, 1987. F-11 Hydrocarbon Processing, September 1981, pages 131-133 F-12 Telcon Bouck/Reay on 13/9/93 in connection with Alliance VDU revamp in a 28ft

col) F-13 SET Report Simulation of vapour flows through chimney tray risers.

Report No 1994-221864 February 1994 F-14 RCS Report An experimental evaluation of the performance of liquid collection

devices for distillation duties, Report No 123625 dated 27 June 1989 F-15 D Reay communication with Bulwer Island, 1991 F-16 Sewell. The Chemical Engineer, July/August 1975, pp442-445.

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BP Oil Distillation Handbook Rev: 0 Volume 1 General Design & Column Internals Date: Mar-95 Section G Other Column Internals Page: 1G.1

Section not yet available

Will include - mist eliminators

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Documents

Distillation Handbook BP Oil.pdf